Catalyst Deactivation 1997

Studies in Surface Science and Catalysis 111 CATALYST DEACTIVATION 1997

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Studies in Surface Science and Catalysis A d v i s o r y E d i t o r s : B. D e l m o n and J.T, Yates Vol. 111

CATALYST DEACTIVATION 1997 Proceedings of the 7th International Symposium Cancun, Mexico, October 5-8, 1997 Editors

C.H. Bartholomew Brigham Young University, Provo, U.S.A. G.A. Fuentes A. Metropolitana-lztapalapa University, Mexico, Mexico

1997 ELSEVIER A m s t e r d a m m Lausanne m N e w York m O x f o r d - - S h a n n o n - - S i n g a p o r e - - Tokyo

ELSEVIER SCIENCE B.V. Sara Burgerhartstraat 25 P.O. Box 211, 1000 AE Amsterdam, The Netherlands

ISBN 0-444-82603-3 91997 Elsevier Science B.V. All rights reserved. No part of this publication may be reproduced, stored in a retrieval system or transmitted in any form or by any means, electronic, mechanical, photocopying, recording or otherwise, without the prior written permission of the publisher, Elsevier Science B.V., Copyright & Permissions Department, P.O. Box 521, 1000 AM Amsterdam, The Netherlands. Special regulations for readers in the U.S.A.- This publication has been registered with the Copyright Clearance Center Inc. (CCC), 222 Rosewood Drive, Danvers, MA 01923. Information can be obtained from the CCC about conditions under which photocopies of parts of this publication may be made in the U.S.A. All other copyright questions, including photocopying outside of the U.S.A., should be referred to the copyright owner, Elsevier Science B.V., unless otherwise specified. No responsibility is assumed by the publisher for any injury and/or damage to persons or property as a matter of products liability, negligence or otherwise, or from any use or operation of any methods, products, instructions or ideas contained in the material herein. This book is printed on acid-free paper. Printed in The Netherlands

CONTENTS

Preface

xiii

Review Articles (Plenary and Award Lectures) Roles of Acidity and Pore Structure in the Deactivation of Zeolites by Carbonaceous Deposits (Plenary lecture) M. Guisnet, P. Magnoux, and D. Martin Impact of Sulfur on Three-Way Automotive Catalyst Performance and Catalyst Diagnostics (Plenary lecture) D. D. Beck

21

Solid State Reactions in Catalysts: An Approach to Real Active Systems and their Deactivation (Founders award lecture) B. Delmon

39

Coke Formation in Catalytic Processes: Kinetics and Catalyst Deactivation (Founders award lecture) G. F. Froment

53

Catalyst Deactivation: How We Cannot Yet Subvert Nature (Maxted award lecture) J. B. Butt

69

Catalyst Deactivation: Opportunity amidst Woe (Maxted award lecture) E. E. Petersen

87

Topical Articles (Oral and Poster Presentations) Carbon Deposition and Coking The Relationship Between Metal Particle Morphology and the Structural Characteristics of Carbon Deposits R. T. K. Baker, M. S. Kim, A. Chambers, C. Park and N. M. Rodriguez

99

Self-Poisoning and Aging of Pd-Ag/AI~O~ in Semi-Hydrogenation of 1,3Butadiene: Effects of Surface Inhomogeneity Caused by Hydrocarbonaceous Deposits A. Sarkany

111

Influence of the Support on the Deactivation of Nickel Zeolite Catalysts During the Conversion of Phenylacetylene E. D. Gamas and I. Schifier

119

Activation and Deactivation of the Zeolite Ferrierite for Olefin Conversions K. P. de Jong, H. H. Mooiweer, J. G. Buglass and P. K. Maarsen

127

Deactivation of Ferrierite during the 1-Butene Skeletal Isomerization R. A. Comelli, Z. R. Finelli, N. S. Ffgoli and C. A. Querini

139

vi

Alcohol Dehydration Reactions as Chemical Precursors for Coke Formation and Acidity Probes in Tungstated Zirconia Catalysts G. Larsen, E. Lotero, M. Nabity, L. Petkovic and C. A. Querini

147

Determination of Coke Deposition on Metal Active Sites of Propane Dehydrogenation Catalysts P. Praserthdam, T. Mongkhonsi, S. Kunatippapong, B. Jaikaew and N. Lim

153

Chemicals

The Role of Coke Deposition in the Conversion of Methanol to Olefins over SAPO-34 D. Chen, H. P. Rebo, K. Moljord and A. Holmen

159

Deactivation and Regeneration of Alkane Dehydrogenation Catalysts S. D. Jackson, J. Grenfell I. M. Matheson, S. Munro, R. Raval and G. Webb

167

A Novel Mechanism of Catalyst Deactivation in Liquid Phase Synthesis Gas-to DME Reactions X. D. Peng, B. A. Toseland and R. P. Underwood

175

Activity, Selectivity and Coking of Bimetallic Ni-Co-Spinel Catalysts in Selective Hydrogenation Reactions. J. C. Rodrfguez, C. Guimon, A. J. Marchi, A. Borgna and A. Monz6n.

183

Stability and Regeneration of Supported PtSn Catalysts for Propane Dehydrogenation C. L. Padr6, S. R. de Miguel, A. A. Castro and O. A. Scelza.

191

Deactivation and Shape Selectivity Effects in Toluene Nitration over Zeolite Catalysts J. M. Smith, H. Liu and D. E. Resasco

199

Regeneration of VPMeO Catalysts for n-Butane Oxidation by Means of Mechanochemical and Barothermal Treatments V. Zazhigalov, A. Kharlamov, J. Haber, J. Stoch, V. Yaremenko, L Bacherikova, L. Bogutskaya

207

Environmental

Sulfur Tolerance of Cu- and H-Mordenite Zeolite Catalysts for the Reduction of NO by Hydrocarbons M. H. Kim, L-S. Nam and Y. G. Kim

213

Deactivation of Cu-ZSM-5 during Selective Catalytic Reduction of NO by Propane under Wet Conditions A. Marffnez, S. A. G6mez and G. A. Fuentes

225

o~

VII

H-Mordenite Deactivation during the SCR of NOx. Adsorption and Diffusion of Probe Molecules on Fresh and Deactivated Catalysts E. E. Mir6, L. Costa, J. M. Dereppe and J. O. Petunchi

231

Deactivation of Pt/Alumina Catalysts for the Hydrodechlorination of 1,1,1Trichloroethane K. A. Frankel, B. W.-L. Jang, G. W. Roberts, J. J. Spivey

239

Understanding Claus Catalyst Deactivation Mechanisms" Optimization of Alumina Using Physico-chemical Parameters C. N~dez and J.-L. Ray

251

Deactivation by Poisoning and the Improvement of Three Way Catalysts for Natural Gas-Fueled Engines T. Tabata, H. Kawashima and K. Baba

259

Deactivation of Manganese-Cerium Oxide Catalysts during Wet Oxidation of Phenol P. Oelker, A. Bernier, P. Ruiz, B. Delmon and P. lsnard

267

Modeling and Kinetic Studies Coke Deactivation of Hydrotreating Catalysts: A Variable Site Model F. E. Massoth

275

On the Metal Deposition Process during Hydrodemetallation of VanadylTetraphenylporphyrin J. P. Janssens, R. M. de Deugd, A. D. van Langeveld, S. T. Sie and J. A. Moulijn

283

Coke Formation in Fluid Catalytic Cracking M. A. den Hollander, M. Makkee and J. A. Moulijn

295

Mathematical Modeling of Deactivation by Coke Formation in the Cracking of Gasoil H. S. Cerqueira, E. C. Biscaia Jr. and E. Falabella S.-Aguiar

303

Catalyst Decay by Simultaneous Sintering and Poisoning: Effects of Intraparticle and Interfacial Gradients M. Chocron, N. Amadeo and M. Laborde

311

A Model for Catalyst Deactivation in Industrial Catalytic Reforming L. M. Rodrfguez Otal, T. Viveros Garcia and M. Sdnchez Rubio

319

Petroleum Effects of the Metal-Metal Interactions on the Stability of Pt-Re/A1,0cC1 Reforming Catalysts J. Barbier, P. Marecot and C. L. Pieck

327

Temperature Programmed Oxidation of Deactivated Pt/Nb,O, Catalysts D. A. G. Aranda, J. C. Alfonso, R. Frety and M. Schmal

335

viii Effects of Sulfidation of Mo Nitride and CoMo Nitride Catalysts on Thiophene HDS S.-K. lhm, D.-W. Kim and D.-K. Lee

343

Catalyst Deactivation by Metals and Coke During Hydrodemetallation M. Nu~ez, E. G6mez, C. Suarez and E. Carmona

351

Deactivation of Pt-Sn/A1,O, Catalysts by Coking: Influence of the Preparation Method. G. Corro, P. Marecot and J. Barbier

359

New Developments in FCC Catalyst Deactivation by Metals: Metals Mobility and the Vanadium Mobility Index (VMI) L. T. Boock, J. Deady, T. F. Lim and G. Yaluris

367

Stability of an FCC Catalyst Matrix for Processing Gas Oil with Resid P. Gamero M., C. Maldonado M., J. C. Moreno M., O. Guzman M., E. Mojica M. and R. Gonzalez S.

375

Kinetics of Coke Combustion during Temperature-Programmed Oxidation of Deactivated Cracking Catalysts C. L. Minh, C. Li and T. C. Brown

383

Activity, Selectivity and Deactivation of Y Zeolite: A Compromise in Fluid Catalytic Cracking Catalysts C. Maldonado M., P. Gamero M., F. Hernandez B., J. A. Montoya F., J. Navarrete, J. A. Toledo A., A. Vazquez R. and A. Vargas

391

M/H-MFI Catalyst Deactivation and Reaction Kinetics for n-Octane Transformation J. Papa, F. Santos, M. J. Becerra, G. Giannetto, F. Yanez and L. Garcia

399

Isobutane Alkylation with C, Olefins: Low Temperature Regeneration of Solid Acid Catalysts with Ozone C. A. Querini, E. Roa, C. L. Pieck and J. M. Parera

407

Influence of Chloride during Coke Burning of a Naphtha Reforming Catalyst Coked in a Commercial Cycle M. C. Rangel, M. N. M. Barbosa, C. L. Pieck and N. S. Figoli

415

Deactivation of Pt-Au/A1,O, Catalysts Prepared by Surface Redox Reaction: Effects of Sulfur and Coke Deposition G. Espinosa, G. Del Angel J. Barbier, P. Marecot and L Schifier

421

Nature of Coke Formed from Tri-isopropylbenzene Over USY Zeolites with Different Rare Earth Content C. A. Henriques, E. Falabella S.-Aguiar, M. L. Murta Valle, S. Varela and J. L. F. Monteiro

427

Differential Effect of Coke Burning with Oxygen or Ozone on Pt-Re Interaction in Pt-Re/A1,O, C. L. Pieck and J. M. Parera

433

ix Hydrodesulfurization of Dibenzothiophene on a Nitrided Supported Molybdena-Alumina Catal st M. Nagai, Y. Goto, H. Sasuga and S. Omi

439

Poisoning Industrial Evaluation of Selective Hydrogenation Catalyst Poisoning B. Didillon, J. Cosyns, C. Cameron, D. Uzio, P. Sarrazin, J. P. Boitiaux

447

The Mechanism of Metal Poisoning by Cyclic Deactivation in Fluid Cracking Catalysts F. Hern6ndez, R. Garcfa de Le6n, E. M6gica, J. C. Moreno, R. Gonz61ez and E. Garciafigueroa

455

AFM and XPS Studies of Thiophene and 1-Butanethiol Deactivation of Pd/AI~O, Model Catalysts During 1,3-Butadiene Hydrogenation K.-H. Lee, R. Catani, R. Miglio and E. E. Wolf

463

Sulfur Poisoning of Nickel-based Hot Gas Cleaning Catalysts in Synthetic Gasification Gas J. Hepola and P. Simell

471

Effects of H~S on Bifunctional Catalysts C. Flego, L. Galasso, S. Vidotto and G. Faraci

479

Performance of Ni/AI~O, Pellets Poisoned by Thiophene D. Rusic and S. Zrncevic

487

Study of the Simultaneous Deactivation by Coke and Sulfur of Naptha Reforming Catalysts Using a Bifunctional Test Reaction A. Borgna, T. F. Garetto and C. R. Apesteguia

495

Syngas Conversion Deactivation and Attrition of Iron Catalysts in Synthesis Gas N. B. Jackson, A. K. Datye, L. Mansker, R. J. O'Brien and B. H. Davis

501

Temperature-Programmed Reaction Study of Carbon Transformations on Iron Fischer-Tropsch Catalysts During Steady-State Synthesis S. A. Eliason and C. H. Bartholomew

517

Deactivation of Iron-based Catalysts for Slurry Phase Fischer-Tropsch Synthesis A. P. Raje, R. J. O'Brien, L. Xu and B. H. Davis

527

Effects of Reduction and Regeneration Conditions on the Activity of CuOZnO Catalysts C. E. Quincoces, N. E. Amadeo and M. G. Gonz6lez

535

Effects of Promoters and Supports on Coke Formation on Pt Catalysts during CH, Reforming with CO~ S. M. Stagg and D. E. Resasco

543

Ni-based Catalysts for Methane Conversion with Carbon Dioxide. Catalyst Pellet Deformation Induced by Coking and Gasification L. A. Rudnitsky

551

Carbon Formation from Decomposition of CH, on Supported Ni Catalysts R. LCdeng, M. Barr~-Chassonnery, M. Fathi, O. A. Rokstad and A. Holmen

561

On Limitations of Regenerating an HZSM-5 Catalyst for the MTG Process A. T. Aguayo, A. G. Gayubo, J. M. Ortega, A. L. Mor6n and J. Bilbao

567

Thermal Degradation Asymptotic Behavior during Sintering of Supported Catalysts G.A. Fuentes and E. Salinas-Rodriguez

573

Sintering and Redispersion of Supported Metals: Perspectives from the Literature of the Past Decade C. H. Bartholomew

585

The Existence of an Unusual Reversible Deactivation Phenomenon Associated with Preferential Surface Segregation in Bimetallic Systems R. T. K. Baker, A. Chambers, C. Park and N. M. Rodriguez

593

Vaporization-Assisted Degradation of High Temperature Combustion Catalysts J. G. McCarty, K.-H. Lau and D. L. Hildenbrand

601

Deactivation by Sintering and Coking of Sol-Gel NiO-AI~O,-TiO~ Hydrogenation Catalysts J. C. Rodrfguez, T. Viveros and A. Monz6n

609

Sintering of Ni/SiO~ Catalysts Prepared by Impregnation and DepositionPrecipitation during CO Hydrogenation G. Martra, H. Swaan, C. Mirodatos, M. Kermarec and C. Louis

617

Techniques An Experimental Protocol for Studying Kinetics and Catalyst Deactivation: Application to Heptane Reforming on Pt-Re/AI~O, K. Liu, S. C. Fung, T. C. Ho and D. S. Rumschitzki

625

Characterizing the Framework Demetallation of Environmentally Relevant Zeolites Using IR, NMR and Neutron Diffraction Techniques M. T. Paffett, J. Szanyi, R. M. Jacubinas, K. C. Ott, R. Von Dreele, C. D. Hughes and W. L. Earl

639

xi Study of the Deactivation of an HY Zeolite Pellet Using 129XeNMR Spectroscopy and ~H NMR Imaging T. Domeniconi, J.-L. Bonardet, M.-A. Springuel-Huet, J. Fraissard and J.-M. Dereppe

647

AFM Study of Carbon Formation on a Manganese Oxide Catalyst N. Batina, L. M. Ioffe and Y. G. Borodko

655

Novel Regeneration Method for Deactivated Noble Metal Catalysts L. A. Camacho, C. Park and N. M. Rodriguez

665

Estimation of Reversible and Irreversible Coke by Transient Experiments M. Larsson, N. Henriksson and B. Andersson

673

Deactivation Studies of Ni/AI~O, ZSM-5 Catalysts: Effect of Nickel Incorporation J. R. Grzechowiak, A. Masalska, L. Kepinski and J. Rynkowski

681

Author Index

687

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xiii

PREFACE We are pleased to present the Proceedings of the 7th International Symposium on Catalyst Deactivation held in Cancun, Mexico on October 5-8, 1997. This 3-day symposium continued in the tradition of the high-quality symposia on Catalyst Deactivation held previously in North America (Berkeley and Evanston) and Belgium (Antwerp and Ostend); it featured 7 plenary and award lectures, 60 oral presentations, and over 20 poster papers. Most of the papers presented at the symposium are contained in this volume. Catalyst Deactivation 1997 focused on 9 key topical areas: carbon deposition and coke formation, chemicals, environmental catalysis, modeling, petroleum processing, poisoning, syngas conversion, techniques, and thermal degradation. All of these areas were well represented at the meeting; moreover, several review articles were presented that provide perspectives on new research and development thrusts. The strong participation by authors and attendees provides evidence that the field of Catalyst Deactivation (despite its rather morbid connotation) is alive, well and growing. An innovation this year was the presentation of Founders Awards to Professors Alexis T. Bell, Benard Delmon, and Gilbert F. Froment who organized the first four symposia in this series and Maxted awards to Professor John B. Butt and Eugene E. Petersen for their important contributions to the field of Catalyst Deactivation. The proceedings of the meeting are organized with six review and award articles at the front of the volume followed by topical articles in alphabetical order by topic. Under each topical area a keynote, 5-6 oral, and 2-3 poster papers of 12-14, 8-10, and 6 pages each follow in that order. A list of authors is provided at the end of the book. It should be emphasized that all of the papers were ranked and reviewed by members of Scientific Committee, whose names are listed on the page following. In most cases authors were asked to respond with a revision of their paper. Papers were also carefully edited, and more than 20 papers were largely rewritten and retyped by the Editors and their staff. We are grateful to the members of the Scientific committee for their help in the difficult task of selecting from submitted contributions, the papers assembled here and for their dedicated efforts in reviewing and in some case partly editing the manuscripts. We also thank our secretarial/editorial staff (especially Melissa Ellis, Beth Graham, Allyson Kauwe, Miriam CanalLee, Matthew Reynolds, and Lance Woolley) for many hours spent in assembling and correcting manuscripts and our wives for their patience regarding time spent away from our families. CALVIN H. BARTHOLOMEW GUSTAVO A. FUENTES

Brigham Young University Universidad A. Metropolitana-Iztapalapa

xiv

SCIENTIFIC COMMITTEE MEMBERS DR. CARLOS R. APESTEGUIA Instituto de Investigaciones en Catalisis y Petroquimica (INCAPE) PROFESSOR DADY B. DADYBURJOR West Virginia University DR. SHUN C. FUNG Exxon Research & Engineering Co. DR. JAN J. LEROU E. I. Du Pont De Nemours & Co., Inc. DR. JON G. MCCARTY SRI International PROFESSOR JACOB A. MOULIJN Delft University of Technology DR. JAMES. J. SPIVEY Center for Process Research, Research Triangle Institute PROFESSOR EDUARDO E. WOLF University of Notre Dame

9

Elsevier Science B. V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

R o l e s o f Acidity and Pore Carbonaceous Deposits

Structure

in the

D e a c t i v a t i o n o f Zeolites by

M. Guisnet, P. Magnoux and D. Martin Laboratoire de Catalyse en Chimie Organique, CNRS - Universit6 de Poitiers UMR 6503, Chimie 7, 40, avenue du Recteur Pineau, 86022 Poitiers Cedex, France Tel.: (33) 5.49.45.39.05 - Fax: (33) 5.49.45.34.99- E-mail: [email protected]

Deactivation of zeolite acid catalysts during hydrocarbon transformations is mainly due to the formation and the retention inside the pores of heavy secondary products. The effects of pore structure of zeolites and of their acidity on the rate of formation of these carbonaceous compounds, on their composition and on their deactivating effect are examined. The roles of strength and density of active acid sites (often protonic sites) are generally limited in comparison to the roles played by size and shape of cavities (or channel intersections) and by the size of their apertures. The formation of carbonaceous compounds (coke) and the deactivation that they cause are clearly shape selective processes. Their formation occurs through a nucleation-growth pathway. In the range of temperatures of refining and petrochemical processes nucleation is due to trapping in the cavities of coke precursors. Condensation and hydrogen transfer reactions are involved both in the formation of coke precursors and in their growth. Four modes of deactivation can be distinguished : (1) limitation or (2) blockage of the access of the reactant to the active sites of a cavity in which one coke molecule is located, due to steric reasons or to a competition for adsorption between reactant or coke molecules and (3) steric limitation or (4) blockageof the access of the reactant to the active sites of cavities or of channels in which no coke molecule is located. The deactivation of monodimensional zeolites and of zeolites with trap cavities (large cavities with small apertures) only occurs through Modes 3 and 4, hence is very rapid. With the other types of zeolites, deactivation passes successively from Mode 1 to Mode 4 as the coke content increases; Modes 3 and 4 are due to the coke molecules which overflow onto the outer surface of zeolite crystallites.

1. INTRODUCTION Zeolites have replaced corrosive and polluting acids (H2SO4,A1C13, chlorinated alumina) or less acidic solids (e.g. silica alumina) as catalysts in many refining and petrochemical processes [ 1-2] and should substitute for them in the synthesis of fine and specialty chemicals [3-5]. The choice of zeolite catalysts is firstly due to their remarkable acid properties. Indeed the density and the strength of their acid sites can be varied on a large scale and can thus be adjusted to the desired catalytic reactions. Redox components can be easily introduced in the pores of acid zeolites in a well dispersed form, which leads to bifunctional catalysts which are of great interest for catalyzing transformations requiring successive acid and redox sites

(hydroisomerization, hydrocracking, etc). The high thermostability of zeolites is another characteristic which renders them particularly attractive for processes requiring repetitive regeneration steps at high temperatures such as FCC. Another major advantage of zeolites is their well defined pore structure, with apertures and cavities of approximately the size of organic molecules. Therefore, the selectivity of these molecular sieves depends not only on the intrinsic properties of the active sites but also on the pore structure. This so called shape selectivity of zeolites has stimulated research on the synthesis of new molecular sieves and brought about the development of various commercial processes, e.g. selectoforming, dewaxing, toluene disproportionation into paraxylene, etc. [6]. Separation processes based on these molecular sieve properties of zeolites (e.g. separation of iso and n-alkanes) have also been developed [7]. There are different reasons for deactivation of zeolite catalysts during the commercial processes: a. Poisoning of the active sites either by feed components and impurities or by carbonaceous deposits, b. Limited access of the reactant to the active sites due to partial or complete blockage by carbonaceous deposits or by extraframework species resulting from dealumination etc., c. Framework alteration, d. Sintering of supported metals (in bifunctional catalysts). Carbonaceous deposits are the principle cause of zeolite deactivation in processes involving hydrocarbon conversions. Firstly they can poison active sites or block their access. Secondly, their removal, which is carried out through oxidative treatment at high temperatures, has detrimental effects i.e., dealumination and degradation of the zeolite and sintering of supported metals. In most commercial processes the cost of catalyst deactivation is very high. Hence, facilitating catalyst stability and optimizing regeneration have become at least as important as controlling the activity and the selectivity. Therefore, while industrial laboratories try to find technical solutions, academic laboratories should establish the conceptual background indispensible to the understanding of the concomitant problems. This latter activity is largely carried out in the case of zeolite catalysts owing to the development of an efficient method for determining the composition of the carbonaceous deposits responsible for their deactivation [8]. Indeed, contrary to what is found with conventional catalysts, the carbonaceous compounds, being formed in pores of molecular size, are not very bulky; hence they can be identified and their distribution quantitatively established by GC/MS coupling. The aim of this paper is to show how pore structure and acidity of zeolites determine the rate of formation of carbonaceous deposits, their composition, their location and their deactivating effects. The characteristics of active acid sites and of the pore structures of the zeolites used in this work will be presented first.

2. THE ORIGIN OF ACTIVITY AND SELECTIVITY OF ACID ZEOLITES 2.1. Acid sites and activity In the transformation of hydrocarbons via acid catalysis, the activity, stability and selectivity of zeolites are obviously determined to a large extent by the characteristics of the acid sites, i.e. chemical structure, strength, density and accessibility. For skeletal

transformations of hydrocarbons, the rate depends essentially on the Br~insted acidity [9], as shown by the good correlations found between the concentration of Br6nsted acid sites and the rate of various reactions, e.g. cumene dealkylation [ 10], xylene isomerization [ 11 ], toluene and ethylbenzene disproportionation [ 13], and n-hexane cracking [ 14]. Apparently, isolated Lewis acid sites are not active in these reactions. However it is well known that Lewis acid sites located in the vicinity of protonic sites can increase the latter's strength and consequently their activity [15]. Thus in HEMT zeolites (Figure 1), an IR band corresponding to OH groups with very strong acidity appears at 3600 cm 1 in samples mildly dealuminated by steaming. These hydroxyl groups would correspond to a zeolite OH of high vibrational frequency interacting with cationic extraframework aluminium (EFAL) species located in sodalite cages. The corresponding protonic acid sites are at least 3 times more active for mxylene transformation than those which do not interact with EFAL species [16]. The acid sites (Lewis and/or Br~nsted) and base sites, e.g. oxygen atoms of the framework, can participate together (acid-base bifunctional catalysis) in various reactions such as cis-trans and double-bond-shift isomerization of alkenes [9]. While the OH groups in amorphous aluminosilicates are terminal, those of zeolites are primarily bridging, which explains their stronger acidity [17]. Furthermore, it is well recognized that the number of neighboring aluminium atoms (hence of OH groups) determines the strength of a given acid site, the strongest acid sites corresponding to isolated OH groups [18]. The minimum strength that the acid sites must have in order to be active depends very much on the reaction considered [9]. Generally the more difficult the reaction (hence the lower its rate) the greater the minimum acid strength required for its catalysis [19]. Thus, the difficult n-hexane cracking requires very strong acid sites (able to retain pyridine adsorbed above 520~ while the very facile skeletal isomerization of 2,2- and 2,3-dimethylbutenes can be catalyzed by very weak acid sites (able to retain pyridine adsorbed above 220~ [19]. Another parameter, the distance between acid sites, could play a role in bimolecularreactions (e.g. hydrogen transfer [20,21 ]), most probably because their catalysis involves several acid sites. The location of protonic sites in the zeolite pores can also affect the catalytic activity. Indeed, some of the sites can be inaccessible by the reactant (e.g. OH groups located in the sodalite cages of Y and EMT zeolites) or accessible with difficulty. 2.2. Pore structure of zeolites Zeolite structure types are described in the well-known Atlas of Zeolite Structure Types [22]. They can be classified into three categories depending on the number of oxygen atoms (or of Al+Si) in the pore apertures. Small pore zeolites (openings of 3.5-4.5 A formed by tings of 8 oxygen atoms) such as A zeolite (LTA) and erionite (ERI) can sorb only linear aliphatic molecules. Those with average pore size (4.5-6.0 A, 10 oxygen atoms) such as ZSM5 (MFI) can sorb branched aliphatic compounds and alkyl monoaromatics while those with large pores (6.0-8.0 A, 12 oxygen atoms) such as zeolites Y (FAU) and EMT can accept very bulky compounds. Zeolites with more than one pore system are classified according to the aperture of the largest pore e.g. offretite (OFF) is classified as a large pore zeolite.

2.9 x 5.7 A

Iii~~ i i!ii!i~ lii iiiiiilI

~~--~ 6.3 A

615 x 7.0 A

Ii. . .i. .~1 lii!!!!iiii . . . .!!! ... I

l I i!iiii!i!!ii I liii~!!!i!ii!!iiiiiiiiiiI t!iiil!iiiiir!i!lli!iiiiiiiii~

a) FAU (Y)

b) MOR

d) EMT

hypercage

c) OFF

hypocage

13A

e) M H Figure 1

f) ERI

g) LTA (A)

Schematic representation of the pore structure of various 'zeolites with large pore apertures 9FAU (Y), MOR, OFF, EMT, with average pore apertures: MFI (ZSM5), with small pore apertures: ERI, LTA (A).

Other structural parameters can play an important role in catalytic reactions: a. The existence of cages or channel intersections (in which the acid sites are located) and their size and shape. See for instance in Figure 1 the supercages of Y zeolites, the hypercages and hypocages of EMT zeolites, the large cages of ERI, and the channel intersections of MFI. b. The dimensionality of the pore system accessible to a given molecule. The pore system of Y zeolites is tridimensional, that of MOR is monodimensional for organic molecules (they can access only to the large channels) and bidimensional for nitrogen and ammonia which have access to the large and small channels, c. The existence of secondary porosity such as mesopores created during the dealumination of zeolites. Thus mesopores created in the crystallites of mordenites can render tridirectional the circulation of organic molecules [23]. As most of the acid sites are located in pores of molecular size the rate and the selectivity of catalytic reactions depend not only on the intrinsic properties of the sites but also on the pore structure. A zeolite catalyst selects the reactant or the product by their ability to diffuse to and from the active sites (reactant and product selectivity). Steric constraints in the environment of the sites limit or inhibit the formation of intermediates or transition states (restricted transition state selectivity) [24,25]. The strong polarizing interaction between zeolite crystallites and adsorbed molecules leads to an unusually high concentration of the reactants in the pores. This concentration effect causes an enhancement of the rates of bimolecular reaction steps over monomolecular reaction steps [26].

3. RATE OF F O R M A T I O N OF CARBONACEOUS DEPOSITS Carbonaceous deposits result from the transformation of reactants, reaction products, impurities of the feed, etc., on acid sites through various successive bimolecular steps: condensation, hydrogen transfer, etc. Therefore, their rate of formation depends on the following parameters which usually affect the rate of catalytic reactions, namely: a. The nature of the reactants and products, b. The operating conditions: temperature, pressure, concentration of the reactants, c. The characteristics of the acid sites: chemical structure, strength, density and number, d. The characteristics of the zeolite pore structure and especially the size and shape of the cavities (or channel intersections) in which are located the acid sites and which can therefore be considered as mieroreaetors. However, carbonaceous compounds have the peculiarity of being non desorbed products, hence must be retained inside the zeolite pores or on the outer surface of the crystallites. This retention is due to the strong adsorption of the carbonaceous compounds on acid sites and/or to their low volatility. It may also be due to their molecular size being greater than the pore aperture (causing trapping in the cavities or at channel intersections). Therefore, the rate of formation can depend on various parameters such as 9 a. The strength of the acid sites. The stronger the acid sites the less reversible the adsorption of the carbonaceous compounds, b. The characteristics of the diffusion pathway of organic molecules. The longer and the narrower this diffusion pathway the more difficult the desorption of bulky carbonaceous compounds,

c. The difference between the size of the cavities and the size of the pore apertures. Indeed the greater the difference the easier the trapping. However, while from these general considerations it is clear that both acidity and pore structure of a zeolite affect the rate of formation of carbonaceous compounds (other factors being held constant), it is generally impossible to quantify the effect of each of these parameters because of the difficulty in obtaining zeolite samples with identical acidities and with different pore structures or vice-versa. Having said that, some examples are given below which illustrate within these limitations the respective roles of acidity and of pore structure. Effects of acidity on coking rates have been generally investigated using a series of large pore zeolite catalysts: HY [27,28], HEMT [29], HMOR [30], cracking catalysts [31,32], etc., with different framework Si/A1 ratios. For example, the initial rates of coke formation from propene were determined over a series of commercial HY samples (Table 1) under the following conditions : fixed bed reactor, 450~ Ppropcne= 30 kPa, Pnitrogen = 70 kPa, 0.72 g of propene introduced per gram of zeolite per hour. With the more dealuminated samples (framework Si/A1 ratios from 16 to 100) the coking rate is roughly proportional to the density of the protonic sites (Figure 2) with turnover frequency (TOF) values in the range of 6 - 8 h 1. This quasi constant value of TOF can be related to the similarity in strength of the isolated OH groups of these zeolites [18]. The slightly greater TOF value found for the most dealuminated sample could be due to the large amount of the extraframework aluminium atoms (Table 1) in this sample [27]. Values three times lower are found for the non isolated sites (hence weaker sites) of the less dealuminated samples. Inaccessibility of part of the protonic sites (located in the sodalite cages)of these slightly dealuminated zeolites [16] is also responsible for the low TOF values.

Table 1 Characteristics of the commercial HY samples (from PQ zeolite) used in the formation of coke from propene at 450~ [27]. Unit cell formula

NEFAL/U.C.

700

Nao.4H38.sA139.2S i 152.80384

9.6

4.3

600

N ao.4H35.8A136.2S i 155.80384

6.1

16.5

560

Nao.3Hll.oAlll.3Si180.70384

2.4

Name

(Si/A1)fram.

CBV 500

3.9

CBV 500 D CBV 720

Na (ppm)

CBV 740

19.3

100

N ao.o5H9.45A19.5S i 182.50384

0.1

CBV 780

30

300

Nao.15H6.05A16.28i185.80384

0.0

Nao.5nl.4All.9Sil90.10384 10.0 CBV 760 100 1000 (Si/A1)f~am.: atomic framework Si/AI ratio, NEFAL/U.C.:number of extraframework aluminium atoms per unit cell.

7

..--00

/

0 Figure 2

'

I

10

'

;

"

;

'

;

20 30 40 Nn+/U.C. Coke formation from propene at 450~ over a series of commercial HY zeolites [27]. Initial rates of coke formation as a function of the number of protonic sites per unit cell (Nn+/u.c.).

Similar results were obtained for a series of HEMT samples prepared by steaming of a NH4EMT zeolite (framework Si/A1 ratios from 4 to 70). However under identical conditions the TOF values for coke formation were found to be 3 times higher with HEMT samples than with HY samples [29], which can be related to the stronger acidity of the protonic sites of the HEMT zeolites [ 16]. This large difference in TOF values could also be due to differences in the pore systems of Y and EMT zeolites. However this seems unlikely, as the sizes of the pore apertures are quite similar; moreover, a lower rate of coking would be expected with EMT owing to the presence of hypocages (Figure 1). On the other hand, in the case of 5A zeolites the very rapid formation of coke is due to the presence of trapping cavities (0~ cages of 11.4 A with pore apertures of only 4.2 A (Figure 1)). Thus, coke formation from propene at 350~ is 20 times faster with a 5A zeolite [33] than with an HZSM5 zeolite (Si/A1 = 15) [34], despite the fact that the acid sites of this latter zeolite are much stronger than those of the 5A zeolite, i.e., HZSM5 retains ammonia adsorbed above 400~ whereas in the case of 5A zeolites no ammonia is retained above 300~ Coke formation from compounds such as alkanes, which do not lead directly to coke was also investigated. As coking is successive to the main reaction the selectivities (e.g. the ratio of rates of coking and of the main reaction) are important to consider as well as conversion. This was carried out [28] for n-heptane cracking at 350~ over two series of dealuminated Y samples, the first one prepared by steaming of a NH4NaY zeolite (hence with a large amount of extraframework aluminium (EFAL) species), whereas the second resulted from treatment with a dilute HC1 solution of the samples of series 1 (in order to eliminate the EFAL species without modifying the framework composition). Figure 3 shows that for both series the selectivity to coke passes through a maximum for a certain value of the theoretical concentration of protonic sites. EFAL species have a promoting effect on the selectivity, particularly with the less dealuminated samples. This promoting effect is most likely due to an enhancement of the acid strength of framework Brrnsted sites owing to their interaction

with the EFAL species. Indeed a new IR band appears at 3600 c m "1 in the less dealuminated samples [ 16], which corresponds to very strong acid sites (able to retain pyridine adsorbed above 450~ This confirms the positive effect of the acid strength of zeolites on the rate of coke formation. n-Heptane cracking was also carried out at 450~ over four protonic zeolites with different pore structures [8]. The degree of protonic exchange was chosen so that the initial cracking activities were similar. The initial coking/cracking rate ratio was found to be close to 1 for HERI and HMOR, 4 times lower for HY and 1,000 times lower for HZSM5. The large selectivity to coke of the first two zeolites was attributed to rapid blockage of coke precursors in their trap cavities (HERI) or in their monodimensional channels (HMOR) [8]. On the other hand the very slow rate of coking found with HZSM5 cannot be explained by the single effect of its pore structure. This pecularity is due rather to the 5 to 8 times lower density of these acid sites. Indeed, as was indicated in Section 2 the rate of bimolecularreactions such as those involved in coke formation (especially hydrogen transfer) is extremely dependent on the acid site density. In agreement with that observation, close values of the coking/cracking ratios are found with HZSM5 and with a dealuminated HY zeolite having similar densities of strong acid sites [8].

16, 12-

d'8-

~

,,

0

'

0

Figure 3

l

10

r

'

s o

;

"

;

'

;

'

:

20 30 40 50 Nn+/U.C. Coke formation during n-heptane cracking at 350~ over dealuminated HY zeolites. Ratio (mk/rnc) of the amounts of n-heptane transformed into coke and into desorbed products for 70 min reaction as a function Of NH+/U.C., the number of protonic sites per unit cell. (a): samples with a large amount of extraframework aluminium atoms; (b): samples with a small amount.

However, while acidity can play an important role in the selectivity to coke, the results obtained by Rollmann and Walsh [35] during the transformation of an equimolar mixture of nhexane, 3-methylpentane, 2,3-dimethylbutane, benzene and toluene over a large series of zeolites demonstrate that in this case the size of the pore apertures is the determining parameter. Indeed, the selectivity to coke (expressed by the ratio between the amount of coke

formed and that of alkane converted) is approximately 1 with all large-pore zeolites, 0.1 with medium pore zeolites and 0.05 with small-pore zeolites [35].

4.

COMPOSITION FORMATION

OF

CARBONACEOUS

DEPOSITS

AND

MODE

OF

The effects of zeolite structure and acidity on "coke" composition are more pronounced at high rather than at low reaction temperatures. 4.1. Low temperature coke Thus, at 120~ irrespective of zeolite composition and structure, the H/C ratio of carbonaceous compounds is close to that of reactants, which indicates that these compounds result mainly from condensation reactions. At 120~ the carbonaceous compounds formed on HY, HZSM5 [34] and 5A [36] zeolites from propene, are branched aliphatic or alicyclic hydrocarbons (no aromatics are observed). Olefin oligomerization is the main reaction involved in their formation. Most of these compounds are located inside the zeolite pores, as less than 5 % of them can be recovered through a direct soxhlet extraction of the coked zeolite samples with methylene chloride. With large-and intermediate-pore zeolites (HY, HZSM5) carbonaceous compounds are retained inside the zeolite pores because of their low volatility (their boiling point is much higher than the reaction temperature (Table 2)), whereas, with the 5A zeolite the retention of these branched compounds is due to their steric blockage in the large ct cages (11.4 A) with small apertures (4.2 A).

Table 2 Composition of the carbonaceous deposits formed from propene at 120~ with HY, HZSM5 and 5A zeolites. Atomic H/C nr ni +

Cy

Branching Size (A)

USHY

HZSM5

5A

1.8

2.0

2.0

25-40

10-35

6, 9, 12, 15, 18,21

4, 3, 2,1 ,0

2, 1, 0

2, 1, 0

1.5

0.25

1.1

6.0 x 25-40

4.3 x 10-35

6 x 9-25

bp76o (~ 350-530 175-490 60-340 no: number of carbon atoms; ni + Cy" number of unsaturated bonds + cyclic compounds; Branching: ratio between the number of protons of CH3 groups and that of CH2 + CH groups.

Differences observed in the number of carbon atoms, in the extent of branching and in the chemical nature of carbonaceous compounds trapped in the zeolite pores result mainly from differences in size of the cavities (or of the channel intersections) and from the differences in the strength or in the density of the acid sites. Indeed, differences in pore structure combined with differences in acidity affect the rate of the main reaction (oligomerization) relative to side

10 reactions, e.g. cracking, isomerization and hydrogen transfer. As is the case with coking rates it is often difficult to discriminate between the effects of the pore structure and acidity, both having at times the same effect. Thus, the lower rate of hydrogen transfer on HZSM5 compared to HY (compounds with more than one unsaturated bond or cyclic compounds are not observed with the former zeolite) can be explained either by the smaller size of the space available near the acid sites: (8.5 A for the channel intersections of HZSM5 against 13 A diameter for the Y supercages) or by the 10 fold lower density of the acid sites for HZSM5. The explanation of differences in reaction and coking rates for 5A zeolite and other zeolites is more simple. In this case the quasi absence of secondary transformations of propene oligomers (no cracking and practically no isomerization and hydrogen transfer) on 5A is clearly related to the very weak acidity of this zeolite, while the presence inside the pores of 5A of relatively volatile compounds (e.g. C6 branched alkenes) is due to the impossibility of desorbing these branched compounds from this small pore size zeolite.

4.2. High temperature coke While at low temperatm'es the composition of carbonaceous deposits formed in acid zeolites is extremely dependent on the reactant, this is no longer the case at high temperatures (i.e. the temperature range of most of the refining and petrochemical processes). Thus, with HZSM5 at 450~ methylpyrenes are the main components of the coke (1.5-3.0 wt %) formed from propene, toluene and a propene-toluene mixture. However, different reaction pathways are involved in the formation of methylpyrenes from these different reactants: Pathway 1 from toluene and Pathway 2 from both the toluene-propene mixture and propene [37]. The fact that with both pathways methylpyrenes are the main coke components, suggests that the pore structure rather than the acidity plays a determining role in the growth of coke precursors.

Pathway

Pathway 2:

1:

2C3 =

ol

~ C6=

~

isom

AaJk~C,

HT

A alk: aromatic alkylation ; DC: dehydrogenative coupling ; trans: transalkylation ; isom: isomerization ; HT: hydrogen transfer ; ol: oligomerization ; alk: alkylation ; cycl: cyclization.

The essential role of the pore structure was confirmed by investigating as a function of time-on-stream the formation of coke during n-heptane cracking at 450~ over various zeolites (HY, HOFF, HMOR, HZSM5 and HERI), with similar initial cracking activities. C3 and C4 were the major products for all the zeolites, the olefin/alkane ratio being smaller than 1, which indicates secondary transformations of olefinic cracking products into non-desorbed products (coke): coking is consecutive to n-heptane cracking.

11

The composition of coke depends both on time-on-stream (hence on the coke content) and on the zeolite [8]. With all the zeolites carbonaceous compounds are initially formed inside the zeolite pores and, except for mordenite, are soluble in methylene chloride after dissolution of the zeolite in a hydrofluoric acid solution. The yield in soluble coke decreases with increasing coke content. The analysis of the soluble part of coke by GC/MS shows that the size and the degree of aromaticity of the coke components increase with coke content (see for example in Table 3 the change in the amount of the various coke families found with HZSM5 as a function of coke content).

Table 3 Distribution of the main families of coke components (wt %) during n-heptane cracking at 450~ over HZSM5 as a function of total coke content. (wt %) 1.0 3.5 4.5 7.0 Coke families - ~

~

~

'--Rx

0.7

(CnH2n-6)

0.3

(CnH2n-14)

~

~---R

3.0

(CnH2n-12)

~

R

0.5

(CnH2n-16) ~ R (CnH2n-22) Insoluble coke

-

-

1.0

0.5

2.2

1.9

1.3

4.6

For similar coke contents the composition of coke strongly depends on the pore system of the zeolite, i.e., coking is a shape-selective process. At low coke contents, the major components are alkylbenzenes in HZSM5 and HERI, alkylnaphthalenes or fluorenes in HOFF, alkylbenzopyrenes in USHY and very polyaromatic compounds (insoluble in methylene chloride) in HMOR. At high contents, the coke components are polyaromatic in all zeolites. However, the molecules of components soluble in methylene chloride are very different as they take the size and shape of the cavities or of the channel intersections [38], as is shown in Figure 4. Whatever the coke content soluble coke molecules are trapped in cages

12 (or at channel intersections). Indeed their molecular size is between that of the pore apertures and the cages [8].

i

b)

Figure4

Location of the main coke molecules formed at high coke contents in the pores of a) HZSM5 (methylpyrene) and b) ERI (methylchrysene).

With all the zeolites (except mordenite), insoluble coke molecules result from the condensation of soluble coke molecules. Indeed insoluble coke molecules are not directly formed and the amount of soluble coke molecules passes through a maximum (Figure 5). Coke formation occurs therefore according to the following scheme :

n-heptane ~

1

2 alkenes ~ + alkanes

soluble coke ~

insoluble coke

Alkenes resulting from n-heptane cracking (Step 1) are transformed into soluble coke molecules which are trapped in cages and pores. Soluble coke molecules are transformed into

13 insoluble molecules(Step 3), that according to TEM analysis [39], overflow onto the outer surface of the zeolite crystallites and into the inner mesopores created by dealumination. Several reactions: alkylation, cyclization and hydrogen transfer could be involved in Steps 2 and 3.

n k

1.6"

la) HYI

O 9

9

ln~ Ib)HZSM5I 0.8--

1.20.60.8A

w

.4"

v

0.4-

0.2"

0, 0.0r 0

Figure 5

9

o

o

~._J

n

z~

9149

i

n

n m

8

n

m m

m

mm

12 %C 16

0.

'

~ 0

'

' 2

4

: ' 6 %C ~

Change in the concentration of soluble (e) and insoluble (O) coke molecules (nk, 1020g"1) as a function of the percentage of coke formed during n-heptane cracking at 450~ on HY (a) and on HZSM5 (b).

However, the existence of another mode of formation of insoluble coke molecules, the dehydrogenative coupling of polyaromatic species located in adjacent cavities has recently been proposed to explain the high selectivity to insoluble coke of HY [27] and HEMT [29] zeolites with high framework Si/A1 ratios (hence with low acid site densities and low coking activities (Figure 6)). This dehydrogenative coupling of soluble coke molecules was demonstrated by the decrease in the H/C ratio of coke and by the large increase in the percentage of insoluble coke caused by the treatment of a coked HY sample with a nitrogen flow at the coking temperature (450~ for a long time (6 hours). The relative significance of the two modes of growth of coke molecules is determined by the residence time of coke molecules in the zeolite pores. Indeed, the longer this time, the greater is the probability of coupling between the molecules of coke or coke precursors, which explains the high selectivity to insoluble coke of zeolites with low coking activities [38]. This dehydrogenative coupling could also explain the very high selectivity to insoluble coke found with HMOR, the residence time of coke precursors in the channels of this monodimensional zeolite being very long [8]. When different types of zeolite pores are accessible to the reactant, the formation of coke (reaction + retention)can occur preferentially in one of them. This was shown in the case of coking from n-heptane on H-Offretite [40], a zeolite consisting of rectilinear cyclindral channels (diameter 6.3/~) interconnected through gmelinite cages (apertures of 3.6 x 5.2 A

14 25

28

-N

,o

HY 4 %

24-

I

HEMT 4

20""

-

,c %

20,,-,,,i

~15-or.,r HY 100 n .=. ~10-

ID

216-

0t

O

;s12 . 73

9

5"

4 | |

0

Figure 6

5

10

15

|

%C

| |

20

|

9

0d

25

,

0

.-j

4

: . : .....

8

:.

12 16 20 24 28 %C

Formation of insoluble coke from propene at 450~ over (a) HY and (b) HEMT zeolites with low and high framework Si/A1 ratios as a function of the total amount of coke.

diameter). The first molecules of coke which are formed are alkylnaphthalenes and fluorenes, (Families 1 and 2), i.e. non bulky and highly volatile compounds which, consequently, cannot be retained in the large channels. These coke molecules are therefore formed and trapped in the grnelinite cages (Figure 7). Table 4 shows that they transform into alkylphenanthrenes and alkylchrysenes (Families 3 and 4) and then into insoluble coke molecules. The compounds of Family 4 (Figure 7), too bulky to be entirely located in the gmelinite cages, overflow into the large channels. This location of coke molecules was confirmed by comparing the adsorption of n-hexane which has access to both gmelinite cages and large channels, to that of methyl-pentane which has access only to the large channels [41 ].

5. D E A C T I V A T I N G

EFFECTS

OF COKE MOLECULES:

MODES

OF

DEACTIVATION Thanks to the method developed for establishing the composition of coke it is possible to express the deactivation as a function of the number of coke molecules, which obviously gives more information on the mode of deactivation than an expression involving a function of total coke content. The deactivating effect of the coke molecules (their toxicity) depends very much on zeolite acidity and structure, even when the initial activities of the zeolite are similar. Thus, during n-heptane cracking at 450~ the number of coke molecules which causes complete deactivation was found for HY and HZSM5 to be close to the number of strong acid sites (i.e. active sites [8]) while it was 7-8 times smaller for HMOR and HERI (Table 5). Therefore deactivation of HMOR and HERI is probably due to blockage (by coke) of the access of the

15 Table 4 Formation of carbonaceous deposits (coke) during n-heptane cracking at 450~ over HOFF. Distribution of the main families of coke components as a function of the coke content. (wt %) ~ e 0.8 1.5 2.3 4.5 10 Coke families ~~---

R

0.4

(CnH2n-12)

~

R

0.4

(CnH2n-16)

~

R

0

1.0

1.0

1.0

1.0

0

0.5

0.6

1.0

4.5

0

0

0.7

2.5

5.0

(CnH2n-18)

~

R

(CnH2n-24)

Insoluble coke

Stv--. --

j

l gmelinit~ cage

9 large annels

a) b) c) Figure 7 Formation and growth of coke molecules in the pores of HOFF during n-heptane cracking at 450~ (a) naphthalene, (b) methylphenanthrene and (c) methylchrysene.

16 reactant to the acid sites. This blockage is furthermore confirmed by adsorption experiments

[8]. Table 5 Toxicity of coke molecules: number of strong acid sites that one coke molecule renders inactive. A: initial values ; B: values at complete deactivation. A

B

USHY

5

1.3

HMOR

25

7.5

HMFI

0.25

1.0

HERI

30

7.0

The effects of pore structure on coke toxicity can also be quantitatively elucidated by comparing the number of sites that one coke molecule renders initially inactive (for this estimation it was supposed that all the acid sites had the same cracking activity): less than one for HZSM5 (0.25), 5 for HY and more than 25 for HMOR and HERI (Table 5). While the very high toxicity of coke molecules found for these last two zeolites is typical for pore blockage, the average value found for HY is most likely due to the heterogeneity in acid strengths of this zeolite. Coke molecules are preferentially formed on the strongest (hence the more active) acid sites and thus preferentially deactivate these sites [8]. The creation of a secondary porosity can significantly decrease the toxicity of coke molecules. This has been shown for the transformation of methanol into short-chain alkenes over HMOR catalysts i.e. the toxicity of coke molecules is more than 10 times lower in dealuminated samples, as mesopores created by dealumination allow a quasi-tridirectional diffusion of organic molecules [23]. This positive effect of mesopores was also observed in liquid phase alkylation of toluene with 1-heptene [42]. The wide-ranging effects of pore structure on coke toxicity lead us to define four modes of deactivation (Figure 8) instead of the two (site coverage and pore blockage) which are generally proposed. Thus, deactivation could be due to : (1) Limitation of the access of the reactant to the active sites of a cage or a channel intersection in which a coke molecule is located, (2) Blockage of this access, (3) Limitation or (4) blockage of the access of the reactant to the active sites of cavities, of channel intersections or of parts of channels in which no coke molecules are located. In Modes 1 and 2 the limitation or the blockage is due to chemical reasons i.e., the coke molecules are (1)reversibly or (2) quasi-irreversibly adsorbed on the acid sites (site poisoning or site coverage) and/or to steric reasons: the diffusion of reactant molecules through the cavity or the channel intersection is (1) limited or (2) blocked. With these modes the toxicity of the coke molecules is low as only the sites located in the cavity or at the channel intersection, often only one site, are partially (Mode 1) or totally (Mode 2) deactivated. In Modes 3 and 4 the limitation or the blockage is due only to steric reasons; the diffusion of reactant molecules to the acid sites of the inner pores is (3) limited or (4) blocked. With these modes the toxicity of the coke molecules is very high, as a large number of active sites are generally located in these inner pores. The effects of the pore structure on the toxicity of coke molecules can be explained by differences in the significance of the modes of deactivation from one zeolite to another : a. With tridimensional zeolites having no trap cavities (large cages with small apertures) deactivation at low coke contents occurs through Mode 1 then Mode 2 ; at high coke contents deactivation occurs through Modes 3 and 4, caused by coke molecules overflowing onto the outer surface of the crystallites (Figure 8). Therefore, when the acid sites are of similar strength as in the case of HZSM5, the greater the coke content the higher the toxicity of coke molecules.

17

.......

17

/ Mode 1

Figure 8

Mode 2

Mode 3 Mode 4 Schematic of the four possible modes of deactivation by carbonaceous deposits in HZSM5.

The situation is more complex when the acid sites are of different strength as is the case with HY zeolites. Indeed the strongest acid sites and hence the most active are the first deactivated with a consequently high toxicity of the coke molecules at low coke content [8]. b. With monodimensional zeolites such as HMOR or with zeolites containing trap cavities such as HERI the deactivation occurs only through Modes 3 and 4. In the case of HMOR, at very low coke contents, one coke molecule located in a large channel (most likely at one intersection of this channel with the narrow pores inaccessible to the organic molecules) is able to limit or block the access of the reactant to all the active sites located in the channel, with consequently a high toxicity of the coke molecules. This high toxicity is also observed with HERI because the coke molecules initially formed, trapped in the cages near the outer surface of the crystallites, limit or block the access of the reactant to the active sites of the inner cages [8]. 6.

CONCLUSIONS

The acidity and pore structure of zeolites play significant roles in their deactivation by carbonaceous deposits ("coke"). This is not surprising, as the formation of coke involves reactions catalyzed by acid sites located inside the pores and also requires the retention of coke molecules by adsorption on the acid sites or by condensation or by steric blockage in the pores. Although it is often difficult to estimate quantitatively and separately the impacts of the acidity and of pore structure, it is clear that it is the latter characteristic which plays the greater role. The development of a method for establishing the composition of zeolite coke has led to significant progress in the understanding of the modes of coking and deactivation of zeolites

18 and more generally of porous catalysts. At high temperatures (> 300~ coking can be considered as a nucleation-growth process. The key step is the trapping of heavy product molecules in the micropores (nucleation). This trapping occurs in cavities or at channel intersections. Obviously, this process is favoured more and more as the difference between the sizes of the pore openings and of the cages increases. Identical reactions are generally involved in the nucleation and growth steps i.e. auto or inter condensation of coke precursors such as alkenes, alkylaromatics and polyaromatics and hydrogen transfer. Four modes of deactivation can be distinguished. The first one corresponds to a limitation of the access of the reactant to the active sites of a cage or of a channel intersection in which is located a coke molecule, the second one to complete blockage of this access. This limitation or this blockage may be due to the adsorption of coke molecules on the acid sites or to steric reasons. The other two modes correspond to steric limitation (Mode 3) or blockage (Mode 4) of the access of the reactant to the active sites of cavities or of channels in which no coke molecules are located. The deactivating effects of coke molecules increase from Mode 1 to Mode 4. The relative significance of these modes of deactivation depends on the degree of coking and on the zeolite pore structure. The deactivation of monodimensional zeolites and of zeolites with trap cavities (large cages with small apertures) only occurs through Modes 3 and 4. These zeolites, which are consequently very sensitive to coke, could be used as catalysts only in processes in which coke formation is very slow (e.g. bifunctional catalysis). For other types of zeolites, at low coke content deactivation occurs through Mode 1 then through Mode 2, while at high coke contents Modes 3 and 4 are predominant owing to coke overflowing onto the outer surface of the crystallites.

REFERENCES 1. J.A. Rabo, Zeolite Chemistry and Catalysis, ACS Monograph 171, American Chemical Society, Washington, 1976. 2. I.E. Maxwell and W.H.J. Stork, in Introduction to Zeolite Science and Practice, H. van Bekkum et al. (Eds.), Studies in Surface Science and Catalysis, Elsevier, Amsterdam, Vol. 58 (1991) 571. 3. P.B. Venuto, Microporous Materials, 2 (1994) 69. 4. W.F. Hotlderich, in Zeolites : Facts, Figures, Future, P.A. Jacobs and R.A. van Santen (eds.), Studies in Surface Science and Catalysis, Elsevier, Amsterdam, Vol. 49 (1989) 69. 5. G. Perot and M. Guisnet, J. Mol. Catal. 61 (1990) 173. 6. N.Y. Chen, W.E. Garwood and F.G. Dwyer, in Shape Selective Catalysis in Industrial Applications, Marcel Dekker, Inc., New-York and Basel (1989). 7. J.A. Johnson and A.R. Oroska, in Zeolites as Catalysts, Sorbents and Detergent Builders, H.G. Karge and J. Weitkamp (eds.), Studies in Surface Science and Catalysis, Elsevier, Amsterdam, Vol. 46 (1989) 451. 8. M. Guisnet and P. Magnoux, Appl. Catal. 54 (1989) 1. 9. M.R. Guisnet, Acc. Chem. Res., 23 (1990) 392. 10.P.A. Jacobs, H.E. Leeman, J.B. Uytterhoeven, J. Catal. 33 (1974) 17. 11.J.W. Ward, R.C. Hansford, J. Catal. 13 (1969) 154. 12.T. Aonuma, M. Sato, T. Shiba, Shokubai 5 (1963) 274. 13.H.G. Karge, K. Hatada, Y. Zhang, R. Fiedorow, Zeolites 3 (1983) 13. 14.J.W. Ward, J. Catal. 10 (1968) 34. 15.C. Mirodatos and D. Barthomeuf, J. Chem. Soc. Chem. Commun. (1981) 39. 16. S. Morin, unpublished results. 17.J.A. Rabo and G.J. Gajda, in Guidelines for Mastering the Properties of Molecular Sieves ; D. Barthomeuf et al. (Eds.), NATO ASI Series B, Plenum Press, New-York and London, 221 (1990) 273. 18.D. Barthomeuf, Mater. Chem. Phys. 17 (1987) 49.

19

19. G. Bourdillon, C. Gueguen, and M. Guisnet, Appl. Catal. 61 (1990) 123. 20.W-C. Cheng and K. Rajagopalan, J. Catal. 119 (1989) 354. 21 .E. Jacquinot, F. Raatz, A. Macedo and Ch. Marcilly, in Zeolites as Catalysts, Sorbents and Detergent Builders, H.G. Karge and J. Weitkamp (eds.), Studies in Surface Science and Catalysis, Elsevier, Amsterdam, 46 (1989) 115. 22.Atlas of Zeolite Structure Types, Zeolites 12 (1992). 23.N.S. Gnep, Ph. Roger, P. Cartraud, M. Guisnet, B. Juquin and Ch. Hamon, C.R. Acad. Sci., 309 (1989) 1743. 24.P.B. Weisz, V.J. Frilette, R.W. Maatman and E.B Mower, J. Catal. 1 (1962) 307. 25.S.M. Csicsery, J. Catal. 19 (1970) 394 and J. Catal. 23 (1971) 124. 26.J.A. Rabo, P.D. Bezman and M.L. Poutsma, Proceedings of the Symposium on Zeolites, Szeged, Acta Physica et Chemica 24 (1978) 39. 27.K. Moljord, P. Magnoux and M. Guisnet, Appl. Catal. A : General 122 (1995) 21. 28.Q.L. Wang, G. Giannetto and M. Guisnet, J. Catal. 130 (1991) 471. 29. G.A. Doka Nassionou, Ph.D. Thesis, Poitiers (1997). 30.H.G. Karge, in Introduction to Zeolite Science and Practice, H. van Bekkum et al. (Eds.), Studies in Surface Science and Catalysis, Elsevier, Amsterdam, Vol. 58 (1991) 531. 31. K. Rajagopalan and A.W. Peters, J. Catal. 106 (1987) 410. 32.T. Ino, S. A1-Khataff, Appl. Catal. A: General 142 (1996) 5. 33.M. Misk, G. Joly, P. Magnoux and M. Guisnet, Zeolites 16 (1996) 265. 34.B. Dimon, P. Cartraud, P. Magnoux and M. Guisnet, Appl. Catal. A: General, 101 (1993) 351. 35.L.D. Rollmann and D.E. Walsh, J. Catal. 56 (1979) 139. 36. Y. Boucheffa, C. Thomazeau, P. Cartraud, P. Magnoux, M. Guisnet and S. Jullian, Ind. Eng. Chem. Research, accepted for publication. 37.P. Magnoux, F. Machado and M. Guisnet, In New Frontiers in Catalysis, L. Guczi et al. (Eds.), Proceedings of the 10th International Congress on Catalysis, Budapest (1993) 435. 38.M. Guisnet, P. Magnoux and K. Moljord, in Deactivation and Testing of HydrocarbonProcessing Catalysts, P.O'Connor et al. (Eds.), ACS Symposium Series 624, ch. 5 (1996) 77. 39.P. Gallezot, C. Leclercq, M. Guisnet and P. Magnoux, J. Catal. 117 (1989) 100. 40. P. Magnoux, M. Guisnet, S. Mignard and P. Cartraud, J. Catal. 117 (1989) 495. 41. S. Mignard, P. Cartraud, P. Magnoux and M. Guisnet, J. Catal. 117 (1989) 503. 42.P. Magnoux, A. Mourran, S. Bernard and M. Guisnet, Proceedings of the D.G.M.K. Conference "Catalysis on Solid Acids and Bases", J. Weitkamp and B. Lticke (eds.) (1996) 49.

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9 Elsevier Science B.V. All rights reserved Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

21

Impact of Sulfur on Three-Way Automotive Catalyst Performance and Catalyst Diagnostics D. D. Beck GM Research and Development Center 30500 Mound Rd. Warren, M148090 USA In an effort to reduce emissions from mobile sources to help address acute ozone nonattainment problems, the State of California adopted a Low Emission Vehicle/Clean Fuel program in 1990. Low Emission Vehicles, or LEVs, are designed to meet significantly stricter HC, CO and NOx standards than vehicles sold in the remainder of the U.S., and are designed to operate on a strictly controlled reformulated fuel to achieve low emissions over the vehicle's lifetime. One of the characteristics of this fuel is a low sulfur content, which averages 30-40 ppm or less. Data provided by the Auto Oil Air Quality Research Program and other studies were instrumental in the development of the low emission properties of this fuel, and vehicle manufacturers now use this fuel in the development, calibration, and certification process for Low Emission Vehicles sold in California. Such vehicles may also be introduced federally through a national low emission vehicle program or by individual states adopting the California program, but there is concern that the low emission benefits obtained in California may not be achieved federally because of the variability in fuel properties, including sulfur content, which may be present in concentrations as high as 1000 ppm. This has caused increased interest on the impact of sulfur on the performance of automotive catalysts, particularly catalyst systems which will be used in Low Emission Vehicles. The impact of sulfur on catalyst oxygen storage is also of interest, due to the requirement of an on-board catalyst performance diagnostic which uses the loss of catalyst oxygen storage as an indicator of catalyst deterioration. This paper will review both laboratory and vehicle studies performed at GM and elsewhere which focus on the impact of sulfur on the performance and oxygen storage capacity of catalyst systems which are anticipated to be used in low emissions vehicles, and notably catalysts which incorporate palladium, and compare the magnitude of these impacts to those observed on earlier generation automotive catalyst systems. This paper will also review how these observations are consistent with conclusions obtained from fundamental studies of the interaction between sulfur-containing gas species in vehicle exhaust and the catalytic and oxygen storage sites in the catalyst.

1. I N T R O D U C T I O N All commercially produced gasoline blends which are intended for vehicle use contain organo-sulfur compounds in concentrations ranging from several parts per million (ppm) to 1000 ppm [1]. These compounds are present in the crude oil feedstocks [2] and can be partially removed by a hydrodesulfurization process in refining [3,4]. Removal of nearly all of the sulfur from gasoline is not always economical due to a number of factors including the variability in feedstock properties and the cost of deep hydrodesulfurization (HDS)

22 processing (HDS catalysts must be periodically regenerated or replaced since they are deactivated by coke and sintering in use). An example of the variability in sulfur content of regular octane grade gasolines in different markets of the U.S. is shown in Figure 1. Fuels in some urban areas are now required to conform to stricter specifications under the Reformulated Federal Gasoline (RFG) program, but even under the more strict phase of this program to be implemented later, the sulfur concentration in these fuels can be as high as 500 ppm [5]. Thus, the sulfur content of RFG fuels will not be controlled to the same strict specifications of California Phase II fuels, which began distribution in March of 1996 and which require refinery product not to exceed 30-40 ppm and retail product not to exceed 80 ppm sulfur.

E"

1,400

-

I/

1,~00 . . . . . ~' 1,000- . . . . ~)

tT

High Average

t!'

L~ ....

1....................... ~---I ...................

800 . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .

8 ~

,.

o---.-t-.-'-i-,-.-,-/,-.-,-.-,-.-,-.-,-.-,l:.-,-i-,Figure 1. Sulfur levels in federal fuels (results of the AAMA fuel survey conducted during the Winter of 1994). Sulfur compounds in gasoline are converted during combustion primarily into sulfur dioxide, and given the proper conditions can further react over the emission control catalyst to form hydrogen sulfide, carbonyl sulfide or even sulfurous and sulfuric acid [6]. For threeway catalyst applications, the exhaust air/fuel ratio is controlled near the stoichiometric point such that carbonyl sulfide and the acids are not favored. Although emission of hydrogen sulfide has been noticeable with the increase in the ceria content in these catalysts over the last two decades, methods to suppress the emission of hydrogen sulfide have been successfully employed. Generally, though, sulfur dioxide is of an emissions concern because it has been demonstrated to deactivate vehicle exhaust catalysts. Laboratory and vehicle

23

studies were carried out to first investigate the effects of sulfur on Pt and Pd oxidation catalysts in the 1970's [6-8], and on Pt-Rh three-way catalysts from the late 1970's until present [9-26]. These studies and discussions of the mechanism responsible for sulfur poisoning of vehicle emission catalysts has been discussed previously [ 18]. When the State of California adopted requirements that vehicle manufacturers introduce Low Emission Vehicles (LEVs) in 1990, increased attention was placed on ways to achieve rapid activation of the catalytic converter including the use of Pd lightoff catalysts to comply with the very strict hydrocarbon standards for these vehicles. The use of Pd in such applications is considered attractive because of its ability to catalyze the oxidation of hydrocarbons at temperatures significantly below those of catalysts containing Pt, Rh or both, and because of the high temperature durability of Pd which is needed as these catalysts are close-coupled to the exhaust manifold for fast lightoff. However, earlier generation Pd catalysts have been found to be more severely affected by sulfur than other noble metal catalysts, leading to concerns of greater sulfur impacts on emission control systems on LEVs than on less strictly regulated Tier 0 or Tier 1 vehicles (Table 1). Therefore, the anticipated emissions benefit derived from the sale of LEVs federally might be at least partially compromised. However, continued developments in Pd catalyst formulations have improved their high temperature durability, as well as NOx activity and tolerance for sulfur poisoning [27-29]. These improvements, coupled with better optimization of the emission control system in low emission vehicles, offer hope that such vehicles may achieve in actual use the low emission standards for which they were designed. In this paper, then, we will discuss results of some investigations of the impact of sulfur on catalysts which are anticipated to be used in low emission vehicles, and compare the results with similar studies using earlier generation catalysts. We shall also discuss how these observations are consistent with what is known from fundamental studies about the interactions between sulfur and the individual catalyst components. Table 1 Emission Standards for Federal and California Low Emission Vehicles. (Passenger Car Category, Standards at 50,000 miles, in g/mi) THC Federal Tier 0 Federal Tier 1 California Transitional Low Emission (TLEV) California Low Emission (LEV) California Ultra-Low Emission (ULEV)

0.41 0.41

NMHC or NMOG*

CO

NOx

0.25

3.4 3.4

0.4

0.125

3.4

0.4

0.075

3.4

0.2

0.040

1.7

0.2

1.0

* Non-methane hydrocarbons or non-methane organic gases. In addition to the adoption of a Low Emission Vehicle/Clean Fuel program, the California Air Resources Board (CARB) has adopted regulations which call for comprehensive on-board monitoring of seven components and/or events related to emissions of exhaust, fuel, and air conditioning heat-transfer agents for vehicles sold in California. This legislation is called "On-Board Diagnostics-Phase II" or "OBD-II" [30]. These systems were installed on all vehicles sold in California in the 1996 model year. The Environmental

24 Protection Agency (EPA) has adopted similar regulations. One of the parameters to be monitored by the OBD-II system is catalytic converter performance. Performance which falls below an acceptable level defined by the regulations must be identified as malfunctioning by activation of an indicator lamp and storage of a fault or trouble code in the engine control module which can be read by a service technician through a serial link provided in the vehicle. When the fault code is set, it can only be cleared during servicing. The definition of a malfunctioning converter depends on the emissions category of the vehicle and has involved a gradual phase-in of standards, but generally requires identification of a failed converter when it exceeds 1.5x the applicable HC standard. Current catalyst diagnostic systems use conventional oxygen sensors placed in front of and behind each catalyst to infer catalyst efficiency from a measure of catalyst oxygen storage [31-33]. There has been much discussion about the advantages and disadvantages of this strategy, including continued discussion about the validity of the correlation between catalyst oxygen storage and catalyst performance [33]. For example, it has been demonstrated that non-Ce containing catalysts as well as close coupled small volume catalysts which have been operated at very high temperatures can be mis-diagnosed as failing by the dual oxygen sensor method when they in fact easily meet emission standards [34]. This is an example of an incorrect diagnosis event which could lead to an undesirable high number of warranty replacements. To minimize the risk of such events, the dual oxygen sensor method of catalyst monitoring continues to be further developed and improved to minimize incidences which can give rise to incorrect diagnosis of the converter. It is well known that organo-sulfur compounds present in the fuel can react with the ceria and alumina components of the catalyst to form sulfites and sulfates, [18 and references therein] thereby interfering with the ability of the catalyst to easily store or release oxygen. Laboratory studies have been conducted which indicate that the presence of sulfur dioxide in relatively low concentrations can result in near complete loss of oxygen storage capacity in production three-way catalysts [35,36], which can influence the OBD-II catalyst monitoring system. Under specific conditions, it may also be possible to poison the exhaust-side noble metal element in the post-catalyst oxygen sensor, which has been observed to retard the response of the sensor to rapid changes in air/fuel ratio. The net result from the inability of the post-catalyst oxygen sensor to quickly respond to these changes is that the signal is mistaken for that of a good converter, even if the emissions exceed the standard due to converter failure. Since the dual oxygen sensor method of catalyst monitoring depends on the measurement of oxygen storage capacity of the monitored catalyst or catalysts and any connection to actual HC conversion performance is inferred, any variable which can impact either the oxygen storage capacity of the catalyst and/or the measurement device is expected to impact the reliability of the method. This is of particular concern for low emission vehicle classes for which the OBD-II catalyst monitor must be able to detect small changes in the emissions performance in order to identify when the levels exceed 1.75x the applicable FTP standard. Using a low emission vehicle as an example, the OBD-II system must be able to detect and indicate failure of the catalyst system when the HC conversion efficiency falls below an average of--93% on an FTP test, assuming typical engine-out hydrocarbon levels [34]. Thus, any variable which may have a relatively small impact on emissions performance and/or the OBD-II catalyst monitor may potentially cause incorrect diagnosis. For a California LEV operating in California there are a number of factors which can influence these measurements. Fuel is not expected to be a significant factor, however, since the regulations for phase II reformulated gasoline should result in state-wide distribution of low sulfur content unleaded gasoline with relatively uniform properties. Conversely, for a California LEV operating outside California, fuel properties vary widely, including sulfur content, thus presenting a potentially significant influence on the emissions performance and the OBD-II catalyst monitor in these vehicles [37]. In this paper, then, we will also briefly discuss a number of laboratory and vehicle studies which address the extent and nature of the impact of sulfur on both catalyst performance and on OBD-II.

25

2. LABORATORY STUDIES ON THE INTERACTION BETWEEN SULFUR IN VEHICLE EXHAUST AND VARIOUS CATALYST COMPONENTS As mentioned above, sulfur originating from organo-sulfur compounds in the fuel reaches the catalytic converter chiefly in the form of SO/. Previous laboratory studies led to the conclusion that SO 2 inhibits catalytic activity in commercial three-way catalysts primarily under fuel-rich, oxygen-poor conditions, and particularly over Pt and Pd catalysts [9-11 ]. This is consistent with data showing that sulfur inhibits steam reforming and water-gas shift reactions which play an important role in the removal of CO and hydrocarbons under fuelrich conditions [11-12]. It has been shown that in catalysts with high levels or ceria, these two reactions are promoted, but in the presence of sulfur, these two reactions are even more severely inhibited than if the ceria was not present [ 18,26], suggesting a strong interaction between sulfur and the ceria component, which we will discuss later in this section. A number of surface science studies have been conducted which have shown that SO 2 adsorbs dissociatively on noble metal surfaces, forming adsorbed oxygen atoms and sulfur atoms: SO2(g ) --

SO2(ads)

"-

S (ads) + 2 0

(ads)

(1)

Whereas the oxygen atoms react with reducing species such as CO, H 2 or hydrocarbons at elevated temperatures, the adsorbed sulfur atoms are very strongly bound by comparison and can only react with an excess of oxygen present in the gas phase at temperatures greater than those required to remove the adsorbed oxygen, or they can react with hydrogen at even higher temperatures (sometimes as high as 700~ to form HzS. It follows, then, that a coverage of sulfur atoms on a noble metal surface is most favored under fuel-rich conditions, leading to the observation that sulfur poisoning is most prominent under fuel-rich conditions. Even early vehicle studies have shown that the importance of sulfur poisoning was dependent on the air/fuel calibration [4-6]. Laboratory studies also suggested that very small amounts of sulfur in exhaust could lead to significant impacts on catalyst performance [18]. For example, studies of SO 2 dissociation on well-characterized noble metal surfaces showed that relatively small surface coverage of adsorbed sulfur, on the order of 0.1 monolayer, could effectively block nearly all of the noble metal surface from adsorbing other reactant molecules such as H 2 or CO. This was attributable to a strong electronic effect of adsorbed sulfur in the donation of electron density to the noble metal surface, with the result being the inhibition of adsorption of other gas phase species [38,39]. Over the last two decades, the ceria content in catalyst formulations was increased, resulting in an increased H2S emissions. New studies were performed to investigate H2S formation over three-way catalysts and ways to control it. These studies showed that significant amounts of sulfur were stored in the catalyst under oxygen-neutral or oxygen-rich conditions, and were then released during fuel-rich excursions [21]. In oxygen-rich, or lean conditions, storage involves the interaction between sulfur dioxide in the gas phase and the base metal oxides in the catalyst to form sulfate and sulfites on the surface of their particles [18]: SO 2 +

1/2 02 = SO 3

(2)

3 SO 3 + A1203 = A12(SO4) 3

(3)

2 S O 3 + CeO 2= Ce(SO4) 2

(4)

3 SO 2 + 6 C e O 2 = C e 2 ( 5 0 4 ) 3 + 2 C e 2 0 3

(5)

The noble metals may assist in this process, since they are well known for catalyzing the oxidation of sulfur dioxide to sulfur trioxide species in an oxygen-rich environment

26

[ 1,14,15]. Fundamental studies have shown that SO 2 adsorbs strongly on alumina and other base metal oxides, which may be an intermediate step to the formation of surface aluminum sulfates. These sulfate species are favored between 600 and 800 ~ C, which is a typical operating range for an exhaust three-way catalyst [22,23]. Most of the storage, however, may involve interaction with ceria to form cerium sulfate and sulfite species, as it has been demonstrated that there is a direct correlation between the cerium content in catalysts and the amount of sulfur stored [24]. Operation under fuel-rich conditions following a prolonged period of operation in oxygen-rich conditions can lead to a rapid release of an undesirable amount of H2S, leading to an exhaust odor problem. Inhibition of this rapid release has been addressed by several strategies. One strategy involves sintering of the ceria particles in early stages of catalyst formulation to decrease the active ceria surface area and thus decrease their capacity for sulfate storage. A second strategy involves incorporation of a sulfur scavenging agent such as nickel or copper, which reacts with gas phase H2S in fuel-rich conditions to form nickel or copper sulfide. In subsequent oxygen-rich conditions, the sulfide compound is oxidized to form gas phase SO 2 and nickel or copper sulfate [24-26]. Finally, we note from very early studies of oxidation catalysts that in oxygen-rich environment and in the presence of water, any sulfur trioxide which is formed from the oxidation of sulfur dioxide can react further to form sulfurous and sulfuric acids, which can further react with alumina to form aluminum sulfate [40] which is stable from 100~ to 650~ Decomposition of this compound at higher temperatures or even interaction of the acid species with alumina can result in formation of surface acid sites which are very active for the oxidation of alkane hydrocarbons, as observed in oxygen-rich conditions when sulfur is present in the exhaust gas [24]. Formation of these surface acid sites in the catalyst, however, can lead to coking [41]. As the need for more effective emission controls became evident through the Clean Air Act Amendments of 1990, attention turned to the use of Pd catalysts to achieve rapid activation, or lightoff, during the early phase of the FTP test, where most of the hydrocarbon emissions are produced. Concerns have arisen, however, concerning the particular sensitivity of Pd catalysts to sulfur poisoning and a sulfur "memory" effect. Of the noble metals used in three-way automotive catalysts, Pd is most likely to exist in an oxide form [42], at least under oxygen-rich conditions, leading to the possible formation of a Pd sulfite or sulfate. In addition, it has also been suggested that Pd in the metallic form is uniquely capable of storing sulfur atoms in the bulk of the metal particle, which would also provide a rational explanation for the sensitivity to sulfur poisoning and memory effects. Indeed, surface science studies have provided evidence for such a phenomenon [43]. In summary, these concerns about Pd catalysts have led to a number of studies comparing its sensitivity to sulfur poisoning to that of other noble metals, as well as further development of ways to reduce this sensitivity, as we will discuss in the following section. 3. T H E I M P A C T OF SULFUR ON PALLADIUM L I G H T O F F CATALYSTS L A B O R A T O R Y STUDIES In this section we will discuss an example of recent laboratory work on the impact of sulfur on palladium catalysts formulations compared to other noble metal formulations. The laboratory studies discussed here have been performed using an apparatus which simulates the exhaust gas generated from a vehicle under several operating modes. The simulated exhaust gas is then heated to a controlled level and directed to a sample core taken from a commercial automotive converter. Quantitative analysis of pollutant and other gas species (CO, HC, NOx, CO 2 and 02) is performed using gas bench analyzers prior to and following the catalyst sample to determine the conversion efficiency for HC, CO and NOx. One of the tests used in analyzing catalyst performance involves a determination of the temperature at which the catalyst becomes active (sometimes called a lightoff temperature). A second test involves varying the average air/fuel ratio to determine the performance of the

27

catalyst as a function of gas stoichiometry, which we designate as an air/fuel sweep test. A third analysis involves operation of the catalyst at a constant temperature while varying the sulfur concentration to determine its impact. The reactor used to generate these tests as well as a more complete description of the tests has been discussed elsewhere [44]. We will first compare the results of tests conducted on Pd and Pt/Rh formulations. Both catalysts were obtained as commercial products in 1993, the Pd formulation with a loading of roughly 0.29%, and the P t ~ h formulation with a loading of 0.13 wt. % Pt, 0.0093 wt. % Rh. Both formulations contained about the same amount of ceria. Sample cores were taken from whole monolith converters of each formulation. Each sample core was aged using a laboratory reactor to an equivalent of 30,000 to 50,000 miles aging on a vehicle. The impact of sulfur on catalyst lightoff was determined first. For these tests, the catalyst sample temperature was varied from 100~ to 600~ using a constant rate of temperature increase. The sulfur level was held constant during the generation of each lightoff curve, and this was repeated using several different sulfur levels representative of the field. The effects of sulfur dioxide on lightoff at stoichiometry of the thermally aged 0.29 % Pd catalyst are shown in Figure 2. Catalyst lightoff profiles were obtained for sulfur levels simulating 0 to 450 ppm in fuel. Each test was followed by a temperature excursion to 700 ~ C in order to restore the catalytic activity to a level equivalent to 0 ppm sulfur in the feedstream. The results consistently show an increase in lightoff temperature, and therefore decreases in activities to remove HC, CO and NOx as the sulfur content in the feedstream is increased. The increase in the lightoff temperatures of HC and CO is non-linear with increasing sulfur concentration such that most of the increase in the lightoff temperature occurs at the lowest

100

Pd Catalyst

t

r

50 r. ! ' - t [ I ".:! }- I/."i

v

0 r (D o uJ tO O3

> 0 0

0

Pt-Rh Catalyst

-++~--~-:.--~ "

-

1"-15oppm I --- 300 ppm I --- 450 ppm

t ...... 0 p p m a f t e r 4 5 0 p p m

100

! |

1

/

.,-.-,--..'...,..,+ .................

50 co

+0I

O0

,~..... i . . . . . . .

....

20'() ..............300

NSx.] 400

Temperature (C)

500 200

300

400

500

Temperature (C)

Figure 2. Impact of sulfur on lightoff profiles of aged Pd and Pt/Rh commercial catalysts determined in a laboratory study.

28 sulfur levels corresponding to 0 to 150 ppm, while a smaller increase in the lightoff occurs with further increase in the SO 2 concentration from 150 ppm to 450 ppm Typical of a Pd catalyst of this generation, the NOx activity is markedly lower than HC or CO, and adding even small amounts of SO 2 results in the failure to reach the 50% conversion efficiency over the temperature range tested. Although not indicated in this figure, the magnitude of the effect of SO 2 on lightoff activity is larger than the effect of the thermal aging procedure on activity. The thermal aging procedure used in this study has been found to generally cause a 30-35~ increase in CO and HC lightoff temperature, while the impact of a sulfur increase form 0 ppm to 450 ppm results in an increase of 40-50~ in CO and HC lightoff temperature. Similar lightoff tests performed with the Pt/Rh catalyst are also shown in Figure 2. Note again that the HC, CO and NOx lightoff temperature increases as the sulfur content in the feedstream is increased. The magnitude of the increase in lightoff is on the order of 35 ~ as the SO 2 content is increased from 0 ppm to 450 ppm, somewhat smaller than for Pd, but the magnitude of the impact is non-linear in the same way as the Pd catalyst. Although not indicated in this figure, the impact of increasing the sulfur level from 0 to 450 ppm is roughly comparable to the effect of the thermal aging treatment used in this study. We note in comparing the two catalyst technologies for the case with no sulfur present in the feedstream, Pd holds a clear advantage over Pt/Rh for reducing cold-start alkene hydrocarbon and CO emissions; this is one of several reasons why Pd-based catalyst technology is used in closecoupled converter applications. The non-linear relationship between the SO 2 concentration and the magnitude of the impact on hydrocarbon performance for Pd and Pt-Rh based catalysts is characteristic of the effect of sulfur on monolithic three-way catalysts observed in previously reported laboratory [ 18] and vehicular studies (Figure 3) [20]. Both studies show that as the sulfur content is increased, the degradation in the lightoff or warmed-up emissions increases at a high rate initially, but becomes more gradual at higher sulfur levels, suggesting that very small amounts of sulfur in the exhaust can have a significant effect on emissions. This is consistent with our earlier discussion of surface science data showing that low surface coverages of SO 2 can inhibit the adsorption of other gas phase species such as CO, NOx or HC on noble metal surfaces. For all of the aged catalysts, additional lightoff tests were performed with no SO 2 in the feedstream following the lightoff experiment using 450 ppm SO 2. The results (dashed lines in Figure 2) show a decrease in the lightoff temperatures from the 450 ppm SO 2 experiment for both Pd and Pt-Rh type catalysts, but only partial recovery of activity (relative to the catalyst prior to exposure to SO2) has taken place with the Pd catalyst, whereas nearly complete recovery of the original activity has taken place with the Pt-Rh catalyst. This irreversible sulfur poisoning has been observed elsewhere for Pd catalysts in particular [20,41 ]. It has been suggested that this apparent irreversibility under typical catalyst operating temperature (400-600~ is related to a direct reaction between SO 2 and supported Pd metal, perhaps leading to the formation of a surface PdS-like species on the surface of the Pd particles [43]. However, as stated earlier, evidence of the ability of sulfur atoms to migrate into the bulk of Pd metal has been found recently, a phenomenon not observed for other noble metals. Additional tests with both aged catalysts were performed in which the net stoichiometry was scanned from a net reducing feedstream (A/F = 14.1) through stoichiometry (A/F = 14.6) to a net oxidizing feedstream (A/F = 15.1) at a constant temperature of 500~ During this test, a cycling amplitude of +0.3 A/F and a cycling frequency of 2.0 Hz was used. This A/F ratio sweep test closely models the exhaust of a late model 3.8 L V-6 engine operating under closed-loop control, and was used to determine the impact of sulfur on catalyst activity as a function of feedstream stoichiometry by first collecting data using no sulfur in the feedstream, and then repeating with 450 ppm sulfur in the feedstream. Figure 4 shows that in the oxidative, or "lean" exhaust environment between mean A/F ratios of 15.1 to 14.7, the presence of SO 2 results in a slight decrease in the conversion efficiency for HC, CO and NOx. Below a value of 14.7, the presence of sulfur results in a slight decrease in the CO act-

29 iii

im

E

o~

O

-i-

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I I I I 100 200 300 400 Sulfur Concentration (ppm)

500

Figure 3. Impact of fuel sulfur on vehicle emission performance in an FTP test for a 1989 model year vehicle. ivity which continues until a rich extreme of 14.1 in A/F ratio. The effect of SO 2 on the conversion efficiencies of HC and NOx, however, is significant below an A/F ratio of 14.7 and increases with decreasing mean A/F until the "rich" extreme of 14.1 is reached. Thus, the degree of the impact of sulfur on HC and NOx conversion efficiency is greatest under stoichiometric and "rich" operating conditions. The maximum impact of 450 ppm SO 2 in terms of breakthrough in HC and NOx occurs at a mean A/F ratio of 14.3 to 14.2. At this point, the HC breakthrough increases by 600-700% while the NOx breakthrough increases by 400-500%. This behavior has been observed in other laboratory studies [26] and is consistent with suggestions that SO 2 poisons the noble metal surface under rich conditions [18]. In the case of the thermally aged 0.13% Pt, 0.0093% Rh production catalyst (Figure 4), the effect of adding 450 ppm SO 2 is generally similar to the Pd catalyst in that the HC and NOx conversion efficiencies are significantly decreased in rich conditions, while the CO conversion efficiency is only moderately decreased under rich conditions. With no sulfur present, propylene conversion efficiency is slightly lower for the Pd catalyst in comparison to the Pt-Rh catalyst in lean or stoichiometric conditions, but higher for the Pd catalyst under rich conditions. This hydrocarbon activity advantage of Pd over Pt-Rh under rich conditions

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Figure 4. Impact of sulfur on performance of aged Pd and Pt/Rh commercial catalysts under warmed-up conditions as a function of air/fuel ratio. is completely lost when 450 ppm S O 2 is present in the feedstream. This phenomenon becomes a disadvantage when sulfur is removed from the feedstream, since the Pd catalyst does not recover all of its original activity, but the Pt-Rh catalyst does. With an increase in the sulfur content in the gas feedstream, the consequent decrease in the NOx activity of the Pd catalyst was larger than the Pt-Rh catalyst. This result is consistent with the reported relative resistance of Rh to sulfur poisoning [ 13-18]. Further tests with the Pd catalyst showed that the deterioration in HC and NOx performance at stoichiometry and at 600~ due to sulfur is fully reversible when the sulfur concentration is increased from 15 ppm to 75 ppm and returned to 15 ppm. However, the deterioration of HC and NOx performance that occurs when the sulfur concentration is increased from 15 ppm to 450 ppm is not completely reversed when the sulfur level is subsequently reduced to 15 ppm unless the catalyst is operated at or above 700~ (Figure 5). The impact of sulfur on warmed-up HC efficiency was minimal for operation above 700~ at either rich or stoichiometric conditions, which again is consistent with our findings that sulfur species can be removed from noble metal surfaces by reaction above this temperature under both oxygen-rich and oxygen-poor conditions. To summarize the laboratory data comparing Pd and Pt/Rh commercial catalysts, we found that an increase in the sulfur content in simulated exhaust results in a loss of both lightoff and warmed-up activity. The lightoff activity of the Pd is generally better than the Pt-Rh catalyst regardless of the sulfur content in the feedstream. Under warmed-up conditions, the loss of activity for HC, CO and NOx in the presence of sulfur was greater under slightly rich conditions than under lean conditions for both Pd and Pt-Rh catalysts,

31

100

Recovery of Activity at Stoichiometry

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.

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.

.

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.

.

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500 C 500 C 600 C 700 C no sulfur no sulfur no sulfur no sulfur 500 C 550 C 650 C 450 ppm no sulfur no sulfur sulfur

Figure 5. Reversibility of performance in Pd catalysts as a function of operating temperature and stoichiometry. while the magnitude of the impact on HC and NOx activity under warmed-up stoichiometric conditions was significant greater for the Pd catalyst than for the Pt-Rh catalyst. The change in the relative impact as a function of air/fuel ratio observed from these tests predicts that engine management can significantly influence the resulting magnitude of the sulfur impact on the emission control system, as discussed in the next section. Finally, it was found that the effect of SO 2 on the activity of the Pd catalyst was partly irreversible under the conditions used in this study, while the effect on Pt-Rh was completely reversible. Part of the irreversible poisoning is attributed to a reaction between SO 2 and Pd. Although the impact of sulfur on Pd catalyst performance as shown by these data is dramatic, recent advances in Pd catalyst technology have led to improvements in resistance to sulfur poisoning. Figure 6 shows a comparison between the catalyst used in the study just mentioned and a more recently developed Pd catalyst formulation having the same metal

32

-~

450

co

0

100

Improved pd Formulation

.

O

,-"

0-~

50

x

g 0

--

0

10

- - -

I

20

I

"

30

!

40

'

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Time (minutes) Figure 6. Comparison of sulfur impact on two Pd catalyst formulations. loading. In both cases, cores from both catalysts were aged to an equivalent level and analyzed on the laboratory reactor system under constant temperature and using the same simulation of a mean air/fuel ratio of 14.6 with a cycling amplitude of _+0.3 air/fuel units and a cycling frequency of 2 Hz. The initial activity of the newer formulation is higher than the earlier generation formulation, and the impact of 450 ppm sulfur is smaller. This improved activity as well as improved resistance to sulfur poisoning can be attributed to the use of promoters and stabilizers such as zirconia and lanthana, as well as improvements in washcoat processing. Some of these improvements have been discussed in the literature recently and involve incorporation of sulfur "guards" or scavengers, as well as modifications to the ceria and the ceria-Pd interface [29]. 4. T H E I M P A C T OF S U L F U R ON A D V A N C E D A U T O M O T I V E C A T A L Y S T S : R E C E N T V E H I C L E STUDIES A number of vehicle studies have been conducted recently on the impact of sulfur on emissions, and recent emphasis has been placed on determining whether newer technology catalysts and emission control systems can be devised with greater tolerance to sulfur poisoning, particularly in light of increased strictness on emission standards. Although we will not be able to include all of the studies conducted recently, we will discuss some examples from the literature. As mentioned previously, there has been particular concern about the sensitivity of Pd and therefore Low Emission Vehicle (LEV) catalyst systems to sulfur poisoning. Indeed, studies of LEV candidate systems compared to earlier technology catalysts appears to verify these concerns (Table 2) [45]. According to these results, LEV catalysts on the whole perform better than earlier generation catalysts when operated on the

33 Table 2 Impact of Sulfur on Low Emission Vehicle Catalysts. Sulfur Level

Catalyst Efficiency (Warmed-Up Conditions, at Stoichiometry) HC (%) CO (%) NOx(%)

Car 1:2 converters, Pt/Rh lightoff and Pt/Rh underbody 40 ppm S 99.7 96.9 300 ppm S 98.5 93.3 600 ppm S 96.1 83.8

88.1 81.5 86.1

Car 2:1 converter, underbody, Pd front monolith, Pt/Rh rear monolith 40 ppm S 97.8 97.4 300 ppm S 94.5 81.6 600 ppm S 93.3 80.5

88.9 69.7 66.8

Car 3:1 converter, underbody, Pd front monolith, Pt/Rh rear monolith 40 ppm S 98.7 85.6 300 ppm S 96.0 77.8 600 ppm S 94.5 79.9

95.4 86.6 79.6

Fleet average from a set of 1989 model year vehicles: 40 ppm S 91.4 81.6 71.8 300 ppm S 89.3 74.1 69.3 600 ppm S 89.0 72.0 67.0 Note: all catalysts aged on an engine dynamometer to simulate 100,000 miles before testing. low sulfur content fuel (40 ppm S), but increases in the sulfur content result in larger relative decreases in catalyst HC efficiency for the LEV catalysts. The magnitude of the sulfur impact varies among the three vehicles, which have different catalyst configurations as described in the table. As discussed later in this section, the extent of the sulfur impact is vehicle and system dependent. Other studies have also been conducted recently that indicate that Pd catalysts may be more sensitive to sulfur poisoning. One recent study compared several catalyst technologies including Pd and Pt/Rh catalysts [28]. Part of the data from that study is summarized in Table 3. At stoichiometry, the performance of the Pd catalyst is less impacted than the Pt/Rh catalyst. Although this result differs from the laboratory studies mentioned earlier, it may be partly attributable to 1) different washcoat technologies being used for the two formulations and 2) the same aging conditions being used for the two technologies (it is well known that Pd and P t ~ h catalysts respond differently to the same laboratory or engine dynamometer aging condition. When tested using a fuel-rich condition with little sulfur present, the Pd catalyst was more active than Pt/Rh, but as the sulfur content was increased, performance deteriorated more rapidly for the Pd catalyst than for the Pt/Rh catalyst. The engine control systems used in this recent study [28] were operated in a fuel-rich mode during the initial warm-up phase of the federal test procedure (FTP), where most of the hydrocarbon emissions are emitted from the tailpipe, which is consistent with the observation that Pd is more severely deteriorated by sulfur than Pt/Rh. Other studies have been performed to compare how catalyst technologies and indeed entire vehicle emission systems respond to increased levels of fuel sulfur [46-48]. For some test vehicles, catalyst HC performance was particularly affected by fuel sulfur when Pd was involved in the catalyst. In other studies using TLEVs a distinct advantage of Pd catalysts

34 Table 3 Impact of Fuel Sulfur on Three-Way Catalyst Performance (warmed-up conversion efficiency at conditions shown during exposure to sulfur at the concentrations shown). Catalyst

Sulfur Level (ppm)

Reaction Condition

HC (%)

CO (%)

NOx (%)

Pd

14 90 500

stoichiometric

96 95 95

95 94 94

94 90 88

Pd

14 90 500

fuel-rich

73 71 59

37 24 12

94 65 43

Pd/Rh

14 90 500

stochiometric

96 94 93

95 87 82

93 86 80

Pd/Rh

14 fuel-rich 54 32 94 90 55 30 92 500 56 28 91 Note: Both catalysts operating at 450 C. Both catalysts had been aged on an engine dynamometer to simulate -50,000 miles. for the reduction of HC emissions was observed under all conditions, although high sulfur fuels caused significant deterioration in CO and NOx activity. Low temperature aging coupled with high sulfur exposure resulted in particularly poor performance for Pd catalysts when compared to P d ~ h and Pt/Rh catalysts [49]. For nearly all of these studies, the authors concluded that the extent of the sulfur effect was greatly influenced by the engine management strategy. Furthermore, the results suggest that emissions could be improved by advances in the emission control system technology. On the basis of these studies, it was also stated that some newly developed catalysts may have less sensitivity to sulfur than earlier generation Pt/Rh and Pd/Rh catalysts (Table 4). The resulting improvements observed in vehicle studies are consistent with laboratory tests mentioned earlier indicating greater sulfur tolerance in newer technology catalysts employing multilayer washcoats, sulfur "guards" and promoters and stabilizers. Table 4 Reduction in Vehicle Emissions Resulting From Reduction in Fuel Sulfur over Three Different Catalyst Technologies (Average Results Using Three Different Vehicles, European Test Procedure) Catalyst Technology CT 1 (Production Pt/Rh) CT2 (Pd/Rh, new technology) CT3 (same as CT2, but increased metal loadings)

HC (g/mi) CO (g/mi) NOx (g/mi) 0.010 0.007 0.015

Note: Results from reduction in fuel sulfur from 397 to 100 ppm.

0.117 0.140 0.030

0.043 0.030 0.030

35

To summarize, recently published vehicle studies suggest that Pd-only catalysts may under some conditions and in some vehicle applications be more sulfur-sensitive than Pt/Rh, Pd/Rh or Pt/Pd/Rh catalysts. However, recent improvements in Pd-containing catalyst technology may lead to better tolerance to sulfur poisoning, but improvements in catalyst performance and in sulfur tolerance may also depend on the vehicle and its emission control system properties. 5. T H E I M P A C T OF S U L F U R ON OBD-II C A T A L Y S T M O N I T O R There have also been some concerns about the impact of fuel sulfur on the operation and durability of on-board diagnostics (OBD-II) which are now required on all light and medium duty vehicles and which must assess the performance of a number of emission controls on the vehicle, including the catalytic converter system and oxygen sensors. Converter performance which falls below an acceptable level defined by the regulations must be identified as malfunctioning by activation of an indicator lamp and storage of a fault or trouble code in the engine control module (ECM) which can be read via access to the vehicle's assembly line data link (ALDL). When the fault code is set, it can only be cleared during servicing. The regulatory agencies have advocated the use of two conventional oxygen sensors placed in front of and behind each catalyst or catalyst system to determine catalyst efficiency [31-33]. This measurement method uses the air/fuel ratio "dithering" activity detected in the post-catalyst sensor either in an absolute way or relative to the dithering activity in the pre-catalyst sensor as an indication of loss of catalyst performance. Other data suggest that this measurement is correlatable to the relative oxygen storage capacity of the catalyst and not to catalyst hydrocarbon conversion efficiency directly [3334]. It has also been demonstrated that this measurement can be significantly affected by a variety environmental and operational conditions which cannot be controlled or are difficult to use as enabling criteria for measurement. This results in a high variability in the oxygen storage measurement itself, particularly for severely aged or damaged catalysts which are near the diagnostic failure criteria, which in turn can lead to a probability of misdiagnosis. One of the factors which can influence the measurable amount of oxygen storage is the sulfur content in the fuel, which in the form of sulfur dioxide in exhaust can react with the oxygen storage component in the catalyst (ceria) as we have mentioned in an earlier section. Some early studies which examined the impact of sulfur on the titratable oxygen storage in a catalyst showed a dramatic impact [35,36], although such studies were conducted in a laboratory using small cores from a production catalysts. The significance of the sulfur impact on OBD-II has been discussed elsewhere [50] and there is considerable disagreement about the extent of its expected impact in the field. Although few vehicle studies have been conducted [37,50,51 ], the available data verify sulfur can affect the OBD-II catalyst monitor, although such studies conclude that only catalysts which have been severely aged or have seen severe service may be susceptible to significant degradation by sulfur on oxygen storage capacity, as shown in the example data in Figure 7 [51 ]. Such catalysts perform adequately using a fuel with relatively low sulfur content, and thus should not be diagnosed as having failed, but when operated on a higher sulfur content fuel, they are diagnosed as having failed by the OBD-II system. Since these situations are not expected to occur until OBD-II equipped vehicles accumulate a high level of mileage (above 50,000 to 100,000 miles), and since improvements and modifications continue to be made in the OBD-II systems themselves, it is not known whether a significant number of these situations will occur in the field. However, there is also concern about the direct impact of sulfur on the operation of oxygen sensors, since they in part depend on the catalytic action of a Pt electrode which is exposed to the exhaust gas. Under certain operating conditions, it may be possible for the sulfur to poison the operation of the sensor itself, and thus interfere with the measurement of oxygen storage [51 ]. All of these issues continued to be studied.

36

"F: Maximum ::~ Value Im

-O ..,

(O 0

=E'E x 0 a)=E

"0

C

..,

(n

U)" -;" Cl Minimum Ill Value 0 0

9. . . .

e, .....

I. . . . .

200

,I .

.

400

.

.

600

.

.

800

1000

Fail Threshold

Sulfur Content in Fuel (ppm) Figure 7. Impact of sulfur on the OBD-II catalyst monitor in a production Transitional Low Emission Vehicle (TLEV). 6. SUMMARY In this discussion, a few examples of recent studies of the sulfur impact on performance of three-way catalysts and on catalyst oxygen storage have been discussed. Although this is not important in such locations as California or Japan where there are strict controls on the sulfur content in fuel, it is important in locations where the sulfur content in fuel is not strictly controlled and where advanced emission control standards have been adopted or are being considered. Although the use of Pd is one of the chief strategies for achieving rapid catalyst activation and for achieving improved HC activity, there is concern about its sensitivity to sulfur poisoning. Recent improvements in Pd and other catalyst technologies have resulted in improved sulfur tolerance, but the bulk of the data suggests the magnitude of the sensitivity of a catalyst to sulfur may depend heavily on how the catalyst is integrated with the vehicles' emission control system and engine management strategy. There are also concerns about the impact of sulfur on OBD-II systems and particularly on catalyst oxygen storage and the sensors used to measure oxygen storage capacity. The available vehicle data indicates sulfur can affect the measurement of oxygen storage capacity and thus impact the proper diagnosis of a converter, but only for severely aged converters. Since OBD-II is still in relative infancy, and it is anticipated that these systems will be improved over the next decade, additional studies will be needed to verify the extent and significance of the sulfur impact on catalyst performance and on OBD-II systems.

37 REFERENCES

1.

"National Fuel Surveys: Gasoline and Diesel Fuel, Winter 1994", American Automobile Manufacturers Association, 1995. 2. H.V. Drushel, Am. Chem. Soc., Div. Pet. Chem., Prepr., 15(2) (1970) C13. 3. O. Weisser and S. Landa, Sulfide Catalysts, Their Properties and Applications, Pergamon Press, New York, 1973. 4. C.N. Satterfield, AIChE J., 21 (1975) 209. 5. EPA Final Rule on Regulation of Fuels and Fuel Additives: Standards for Reformulated Gasoline, Federal Register Notice, July 20, 1994, p. 59FR36944. 6. K.C. Taylor, Ind. Eng. Chem., Prod. Res. Dev., 50 (1976) 264. 7. J.C. Summer, Env. Sci. Tech., 13 (1979) 321. 8. H.C. Yao and J. Yu, J. Catal., 36 (1975) 266. 9. J.C. Summers and K. Baron, J. Catal., 53 (1979) 380. 10. W. B. Williamson, H. S. Gandhi, M. E. Heyde and G. A. Zawacki, SAE Paper No. 790942 (1979). 11. G. J. Joy, G. R. Lester and F. S. Molinaro, SAE Paper No. 790943 (1979). 12. J. C. Schlatter and P. J. Mitchell, Ind. Eng. Chem., Prod. Res. Dev., 19 (1980) 288. 13. U. Kohler and H.W. Wassmuth, Surf. Sci., 126 (1983) 448. 14. O. K. T. Wu and R. P. Burns, Surf. Int. Anal., 3 (1981) 29. 15. M. L. Burke and R. J. Madix, Surf. Sci., 194 (1988) 223. 16. R. C. Ku and P. Wynblatt, Appl. Surf. Sci., 8 (1981) 250. 17. St. Astegger and E. Bechtold, Surf. Sci., 122 (1982) 491. 18. D. D. Beck, M. H. Krueger and D. R. Monroe, S AE Paper No. 910844 (1991). 19. R. L. Furey and D. R. Monroe, SAE Paper No. 811228 (1981). 20. W. J. Koehl, J. D. Benson, V. R. Burns, R. A. Gorse, Jr., A. M. Hochhauser, J. C. Knepper, W. R. Leppard, L. J. Painter, L. A. Rapp, R. M. Reuter and J. A. Rutherford, Effects of Gasoline Sulfur Level on Exhaust Mass and Speciated Emissions: The Question of Linearity - Auto/Oil Air Quality Improvement Program, SAE Paper No. 932727 (1993). 21. A. F. Diwell, C. Hallett and J. R. Taylor, SAE Paper No. 872163 (1987). 22. A. V. Deo, I. G. Dalla Lana and H. W. Habgood, J. Catal., 21 (1971) 2710. 23. A. Datta, R. G. Cavell, R. M. Tower and Z. M. George, J. Phys. Chem., 89 (1985) 443. 24. H. G. Henke, J. J. White and G. W. Denison, SAE Paper No. 872134 (1987). 25. J. C. Summers, J. F. Skowron and W. B. Williamson, SAE Paper No. 920558 (1992). 26. D. R. Monroe, M. H. Krueger, D. D. Beck and M. J. D'Aniello, Jr., Stud. Surf. Sci. Catal., 71 (1991) 593. 27. H. Muraki, SAE Paper No. 910842 (1991). 28. J. C. Summers, J. F. Skowron and M. J. Miller, SAE Paper No. 930386 (1993). 29. P. Burke, Presented to the American Chemical Society Fall National Meeting, August 1996, Orlando, FL. 30. California Code of Regulations, Section 1968.1, Title 13, Adopted September 14, 1989, Effective Sept. 26, 1990. 31. W. B. Clemmens, M. A. Sabourin, and T. Rao, SAE Paper No. 900062 (1990). 32. J. W. Koupal, M. A. Sabourin, and W. B. Clemmens, SAE Paper No. 910561 (1991). 33. J. S. Hepburn and H. S. Gandhi, SAE Paper No. 920831 (1992). 34. G. B. Fisher, J. R. Theis, M. V. Casarella and S. T. Mahan, SAE Paper No. 931034 (1993). 35. D. D. Beck, J. W. Sommers and C. L. DiMaggio, Appl. Catal. B: Environmental, 11 (1997) 273. 36. D. D. Beck, D. R. Monroe, C. L. DiMaggio and J. W. Sommers, SAE Paper No. 952416 (1995). 37. D. D. Beck, T. W. Silvis and S. T. Mahan, SAE Paper No. 941054 (1994). 38. C. R. Apesteguia, C. E. Brema, T. F. Garetto, A. Borgna and J. M. Parera, J. Catal., 89

38 (1984) 52. 39. J. M. MacLaren, J. P. Pendry and R. M. Joyner, Surf. Sci. Lett., 165 (1986) L80. 40. W. W. Decker and J. R. West (eds.), The Manufacture of Sulfur Acid, Robert Krieger Publishing Company, Huntington, New York, 1974. 41. D. D. Beck and J. W. Sommers, Stud. Surf. Sci. Catal., 96 (1995) 721. 42. R. J. Farrauto, J. K. Lampert, M. C. Hobson and E. M. Waterman, Appl. Catal. B: Environmental, 6 (1995) 263. 43. C. L. DiMaggio and D. D. Beck, "Interaction Between Sulfur and Pd Surfaces", Surf. Sci. Lett (submitted). 44. D. D. Beck, M. H. Krueger, D. R. Monroe, D. J. Upton and J. M. Lendway, SAE Paper No. 920099 (1992). 45. F. S. Gerry, R. A. Gorse, Jr., W. J. Bandy, D. D. Beck, V. Bums, H. Doherty, G. Herwick and J. Rutherford, CRC Sulfur/OBD-II Laboratory Research Program: Executive and Technical Final Reports, Coordinating Research Council Report No. 602 (1997). 46. P. Beckwith, P. J. Bennett, C. L. Goodfellow, R. J. Brisley and A. Wilkins, SAE Paper No. 940310 (1994). 47. P. J. Bennett, P. Beckwith, S. D. Bjordal and C. L. Goodfellow, SAE Paper No. 961901 (1996). 48. S. D. Bjordal, C. L. Goodfellow, P. Bennett and P. Beckwith, SAE Paper No. 961902 (1996). 49. J. E. Thoss, J. S. Rieck and C. J. Bennett, SAE Paper No. 970737 (1997). 50. L. H. Browning and C. B. Moyer, SAE Paper No. 952422 (1995). 51. D. D. Beck and W. A. Short, SAE Paper No. 952424 (1995).

9

Elsevier Science B. V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

39

Solid State Reactions in Catalysts: A n A p p r o a c h to Real Active Systems and t h e i r Deactivation

Bernard Delmon Universit6 catholique de Louvain, Unit6 de catalyse et chimie des matrriaux divisrs, Place Croix du Sud, 2/17, B- 1348 Louvain-la-Neuve (Belgium).

Solid state reactions are involved in many stages of catalyst life, from preparation to deactivation. Reactions taking place at the catalyst surface during catalysis may relate to those occurring during preparation. They are the origin of several deactivation processes. The present contribution is aimed at enlarging the concept of deactivation by solid state reactions by considering all those involved in the bulk and surface over the entire life of the catalyst. The main example will be that of the behaviour of molybdenum oxides in selective oxidations. It will be shown that the formation of shear planes explains (i) the activation of MoO3 to the very flexible structure of the active suboxide (Mo18052 or Mo8023), (ii) the exceptional catalytic activity of these two intermediary oxides and (iii) the deactivation due to the formation of Mo4Oll. This conclusion concerning the similitude of solid state reactions at various stages of catalyst life can be extended to molybdates of various metals and other catalysts active in selective oxidation. This suggests that a comprehensive view of the whole catalyst life may enable a better control of useful catalytic properties. The beneficial role of spillover oxygen, adequate supports, or the advantage of alternate reduction and oxidation stages in the industrial process (e.g. the so-called "riser" reactor) are, among others, examples of useful, practical applications which emerge when this comprehensive view is adopted.

1.

INTRODUCTION

This contribution will address some issues which are not examined usually in meetings on catalyst deactivation. We will look at catalyst deactivation in a broader context, namely the modifications that solid catalysts (and probably also homogeneous ones!) continuously undergo while they are acting catalytically under reaction conditions. As underlined in our previous contributions, this is particularly relevant when solid-state reactions are considered, because solids have a sort of memory (1,2). Their transformations at a given moment are strongly influenced by the whole succession of conditions they have been subjected to. If we adopt this global view, another question arises. Catalysts, as they work in catalytic reactors, are just the intermediary product of a solid state reaction chain which starts with preparation and activation, before the catalysts is contacted with the feed, and is followed by deactivation and, ultimately, "death". The question is: what are the relationships among solid-state reactions occurring during the initial activation, catalytic reaction and deactivation?

40 For many catalysts, the technical challenge is not only attaining higher activity or better selectivity, but also achieving more resistance to ageing and deactivation. The second goal is even predominant in the case of steam reforming catalysts, some water-gas-shift catalysts, and methanation catalysts, especially in coal gasification. And the demands on catalysts are increasingly more severe. Even in hydrotreating, which had been a long-time work horse of catalysis, extreme activity and selectivity are now necessary for producing fuels increasingly less harmful to the environment. Selective oxidation (SO) is coming to the forefront of catalysis. For paraffins, activity and selectivity are crucial to the success of SO processes. Let us consider a classical example, the selective oxidation of butane to maleic anhydride (MAA). The best catalysts give, let us say, 70% selectivity to MAA in laboratory experiments. In fixed bed tubular reactors, it may take up to two weeks for freshly prepared catalysts to reach optimal activity and selectivity. But until recently, the selectivity in industrial reactors declined progressively to less than 55% over a period of 2 to 3 months, and many plants reported even 45% selectivity to MAA. It is now known that almost all these changes involve solid state reactions. This clearly suggests that maintaining high performance demands stimulating favourable reactions leading to selective phases and repressing those causing deactivation. In the previous case (SO to MAA), both activation and deactivation are reactions of solids with butane and oxygen. They are necessarily linked to each other. Hence, it follows that an effort aimed at maximising the active and selective phase must take into account all solid-state reactions from preparation to death. Making an analogy with ordinary kinetics, the problem is to maximise an intermediate product in a system of consecutive reactions. As regards the catalysts active in the reaction of butane to MAA, the beautiful work of J.C. Volta et al. gives clues to the activation reactions (3,4). A challenge is now to disentangle the solid state phenomena leading to deactivation. The above line of thought suggests that, where solid state reactions and deactivation are concerned, a closer look at the active state of the catalyst is absolutely necessary. This is done in this contribution. This comprehensive approach is probably not as important when other causes of deactivation are concerned, namely poisoning and coke formation. Nevertheless, in pursuing a comprehensive approach, it should be remembered that the various types of deactivation processes occurring in a given catalyst are far from independent. This conclusion is strongly supported by the fact that a large proportion of the authors in the series of Symposia on Catalyst Deactivation had to consider the coming into play, simultaneously, of more than one process, in the deactivation phenomena they described. These processes often belonged to different categories (coking, poisoning or solid state transformation). In this contribution, we shall first provide a backdrop by addressing the role of chemical structure in the working catalyst and the problem posed by the fact that the real working catalyst largely escapes physico-chemical characterisation. The subsequent section presents in detail the behaviour of MoO3 in selective oxidation catalysis. This is a relatively simple illustrative example of the comprehensive approach which is necessary for correctly understanding the problem and correlatively finding practical solutions. In the last section, a few examples are outlined where the approach presented earlier can be used with profit.

2.

SOLID STATE R E A C T I O N S IN W O R K I N G CATALYSTS

Scientists are now convinced that high activity and selectivity demand that the catalytically active atoms be in a very precise coordination. Well defined geometrical arrangements of surface

41 metal and oxygen atoms are certainly necessary in selective oxidation (especially in the case of the butane to maleic anhydride reaction, where 7 oxygen atoms must be extracted almost simultaneously from the vanadium phosphate lattice in a concerted mechanism). In hydrotreating, the hydrogenation sites consist of molybdenum atoms in a special coordination at the edges of MoS2 (a threefold coordinatively-unsaturated Mo). The hydrodesulfurization sites are very likely made of an ensemble of 2 Mo vicinal atoms in two different coordination environments. Nevertheless, the fundamental approach to heterogeneous catalysis still rests heavily on correlations between catalytic properties on the one hand and the coordination of surface atoms in the metal, oxide or sulfide phases as prepared by the experimenter. The control of this coordination is exerted via the choice of matrices with a selected structure, the use of grafted organo-metallic compounds of known configuration, or innumerable special fabrication techniques. This approach is observed to be unsuccessful in the vast majority of cases, because it is based on the implicit assumption that catalyst surface is not modified under reaction conditions. Actually, the surfaces of catalysts undergo continuous changes. The coordination achieved at the preparation stage is extensively modified under reaction conditions. Considering any given point of the surface, this coordination varies cyclically under the influence of the dynamics of the catalytic reaction. This may bring about changes in the course of time. It should be remarked incidentally that very few methods have been envisaged to achieve control of coordination at a later stage of catalyst life, namely when the catalyst is in a state closer to the functioning form. The mere adsorption of molecules modifies the atomic arrangement at the surface of solids. This was mentioned by Boreskov and other groups more than 20 years ago (5-7). Later, surface science experiments demonstrated this unequivocally. Adsorption brings about a reconstruction of surfaces. Many examples concerning metals have been published (8,9). The changes concern the coordination of a large proportion of surface atoms (often all of them). One or two layers below the surface are also affected. More profound changes in the coordination of surface atoms must occur when catalytic surfaces are contacted with reacting molecules. These changes are also demonstrated in surface science experiments on metals. When deactivation by solid state reactions is considered, the structure or, more precisely, the dynamic (one could say "wobbly") structural arrangement of the working catalyst, may be the starting point of further solid-state reactions. It should be emphasised that catalysis is a dissipative phenomenon occurring at surfaces (in the sense of the so-called "irreversible thermodynamics"). Energy is continuously dissipated. This facilitates the movement of surface atoms. This explains why loose or flexible structures form during the catalytic process. The excess free energy of these structures and the corresponding high reactivity make them inevitably excellent precursors for new, more stable, phases. These phases of higher stability are likely to be less active. The consequence is that the changes occurring continuously in the dynamic state trigger a part of, if not all of, the solid state reactions that lead to deactivation. At first glance, the above remarks do not seem to bring much new insight. In the field of oxidation, it has always been considered that the reduction-oxidation state of the working catalyst determined the rate of the deactivation due to solid-state processes. It was widely recognised, in particular, that progressive reduction of a catalytic phase during the successive catalytic cycles could lead to deactivation by loss or segregation of components. The nuance introduced by the above brief analysis is that the reactions and mechanisms determining activity and selectivity may be the same as those triggering deactivation. This justifies the present attempt to undertake a comprehensive approach. Selective catalytic oxidations provide examples which can illustrate this

42 new approach. This is the reason why a substantial part of this work deals with this group of reactions.

3. MoO3 AS O X I D A T I O N C A T A L Y S T : T H E S T A R T I N G O X I D E , T H E WORKING CATALYST, THE DEACTIVATED STRUCTURE

3.1.

General mechanism

Molybdenum oxide is present, together with molybdates of various metals (e.g. iron, cobalt or bismuth), in many selective oxidations catalysts (e.g. catalysts for oxidation of olefins and oxidative dehydrogenation of alcohols). Structurally, much is known concerning MOO3. Its catalytic activity has been extensively investigated. It will serve here as an illustration of the close relationships between the mechanisms explaining catalytic activity and deactivation. In doing so, account will be taken of results presented during the previous Symposium on Deactivation (10). In a sense, this section constitutes a logical continuation of this previous work. The catalyst as prepared by precipitation and calcination is fully oxidised (MOO3). Structurally, it consists of corner-sharing octahedra (Fig. 1a). The conventional picture of Fig. lb will be used for representing these MoO6 octahedra in several figures. Let us recall that the catalytic cycle in selective oxidation consists in a reduction of the catalyst by the hydrocarbon and the subsequent reoxidation by molecular oxygen. In the case of the so-called aUylic oxidation, e.g. propene to acrolein, 2 atoms of oxygen from the catalyst are used, one for removing 2 hydrogen atoms from the olefin, and the other one for forming the unsaturated aldehyde: Reduction: CH2 = CH - CH3 + MonO3n --~ CH2 = CH - CHO + H20 + MonO3n-2

(1)

Oxidation: MonO3n-2 + 02 --~ MonO3n

(2)

a

b

Figure 1. MoO6 octahedra in MoO3 (in the full 2-dimensional structure, each O atom is shared between 2 octahedra). The distortion of the real structure (a) is not shown in the conventional representation (b). Q Mo atom, O O atom

43

3.2.

Structural aspects

The octahedra in MoO3 can also adopt an arrangement different from that of Fig. 1. By losing one oxygen, they form edge-sharing pairs. The cyclic reduction-reoxidation of the surface of MoO3 therefore corresponds to the cyclic transformation of twin adjacent pairs of corner-sharing octahedra losing 2 oxygen atoms to form 2 adjacent pairs of edge-sharing octahedra. This is represented in Fig. 2.

c.2:c.-c. 3

CH2 = CH- CliO

..2o

| o2 Figure 2. Schematic representation of the cyclic reduction/oxidation of twin pairs of MoO6 octahedra between the comer and the edge-sharing arrangements (Boxes represent MoO6 octahedra with sharing of oxygen atoms at comers for MoO 3 or edges for MOO2). The figure is not completely accurate, because it cannot take into account the fact that the arrangements are not perpendicular to the main axes of the lattice. It turns out that the reduction by the olefin is more rapid than the reoxidation by 02 when pure MoO3 is used under ordinary oxidation conditions (this is suggested by the thickness of the arrows in Fig. 2). This occurs in spite of the fact that oxygen is in stoichiometric excess with respect to the selective oxidation. This imbalance in the kinetics results in a reduction of the surface of MOO3, which triggers a reduction in the bulk. Recent results obtained by Gai and Boyes (see Gai, ref. 11) show that a flexible structure which constitutes the most active and selective phase could be Mo18052. This is slightly reduced compared to MOO3; the presence of oxygen vacancies permits a smooth reduction-oxidation cycle. In Mo18052, two oxygens are removed out of three in some oxygen rows (one row out of 18). Each missing atom corresponds to replacing one comer-sharing arrangement of octahedra by one edge-sharing link. Thus, in Mo18052, this occurs for one Mo atom of the structure out of 18. Fig. 3a and 3b represent schematically the structure of fully oxidised MoO3 and MO18052. Considering in a general way the sub-oxides MoO3_x, the edge-sharing pairs assemble in shear planes going obliquely across the structure, as suggested by the arrows in Fig. 3. This is very close to our finding that the active structure might be Mo8023, which we detect in substantial quantities after work in catalysts having reached a high activity, and protected from overreduction thanks to the action of spillover oxygen during the catalytic reaction (12). Mo18052 and Mo8023 do not seem to differ very much in their behaviour in reduction and reoxidation.

44 If the reoxidation to MoO3 (Fig. 1) is too slow, further reduction occurs, leading t o other suboxides of molybdenum. These are less selective catalytically. Although not proven unequivocally, these suboxides are more difficult to reoxidise in the bulk (especially the ultimate term of the series, MO4Oll). This brings about a deactivation difficult to reverse. As will be mentioned later, the deactivation effect is still more catastrophic in binary oxides (molybdates) than in MOO3. Fig. 3 strikingly illustrates the fact that exactly the same reaction mechanism is responsible for activation (transforming MoO3 to Mo18052 or Mo8023) and deactivation (Mo18052 or Mo8023 to Mo4Oll represented in Fig. 3c). Generally speaking, avoiding deactivation implies that the reduction and reoxidation rates of Fig. 1 are perfectly balanced. Our contribution in a previous symposium on Catalyst Deactivation was to show that spillover oxygen produced by a "Donor" (e.g. Sb204) could re-establish this balance (10). Another possible strategy would be to exert an adequate strain to the catalytic oxide lattice to stabilise the loose, unstable state. This is probably the action that TiO2 supports exert on V6013 islands (13-15) and this could explain the unique role of TiO2 (anatase) in V205/TiO2 catalysts active in the oxidation of o-xylene to phthalic anhydride. I Ix3 b(l IX3 [x~ ~ ~ Ix3 ~
e

I~ ~ cx3~ cx3~ lx~~ ~ ~ ~ ~ ~ ~ ~ ~ I I IXi IX] IXI IXI IXI 1XI IX] IX] [Xl IXl IXl IX] 1Xl 1x] IX] IX] I I [ X ] IXl IX] IX] I ~ IX] IX] Ix]

3.3.

Mesoscopic

changes

The fact that the really active oxides, namely Mo18052 or Mo8023, correspond [XI ~ t~ ~ I~ ~ ~ 1 to loose, flexible structures, in which I ~ ~ ~ ~ g~ ~ ~ t~ ~ some small domains continuously change M~ 3 ~-- their structural arrangement, has still ~I q ~ other consequences than just the I = , = N IX! ~ I ~ ~ ~ ~ ~ ~ i ~ ~ ~x~ ~ structural modifications outlined above. g~ ~ a ~ ~ ~ ~ ~ ~ t~ ~ a ~ ~ I We have postulated that the continuous I IX] I IX] IXI IX] IX:] IX] IXI IX] IX] =.~a iX] IXI ~a ~ ~ ~ ~ ~ ~ ~ ~ 1 ~ ~ I changes between corner-sharing and b 1 IX] ~z,a IX] I>q IX] IXI IX] IXI IX] ~ IXI ~ t edge-sharing arrangements of MoO6 IX] IX] I IX] IXl IX] I ~ IX] IXl IXI IXl ~ IX! I I IX] r IXI IXI IX] IX] IXI IXI IXI IXI I IX] IXl octahedra could permit the progressive IX] IXl txBa IXI IX11XI IXI I,X] IX] IXl t~x= IX] I reconstruction of the surface withmhopeI IX] IXI I IX] IXI IX] IX] IXl IX] IXI IX] =~x= IXI fullymthe consequence that the number ~__ 18Mo "columns" ~_ ~ of active sites would increase. We Mo1aOc;9 (or Moa 09-~, with 8 columns) therefore predicted that the topography of the surface of MoO3 should undergo I ix=x= IXl I IX] =,:=a = ~ a IXI tXi ix=x= =a=~ IX] I IX] =,a~ I changes at the nanometer or micrometer I IX] I IX] = = a z.=xs IX] I IXI scale ("mesoscopic" scale) in catalytic IXI =~x= IX] I ~ r r I c I-IX] =~=a =xma IXI I IX] =mix. conditions where the loose, unstable R lXl I IXI r =':=a IXI I I =xmq ix=x= IXI ! IX] =xma r structure exists. Our experiments show I I IXI ==a ~cmx= IXI I IXI ~ ~=,a IX] I IXl DaDqf =z,~ ]XI I that. At the micrometer scale, steps appear on MoO3 crystallites (12,16). At the nanometer scale, in addition, the M~ surface develops fractures or, more Figure 3. Schematic representation of the structure of precisely, pits (17). This takes place only MOO3, Mo18052 and Mo4Oll. The same inaccuracy as in Fig. 2 impairs the present picture. The shear planes in when the dynamic, loose structure is Mo18052 and Mo4Oll are represented by the oblique maintained. Spillover oxygen maintains arrows. (Boxes with an "X" represent MoO5 octahedra), the catalyst active (i.e. in the dynamic

45 state) in conditions where the kinetic imbalance between reduction and oxidation would reduce MoO3 in the absence of spillover oxygen. This is possible because oxygen is always in excess in the gas phase as long as the reaction is selective and does not give too large amounts of complete oxidation products (H20 and CO2). This allowed a comparison and produced this striking result. Although our analysis of the dynamic phenomenon is not as complete as that involving changes in structure, we can speculate on the results obtained. This analysis strongly suggests another relation between the "life" of the real working catalyst and deactivation. To understand this relationship, we need a more detailed representation of the structure. This is presented in Fig. 4. A better understanding of the structure of MoO3 is obtained by comparing the relatively crude model of Fig. 3 and a more accurate one (Fig. 5). This latter figure makes 010~ netolane the MoO6 octahedra more apparent than in ~'Fig.. 4, and shows a shear plane (18). Returning to Fig. 4, the (010) face is totally non selective in the oxidation (100) netplane of olefins, producing only CO2 and H20. In addition, when crystalMO lographically perfect, this face has poor activity O (12,17). The active sites are situated on the (100) faces. Pits formed during the catalytic reaction are ~,, visible when looking at the (101) surface (nameMolybdenum trioxide, MoO 3 ,,~, ly looking vertically from the top in Fig. 4: this is Figure 4. Perspective view of the lattice of MoO3 (Balsac catalogue),

shown in Fig. 6). The pits (Fig. 6b) are holes or small crevices, the walls of which are composed of (100) planes (or other planes nearly perpendicular to the (101) face). The creation of pits does not occur under our reaction conditions in the absence of spillover oxygen, namely under conditions where MoO3 is overreduced (Fig. 6a). This can be understood when looking at Fig. 3 which shows that Mo4Oll, for example, has no twin pairs of comersharing octahedra. The dynamic reduction-reoxidation process presented schematically in Fig. 2 cannot take place in this case. The strong covalent links in the twin pairs of edge-sharing octahedra rigidify the structures. Incidentally, when looking at Fig. 5, it is apparent that shear planes contain more Mo-O-Mo bonds per unit volume, this suggesting intuitively that they make the structure less flexible.

46

o

b%

qB

i Figure 5. Representation of the layered structure of MOO3_x seen from the side. The shear structure shown (b) suggests that this induces the formation of steps or corrugations in the layers (18).

The discovery that pits form, namely at near-surface zones (we speculate the pits are about 5 nm deep) raises questions concerning deactivation. It is known that Mo-containing oxidation catalysts are mechanically fragile. Pit formation may trigger a fragmentation of MOO3, and consequently initiate deactivation by mechanical decay. But we must emphasise that this is still speculative. If this were the case, a second type of solid state reaction (admittedly linked to the first) would be involved both in activation (extension of (110) faces thanks to pit formation) and deactivation, supporting our view that the solid state reactions leading to the creation of catalytic sites (the creation of pits), permitting a steady state catalytic process and causing deactivation, are similar or identical.

3.4.

Possible consequences, possible remedies For many reasons, but principally in view of the fact that the highly efficient industrial ironmolybdate based catalysts for the oxidation of methanol to formaldehyde are composed of several phases and contain a very large amount of MoO3 besides Fe2(MoO4)3, we wondered whether MoO3 contributes significantly to the selective reaction. In view of our previous experiments, it could have been speculated that MOO3, by itself, is overreduced in the methanol-air mixture, and therefore deactivates quickly. We therefore tested MoO3 mechanically mixed with a-Sb204. Table 1 shows that, at 500~ the performance (at the laboratory scale) is equivalent to that of the industrial Fe2(MoO4)3-MoO3 catalyst (19). We attribute this spectacular result to the fact that the addition of a-Sb204 strongly retards the solid state reaction leading to deactivation of MOO3, very likely keeping the MoO3 surface in a higher oxidation state, thanks to spillover oxygen produced by aSb204.

Table 1 Activity of MoO3 and a MoO3 + a-Sb404 mixture (MM) in the selective reaction of methanol to formaldehyde (formaldehyde ~r conversion, and selectivity). Temperature ~

Yield % MM MoO 3

Conversion % MM MoO 3

350 42.4 60.5 42.3 82.0 400 68.1 67.8 68.8 78.6 500 93.5 63.8 94.0 82.1 Gas feed: methanol 6.5%, oxygen 19.6%, nitrogen 73.9%, gas hourly MoO3 174 mg + a-Sb204 174 mg; MoO3 174 mg. (19)

Selectivity % MM MoO 3 100 73.8 99.0 86.3 99.5 77.7 space velocity 26,220 h -1, MM:

47

Fr-e',...sh r t o O 3 - U

(Middle

o.f

the

crys:tal)

.8O0

.6O0

400

0

L:~Q

r~~.::~i~]ii!~-~. ~ + S b 2 0 4 .~,

Figure 6.

~

403

a ~"t, e r

600

r e a c : t, I or~

(. c,d.qe ..... r "~ ,. . . . .

. ....

(contact) pictures of (a) pure MOO3, (b) M o O 3 in an (~-Sb204 mixture ((z-Sb204

crystallites are not visible because they were swept away by the AFM tip) after the catalytic reaction (T 420 ~ isobutene:O2:He = 1:2:7; 3 hr). The pure sample is unchanged after the reaction with the isobutene - 02 - He mixture. The addition of a-Sb204 to MoO3 does not alter the surface of MoO 3 before the reaction (17). There is good reason to believe that similarly interesting effects could be observed with molybdates of various metals, which are widely used as selective oxidation catalysts. The postulated solid state reactions of MoO3 can serve as a relevant model for the phenomena occurring in these catalysts. The structures of many molybdates are very similar to that of MOO3, including the possible presence of shear planes (for a rich source of information on this, see, for example, several chapters in Ref. 20). Although activation, catalytic reaction and deactivation mechanisms have not been studied in as much detail as those concerning MOO3, analogies are apparent.

48 With respect to structural changes in molybdates, we examined Fe2(MoO4)3 in a previous Symposium of this series (1). Too deep of a reduction of the "active state of the catalyst" leads to deactivation by segregation to MoO3 and FeMoO4. In a catalyst maintained at a higher level of oxidation (in our case, by using spillover oxygen), selectivity is much better and segregation does not occur. In terms of the mesoscopic changes, the main cause of deactivation of the Fe2(MoO4)3 based catalysts for the oxidation of methanol to formaldehyde is the formation of fines which progressively plug the reactor (21,22), a phenomenon probably linked to a formation of pits similar to those occurring in MOO3. The conclusion concerning the close similarity of solid state reactions explaining the selective catalytic action of certain phases and at the same time their deactivation should not be viewed negatively. On the contrary, the very positive aspect is that this concept should serve to focus future research on the fundamental problem, namely protecting the loose, dynamic state, keeping it "wobbly". This can be accomplished through relatively minor measures, like creating some 80 epitaxy, adequate doping or the use of spillover species and probably other means. A laboratory 60 result justifies this point of view. We selected #, BiPO4 to produce spillover oxygen and compared the behaviour of pure Fe2(MoO4)3 ~o over a long run with that of a sample containing BiPO4 introduced by impregnation of 0,,-Fe2(MoO4)3 (Fig. 7). The stabilising effect by 20 80 this route on Fe2(MoO4)3 is not of long enough duration to produce a practical catalyst, but this example shows that ageing can be made less rapid. More generally, this strongly suggests g 65 J that lines of research based on the identity of "'o--.--r o,... solid-state processes in the various stages of catalyst life are relevant to the problem of catalyst deactivation. 0 " I I 0

20

z,O

60

80

Figure 7. Comparison of activity and selectivity 4 . A BRIEF L O O K AT O T H E R of FeE(MoO4)3 (8.1 meg-1) (O) with the same SIMILAR CASES catalyst impregnated with BiPO4 (19 wt.%) (O) in the oxidation of methanol to formaldehyde. BiPO4 alone is completely inactive (Fe2(MoO4)3 The above discussion, centered on MOO3, weight 350 mg in both cases; T = 280~ highlights the role of reduction in deactivation. CH3OH:O2:N2 = 4.90:2.38:92.72 (vol.); feed Other types of solid-state reactions may take rate 50 cm 3 mm-1), place in a catalyst under working conditions. An instructive case concerns mixtures of t~-Sb204 and Fe2(MoO4)3. A strong synergy between these phases is observed in the selective oxidation of isobutene to methacrolein (23,24). No detectable mutual contamination of the surface of the partners, nor solid state reaction occurs during the first few hours of contact with the reacting gases. But additional phases appear after a few days (FeSbO4 and MOO3) leading to new synergies with the remaining unreacted phases, namely Fe2(MoO4)3 and a-Sb204 (25,26). In these experiments, the formation of the 4-phase system

49 corresponds to an activation stage. Because nucleation of the new phases plays a crucial role in the kinetics of the solid state transformation, the overall surface area of the catalyst is increased. There is a possibility that a longer experiment would show further rearrangements and, possibly, a loss of surface area corresponding to deactivation. The 4-phase system clearly needs further investigation. Nevertheless, the example suggests at this stage that solid state reactions leading to new phases could play a role in the deactivation of selective oxidation catalysts and that the phenomenon could be narrowly related to the genesis of phases which interact synergetically with each other in catalysis. A recent example also highlights a deactivation which is clearly related to the formation of a new phase. This concerns the oxidative dehydrogenation of propane to propene. Mg2V207 reacts with the donor of spillover oxygen tx-Sb204 to make MgSb206 which has low activity (27). The transformations of vanadium phosphate ("VPO") catalysts active in the oxidation of butane of maleic anhydride can also trigger interesting speculations. Recently, the evolution of VPO phases over several hours of contact with reacting mixtures of butane and oxygen has been magnificently described (3). This is very important, in view of the fact mentioned in the introduction that the catalysts gain activity and selectivity during this period. It is a pity that adequate detailed studies of the deactivation period have not yet been published. This would permit a comparison. However, we would like to risk a speculation with respect to the phenomena and their consequences. The fact that a strong deactivation occurs, with a dramatic loss of selectivity, strongly suggests that the reactions initiated during the activation stage do continue. An optimum balance in the cooperation of the various phases probably occurs after some time on run, and further evolution is probably detrimental. A pertinent question can then be answered: why is the so-called "riser" technology in butane oxidation so attractive? In this technology, a discontinuous regime is established in which the vanadium phosphate catalyst is alternatively contacted with butane, which gets oxidised (while the catalyst gets reduced) and in this technology, a discontinuous regime is established in which the vanadium phosphate catalyst is alternatively contacted with butane, which gets oxidised (while the catalyst gets reduced) and with oxygen. The reoxidation step in this cycle could perhaps bring back the catalyst to a phase composition just slightly before the optimum (similar to that at the end of the activation stage). The other part of the cycle, namely the contact with butane, is probably terminated just after the optimum. This is just mere speculation, but a solid state chemist could find it reasonable. It has been repeatedly postulated that various phases in VPO catalysts are in epitaxy with each other. Direct experiments have failed to prove this postulate. But we can reason that the key to success for the "riser" technology is that nuclei of all the useful phases are kept during the reduction-oxidation cycle. The fact the X-ray diffraction indicates these phases in small quantities might have been mistaken for epitaxy. Very likely, these nuclei would be in close contact (if not really epitaxy) and enable a harmonious cyclic operation.

5.

CONCLUSIONS AND O U T L O O K

The speculative remarks we make here are aimed primarily at suggesting perspectives based on the theme of the present contribution, namely (i) that there is continuity in the solid state reactions taking place during activation, catalytic work and deactivation, and (ii) that we can take advantage of this continuity to achieve both more selective reactions and to limit the inconvenience of deactivation.

50 The line of reasoning we have outlined above can certainly be applied to other cases. Some work in progress suggests that the problem of loss of active material in selective ammoxidation catalysts (e.g. TeO2 in MoTe catalysts) can be approached similarly (28). This might also be the case of bismuth-promoted palladium catalysts for the selective oxidation of glucose to gluconic acid (29). Work carded out in other laboratories especially with heteropolyacid points in the same direction. This general approach can also be taken in the case of hydrotreating catalysts. Just to cite one example, the "life" of the MoS2 sulfide crystallites in hydrotreating certainly explains both activity and loss of surface area, this in turn leading to deactivation (30). In conclusion, we propose that solid-state reactions occurring at all stages of catalyst life should be placed in a comprehensive frame. There are many proofs that these solid state reactions are identical or similar. Funding of practical solutions to improved activation, better catalyst performance in terms of activity and selectivity and longer catalyst life can be justified more easily if the corresponding problem is put in the general perspective we propose.

ACKNOWLEDGEMENTS

The author thanks Dr. P. Ruiz for his advice during the preparation of this manuscript.

REFERENCES

1. 2.

3. 4. 5. 6. 7. 8. 9. 10. 11. 12. 13. 14. 15. 16.

B. Delmon, P. Grange, in "Catalyst Deactivation" (B. Delmon, G.F. Froment, eds.) Elsevier, Amsterdam, 1980, 507-543. B. Delmon, P. Grange, in "Progress in Catalyst Deactivation" (J.L. Figueiredo, ed.) (NATO Advanced Study Institute Series, Series E, no. 54), Martinus Nijhoff, The Hague, 1982, pp. 231-280. C.J. Kiely, A. Burrows, S. Sajip, G.J. Hutchings, M.-T. Sananes, A. Tuel, J.C. Volta, J. Catal., 162 (1996) 31. J.C. Volta, Catal. Today, 32 (1996) 29. G.K. Boreskov, Kinet. Katal., 21 (1980), 15. F.J. Barry, D.J. Smith, J. Catal., 88 (1984), 107. B. Delmon, J. Mol. Cat., 59 (1990), 179. M.A. Van Hove, in "The Nature of the Surface Chemical Bonds" (T.N. Rodin, G. Ertl, eds.), North Holland, Amsterdam, 1979, pp. 277-311. G. Somorjai, M.A. Van Hove, Progr. Surf. Sci., 30 (1989), 201. B. Delmon, in "Catalyst Deactivation 1994" (B. Delmon, G.F. Froment, eds) Elsevier, Amsterdam, 1994, pp 113-128 P.L. Gai-Boyes, Catal. Rev. - Sci. Eng., 34 (1992) 1. E.M. Gaigneaux, P. Ruiz, B. Delmon, Catal. Today, 32 (1996) 37. T. Machej, M. Remy, P. Ruiz, B. Delmon, J. Chem. Soc. Faraday Trans., 1990, 86, 715722. T. Machej, M. Remy, P. Ruiz, B. Delmon, J. Chem. Soc. Faraday Trans., 1990, 86, 723730. T. Machej, P. Ruiz, B. Delmon, J. Chem. Soc. Faraday Trans., 1990, 86, 731-738. B. Delmon, Chimia, 51 (1997) 31.

51 17. E. Gaigneaux, E.E. Wolf, P. Ruiz, B. Delmon (6th Iketani Conf., Tokyo, Nov. 25-27, 1996), accepted, Appl. Suf. Sci. 18. "Molybdenum. An outline of its chemistry and uses" (E.R. Braithwaite, J. Haber, eds.), Elsevier, Amsterdam, 1994. 19. R. Castillo, K. Dewaele, P. Ruiz, B. Delmon, Appl. Catal. A, accepted. 20. Solid State Chemistry in Catalysis (R.K. Grasselli, J.F. Brazdil, eds.), ACS Symp. Series 279, Am. Chem. Soc., Washington, D.C., 1985. 21. N. Burriesci, F. Garbassi, N. Petrera, G. Petrini, N. Pernicone, in "Catalyst Deactivation" (B. Delmon, G.F. Froment, eds.), Elsevier, Amsterdam, 1980, pp. 115-126. 22. N. Pernicone, Catal. Today, 11 (1991) 85. 23. X.L. Xiong, L.T. Weng, B. Zhou, B. Yasse, E. Sham, L. Daza, F. Gil-Llambias, P. Ruiz, B. Delmon, in "Preparation of Catalysts V" (G. Poncelet, P. Grange, P.A. Jacobs, B. Delmon, eds.) Elsevier, Amsterdam, 1991, pp. 537-546. 24. Y.L. Xiong, R. Castillo, Ch. Papadopoulou, L. Daza, J. Ladriere, P. Ruiz, B. Delmon, in "Catalyst Deactivation 1991" (C.H. Bartholomew and J.B. Butt, eds.) Elsevier, Amsterdam, 1991, pp. 425-432. 25. L.E. Cadus, F.J. Gotor, D. Acosta, J. Naud, P. Ruiz, B. Delmon, Solid State lonics, 1993, 63-65, 743-747. 26. L.E. Cadus, Y.L. Xiong, F.J. Gotor, D. Acosta, J. Naud, P. Ruiz, B. Delmon, in "New Developments in Selective Oxidation Ir' (v. Cort6s CorberS.n, S. Vic Bell6n, eds) Elsevier, Amsterdam, 1994, pp 41-54 27. S.R.G. Carraz~in, C. Peres, M. Ruwet, P. Ruiz, B. Delmon, Proc. XIIth Int. Symp. Reactivity of Solids, Hamburg, Preprints, 5PO 304, submitted for publication. 28. S. Zeyss, P. Ruiz, B. Delmon, submitted. 29. M. Wenkin, R. Touillaux, P. Ruiz, B. Delmon, M. Devillers, Appl. Catal. A, 148 (1996) 181. 30. B. Delmon, Bull. Soc. Chim. Belg., 1995, 104, 173-187.

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~ Elsevier Science B.V. All rights reserved Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

53

C o k e F o r m a t i o n in Catalytic Processes: Kinetics and Catalyst D e a c t i v a t i o n G.F. FROMENT Laboratorium voor Petrochemische Techniek Universiteit Gent Krijgslaan 281 9000 Gent, Belgium.

1.

INTRODUCTION

In spite of its far reaching industrial implications catalyst deactivation by coke formation has been studied with insufficient depth and rigor. A better understanding of the interaction between the hydrocarbons and the catalytic sites leading to coke formation and a more detailed modeling of its rate might contribute to the development of catalysts generating less coke and/or less subject to deactivation. To reach this goal it has to be kept in mind that the study of coke formation and the associated catalyst deactivation is not different from the study of any catalytic process; coke is formed through one or more reactions, involving intermediates also acting in the "main" reactions. The influence of the coke on the reaction(s) causing catalyst deactivation is then naturally dealt with in terms of the real variables and the elementary steps. Presently this is still done in an empirical way by multiplying the rate of the main reaction(s) by a deactivation function which is often related to time, tO =fit)

(1)

where e.g. j'( t) = e -at

although it would be preferable to relate it to the true deactivating agent, coke itself : tp =f(Cc)

(2)

where e.g. f ( C c ) = e -ace

An exponential function of the coke content has been derived for a variety of processes [De Pauw & Froment, 1975; Dumez & Froment, 1976; Hatcher, 1985; Beirnaert et al, 1994]. The approach symbolized by (1) is easier since it does not require measuring the coke content of the catalyst, as (2) does. However, measurement of coke content is no longer a problem; adequate equipment has been developed that permits the simultaneous study of the main reaction(s) and coke

54 formation, namely an electrobalance with or without recycle of the products and the TEOM [Beirnaert et al, 1994; Chen et al, 1996]. The shortcomings of the approach symbolized by (1) have been extensively discussed [Froment, 1976]. Briefly, if coke is not measured, the deactivation of the coking reaction itself can not be accounted for, so that (1) is really a very biased equation. This approach cannot predict the coke profile in a reactor, although the knowledge of this profile in an industrial reactor is of key importance for the operation of the catalyst regeneration. When the coke content of the catalyst is measured a distinction can be made between the deactivation function for the main reaction, CpA,and that for the coking reaction Cpcdefined by : FA - " FA(DA

(3)

= rc%Pc

(4)

and rc

where rA~ and rc ~ are the initial rates, i.e. the rates at zero coke content. These are modeled in terms of the Hougen and Watson approach, accounting for the competitive adsorption of the reacting components and coke. Any progress in understanding the effects of coke on the rates of the reactions beyond the empirical level of the type of equation (2) requires more information on the structure of the coke and of the catalyst. In other words: is the effect of coke limited to the coverage of sites? Or can it grow to a size blocking the pores? What is the pore size distribution of the catalyst? Are the pores interconnected? The present paper aims at reviewing progress in the modeling of coke formation and catalyst deactivation along the lines set in Table 1, although it stresses to a large extent the kinetic formulation, an area in which considerable progress has been achieved. Table 1. Classification of phenomena occurrin~ in coke formation at various levels. Micro-level Phenomena 1) Uniform sites coverage (+ growth) 2) Clusters of uniform sites coverage (+ growth) 3) Sites of different nature coverage + migration Meso-level 1) Pores 2) Networks of pores

coverage + growth + pore blockage + diffusion in the structure

Macro-level Reactor

+ flow + mass & heat transfer

55

2. C O K E F O R M A T I O N ON T H E M I C R O - L E V E L : S I T E C O V E R A G E AND COKE GROWTH

2.1. Rate equations for simple, ideal cases For a simple reaction A ~ B occurring on single sites and with the reaction between adsorbed species as the rate determining step the classical Hougen & Watson approach leads to [Froment & Bischoff, 1990]:

Ct_Cc

l kACtKA( Ca - ~ )

ra =

=r

C,

(5)

1 + K AC A + KBC B

and for the rate of irreversible site coverage by coke formed out of A (parallel coking)

rcl = Ctr s =

kcCtKACA

C t --Ccl Ct

1 + K AC A -~- K s C e

=

rpr~

(6)

where KA is replaced by KB in the numerator of (6) and KACA in the denominator by KBCB when the coke precursor is formed out of B (consecutive coking). The coke precursor is defined as the first of the components in the coking sequence which is irreversibly adsorbed. The rate of coke formation is related to the rate of irreversible fractional site coverage, rs, through:

rc = e~CtMcr~ = dC~ dt

(7)

provided there is no multilayer growth of the coke. For a given gas phase concentration integration of (7) yields:

(8)

Cc = 6 C t M c ( 1 - e x p ( - r ~ revealing that ~ in the empirical relation (1) is not a constant. Also, q)A =rPc = e x p ( - r ~

C~

= 1 - ~

~Mc

(9)

For dual site reactions the relation between {p and Cc is non linear. Even with single site reactions a non linear relation would be obtained if there were a distribution in the strength of the sites. The deactivation functions {PAand {Pc would not be identical if a different number of sites were involved in the main and in the coking reaction.

56

2.2. Catalyst deactivation by site coverage simultaneous with coke growth The main reaction is affected only through the site coverage. For given PA, PB.... the following formula was derived by Beeckman & Froment [1979] for the deactivation function of the main reaction: ~Pa : e x p ( - r o t )

(10)

When the coke grows through some polymerization process at a rate rp which is independent of the degree of polymerization: (11)

r

+

r:

(1_ exp(_rOt) )

(12)

,Mr~

9A differs from 9c and is no longer a linear function of the coke content Cc. Whereas 9A is a measure of the coverage and always decreases, 9c may exceed one, depending upon the value of the ratio rp~ ~ Only when all the coke molecules have the same size can 9c be directly related to the site coverage.

2.3 Rate equations for complex industrial processes In steam reforming for synthesis gas production from natural gas on a Ni/alumina catalyst, coke is mainly formed by the cracking of hydrocarbons. For a methane feed the main reactions and the coking reactions can be represented by: CH4 + H20 ~ C H 4 + 21120 ~

CO + 3 H 2

(13)

C 0 2 + 4H2

CO + 1-120 ~ C02 + 112 CH 4 ~ C + 2H 2

(14)

Fig. 1 represents the reaction mechanism involved in H 2, CO and CO2-production derived by Xu and Froment [1989] and by Snoeck and Froment [1997] for coke formation. In the steam reforming sequence (2) - (9), leading to CO, the rate determining step is (8), in the sequence (2) - (7), leading to CO 2 it is step (7), in the water gas shift step (10). In the coking sequence (2) - (3) - (4) - (12) - (13) the rate determining step is the first H-abstraction, (3). By way of example the following rate equation for the formation of CO was obtained by Xu and Froment [ 1989]:

'k' Ken, "KH~o( rco

-

P~Pco )

H, K1 ( Pn~~ 2 p2n"~ 1+ KcoPco + Kn, pn' + Ken, Pen' + Kn~o Pn2 )

(15)

57

Through an analogous procedure Snoeck and Froment [ 1997] obtained the following equation for the coke production by methane cracking 9 kn re=

KCH4Pen,

- kzKcKn2 Cc~,.i "P2H2

(1+ ~cCc~,,. + KcCc~,.,. p~,'~ + ~c,, pc,, )

(16)

2

where CCNi,f is the coke content of a further evolution of C-1 in Fig. 1, to be defined below.

IH2OI

I CH4 ]

Ol+l

| +l _.,+ 1 H2-l ~

IL

2H+I

+ O-I

CH4-I

| CH3-I + H-l |

I

CH2-I + H-l o-,

|

O-I

H-I + CO2-I

r

|

|

CH-I +

H-I

CH20-I + l /

CHO-I + H-l

| CO-/+ H-l

O-I

QI"

(~)

CO2-1 + l

--O+l

+l

Figure 1. Reaction mechanism of the conversion of methane into H 2, CO, CO 2 and coke Clearly the rate equation for the cracking of methane, i.e. for the coke formation is not fundamentally different from that of one of the main reactions (12). What remains to be done is to link the coke content of the catalyst to the rate of the main reactions. Thereby a specific aspect of coke formation on Ni/alumina catalysts has to be accounted for, namely whisker formation. The rate equation (16) is not directly applicable because it contains the concentration of coke adsorbed on the Ni-surface, which is not accessible, just like Celia.l, CH_1.... The latter are eliminated through adsorption-isotherms in favor of the measurable gas phase partial pressures PCH4, Pn2, but this is not possible for coke. What is done in the derivation of (5) and (6), where the same problem is already encountered, is to manipulate the expressions so as to factor out the Ccl, thus yielding the

58 ratio ( C t - C c l ) / C t , defined as the deactivation function [Froment & Bischoff, 1990]. This cannot be done here because of an additional complication: carbon segregates on the side of the Ni-particle facing the gas, where its concentration is represented by CcNif, dissolves in the Ni, diffuses through it and precipitates at the rear of the particle, lifting it 'progressively from the alumina support, thus leading to a whisker carrying the Ni-parficle at its top. This was also accounted for by Snoeck and Froment [1997] who obtained an implicit rate equation for r c containing the diffusivity of carbon in Ni and the saturation concentration of carbon in Ni, assuming that the supersaturation is low. Therefore, at steady state there is no deactivation. The latter is only observed when the Ni is encapsulated by coke growing on coke, but the modeling of this aspect remains to be developed. Chemical processes encountered in petroleum refining are far more complex than steam reforming. The feedstocks are complex mixtures of hydrocarbons of various types each of which is converted along intricate pathways. So far these processes have been modeled in terms of lumps. In catalytic cracking the 3-lump and the 10-lump model [Nace et al, 1971; Jacob et al, 1976] are still widely used although the lumps are based on boiling ranges rather than on chemical nature. These models contain in general only one deactivation function of an empirical nature for the reactions of the various lumps. In their study of the catalytic cracking of n-hexane on a US-Y-zeolite in an electrobalance with recycle Beimaert et al, [1994] derived an empirical deactivation function of the type (2) for the various reactions, but with different o~-values, as illustrated in Table 2 for the isomerizations. Table 2 Estimated r

of the deactivation functions of some isomerizations in n-hexane cracking.

Product

r

95 % confidence interval

2-Me-pentane

1.65

1.44 < tz < 1.83

3-Me-pentane

1.76

1.51 < r < 2.02

2,3-di-Me-butane

1.81

1.57 < 0~ < 2.04

2,2-di-Me-butane

2.62

2.36 < tx < 2.89

Coke

0

Note that the coke formation itself is not deactivated. This is probably due to the growth of coke, which is not hampered yet in this zeolite with its relatively large pore size, unlike that observed on ZSM-48 [Froment et al, 1990]. Fig. 2 shows a reaction scheme for the cracking of n-hexane on a US-Y-zeolite, illustrating again how the main reactions and the coke formation are closely related and involve similar reactions [Froment, 1996]. The monomethylpentanes and also the dimethylbutanes are rigorously lumped, since the equilibrium between the components is rapidly established. The alkylation of nC 6 and its isomers leads to a lump C k which cracks or cyclicizes and dehydrogenates into aromatics and into coke, either directly or through alkylation of the aromatics. The network contains 5 reacting components involved in 14 reactions belonging to 6 types (4 for the main reactions, 2 for the coke formation) so that 6 rate coefficients have to be determined. If all the reactions occur on the same type of sites and involve the same number of sites all the rate equations have the same denominator and the number of adsorption equilibrium constants amounts to 14 (3 for the hexane isomers, 6 for the saturated C 3 to C 5 molecules and their corresponding olefins, 1 for the Ck-lump and 1 for the aromatics lump). When the deactivation is expressed in terms of an

59 exponential function with respect to the coke content, 6 additional parameters are introduced, one for each type of reaction, excluding the coke formation which is not deactivated. H2,

Cl,

C2, C3

M e - C~

C~ + C3

C~

diMe - C7

(6)

(3)

(I)

n - Ca

(8)

(7) (2)

M e -C s

diMe - C(

(9)

(10)

Ck

(13)

(12)

(11) C7 +

Ck.~

(14)

coke

Aromatics

Figure 2. Molecular reaction scheme for the catalytic cracking of n-hexane. c~

C~ + H 2, C I, C2, C3

(4)

Me- C;

MeCs + C~.

C; + C;

(6) 1

(7)

1(8)

(2)

Me-C;

(3)

diMe- C~

diMe - C4 + C~ (IO)

(9) 1

(5)

n-C6T~ C, +n-C~

(11)

(14)[

(12)

diMe-C:

(13)

c~ (15)I

(c;), + (c;). (= coke) (16)]

[ (17)

Aromatics

Figure 3. Reaction network in terms of elementary steps of carbenium-ion chemistry for the catalytic cracking of n-hexane. Clearly, applying this "molecular" approach to the cracking of a complex mixture like vacuum gas oil is an impossible task: the number of parameters would be overwhelming. Indeed, even in a homologous series of n-paraffins each molecule reacts at a different rate; the rate coefficient for the disappearance of the molecule increases with chain length.

60 The only way out is to express the transformation of the molecules in terms of the elementary steps of carbenium-ion chemistry: protonation, H- and Me-shifts, protonated cyclo-propane-type isomerization, 13-scission ... [Froment, 1996]. In Fig. 3 the cracking of n-hexane is represented in terms of the single event approach introduced by Froment and co-workers [Baltanas et al 1989; Vynckier and Froment, 1991; Froment, 1991; Svoboda et al, 1995]. Excluding CH3 + and primary carbenium ions which are far less stable than secondary and tertiary carbenium ions, leads to 17 rate coefficients for paraffin cracking in general and 26 adsorption equilibrium constants, therefore a total of 43 parameters, which is more than the 26 parameters in the molecular model. The superiority of the single event approach becomes evident when mixtures have to be dealt with. The rate coefficients of the elementary steps of the carbenium ions may be assumed to be independent of the chain length, so that the number of parameters does not increase for complex mixtures, unlike what happens with the molecular approach. Analogous assumptions are made within the naphthenes- and aromatics-families so that the total number of parameters entering in the kinetic modeling of the cracking of a vacuum gas oil and the associated coke formation is of the order of 40 to 50. It should be added here that these huge reaction networks have to be generated by computer [Baltanas & Froment, 1985]. Booleans relation matrices were used to describe the molecules and carbenium ions. Since a component analysis of a vacuum gas oil is not entirely feasible some lumping is inevitable but the rate coefficients for the reactions between the lumps can be constructed from those of the single events entering in the reactions of the components of the lump [Vynckier & Froment, 1991 ]. 2.4. Clusters of uniform sites For multi-site reactions a site is only active when it is surrounded by the required number of vacant sites. The probability of such a configuration decreases as more sites become irreversibly covered. Also, certain reaction steps are only possible when the sites are covered by a specific set of molecules. These aspects are not accounted for in the phenomenological approach introduced in Section 2.1. Beeckman et al. [ 1987], pursuing work by Nam and Froment [ 1987], developed an approach accounting explicitly for the topology of a cluster of sites occupied by reacting species. Using probability calculus they derived expressions for % and qu in terms of the degree of coverage, co. When the coking reaction and the main reaction involve more than one site the deactivation function develops a dependence on the number of sites in the cluster. Both reactions now become structure sensitive. When the number of sites required by the coking reaction is higher than that of the main reaction a high metal dispersion- a small cluster size - decreases the selectivity for coke formation. Such a trend is well known in metal supported catalysis. The selectivities of the various main reactions can also be influenced as illustrated by Fig. 4. The topological model provides guidelines for optimizing the cluster size so as to achieve maximum selectivity for the main reaction and a maximum resistance to coke formation. 2.5. Sites of different nature: Dual function catalysts The catalytic reforming of naphtha and the hydrocracking of heavy oil fractions are based on dual-function catalysts. These contain a metal component (Pt, Pt~e, Pt/Sn in reforming, Pt, Pd, Ni/Mo in hydrocracking) and an acidic component, (alumina in reforming, alumina or zeolite Y in hydrocracking). (De)hydrogenation occurs on the metal component, isomerizations and cracking on the acid component. In dual-function catalysis species formed on one type of site have to move to the other before the next step of the reaction sequence can take place. Evidently, proximity of both types of sites is required, although it has been shown that even a physical mixture of the two catalyst components can achieve the desired result. As sites are being covered by coke produced by side reactions the rate of the reactions decreases, but since the distance between sites remaining active increases, an additional effect has to be accounted for. Singh and Froment [1997] recently developed a probabilistic model for the deactivation of a dual function catalyst by coke formation accounting for reaction and surface migration. The reaction sequence was the isomerization of n-

61

into i-pentane on Pt/alumina, studied in an isothermal tubular reactor by De Pauw and Froment [1975] and represented schematically in Fig. 5 in which ~ stands for Pt-sites, x for acid alumina sites and coke is formed on x-sites only. All the reactions are assumed to occur on single sites. 1.0

0.8

0.6

% (dual) (single) 0.,

q)A

0.2

N=2

;

oj2

.~=3 ,..,

0:,

o.~

N..4

~=5

o.B

(.O Figure 4. Evolution of the ratio of rates of the main reactions on dual and single sites with degree of surface coverage by coke. i-C5

n-C5

n-Cs.o"

-- n-C~.o" i i

,_

n-Cs.'r i ,

T

C

- i-C~.'r i !

T

,

i-C~.a

-'- i-Cs" o"

i i

C

Figure 5. Schematic representation of a process consisting of various adsorptions, desorptions, reactions and migrations. A square grid of x and o-sites was selected with only one site on each node of the grid. The migrating species randomly chooses its direction upon encountering a site which is covered already by a reacting component or by coke. Each migration step is limited to the adjacent site only. The expected distance for the migration of n-C5-- can be written:

L1

=(1-Ps''/dPs,,)

(17)

62 where d(A) is the length of a segment and Pf~s is the conditional probability that a site encountered by n-C: is a free x-site. A similar formula is derived for L2, the distance of migration of i-C5 =, with replacing:. The conditional probability Pewscan be written 9

p., , =

P//,~P~ + Po/oPo + P//,P, + Po/,P~ + Pd/,P,

(18)

where Pf/,, Po/~ are the probabilities that a ~-site is free or occupied and Pf/~, Po/~, Pd/~ are the probabilities that a x-site is free, occupied by the reacting species or by coke and, therefore, deactivated. The probability that a site randomly located on the catalyst surface is a x-site e.g. is given by: where Ct is the total concentration of both ~- and x-sites and C,, is the total concentration of x-sites. The probability that this x-site is free is given by: p, = C,~

(19)

ct P:/, = ~

(20)

and the probability that it is occupied by the reaction species:

Po,,

Cc~" + C/c~ -

-

(21)

The probability of a x-site being deactivated by site coverage is calculated from [Beeckman & Froment, 1979]"

Pd/, = [1-- exp(--r~

(22)

Assuming coke formation from i-Cs--only, the parameters were estimated from the data of De Pauw and Froment [ 1975]. Fig. 6 shows the evolution with the coke content of the migration distances in a point of the reactor where the partial pressure of n-pentane has decreased from its inlet value of 0.757 atm to a value of 0.675. Initially, i.e. at zero coke content L~ and L 2 amount to 3.31 and 2.47 ]k respectively. The migration distances vary with the coke content as a consequence of two opposing effects: the coverage of x-sites by coke formed from i-C5-- that increases L l and L 2 and the deactivating effect of coke on the reaction n-C:x ---) i-C:x that decreases the coverage of x-sites by i-Cs--, thus reducing L 1 and L 2. The magnitude of the variations of L 1 and L 2 depends on the local concentrations of n-C 5 and i-C 5 in the reactor. 3. COKE F O R M A T I O N ON THE MESO-LEVEL The case of site coverage and coke growth was already discussed in Section 2.2. Coke can grow to a size sufficient to block even mesopores [Levinter, 1967]. Since there is no preferential location for the coverage of a site and the subsequent blockage of the pore Beeckman & Froment [ 1979] used a probabilistic approach. The deactivation function is nothing but the probability that a

63 site is still active, q) equals the product of the conditional probability S that the site is not covered with P, the probability that it is accessible: = PS

(23)

,,~

I- I

%

.,,,~

~

2,

o

o.ooG$ o.r

o.ooIs

o.o(3:1 o.co2s

o.o~3 0.0035 o.oG4

coke content (gcoke/g,,) Figure 6. Evolution of migration distances L~ and L 2 with coke content on a dual function catalyst.

When the growth of the coke is instantaneous, which is the case when in (12) rp~ > 8CtMRs ~ in other words when the rate determining step in the coking sequence is the production of the coke precursor, all the coke has the same size, determined by the pore diameter. Blockage of the pore causes coke profiles even when there are no concentration gradients of the reacting species. Also, q)A : ( P C "

In the more general case when the potential rate of polymerization is of the same order of magnitude as the potential rate of site coverage, two periods have to be distinguished: a first ranging from t = o to t = tb, the time required to reach a size sufficient to block the pore and during which only site coverage and cokes growth occur and a second, starting at t = t b, during which blockage occurs and the deactivation also depends upon the site density, as shown in Fig. 7. It is very questionable if a single pore is sufficiently representative for a catalyst particle in which the pores are interconnected. Beeckman and Froment [1979] modeled a particle of an amorphous catalyst in terms of a Bethe tree, which is a branching structure without closed loops. The type of tree they considered has a coordination number of three, while the pores are characterized by two stochastic variables: their length and their diameter. The probability P in (23) now becomes a much more complex function. The model was applied to a meso-macro porous structure representative of a chromia-alumina catalyst used in butene dehydrogenation and for single site coking and instantaneous growth of coke. CpAthen equals (Pc but their value drops to zero long before all the sites are covered with coke, because of pore blockage. The approach was applied to the butene dehydrogenation data of Dumez and Froment [1976] by Matin et al [1986]. Beeckman and Froment [ 1982] also extended the approach to account for diffusional limitations. For consecutive coking a saddle-like coke profile through the particle was predicted.

64

,oo\

1.00

~

parameter

: o" L - 0.80

- 0.60

0.40

0.20

0.00 -

0.00

i

i

0.16

i

i

i

i

032 0.48 x I0=

i

1

0.64

!

tO.O0

0.80

cc Figure 7. Averaged deactivation function for the main reaction as a function of averaged coke content for several site densities ~ in a single ended pore with length L. A Bethe-tree is a particular case of more general networks considered in percolation theory. Sahimi and Tsotsis [ 1985] applied percolation theory and Monte Carlo simulation to deactivation in zeolites, approximated by a simple cubic lattice. Beyne and Froment [1990, 1993] applied percolation theory to reaction, diffusion and deactivation in the real ZSM-5 lattice. The finite rate of growth was described in terms of a polymerization mechanism. Pore blockage was reached in this small pore zeolite. It also affects the path followed by the diffusing molecules that becomes more tortuous, so that the effective diffusivity has to be expressed in terms of the blockage probability.

0.75

~

.

i

~.

0.5 0.25

00

........

0.02

'

'

' 0.04i

,

,

. 0.06

A

Cc (kg coke/kg cat) Figure 8. Variation of (q)A)d and (tpc)d with the coke content for consecutive coking in a ZSM-5 particle. Fig. 8 illustrates the variation of (q)A)d and ((pc)d with the coke content for consecutive coking. The index d indicates that, besides deactivation, diffusional limitations also affect the rate of transformation of A. The steep drop in activity corresponds to a coke content at which the

65

percolation threshold has been reached. Curves of this shape have been reported in the literature [Dejaifve et al, 1981]. Recently Gottifredi and Froment [ 1997] dealt with reaction, diffusion and coke formation in a catalyst particle, mainly focusing on the derivation and application of approximate but accurate formulae for the calculation of transport fluxes and deactivation functions. In the case of parallel coking and an exponential deactivation function the time behavior profile of the rate of the main reaction can be represented as in Fig. 9. Whereas for short times the highest rate is observed at the surface, with increasing reduced time, x, an internal maximum develops, the location of which shifts towards the center of the particle. Indeed, for parallel coking the coke content is always highest at the surface, since the diffusion limitation leads to a concentration profile of A which decreases towards the center of the particle. The deactivating effect of the coke is strongest at the surface. Inside the particle the coke content is lower and the rate is less affected so that a maximum in rA~ develops. The figure is similar to that derived by Froment and Bischoff (1961) for deactivation by coke formation in a tubular reactor without diffusional limitations inside the catalyst particles, however.

rA~

h2o- 1O0

Figure 9. Time behavior of the profile of the effective rate of the main reaction inside a catalyst particle for parallel coking. 4. THE MACRO-LEVEL:

DEACTIVATION IN A REACTOR

Beyne and Froment [ 1993] simulated a tubular reactor with plug flow and diffusional limitations inside the catalyst for the process discussed already in Section 3. The main reaction is of the type A ~ B and coke is formed through a polymerization mechanism from a site covered by coke

66 precursor onwards. Transport, site coverage and blockage in the ZSM-5 catalyst is dealt with through percolation theory and effective medium approximation. Fig. 10 shows how for parallel coking the prof'lle of the effective rate of reaction of A through the reactor evolves with time. Initially the highest effective rate (the rate in the presence of diffusional limitations) is located at the inlet, but as time evolves a maximum develops that moves downstream through the reactor. For parallel coking the coke content of the catalyst is always highest at the inlet of the reactor and that is where the rate of disappearance of A is more affected, leaving a larger fraction of A to be converted in downstream zones as the run length increases. Such a wave does not develop with consecutive coking. As shown in Fig. 11 in such a case and pure A feed the coke content increases through the reactor, thus enhancing the importance of the upstream zone. With increasing time a smaller fraction of the reactor is participating in the reaction. Such a behavior was predicted by Froment and Bischoff [ 1961], but without including diffusional limitations inside the catalyst particle.

0.75

..............

r* A,, e l f '0.5 r A,, o f f

1:=0 z=0

0.25

0

0

0.4

0.8

1.2

z(m)

i.6

Z

Figure 10. Evolution with time of the effective rate profile of the main reaction in a plug flow reactor. Parallel coking. Diffusion-limited process on a ZSM-5 type catalyst. 5. C O N C L U S I O N S Studies of catalyst deactivation by coke formation should be conducted along the lines followed in the study of the main reactions of the process. By doing so advantage can be taken from the analogy between steps pertaining to both the main- and coking reaction pathways. The insight gained by such an approach will also permit one to distinguish between the deactivating effect of coke on various types of reaction and to model the evolution of the selectivity in terms of the increasing coke content. Catalyst deactivation has such an impact on the design and operation of plants that studies along the approach outlined here are fully justified. They should be an integrated part of the process development.

67

\

:

t (iO

si)

i

t

:

!

i

i

i

i

0.75

i

i

i

9

1

1

i

i

!

i ......................

i ...........................

!

r

A,

A, e l f

elf

I

0.5

1

:.

I t=0

I Z=0 0.25 ....................

.........................

| ,

.

,

i

l

o.s

: i

,

.

,

I

t

: ,

i

.

,

i

t.$

i

i

|

t

2

zfa~ Figure 11. Evolution with time of the effective rate prof'de of the main reaction in a plug flow reactor. Consecutive coking. Diffusion limited process on a ZSM-5 type catalyst.

REFERENCES Baltanas M.A., K.K. Van Raemdonck, G.F. Froment and S. Mohedas, L & E.C. Res. 28, 899910, 1989. Baltanas M.A. and G.F. Froment, Comp. Chem. Engng. 9, No. 1, 71, 1985. Beeckman J.W. and G.F. Froment, L & E.C. Fund. 18, 245-256, 1979. Beeckman J.W. and G.F. Froment, Chem. Eng. Sci. 35, 805, 1980. Beeckman J.W. and G.F. Froment, L & E.C. Fund. 21,243-250, 1982. Beeckman J.W., In Sik Nam and G.F. Froment, Catalyst Deactivation. B. Delmon & G.F. Froment (Eds), Elsevier Sci., Amsterdam, 1987. Beirnaert H.C., R. Vermeulen and G.F. Froment, Proc. 6th Int. Congress on Cat. Deact. 1994. Eds. B. Delmon & G.F. Froment. Studies in Surf. Sci. and Catalysis, 88, 97-112, Elsevier Sci., 1994. Beyne A.O. and G.F. Froment, Chem. Eng. Sci., 45, 2089-2096, 1990. Beyne A.O. and G.F. Froment, Chem. Eng. Sci., 48, 550, 1993. Chen D., A. Gronwald, H.P. Rebo, K. Moljard and A. Holmen, Appl. Catal. A, 137, 11-18, 1996. Dejaifve P., A. Auroux, P.C. Gravelle, J.C. Vedrine, Z. Gabelica and E. Derouane, J. Catal. 70, 123, 1981. De Pauw R.P. and G.F. Froment, Chem. Eng. Sci., 80, 789, 1975. Dumez F. and G.F. Froment, L & E.C. Proc. Des. & Devpt. 15, 291, 1976. Froment G.F., Proc. 6th Int. Congr. Cat. 1, 10-31, Chem. Soc. London, 1976. Froment G.F., Rev. Inst. Fran~ais du P~trole, 46, 491, 1991. Froment G.F., Chemreactor-13, June 1996, Novosibirsk. Froment G.F. and K.B. Bischoff, Chemical Reactor Analysis and Design, 2nd Ed., 1990, J. Wiley, N.Y. Froment G.F. and K.B. Bischoff, Chem. Eng. Sci., 16, 189, 1961. Froment G.F., J. De Meyer and E.G. Derouane, J. of Catal., 124, 391, 1990. Gottifredi J.C. and G.F. Froment, Chem. Eng. Sci., 1997.

68 Hatcher W.J., Ind. Eng. Chem. Prod. Res. & Devpt., 24, 10-15, 1985. Jacob S.M., B.E. Gross, F.E. Voltz and V.W. Weeckman, A.LCh.E.J., 22, 701-713, 1976. Levinter M.E., G.M. Panchenkov and M.A. Tanatarow, Int. Chem. Eng. 1, 23, 1967. Marin G.B., J.W. Beeckman and G.F. Froment, J. of Catal., 97, 416, 1986. Nam I.S. and G.F. Froment, J. of Catal., 108, 271-282, 1987. Reyes S.C. and L.E. Scriven, I & E.C. Res., 30, 71,1991. Sahimi M. and T.I. Tsotsis, J. Catal., 96, 552, 1985. Singh S. and G.F. Froment, Kinetics of Her. React. & Dynamics of Surfaces, Eds. G.F. Froment & K.C. Waugh, Elsevier, 1997. Snoeck J.W. and G.F. Froment, J. Catal., 1997. Svoboda G.D., E. Vynckier, B. Debrabandere and G.F. Froment, L & E.C. Res., 34, 22, 25282530, 1995. Vynckier E. and G.F. Froment, Kinetics & Thermodynamic Lumping, Eds. G. Astarita & S.I. Sandier, Elseviers, Amsterdam, 1991. Xu J. and G.F. Froment, A.L Ch.E.J., 35, 88, 1989. NOTATION concentration of reacting species, kmol/mag coke content of catalyst, kg coke/kg cat total concentration of active sites, kmol/kg cat.h rate coefficient, kmol/kg cat.h adsorption equilibrium constant, m3g/kmol or bar~ migration distances,/~ molecular weight of coke, kg/kmol partial pressure of component i, bar probability that a site is accessible rate of reaction of A, kmol/kg cat.h rate of coke formation, kg coke/kg cat.h rate of growth of coke, kg coke/kg cat.h fractional rate of site coverage, h-1 probability that an accessible reaction site is active time, h

C A, C B

Cc Ct k K A ' "'" L I , L 2 , ..

Mc Pi P rA rC rp rs

S t GREEK

LETTERS

(%)d'(%)d G CO

deactivation constant conversion factor, kmol coke/kmol sites deactivation function for main and coking reaction deactivation functions in the presence of diffusional limitations (de)hydrogenation sites in dual function catalyst isomerization sites in dual function catalyst degree of surface coverage

9

Elsevier Science B.V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

69

Catalyst Deactivation: How We Cannot Yet Subvert Nature John B. Butt Department of Chemical Engineering, Northwestem University, Evanston, IL 60208 USA This paper in many ways presents the author's 35 year encounter with catalyst deactivation in various forms and disguises. It is a combination of personal experience with historical development and scientific progression. Overall developments are surveyed, and in a way how they impacted one worker in the field. Old results are reported in an historical context, but some new ones are also given. There are some lessons: 1) It is impossible to understand catalyst deactivation without knowing the chemistry of the reactions and the catalysts involved. 2) It is impossible to understand catalyst deactivation without knowing something about the interactions of reaction and mass transport kinetics. 3) It is impossible to understand catalyst deactivation without knowing something about chemical reaction and chemical reactor dynamics. Number 3) above renders us into the uncomfortable realm of unsteady-state phenomena - we must realize that this is quite long-term as well as short-lived. "Even with Chee-Chee's help as a translator, I couldn't keep track o f it all."

1. INTRODUCTION "A long, long time ago, in a Galaxy far, far away ..." starts off George Lucas' classic STAR WARS trilogy. Sometimes I feel the same way about what we are here to discuss. In some ways, through the years of experience, I think that this whole business of catalysis, and of catalyst deactivation, reflects something of a " ... long, long time ago ... ". Yet, by our very presence, it is not; the arenas of inquiry are solid and growing, so I will claim that we are very much in the present. In fact, detailed studies of catalysis, and especially of catalyst deactivation, belong pretty much to the 20th Century. How much will this extend to the 21st? Thus far it is true that " ... he who ignores history has to learn it all over again...". In this sense, please allow some personal recollections. I did my PhD Dissertation on the use of alumina as a catalyst (don't laugh; this was 1956), particularly as applied to the dehydration and dehydrogenation of alcohols. Since methanol was uncooperative, and propanol too complicated (pre any good chromatography), I settled upon man's most beloved compound - ethanol. The first thing I found out, after two years of struggle in building a recycle reactor (maybe a predecessor of a modem CSTR), was that the alumina at about 300 ~ or so was super at removing water from ethanol, giving lots of ethylene and small other bits of strange things. I also found out, though, that the alumina became fired after a while and the ethylene yields decreased substantially. What was this? I mean, this is a chemical reaction and chemical reactions do what chemical reactions are supposed to do! Maybe it was obvious that something else than ethanol, water and ethylene was involved, but I guess I was blind to anything else at the time. Let us revisit 1956. Wheeler had published his great paper on deactivation and diffusion in 1955 far too soon for it to penetrate into the dusky laboratories of Yale University. Professor Herman

70 Pines had done much work on properties of alumina and modifications thereof as catalysts - but that was chemistry, not chemical engineering. In justification, if for no other reason, I would like to ask how many chemical engineers in 1956 read the JOURNAL OF THE AMERICAN CHEMICAL SOCIETY? Zero might be close. The first thing I had to learn was " ... what is a Lewis Acid ... "? That meant I had to read some chemistry - really serious stuff. Then I found out that the term "alumina" does not define a specific, unalterable substance. To say "alumina" is like saying "people". At that time Professor Pines was perhaps 60+ years and I was 20+; he was in Evanston and I in New Haven; however the measure of the man is that he was kind enough to spend much time in correspondence to explain to me all those papers in JACS and to tell me what a Lewis Acid is - and maybe even a Lewis Base. I did learn, though, that what the "alumina" did to the ethanol depended upon the "alumina". As Professor Burwell told me much later, and which I think I understand only now, " ... these are not compounds, they are materials ... ". Any catalyst, defined as you like in chemical terms, is a MATERIAL, not a well-defined chemical substance. It is therefore subject to any number of events, physical or chemical, that can alter its behavior. My old work with ethanol employed a commercial alumina, HARSHAW A1-0104 | and I had no idea as to the exact chemical composition. When I started making my own alumina, things became clearer because I could observe, at the least, different appearing materials resulting from variations in preparation - thus, maybe not all the same stuff!.. (This is not to denigrate HARSHAW; their alumina was a respected and widely used material, but it was a MATERIAL). Where does all this little bit of history lead us? It is certainly not just to say that some materials called "alumina" can catalyze the dehydration of ethanol to produce ethylene. It is to say that product distributions with the various alumina preparations were different, and changed in different ways as the time of operation increased. In the long run the lesson learned from this work was to recognize, however imperfectly that catalysts are materials, they undergo solid state transformations, they are subject to chemical attack from feed stream impurities, they suffer thermal stress, maybe they are susceptible to cancer- at any event, most of them eventually die. It's like a medical history: why and how?

2. SOME BACKGROUND I think, personally, that it is impossible to understand catalyst deactivation in its many forms and disguises without knowing first off about diffusion and reaction in catalyst particles, and knowing about chemical reactor analysis. The Laws of Nature make all diffusion-controlled reactions want to look like first-order, however complicated their real kinetics might be. As a good catalyst designer one likes to utilize all the interior surface of the material; this strategy may end up with operation close to diffusion control, but not always. From this we have many lessons; for example in the areas of reforming and hydrotreating catalysis. It is, though, in this proximity to optimization of the use of materials where we seem to become perplexed when deactivation inevitably occurs. One of the worst maxims, repeated in almost every introductory physical chemistry text is that" ... a catalyst is a substance that can promote the rate of a chemical reaction without itself being changed ...". That is like saying that L. Borgia is now a Saint. Catalysts do chemistry, for sure, but they HAVE chemistry as well. But, like Lucretia, they are fairly complicated. Maybe the best way to start here is with a good empirical observation of VOORHIES [ 1], who noted that in the deactivation of various catalysts used for petroleum processing the activity history was well tracked by an expression of the form: (Activity) = A ( t f = a

(1)

71 By now this basic formulation has had many interpretations. For example: (Activity) has been used to refer to coke-on-catalyst, amine index of the material, reference to conversion in some specific chemical test; who knows what else. The value of n, reported in various studies as ranging from 0 to 12, has been represented to indicate diffusion control (0.5) up to essentially " ... we don't know what is going on here ...(12) ... ". The factor A is a proportionality constant specific to catalyst, operating conditions and chemical reaction. Voorhies' model, based on time-on-stream observations, is obviously not general, but it is a good place to start.

3. O N E W A Y TO L O O K AT IT A reasonable and useful definition of activity is: Activity = (Rate of reaction of Main Reactant)/(Rate at t = 0) = a

(2)

which fits in with the Voorhies relationship. If one is allowed this latitude, then it is possible to try exploring a complex set of kinetics, such as the hydrogenation of benzene on Ni or Pt. This reaction generally is known to be described by the rate equation: (3)

(Rate)B7 = (kPHPBz)/(1 + KBzPBz) at low hydrogen partial pressures. If we take a literal interpretation of eqn. 2, then one can write:

(4)

(Rate)Bz = (a)(kPuPBz)/(1 + KBzPBz)

Then, all that is necessary is to take care of a. That tums out not easy to do because, like Lucretia, a is a product of chemistry. We probably can guess, knowing about catalysts, that a starts at unity, and eventually goes to zero. How it does so, and at what rate, all depends on the chemistry of the reaction, the chemistry of a, and the reaction conditions. There are few general rules.

4. S O M E IDEAS A B O U T "a"

Many ideas come to mind when one tries to think about this elusive variable. Professor Petersen and I wrote a book about it, and we thought we really had an idea. Unfortunately, the older I get, the more fleeting the idea seems. However, I do think that we can pin down the elusive a in some cases. One case is deactivation by poisoning. This is because the catalyst is subject to a specific chemical deactivation; thus for the classical academic reaction A --+ B we have a parallel scheme such as A+S-+B+S

(Main Reaction, k)

P + S--+P.S

(Deactivation Reaction, kd)

where P is a poison, S some active catalytic site, and P.S the poisoned (deactivated) site. Following simple kinetics in each case and utilizing the definition of a in eqn. (2) we have rA = -kCgCs - -k_CAa

(Main)

72 r,~r

=

(Deactivation)

-k,CpC~ -kdCra =

or, for the deactivation (da/dt) = -kuC~

(5)

where it is assumed that the rates of both the main and deactivation reactions are dependent upon a. We are not limited to first order kinetics so that more generally it is possible to have a deactivation rate equation of the form (da/dt)

=

-lq(Cp)m(a)w

(6)

with m and w the orders with respect to poison and current activity, respectively. For example, sulfur poisons platinum and nickel pretty much in accord with the kinetics of eqn (5). In cases of deactivation by coke formation we may also have a parallel mechanism in which A is the coke precursor (these are limiting cases, of course). Correspoonding deactivation rate equations follow from development similar to that for eqns (5) and (6). Let us now retreat into another classical result, that of Bohart and Adams [2], that sort of indirectly addresses the parallel poisoning mechanism above. They really did not have catalyst poisoning in mind, rather adsorption in fixed beds, but the situation is analogous in many respects. With parallel poisoning (or coke formation) there can develop activity waves - interior regions of active and inactive catalyst that pass through the bed - and the entire process is unsteady state. Such waves are shown in Figure 1; however the observer standing at the bed exit measuring conversion vs. time of reactor operation

1.0

(a)

1.0

t---~

(b)

Deactivated

_

T

Q.

J~~eakthrough

o" 0 . 5

Quasi - Steady-State

Position

Time

Figure 1. (a) Bohart-Adams poisoning wave profiles, and (b) Conversion-time behavior for a sharp profile. would not know that there was a wave going through the bed until it reached the end, and then might be very confused. A very simple and elegant example of this has been provided some years ago by Froment and Bischoff[3]. Here we take a reaction, parallel in deactivation, and examine the evolution of activity within the bed. One sees in Figure 2 the evolution of an activity maximum that develops and passes through the bed, and a puzzling maximum conversion that comes and goes with time-on-stream Now this does not correspond exactly with the Bohart-Adams analysis, but in reality it is not all that different. In fixed-bed catalyst deactivation, there are waves of all kinds.

73

Length

Time

Figure 2. Rate surface for a parallel deactivation model. Now Bohart and Adams and Froment and Bischoff were concemed mostly with isothermal conditions in their studies. What if we have an exothermic reaction? In that case there will be thermal waves as well as concentration waves. Here there can be a sharp reaction front, riding the concentration wave as it were, giving a very pronounced thermal wave, almost like a spike, passing through the bed. An example of this, taken from our studies with benzene hydrogenation on supported Ni (Weng, Eigenberger and Butt [4]) is shown in Figure 3.

!

--e--140

P --- I 0 0 1-

!

!

!

!

G3

180 -

.

(o)

meosured

--e--

colculoted

i40

- ' ~ -

~

|t

w

2

!

w

(b).

meosurld colculoted 1

-

~

N

=--- I O 0

1-

60

-

60 --

20

G3

;--

I

I

inert

I

-~:

..

I

active

-.. i

~;:

4t

-

20

5 !

inert-~t

I

1

inert

:I~

I

octive

!

N: i n e r t - - ~

Figure 3. Profile simulation for hydrogenation of benzene in a nonisothermal fixed-bed. It should be remembered that these thermal waves can be large enough to cause catalyst sintering, thus activating another mechanism of deactivation. Did the observer at the end of the bed see any evidence of deactivation? Of course not, because the active reaction zone was actually confined to a small part of the bed, conversion is very good there, and until the thermal wave passes out of the end of the reactor, there is no evidence as to what is going on. Then immediately conversion goes to zero and the unprepared observer may find it necessary to seek some other sort of employment.

74 5. THE ROLE OF DIFFUSION There was a brief mention earlier about the role of diffusion in all this, and indeed it is very important. Mass transport can well limit the rate of main reaction-transformations, but it can also limit the rate of deactivation, particularly in parallel mechanisms of both poisoning and coking reactions. Diffusion/reaction has been a favorite topic for a long time, but many forget one simple thing: mass transfer is FIRST ORDER. Complicated reactions that are first order are often diffusion obscured. Poisoning or parallel coking reactions, however, are often first order. One can bring forward many more complicated schemes, particularly with coking reactions, but we do not try to do that now. How does diffusion influence the course of deactivation? This is fairly simply stated in the literature, but it depends upon the mechanism of deactivation and the extent of diffusional resistance. If we have a strong poison, one that is very rapidly taken up on the catalyst surface, pore diffusion often controls the overall rate and there is "pore mouth poisoning". If there is one not so strongly controlled by diffusion, then we have "uniform poisoning". If there is a substance that penetrates completely to the mid of the catalyst particle, then we have "core poisoning". All of these are very different, as shown in Figure 4. "Uniform Poisoning" is probably the best example. Let us assume that we

Pore Mouth

I

I

Pore Center

Pore Mouth

Figure 4. Distributions of foulant in the catalyst pore. have already a diffusion-controlled reaction (or near one, for a good catalyst designer); then Effectiveness Factor = rl = tanh r

(7)

But, this is not enough, since the reaction is subject to poisoning. We assume that the reaction is of a case where the intrinsic activity is related to fraction unpoisoned, thus: a=f

(8)

and that value should be used for the first-order rate constant of the unpoisoned rate constant.

75

Thus, for the poisoned case we obtain (9)* We can now get the activity as a function of fraction poisoned as before,

a

=

(O'~(tanh(r

r

(10)

Eqn (10) is obtained by taking kf as the rate constant for the poisoned catalyst. This is an important result, and is shown in Figure 5 for a range of parameters. One sees the linear variation of a with f for uniform poisoning at low ~, and significant variations from linearity for larger ~. Note that in the latter case a is higher at a given value of f

. o ~ r------t

i

" .~

r-------T

T

0.6

0.4

0.2 O

0

0.2

0.4

0.6

"0.8

1.0

Fraction Poisoned, I - f

Figure 5. Decrease in activity for different types of poisoning. (a) Uniform poisoning, low e; (b) Uniform, high o; (c-e) Core poisoning, increasing o; (f-g) Pore-mouth poisoning. Pursuing an analysis similar to that for the uniform poisoning case one can obtain for activity in the case of pore mouth poisoning ap = [tanh(fr162

[ 1 + (1-f)r162

1

(11)

For core poisoning, one obtains ao = [tanh(fr162

(12)

where ap and ac are acfm'ties for pore mouth and core poisoning, respectively. The important message in this treatment is that the apparent activity of a pellet can vary greatly with the type of poisoning taking place if diffusion rates influence the overall reaction rate. In fact, " dpis defined here as the ratio = L(2k/rD)1/2,where L is the diameter of the pellet, k is a first-order rate constant, r is the radius of the pellet, and D is some effective diffusion coefficient. Effective diffusion parameters are another thought altogether,but beyond the scope of the present discussion.

76 the activity relationship to (l-f) can be anywhere in the scale range 0-1 depending upon the extent of diffusion and the type of poisoning, as shown in Figure 5. There have been many approaches to this problem; our discussion above has concentrated on poisoning as the mechanism mostly because of space limitations in a presentation such as this, however it is important to recall the work of Masamune and Smith [5] treating examples of coke formation in diffusion-limited pellets according to both parallel and series mechanisms. They solved the appropriate conservation equations numerically and were able to demonstrate in some detail the evolution of coke as well as activity profiles within the pellet as a function of time of operation. The types of profiles they demonstrated for both parallel and series coke formation mechanisms, and their evolution with time, are shown in Figure 6 for ~ = 5, in the region of significant diffusional limitation; the profiles are indeed steep. Note that the parallel mechanism leads to profiles akin to pore mouth poisoning, 2.0

1.8

~, 1.6 8 1.4

" " :

"

I.O 0 . 8 ~

0.8 0.6 0.4 0.2

0

0.6 f 0.4 0.2 0

1.2 1.0

1.0

f I 0.5 0

0.5

1.0 0.8 0.6 0.4 0.2 0

I

1.0 0.8 0.6 0.4 0.2 0

0.5 0 0.5

Figure 6. Profiles for self-fouling, parallel and series mechanisms. and the series mechanism to profiles suggestive of core poisoning. Of course, at lower values of the Thiele Modulus these steep gradients will level out and the pellet, for either mechanism of coking, will become more uniformly deactivated. The inclusion of the time of operation by Masamune and Smith strongly indicates the unsteady-state nature of deactivation, was recognized in the early work of Voorhies. The time scales and corresponding real times-on-stream are discussed by Butt and Petersen [6]. A final topic we shall mention concerning the deactivation of isothermal pellets is the effect of nonuniform distribution of the active ingredient within the pellet. It has been known for sometime that this is possible - wittingly or unwittingly. This is particularly so for metals supported on porous oxides, e.g., alumina. We will take one simple example, the analysis of Shadman-Yazdi and Petersen [7] for self-fouling by a series mechanism in which the concentration of catalytic material was highest on the exterior of the pellet and fell to zero at the center according to the equation

~(~) = a0~5

(13)

where a0 is the activity at the exterior of a spherical pellet, ~ is the dimensionless radius, and 6 a distribution parameter. We normalize the distribution such that

77

~.o1 a ( ~ ) d ~

which yields

= 1

k = k~(a)(~)

(14)

where k~ is the average rate constant. The rate of reaction at any time, OB, is given by r(0B) = Vo~kaCa0rl

(15)

which reduces to r(0) = VpskaCA0n

(16)

where V is pellet volume, p, the surface area per unit volume, 1] the effectiveness factor of the unfouled catalyst, and the apparent activity of the pellet. Values of and 1"1 as a function of the Thiele Parameter ~ are shown in Figure 7 for two differing distributions. Here ~ is defined as (Lpsk/D) v2 and 0B as k~CA0t. These results indicate that distributing the active ingredient towards the exterior improves the resistance of the catalyst to poisoning if a series mechanism operates, since both the effectiveness factor and average activity are higher with higher loadings near the surface (dotted line). One might expect the opposite to be the case for a parallel mechanism, and in general intuition serves us well for these simple models, c.f other early work of Corbett and Luss [8] and Becker and Wei [9]. There have been many studies since then. Experimental verification of such analyses for parallel poisoning has been reported by Au et al. [ 10]. It should be remembered, however, that intuition concerning distributions may not serve one well in selectivity problems involving complicated sequences of reactions.

1.0

/N

0.6

s

0.4

,5

I-"---.._

-------4.__'_

I

'

I , l

I

~.o < e

0.2

0

I

,

It

I,I

,

I'Otr ~ ~ - - . ,o

0.81.. 0.6

.e_>

0.4

----o(~) = 3

0.211

, I~l~l 0.40.6

i

' , ',,

,

'

"~ ~ . . . ~-~

I

I 2

L

* 4

, I 6

810

Thiele Porometer: 6

Figure 7. Deactivation for nonuniform impregnation, series mechanisms.

78 6. SOME EXPERIMENTS A considerable amount of experimental information has been obtained for diffusion-reactiondeactivation in pellets by the use of so-called "single pellet diffusion reactors". These come in two forms: the first, originally developed by Balder and Petersen [11] ("Petersen's Pellet Poisoner"), relies upon the analysis of concentrations of reactants and their variation with time; the second, from Kehoe and Butt [12] ("Kehoe's Katalyst Killer") involves measurement of temperature profiles within the pellet and their variation with time. The first is useful for deactivation in isothermal systems; the second, at the expense of more complexity, can be used for both isothermal and nonisothermal systems. It is not possible to go into a long discussion of these systems here, but some discussion of the more simple reactor, for isothermal systems, will be useful as an example. The isothermal diffusion reactor provides a means for measuring the composition of the reaction mixture at the center of the pellet and near the external face. The reactor is designed to

R..ct..t. ~ Co_

R.=,..,. ~ Co-

r

I

I

X+AX

c,o..c,o.. I

X*AX

(a)

(b)

Figure 8. Comparison of Thiele-Zeldovich model (a), with the single pellet diffusion reactor (b).

mimic the original Thiele-Zeldovich model of diffusion and reaction in a porous pellet, and a comparison of the model with the reactor is shown in Figure 8. In the reactor the centerplane chamber contains no catalyst thus the concentration gradient is forced to zero, which corresponds to the mid-point of the pellet in the model, L. Now, when deactivating pellets are studied, the centerplane concentration will change with time, as shown in Figure 9. C(O,O)

Progressive

c(t z, t) c(t s , i) t:O

~ I 0

0

C(O, I)

Dim,nsionless Distance Into Pore. F : X / L

Figure 9. Changes in certerplane concentration with progressive deactivation.

79 Information on deactivation can be obtained by defining a dimensionless centerplane concentration: n = [~(z, 1)-qt(0,1)]/[ 1-~)(0,1)]

(17)

where ~(z,1) and ~(0,1) are centerplane concentrations at times x and zero, respectively. ~ is useful because it can be measured without explicit reference to time, and because analytical expressions can be derived from various pore poisoning models as those discussed above. For example, for uniform fouling one can obtain U~= [ 1/cosh(r

1/cosh~)]/[ 1-1/cosh~]

(]8)

and since we know the relationship between and f, then can be represented directly as a function of ft. Other relationships can be obtained, for example, for pore mouth poisoning, and then a simple comparison of f, or , behavior with fl gives a good idea of what is going on. An example is shown in Figure 10 for deactivation in the hydrogenolysis of cyclopropane on Pt/alumina. It is easy to see that this deactivation is somewhere intermediate between pore mouth and uniform fouling. To read a little more from the result, one might infer that the cyclopropane, rather than some product, is responsible for the deactivation. Obviously, by combinations of experiments at different conditions and with different particle dimensions, one might be able to identify a rather comprehensive picture of the deactivation sequence involved in this reaction system. This method has subsequently been refined in the work of Hegedus and Petersen [ 13]. 1.0 L

,

,

,

,

'~ : 3.2 Exper,mentol Doto _

0 . 8 ~

0.6 ~ - ~ h ~ <0>

0

0

P~176 0.2

I 0.4

0

[ 0.6

O~

1.0

Figure 10. Comparison of experiment with uniform and pore-mouth poisoning. In practice, however, many reactions involve significant heats of reaction, and analysis based on assumption of isothermality is only approximate. In principle, a concentration decrease in the pellet is always accompanied by a temperature change unless the reaction is absolutely thermally neutral. An estimate of the maximum influence of nonisothermality can be obtained from the relationship of Prater [ 14] (AT)max = (-AH) (Do~)Co/K~

(19)

where Co is the reactant concentration at the external surface of the pellet, D~r the effective diffusion coefficient within the partite, and Kar some effective thermal conductivity of the particle. Now, these thermal effects can indeed be important. Figure 11 shows some results from the hydrogenation of benzene on a Ni/kieselguhr catalyst in a single pellet reactor at various

80 concentrations of benzene in the feed and at several different flow conditions. The conditions of the experiment are specific, of course, but one can see that intraparticle temperature gradients of 4050~ are not unusual. One would have to conclude that much of the original theory of reaction/diffusion/deactivation developed on the basis of isothermal theory might have to be rethought (and much of it has been). Sagara et al. [15] redid the work of Masamune and Smith using similar kinetic/deactivation models, but admitting for intraparticle temperature effects. The results are difficult to describe simply because the parameterization of the analysis increases significantly. We now have activation energy parameters for both main and deactivation reactions, as well as a quantity normally termed the Prater Number: 13= (-~-t)D,~Co/I~frTo

(20)

However, it is apparent from the results of Sagara et al., and from subsequent experiments that nonisothermality has a profound influence upon the development of intraparticle profiles in the presence of deactivation, as will be seen below.

Ioo-

~

I O0 - ~ -

~

T,*C 8 0

SL/mi,

.

4

" ' -

13.o3"t,_ -

6

3.90

I00 _lcl

1:3.48~"

80 . -

.

i

Bulk 1.0

i

0.8

t.

1

i

1

0.6 0.4 = r/r 0

t

I

0.2

0

Figure 11. Intemal and extemal temperature profiles for benzene hydrogenation in a single pellet reactor. One of the nice things about thermal measurements in a single pellet reactor is that they allow one to observe transients. The transient behavior of particles upon initiation of a poison-influenced reaction provides a particularly good example of this. Figure 12 shows some results. These again are for the benzene hydrogenation poisoned by thiophene and show the responses for three catalysts of differing initial activity levels. Benzene and hydrogen are introduced to the catalyst (remember these are all single pellet experiments) at time zero and 200~ The maximum activity catalyst, (A), shows the development of an internal temperature that grows to about 210~ from the initial 100~ in about 15 min. The 40% activity catalyst, (B), attains an internal temperature of about 165~ on the same basis, but the steady state time, 15 min is the same. Even the low activity, (C), catalyst, at 15%, appears to approach steady state within the same time frame, with an observed temperature of

81 135~ vs. the start at 100~ Thus we have a significant effect of the level of activity upon the magnitude of intraparticle temperatures, but not upon the time response. None of this behavior shows up in the Prater Number. The key to understanding time response comes from a detailed solution of the transient heat transfer equation. The dimensionless time will take the form (Lee et al. [16]): KenC/pC0(r0)2 = 1

(21)

where t is the characteristic time of response for the system and p the catalyst density. It is tempting to conclude, then, that the level of deactivation under transient conditions significantly affects the magnitude of the change (we look here at temperature; presumably concentration would be similar), but has minor effects on time scale. This needs more study. Not shown here, but of importance, is the fact that the magnitude of the ratio (K~dDon) is very significant in determining both the level of deactivation and the thermal response of individual catalyst pellets. Particularly for exothermic reactions, this will define where the final intraparticle exotherm resides, and since it is a ratio of quantities, it is difficult to discern the contribution of individual factors in general. This is shown in detailed calculations by Lee et al. Finally, here, one can look at intraparticle waves. A close look at Figure 12 shows that there are certain maxima ("hot spots") in the thermal profiles that develop and then disappear as steady state is

:

,.ovL- t

[ o.o

o

o

Figure 12. Effect of deactivation on startup transients for benzene hydrogenation. approached. In fact, there are all sorts of thermal waves that have been observed in transients with particles subject to deactivation. Some of these start at the pellet exterior, migrate to the interior and then remain there. Others, as shown in the results of Figure 12, stay close to the exterior of the pellet, and so on. The motion of these transient hot spots has been found to be a strong function of the Damkohler Number of the system. This is defined as noND~ = Rd(k~a)Co

(22)

where 110 is an overall effectiveness factor as defined by Carberry [ 17], 1% is the observed rate of reaction, kga the mass transfer coefficient extemal to the catalyst pellet surface, and Co the reactant concentration. Results of this for experiments with benzene hydrogenation are shown in Figure 13.

82

6 .c ,,,,r | 4 E I-

01.0 0.8

0.6

0.4

0.2

0

Figure 13. Variations in hot spot migration patterns on startup. One sees that there are really wandering hot spots dependent upon the parameters of the reaction system. I prefer to think of these as various types ofthennal waves within the partide. There is one interesting fact; however the transients set up operation, and however one obtains steady state, under deactivation conditions, with time, most pellets follow the Voorhies correlation, more or less. In view of the catalyst particle transients, it is not then unusual that one might look for waves in reactor operation, our final topic.

7. DEACTIVATION IN FIXED BEDS

We have already touched upon this topic in a qualitative manner in mentioning the work of Bohart and Adams and Bischoff and Froment. It is may be worth some additional effort to consider the origin of these results while minimizing as best we can the exposure to deadly equations. Wheeler and Robell [18] were the first to make use of the Bohart-Adams model which, as we have said, was developed for a fixed-bed adsorption problem. They used the adsorption model to establish the shape and motion of the poison profiles (c.f. Figure 1) within the bed and then essentially solved a reaction/deactivation problem based on these profiles (in their case a simple first order system). However, it was also necessary to specify how the reaction rate constant, k, varied with the concentration of poison on the catalyst. In the simplest case we can choose a linear relationship (nonselective poisoning) as they did and write = a = 1- Cp/Cp,tot

(23)

where ko is the initial first order rate constant, C~ is the concentration of poison on the catalyst, and Cp,tot is a limiting capacity of the catalyst for poison, where a = 0. Upon solution of the conservation equations one obtains the relationship between reactant concentration and position/time-on-stream At the bed exit, L ln(CL/Co= (-ko/lqd~)ln{1-exp(-Nt0/0t)+exp[Nt(1-0/0t)] }

(24)

83

where Nt = (kadsL/V) is the number of adsorption transfer units in the reactor, 0t = pBCp,totL/Mv(Cp, g)o is the ratio of the total capacity of the catalyst in the bed for poison to the rate of introduction of poison, v the feed velocity, OB catalyst bulk density, M molecular weight of poison and (Cp,g)o entering concentration of poison in the gas (fluid) phase. This result was also generalized by writing eqn (23) in terms of the selective poisoning model of Wheeler [ 19], giving an equation for conversion too long to be comfortable on these pages, and further generalized by Haynes [20] for diffusion-controlled poisoning. It should be recognized, however, that useful as these resuks may be, they do not represent a solution to the simultaneous reaction/deactivation problem but a superposition of the chemical events on the deactivation events. Thus the poison waves are always the result of the BohartAdams theory. The next step, of course, is to write a unified model, as was done by Froment and Bischoff. Again we deal with a one-dimensional, plug flow, isothermal reactor, but here treat the case of deactivation by coke formation where either the reactant or the product is the coke precursor. Using the notation of the original work, the conservation equation for reactant A, mole fraction x, is 0x/0"t = -(s

(25)

x = (F/epAf~dp)0, z = w/dp , Z = W/dp where F is the feed rate in mass/time, e reactor void fraction, dp catalyst particle diameter, f~ total reactor cross section, 0 time, W reactor length, PA mass density of reactant A, OBbulk density of the catalyst and rA the rate of reaction of A. The rate of deactivation is proportional to the amount of coke on the catalyst and is slow compared to the rate of the main reaction, and rates of main and coking reactions are first order, in the latter dependent upon either reactant or product precursor, whatever the case may be. Now we still have to specify a relationship between the coke (poison) on the catalyst and the catalyst activity. Several relationships were considered in the original work; our illustration is based on an exponential relationship between activity and coke concentration k_g = (k_A)o(a)

(26)

a = exp(-ot2C,)

(27)

where ct2 is an experimentally determined scaling factor. When the time scale of deactivation is slow compared to the main reaction, initial coke is zero and the reactant is the precursor one can obtain a relatively simple analytical solution x = { 1+exp(-ot2bq_)[exp(az)- 1] }-1

(28)

exp(-cz2Cc) : { 1+exp(-az)[exp(ot2bvl_)- 1]}-1

(29)

with a and b groupings of reactor parameters, ~pBdpP(k_A)0~ and ~'~pAt~dpP(k_A,d)0/F, respectively. When the coke formation is via a predominant reactant precursor, the coke concentration on catalyst generally decreases with reactor length regardless of the relationship between Cc and a; conversely, for product coke precursor coke on catalyst increases with reactor length. In the real world there is very likely a mix of the two mechanisms; for a dose of reality see the work of Dumez and Froment [21 ] on n-butene dehydrogenation. The nature of these coke profiles helps one to understand the wave motion illustrated in Figure 2. Under initial conditions no coke is present and the reaction rate will be highest at the entrance of

84 the bed where the reactant concentration is highest. On the other hand, the rate of coke formation and deactivation will also be highest where the reactant precursor concentration is highest, so the entrance to the bed will also deactivate at the highest rate. Following the progression with time, it is not difficult to understand, then, that the preferential bed entrance deactivation leads to a higher concentration of reactant (precursor) penetrating deeper into the bed and the subsequent development of an activity wave as shown. It is notable that in Figure 2 the final condition shown, with activity highest at bed exit, is exactly the opposite from initial conditions with highest activity at entrance. There is a Table in the text of Butt and Petersen (Table 10-2, Ch. 10) listing ten substantial studies of isothermal fixed bed deactivation up through about 1989 (some of the studies involving multiple references themselves). While I have not attempted an accurate count, there must be on the order of 5-10 publications of various sorts on this topic in each of the past six years, so we by now have a substantial literature to build upon, even if all the problems have not been resolved. Deactivation-induced wave motion can be very prominent in nonisothermal-nonadiabatic reactors, as seen in Figure 3. The simulation of these reactors is a much more difficult problem than isothermal cases, for while often a plug flow model is adequate for mass conservation equations alone if selectivity functions or very short beds are not involved, in the energy conservation equation bed thermal conductivities are often large enough that a dispersion term is needed. This is particularly so in deactivation problems where steep gradients, moving with time as in Figure 3, are encountered. Short of writing another paper there is not much to be gained by going into a detailed discussion of conservation relations and deactivation equations for the nonisothermal case. A typical development is that of Weng et al., and there are many others similar. One prominent feature of the analysis of fixed-bed reactor deactivation is the (unavoidable) large number of parameters involved, at least thirteen or so without including pore diffusion effects and not counting kinetic constants of the reaction and deactivation rate equations. These thirteen parameters, of course, are not all equal in their influence (or, in the Orwellian sense "... all parameters are equal but some are more equal than others ...") and include very sensitive quantities such as wall heat transfer coefficient, bed effective thermal conductivity, and thermal terms related to heats of reaction and activation energies. Aside from trying to cope with the large number of parameters, most workers have struggled somewhat with methods for numerical solution of this type of problem. That also is another story all by itself. Some approaches have included a variable space-step method (Eigenberger and Butt [22]) and collocation (Carey and Finlayson [23], Kam and Hughes [24]). There still seems to be a good distance to go in this area, but with the advent of powerful desktop machines, computer time seems not so precious a commodity as it was not too long ago.

8. CONCLUSIONS It should be apparent by now that, following the "lessons" provided in the Abstract, indeed the three areas of chemistry, reaction and mass/heat transport kinetics, and unsteady-state analysis all play important roles in the understanding of catalyst deactivation. Unfortunately, few of us are blessed with expertise in all of these areas, so there is some compartmentalization in the literature of this field. Fundamental knowledge of poisoning, particularly of metal surfaces, and its relation to surface 'site' geometry has been increased greatly via techniques such as IR and XPS, but it is still a long way, for example, from the operating conditions of an XPS machine to a commercial hydrogenation reactor. Significant strides, however, have been made in recent years in bridging this gap. Similarly, our knowledge of oxide catalysts is much advanced from even a decade ago, but the end of the road is yet not in sight. Advancing as well is our understanding of the many mechanisms of coke formation, but these are so reaction/catalyst-specific that we shall for the foreseeable future still find it necessary to

85 determine individual correlations between coke on catalyst and catalyst activity. This is especially so because in long operating-cycle processes such as hydrotreating the nature of the coke changes with time-on-stream. Solid state transformations like sintering or phase changes, which we have only mentioned in this report, remain probably the least well understood of the mechanisms of deactivation. One would hope that the next decade will see advances here comparable to those made with respect to poisoning in the last. As a first estimate, it would appear that most combinations of deactivation mechanism and intraparticle transport processes have been investigated, at least theoretically. We still would benefit from additional experimental studies in this area, especially in systems where thermal effects are of importance. We also tend to forget that sometimes important physical properties such as porosity and effective diffusivity change significantly with time-on-stream. The basics of such studies, however, are well-established from work dating back even twenty plus years, as seen here, so unfortunately research of this kind is probably not considered very glamorous. In addition, although the topic has not been discussed here, more work on reaction/diffusion/deactivation in biochemical systems, in many instances similar to that done for chemical systems, is much needed and is an important area for future study. Reactor modeling for reactions subject to deactivation would stand at approximately the same stage of development as the intraparticle problem. The large number of parameters and the parametric sensitivity involved in detailed models makes Occam's razor look a little bit rusty. The difficulty doesn't just go away, though, and in this sense we are truly not able to subvert nature. Concern with the development of super-efficient computational algorithms seems to have diminished in view of hardware developments in recent years, although the use of series mixing cell sequences as a reactor model basis has not been exploited as much as it could have been. Of course, the parameterization still does not go away. Again, applications to biochemical systems will be an important area in the future. Overall, I have enjoyed my journey through the mountains, valleys, plains and deserts of catalyst deactivation. The combinations of physical/chemical phenomena are almost unique, the chemistry often complex and somewhat obscure, and the modeling problems formidable. Who could ask for anything more? REFERENCES 1. 2. 3. 4. 5. 6. 7. 8. 9. 10. 11 12. 13. 14. 15. 16. 17.

A.Voorhies, Jr., Ind. Eng. Chem. 37 (1945) 318. G.Bohart and E. Adams, J. Am. Chem. Soc. 42 (1920) 523. G.F.Froment and K.B. Bischoff, Chem. Eng. Sci. 10 (1961) 189; 17 (1962) 105. H.S.Weng, G. Eigenberger and J.B. Butt, Chem. Eng. Sci. 30 (1975) 1341. S.Masamune and J.M. Smith, AIChE J. 12 (1966) 384. J.B.Butt and E.E. Petersen, "Activation, Deactivation and Poisoning of Catalysts", Academic Press, Inc., 1988. F.Shadman-Yazdi and E.E. Petersen, Chem. Eng. Sci. 27 (1972) 227. W.E.Corbett, Jr. and D. Luss, Chem. Eng. Sci. 29 (1974) 1473. E.R.Becker and J. Wei, J. Catal. 46 (1977) 372. S.S.Au, J.S. Dranoffand J.B. Butt, Chem. Eng. Sci. 50 (1995) 3801. J.R.Balder and E.E. Petersen, Chem. Eng. Sci. 23 (1968) 1287. J.P.G.Kehoe and J.B. Butt, AIChE J. 18 (1972) 347. L.L.Hegedus and E.E. Petersen, Catal. Rev. - Sci. Eng. 9 (1974) 245. C.D.Prater, Chem. Eng. Sci. 8 (1958) 284. M.Sagara, S. Masamune and J.M. Smith, AIChE J. 13 (1967) 1226. J.W.Lee, J.G. Butt and D.M. Downing, AIChE J. 24 (1978) 212. J.J.Carberry, "Chemical and Catalytic Reaction Engineering", McGraw-Hill, Inc., 1976.

86 18 19. 20. 21 22. 23 24.

A.Wheeler and A.J. Robell, J. Catal. 13 (1969) 299. A.Wheeler, Adv. Catal. 3 (1950) 307. H.W.Haynes, Jr., Chem. Eng. Sci. 25 (1970) 1615. F.J.Dumez and G.F. Froment, Ing. Eng. Chem. Process Des. Dev. 15 (1976) 291. G.Eigenberger and J.B. Butt, Chem. Eng. Sci. 31 (1976) 681. G.F.Carey and B.A. Finlayson, Chem. Eng. Sci. 30 (1975) 587. E.K.T.Kam and R. Hughes, Chem. Eng. J. 18 (1979) 93.

~ Elsevier Science B.V. All rights reserved Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

87

Catalyst Deactivation: Opportunity amidst Woe Eugene E. Petersen Department of Chemical Engineering University of California, Berkeley, CA 94720

I would like to thank the organizing committee of 7 th ISCD for awarding John Butt and me the Maxsted Award for our joint authorship of the book, Activation, Deactivation and Poisoning of Catalysts. The collaboration between John and me was both educational and satisfying for me and it is noteworthy to mention that our personal friendship has survived the collaboration. My goal xn this talk is certainly not to summarize the topics in the book as that would be neither possible nor appropriate. Rather, as the title suggests, I would like to look at one aspect of the subject the importance of which I have a rather well-developed bias. This has to do with the utilization of reaction and deactivation data together to determine reaction rate and deactivation rate expressions as well as to design reactors. First let's review the situation facing anyone wanting to use the deactivation technology. For most chemical reactions, we will have at best only incomplete rate information available to be interpreted in terms of a theoretical framework that is overly flexible and undersensitive diagnostically. Superimposed upon this we usually have insufficient deactivation data to be interpreted in terms of a theoretical framework having the same shortcomings as the kinetic rates of the main chemical reactions. We face a situation somewhat similar to that depicted in Samuel Hoffenstein"s quip: Little by little we subtract Faith and fallacy from fact, The illusory from the true, And starve upon the residue. Perhaps this characterization is a bit pessimistic, but it emphasizes how important it is for the reactor designer to know the limitations of the theoretical framework in which his data must be interpreted. The main theme of this talk explores an example of the use of deactivation information to learn more about the kinetics of the main reaction. This idea is not new. Noteworthy is the pioneering paper of Herrington and Rideal [ 1] who demonstrated theoretically that a main reaction requiting more than one surface site loses activity faster than proportional to the number of unpoisoned sites, i.e., poisoning one quarter of the surface sites reduces the number of multiple site ensembles by more than twenty-five per cent. The number of six-site ensembles falls off very rapidly with the extent of poisoning. Their paper is all the more remarkable because the numerical work was done before the availability of digital computers. Accordingly, the relationship they found was based upon a relatively small section of a (111) plane surface. Never-the-less, their results are qualitatively correct.

88 Catalyst developers experience woe in their quest to find a practical catalyst owing to the very general tendency of catalysts to deactivate during use onstream. To live with this phenomenon, a developer must know the time scale of the deactivation process in order to know what kind of reactor to use. If development schedules permit, the developer may even have the luxury of making a detailed study of the deactivation process in the laboratory or in a pilot plant to find conditions to increase the life of the catalyst on stream. I have always been impressed by the success that an intelligent engineer can achieve with limited data when it is supplemented by an undefinable intuition and insight and a few model calculations. This works largely because the magnitude of the reactor design integral is not terribly sensitive to the exact shape of the rate function so long as it has approximately the right general shape. As a consequence one can often use a lumped parameter to approximate a phenomenon in place of the detailed parameter behavior. For example, in some cases it is possible to replace an array of several similar parallel chemical reactions by a single reaction of averaged properties. The trick is to choose the parameters of the single reaction to behave close enough to the real system to be useful. The chosen parameters would then be designated as lumped parameters. Of course a word of caution in using this approach is that one must constantly check the results against experimental data because lumps are not fundamental quantities. A rather nice story to serve as a case in point is the work of Weekman and his co-workers [2] at Mobil to describe the behavior of fixed, moving and fluidized bed reactors to catalytically crack feed streams under deactivating conditions. Both the main reaction and the deactivation reaction are exceextingly complex and to describe them precisely is well beyond our knowledge to date. The woe is that these processes are industrially very important and must be used to produce gasoline. So it is prudent and necessary to try to describe them as well as possible in the absence of complete information. The approach of Weekman and his co-workers was to assume that the array of chemical compounds in the gas-oil feed fraction can be lumped as a single chemical identity which under cracking conditions forms chemical products; gasoline (C5)-----410~ F and other fractions consisting of Ca's, dry gas and coke. The deactivation process must also be handled as a lumped process. Most descriptions are of the Voorhies type that relates activity as inversely related to the coke content. Care must be exercised in making this assumption because very definitive experiments have demonstrated that coke per se is not the deactivating substance but rather specific compounds, such as organic bases, cause deactivation. Useful lumping results when the specific deactivating compounds are produced approximately in proportion to the coke. Now we make the assumption that the rate of chemical reaction under deactivating conditions can be expressed as the product of two quantities: a rate expression for an unpoisoned catalyst and a proportionality constant which is defined as the activity. This is the so called separable form proposed by Szepe and Levenspiel [3]. Accordingly, if we say that the deactivation rate is proportional to the reaction rate, or more specifically, the rate of change in activity is proportional to the activity, then we arrive at an exponential decay in activity with time of the form activity = exp(-kdt)

,

(1)

where kd is the deactivation rate constant which must be evaluated for feedstocks, reactor conditions, catalyst used and a whole host of variables specific to a particular cracking operation. This is the Mobil approach to commercial deactivating reactors and it apparently works well when attention is focused upon the lumping parameters and their meaning. I would like, now, to review in a limited way some work of some of my former students and

89 myself-- Larry Jossens, Bruce Isaacs and in particular Mike Pacheco--- to try to understand the role that rhenium plays on the platinum catalyst used to reform naphtha streams. I believe that the study of the deactivation process offers some opportunities to understand both the main reaction and the deactivation reactions better.

4r•

r. d.s. ~

Fouled Sextet

lie

~

.

.

.

.

.

.

.

Main Reoction

"It" Figure 1. Six-site fouling mechanism as a parallel pathway to the main reaction mechanism (* represents a metal site) Ref [4]. Reproduced by permission of Academic Press, Inc.

The deactivation process on this catalyst consists of progressive dehydrogenation, isomerization and polymerization on the metal catalytic sites [4] with paraffinic and aromatic molecules both contributing to the loss of catalytic activity. We start with a mechanism involving a paraff'mic molecule sorbed on the surface losing hydrogen atoms in a series of reversible dissociative steps. The multiply-bound molecule then reacts in a Rideal-Ely type mechanism that initiates the site-consuming polymerization process. Specifically, the methylcyclohexane loses all 11 of the hydrogens in the ring and the dehydrogenated surface species reacts with toluene by the process depicted on Figure 1. The main reaction is similar to other proposed mechanisms [5]. If we assume that the reaction between the surface bound molecule and toluene is rate determining step, we arrive at the expression shown in Equation 2 for the deactivation kinetics.

dSt _6ktK6S6 [MCHI[TOL] dt = [H2]~t

(2)

where St is the number of unfouled sites and So is the number of vacant sites. To use this equation to get a general relationship for how the activity changes with time requires that some relationship between St and Sv be established. In kinetic studies of the main reaction, product inhibition was proportional to toluene concentration to the - 8 9power. This leads to

St - K5Sv[TOL]~ .

(3)

90 Substitution of Equation 3 into Equation 2 reveals that the time derivative of St is proportaional to St6. Integration for constant concentrations and temperature and a boundary condition of St=Sto at t=0 yields

st = (1 + kt)Sto

(4)

where

k = 30kIS5~ K~[TOLI[H2I~

(5)

Since the main reaction is assumed to require one site, the left side of Equation 4 is equal to the activity and 1

Activity = a = (1 + kt)-~

(6)

At this point it should be emphasized that one cannot legitimately carry out the above integration unless the concentrations of all species are held constant throughout the deactivation process. This can be accomplished easily for a reaction involving one independent stoiehiometric equation by varying the flow rate to the reactor to keep one concentration constant [6]. In the case of more than one independent stoichiometric equation, maintaining constant concentrations is a difficult control problem. We can now compare Equations 5 and 6 with the data on the dehydrogenation of methylcyclohexane to toluene over a P t - A1203 catalyst shown in Figure 2 where the abscissa has been scaled with the corresponding k values for each experiment. The slope of the line becomes approximately proportional to the -0.2 power after an induction period characteristic of Equation 6, but at longer times on stream the data systematically deviates from the model line with an increasing slope. In keeping with the interpretation that led to Equation 4, we infer from this that the order of the deactivation process decreases with time on stream. This is consistent physically with the deactivation process using fewer sites as time passes. To explore this kind of behavior more quantitatively, we extend the model to allow deactivation using two, three, four and five sites mechanisms. This model takes the mathematical form

dSt

6

=- k so

(sv)Cm-D

(7)

m"-2

If we now define the following dimensionless quantities:

Sv _ number of vacant sites St _ number of active sites - St - number of'active sites ' f -- S t o - number of active sites initially T - - 5 k 6 v 6 t , k m = Amexp ( ~Tm ) Substituting these quantities into Equation 7, we get

d'-'t = - 5 m=2

(A.) ~

~,rn-6fmexp

RT

(8)

91 ~

L0

- ..... I

........

I

........

I

........

I

........

9

9

O

9

t>

O

a

,~q:>

| A

o..1

Model

(C) <J

o

o 01,"1 I J

C, 0 a

j J I

B v |

MCH Tol.

352

352

352 352 351

12

,105

1.5

1'

413

2.3

32

19 12 2

1' 1'

405 405 220

18

T

405

18 18

I ,o ~

,~'

11 34

4OO

o o.1 . . . . . . . . .

(mio ] )

18

34

377 352

H,

........

I ,o'

403 405

-a.-~~ "

2.5 4.6

O~

I ,o ~

J

:8

I ..J

~ 9

........

I ,d

I

-

9 #~

30. 33. 45. 150.

........

--

........

J j J

o~v 9 9 ~ I ,0 4

j j J

........

, ,o s

Fouling Time, k t Figure 2. Fouling data for dehydrogenation of methylcyclohexane to toluene over a platinum-alumina catalyst Ref [4]. Reproduced by permission of Academic Press, Inc.

I0.~ =i,=-

._>

X = _>B kcal/mol

9

I0 - I -

~ >,

o 0

1 ~,~

_

:

I0 :z 1

-

3

basis: 4 0 0 " C

: :_

0 I 0 -I

--

Empirical Correlation - - Multiplel Model -

3

~

I J___ 10 3

I01

"~ 10 5

.~ 107

Fouling Time, kft Figure 3. Comparison of multiplet model with empirical fouling correlation for values of A. Ref [6]. Reproduced by permission of Academic Press, Inc.

Finally, we assume that .for all values of m (9)

( E r a - E6) = A(6 - m)

(i0)

Equation 9 assumes that the product of the pre-exponential factor and the activity for all multi-site reactions are equal. Differences in the rates of the multi-site reactions obtain from

92 the values of their respective activation energies. The assumption of Equation 9 minimizes the number of adjustable parameters. Equation 10 is equivalent to saying that each bond to the catalyst surface lowers the activation energy by an amount )~. Figure 3 shows several solutions to Equation 8 subject to the assumptions of Equations 9 and 10 for various values of ,k. A value somewhere between 2 and 3 kcal/mol has the approximate shape of the empirical correlation. Although this value of ~ is smaller than expected for the interaction energy per bond to the metallic surface, it is influenced by the assumption of Equation 9. A reasonable conclusion from this figure is that the multiplet model which is physically sensible also has the fight quantitative order of magnitude and effect. Of course one could play with Equations 9 and 10 to modify the assumptions to get a better fit, but in my opinion, as it stands, it lends support and utility to the idea that the deactivation process involves ensembles of sites containing progressively fewer sites as the deactivation process continues. When the numbers of six-site ensembles is large, the deactivation process proceeds rapidly, but as a rather small amount of surface sites are covered, the population of six-site ensembles decreases disproportionately much faster and further deactivation must proceed by parallel reactions involving ensembles containing fewer than six-sites. It is well known in refinery practice that sulfiding a P t - Al203 catalyst results in a more stable although somewhat less active catalyst. If it is assumed that the sulfiding simply deactivates single platinum atoms on the surface, the multiplet theory can easily account for the behavior. I

IO 0

"'

a o = O.G '

o....

= - 9, - -

I

..,,..

Im i

im i.

i

i

l

i . . . . . .

I

,,...., ~

~ " ' w

,,ram

--

- "='1-- ~,..~,/Empiricol

Correlation

IO-I

. ..,.,.,

o o --0.1

U

10-2 (kf)ptoS = 2 0 x (kf )Pt + Re S

0 0

10 -3

....

I I 0 -I

l

[,.. I01

!

1. 10 3

%, %" %"

.1

I I0 5

l

.,--

10 7

Dimensionless Fouling Time ( k f . t ) Figure 4. Typical fouling data for sulfided reforming catalysts. Shaded symbolsmPt, open symbols---Pt-Re Ref [7]. Reproduced by permission of Academic Press, Inc.

Sulfiding, therefore, reduces the activity of the catalyst for the main reaction and at the same time stabilizes the catalyst by reducing the number of six-site ensembles and, most importantly, it reduces the initial rate of deactivation because now it can occur only by a mechanism involving fewer active sites which is a slower process. In so doing note that by reducing the number of six-site ensembles before any deactivation takes place on stream results in a decrease in the deactivation rate at any given activity level as compared with the unsulfided catalyst. For a main reaction involving only one surface site, the activity itself is a measure of the number of unpoisoned active active centers on the surface. But the distribution of deactivated sites in a catalyst that has been sulfided is different than a catalyst deactivated by a six-site mechanism--the former being essentially a deactivation involving fewer than six

93

10-3

I

I

I

? ..=. 0

E

c 0

L

== v

10 . 4

U "10 0

n

E e) 0

10. 5

I-

Y I n t i m a t e Mixture A Physical Mixture 9Bimetallic

t:}

n,, 10 . 6 I 0 -I

I I00

! I01

! 102

I

03

Reaction Time (min) Figure S. Deactivation profiles for Pt and Pt-Re catalysts. Ref [7]. Reproduced by permission of Academic Press, Inc.

site.. As a consequence, at any comparable activity level, the number of remaining six-site ensembles is always greater for the unsulfided catalyst. Figure 4 shows how sulfiding affects the kinetic behavior of Pt-A1203 and P t - R e A l 2 0 3 catalysts when compared with a fresh unsulfided P t - A1203 catalyst. The activity of the P t - Re catalyst is reduced much more than the Pt catalyst with the same H2S treatment. As stated above, both catalysts are stabilized and their corresponding decay constants are reduced. "Significantly, the sulfided P t - Re catalyst's decay constant is 20 times smaller than the sulfided Pt catalyst. Integration of Equation 7 assuming that the initial activities are 0.6 and 0.1, respectively, for the Pt and P t - Re catalysts, yields the solid lines on Fig.ure 4. This figure may be misleading because time on the abscissa is dimensionless. A pomon of these data on a real time axis is shown on Figure 5. On this figure we compare the activitytime characteristics of the sulfided Pt and P t - Re catalysts. These curves clearly show the enormous stability of the P t - Re catalyst that is not so apparent in the dimensionless form. These data show that physical mixtures of R e - Al203 and P t - Al203 do not give the same deactivation behavior as a P t - ReAl203 catalyst. In fact the physical mixture is indistinguishable from the P t - A1203 catalyst alone. This point is made to demonstrate that our data do not support a proposed theory that Re in the bimetallic catalyst acts as a scavenger for deactivation precursers. These stu~es on the role of rhenium in the bimetallic reforming catalyst can be summarized on Figure 6. The PtAl203 catalyst appears to deactivate during its initial period by a mechanism involving six surface sites. As the deactivation progresses, the apparent order decreases to a value somewhere between two and three as shown on the figure. Sulfiding the PtAl203 catalyst results in an initial loss of activity but stabilizes the catalyst by lowering the magnitude of its decay constant by greatly reducing deactivation by a six-site mechanism. A similar but more pronounced effect is observed with the P t - ReAl203 catalyst. To me this, perhaps over-simplistic, interpretation gives us a way of thinking about how to improve the stability of catalysts by trying to understand the mechanisms by which they deactivate and try to interfere

94 with these processes in order to slow them down. I believe these kinds of approaches are often overlooked by many who are preoccupied by the woes of catalyst deactivation-- maybe a situation related to the yin and the yang of traditional Chinese cosmology. It is always satisfying when a theory that evolves from the interpretation of data explains an observed phenomenon that was not used in its development. Such is the case of using the multiplet theory to explain shape of the characteristic temperature-time curves for constant conversion in an industrial reactor. That is, to compensate for the inevitable decrease in conversion that would result as the catalyst in a reactor deactivates onstream, and operator raises the inlet temperature to maintain constant conversion. These curves are of three types: A curve whose slope that monotonically increases with time, a curve whose slope monotonically decreases with time and a curve whose slope decreases with time and then increases after passing through a inflection point. The last shape is very common. An elementary analysis of the problem that follows reveals that it is not easy to explain the inflection point in the curve. Jo ~

,.

~

J

I

I

\ % = 1.0

~,9

-

u

-I

v ~ ~ ~

---

!

~

AlexA

fresh Pt/AI 20

-----

io-Z

I

_

~

I

I

I

~\d:?_

~

\

Single-order Model ' Mult=ple-order Model

io-~

!

I I I Io z io 3 io 4 Dimensionless Fouling Time, k d , t

io o

I

io ~

'il ~'= I io 5

I

io 6

I

Figure 6. Dimensionless fouling curves for Pt and Pt-Re catalysts for the dehydrogenation of methylcyclohexane. Ref [7]. Reproduced by permission of Academic Press, Inc.

A well-stirred reactor illustrates the general behavior of deactivating systems. The usual design equation for such a reactor shown in Equation 10, = a.Vk.(To)C

where: v = volumetric flow rate C o - inlet concentration of reactant C e - outlet concentration of reactant a , = separable form of activity for main reaction V - reactor volume kn(T) - rate constant for main reaction at T q - o r d e r of the main reaction with respect to the limiting reactant.

(II)

95 Equation 11 can be rearranged to the form

k,.,(T) ankn(To)

Xr - -

(1 - x e ) q D a

(12)

where: x.

-

(Co - c D I C o

D a - Vk,~(To)Cqo-1/v, a DamkOhler number. The only variable on right-hand side of Equation 12 is the conversion, Xe, which is to remain constant by raising the inlet temperature of the feed mixture as a function of time. Therefore, the right-hand side of Equation 11 is a constant. But initially the left-hand side of Equation 11 is equal to one because the initial activity, an, is by definition unity. It follows, then, that the left-hand side of Equation 11 is not only constant but it is equal to one and remains equal to one throughout the deactivation process. If we replace the activity of the catalyst, a,~, by og'~, where og is the fraction of the sites that are active (not deactivated) and n is number of sites used by the main reaction and substitute the Arrhenius form for the rate constant in terms of the activation energy into the left-hand side of Equation 12, we get

oLnexp Y = 1

(13)

where:

Y - ,y.,.,(e/(l+ e ) ) e - (T-

To)IT,,)

% - E,URTo

En - activation energy for the main reaction. If we treat the deactivation reaction in a manner similar to that used previously, we get dog d"t

=

ogmkm(T)C~'

(14)

where a m is the activity for the deactivation reaction assuming that m sites are involved. Substituting the Arrhenius form for km in Equation 13 and simplifying we get

, ~da

= og"%xp(eY)

(15)

where: e - E m / E n and r - nkm(To)CPo(1 - Xe)Pt Using Equation 13 to eliminate a in Equation 15 yields 9

dY = exp(6Y) dr

(16)

96 where:

')

(17)

What we seek from Equation 16 is shape of the O versus T curve which would reflect the inlet temperature vs time curve. An efficient way of doing this is to look at the sign of the second derivative, d20/dr 2. Differentiating Equation 16 yields

1+0 The only term in Equation 17 that can be negative is enclosed in brackets which we will define here for convenience as A. In terms of real quantities

A = 7n

(E.~

En

m-

n

1

2T ) + 7nTo

(19)

There are three possible cases: A < 0: The slope of the T vs t curve increases monotonically (positive curvature). A = 0: The slope of the T vs t curve is constant. A > 0: The slope of the t vs t curve decreases monotonically (negative curvature). l.x~king at the quantifies in A, we observe they are all constants except m, n, and T. If we assume that the mechanisms of the main reaction and the deactivation reaction remain constant, then T is the only quantity that varies. The term 2T/"/nTo has normal values of about 0.1, i.e. T/To changes from 1 to about 1.2 during the entire period of deactivation and 7n is of the order of 20. If E m / E n ~ 2 - 4, m ~ 0 - - 5 and n ~ 1 - 2; then temperature variation in the 2T/%.,To is most unlikely to cause a change in the sign of the second derivative. Hence the often observed inflection in the T vs t curve is difficult to explain unless we allow the value of m to change with time. But in the early part of this paper we observed apparent changes in value of m with progressive deactivation. Accordingly, if we substitute the multiplet deactivation model into the theory, clearly there is opportunity for an inflection point because m - 1/n starts out at about 5 and decreases to about 2 and at the same time Em increases. Figure 8 shows the shape of Y vs r for various values of the ratio of the activation energy for deactivation to the activation energy of the main reaction. The model has the inherent characteristics to give the inflection point. It is not surprising that similar modeling of a plug flow reactor gives a similar result. Although the details are well beyond what I want to present here, suffice it to show that the results of the T vs t of a plug flow reactor are shown and compared to a typical T vs t for a commercial hydrodesulfurization reactor in Figure 9. Superimposed are the corresponding single-order deactivation models.

97 ,5.0

v

I

v

I

I

'1',

I

"

v i Vl , i v | i i v v- i i r i Bosi.....~s: T o 6 2 3 K ( 3 5 0 " C , 6 6 2 " F )

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_

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_

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|

| +

v 1 v i

/

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I

/

/

i X T y p i c o I HDS d a l o I . E . C . ( 1 9 8 1 ) Torero. Hornsberger ond B r i d ~ s ~ x

h~ O . I C

/

iii

iI

_

| >c 2.0

3

E

>-

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4

1.0

0

.g_ 0

I

1

1

~

0.2

0.4

i

l

O

50

0.6

I00

150

Dimensionless

200 Time,

250

5OO

k d ( T O ) 9!

T

Figure 8. Characteristicplot for a multiplet model in a CSTR. Ref [9]. Reproducedby permissionof Academic Press, Inc.

Figure 9. Typical T versus t curve for various kinetic models and a commercialI-IDSreactor. Ref [9]. Reproduced by permissionof AcademicPress, Inc.

Few people involved with catalysis find deactivation process anything but undesirable, more likely a nuisance, except for a few of us who play with the phenomenon as a part of.their livelihood. The main message in this paper is that those stuck with a catalystthat deactivates under operating conditions might profitably spend a little time trying to understand something about the mechanism of process. The rewards can be anything between marginally better design information to a whole new way of thinking about the catalytic process that leads to better catalyst design. I am sure that many of you are aware that the bimetallic reforming catalyst evolved as the result of an experiment using rhenium on the catalyst as a possible means to suppress sintering. The enhanced stability of the sufided form to deactivation by coke was a serendipitous result. These fortuitous events have led the way to a whole series of bimetallic and even trimetallic catalysts with increasingly superior deactivation properties. The work discussed herein is an a p o s t e r i o r i explanation of the enhanced stability for the P t - ~ catalyst based upon the idea that a well distributed sulfided Re surface sites markedly reduce the number of six-site ensembles---the ensembles that contribute to rapid initial deactivation. The result is in itself important for an understanding of the deactivation of this important industrial catalyst. But more important, it may serve as a model to understand the deactivation of of other supported metal catalysts to carry out other hydrocarbon reactions where polymerization type processes cause deactivation.

NOMENCLATURE a -- activity of catalyst for the reaction of MCH to TOL. an -- separable activity for main reaction. Am = preexponenti.al factor for the m th reaction. Co - concentration of reactoant entering reactor. Ce - concentration of reactoant leaving reactor. D a - Damkohler Number, see Equation 12. Em = activation energy for the m th main reaction. En =- activation energy for the n th deactivation reaction. f _= fraction of surface sites that are active, see Equation 7. [H2] = hydrogen concentration k defined by Equation 5. k,~ - deactivation rate constant, see Equation 1.

98

km =-- deactivation rate constant for m th reaction kt -- deactivation rate constant MCH to TOL reaction K5 ----- equilibrium constant for inhibition MCH to TOL reation. K 6 ------equilibrium constant for dehydrogenation of MCH on surface. [MCH] - methylcyclohexane concentration. m is subscript pertaining to a deactivation reaction. n is subscript pertaining to a main reaction. St - number of active (unpoisoned) sites on surface at time=t. Sto =- number of active (unpoisoned) sites on surface at time=0. Sv -= number of vacant sites on surface at time=t. t - time on stream for deactivation. [TOL] - toluene concentration. v - volumetric flow rate to the reactor. X e - (Co - O e ) / G o , the conversion in the reactor. Y - 7 n O / ( 1 - O), see Equation 13. a n - activity for an n th order surface main reaction. a m -- activity for a n n t h order surface deactivation reaction. c

-

E~/E,,

v 7A ~5 A

-- defined by Equation 7. - defined by Equation 7. -- bond energy to surface. -- see Equation 16. -- see Equation 19. 0 = ( T - T o ) / T o , a dimensionless temperature. 7 - E/l~o, an Arrhenius Number.

REFERENCES 1. E . E G : Herrington and E.K. Rideal, Trans. Faraday Soc.40 (1944) 505. 2. J.B. Butt and E.E. Petersen, Activation, Deactivation and Poisoning of Catalysts, Academic Press, 1988. 3. S. Sz6pe and O. Levenspiel, Chem. React. Eng., Proc. Eur. Symp.. 4th, 1968. 4. M. A Pacheco and E. E. Petersen, J. Catalysis 86 (1984)75. 5. B.C. Gates, J.R. Katzer and G.C.A. Schuit, Chemistry of Catalytic Processes, McGraw-Hill, 1979. 6. E.E. Petersen and M.A. Pacheco, ACS Syrup. Series 237 (1984) 363. 7. M. A Pacheco and E. E. Petersen, J. Catalysis 88 (1984) 400. 8. M. A Pacheco and E. E. Petersen, J. Catalysis 96 (1986) 499. 9. M. A Pacheco and E. E. Petersen, J. Catalysis 98 (1986) 380.

9

Elsevier Science B. V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

99

The R e l a t i o n s h i p B e t w e e n Metal Particle M o r p h o l o g y and the Structural Characteristics of C a r b o n Deposits R. T. K. Baker a, M. S. Kim b, A. Chambers a, C. Park a and N. M. Rodriguez a aDept, of Chemistry, Northeastern University, Boston, MA 02115. bDept, of Chemical Engineering, Myong Ji University, Kyonggi-do, Korea 449-728.

The understanding of the factors controlling the deposition of carbonaceous solids resulting from the decomposition of hydrocarbons over metal particles has a considerable impact on a number of commercial processes including: catalytic steam reforming of methane, catalytic reforming and systems involving carbon monoxide disproportionation reactions. The highest catalytic activity for carbon deposition is exhibited by iron, cobalt and nickel, and alloys containing these metals. During the past few years we have pursued a program designed to achieve conditions where it is possible to control the catalytic properties of a given metal by inducing perturbations to the reactive surfaces of the crystallites. One of the consequences of such an action is to enable one to alter the catalytic reactivity in such a fashion so as to optimize the performance for a desired reaction pathway, while simultaneously suppressing the rate of detrimental side reactions, such as encapsulating forms of carbon deposition. Our strategy has centered around a study of the effect of introducing selected adatoms into the host metal and using the decomposition of ethylene to probe the manner by which the chemistry of the various faces of the crystallites is modified. In this paper a discussion of the information obtained from the use of a combination of controlled atmosphere and high resolution transmission electron microscopy techniques to study the impact of metal particle morphology on the characteristics of the carbon deposit will be given.

1. I N T R O D U C T I O N It should be recognized that the potential for carbon deposition exists in many systems in which hydrocarbons or carbon monoxide undergo decomposition over heated metal surfaces. Several reviews have highlighted the complex nature of carbonaceous deposits (1-6), which can be divided into three main types: amorphous, filamentous and graphitic shell-like structures. During a traditional routine analysis of a spent catalyst, these three forms of carbon would not be necessarily distinguished, but merely referred to collectively as "coke". Available evidence indicates that the amorphous carbon component is formed via condensation and polymerization reactions and this material originates from thermal processes. It is conceivable that a significant amount of hydrogen is retained in the deposit, however, as the temperature is raised dehydrogenation reactions will tend to reduce the hydrogen content. There is now a general consensus that the formation of the filamentous and graphitic forms of carbon require the participation of a catalytic entity. During the reaction, the metal particles adopt well defined geometries with the carbon-containing gas molecules being adsorbed and decomposed on certain faces of the metal, this process being followed by diffusion of carbon atoms through the catalyst particle

100 to precipitate at another set of faces in the form of a fibrous structure. It is well established that carbon diffusion through the particle is the rate determining step in the growth process. Depending on the chemical nature of the catalyst and the conditions at which the reaction is performed, assorted filamentous structures with various morphologies and different degrees of crystallinity can be produced (7,8). While the rudiments of the mechanism for the formation of the graphitic shell-like deposit have not received the same attention as that devoted to the generation of carbon f'llaments it is probable that many of the steps outlined above are also operative in the growth of this form of carbon (9,10). The major difference between the characteristics of the two catalytically formed types of deposit may lie in the number of the faces that perform the dissociative chemisorption of hydrocarbons compared to those that only allow for the carbon precipitation mode. The modification of the catalytic behavior of a given metal by the addition of controlled amounts of other metals has been used extensively in order to alter both the activity and selectivity as well as to control the deactivation of active surfaces with respect to poisoning and deactivation due to carbon deposition (11-15). It is the understanding of this latter aspect that has eluded researchers in this area and presents challenges to the various mechanisms that have been proposed to account for carbon deposition on metal surfaces. McCarty and coworkers (16,17) used temperature programmed reaction techniques to establish the existence of several forms of carbon on deactivated nickel catalysts. From these studies graphitic oveflayers and filamentous carbon growths were identified as being the most prevalent types associated with metals. The graphitic form of deposit tends to encapsulate the particle surface in contact with the reactant gas and thus normally results in rapid deactivation of the catalyst. In contrast, filamentous carbon is produced by a process in which carbon diffuses through the particle and precipitates at the rear faces, thereby leaving the exposed faces free to undergo continued reaction. The net result of this behavior is that the catalyst system can accumulate large amounts of carbon and maintain activity for prolonged periods of time (18). It is now being recognized that the incorporation of controlled amounts of certain non-metallic atoms into metals, which have hitherto been regarded as poisons, can actually function as catalyst promoters. This approach is predicated on the assumption that the poison atoms are preferentially chemisorbed on sites that are active for undesirable reactions, while those sites which perform the desired reactions are preserved in an unadulterated state (19,20). Previous work from this laboratory (21-23) has demonstrated that either pretreatment or continuous addition of 5 to 10 ppm H2S to iron, nickel or cobalt undergoing reaction with ethylene had a significant impact on the catalytic activity of these metals, particularly with respect to the enhancement in yields of carbon filaments. We have extended this investigation to cover the influence of chlorine on the same reactions (24). Careful poisoning with chlorine has been found to be a very effective method of altering the activity and selectivity of catalysts and it has been suggested that such behavior could in part be attributable to induced electronic perturbations in the metal (25). In catalytic reforming, the chlorine content derived from the metal precursor salt is thought to be directly related to the amount of carbon that accumulates on the catalyst (26). 2. E X P E R I M E N T A L 2.1. Materials Bimetallic powders consisting of combinations of iron, nickel, cobalt and were prepared by coprecipitation of the respective metal carbonates from metal solutions, Fe(NO3)3 9H20 (98%, Fisher Scientific), Co(NO3)2.6H20 (99% Scientific), Ni(NOa)E.6H20 (99% Fisher Scientific) and Cu(NO3)2 2.5H20 (99.2%,

copper nitrate Fisher Fisher

101 Scientific) using ammonium bicarbonate at room temperature and a pH of about 9.0, followed by oxidation to metal oxides and finally, reduction in 10% hydrogen/helium at 500 ~ for 20 h, as described previously (27). The reduced catalyst was cooled to room temperature under helium, then passivated by a 2% air/helium mixture for 1.0 h and stored in sealed containers. X-ray diffraction analysis was performed on all of the bimetallic powders at room temperature and only peaks corresponding to metallic iron, nickel, cobalt and copper were evident in these samples. It should be recognized, however, that upon heating to 600~ the phase diagrams for these mixed metal systems shows that in all cases alloy formation occurs. Pure iron, nickel, cobalt and copper powders were also prepared according to the above procedure. The gases used in this study, carbon monoxide (99.99%), hydrogen (99.999%), helium (99.99%) and ethylene (99.5%) were obtained from MG Industries and were used without further purification. Approximately 50 mg of each sample of all these powders were initially given a reduction treatment in 10% H2/He at 500~ and then reacted in a horizontal flow reactor system in the presence of either ethylene/hydrogen (4:1) or carbon monoxide/hydrogen (4:1) mixtures at temperatures over the range 450 to 800~ for a period of 2.0 h. Gas flow rates were maintained at 100 cc/min and regulated with MKS mass flow controllers. At the conclusion of each experiment the solid carbon product was carefully removed from the quartz reactor, weighed and then stored for subsequent structural analysis. In a final series of experiments, iron-nickel powder was reacted in the presence of ethylene/hydrogen (4:1) at temperatures in excess of 800~ in order to investigate the formation of shell-like deposits that are known to form under these conditions.

2.2 Techniques The details of the structural characteristics of individual constituents in the various carbon deposits were obtained by examination of a number of specimens from each experiment in a JEOL 100 CX transmission electron microscope that was fitted with a high resolution pole piece, capable of 0.18 nm lattice resolution. Suitable transmission specimens were prepared by applying a drop of an ultrasonic dispersion of the deposit in iso-butanol to a carbon support film. In many cases the solid carbon product was found to consist entirely of filamentous structures. Variations in the width of the filaments as a function of both catalyst composition and growth conditions were determined from the measurements of over 300 such structures in each specimen. In certain samples evidence was found for the existence of another type of carbonaceous solid, a shell-like deposit in which metal particles appeared to be encapsulated by graphitic platelet structures. Selected area electron diffraction studies were performed to ascertain the overall crystalline order of the carbon filaments and the shell-like materials produced from the various catalyst systems. The controlled atmosphere electron microscopy (CAEM) experiments were carried out in a modified JEOL 2000EXII TEM instrument (28). This instrument is equipped with a custom designed environmental cell, which accommodates a heating stage. With this arrangement it is possible to continuously observe changes in the appearance of a specimen as it undergoes reaction in a gas environment at temperatures up to 1000~ The dynamic events occurring during the reaction are captured on video-tape and replayed later for detailed analysis. We estimate that the point-to-point resolution achieved on the TV monitor is of the order of 0.4 nm under sufficiently stable conditions. A recent development with the technique is the capability of performing in-situ electron diffraction analysis of supported small particles to establish the chemical state of the reacting specimen (29). This technique was used to determine the morphological characteristics of both the pure metal and bimetallic particles supported on single crystal graphite surfaces. The samples used in these experiments were prepared by two methods: (a) single metal/graphite specimens were made by evaporation of the metal in the form of a wire from a tungsten

102 filament at 10 -6 Torr onto transmission sections of single crystal graphite; (b) graphite supported bimetallic specimens were produced by introducing the components mixed in the desired ratio onto the substrate medium, as an atomized spray from an aqueous solution of the respective metal nitrates. Prior to reaction in the hydrocarbon environment both these types of systems were pretreated in 1 Torr hydrogen at 400~ for 3 hours.

3. RESULTS AND DISCUSSION 3.1.

Dynamic Behavior of Graphite Supported Metal Particles

When samples of either the pure metals or bimetallics dispersed on single crystal graphite were examined in the CAEM following the initial treatment in hydrogen it was evident that the evaporated films had undergone nucleation to form discrete particles that were of the order of 2 nm in diameter and tended to adopt a spherical shape. When such specimens were subsequently heated in either 2.0 Torr ethylene/hydrogen or carbon monoxide/hydrogen mixtures some dramatic changes in both size and morphology were observed. For example, in the presence of an ethylene/hydrogen (4:1) mixture at temperatures approaching 500~ particles of iron-nickel started to undergo reconstruction from a globular to a square form. This aspect is shown in the schematic diagram, Figure 1a, which taken under reaction conditions. In experiments where the temperature was raised to temperatures in excess of 750~ these particles were observed to gradually transform and adopt a hexagonal morphology. Inspection of copper-iron particles when heated in a CO/H2 (4:1) reactant revealed that in this case, the particles acquired a rectangular morphology as the specimen temperature was increased to 475~ Figure lb. In a further example, coppercobalt particles exhibited a somewhat fascinating transformation when the system was heated to 525~ in an ethylene/hydrogen (4:1) mixture, the initial spherical-shaped particles transforming into well-defined fiat hexagonal-shaped crystallites, as seen in Figure l c. It should be mentioned that even on continued heating up to 650~ no carbonaceous deposits were formed at the very low pressures at which the CAEM experiments were performed with these particular gas mixtures.

(a)

(b)

(c)

Figure 1. Schematic representation of the observed morphological characteristics of various bimetallic particles during interaction with selected gas environments. (a) an Fe-Ni (5:5) particle undergoing reaction in C2Hn/H2 (4:1) at 500~ (b) a Cu-Fe (2:8) particle heated to 475~ in the presence of CO/H2 (4:1), and (c) a Cu-Co (1:3) particle treated in C2H4/H2 (4:1) at 525~ It was significant to find that in many of these systems the geometric shapes of the particles were not retained when the specimen was slowly cooled to room temperature and the reactant gas flow switched off. Based on these studies it is clear that under reaction

103 conditions metal particles can adopt well-defined faceted forms that are dependent not only on the nature of the supporting medium, but also chemical composition of the reactant gas. Experimental evidence indicates that it is highly probable that many of these features would not be revealed in a post-reaction electron microscopic examination, since the strength of the metal/support interaction is not always maintained in the absence of a reactant gas and particles appear to relax to the more energetically favorable form of a sphere under these circumstances. 3.2. Influence of metal particle morphology on the structural characteristics of carbon filaments.

Transmission electron microscopy examinations of the solid carbon produced from the interaction of ethylene/hydrogen or carbon monoxide/hydrogen with a number of pure metals and bimetallics containing copper mixed with iron, cobalt and nickel revealed that the deposit consisted almost entirely of filamentous carbon when such reactions were performed at temperatures up to 650~ The amount of solid material produced from the growth of this type of deposit is generally quite high and the presence of the carbon structure serves as a means of "freezing" the shape of the catalyst particle in its reactive state even when the specimen is cooled to room temperature and the reactant gas mixture removed from the system. This aspect can be seen very clearly in the set of electron micrographs, Figures 2 to 4, where the original shapes of the particles shown in the corresponding set of diagrams, Figures 1a to lb are preserved during the catalytic growth of carbon filaments.

lOOnm Figure 2. Appearance of a carbon filament and the associated "diamond" shaped Fe-Ni (5:5) catalyst particle after heating to 600~ in CzH4]H2 (4" 1) at 600~ Inspection of the conformational characteristics of the filaments produced from these three catalyst systems shows that in all three cases, the growth occurs from more than one face of the metal particle. Filaments generated from the interaction of copper-iron with C2H4/H2 (4:1) and iron-nickel particles with CO/H2 (4:1) grow via a hi-directional mode and tend to exhibit a high degree of crystalline order. In contrast, the structures created from the reaction of C2H4/H2 (4:1) with the hexagonal shaped copper-cobalt crystallites are multidirectional with several filaments emanating from a single catalyst particle. As with the previous systems, these carbon structures were also found to be highly crystalline in nature.

104

Figure 3. Appearance of a carbon filament generated from the interaction of Cu-Fe (2:8) with a CO/H2 (4:1) mixture at 600~ Note the geometric shape of the catalyst particle. . . . .

172~.-'.

:~ 9 '='~"

.~i~"'~ ~ ~ . .

.

:~'~,~

lOOnm Figure 4. Appearance of multi-directional filaments formed at the faces of an hexagonalshaped Cu-Co (1:3) particle after reaction in C2H4/H2 (4:1) at 575~

105 In a further series of experiments the filaments formed from these three catalyst systems were studied by high resolution transmission electron microscopy. From the lattice fringe images it was apparent that the structures consist of graphite platelets that are oriented in various directions with respect to the filament axis. A schematic representation showing the structural relationships between the catalyst particles and the precipitated graphite platelets formed during the growth of the filaments from the three systems is presented in Figure 5. It is apparent from these models that in the case of copper-iron/C2H4/H2 (4:1), Figure 5(a), the graphite platelets constituting the filaments adopt a "herring-bone" arrangement and under these circumstances two faces of the catalyst particle are responsible for the dissociation chemisorption of the hydrocarbon and four faces that participate in the precipitation of graphite. When the filaments are produced from the interaction of iron-nickel particles with CO/H2 (4:1), Figure 5b, the graphite platelets are stacked in the form of a "deck of cards" and examination of the associated catalyst particle suggests that at least four faces are involved in the decomposition of the reactant gases and precipitation of graphite occurs at only two faces. Finally, when C2Hn/H2 (4:1) is allowed to react with copper-cobalt several filaments are generated from each particle and the individual structures exhibit identical characteristics where the graphite platelets are aligned in a direction parallel to that of the face where the carbon precipitation step occurs, Figure 5c. Once again, it is possible that in this system there are two faces of the catalyst particle involved in the hydrocarbon adsorption and decomposition reaction and another set of six faces that are responsible for carbon deposition.

a

b

c

Figure 5. Schematic representation of the different arrangements of graphite platelets generated from the three systems shown in Figures 2 to 4, respectively. 3.3.

Formation of graphitic shell-like deposits on iron-nickel particles at >800~

Transmission electron microscope examination of the solid carbon produced on ironnickel after reaction in C2H4/H2 (4: I) revealed the existence of two types of material; carbon filaments and another type of carbonaceous solid, a shell-like deposit in which metal particles appeared to be encapsulated by graphitic platelet structures, Figure 6. This latter form tended to predominate at 825~ Close inspection of many examples of the shell-like deposit failed to reveal any definite correlation between the width of a given catalyst particle and that of the surrounding graphite structure. Inspection of the metal particles showed that they adopted a faceted outline, which showed a close correspondence to that of the carbon deposit that was generated at the catalyst surfaces. High resolution studies revealed that the platelets were very thin, as evidenced by the fact that it was possible to frequently observe the morphological features of the underlying substrate through the graphitic structures. It was also apparent that as one scanned across any of the platelets the electron density remained

106 relatively constant, suggesting that these structures were flat rather than spheroidal in nature. A further feature of note was the finding that the lattice fringes of the deposit were oriented in a direction parallel to the faces of the catalyst particles associated with these platelet structures and had a spacing of 0.34 nm, indicative of a highly graphitic material.

/

200nm ~ - , r

". . . .

~

........

~"~

Figure 6. Electron micrograph showing the shell-like appearance of graphite deposits that encapsulate iron-nickel particles when heated at temperatures of about 825~ in a C~Hn/H2 (4:1) mixture. As mentioned previously, the key steps and factors controlling the growth of the carbon filaments on metal particles are well understood. The range of crystalline order of the carbon deposit generated at the precipitating faces of the particle is controlled to a large degree by the extent of wetting exhibited between the metal and graphite. However, the geometric alignment of the precipitated graphite platelets in the filament structure as well as their degree of crystalline perfection is ultimately determined by the crystallographic orientation of the metal faces in contact with the solid carbon deposit (30-32). While the basics of the formation of the graphitic shell-like deposit have not received the same attention as that devoted to generation of carbon filaments it is probable that many of the steps outlined for the growth of these structures are also operative in the growth of the shell-like form of carbon. In recent years there have been reports of the existence of graphite shell-like structures on supported nickel and cobalt particles following reaction in CO at temperatures over the range 400 to 650~ (9,33). It was claimed that carbon "shells" were the exclusive form of deposit when hydrogen was not present in the gas phase (9). Furthermore, it was stressed that the thickness of the shell surrounding the metal appeared to reach a limiting thickness of about 30 graphite layers at which point the particles were deactivated. This

107 observation leads to the conclusion that the growth of graphite layers was not occurring from the surface of the metal, however, no alternative mechanism was suggested (9). One of the major features to emerge from the present work is the finding that no definite correlation exists between the thickness of the graphitic apron and the size of the associated catalyst particle. A further aspect revealed from this investigation is that the metal particles adopt a well defined faceted structure that appear to function as a template for the growth of graphite platelets. Previous work has indicated that during carbon deposition reactions, catalyst particles acquire well defined crystallographic orientations, where some of the faces will exclusively decompose the hydrocarbon whereas others will only precipitate carbon since they are unable to dissociate carbon-carbon bonds (34).

graphite platelets

(a)

(b)

Figure 7. Schematic rendition of a possible growth mechanism of the shell-like graphite deposits, where the catalyst particle shape changes, but the overall volume remains constant.

a

b

Figure 8. Schematic rendition of a possible growth mechanism of the shell-like graphite deposits, where the catalyst material undergoes a wetting and spreading action with the platelet structures and the particle is progressively depleted in size as the reaction proceeds.

108 Based on the observations of the current investigation we suggest two possible models that should be considered to account for the formation of the graphitic shell-like deposits. The first, which is outlined in the schematic diagram, Figure 7, assumes that during the graphite precipitation step the catalyst particles undergoes a progressive transformation in shape, whilst at the same time maintaining a constant volume. This condition is forced on the system since carbon is being continuously formed on the hydrocarbon decomposing faces, followed by dissolution and diffusion through the particle and eventually precipitating as rigid graphite platelets at other faces. Since the precipitation step imposes a constraint on the particle geometry at any given point in time, the only way for the process to continue in an uninterrupted manner is for the particle to progressively shrink in width in order to accommodate this restriction. A variation on this theme is presented in the model, Figure 8, where the only difference is that in this case, reorganization of the catalyst particle is accompanied by an overall decrease in volume. Indeed, the establishment of a strong interaction of the metal with the edges of the carbon will be manifested by a wetting and spreading action of material along the graphite edge regions. Under these circumstances, it is possible that the volume of the particle is progressively decreased due to the loss of metal as it leaves a monolayer coverage on the graphite edge regions and the catalyst particle can readily conform to the geometric limitations imposed by the platelet structures.

ACKNOWLEDGMENTS Financial support for this work was provided by the Department of Energy, Basic energy Sciences Grant No. DE-FG02-93ER14358.

REFERENCES

1. 2. 3. 4. 5. 6. 7. 8. 9. 10. 11. 12. 13. 14. 15.

L . J . E . Hofer, in "Catalysis" (P. H. Emmett, Ed.) Reinhold Publ. Co. 4 (1956) 373. H.B. Palmer, and C. F. Cullis, in "Chemistry and Physics of Carbon" (P. L. Walker, Jr., Ed.) Marcel Dekker, New York, 1 (1965) 265. J. R. Rostrup-Nielsen, "Steam Reforming Catalysts", Tekorisk Forlay A/S (Danish Technical Press), 1975. D.L. Trimm, Catal. Rev.-Sci. Eng. 16 (1977) 155. R. T. K. Baker, and P. S. Harris, in "Chemistry and Physics of Carbon" (P. L. Walker, Jr., and P. A. Thrower, Eds.) Marcel Dekker, New York, 14, (1978) 83. C.H. Bartholomew, Catal. Rev.-Sci. Eng. 24 (1982) 67. Baker, R. T. K. in "Carbon Fibers Filaments and Composites" (J. L. Figueiredo et al. Eds.) NATO ASI Series, Kluwer Academic Publ. Dordrecht, 177 (1990) 405. N . M . Rodriguez, J. Mater. Sci. 8 (1993) 3233. P.E. Nolan, D. C. Lynch, and A. H. Cutler, Carbon 32, (1994) 477. N. M. Rodriguez, M. S. Kim, F. Fortin, I. Mochida, and R. T. K. Baker, Appl. Catal. in press. J. H. Sinfelt, Accounts of Chem. Res. 10 (1977) 15. W. M. H. Sachtler, and R. A. van Santen, Adv. Catal. 26 (1977). 69 V. Ponec, Adv. Catal. 32 (1983) 149. M. J. Kelley, and V. Ponec, Prog. Surf. Sci. 11 (1981) 139. J. H. Sinfelt, "Bimetallic Catalysts", an Exxon Monograph, John Wiley & Sons, New York, (1983).

109 16. J. G. McCarty, and H. Wise, J. Catal. 57 (1979) 406. 17. J. G. McCarty, P. Y. Hou, D. Sheridan, and H. Wise, in "Coke Formation on Metal Surfaces" ACS Symposium Series 202 (L. J. Albright and R. T. K. Baker eds.) Washington D.C. (1982) p. 253. 18. N. M. Rodriguez, M. S. Kim, and R. T. K. Baker, J. Phys. Chem 98 (1994) 13108. 19. J. R. Rostrup-Nielsen, in "Catalyst Deactivation" (C. H. Bartholomew and J. B. Butt Eds.) Elsevier Sci. Publ. Amsterdam, p.85 (1991). 20. J. R. Rostrup-Nielsen, J. Catal. 85 (1984) 31. 21. M.S. Kim, N. M. Rodriguez, and R. T. K. Baker, J. Catal. 143 (1993) 449. 22. W. T. Owens, N. M. Rodriguez, and R. T. K. Baker, Catal. Today, 21 (1994) 3. 23. W. T. Owens, M. S. Kim, N. M. Rodriguez, and R. T. K. Baker, in "Catalyst Deactivation 1994", (B. Delmon and G. F. Froment, Eds.) Elsevier Sci. Publ. Amsterdam, p.191 (1994). 24. A. Chambers and R. T. K. Baker, J. Phys. Chem. in press 25. X. Wu, B. C. Gerstein, and T. S. King, J. Catal. 135 (1992) 68. 26. R. J. Verderone, C. L. Pieck, M. R. Sad, and J. M. Parera, Appl. Catal. 21 (1986) 239. 27. J. H. Sinfelt, J. L. Carter, and D. J. C. Yates, J. Catal. 24, (1972) 283. 28. N. M. Rodriguez, S. G. Oh, W. B. Downs, P. Pattabiraman, and R. T. K. Baker, Rev. Sci. Instrum. 61 (1990) 1863. 29. N. M. Rodriguez, S. G. Oh, R. A. Dalla-Betta, R. T. K. and Baker, J. Catal. 157 (1995) 676. 30. M. Audier, A. Oberlin, M. Oberlin, M. Coulon, and L. Bonnetain, Carbon 19 (1981) 217. 31. R. T. Yang, J. P. and Chen, J. Catal. 115 (1989) 52. 32. M. S. Kim, N. M. Rodriguez, and R. T. K. Baker, J. Catal. 134 (1992) 253. 33. M. Audier and M. Coulon, Carbon 23, (1985) 317. 34. N. M. Rodriguez, A. Chambers, and R. T. K. Baker, Langmuir 11 (1995) 3862.

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Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

111

Self-Poisoning and A g i n g of Pd-Ag/A1203 in S e m i - H y d r o g e n a t i o n o f 1,3Butadiene: Effects o f Surface Inhomogeneity Caused by H y d r o c a r b o n a c e o u s Deposits A. Sarkany Institute of Isotopes of the Hungarian Academy of Sciences, H-1525, Budapest, Hungary Effects of hydrocarbonaceous deposits have been investigated in hydrogenation of 1,3butadiene over Pd-Ag/ALO3 catalysts. Studies on the effect of various treatments on selfpoisoned samples revealed very different behavior of the surface deposits. Because of hydrogen's low surface fugacity, hydrogen treatment was not sufficient to remove surface species at T<373 K. Vacuum or 02 treatment increased the activity of the self-poisoned samples. The observation has been interpreted in terms of surface restructuring of the firmly held species. Deliberate poisoning - amount of deposits 1-20 mg/gcat- increased the selectivity of n-butane formation. Presence of deposits seems to decrease the effective surface concentration of butadiene. The results have been interpreted as slow diene transport allowing saturation of n-butenes on a certain fraction of sites. 1. INTRODUCTION Pd-based catalysts are widely used in selective hydrogenation of acetylene in ethene or butadiene in C4 alkene cuts [1]. The objective of hydrogen addition is to decrease the concentration of these highly unsaturated hydrocarbons to ppm level without the hydrogenation of alkenes. Because of low HJHC ratios and the high reactivity of acetylene (diene), formation of oligomers and surface deposits is almost inevitable in these processes. Considerable attention has been focused on the understanding of the surface composition of the working surface and the role of hydrocarbonaceous materials in the hydrogenation process. Surface oligomers [2-6] and various multiply bonded so-called spectator species [79] are proposed to participate in formation and stabilization of reacting sites [10], and in altering hydrogen-storing and hydrogen-transfer capacities [ 11,12]. They seem to modify the strength of diene (acetylene) complexation relative to alkene, alter the intrinsic selectivity of acetylene hydrogenation and create hydrogen transfer sites on the support [2,4,11,12]. The present paper investigates the effect of hydrocarbonaceous materials on hydrogen addition using gas phase hydrogenation of 1,3-butadiene as probe reaction. Two different types of deposit formation were tested: (i) initial stage of deposit formation which occurs during self-poisoning and (ii) aged stage when the catalyst is deliberately covered with oligomers. 2. EXPERIMENTAL

2.1. Catalyst preparation and characterization Catalysts used in the present study are listed in Table 1. Pd-Ag supported catalysts (Catalyst A, C and D) were prepared using low surface area alpha alumina.(A13980-T and

112 A1 1008-T, respectively). The support was impregnated with a solution of Pd(NO3)2 and AgNO3. The precursor was dried at 370 K overnight and reduced in stream of hydrogen at 723 K. Pd and Ag content of the precursors were determined by X-ray fluorescence. Table 1 Catalysts prepared and results of characterization Cat. wt%Pd A 0.11 C 1 D 1 Pd powder

wt%Ag 0.17 0.5 0

S(BET) 4 5.5 5.2 7.7

Vp(cc) 0.27 0.3 0.3 -

CO (umol/gcat) 0.44 3.5 12.1 183

Palladium powder was prepared by mixing solutions of PdC12 and hydrazin and adding solution of NH,OH. The sample was carefully washed and treated in 02 at 523 K in order to stabilize the sample. The reduced precursors were characterized by CO (H2) adsorption measurements performed in a small reactor attached to a quadrupole mass spectrometer and UHV system. The reduced samples were saturated with CO (H2) at 298 K (1.33 kPa CO), the reactor was then evacuated (10 "3Pa) and the next thermal desorption test was applied. The experimental setup was also used to follow hydrogen and hydrocarbon desorption from aged catalysts. The carbon content of the poisoned samples (self-poisoning and deliberate poisoning at T>373 K in diene or acetylene)was measured by TPO (1.3 kPa 02, flow and/or static systems, respectively; CO2 (m/z=44) and CO (m/z=28) were detected by mass spectrometer (AEI MS 20 and Balzers QMG 511). DSC thermograms (Perkin Elmer DSC-4) were recorded in 02 stream (40 K/min ramp from ambient to 773 K). Formation of Pd-Ag alloy (Cat. C) and the presence of PdCx phase after poisoning were analyzed by X-ray diffraction. CuKa X-ray powder diffractograms were recorded with a Guinier camera equipped with curved quartz monocromator. 2.2 Conditions of hydrogenation Hydrogenation of 1,3-butadiene (BD) was tested in a static circulation system (batch reactor) of 0.165 dm 3 total volume attached to a high vacuum system [13]. Gas circulation was accomplished with a pump driven with solenoid coils. The pumping speed at atmospheric pressure was 2.52 dm3/min. Hydrogen was purified by diffusion through Pd-Ag thimbles. Deuterium (99.7 %) was purchased from Linde. Diene was distilled prior to use and stored in a bulb. It contained 0.12 % cis-2-butene which was considered in the calculation of the activity and the selectivity values. The experiments were performed at 284 K in the presence of 1.33 kPa BD and only 2.99 kPa H2 to facilitate poisoning of Pd sites. Selectivity was defined in accordance with previous papers [3,13]; rate of product formation was divided with the rate of diene consumption. All through the text conversion refers to transformation of butadiene. Rate is given in %/min units. Conversion was measured as a function of time; in each experiment the initial rate was calculated from the inlet and the first two analysis at <5% conversions, subsequent rate was determined from the next three measurements and we also presented an "average"value calculated from the last point and the inlet composition. Transformation of BD was measured up to 55 % conversion. The catalysts were aged by performing successive hydrogenations. Between activity tests standard regeneration treatment was applied unless otherwise stated. The standard treatment consisted of the following steps: first the hydrocarbon was frozen out in a liquid N2 trap, and next the hydrogen partial pressure was increased to 3.3 kPa and circulated at 284 K for 20 min. After that the reactor was separated from the pumping part, the liquid nitrogen trap

113

was removed, the hydrogen-hydrocarbon mixture was evacuated from the pumping part and finally the hydrogen was removed from the reactor tube before beginning a new experiment. 3. RESULTS Successive experiments performed at 284 K with H/BD=2.2 mixtures resulted in severe poisoning of the reaction sites in spite of hydrogenation (standard treatment) between the experiments. Typical results measured on Cat. A are presented in Fig. 1.

lO01

E v

lOt

~f

A B

I 1_

+'

C D '

E

+_

'

+ 1

0

n 11 0".'011

lllli'i+llllI, I IIII|

h.

tl.

" "I Jl w3

9

================================

10

20

Successive experiments

30

Figure 1. Rate of BD consumption in successive runs (H/BD=2.2, T=284 K, 54.3 mg cat). Between rtms Cat. A was hydrogenated at 284 K with the exception of the experiments indicated by the arrows. A: O5 treatment at 388 K, B: evacuation at 383 K, C: H2 at 371 K for 40 min, D: H~ at 429 K for 10 min, E: evacuation at 403 K for 10 min. Symbols: o initial rate; FI subsequent rate; O "average". Under the reported conditions the activity stabilized at about 1% of the initial value. Replacement of H2 with D2 (D2 in Fig. 1) decreased the rate of hydrogen addition, and for self-poisoned (stabilized) samples R"/R~ for consumption of butadiene was observed to be 2.3-2.4 at 284 K. 02 (arrow A) or vacuum (arrow B) treatment at T<413 K of the selfpoisoned samples increased the hydrogenation activity as shown in Fig. 1. Smaller effects were observed upon H2 treatments (arrows C and D). Similar aging curves were observed with other catalysts. The selectivities of n-butanes over Cat A and C were initially less than 0.05% at 40-55% conversion and were not affected by the above treatments. Over Cat D and Pd powder 0.5-1% n-butane appeared in the first experiments but their selectivities gradually decreased with aging. The hydrocarbon content of the samples were measured after H2 treatment at 298 K. The C/CO,d, ratios measured (the number of successive runs are in brackets) are as follows: Cat. A (5) 3.3, Cat. B (31) 5.8, Cat. C (5) 2.7. After '5-10 successive experiments further aging of the self-poisoned samples was slow; therefore we deliberately poisoned our aged catalysts with acetylene and diene in the absence of H2 at T>373 K. Typical results for Cat. A are presented in Figure 2; results for Cat. C are in Table 2. As shown in Figure 2 poisoning at T>373 K first increases the hydrogenation activity. A further interesting feature of the experiments is that after deliberate poisoning the n-butane selectivity, S(nB)%, gradually increases and reaches 4-5% over Cat A (Figure 2), 10-11%

114

A~

B

C

D

m

E 0.1 O

IB..

ill

0.01 6 II1 v

o .................. j,,,,,,, .....

...... IIIIII.

10 30 50 Number of successive experiments Figure 2. Effect of diene poisoning on hydrogenation activity (Hz/BD=2.2, T=284 K, 51.4 mg Cat. A). A: evacuation at419 K; B: acetylene treatment at 433 K for 4 hr; C: BD at 455 K for 17 hr and D: BD at 492 K for 17 hr. Table 2 Hydrogenation of 1,3-butadiene and 1-butene over Cat. C (10.3 mg cat, 1.33 kPa HC and 3 kPa H2, T=284 K) Exp

EBD(kJ/mol)

RI-B/RBD

R"/RD(BD)

R"/RD(nB) S(nB)%

5-11

46+13.2

3.2-2.8

2.2-2.4

-

0.04

39-48 49-53* 57-63 65-72 77-87 91-100

41+6.6 40 n.m. n.m. 33+3.3 22+5.4

0.96-0.71 n.m. n.m. n.m. 0.18-0.26 0.09-0.12

1.6-1.48 1.52 1.48-1.52 1.5-1.32 1.46-1.35 n.m.

1.2-1.1 1.35 1.16-1.21 1.28-1.06 1.22-1.20 n.m.

6.5-7.5 6.6 8.2 6.3 9.9-7.6 11.1

=,_

m

5-11, self-poisoning; (38) acetylene (409 K, 17 hr); (56) acetylene (423 K, 15 hr); (65) H2 (467 K, 12 hr); (76) BD (461 K, 16.5 hr); (90) BD (488 K, 19 hr); n.m.: not measured; *experiments at 302 K

115 Table 3 Deposit formation from 1,3-butadiene over Pd catalysts (8-10.5 kPa BD). S(nB)% was measured at 284 K (HJBD=2.2, 1.33 kPa BD). Catal.

T/K

Cat. A Cat C.

t/hour

487 407 446 489 407 487* Cat. D 407 443 489 Pd-black 391 433 *-acetylene treatment

1.5",

H/C

21 21 21 18 21 21 21 21 18 67 21

1.38 1.32 1.31 1.26 1.32 1.15 1.12 1.11 1.03 1.12

mol C/gcat

mg dep/~cat

S(nB)%

1.1xl0 4 1.7x10 4 2.4xl 0 .4 3.2x 10-4 1.7X10-4 5.0xl 0 4 1.5xl 0 -3 2.0x10 3 1.8x10 -3 5.9x10 -4 1.6x10 -3

1.4 2.3 3.1 5.1 2.3 6.1 20.3 26.5 24.2 7.8 20.9

4.7 1.7 n.m. 6.7 1.7 8.8 10.1 15.3 13.5 n.m. n.m.

p/i

QMS: H2x1072 / / / ~ ~ I I

g 1.o ._~

I

l

r

--= 0.5

I

5 .......

373

573 T/K

773

773

9

4..i

~.~ ~ ~

--

\

9#"

I 373

X

I

I

573 T/K

,

I

-,,.

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773

Figure 3. TPD of hydrogen over Cat. C 1- Figure 4. DSC thermograms in 02 (40 regenerated sample (02 at 623 K and H2 at K/min). 1-Pd black (393 K, 4 hr); 2 Cat. C 713 K); 2- after BD poisoning at 489 K for (489 K, 18 hr); 3-Cat C. (489 K, 18 hr) after 18 hr (see Table 3). He at 675 K 4-Pd black (433 K, 21 hr) and 5Cat. D (489 K 18 hr). over Cat. C and 10-15% over Cat D (Tables 2 and 3). The presence of foulant decreases the consumption of 1-butene to a greater extent than hydrogenation of diene; the RI_B/RBDratios are included in Table 2. Over Cat A R,.B/RBDis 2.1 and 0.11 in experiments 9-10 and 59-60, respectively. The activation energy of hydrogenation (BD consumption) also decreases (see Table 2). Over Cat. A the following values are observed (number of run in brackets): 48+8 (27-30), 36+6 (42-45) and 32+4 kJ/mol (55-58). After these experiments the carbon content of the samples was 1.1 mg/gcat on Cat A and 7.4 mg/gcat on Cat C.

116 Data on the formation of carbonaceous materials from diene and acetylene over 373 K and selectivity of n-butane in butadiene hydrogenation (H2/BD=2.2, 1.33 kPa BD, 284 K) are collected in Table 3. The hydrogen content of the deposits was calculated from the amount of the self-hydrogenated products in the gas phase. Reaction of 1,3-butadiene was accompanied with formation of n-butenes whereas in reaction of acetylene 1,3-butadiene was the principal product. Hydrogen content of the deposits was also measured by TPD technique. Desorption of H2 is presented in Figure 3 measured on Cat. C before (1) and after BD poisoning at 489 K for 18 hr (2). Typical DSC curves with diene poisoned catalysts are presented in Figure 4. Over Pd black a single exothermic peak appears between 408-498 K. Over supported samples the high temperature peak (Tmax553-573 K) can be attributed to oligomers on support sites. XRD measurements were performed with Cat C, D and Pd powder treated with butadiene (4-8 kPa) at 453-473 K for 4-17 hr. The diffraction pattern for Cat C has not indicated any convincing changes in the position of Pd-Ag diffractions. Solution of carbon in Pd (Cat D. and Pd powder) resulted in an expansion of the Pd lattice. XRD lines attributable to PdCx phase could be discerned at theta values: "19.538, "22.718, "33.117, *39.763, "41.999. It is intriguing that the dissolution of carbon in Pd was accompanied with only trace CH4 formation. 4. DISCUSSION In the present experiments two rather different stages of catalyst poisoning are reported as indicated by the amount of hydrocarbon left on the surface in course of self-poisoning and after deliberate poisoning with butadiene or acetylene at T>373 K. After 5-8 experiments the hydrogenation activity reached steady state and the hydrocarbon coverage (number of carbon atoms on CO chemisorption sites) was 3.3-5.8 on Cat. A and 2.7 on Cat. C. The deliberate poisoning in the presence of butadiene and acetylene above 373 K further increased the amount of deposit and after completing the experiments it reached 1.1 mg/gcat over Cat. A and 7.4 mg/gcat over Cat. C. Assuming a density of oligomers of ca. 0.8 g/ml at 294 K, the thickness of hydrocarbon layer over BET surface corresponds to about 0.3 nm on Cat. A and 1.7 nm over Cat. C. At such a high hydrocarbon loading the metal surface (or a fraction of it) cannot be reached directly from the gas phase and thus the transport hindrance of the reactants cannot be avoided. The self-poisoning experiments have confirmed that hydrogen treatment at 284 K between the experiments is not sufficient to restore the initial activity. We assume that hydrogen's low surface fugacity is the main reason that the surface is covered by multiply bonded species [9-12] which are "end attached" (butylidyne, 3-butenylidyne) or "flat lying" over metal atoms (di-sigma, quad-sigma, di-sigma/pi) although formation of surface oligomers cannot be entirely disclosed [8]. Formation of C8 oligomers was not detected in our hydrogenations. This result however does not exclude the limited formation of high molecular weight oligomers and their presence on Ag sites or on the support. After 5-6 runs the activity usually stabilized at about 1% of the initial activity. We assmne that the working sites involve Pd atoms of low coordination number on which the accumulation of deposits is prevented by the ease of hydrogen dissociation. This assumption is supported by the rate of diene hydrogenation and the sticking coefficient for hydrogen which are higher on the more open (110) than on the (111) surfaces of Pd [ 14,15]. One of the intriguing observations of the present experiments is that hydrogen treatment of the self-poisoned catalysts (see arrows C and D in Figure 1) is less effective in the regeneration of the active surface than for example the vacuum treatment. In principle both 02 and vacuum treatments facilitate hydrogen loss, desorption and/or surface polymerization in the adsorption phase. We assume that the above processes result in restructuring of the surface hydrocarbons on the metal sites creating thereby sites for hydrogen dissociation. Although our assumption requires spectroscopic verification the TPO measurements have shown that formation of CO2 starts around 373 K but the amount of CO2 desorbed is only

117 about 2-3% of total hydrocarbon coverage. The activity increase is apparently connected with small changes in the concentration of free surface sites. In separate measurements (selfpoisoned Cat. C. heated up to 477 K in a reactor attached directly to quadrupole mass spectrometer) desorption of hydrogen, diene and n-butenes could be detected. Selfhydrogenation or reversal of sigma bonded species into pi-olefin bonded ones are the possible routes for the formation of these hydrocarbons. One might also propose that formation of carbon atoms and their migration into the Pd lattice during high temperature evacuation or hydrocarbon treatments above 373 K provide a route to site regeneration. Formation of PdCx (x<0.15)supersaturated carbon solution in course of deliberate poisoning could be unambiguously detected with Cat D and Pd powder [16,17]. However, absorption of carbon might not be significant in Pd-Ag alloys. In fact, in Cat C we could not detect any PdCx phase. Deliberate poisoning at T>373 K produces unsaturated surface deposits (Table 3) present both on metal particles and support. The molC/gcat values in Table 3 (compare Cat C and D) show that the presence of Ag inhibits formation of deposits. Over both Pd black and Cat D the H/C ratio is less than over Pd-Ag samples. As evidenced by DSC thermograms a significant part of the foulant is present on the support. In contrast to the samples used in acetylene hydrogenations in the present DSC thermograms an initial endothermic peak indicative of the evaporation of low molecular weight oligomers was not observed. Treatment of fouled samples in He decreased the intensity of the high temperature DSC peak which can be interpreted as the evaporation of the oligomers from support sites (compare (2) and (3) in Figure4). Inspection of Figure 2 shows that the presence of oligomers does not result in drastic loss of hydrogenation activity. Actually after the first acetylene or diene poisoning experiments the hydrogenation activity was higher than after self-poisoning. Upon repeated poisoning the activity decreased (see arrows C and D in Figure 2) due to the accumulation of deposits on the catalyst. We assume that the onset of transport control appears in this region. The effect of deposits on the distribution of n-butenes was not significant. Accumulation of deposits apparently markedly affects the formation of n-butane. The selectivity of n-butane increases gradually after diene treatments and finally reaches 5.5 % over Cat. A (Figure 2) and 10-11% over Cat. C (Table 2) which is a tremendous increase if one considers that in the self-poisoning experiments at low conversions (<5%) n-butane can not be detected at all and at 50-65% conversions its selectivity is still less than 0.05%. In previous studies on hydrogenation of acetylene in ethene it has been concluded [2,4] that ethane forms mainly from acetylene but with time on stream due to activation of some special sites hydrogenation of ethene takes place even in the presence of acetylene. Returning to the present study the hydrogenation rate of 1-butene was observed to decrease to a larger extent than the hydrogenation of diene (see R,_JRBD in Table 2). Over fresh catalyst C the above ratio is about 4.5-4.2. After self-poisoning (experiments 9-11) the observed ratio is 3.2-2.8. On catalyst poisoned deliberately by BD and acetylene treatments the ratio is only 0.09-0.12 (experiments 92-95). Similar results were observed over Cat A. This suppression of 1-butene hydrogenation assures that in the presence of butadiene n-butane will not form in any appreciable amount from n-butenes desorbed into the gas phase. It requires further investigations to decide whether the decreasing R,_B/R~D ratio is the consequence of a very different flux (transport rate) of diene and 1-butene through the oligomers (which is unlikely) or that surface sites are less available for 1-butene due to its small complexation strength as observed over self-poisoned Pd-Ni/Nb205 catalysts [ 18]. The drastic increase of n-butane selectivity upon diene or acetylene treatments at T>373 K suggests formation of some special sites on the initially highly selective surface which favor over-hydrogenation of butadiene. Our TPD measurements show (Figure 3) that the foulant supplies only a very limited amount of hydrogen at 284 K. It is tempting to attribute n-butane forming sites to Pd atoms modified by subsurface carbon atoms. It is not obvious, however, why these sites are not present after vacuum treatment of the self-poisoned samples at T>423 K. Formation of n-butane can be accounted for if one assumes that the

118 formation of this product is controlled by transport limitation rather than by the appearance of some special sites which have excellent hydrogen availability and favor the formation of carbene type intermediates. The latter species are proposed precursors in the reaction route leading to n-butane [5]. The appearance of transport limitation is supported by the significant decrease of the apparent activation energy of hydrogenation both over Cat. A (from 48 kJ/mol to 32 kJ/mol) and over Cat. C (from 46 kJ/mol (self-poisoned sample) to 22 kJ/mol. Replacement of hydrogen with deuterium only slightly affected the rate of n-butane formation: R"/RDwas observed to be 1.1-1.2 for n-butane formation and 2.4-1.3 for diene hydrogenation. The results might be interpreted in such a way that due to transport control equilibrium effects compensate for kinetic effects. We assume that the oligomer film itself induces inhomogeneity in the working of the reaction sites. As a consequence of slow diene transport through the oligomers the surface fugacity of diene becomes low, consequently, nbutenes formed by hydrogen uptake are not readily displaced by diene molecules on a certain fraction of sites. Apparently, there is sufficient hydrogen on these sites. One would expect that the solubility of hydrogen in the oligomer layer is less than that of diene [ 19]. It is not clear however how these unknown values are compensated by the transport (diffusion) coefficients and how these values are affected if the oligomers are high molecular weight molecules.

Acknowledgments: The financial support of the Grant OTKA (T 017047 and T 019128) is gratefullyacknowledged. REFERENCES 1. 2. .

4. 5. 6. 7. 8. 9. 10. 11. 12. 13. 14. 15. 16. 17. 18. 19.

J-P. Boitiaux, J. Cosyns, M. Derrien and G. Leger, Hydrocarbon Processing, (1985)51. W.T. McGown, C. Kemball, D.A. Whan and M.S. Scurrel, J.Chem. Soc. Faraday Trans. I., 73 (1977) 632. A. Sarkany, A.H. Weiss, T. Szilagyi and P. Sandor, Appl. Catal., 12 (1984) 373. A. Sarkany, A.H. Weiss and L. Guczi, Appl. Catal., 12 (1984) 373. M. Primet, M. E1 Azhar and M. Guenin, Appl. Catal., 58 (1990) 241. S.D. Jackson and N.J. Casey, J.C.S. Faraday Trans., 91 (1995) 3269. T.P. Beebe and J.T. Yates, Jr., J. Am. Chem. Soc., 108 (1986) 663. T.J. Campione and J.G. Ekerdt, J. Catal., 102 (1986) 64. M.A. Chesters, C. Cruz, P. Gardner, E.M. McCash, P. Pudney, G. Shadid and N. Sheppard, J.Cem. Soc. Faraday Trans., 86 (1990) 2757. I. Yasumori, S. Moriki, I. Kojima and Y. Inoue in Proc. 6th Intern. Congr. Catal., London 1976, The Chemical Society, London, Vol.1, 1977, p. 139. S.J. Thomson and G. Webb, J. Chem. Soc., Chem. Commun., (1976) 526. A.S. A1-Ammar and G. Webb, J. Chem. Soc. Faraday Trans., I, 74 (1978) 1900. A. Sarkany, G. Stefler and J.W. Hightower, Appl. Catal. A, 127 (1995) 966. J. Massardier, J.C. Bertolini and A. Renouprez, in Proc. 9th Int. Congr. Catal., Calgary, Vol.3, 1988, p. 122. R.J. Behm, K. Christmann and G. Ertl, Surf. Sci., 99 (1980) 320. J. Stachurski and A. Frackiewicz, J. Less Comm. Met., 108 (1985) 249. S. Ziemecki, G. Jones, D. Swartzfager, J. Less Comm. Met., 131 (1987) 157. A. Sarkany, Z. Schay, Gy. Stefler, L. Borko and J.W. Hightower, Appl. Catal. A, 124 (1995) L181. R.H. Perry and C.H. Chilton (eds.), Chemical Engineer's Handbook, 5th Ed. McGraw-Hill Publ., New York, 1974.

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Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

119

I n f l u e n c e o f the S u p p o r t o n the D e a c t i v a t i o n o f N i c k e l Z e o l i t e Catalysts D u r i n g the C o n v e r s i o n o f P h e n y l a c e t y l e n e E.D. Gamas, and I. Schifter Gerencia Transformaci6n de Energ6ticos, Subdirecci6n de Protecci6n Ambiental Instituto Mexicano del Petr61eo, Eje Central L~izaro CS_rdenas 152, Col. San Bartolo Atepehuacan M6xico, D.F., 07730 M6xico The low-temperature (20~ deactivation by coke of zeolite-supported nickel catalysts was studied. Using the hydrogenation of phenylacetylene as a model reaction, channel dimension was found to affect both the loss of catalyst activity and the distribution of nickel particles on the active surface; activity was found to decrease in the following order: ZSM5>HUSY>Hmordenite. Selectivity'curves as well as temperature-programmed reduction runs show that the majority (95%) of the nickel particles is deposited on the outside of the channel network of HZSM5 support while for Hmordenite and HUSY supports the majority of the nickel is found inside the channels. In spite of coke being formed at low reaction temperature, temperatures as high as 790~ were required for its removal. Both Hmordenite and HZSM5 also contained substantial quantities of reactive coke that burned at lower temperature, while the HUSY contained substantial quantities of high-temperature coke. 1. I N T R O D U C T I O N One of the major problems in the operation of industrial catalytic processes is the loss of activity and selectivity due to the deactivation of the catalysts by coke. It is also a well known fact that "coke" represents a combination of different carbonaceous materials whose structure depends on the operation conditions of the reaction and on the nature of the catalyst itself. Although most processes involving catalyst deactivation by coke operate at temperatures above 300~ it is also possible to have catalyst deactivation by coke formation even for reactions that occur at ambient temperatures (20~176 [ 1,2]. The nature of coke formed at low temperature must be expected to vary in nature from that of coke formed at higher temperatures. It is generallyagreed that coke is formed via polymerization of reactions of unsaturated hydrocarbons chains. In order to test the influence of steric effects, the effects of acidity and the location of the supported metal, in this work the hydrogenation of phenylacetylene was used to study the deactivation by coke of nickel catalysts based on zeolite supports. Three zeolites were chosen as supports: ZSM-5, ultrastable Y and mordenite; all three zeolites were in their acid form. The ZSM-5 sample (Surface area = 402 m2/g, SiO2/A1203 = 42) was synthesized following procedure #23 outlined by Argauer and Landlot [3]. The USY sample used was a CP100-35 zeolite from Valfour USA (Surface area - 600 m /2 g , SiO2/A1203 = 6.5). Mordenite was obtained from purified samples of a natural mordenite from Tamaulipas, M6xico (Surface area - 495 mE/g, SiO2/AI203 = 5.91). The low temperature reaction as presented in this study yields carbonaceous deposits that required temperatures as high as 790~ for their removal from the catalysts. Nickel seems to be located in the intemal channels as well as on the surface of the supports.

120 Studies of catalysts deactivation by coke are abundant in the literature; most of them are usually conducted at high temperatures (around 500~ using metal catalysts supported on oxides with low surface area such as silica, aluminas or silica-alumina [2 and references therein]. The deactivation by coke of zeolite catalysts has also been studied and such studies have mostly been done for high temperature reactions such as the conversion of n-hexane or the isomerization of xylenes [2,4]. However, low temperature coke formation (20-25~ combining the effect of high acidity and size specificity for a high coking component such as nickel, has not yet been considered from the point of view of the presence of compounded effects of crystalline structure and location of metal particles. 2. E X P E R I M E N T A L Nickel was deposited by impregnation of zeolite powders using the incipient wetness method with a 3% wt solution of Ni(NO3) 2 96H20 (Baker Analyzed). The solids were dried at 100~ in air for 24 hours and calcined in flowing air (2.4 l/h) at 500~ for 2 hours; the nickel content of all samples was 2.5% wt. Before reaction the catalysts were prereduced in flowing hydrogen (2.0 I/h) at 500~ for 2 hours. Conversion of phenylacetylene [Merck, 97.8% wt] was conducted in a continuous microreactor system using 0.05 g of catalyst at a temperature of 20~ and atmospheric pressure. The feed consisted of hydrogen (1.8 1/min) saturated with liquid phenylacetylene at a temperature of 12~ (3.40x10-~ g-mole/h of phenylacetylene). The solids were characterized by X-ray diffraction using a Siemmens D-500 diffractometer equipped with a copper anode and a nickel filter. The deposition of carbon was studied by thermogravimetric analysis (TGA) using a Perkin-Elmer 1700 TGA analyzer. For the TGA runs the sample was heated at 10~ in air (2.0 l/h) from ambient temperature (20~ to 850~ the procedure was repeated to eliminate the effect of adsorbed water. The effect of the support on the reducibility of nickel was measured by temperatureprogrammed reduction (TPR). Samples were pretreated in air at 10~ from ambient temperature to 500~ After purging the lines with helium and cooling the reactor, the reduction process was initiated from ambient temperature to 1000~ at heating rate of 10~ in a flow of 2.0 1/min of 5.0% wt hydrogen in argon. The consumption of hydrogen by the sample was evaluated in a Gow Mac 550 gas chromatograph equipped with a thermal conductivity detector. 3. R E S U L T S AND D I S C U S S I O N Structural characterization of the solids by X-ray diffraction showed that the materials ZSM-5 and USY zeolites studied had a high crystallinity, while a mordenite structure was observed for the natural zeolite; traces of clinoptilolite were observed in this sample. The reaction studied follows the stoichiometric path shown below:

~ ~ ~ . ~ C-- CH "j H2

-- CH2 CH+

CH2- CH3

The results of catalyst testing runs are shown in Figures 1 and 2. With a critical molecular size of 0.5-0.55 nm for all three molecules involved in the reaction [5], the influence of pore dimensions is clearly seen as nickel deposited on a ZSM-5 support (average channel size of 0.55 nm [6]) does not deactivate rapidly, while nickel supported on USY zeolite (average channel size of 0.77 nm [6]) and nickel on mordenite (average channel size of 0.68 nm [6] )

121

have lower initial conversions and deactivate rapidly. Given the low reaction temperature (low molecular excitation) the possibility of having a dynamic molecular diameter for molecular diffusion of reactant into the channels of the zeolites cannot be expected.

40.00 NV[ES

30.00 20.00 10.00

-

Ni/HMor~,~ 0.00

~

T

_ "

i

10

0

-

9

-

20

-

30

40

Time, min Figure 1. Conversion of phenylacetylene for nickel supported on ZSM5, USY and mordenite zeolites.

15.00 O A

10.00

5.00

A

A

-

-

[]

0.00 0

v

A w

m

m

--

9"

10

.

tin, v

A

A

m

m

m

_

~

20

_ytmZSM_...~ Ni/USY m

m

30

40

Time, min Figure 2. Selectivity ratio of styrene to ethylbenzene during phenylacetylene.

the conversion

of

Due to the close match of molecular size of reactants and products with the size of the channels for the supports used, it should be expected that for the smaller-pore ZSM-5 support the reaction takes place outside of the channel network, while for larger-pore USY and mordenite the reaction is more likely to take place inside the channels and cages where the acid sites are not only of stronger acidity but their population is also greater; this results in a greater coke producing activity in comparison with that of external acid sites. The ZSM-5

122 support exerts therefore a reactant selectivity to diffusion which is reflected in the low deactivation observed (see Figure 1). The loss of activity due to coke deposition also has an effect upon the selectivity of the reaction. As seen in Figure 2, for the ZSM5 support a constant selectivity ratio styrene/ethylbenzene is observed; this is consistent with the conversion vs. time curve in Figure 1, which remains almost constant independent of the amount of coke deposited. The constant ratio styrene/ethylbenzene of approximately 2.7 indicates that, for Ni/ZSM5, 75% of the molecules of phenylacetylene go through a single step of adsorption on hydrogenation sites (nickel sites) located on the external surface of the zeolite, i.e., the greater amount of nickel particles are deposited on the outside of the cavities of the crystalline support. The temperature-programmed reduction (TPR) curve for Ni/ZSM5 is given in Figure 3. Four main peaks of hydrogen consumption are seen (maxima at 440, 558, 676 and 1020~ respectively); the first peak (maximum at 440~ corresponds to a typical reduction of supported nickel oxide [7], i.e., supported nickel oxide present on the outer surface of the zeolite. The ratio of areas for the peak with a maximum at 440~ with respect to the peaks observed above 550~ is 1.58; thus, the TPR data suggest that over 50% of the nickel oxide is located on the external surface of the support. One can conclude therefore that the TPR peaks at temperatures above 500~ may correspond to the reduction of nickel particles located inside the cavities of the zeolite, some in the channels and some in the cages of the structure. This conclusion is supported by the observation that the peaks attributed to the reduction of nickel oxide are mostly found at temperatures above 500~ for zeolites with the greater channel opening (Figures 4 and 5). The larger size of the cavities should facilitate nickel nitrate diffusion into the cavities for mordenite and USY. Concerning the distribution of supported nickel either inside or outside the zeolitic structure, Jaeger et al., [8] reported experimental evidence by X-ray Photoelectron Spectroscopy that metal phases in reduced ionexchanged zeolites form preferentially inside the zeolite matrix. Delaffose [9] reported the deposition of nickel inside Faujasite channels and gives a correlation between Ni +2 location at the acid sites of the zeolite (via X-ray Diffraction studies) and the reducibility of the Ni § cation. However, neither of these two papers report quantitatively the extent of the reducibility of nickel via TPR. The similarly large deactivation effect caused by coke deposition for nickel supported on mordenite and on USY zeolitic materials is clearly seen in Figure 1. A basic difference between these two materials is the effect observed on the selectivity ratio styrene/ethylbenzene (see Figure 2). Although both materials have nearly the same deactivation profile, the high selectivity observed in Figure 2 for Ni/mordenite is probably due to shape selectivity. Since USY zeolite consists of supercages and windows while mordenite consists of pores and supercages connected in a straight line [4,10], the availability of alternative exits for reaction products in Ni/USY probably explains its lower selectivity. Moreover, the channels of mordenite only allow one-way molecular traffic. In addition, since the reaction follows a sequential stoichiometric path and given the higher intricacy of the structure~ of USY with respect to that of Mordenite, in terms of residence time, a molecule of phenylacetylene is more likely to proceed through a full completion of the hydrogenation stoichiometry since the probability of hitting a hydrogenation site is higher for the USY than for the mordenite support. Another aspect is the effect of the loss of population of the more active hydrogenation sites associated with nickel in USY compared to those of nickel in the Mordenite channels [9]. The shape of the TPR curves for Ni/mordenite and Ni/USY (Figures 4 and 5) provides evidence that NiO is reducible only at high temperatures, a behavior which is not characteristic of nickel oxide supported on low surface area carders [7]. The conclusion is that for zeolites the majority of nickel particles is located inside the crystalline structure of the supports. This is also supported by the results of coke deposition as studied by TGA. As seen in Tables 1 to 3, the total amount of coke found for each catalyst sample explains the deactivation curves in Figure 1 only if the structural aspect is considered. Nickel supported on USY contains 4.5 times more coke than nickel supported on ZSM5 and 2.6 times more

123 coke than nickel supported on mordenite; however, both Ni/USY and Ni/mordenite deactivate at the same rate. The coke deposited in mordenite is adequate to cause blockage of the parallel channels while for Ni/USY a larger amount of coke is necessary to block all possible 1.2

o -~

0.8 0.6 0.4 0.2 0 -0.2

i

! i i

0

400

200

600

800

1000

1200

Temperature, ~

Figure 3. Temperature-Programmed Reduction pattern for Ni/ZSM-5. 1 0.8 0.6

I

0.4 0.2

o i

-0.2

! 0

200

400

600

800

1000

1200

1400

Temperature, *C

Figure 4. Temperature-Programmed Reduction pattern for Ni/USY. exits in the faujasite structure. For Ni/ZSM5, since the reaction takes place on the external surface of the support where the acidity is low and the promotion of coke-forming polymerization reactions does not occur, a smaller quantity of coke is expected. The relatively large amount of high temperature coke seen for Ni/USY (see Table 2) provides evidence that the acid sites located in the interior of the structure catalyze polymerization reactions that form coke. The low temperature coke, formed at a reaction temperature of 20~ is expected to be promoted by nickel present on the external surface and by the low acidity external sites for Ni/mordenite and also for Ni/ZSM5 (Tables 1 and 3). Table 1. Coke deposition from TGA runs for Ni/mordenite. Temperature, ~

Amount of Coke, m ~ m l ~ ,

Relative % of Total

150-300 300-500 500-790 150-790

3.16 1.60 0.58 5.33

59.14 29.91 10.95

124 Table 2. Coke deposition from TGA runs for Ni/USY. Temperature, ~

Amount of Coke, mg/mg~.,

Relative % of Total

150-300 300-500 500-790 150-790

2.90 1.55 9.17 13.61

21.27 11.36 67.37

Table 3. Coke deposition from TGA runs for Ni/ZSM5. Temperature, ~

Amount of Coke, m ~ / m ~

Relative % of Total

150-300 300-500 500-790 150-790

1.21 1.27 0.33 3.02

40.07 42.04 17.89

1.00

0.80 O

-

0.60 0.40 0.20 -

0.00 0

I

I

I

I

I

I

200

400

600

800

1000

1200

Temperature, ~ Figure 5. Temperature-Programmed Reduction pattern for Ni/Mordenite.

1400

125 4. C O N C L U S I O N S Effects of crystalline structure and acidity differentiate the catalytic behavior of ZSM5, USY and mordenite zeolites. Compounded with the nature of the support, the location of nickel particles leads to very peculiar behaviors in the formation of low-temperature coke during the hydrogenation of phenylacetylene. The principal differences in the high temperature deactivation are determined by the size specificity of the zeolitic supports, and by the high acidity available to the reactant molecules, especially for the USY support. The contribution of nickel to coke formation at low temperatures occurs mainly at the internal surface of Ni/mordenite and Ni/USY and at the external surface of Ni/ZSM-5. This conclusion is supported by the TPR patterns as well as by the relative values of low, intermediate and high temperature coke for each individual support.

REFERENCES

1.

J.A. Montoya, J.C. Rodriguez, I. Schifter, A. Monzon and T. Viveros, in B. Delmon and G.F. Froment, "Catalyst Deactivation", Studies in Surface Science and Catalysis, V88, pp. 531-536, Elsevier Science, B.V., 1994. 2. P.A. Sermon, M.S.W. Vong and M. Matheson, International Symposium on Deactivation and Testing of Hydrocarbon Conversion Catalysts, pp. 370-373, Div. of Petroleum Chem., Inc., 210 t" National Meeting, ACS, Chicago, IL, August 20-25, 1995. 3. R.J. Argauer and G.R. Landolt, U.S. Patent 3702886, assigned to Mobil Oil Corp. (1972). T. Masuda and K. Hashimoto, ibid pp. 359-364. 5. B.C. Gates, Catalytic Chemistry, John Wiley and Sons, Inc. 1982 6. R. Szostak, Handbook of Molecular Sieves, Van Nordstrand Reinhold, NY (1992) 7. E.D. Gamas-Castellanos, Ph.D. Dissertation, Dept. of Chemical Engineering, University of Houston, May 1996. 8. N.I. Jaeger, P.Ryder and G. Schulz-Ekloff, in P.A. Jacobs et al. (Editors) Structure and Reactivity by Zeolites, p. 299, Elsevier Scientific Pub. B.V. (1984). 9. D. Delafosse, in B. Imelik et al. (Editors) Catalysis by Zeolites, p. 235, Elsevier Scientific Pub. B.V. (1980). 10. G.K. Boreskov and Kh. Minachev, Applications of Zeolites in Catalysis, AKADEMIAI KIADO, Budapest, 1979. .

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Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

127

A c t i v a t i o n and D e a c t i v a t i o n o f the Zeolite Ferrierite for Olefin C o n v e r s i o n s

Krijn P. de Jong, Henk H. Mooiweer, John G. Buglass and Paul K. Maarsen Shell International Oil Products B.V., Shell Research and Technology Centre, Amsterdam P.O. Box 38000, Amsterdam, The Netherlands E-mail: [email protected]

Conversion of olefins over the zeolite Ferrierite (FER) has been investigated in the temperature range 300 to 500~ Fast irreversible adsorption of butenes at 350~ completely filling the micropores, is observed using a microbalance. In the skeletal isomerization of butenes at that temperature, rapid initial coke build up is observed which slows down after 100 hours on stream. In the same time period the yield of isobutene produced from 1-butene increases while by-product formation, i.e. oligomers, is reduced. Fresh FER leads to extensive scrambling of a 13C-label during butene isomerization as deduced from GC/MS experiments, whereas spent FER does not give rise to scrambling. Carbonaceous deposits play an important role in modeling for activation/deactivation of FER for olefin conversion. These deposits are initially involved in oligomerization and cracking reactions and their aromatisation leads to coverage of the external surface of the crystallites m as deduced from XPS and TEM data m thus blocking pores and leading to a slow overall deactivation. The stability of FER in olefin processing is further demonstrated for olefin aromatisation at 500~ 1. INTRODUCTION Hydrocarbon conversions over zeolite catalysts play a major role in oil and petrochemical processes. Catalytic cracking and hydrocracking using FAU (zeolite Y) based catalysts were introduced in the 1960's, paraffin isomerization based on MOR around 1970, while shape selective conversions where MFI has an important role evolved around 1980 [1]. In these processes, paraffin and aromatics conversion are predominant. Olefin conversions using zeolites are more demanding since severe catalyst deactivation can occur due to rapid coke formation. With the advent of reformulated gasoline, however, the necessity to convert olefins is growing [ 1-3]. Prime examples are (i) skeletal isomerization of normal olefins to iso-olefins as feedstock for ether manufacture, (ii) alkylation of aromatics with light olefins e.g. to convert benzene, (iii) paraffin alkylation with olefins, (iv) aromatisation of gasoline olefins to enhance octane

128 quality. Low temperature (100--250~ conversion of olefins, notably oligomerization and alkylation, is made possible by involving cracking reactions to clean up the pore system [4] or to operate at very low olefin concentrations to prevent formation of deposits [5], respectively. Processing of olefins in the temperature range of say 300-500~ at high partial pressures of olefins is even more difficult. Skeletal isomerization of olefins and aromatisation of olefins are examples of desired reactions in this temperature range. Following the work of Grandvallet et al. [6] the zeolite FER has attracted much attention for the conversion of normal butenes to isobutene. Reviews in this area have underlined that FER displays unique selectivities to isobutene [7,8]. Studies of a broad range of zeolites with differences in acidity and pore structure have revealed that for this reaction shape selectivity prevails in enhancing isomerization over oligomerization [8-11 ]. The pore system of FER is unique in limiting oligomerization and aromatisation reactions while allowing isomerization reactions. The characteristics of the FER catalyst are very important however, of which acidity is one. Work by Xu et al. [12] and Pellet et al. [ 13] has indicated that a high Si/A1 ratio, the absence of extra-framework aluminum and a high crystallinity are all beneficial for enhancing selectivity to isobutene. Recent work has elucidated that the key to the isomerization of butenes over FER may lie in the subtle effects of carbonaceous deposits inside the unique pore system of the zeolite. It has been shown by Mooiweer et al. [ 14] that the pores of FER are largely filled with carbonaceous deposits under reaction conditions. An extensive study by Xu et al. [15] has shown that at 420~ olefinic deposits are converted to aromatic and polyaromatic material in FER. Seo et al. [16] have elegantly shown that coke deposits lead to higher selectivities for butene isomerization over oligomerization. Guisnet et al. [17,18] have made an intriguing proposal that coke deposits facilitate isomerization of normal butene by opening up a reaction path that circumvents primary carbenium ions. The deposits, which are either of an aliphatic [17] or of an aromatic nature [18], give rise to tertiary carbenium ions, thus enabling isomerization of butenes via alkylation, methyl shift and cracking [17]. Build up of coke thus gives rise to higher selectivities to isobutene since initial conversion involve dimerisation and cracking giving rise to products other than isobutene. In this paper we contribute to clarifying the role of coke in the activation/deactivation of FER for olefin conversions with emphasis on butene isomerization. To this end we establish the nature of the coke deposits by adsorption/desorption measurements. Furthermore, the isotopic scrambling for butene isomerization is determined in support of the Guisnet mechanism. Finally, slow deactivation of FER by coke is studied as this limits the ultimate lifetime of the catalysts. Details of the coke deposits in spent catalysts are revealed using electron microscopy and X-ray excited Photo-electron Spectroscopy. 2. EXPERIMENTAL Ferrierite zeolite samples in their acidic form were prepared in-house or obtained from a commercial supplier. The properties of the base materials are summarized in Table 1. For testing purposes, the zeolite powders were shaped into small particles

129 (30-80 mesh). Catalytic tests were carried out using 5 grams of catalyst diluted with an equal volume of SiC particles in a fixed-bed down-flow gas phase reactor. The off-gas composition was determined by gas chromatography (GC). For butene isomerization experimems 1-butene (99.5% pure) was used as feedstock without dilution at a total pressure of 1.2 bar. The results reported here have been obtained at 350~ and a weight hourly space velocity of 2 kg(butenes)/(kg of catalysVhr). Table 1 Basic data of Ferrierite samples Catalyst code

Si/A1 ratio (at/at)

Source

MPV

MPV

LS-FER

9

Toyo Soda

n.d.

n.d.

HS-FER

57

In house synthesis

0.13

1.4

Notes: n.d. = not determined

MPV is Micropore Volume determined with liquid nitrogen assuming a density of 0.808 g/ml. For olefin aromatisation experiments we used an FCC gasoline (22 %w aromatics and 28 %w olefins) with a boiling range of 85-175~ at process conditions of 500~ and 20 bar total pressure. For the latter reaction the feed and product were analyzed by GC and PIONA-GC (liquid product, TLP). Adsorption and desorption studies were carried out using HS-FER in a static atmosphere using an automated vacuum balance system. The pressure was increased stepwise by dosing with pure 1-butene. A subsequem dose was admitted when the pressure was stable, typically after 1 hour. Similarly, desorption was carried out stepwise. A mechanistic study of butene isomerization was carried out using 13C labelled butene, 13CH2=CH-CH2-CH3 and a GC-MS set up. The GC injector was equipped with a narrow glass tube containing 20 mg of powdered FER catalyst. In a typical experiment a butene sample (1 ktl, 0.9 bar) is injected, the products are separated by GC and then individually analyzed by mass spectrometry (MS) (Kratos Concept). Fresh and spent (but still active for butene isomerization) HS-FER was used for these experiments. Experiments were carried out at a catalyst temperature of 350~ Following extraction with toluene and pentane, the amount of aromatic coke has been determined by Pyrolysis Combustion Mass Spectrometric Element analysis (PCME). In the PCME technique the catalyst was first heated to pyrolyse the coke followed by combustion and determination of the remaining aromatic coke. The structure of the total

130 coke deposits after extraction was studied using X-ray excited Photo-electron Spectroscopy (XPS) with a Kratos XSAM 800 instrument. C/Si signal intensity ratios were thus obtained. Various Electron Microscopy (EM) techniques were used. High resolution Scanning Electron Microscopy (HRSEM) was performed on a Jeol 6000F instrument, operated at 25 kV. High resolution microstructural imaging and Selected Area Diffraction (SAD) were carded out using a Philips EM400T Transmission Electron Microscopy (TEM) instrument operated at 100 kV. Finally, nanoanalysis of surface coke deposits was performed using a VG HB501 Dedicated Scanning Transmission Electron Microscopy (D-STEM) instrument operated at 100 kV. The latter microscope was equipped with a Parallel Electron Energy Loss Spectroscopy (PEELS) facility which permits the efficient detection of carbon and identification of its form, e.g. amorphous or graphitic. % 1OO

80

\ n n i B m ~ ~ n - C4= c o n v e r s i o n (~

60

40

r

20

0

i-C4=yield (~ wof)

0

~JUi~lm~muW m I 200

n 400

C5§ y i e l d (*/owof)

I

I

600

800

T i m e on s t r e a m

I

1000

I

1200

(h)

Figure 1. Skeletal isomerization of 1-butene using HS-FER as catalyst. Product yields and n-butene conversion vs Time on-stream. 3. RESULTS

3.1. Catalytic experiments The stability of FER catalysts for butene isomerization has been studied extensively. In Fig. 1 a typical example of the long term activity and selectivity development of HS-FER is shown. During the first 100 h, a fast drop in the n-butenes conversion is observed which is related to the strong drop in the production of oligomers (C5+),

131 Yield

100

(~ /r

',

90

0 .Gr'" //A

~,

80

Cs+-yield, MFI A

Cs-,--yield , FER

z/

A 70

\

0 _~ 0

60

,

Aromatics,FER '1~, "\

50

0 x

0

0

O \

~'. x x .-~x

40 30

Aromatics, MFI 0

,

I

20

,

I

40

,

I

60

,

I

80

,

I

100

,

I

120

,

140

Time on s t r e a m (h)

Figure 2. FCC gasoline aromatization stability of ferrierite ans ZSM-5.

Butenes 100 Adsorpt ion (mg/g FER)

DESORPTION

80

60

40-

20-

2 0

1

0

200

400

,

I

600

Butenes pressure (mm Hg)

Figure 3. Vacuum balance adsorption/desorption of 1-butene on HS-FER at 350~

800

132 mainly consisting of dimers. Note that the yield of isobutene increases as has been noted before [12,13,15]. Beyond 100 hours on stream, a slow decrease in overall activity is observed but still the isobutene yield increases while the C 5+ yield continues to decline. The yield of isobutene reaches a maximum around 800 hours whereafter it starts to decline slowly. Following this long catalytic test (1280 hours) the carbon content of the catalyst was determined by PCME amounts to be 9.1%w. Aromatisation of olefins as present in FCC gasoline was carried out over both HS-FER and MFI (Si/A1 = 140 at/at). The results with FER (Fig. 2) show a constant production of aromatics at a relatively stable liquid (C5+) yield. For MFI the aromatisation activity initially surpasses that of FER but a strong decline is observed with the former catalyst.

Label

distribution

100

0

80

(*/,)

no label, mass 5 6

[]

1 label, m a s s 57

0

2 label,mass 58

I

3 label ,mass 59

~

Aromatic 1

....... iiiiiiiiiiiil !:!:i:i:i:i:i ii!iiiiiiiiii .............

4o 20

.......

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...........! i

~i~i~i~!~i~i~ i i lili i i

,

........

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o

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iiiiiii i i

0

............ i

....

9

o

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. . . .

,

10100' Run t i m e (h)

Figure 4. 13C-label distribution for iso-

Figure 5. PCME carbon content of

butene as obtained from 13C-labelled 1butene with a fresh and a spent high Si FER catalyst.

spent FER vs time on stream.

3.2. Adsorption/desorption experiments Experiments in the temperature range 50-350~ on HS-FER showed rapid adsorption of 1-butene whereas desorption takes place to a very limited extent only. A typical plot of an adsorption followed by desorption experiment is shown in Fig. 3. Data were obtained at 350~ which is the reaction temperature of the catalytic

133

experiments (Fig. 1). Clearly, the adsorption of butenes on FER is irreversible in the temperature range studied.

3.3. Isomerization experiments with labelled 1-butene The mass spectrum of the GC/MS experiment with an empty reactor provided us with the fragmentation pattern of 13C-labelled 1-butene. Subsequently, isomerization experiments were carried out using fresh and spent HS-FER catalysts. The MS spectra corrected for the fragmentation pattern of the butene in question provided us with the label distribution in butenes. In Fig. 4 the results of the label distribution for isobutene, as obtained from isomerization at 350~ over the catalysts, are displayed. The scrambling of 13C over isobutene as observed with the fresh HS-FER after correction for natural abundance of 13C, is close to the values expected from complete (statistical) scrambling. Clearly, with the spent Ferrierite catalyst, which is still active for butene isomerization, scrambling has not taken place either in isobutene, or in the other butenes. Very similar results were obtained using 13C-labelled isobutene instead of

1-butene.

2.5

XPS C/Si r a t i o ( a t / a t )

/ surface coking

10

/I/ /

1.5

00~ 9 9

....//

9

0.5 ....

0

i

I

2

,

channel

4

/' 9 // / /

9

0

Figure 6. Variation of the XPS C/Si ratio as a function of carbon content (PCME) in various spent FER catalysts. Theoretical curves for channel filling and surface coking are included.

/

filling ~

i

6

a

I

8 Aromatic

i

i

10

i

12

carbon (*/.wt)

3.4. Characterization of carbonaceous deposits The build up of carbon deposits as a function of time on stream is shown in Fig. 5. The data were obtained at 350~ and WHSV of 2 kg/(kg 9h) for both LS-FER and HS-FER. Ferrierite having higher aluminum content is more prone to coke deposition in line with a faster deactivation (data not shown). The PCME data indicate that the H/C ratio of the aromatic coke decreased with time on stream (from 0.4 to 0.2 at/at).

134 For a variety of spent FER catalysts the C/Si signal intensity ratio was determined by XPS (Fig. 6). The experimental data points were compared with predicted values (dotted lines) obtained from the theory developed by Sexton et al. [19]. This theory gives a description of the attenuation of the Si signal due to carbon deposition inside the pores (channel filling, weak attenuation) followed by carbon built up at the external surface of the zeolites (surface coking, strong attenuation). For all coke contents the XPS C/Si intensity ratio is above the line expected for channel filling only. Above a carbon content of 6 %w, surface coking predominates. Electron microscopy was used to provide more details on the nature of the coke deposits as well as their location. HRSEM reveals the HS-FER crystallites to be platelets, typically 250 nm in size, with a mean thickness of only 20 nm. For a spent HS-FER containing 9.1%w of coke TEM photographs are displayed in Fig. 7. Careful inspection of the FER platelets reveals that they are bordered by a layer of amorphous material (labelled) which is probably coke. Also at the edge of the platelets there are layers of similar width (picture corresponds to a view down the (100) surfaces as determined by a selected area diffraction study). This suggests that all the external surfaces, both porous and non-porous, contain a layer of coke. To confirm that the amorphous layers indeed consist of carbonaceous material, PEELS spectra were collected which closely resemble that of the amorphous carbon film supporting the sample crystallites. With samples containing less carbon (below 7 %w), TEM does not reveal a coke layer at the surface of the crystallites, but rather reveals discrete coke clusters of 1.0-1.5 nm. 4. DISCUSSION From the adsorption/desorption experiments (Fig. 3) it appears that butene adsorbs irreversibly on FER in the relevant temperature range for butene isomerization. Since the heat of adsorption of butenes in silicon-rich FER is estimated to be 10 kcal/mol, desorption should be facilitated at about 350~ The low rate of desorption at 350~ proves, therefore, that butenes have reacted to higher molecular weight material, presumably oligomers, inside the pores of FER. From the results shown in Fig. 3 it is estimated that at 350~ about 90 mg of hydrocarbons per gram of FER core are adsorbed. Taking the micropore volume of FER as 0.13 ml/g (Table 1), the apparent density of the hydrocarbons is 0.69 g/ml, assuming complete filling of the micropores. Note that the density of liquid n-butenes is 0.60 ml/g, while that of methylheptenes is 0.71 g/ml [20]. In view of the irreversibility of the adsorption of butenes in combination with the apparent density, it is strongly suggested that butene oligomers are formed inside the pore entrance. In contrast to this model, Xu et al. [ 15] have suggested that the apparent high density of butenes adsorbed in FER is

135

~ - . ~ , . i

" ~i~r ~.~;..ii ~~ ~ ~ ~ . ~ ~.~

9

9

~

";

o

~

i.

o~. ,(

Figure 7. TEM images of spent HS-FER (9.1 wt% carbon). (A) Close to <100> and (B) imaged edge-on (the external coke layers on (100) surfaces are labelled).

136 due to butenes entering the 8-membered pore, whereas nitrogen m used to determine the micropore volume m cannot enter these pores. Given, however, the similar molecular dimensions of butene and nitrogen, this seems highly unlikely. Moreover, the pore volume determined by N2 is in good agreement with the theoretical pore volume on basis of the pore structure (0.15 ml/g). In summation, the 10- and 8-membered pores must be (almost) completely filled with butene oligomers following adsorption of the monomer. Although aliphatic coke (butene oligomers) fills up the pores immediately, the buildup of aromatic coke occurs much more slowly (Fig. 5). Combining the results of Figs. 3 and 5 reveals that from a pool of aliphatic material a slow transformation to aromatic coke takes place. This transformation is relatively fast during the first 100 hours, slowing down thereafter. GC/MS experiments using labelled butene show that extensive scrambling occurs with a fresh FER catalyst (Fig. 4). Since the adsorption measurements show that pores fill with oligomers on a short time scale relative to the time scale of the isomerization experiments, it is concluded that oligomerization is accompanied by cracking at this initial stage. This explains the scrambling: oligomerization, methyl shifts and cracking will give rise to a statistical distribution of the labelled carbon over the product molecules. In the case of spent FER, however, scrambling is no longer observed (Fig. 4). Apparently, the isomerization reaction takes place without the possibility of labelled carbons being exchanged between molecules. In our view two aspects play a role, viz. the absence of strong acid sites and the occurrence of 'fixed' coke after a longer time on stream. The strong acid sites initially catalyze cracking of oligomers, which 'mobilize' coke deposits. Slow aromatisation of coke inhibits the possibility of cracking as a mechanism for coke removal. It is essential to recognize that the initial activation of FER for butene isomerization is accompanied by a drop in C 5+ production and is thus a more 'label selective' mechanism. During the initial period larger molecules can escape from the pores of FER, thus giving rise to large amounts of C5+ products. Later, the 'fixed' coke restricts escape of larger molecules from the pores. The slow deactivation of FER after 100 h on stream (Fig. 1) is now addressed. In view of the above results, we propose that slow aromatisation of the carbonaceous deposits takes place both inside the pores and at the external surface of the crystallites. The density of the aromatic coke is higher than that of the initially formed aliphatic coke, thus allowing slightly more hydrocarbons to enter the pores. The pores would then fill until butenes can no longer enter them to isomerize. Ultimately, coke will block internal and external surfaces, the first signs of which appear in Figs. 6 (above 6 %w coke) and 7. For the catalyst containing 9.1%w coke (Fig. 7) estimates from XPS and TEM indicate that the coverage of the FER crystals by coke is around 50% with a layer thickness of 1-2 nm.

137 The key to the success of FER in olefin conversion resides with the fact that carbonaceous deposits quickly form without blocking the catalytic action. According to Guisnet et al. [17] the deposits in fact assist the catalysis. Aromatisation of the coke being slow is crucial to preventing ultimate pore blocking. Even under conditions where aromatisation of olefins occurs quickly (Fig. 2) it is Clear that FER is stable. Under these conditions of high temperature, restricting the conversion of mono-aromatics to polyaromatics is most likely the key to enhancing catalyst stability. ACKNOWLEDGMENTS

The authors express their gratitude to Dr. Herman P.C.E. Kuipers for stimulating discussions and to Dr. Andreas K. Nowak for critically reviewing the manuscript. REFERENCES

1.

2. 3. 4. 5. 6. .

8.

9. 10.

11. 12. 13. 14.

J.E. Naber, K.P. de Jong, W.H.J. Stork, H.P.C.E. Kuipers and M.F.M. Post, in: J. Weitkamp, H.G. Karge, H. Pfeifer and W. Holderich (eds.), 'Zeolites and Related Microporous Materials: State of the Art 1994', Elsevier, Amsterdam, p. 2197. I.E. Maxwell and J.E. Naber, Catalysis Letters, 12 (1992) 105. K.P. de Jong, W. Bosch and T.D.B. Morgan, in: A. Frennet and J.-M. Bastin (eds.), 'Catalysis and Automotive Pollution Control III', Elsevier, Amsterdam, 1995, p. 15. N.Y. Chen, W.E. Garwood and F.G. Dwyer, 'Shape selective catalysis in industrial applications', Marcel Dekker, New York, 1989, p. 67. K.P. de Jong, C.M.A.M. Mesters, D.G.R. Peferoen, P.T.M. van Brugge and C. de Groot, Chem. Eng. Sci., 51 (1996) 2053. P. Grandvallet, K.P. de Jong, H.H. Mooiweer, A.G.T.G. Kortbeek and B. Kraushaar-Czarnetzki, European Patent 501,577 (1992) to Shell. A.C. Butter and C.P. Nicolaides, Catal. Today, 18,(1993)443. J. Houzvicka and V. Ponec, Catal. Rev., Sci.-Eng., submitted. J. Houzvicka and V. Ponec, J. Catal., in press. P. Meriaudeau, T. Vu Anh, H. Le Van and C. Naccache, in: H.Chon, S.-K. Ihm and Y.S. Uh (eds.), 'Progress in Zeolite and Microporous Materials', Elsevier (1997), Amsterdam, p. 1373. G. Seo and H.S. Jeong, Catal. Lett., 36 (1996) 249. W.-Q. Xu, Y.-G. Yin, S.L. Suib, J.C. Edwards and C.-L. O'Young, J. Phys. Chem., 99 (1995) 9443. R.J. Pellet, D.G. Casey, H.-M. Huang, R.V. Kessler, E.J. Kuhlman, C.-L. O'Young, R.A. Sawicki and J.R. Ugolini, J. Catal., 157 (1995)423. H.H. Mooiweer, K.P. de Jong, B. Kraushaar-Czarnetzki, W.H.J. Stork and B.C.H.

138

15. 16. 17. 18.

19. 20.

Krutzen, in: J. Weitkamp, H.G. Karge, H. Pfeifer and W. Holderich (eds.), 'Zeolites and Related Microporous Materials: State of the Art 1994', Elsevier, Amsterdam, p. 2327. W-Q. Xu, Y.-G. Yin, S.L. Suib and C.-L. O'Young, J. Phys. Chem., 99 (1995) 758. G. Seo, H.S. Jeong, D.-L. Jang, D.L. Cho and S.B. Hong, Catal. Lea., 41 (1996) 189. M. Guisnet, P. Andy, N.S. Bnep, C. Travers and E. Benazzi, J. Chem. Soc., Chem. Commun. (1995) 1685. M. Guisnet, P. Andy, N.S. Bnep, C. Travers and E. Benazzi, in: H.Chon, S.-K. Ihm and Y.S. Uh (eds.), 'Progress in Zeolite and Microporous Materials', Elsevier (1997), Amsterdam, p. 1365. B.A. Sexton, A.E. Hughes and D.M. Bibby, J. Catal., 109 (1988) 126. CRC Handbook of Physics and Chemistry, 73rd Edition.

~ Science B.V. All rights reserved Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

139

Deactivation o f Ferrierite during the 1-Butene Skeletal Isomerization R.A. Comelli, Z.R. Finelli, N.S. Figoli, C.A. Querini Instituto de Investigaciones en Cat~ilisis y Petroquimica- INCAPE (FIQ-LrNL, CONICET). Santiago del Estero 2654. 3000 - Santa Fe, Argentina The influence of temperature, 1-butene partial pressure, weight hourly space velocity (WHSV), and pretreatment conditions on the activity, selectivity and stability of ferrierite during the 1-butene skeletal isomerization has been studied in order to understand the process of coke formation. Coke is characterized by temperature-programmed oxidation (TPO) and Fourier transform-infrared diffuse reflectance spectroscopy (DRIFTS). The catalytic behavior is strongly affected by the intrinsic properties of the zeolite, showing a significant deactivation associated with coke formation. The operational conditions, such as temperature, 1-butene partial pressure and WHSV, play an important role in the amount of coke produced, the temperature being the most important factor affecting the amount of coke. These conditions also modify the nature of the coke. A significant fraction of coke burns at temperaturehigher than 600~ 1. INTRODUCTION The skeletal isomerization of n-butenes is an attractive route for the production of isobutene which is an important raw material for the synthesis of methyl tert-butyl ether (MTBE), currently used as a major booster of the octane number for reformulated gasoline. Amorphous materials such as halogen-modified aluminas are active in the n-butene skeletal isomerization (1). Environmental and stability considerations do not favor the usage of these catalysts. Zeolites also might catalyze this isomerization if both acidity and pore size are suitable to achieve desirable activity and selectivity (2-4). Ferrierite, a small pore zeolite, is of current interest since it shows exceptional selectivity for the 1-butene skeletal isomerization (5). The ferrierite structure (6) has a two-dimensional pore system with tenmembered tings (4.2 x 5.4 A) intersected by eight-membered rings (3.5 x 4.8 A). This pore structure is related to the high isobutene selectivity and is more important than the acidity, as has been recently reported (7). The ferrierite is one of the best catalysts known for this reaction, due to a very good selectivity. However, initial selectivity to isobutene is low (8), and the catalyst displays better performance alter some time-on-stream when coke has been deposited. This behavior has been related to the presence of two kinds of active sites (9), the external acid sites which are non-selective sites for skeletal isomerization and the internal acid sites selective for this reaction. A monomolecular mechanism is reported over the latter sites, while a bimolecular mechanism occurs over the non-selective acid sites (10). The deactivation and coke formation over ferfierite was recently related to the crystallinity degree of the zeolite (9). In this study, coke formation on ferrierite during the 1-butene skeletal isomerization was investigated. The influence of temperature, 1-butene partial pressure, WHSV, and pretreatment conditions on the activity, selectivity and stability was analyzed in order to understand the process of coke formation. The coke was characterized by TPO and DRIFTS. 2. EXPERIMENTAL Ammonium ferrierite was provided by TOSOH, Japan (sample number HSZ-720NHA, Lot N ~ D 1-51). The SIO2/A1203 molar ratio was 17.8, with Na20 < 0.05% and K20 < 0.1%.

140 The crystalline structure was characterized by X-ray diffraction (XRD) using a RichSeifert Iso-Debyeflex 2002 diffractometer. Radiation was Cu K(~, filtered with Ni; the diffraction spectrum range was 0 < 2( < 60 ~ The catalytic behavior during the 1-butene skeletal isomerization reaction was studied using a continuous flow fixed-bed quartz tubular reactor operated at atmospheric pressure. For the test, 500 mg of catalyst sieved to 35-80 mesh were heated in a nitrogen stream (60 ml min -~) from 25~ to 550~ in 30 min and held at this value for 30 min. Some samples were pretreated in nitrogen saturated with water or in hydrogen streams. After heating, the samples were cooled to reaction temperature; then a nitrogen stream and pure 1-butene were fed for 180 min. The reactants and reaction products were analyzed by on-line gas chromatography, using a 30-m-long, 0.54-mm-O.D. GS-Alumina (J&W) megabore column, operated isothermally at 100~ From these data, catalytic activity, selectivity to isobutene, and isobutene yield, were calculated on a carbon basis. The catalytic activity is expressed as n-butene conversion. For this calculation the three isomers of n-butene are grouped together; this consideration is based on the fact that the isomerization of 1-butene to 2-butene rapidly reaches the equilibrium via double-bond migration under reaction conditions. TPO analysis of used catalysts was performed using specially designed equipment (11) in order to improve both sensitivity and resolution. The unit converted the carbon dioxide generated during coke combustion to methane, using a nickel catalyst. The methane was then continuously analyzed with a flame ionization detector. The experiments were performed using a 6% oxygen in a nitrogen stream (20 ml min-~), heating the sample at 12~ min -~. The sample weight was about 0.01 g. In order to obtain additional characterization of the coke deposited on the catalyst, DRIFTS measurements were performed using a SHIMADZU 8101M spectrometer. Coked samples were diluted to 5% in KBr. 3. RESULTS AND DISCUSSION The catalytic performance of ferrierite during the 1-butene isomerization at reaction temperatures between 300 and 430~ was determined. Figure 1-A displays the catalytic activity (expressed as n-butene conversion) and the selectivity to isobutene. A similar qualitative behavior is observed at any temperature: a high activity with low isobutene selectivity is reached after a short time on-stream, and then a rapid decrease in activity takes place, with an increase in the isobutene selectivity. By increasing temperature, catalytic activity increases and stability improves, while the selectivity to isobutene decreases. Figure 1-B shows the isobutene yield at several temperatures. The higher the reaction temperature, the larger the isobutene yield and the higher the stability. The effect of the 1-butene partial pressure over the catalytic behavior of ferrierite at 350~ can be seen in Figures 2-A and 2-B. These experiments were carried out at atmospheric pressure, and the 1-butene partial pressure was modified by changing the nitrogen flow rate. The same qualitative behavior that was observed when changing the temperature is observed when changing the 1-butene partial pressure. By decreasing the 1butene partial pressure, catalytic activity increases and stability improves, while the selectivity to isobutene decreases slightly, as shown in Figure 2-A. The lower the 1-butene partial pressure, the larger the isobutene yield and yield stability (see Fig. 2-B). The effect of the 1-butene partial pressure on the catalytic performance is less noticeable than the temperature effect. The effect of the WHSV was also studied by modifying the 1-butene flow rate and maintaining at the same 1-butene partial pressure. These results, shown in Figures 3-A and 3B, are similar to those obtained by changing the 1-butene partial pressure. At low 1-butene space velocities, higher conversion and yield are obtained with lower isobutene selectivity. At all operational conditions a high n-butene conversion and poor isobutene selectivity was observed at short time on-stream, followed by a decrease in conversion and an increase in the isobutene selectivity with increasing time on-stream. This behavior is similar to data previously reported (8-10). Our results show a decrease in the isobutene yield at short times

141 on-stream, as also found previously (9) when the silicon to aluminum ratio is low. The morphology, crystallinity and pore size distributions of ferrierite change when the silicon to aluminum ratio decreases (9). Comparing the X-ray diff action pattern of our sample with those showed by Xu et al. (9), it was found that the ferrierite used in our study was similar to their sample having the lowest crystallinity, although the silicon to aluminum ratio of both samples were different. Therefore, the initial decrease in both activity and isobutene yield, shown in Figures 1-3, can be explained because of the low crystallinity of our sample. .<

30

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0

0

40

80

120

160

0

200

0

40

80

120

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TIME, min

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Figure 1. (A) n-Butene conversion (solid lines) and selectivity to isobutene (dashed lines) as a function of time on-stream at different temperatures. (B) isobutene yield at different temperatures. Pica= = 0.14 atm; WHSV = 0.184 h -]. IO0

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Figure 2" (A) n-Butene conversion (solid lines) and selectivity to isobutene (dashed lines) as a function of time on-stream at different 1-butene partial pressure. (B) isobutene yield at varied 1-butene partial pressure. T = 350~ WHSV =0.184 h l. Moreover, it has been reported that the presence of Na ions in the zeolite structure is detrimental to isomerization reactions since they contribute to channel blocking, limiting the diffusion ofbutene molecules, while coke buildup causes the isobutene yield to decrease with

142 time on-stream (9). Deactivation due to channel blocking by sodium cannot be neglected in our case because sodium was detected by elemental analysis. However, since the sodium content in our sample is rather low, it can be expected that this effect may not be very important. The increase in isobutene selectivity with time-on-stream is a particular property of the ferrierite. This zeolite has two types of active sites: the external sites (on the external surface of the zeolite crystallites) which are non-selective for skeletal isomerization and the internal sites (inside the zeolite pores) which are selective for this reaction (9). The changes observed on the selectivity have been associated with modifications of pore shapes through coke deposition that favor reactions involving small molecules, such as n-butene to isobutene isomerization (8). More recently, it has been reported that a bimolecular mechanism takes place at the non-selective acid sites, while a monomolecular mechanism occurs on the selective sites (10), the coke deposition being necessary in order to poison, block, and modify the non-shape selective acid sites. The catalytic performance of ferrierite during 1-butene skeletal isomerization is essentially unchanged as pretreatment conditions are varied. The initial heating step was made in nitrogen, nitrogen saturated with water, or hydrogen streams; a slightly lower n-butene conversion and isobutene yield was only observed when hydrogen was used.

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Figure 3. (A) n-Butene conversion (solid lines) and selectivity to isobutene (dashed lines) as a function of time-on-stream at different WHSV. (B) isobutene yield as a function of time onstream at different WHSV. T = 350~ Plc4= = 0.14 atm The amount of coke deposited on the ferrierite at the end of the catalytic test reaches levels between 5.7 and 9.1% for different temperatures, 1-butene partial pressures, and WHSV, as can be seen in Figures 4A-C. Upon increasing the temperature, the amount of coke reaches a maximum at 380~ Increasing the WHSV increases the amount of coke deposited. Lowering the 1-butene partial pressure, decreases the amount of coke formed. The high amount of coke formed under the different operational conditions studied could be related with the capacity of the ferrierite sample to form coke on the external acid sites where dimerization followed by cracking can take place. According to these results, coke is formed by a parallel mechanism, i.e., coke is formed from the 1-butene. The relationship between the amount of coke and both the space velocity (WHSV) and the 1-butene partial pressure points towards such a coking mechanism. When the space velocity is increased, the n-butene

143 conversion decreases, and therefore the average n-butene concentration along the reactor is higher. In these experiments, the higher the n-butene concentration, the higher the amount of coke. Similar results are obtained when increasing the partial pressure of 1-butene. Therefore, coke formation correlates with n-butene concentration. This indicates that coke formation follows a parallel mechanism.

10

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t

400

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Figure 4. Coke content on catalysts; (A) as a function of reaction temperature, P1c4= = 0.14 atm; WHSV = 0.184 hl; (B) as a function of 1-butene partial pressure at T=350~ WHSV = 0.184 hl; (C) as a function of WHSV; T = 350~ Pica= = 0.14 atm. The influence of temperature on the amount of coke is quite complex. It involves the competition between n-butene adsorption, coke formation and cracking of adsorbed molecules. The n-butene conversion increases when the temperatm'e is increased in the 300430~ range, but the selectivity decreases because the amount of C1-C3 products is higher. Therefore, one possible explanation for the observed decrease in the amount of coke at 400~ (compared to 380~ is the higher cracking activity. Another possibility is that the coke deposition rate at 400~ is higher than that at 380~ but fast pore mouth plugging occurs and as a consequence, the total amount of coke is lower. However, since the stability is improved at the higher temperature (see Figs. 1A-B), the former explanation seems to be the correct one. This means that the acidity of ferrierite is adequate to crack the adsorbed molecules without a significant oligomerization of such compounds, in the upper range of temperature studied. In fact, it can be expected that at lower reaction temperature and as a consequence of a lower cracking activity, oligomer formation on the outer site where bimolecular mechanism takes place, leads to a fast coke deposition and therefore a pore mouth plugging occurs. This results in a lower stability and lower coke formation at 300 ~ than at 380~ (see Figures 1 and 4). Figure 5 shows the TPO profiles for the catalysts after reaction at different temperatures. These TPO show two well defined peaks. The first peak is between 250~ and 400~ and the second one is at very high temperatures, between 550 and 750~ As reaction temperature increases, the coke that bums at high temperatures increases, while that burning at low temperatures, decreases. The second peak most likely comes from the reorganization of the coke that occurs during the TPO at temperatures higher than the reaction temperature. After the reaction the catalysts are light gray. One experiment was carried out by heating a coked catalyst up to 450~ in a 6% O2/N 2 stream. After this treatment the catalyst was black. This indicates that above the reaction temperature the coke modifies its nature, changing its structure to an aromatic one. The first peak in the TPO can be related to coke with an aliphatic structure, and the second peak to coke that was reorganized into an

144 aromatic structure. The higher the reaction temperature, the higher the proportion of this type of coke. However, as discussed below, coke aromatization also occurs during reaction at high temperature. It is expected that the regeneration of this catalyst is severely complicated by the presence of this coke that bums at high temperatures. In order to characterize the coke deposited on catalysts after reaction, DRIFTS measurements were made on the used catalysts, as well as

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800

TEMPERATURE, *C

Figure 5. TPO profiles of ferrierite samples tested at different reaction temperatm'es. the fresh ferrierite. Little information about coke can be obtained below 1250 cm l , mainly because strong bands characteristic of the ferrierite appear. Other characteristic bands of the fresh catalyst are centered at about 1633, 3649 and 3738 cm "l. The ones above 3500 cm l are associated with different hydroxyl groups. Coked samples obtained either at 300 and 350~ or at different WHSV, showed some differences from the fresh catalyst in the region corresponding to OH vibrations. The band centered at about 3738 cm 1 disappears while the intensity of the band centered at 3649 cm -l decreases upon increasing either WHSV or temperature - conditions under which the amount of coke deposited increases, as shown in Fig. 4. According to data previously reported (12,13), bands centered at about 3650 and 3740 cm" correspond to different hydroxyl stretching vibrations. More recent data (9) suggest that bands at 3742 and 3601 cm" on ferrierite are related to hydroxyl groups of terminal silanol and hydroxyl groups associated with the framework silicon and aluminum ions, respectively. Considering our results, the disappearance of the band at 3738 cm l could be associated with the deactivation of non-selective acid sites which takes place at short time-on-stream, while the band at 3649 cm -l could be related to the selective acid sites and its intensity decreases as the amount of coke increases. Other bands in the DRIFTS spectra centered at about 1387, 1411, 1508, 1616, 2875 and 2930 cm -l allow characterization of coke. The behavior is qualitatively similar when increasing either WHSV or temperature. A slight increase in the intensity of bands at 1387 and 1411 c m l is observed while the band at 1508 crnl is resolved into two bands centered at about 1502 and 1512 cml; the other bands do not show significant changes. Bands at about 1620 cm l are assigned to C=C stretching vibrations of butenes (14, 15). Bands in the 1370-1394 cm l [assigned to CH3 and CH2 bending vibrations (13)] and 2870-2930 cm l [assigned to CH3 and CH2 stretching vibrations (16)] regions correspond to oligomers (15), while a band at 1520 cm" has been assigned to C=C stretching vibration in aromatics (9,14). Considering these assignments, our results for 300 and 350~ and different WHSV show the presence of oligomers on the surface, in accordance with the

145 TPO results, and a low aromaticity of the coke. Furthermore, the increase in the intensity of the band at about 1510 cm l as reaction temperature increases indicates a higher proportion of coke having aromatic nature, in accordance with data previously reported (9). 4. CONCLUSIONS The catalytic behavior of ferrierite during the 1-butene skeletal isomerization is strongly affected by the intrinsic properties of the zeolite, showing a significant deactivation associated with coke formation. The operational conditions (mainly temperature rather than 1-butene partial pressure and WHSV) play an important role in the amount of coke produced. Moreover, these conditions also modify the nature of the coke, i.e., the proportion of coke burning at low and high temperatures. A significant fraction of coke bums at temperatures higher than 600~ suggesting that regeneration should be carefully studied. 5. ACKNOWLEDGMENTS The authors greatly thank JICA (Japan International Cooperation Agency) for their assistance to this project. We acknowledge TOSOH for the provision of a ferrierite sample. REFERENCES

.

.

.

6. ,

8.

9. 10. 11. 12. 13. 14. 15. 16.

A.C. Butler and C.P. Nicolaides, Catal. Today, 18 (1993) 443. S.M. Yang, D.H. Guo, J.S. Lin and G.T. Wang, Zeolites and Related Microporous Materials: State of the Art 1994 (J. Weitkamp, H.G. Karge, H. Pfeifer and W. H/51derich, Eds.), Studies in Surface Science and Catalysis, Vol. 84, p. 1677, Elsevier Science B.V., Amsterdam, 1994. M.W. Simon, W.-Q. Xu, S.L. Suib and Ch.-L. O'Young, Microporous Materials, 2 (1994) 477. L.H. Gielgens, I.H.E. Veenstra, V. Ponec, M.J. Haanepen and J.H.C. van Hooff, Catal. Lett., 32 (1995) 195. H.H. Mooiweer, Eur. Pat. N ~ 0 574 994 A1, (1993). W.M. Meier and D.H. Olson, Atlas of Zeolites Structure Types, 3rd Edition, p. 98, Butterworth-Heinemann, London, 1992. G. Seo, H.S. Jeong, S.B. Hong and Y.S. Uh, Catal. Lea., 36 (1996) 249. W.-Q. Xu, Y.-G. Yin, S.L. Suib and Ch.-L. O'Young, J. Phys. Chem., 99 (1995) 758. W.-Q. Xu, Y.-G. Yin, S.L. Suib, J.C. Edwards and Ch.-L. O'Young, J. Phys. Chem., 99 (1995) 9443. P. Meriaudeau, R. Bacaud, L. Ngoc Hung and Anh. T. Vu, J. Mol. Catal. A: Chemical, 110 (1996) L 177. S.C. Fung and C.A. Querini, J. Catal., 138 (1992) 240. J.P. Blitz and S.M. Augustine, Spectroscopy, 9(8) (1994) 28. J. Datka, Z. Sarbak and R.P. Eischens, J. Catal., 145 (1994) 544. A.K. Ghosh and R.A. Kydd, J. Catal., 100 (1986) 185. C. Flego, Y. Kiricsi, W.O. Parker and M.G. Clerici, Appl. Catal. A: General, 124 (1995) 107. A.G. Gayubo, J.M. Arandes, A.T. Aguayo, M. Olazar and J. Bilbao, Ind. Eng. Chem. Res., 32 (1993) 588.

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Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

147

A l c o h o l D e h y d r a t i o n Reactions as C h e m i c a l P r e c u r s o r s for C o k e F o r m a t i o n and A c i d i t y Probes in T u n g s t a t e d Z i r c o n i a Catalysts Gustavo Larsen a, Edgar Lotero a, Mark Nabity a, Lucia Petkovic a and Carlos A. Querini b aDepartment of Chemical Engineering, University of Nebraska-Lincoln, NE 68588-0126 blNCAPE, Universidad Nac. del Litoral, CC 93, Santa F6 (3000), Argentina The temperature-programmed desorption of low-molecular weight alcohols was studied over tungstated zirconia and acidic zeolites. Tungstated zirconia (WZ) acids have less tendency to deactivate due to carbon deposition and to form oligomeric products. Proton exchange is also less important in WZ. 1. I N T R O D U C T I O N Tungstated zirconia (WZ) catalysts have been proposed as viable candidates for alkane isomerizations, especially those heavier than C4 [ 1]. In this work, alcohol dehydrations were used to rank the acidity of WZ with respect to that of HY zeolite, and to evaluate the ability of these catalysts to resist coking during a reaction that yields an olefin molecule as the primary product. The results from temperature-programmed reaction and infrared spectroscopy studies allowed us to gain some insight into the relative stability of WZ. The higher stability of WZ with respect to sulfated zirconia is an interesting feature [2,3]. A bimolecular mechanism appears to be responsible for the alkane isomerization activity of both WZ and SZ [4,5], but several observations concerning the kinetic behavior of WZ and SZ lead to the belief that carbenium ion-like intermediates in WZ ought to be relatively shortlived [6]. The key issue is that oligomeric species are expected to undergo either B-scission (isomerization) or further condensation to form high molecular weight coke precursors. Thus, the reason why we chose to study a set of alcohol dehydration reactions (with and without deuterium labeling) is that these species are excellent precursors for carbenium-like surface intermediates. In addition, information regarding the acidity of WZ catalysts can also be obtained by such chemical probes. Temperature-programmed reaction and oxidation experiments (TPReac and TPO) and diffuse reflectance infrared spectroscopy (DRIFTS) were used to study the products, and the type and amount of coke formed upon TPReac. 2. E X P E R I M E N T A L

The alcohols used were 1-propanol and tert-butanol (with and without deuteration of the hydroxyl group) from Aldrich | The preparation of WZ (W content: 6 wt%) is described elsewhere [2]. In brief, the catalyst is obtained by aqueous incipient wetness impregnation of hydrous zirconia with a solution containing the appropriate amount of ammonium metatungstate. The sample is subsequently calcined at 1096 K for one hour under flowing air. EXAFS measurements carried out at the National Synchrotron Light Source [7] showed that the WZ sample has an average W-O coordination number and distance of 4.7 and 1.69/~ respectively. The surface area of the calcined material is 36 m2/g. XRD reveals that a small amount of monoclinic zirconia was formed upon calcination. The Y-zeolite catalyst was

148 purchased from Aldrich @in the NH4 + form. The HY sample was prepared by calcination of NH4Y at 873 K for 2 hours in situ, prior to both DRIFTS and TPReac experiments. The WZ sample (150 mg catalyst bed) and the HY reference were calcined in situ prior to the TPReac experiments. The air flow was switched to pure helium and the cell purged for about 15-20 min. prior to injection of 15 l.tl alcohol (liq.) at the reactor inlet at 298 K. An MKS residual gas analyzer was used to follow a total of 16 different m/e values. For simplicity, out of the 16 rn/e values chosen we only show those that gave detectable signals. Our choice of m/e fragments was based on the assumption that the catalysts could only produce water, unreacted alcohol, olefins (both monomer from dehydration and heavier oligomers), aldehydes (for 1-propanol only) and ethers. Representative m/e ratios for each species were then selected from the available literature [8] After alcohol injection the cell was purged for another 15 min. to remove most of the reversibly adsorbed material prior to ramping the temperature at a 3.1 K/min rate under mass-flow controlled He. The final temperature for all TPReac runs was 773 K, but only the useful T-ranges are shown. The DRIFrS studies were carried out at different temperatures on a Nicolet 20 SXB FTIR spectrometer equipped with a commercial DRIFTS catalytic chamber and associated hardware from Spectratech| The DRIFTS cell is a temperature-controlled flow-through reactor. Upon calcination, samples were cooled to room temperature under nitrogen. Subsequently, 25 microliters (liq.) of the chosen alcohol were injected and spectra were collected at 373, 473, 623 and 773 K during the heating cycle. A spectral resolution of 4 cm -1 was used. Temperature-programmed oxidation (TPO) experiments were carried out on samples which had been exposed to the tert-butanol TPReac cycle. In order to measure the low coke content on the catalysts used in this study, a highly sensitive TPO method was employed. Our TPO setup converts the CO2 produced during coke oxidation to CH4, by means of a Ni catalyst. The CH4 is then continuously monitored with a flame ionization detector. This improves both sensitivity and resolution. Details of the technique have been reported earlier [9]. The TPO analyses are done using under 30cc/min of 3% O2/N2 , while ramping the temperature at 12 K/min. Sample weight was variable, depending upon the anticipated carbon content of each sample in order to improve the reliability of the method.

3. RESULTS AND DISCUSSION

. ,...4

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8

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Temperature (K) Fig. 1" TPO of HY and WZ catalysts after t-BuOH TPReac.

Figure 1 shows the TPO results for both HY and WZ upon TPReac with tert-butanol. The blank zirconia support (not shown) gave a total carbon content of 0.25%. Apart from the lower carbon content that may well be due to a lower ability of WZ to retain the alcohol molecules prior to the temperature ramp, perhaps the most salient feature is the larger tendency of the zeolite to form larger relative amounts of apparently "harder" coke. If we arbitrarily assign the low- and hightemperature contributions to the TPO curves to soft and hard coke respectively, it is clear that the zeolite retains a larger fraction of the latter. In fact, a small but significant amount of methanized CO2 was still being detected when the TPO experiment was terminated around 900 K. In order to understand why the zeolite may, under the experimental conditions employed in this work, retain more carbonaceous residues at higher temperatures than WZ, it was necessary to follow a large number of m/e signals during the TPReac

149

experiments to understand the types of olefins and coke precursors formed during the desorption with reaction process. Tert-butanol and 1-propanol were chosen because a marked difference in reactivity between the two alcohols was expected a priori. Transition metal oxides may also catalyze the formation of ketones or aldehydes by oxidative dehydrogenation providing they have a significant number of basic/acid site pairs [10]. With both HY and WZ, the alcohol dehydration to ether and oxidative dehydrogenation pathways were essentially negligible.

d c

b

a

1600 Wavenumbers (cm-1)

1400

1(~oo Wavenumbers (cm-1)

1~,0o

Figure 2. DRIFTS spectra in the coke region Figure 3. DRIFTS spectra in the coke retion of WZ upon tea-butanol exposure, a) 373, on HY upon tert-butanol exposure, a) 373, b) b) 473, c) 623 and d) 773 K. 473, c) 623 and d) 773 K. Figures 2 and 3 show the DRIFTS spectra in the so-called "coke region". The 1580-1600 cm -1 band has been assigned to 8(C-H) vibrations of hydrogen-deficient carbonaceous polyethene- or aromatic-type residues [ 11], and has also been referred to as the "hard coke" band [12,13]. The band around 1540 cm -1 has been assigned to polyphenylene structures, whereas that at 1470 cm -1 is due to paraffmic species [11,14]. The olefinic CC stretching frequency is observed in the 1600-1650 cm -1 range [14]. Interestingly, both catalysts show evidence for both olef'm formation at low temperatures (1600-1650 cm -1 signal) and aliphatic species (1470 cm -1) which progressively yield coke signals at 1580 and 1550 cm -1 for HY and WZ respectively. The DRIFTS results with 1-propanol presented the same trend (not shown), with olefinic CC stretching signals evolving toward broad coke bands around 1590 and 1540 cm -1 for HY and WZ respectively. Clearly, the end result for deposition of carbonaceous residues upon exposure of WZ to alcohols appears to differ from that of HY. Coke residues formed from alcohols on HY are characterized by the observation of hard coke bands very similar to those formed after alkane cracking [ 12,15]. Figures 4 and 5 show the tert-butanol TPReac results over the WZ and HY catalysts respectively. The m/e = 59 is the most important alcohol fragment. The WZ catalyst is characterized by desorption with reaction to olefins almost exclusively (m/e values of 41 and 55). The ratio of masses 41 to 55 is consistent with the isobutene MS pattern throughout the whole tert-butanol TPReac spectrum over WZ. In heavier olefins (e.g., octenes), the m/e = 55 is the dominant fragment, which would have made extensive oligomerization easily detectable. On the other hand, the HY catalyst showed a richer tea-butanol TPReac pattern. Olefin oligomerization over HY was also evident by the asymmetry of the m/e 41 and 55 peaks. Note that if the two signals were only due to isobutene, we should expect them to follow exactly the same trend. However, Fig. 5 shows that m/e = 55 becomes more

150

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./ \ J ......'"

290

310

330

350

Temperature I I~lvin

370

39O

300

400

500 Temperature I Kelvin

600

700

Figure 4. TPReac of tert-butanol over WZ. Figure 5. TPReac of tert-butanol over HY. M/e = 59 (xl0, solid line), 18 (o), 55 (--) and M/e = 69 (x5, fdled triangles), 18 (long 41 (triangles). dashes), 55 (x5, x), 115 (xl0, solid), 41(--), 65 (xl0, o) and 59 ( x 5 , ) . .. ~

]

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Figure 6. TPReac of 1-propanol over WZ. M/e = 18 (solid line), 41 (o), 31 (dashed).

350

400

450 Temperature I Kelvln

500

550

Figure 7. TPReac of 1-propanol over HY. M/e = 41 (long dashes), 31 (10x, short dashes), 55 (xl0, solid), 69 (xl0, triangles) and 18 (squares).

important at higher temperatures, which is in agreement with the idea that heavier unsaturated hydrocarbons are formed at higher temperatures. A number of other heavier olefm masses were followed to conf'Lrm such conclusion. Another interesting feature is the observation of a detectable m/e = 65 signal (cyclopentadienyl cation) in the HY TPReac pattern which suggests that condensed carbon nuclei are formed and even detected in the gas phase. Such an m/e fragment was not observed in WZ. This observation is also consistent with our TPReac results on alcohol dehydration over H-Mordenite [7]. The evolution of the rn/e = 65 signal with temperature in HY coincides with the appearance of the coke band around 1580 cm -1 (see figures 3 and 5). The idea is that both release of oligomers to the gas phase and retention of oligomeric residues (coke precursors, see TPO results above) are less significant in WZ than in acidic zeolites, and that the residue formed at high temperatures in WZ (1550 cm -1 DRIFTS band) is more easily eliminated by TPO. One plausible explanation is that the broad 1550 cm -1 band that appears on WZ upon heating the catalyst with preadsorbed tbutanol arises from many chemically non-equivalent soft coke residues, rather than from polyphenylene structures. Alternatively, if the 1550 cm -1 band in WZ is indeed due to

151

polyphenylene fragments, the latter must be more easily subjected to oxidation than conventional hard coke. Based on an extensive literature search, we found no precedent that this should be the case, although this hypothesis cannot be entirely ruled out at this point. A small m/e = 115 signal was detected on the TPReac of t-butanol over HY which is tentatively assigned to the formation of small amounts of ether. In Figures 6 and 7 we report m/e values detected during TPReac of 1-propanol over WZ and HY respectively. A much larger proportion of the alcohol desorbs without reaction over WZ, due to the lower reactivity of primary alcohols with respect to t-BuOH for dehydration to olefins. No evidence for aldehyde formation (oxidative dehydrogenation) was found for either WZ or HY, consistent with the hypothesis that such a reaction pathway would require the presence of acid-basic site pairs [ 10]. Once again, masses 55 and 69 are more representative of olefin oligomers (Fig. 7). It is interesting to note again the shift toward higher temperatures for olefin oligomeric masses with respect to m/e = 412 (allyl cation, the most important propene and isobutene fragment) during TPReac of the primary alcohol. In general, WZ tends to retain water much less than HTY at higher temperatures. We have previously observed that both 1- and 2propanol release water over WZ at essentially the same time as the m/e = 41 signal [7]. The WZ is a Lewis acid upon calcination at 1096 K. It appears that water of reaction does not hydrate such sites upon alcohol dehydration. Nevertheless, there is still the question of whether some water is sti!l present at the WZ surface and if so, how reactive it would be toward isotopic exchange. 2

20

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~

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400

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450 500 Temperature I Kelvin

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Figure 8. Ratio of masses 41/42 for the tert- Figure 9. Ratio of masses 41/42 for the tertBuOH TPReac experiments 9 Solid: t-BuOH, BuOH TPReac experiments. Solid: t-BuOH dashed: t-BuOD. HY dashed: t-BuOD. WZ Figures 8 and 9 show the ratio of masses 41/42 observed during TPReac of tert-butanol and deuterium-hydroxyl labeled tea-butanol. In principle, if there were no H+/D§ exchange (and if both masses were due to the same product molecule) we should see a constant 41/42 ratio. On the other hand, if the deuterium ion exchanges with surface sites responsible for the dehydration reaction, one should expect incorporation of D atoms into the hydrocarbon skeleton (in this case, into the allyl cation olefinic fragment which is common to both olefin monomers and oligomers). Looking at the solid lines in both Figs. 8 and 9 (unlabeled tertbutanol), it is apparent that both fragments may not belong to a single chemical species. Nevertheless, a few interesting observations can be made in regards to the experiment with tert-BuOD. In the temperature range where m/e 69 and 55 become important in the TPReac

152 of t-BuOH over HY (oligomeric masses, 450-500 K, see Fig. 5), the 41/42 ratio with tertBuOD follows an opposite trend to that observed in the tert-BuOH run. Note that due to the protic nature of HY, we expect alcohols to undergo facile H+/D§ exchange [16]. The residence time of the olefin oligomers that manage to desorb from zeolite pores is expected to be larger than that of isobutene. Hence, the larger proportion of HDC=CH-CH2 § (or H2C=CD-CH2 § in products is believed to be due to secondary species, namely olefin oligomers. A similar behavior has been observed in H-Mordenite [7]. In WZ, the smaller 41/42 ratio observed with tert-BuOD suggests that deuteration and proton exchange indeed occur rapidly prior to isobutene formation. These observations are in qualitative agreement with those of other m/(m+ 1) pairs. 4. C O N C L U S I O N In summary, tungstated zirconia solid acids are more resistant to carbon deposition in alcohol dehydration than HY. The products of this reaction are good olefin precursors and can be used as chemical probes for deactivation from conventional polymerization of carboncontaining molecules. The nature of the coke deposits is also different for the two catalysts, judging from both DRIFTS and TPO observations, these being more easily removable in WZ. This particular issue deserves further study. Reactive desorption is faster in WZ than HY, which in principle suggests that tungstated zirconia is more acidic than HY. This cannot be stated uniquivocally because readsorption phenomena may also play a role. Both WZ and HY catalyze the H+/D+ exchange of product olefins. ACKNOWLEDGMENTS We thank the donors of the Petroleum Research Fund (American Chemical Society, Grant 31067-AC5) for generous support. REFERENCES

1. 2. 3. 4. 5. 6. 7. 8. 9. 10. 11. 12. 13. 14. 15. 16.

S.L. Soled, W. E. Gates and E. Iglesia, U. S. Patent 5,422,327 (1995). G. Larsen, E. Lotero, S. Raghavan, R. D. Parra and C. A. Querini, Appl. Catal. Gral., 139 (1996) 201. G. Larsen and L. M. Petkovic, Appl. Catal. A: Gral, 148 (1996) 155. G. Larsen and L. M. Petkovic, J. Mol. Catal. A: Chem. 113 (1996) 517. V. Adeeva, G.D. LeiandW. M.H. Sachtler, Appl. Catal. A:Gral. 118(1994)Lll. E. Iglesia, D. G. Barton, S. L. Soled, S. Miseo, J. E. Baumgartner, W. E. Gates, G. A. Fuentes and G. D. Meitzner, Proc. 1 lth Int. Congr. on Catal., J. W. Hightower and W. N. Delgass, Eds., Elsevier Sci. Publ., 1996, p. 533. G. Larsen, E. Lotero, L. M. Petkovic and D. S. Shobe, J. Catal., submitted. A. Cornu and R. Massot, Compilation of Mass Spectral Data, Vol. 1, 2nd Ed., Heyden, London, 1975. S.C. Fung and C. A. Querini, J. Catal. 138 (1992) 240. G. A. M. Hussein, N. Sheppard, M. I. Zaki, and R. B. Fahim, J. Chem. Soc. Far. Trans. 1 85 (1989) 1723. H. G. Karge, W. Niel3en and H. Bludau, Appl. Catal. A: Gral. 146 (1996) 339, and references cited therein. D. G. Blackmond, J. G. Goodwin, Jr. and J. E. Lester, J. Catal. 78 (1982) 34. H. G. Karge and E. Boldingh, Catal. Today 3 (1988) 53. J. P. Lange, A Gutsze, J. Allgeier and H. G. Karge, Appl. Catal. 45 (1988) 345. C. Li, Y-W. Chen, S-J. Yang and R-B. Yen, Appl. Surf. Sci. 81 (1994) 465. F. Haase and J. Sauer, J. Amer. Chem. Soc., 117 (1995) 3780.

9

Elsevier Science B.V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

D e t e r m i n a t i o n of Coke Deposition on Metal Active Sites o f Propane D e h y d r o g e n a t i o n Catalysts P.Praserthdam*, T.Mongkhonsi, S.Kunatippapong, B.Jaikaew and N.Lim Petrochemical Engineering Laboratory, Department of Chemical Engineering, Faculty of Engineering, Chulalongkorn University, Bangkok 10330 THAILAND. e-mail : [email protected] Sn and alkali metals (Li, Na and K) can reduce coke covering on the Pt active site of a propane dehydrogenation catalyst, Pt/y-A1203. The role of the alkali metals is to increase excess mobile electrons of the catalyst surface. Sn and Sn-alkali metal promoted catalysts show higher excess mobile electrons than unpromoted ones. The excess mobile electrons enhance hydrogen spillover on the catalyst surface, thus reducing the amount of coke deposits.

I. INTRODUCTION Coking is a common deactivation mode in hydrocarbon conversion processes, involving the deposition of carbonaceous materials on the catalyst surface. Materials deposit may include elemental carbon, high molecular weight polymer and polycyclic aromatics [1,2]. Coke formation involves the metallic and acidic functions of the catalyst with the steps of dehydrogenation, condensation, alkylation and cyclization [3]. The structure of coke is rather complex, containing several different growth forms, which can be grouped into amorphous, filamentous and graphitic platelets [4-6]. The surface on which coke is deposited and their effects on coking can also vary widely. Most metallic catalysts are supported and the metal, the support and metal-support interaction can affect the coking. The thermodynamics of the dehydrogenation reaction of propane to propene are such that it is desirable to operate at high temperature and low pressure. But these conditions are the conditions that favour coke formation. Therefore, there are many attempts trying to improve the performance of the present catalyst, based on Pt/y-A1203, and to develop new catalyst compositions that yield the desired results. Sn and alkali metals are examples of promoter that can increase catalyst resistance to coking [7-11 ]. In the present work, the effect of Sn and the alkali metals (Li, Na and K) is presented. The main objective is to clarify their role in enhancing coking resistance of the resulting catalyst.

2. EXPERIMENT

Pt/y-A1203 (0.3wt%Pt), Pt-Sn/y-Al203 (0.3wt%Pt, 0.3wt%Sn), Pt-Sn-Li/y-A1203 (0.3wt%Pt, 0.3wt%Sn, 0.6wt%Li), Pt-Sn-Na/y-A1203 (0.3wt%Pt, 0.3wt%Sn, 0.6%wtNa), and

154 Pt-Sn-K/T-AI203 (0.3wt%Pt, 0.3wt%Sn, 0.6wt%K) were used in the research. The catalysts were prepared by a conventional dry impregnation method using H2PtCI6, SnCI2 and alkali metal nitrates as salt precursors. All chemicals used are normally analytical grade. Coked catalyst was prepared from the dehydrogenation reaction of C3Hs to C3H6. 0.1 mg of the catalyst was packed in a quartz reactor. 20% C3Hs, balanced with N2 was used as reactant gas. To study the effect of HE, H2 was mixed with the reactant gas at a Hydrogen/Hydrocarbon (H/HC) ratio equal to 1. All gas reactants were supplied by Thai Industrial Gas Co.Ltd. and were passed through oxygen and moisture traps before entering the reactor. The reaction was performed at near atmospheric pressure with gas hourly space velocity (GHSV) 25000 hr-1. Temperature programmed oxidation (TPO) was performed by burning the obtained coked catalyst in 1%02 in an He atmosphere. The heating rate was 10~ CO2 produced was measured using a gas chromatograph equipped with a TCD and an on-line gas sampling valve. To measure the amount of metal active sites, a CO adsorption technique was used. In the case of fresh or coked samples, the measurement was performed by monitoring the amount of CO adsorbed at room temperature. The % active site covered by coke was defined as (active site lost due to coke coverage)/(active site of fresh catalyst) x 100 The electrical conductivities of all catalyst samples were measured using a Philips PM 6303 automatic RCL meter. The catalyst was first ground to a fine powder and packed into a die. Then the sample was reduced with H2 for 1 hour. At the end of the reduction period, the powder was pressed at 13.3 MPa for 5 min. The measurement was performed under these condition. This measurement was used only for qualitative guidance.

3. R E S U L T S A N D D I S C U S S I O N

3.1. Temperature programmed oxidation and metal active site measurement Figure 1 shows the TPO profiles of 0.3%Pt/y-AI203, 0.3%Pt-0.3%Sn/y-AI203 and 0.3%Pt0.3%Sn-0.6%Li/y-A1203 catalysts. Each sample shows a TPO peak around 460~ In addition, a small peak around 100~ was observed for all catalysts used in this work. For the same operating conditions, the amount of coke deposit can be arranged in the following order: Sn-promoted > unpromoted > Sn-Li-promoted. However, when based on propane conversion, the following order was found ; unpromoted > Sn-promoted > Sn-Li-promoted. Sn-Na- or Sn-K-promoted catalysts also have less coke than unpromoted and Sn-promoted. BET surface area and metal active site measured by CO adsorption of the promoted and unpromoted fresh catalysts are shown in Table 1. The addition of Sn significantly reduced the Table 1 BET surface area and metal active site Surface area (m2/g cat) Catalyst 316 y-Al203 366 0.3%Pt/y-A1203 351 0.3%Pt-0.3%Sn/T-AI203 282 0.3%Pt-0.3%Sn-0.6%Li/y-A1203 289 0.3 %Pt-0.3 % Sn-0.6%NaJ~/-A12O3 304 0.3 %Pt-0.3 % Sn-0.6%K/y-Al2O3

Metal active site (site/$ cat) 1.63 x 10 is 0.73 x 10 is 1.17 x 10 Is 1.53 x 10 is 1.50 x 10 Is

155

number of surface Pt atom. In this case, more than half of the surface Pt atoms disappear. However, the incorporation of the alkali metals can the increase number of surface Pt atoms again. The percentage of metal active site covered by coke is shown in Figure 2. The figure demonstrates that in the low reaction temperature region ,i.e. < 550~ Sn does reduce coke deposits on the metal active site. At higher reaction temperatures, however, adding only Sn does not yield any benefit. The addition of alkali metals to Sn-promoted catalysts significantly increases the metal surface available in both low and high reaction temperature regions. 12000 ~Pt =i - - Pt-Sn ~i 9000-. . . . Pt-Sn-Li ---- Pt-Sn-Na .~ 6000 -Pt-Sn-K

/ ta~/ I I

li !

O 3000-0

' 0

- ;--- ,

100 200

,

,

300 400 500 Temp. (C)

- 9 --600

700

Figure 1 Effect of Sn and alkali metals on TPO spectra. Reaction Temperature 600~

.•100 80

~ 60

, ,-~~'"

~ 40t

1/

--Pt-A1203

"~

"/

--- -Pt- - -Sn/AI203 - - -----

~9 2 o

O450

500

550 Temp. (C)

600

650

Figure 2 Total coke deposited on the metal active site of various catalysts at different reaction temperatures. H/HC = 0 The effect of hydrogen partial pressure on % active site covered by coke is shown in Figure 3. For Pt/?-A1203 catalyst, H2 decreases coke deposits only in the initial period. On the other hand, on Sn or Sn-Na promoted samples, higher H2 pressure results in less coke on the metal active sites throughout the reaction period. This underlines the role of Sn and the alkali metals in enhancing the activity of H2 in the coke elimination process. It should be noted here that coke can cover a fraction of the metal active sites, in accordance with the literature [12,13].

156 ~100

-

O O

8

0

-

-

~

.i

~ 60--'

g O

9 40 -~i 9~

-

9-~ 20 o "< 0 0

...

I 100

pt.

H/HC=

1

"-- Pt-Sn 9H/HC = 0 -'- Pt-Sn" H/HC = 1 --,- Pt-Sn-Na 9H/HC = 0 -.- Pt-Sn-Na" H/HC = 1 I I 200 300 Time (min)

400

Figure 3 Coke deposit on the metal active site of 0.3%Pth/-A1203, 0.3%Pt-0.3%Sn/~/-A1203 and 0.3%Pt-0.3%Sn-0.6%Na/~/-A1203 catalysts at reaction temperature 500~ H/HC = 0,1. 3.2 Electrical conductivity Table 2 shows electrical conductivity data of the catalysts and the support. A is the electrical conductivity of alumina. B is the electrical conductivity of Sn and alkali metals promoted alumina. C is the electrical conductivity of Pt catalyst while !) is the electrical conductivity of Pt-Sn catalyst. E is the electrical conductivity of Pt-Sn-Alkali metals catalyst. The data shows that the addition of metal to alumina increases electrical conductivity. The addition of Sn to Pt catalyst augments electrical conductivity approximately three times. Further incorporation of the alkali metals results in an order of magnitude further increases. Since electrical conductivity reflects the mobility of electrons in the bulk solid (14), the data in Table 2 can be used to compare the amount of mobile electrons in each sample. Table 3 shows the amount of excess mobile electrons (in conductivity unit) of the catalysts shown in table 2. The value (B-A) is the electrical conductivity of 0.3wt%Sn added to ~-A1203 support. The value (D-C) is the electrical conductivity of 0.3wt%Sn added to 0.3%Pt/7-A1203 catalyst. If Sn does not have any electronic effect on the Pt site, the value (B-A) should be equal to the value (D-C). The calculation, however, clearly indicates that 0.3wt%Sn loaded on 0.3%Pt/~/A1203 catalyst does provide more mobile electrons to the catalyst than its presence on ~/-A1203 support. The addition of alkali metals also shows an interesting result. The value (E-D) is the increase in electrical conductivity of 0.3%Pt-0.3%Sn/~/-A1203 after 0.6wt% of the alkali metals was added. The result demonstrates that the alkali metals greatly increase the amount of the excess mobile electrons in the bulk catalysts. By decreasing the amount of coke, Sn functions by creating an ensemble effect and forming a solid solution with Pt in electron-rich Pt sites [15]. The proposed synergistic model for Sn addition is exhibited in Figure 4. The presence of Sn on Pt surface results in a dilution in the number of large active ensembles of Pt. Thus, it is more difficult for coke molecules to deposit on the metal surface. In addition, Sn also provides some additional electron to Pt site. Not only do the alkali metals not only act as electron donors to Pt [16], but their addition also decreases the acidity of the catalyst, which results in less coke forming on the support. Moreover, the alkali metals also promote hydrogen spillover, which can eliminate some coke already formed on the metal site. The alkali metals also act as textural promoters by reducing

157

Pt-Sn alloy formation. The synergistic mechanism model for Sn and the alkali metal addition are shown in Figure 5. Table 2 Electrical conductivity data of catalysts and support Code Catalyst Electrical conductivity (Ohm 1c m "l ) A ]r 2.15 X 10 .6 B 0.3%Sn/~/-A1203 3.05 x 10.6 0.6%Li/~-A1203 5.58 x 10-6 0.6%Na/y-A1203 5.23 x 10.6 0.6%K/7-A1203 4.04 x 10-6 C 0.3%Pt/)'-A1203 3.25 x 10.6 D 0.3%Pt-0.3%Sn/),-A1203 9.43 x 10.6 E 0.3%Pt-0.3%Sn-0.6%Li/~/-A1203 20.3 x 10.6 0.3%Pt-0.3%Sn-0.6%Na/~/-A1203 26.6 x 10.6 0.3%Pt-0.3%Sn-0.6%K/~{-A1203 74.7 x 10.6 Table 3 Amount of excess mobile electrons on the surface of bulk catalysts (in conductivity units) Equations excess mobile electrons Catalysts (Ohm-1 c m "1)

(D-C) (E-D) (E-D) (E-D)

0.3%Pt-0.3%Sn/qt-A1203 0.3 %Pt-0.3 % Sn-0.6%Li/~/-A12 O3 0.3 %Pt-0.3 %Sn-0.6%Na/7-Al2 O3 0.3 %Pt-0.3 % Sn-0.6%K/y-AlE O3

-

(B-A) (B-A) (B-A) (B-A)

5.28 7.44 14.1 63.4

x 10-6 x 10.6 x 10.6 x 10.6

C--C

C--C I C [ e-

I C

Hi

'/-A1203 Figure 4 Synergistic mechanism model for Sn addition C--C

I

C

C

I

C--C

C--C

I H/e

I

I

C

e-

fiL

~,-A1203 Figure 5 Synergistic mechanism model for Sn and alkali metal (A) addition

HI

158 4. CONCLUSIONS Test results reveal that the addition of only Sn or Sn + alkali metals (Li, Na and K) can reduce coke covering on Pt active sites of propane dehydrogenation catalyst, Pth/-Al203. Sn reduces the coke coverage area on the metal active site by creating an ensemble effect and providing additional electrons to Pt atoms. The role of the alkali metals is to increase excess mobile electrons of the catalyst surface and reduce Pt-Sn alloy formation. Sn and Sn-Alkali metal promoted catalysts show higher excess mobile electrons than the unpromoted ones. The additional excess mobile electrons enhance hydrogen spillover on the catalyst surface, thus, reducing the amount of coke deposits on the catalyst surface.

REFERENCES

1. F.Caruso, E.L.Jablonski, J.M.Graw and J.M.Parera, Appl. Catal., 51 (1989) 195. 2. B.C.Gates, J.R..Katzer and G.C.Schuit, Chemistry of Catalytic Processes, McGraw Hill, New York, 1979. 3. J.M.Parera, R.J.Venderone and C.A.Querini in Catalyst Deactivation 1987, B.Delmon and G.F.Froment (eds.), p 135, Amsterdam, 1987. 4. H.B.Palmer and C.F.Cullis, Chemistry and Physics of Carbons (vol. 1) p265, Dekker, New York, 1965. 5. R.T.K.Baker and P.S.Harris, Chemistry and Physics of Carbons (vol. 14) p83, Dekker, New York, 1978. 6. B.W.Gainey, U.S. Energy Research and Development Admin. Report. GA-A13982, UC77. 7. J.Barbier in B.Delmon and G.F.Froment (eds.), Catalyst Deactivation 1987 p135, Amsterdam, 1987. 8. J.Barbier, P.Marecot, N.Martin, L.Elassal and R.Maurel in Catalyst Deactivation 1987, B.Delmon and G.F.Froment (eds.),p135, Amsterdam, 1987. 9. R.D.Corright and J.A.Dumesic, J.Catal., 157 (1995) 576. 10. O.A.Barias, A.Holmen and E.A.Blekkan, J.Catal., 158 (1996) 1. 11. S.de Migues, A.Castro, O.Scelza, J.L.G.Fierro and J.Sorin, Catal. Lett., 36 (1996) 201. 12. J.Barbier, G.corro, P.Marecot, J.P.Bournonville and J.P.Frank, React. Kinet. Catal. Lett., 28 (1985) 245. 13. J.Barbier, E.Churin, J.M.Parera and Riviere, React. Kinet. Catal. Lett., 29 (1985) 323. 14. F.W.Sears, M.W.Zemansky and H.D.Yound, University Physics, 5th ed. p784, AddisonWesley Publishing Company 1981. 15. R..Burch, J. Catal., 71 (1981) 348. 16. J.Oudar, in J. Oudar (ed.), Deactivation and Poisoning of Catalysts, New York, Dekker, 1986.

9 Elsevier Science B.V. All rights reserved Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

159

T h e R o l e o f C o k e D e p o s i t i o n in the C o n v e r s i o n o f M e t h a n o l to Olefins over SAPO-34 De Chen a, Hans Petter Rebo a, Kjell Moljord b and Anders Holmen ~ aDepartment of Industrial Chemistry, Norwegian University of Science and Technology (NTNU), N-7034 Trondheim, Norway. bSINTEF Applied Chemistry, N-7034 Trondheim, Norway. The conversion of methanol and dimethyl ether (DME) to olefins (MTO and DTO) over SAPO-34 has been studied using a Tapered Element Oscillating Microbalance (TEOM). The effect of coke formation in the MTO reaction was investigated by coking SAPO-34 to desired coke levels using different probe molecules before exposing the catalyst to methanol. Mainly internal coke was obtained when using propene, while mainly external coke was obtained when using a Ca component too large to enter the pores of SAPO-34. The coke formed from oxygenates, referred to as active coke, promoted olefin formation, while the coke formed from olefins, referred to as inactive coke, only had a deactivating effect. The role of DME during MTO was studied by comparing the MTO and DTO reactions at different coke levels. Apparently, olefins are formed via DME, and the diffusion of DME, which is affected by coke, plays a role in the conversion of methanol to olefins in MTO. The yield of olefins during the MTO reaction was found to go through a maximum as a function of both time and amount of coke. This is partly due to promotion by active coke formed initially, while further coke formation led to deactivation, probably by pore blocking. 1. INTRODUCTION The catalytic conversion of methanol to lower olefins (MTO) is a potential route for converting natural gas to chemicals. The small pore silicoaluminophosphate molecular sieve SAPO-34, which is an isomorph of cabazite [1 ], is a promising catalyst for obtaining high selectivities of ethene and propene from methanol [2, 3]. Many investigations have recently been devoted to the study of the effect of the reaction conditions on the activity and selectivity of SAPO-34 [4,5], and examination of the MTO reaction mechanism [6]. However, most of the experiments have been carried out at 100% methanol conversion due to the very active catalyst. Furthermore, the catalyst deactivates very rapidly due to coke formation, making it difficult to decouple the kinetics of the main reactions from the coke forming reactions. The effect of coke deposition on the activity and selectivity during MTO over SAPO-34 is therefore still not obvious. The Tapered Element Oscillating Microbalance (TEOM) reactor has recently been applied to study deactivation of zeolite catalysts [7,8]. The main advantage of the TEOM reactor is that all gases in the reaction mixture are forced to flow through the catalyst bed as in a conventional fixed-bed reactor. Coupled with on-line gas chromatography, the catalyst activity, selectivity and coking rate can be measured simultaneously as a function of the amount of coke on the catalyst. Hence, the TEOM represents a unique way of studying the effect of coke deposition in detail. The present work deals with the effects of coke deposition in the MTO reaction over SAPO-34 using the TEOM reactor. The effects of the external and internal coke on the activity and selectivity on SAPO-34 are studied by controlled coke formation from

160

different probe molecules. The role of dimethylether (DME) in the reaction network, was studied by comparing the results from conversion of methanol with the results from DME. 2. E X P E R I M E N T A L The reactions were carried out in a TEOM reactor where the weight of the catalyst bed is continuously recorded. The setup is similar to that described previously [8]. The methanol flow was controlled by a liquid flow controller while DME and propene were fed using gas flow controller. The MTO and DTO reactions were carried out at 425~ WHSV=417 h "I and a methanol or DME partial pressure of 8 kPa, with helium as diluent. One DTO experiment was also performed at WHSV=600 h -1 to keep the residence time identical to those from the MTO experiments. Such high space velocity and low partial pressure were used to avoid non-uniform coke distribution through the catalyst bed, and to keep the conversion well below 100% to minimize secondary reactions of olefins. SAPO-34 with a unit cell composition of (Si2.86Al17.49PIs.66)O72 and an average crystal size of 0.25 m was obtained from SINTEF-Oslo. A detailed description of the preparation and characterization of the catalyst can be found elsewhere [9]. The sample was calcined in a quartz reactor at 550~ for 6 hours in diluted air (10% O2). About 5 mg (dry basis) of calcined SAPO-34 was installed in the TEOM reactor and dried in situ at 500~ for 3 hours. Due to rapid catalyst deactivation, the MTO reaction was studied using 1 minute interrupted pulses with GC analysis taken after 1 minute. The time between each pulse was 40 minutes, allowing completion of the GC analysis. It has been found by preliminary experiments that the aging in helium did not significantly influence the coke formation and deactivation. The Cl to C6 hydrocarbons, methanol and DME were analyzed with a HP 5890 GC using a GC-Q capillary column (30 m, 0.543 mm). The amount of water formed was calculated from the mass balance. The yield of olefins was defined as: y=l- XMeOH XDME-XH20, where x is the weight fraction of each reaction component. Additional experiments with controlled coke formation were performed in order to study the effect of coke nature and location on the MTO reaction. Internal coke was obtained using propene while external coke was obtained using i-butene produced from i-butanol, iButanol was introduced by switching to a carrier gas saturated in thermostated bottles containing the liquid. After the desired coke content was reached by using propene or ibutene/i-butanol, methanol or DME was introduced at the standard MTO conditions. The pore volume of fresh and coked catalysts was determined gravimetrically by TPD of catalyst samples saturated with water at room temperature. 3. R E S U L T S AND DISCUSSION 3.1. Coke formation

A space velocity of 417 h -1 always gave a methanol conversion less than 60%. Since methanol was converted not only to olefins but also to DME and water, the yield of olefins was relatively low (maximum yield 16 wt%) at these conditions. This, combined with low catalyst loading (5 mg dry SAPO-34), allowed us to assume uniform coking through the catalyst bed. Interestingly, as coke was formed during MTO over SAPO-34, the yield of olefins went through a maximum at a coke content of about 7 wt% (Fig. 1). When DME was used as the reactant with a similar residence time, the yield of olefins increased only slightly with coke formation and went through a maximum at about 5 wt% coke (Fig. 1). The initial and maximum reaction rates and coking rates for the different reactants are presented in Table 1. The coke formation with time on stream is shown in Figure 2. The MTO and DTO initial reaction rates were identical, but the maximum MTO rate was three times higher than the maximum DTO rate. Olefin distributions were almost identical for MTO and DTO. These results indicate that methanol and DME are probably converted to olefins through the same reaction mechanism. However, the DME conversion can be retarded by such effects as slower diffusion into the pores of SAPO-34 compared with methanol. Experiments with SAPO-34 of different crystal sizes have confirmed that

161

DME diffusion limits the DME conversion. The selectivities of olefins were identical and independent of crystal size for both methanol and DME conversion [ 10]. Table 1 Reaction and coking rates over SAPO-34 Sel coke,ini2 Reactant r ini. 1,3 r max. 1 r coke, ini. r coke, max mol/g,h mol/~,h g/g,h g/~,h wt% 1.15 4.60 0.37 0.82 2.26 Methanol 1.14 1.44 0.38 2.46 DME 0.05 0.02 0.92 Propene 1) Reaction rate based on oxygenates conversion calculated on a CH2 basis [3] at WHSV=417 h 1 for methanol, 600 h ~ for DME and 417 h1 for propene. 2) Selectivity to coke=deposited coke(g)/hydrocarbons formed(g). 3) Initial data was obtained from the first pulse.

16

16

.-,o. "-'12

312

e-

r ~

8

" o o

8

-~ 4 >-

o

o

4

g o "13 .-.

O

0

4

8

12

16

Coke (wt%) Figure 1.Yield of olefins versus coke content during : 9 MTO; A: DTO at WHSV: 417 h-l; 9 : DTO at WHSV: 600 h-1;e:MTO after 3.5 wt% coke from propene.

,

0

v

5

10

15

.

.

.

.

.

20

25

30

Time on Stream (rain)

35

Figure 2. Coke formation with time on stream: m:MTO; o:MTO after 0.63 wt% coke from ibutanol; 0 : propene; A:DTO at WHSV=417 h -~" *: DTO at WHSV=600 h -1.

Similar coking rates and coke selectivities were found for methanol and DME conversion, while the coking rate was about twenty times higher and the initial selectivity was two times higher than for propene conversion. Figure 2 shows an increased coking rate during DTO as a result of lower space velocity or higher conversion. Interestingly, it can be seen that the coking rate at low coke content (< 8 wt%) is much higher than at higher coke content during DTO, while the significant decrease in the coking rate was observed at much higher coke contents (> 13 wt%) during MTO. This might be explained by pore blocking by coke and by a lower rate of reactant DME diffusion at high coke contents. The influence of methanol diffusion during MTO is less important, probably due to the smaller molecular size of methanol in comparison to DME. Studies on SAPO-34 of different crystal sizes [10] have shown that the effectiveness factor decreases significantly with coking during the DTO reaction. In contrast to the results obtained at higher conversions [11], the initial coking rate during relative low propene conversion was about twenty times lower than during the

162 methanol and DME conversions. This indicates that secondary reactions and coke formation from olefins in the gas phase are less important in the presence of oxygenates during the MTO reaction. 3.2. Effect of external coke

The SAPO-34 external surface is relatively large for the small crystal material used in this study. The role of the external surface in the MTO reaction is quite interesting but not clear. Coke deposition on HZSM-5 during methanol conversion was reported to occur essentially on its external surface [12]. Whether coke deposition during MTO on small crystals of SAPO-34 occurs in the same way is less clear. Experiments with controlled coke formation at the external surface were performed using i-butanol as a probe molecule. i-Butanol will dehydrate to i-butene on acidic sites on the external surface, but i-butanol and i-butene are too large to enter the pores of SAPO-34. Therefore, i-butene is a potential coke precursor on those external acidic sites. 20

The coking rate from i-butanol/i-butene was found to be rather low, and only 0.2 wt% o~ coke was formed when SAPO-34 was --16 exposed to i-butanol at a partial pressure of 1.2 kPa for 90 min at 425~ The formation of i-butene was confirmed by GC analysis. .E 12 tim After further exposing this sample to ibutanol at 460~ for 1 h, the coke content 0 .-8 increased to 0.63%. Furthermore, hydroo carbons heavier than C6 and aromatics were "13 absent in the products during MTO and DTO. This suggests that the acidic sites at the external surface are rather weak. If the external acidic sites are strong enough, a nonselective reaction is expected. Since similar 0 4 8 12 16 coking behavior was found for precoked Coke (wt%) samples and samples that had not been pretreated (Fig. 2), it follows that the coke Figure 3. Yield of olefins as function of formed during MTO at 425~ is formed precoke content from i-butanol during MTO: mainly in the intemal pores. However, it =: 0 wt%; e: 0.63 wt%; A:l.5 wt%. should be noted that the ethene to propene ratio increased more rapidly with coking at high coke content (Fig. 5), while the yields of olefins decreased rapidly with coking (Figs. 1 and 3). This indicates extemal coke formation at high coke content, probably occurring from growth of coke molecules originating from intemal active sites. Figures 3 and 6b illustrate that a small amount of coke formed from i-butanol/i-butene slightly increases the yield of olefins and the methanol to DME ratio during MTO. The DME yield was 3 wt% lower for the 0.63 wt% precoked sample compared with the sample without pretreatment for almost all coke levels (2 - 15 wt%). This additional DME is obviously formed at the external surface of fresh SAPO-34 during MTO. However, the role of DME formed at the external surface during the MTO reaction does not seem to be very significant. When the catalyst was exposed to i-butanol at 500~ the coke formation was relatively fast. 1.5 wt% coke was deposited during 10 minutes. The formation of olefins from methanol was lower for this precoked sample. This is probably due to n-butene formation, originating from i-butene isomerization at 500~ n-Butene probably forms inactive coke in the pores or at the pore openings, which results in significant deactivation.

163 12

r.,

~1.6

: I

1.2

8

.m LIE n

0

o

o

IX.

i.#..

"o

4

o 0.8

(11

>.

t-

t-

uJ 0.4 0

5 Coke

10

15

(wt%)

Figure 4. Yield of olefins versus coke content during DTO for different coke contents obtained from propene: &: 0 wt%; ":2.4 wt%; e: 4.7 wt%.

0

4

8

Coke

12

16

(wt%)

Figure 5. Ethene to propene ratio versus coke content during:MTOovercatalysts: (":fresh; *: 4.2 wt% from propene; o: 0.63 wt% and 0: 1.5 wt% from i-butanol); A:DTO.

3.3. Active and inactive coke The effect of coke deposition on the MTO reaction is rather complex. Coke was found to influence not only the external DME formation, but also the DME conversion taking place internally during MTO. However, the effect of coke deposition on the DTO reaction is much more simple and allows us to focus on the effect of intracrystalline coke on the DME conversion. The amount of methanol formed from DME at DTO reaction conditions is quite small since methanol formation requires water produced from olefin formation. Thus, the DTO reaction model appears to be straightforward: DME diffuses into the pores and is converted to olefins on the acidic sites in the cavities of SAPO-34. Figure 1 shows that the DME conversion increased slightly with coke formation up to about 5 wt% coke. Coke appears to have a promoting effect besides a deactivating effect on the reaction forming olefins. A number of reaction mechanisms reviewed by Froment et al. [13] have been proposed for the MTO reaction over zeolites. Although the reaction mechanism is still a matter of debate, one can assume that some kind of surface species which takes part in the reaction is formed on the active sites. Methoxy compounds in coke during methanol conversion on HZSM-5 and H-mordenite have been confirmed by NMR measurements [14]. Formation of such surface species causes mass increase, but the complete weight increase was ascribed to coke in the present work. The surface species formed can react with oxygenates to form olefins, hence creating a promoting effect. Accordingly, if larger amounts of these species are formed the rate of olefin formation is higher. Nevertheless, it must be noted that another type of coke is simultaneously formed from the olefins, when the active coke reacts with oxygenates to form olefins. Although the apparent maximum reaction rate is only 1.3 times higher than the initial rate, the net promoting effect of coke is probably larger than observed, if the deactivating effect is eliminated. A set of experiments with DTO and MTO on SAPO-34 precoked with propene were performed at 425~ to determine the role of coke formed from olefins. Due to a low propene coking rate, it was necessary to use a relatively high propene partial pressure (>30 kPa) to reach coking rates similar to those used for MTO. It is important to keep the time on stream (TOS) similar for the desired coke amounts obtained in both methanol and

164 propene conversion, since TOS might also influence the nature of coke. Propene is expected to form mainly intracrystalline coke in similar locations as in MTO or DTO. The yield of olefins over this precoked sample was lower compared to a fresh sample for the same coke content in DTO (Fig. 4) and in MTO (Fig. 1). This indicates that the nature of the coke formed from oxygenates is different from the coke formed from olefins, and that the coke formed from olefins is inactive for the MTO and DTO reactions.

4o

I

12

(a)

0

(b)

E u

LU

2zo

r~

6

0

0

-a

0 C:

~_1o

3

,e.., (D i

0

4

I

8 C o k e ('r

0

I

12

16

0

4

Coke

8

12 (wt%)

16

Figure 6. Yield of DME (a) and methanol to DME ratio (b) with coke formation after precoking with i-butanol to different levels:" 0 wt%; e: 0.63 wt%; &: 1.5wt%.

Another possible explanation for the lower conversion over precoked SAPO-34 is that the coke formed from propene would be of such a nature that it would result in a more significant decrease in the void space of the cavities. This was examined by porosity experiments using water adsorption. The pore volume found by the saturation of water for a fresh SAPO-34 was 0.30 cm3/g, which is in good agreement with literature data [15]. Although this volume is a bit larger than what was found from nitrogen adsorption (0.26 cm3/g), it is a useful parameter to specify the relative volume changes. The pore volume was found to be 0.21 cm3/g for 3.14 wt% coke during MTO, while it was 0.20 cm3/g for 3.1 wt% coke during propene conversion. The difference in pore volume was quite small, and it can be concluded that the lower activity of precoked SAPO-34 is mainly caused by the inactive nature of coke formed from olefins. Therefore, the coke in the MTO and DTO should be divided in two categories: Inactive coke formed from olefins having a deactivating effect and active coke formed from oxygenates having a promoting effect. The activity at different coke contents depends on the ratio of active to inactive coke. 3.4. Effect of coke on the ethene to propene ratio

The ethene to propene ratio is an important parameter in the MTO reaction. It is shown in Figure 5 that this ratio increases with coke formation. For similar coke contents, the ethene to propene ratio was almost identical for coke formed from both oxygenates and propene. This indicates that the ethene to propene ratio depends only on the coke content, regardless of the nature of coke. It has been found that the ethene to propene ratio was also

165

independent of the crystal size and the transition-state shape selectivity seems to be important in product selectivities. A detailed discussion about the effect of coke formation on the shape selectivity on SAPO-34 is given elsewhere [10]. However, a higher ethene selectivity by deposition of inert material in the pores to reduce the pore volume can be expected. 3.5. Effects of coke on D M E diffusion and conversion

The methanol to DME ratio was found to increase with coke formation on the catalysts (Fig. 6b). This ratio was quite far from the ratio at chemical equilibrium, which was calculated at 425~ to be 0.47. The deviation from chemical equilibrium ratio was found to be larger at higher coke contents and on the externally precoked samples. This indicates that the methanol conversion to DME is not fast enough to reach equilibrium, probably due to the moderate external acidity of SAPO-34 and partly also due to the effect of diffusion of DME at the higher coke contents. MTO can be considered a simple sequence reaction: Methanol~ D M E ~ olefins For low coke contents, the yield of olefins increased (Fig. 1) while the DME yield decreased with coke formation, even for the samples modified by external coking (Fig. 6a), where the effect of external surface was eliminated. The observed decrease in DME yield with coke formation for externally coked samples corresponds with a mechanism where DME is formed, diffuses and is converted in the catalyst pores. The diffusion of DME is expected to largely influence the olefin formation in the pores. The DME diffusivity decreased with coke formation, and the difference in activity with the amount of coke formed on SAPO-34 between MTO and DTO can then be explained by DME diffusion. Furthermore, the decrease in DME diffusivity has the opposite effect for DTO and MTO. Lower diffusivity decreases the DME concentration in the pores in DTO, hence reducing the olefin formation, while it increases the DME concentration in MTO, thereby promoting the formation of olefins. Obviously coke formation has both negative and positive effects during MTO. The positive effects are promotion by active coke and decrease in the DME diffusivity, and the negative effect is poisoning of the active sites resulting in MTO deactivation. The observed yields of DME or olefins at different coke contents represent the net result from negative and positive effects. At high coke contents, the deactivating effect which might be due to pore blocking, was dominant. This resulted in increased DME yield and decreased olefin yield. The increased DME yield is explained by faster deactivation for DME conversion to olefins rather than by DME formation from methanol. 4. CONCLUSIONS The effect of coke formation during the MTO reaction over SAPO-34 at 425~ summarized as follows:

can be

1. Methanol is converted to DME both externally and internally over SAPO-34. The DME formation at the external surface was not significant, probably due to the relatively weak external acidity of SAPO-34. The formation of DME at the external surface gave a slightly lower rate for olefin formation. The methanol to DME ratio found from experiments was far from equilibrium. 2. The coke, which was recorded as the total mass increase in the TEOM reactor, can be divided into active coke formed from oxygenates and inactive coke formed from olefins. The active coke promotes the conversion to olefins and could in fact represent the true surface intermediate of the reaction. 3. Coke reduces the DME diffusivity, which enhances the formation of olefins. 4. The coke formation enhances the ethene formation. The ethene to propene ratio increased with intracrystalline coke content, regardless of the nature of coke.

166 ACKNOWLEDGMENTS

The authors thank the Norwegian Research Council and Norsk Hydro ASA for supporting this work, and Terje Fuglerud, Norsk Hydro for fruitful discussions. REFERENCES

1. 2. 3. 4. 5. 6. 7. 8. 9. 10. 11. 12. 13. 14. 15.

N.Y. Chen, W.E., Catal. Rev. -Sci. Eng. 28 (1986) 185. B.V. Vora, T.L. Marker, P.T. Barger, H.E. Fullerton, H.R. Nilsen, S. Kvisle, T. Fuglerud, Stud. Surf. Sci. Catal., 107 (1997) 87. A.N. Ren6, P.J.J. Tromp, H.N. Akse, Ind. Eng. Chem. Res., 34 (1995) 3808. J. Liang, H. Li, S. Zhao, W. Guo, R. Wang, M. Ying, Appl. Catal. 64 (1990) 31. A.J. Marchi, G.F. Froment, Appl. Catal., 71 (1991) 139. I.M. Dahl, S. Kolboe, J. Catal., 149 (1994) 458. F. Hershkowitz, P.D. Madiara, Ind. Eng. Chem. Res., 32 (1993) 2969. D. Chen, A. Gronvold, H.P. Rebo, K. Moljord, A. Holmen, Appl. Catal., 137 (1996) L1. R. Wendelbo, D. Akporiaye, A. Anderson, I.M. Dahl, H.B. Mostad, T. Fuglerud, in preparation. D. Chen, H.P. Rebo, K. Moljord, T. Fuglerud, A. Holmen, in preparation. A. Gronvold, K, Moljord, T. Dypvik, A. Holmen, Stud. Surf. Sci. Catal., 81, 1994, 399. P. Dejaifve, A. Auroux, P.C. Gravelle, J.C. Vtdrine, Z. Gabelica, E.G. Derouane, J. Catal., 70 (1981) 123. G.F. Froment, W.J.H. Dehertog, A.J. Marchi, Catalysis (Royal Society of Chemistry), 9 (1992) 1. E.G. Derouane, J.P. Gilson, J.B. Nagy, Zeolites, 2 (1982) 42. T. Ishihara, H. Takita, Catalysis (Royal Society of Chemistry), 12 (1995) 21.

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Elsevier Science B. V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

167

D e a c t i v a t i o n and R e g e n e r a t i o n o f A l k a n e D e h y d r o g e n a t i o n Catalysts S. D. Jackson*, J. Grenfell, I. M. Matheson l, S. Munro 2, R. Raval 2, and G. Webb I ICI Katalco, R T and E Group, PO Box 1, Billingham, Cleveland TS23 1LB, UK 1 Dept. of Chemistry, The University, Glasgow, G12 8QQ, Scotland, UK 2 Leverhulme Centre for Innovative Catalysis, Liverpool University, Liverpool L69 3BX, UK

A study of the deactivation and regeneration of a Pt/alumina catalyst used for propane dehydrogenation has shown that initially the deposit is made up of C-1 fragments, which age to form polycyclic aromatics. Both surface hydrogen and carbon can exchange with the gas phase; however, some of the carbon cannot be removed from the metal function by treatment in oxygen at 873 K. As the catalyst ages it becomes increasingly selective to propene, indicating that the carbonaceous deposit has a major role in defining the selectivity and the activity of the catalyst.

1. INTRODUCTION The dehydrogenation of light alkanes to their respective alkene is an established technology [1-4]. The catalytic systems are based either on Pt/alumina or chromiaJalumina. The process conditions are typically atmospheric pressure and high temperature. This is to ensure a satisfactory yield of the alkene as the systems are run at, or near, equilibrium. However, as is common with high temperature hydrocarbon processes, all of the catalytic systems have a wellknown propensity to lay down carbon, indeed all the commercial systems have a regular regeneration sequence to ensure sustained activity [ 1-3]. Hence the nature and reactivity of the carbonaceous deposit is important with respect to optimization of the regeneration cycle. We have also shown previously that the carbonaceous deposit has a role in the generation of selectivity within the Pt/alumina system [5]. To determine the nature and reactivity of the deposit during the course of the reaction, we have used 13C and 2H labeled propane and propene in conjunction with FTIR and GCMS analysis of the surface deposit. We have also investigated the efficacy of the regeneration process in terms of its ability to remove the deposit and to reestablish the activity/selectivity profile of the catalyst. By using this combination of techniques we have been able to specify the nature and reactivity of the deposit in the first few seconds of the catalyst life, during the time-on-stream and at the point before regeneration. 2. E X P E R I M E N T A L Two reactor systems were used in this study. Pulsed reaction studies were performed in a dynamic mode using a pulse-flow microreactor system [6], in which the catalyst sample was placed in a glass-lined stainless steel tube, in a vertical position, inside a furnace. Using this system the catalysts (typically 0.3 - 0.5 g) could be reduced in situ in flowing 5% H2/N2 (50 cm 3 min -1) by heating to 873 K at 10 K min 1 and then holding at this temperature for 0.25 h. After reduction had ceased the catalyst was maintained, at the desired temperature, in flowing helium

168 (70 cm 3 min-1). The reaction gases were admitted by injecting pulses of known size (typically 0.18 cm 3, 0.101 MPa) into the helium carrier-gas stream and hence to the catalyst. A typical reaction cycle was reduction (H2/N2)/reaction (C3H8 pulses)/regeneration (02 pulses). The pulsing of oxygen was continued until no carbon oxides could be detected and the exit oxygen equaled the inlet. Adsorption, desorption, and reaction were followed using a GC. Continuous flow reaction studies were performed in a 0.101 MPa flow microreactor with the gas stream exit the reactor being sampled by on-line GC. The catalyst sample was placed in a glass-lined stainless steel tube, in a horizontal position between two quartz wool plugs, inside a furnace. Using this system the catalysts (typically 0.3 - 0.5 g) could be reduced in situ in flowing 5% H2/N2 (50 cm 3 min -1) by heating to 873 K at 10 K min 1 and then holding at this temperature for 0.25 h. After reduction had ceased the flow was switched to propane (50 cm 3 min "1) and the first analysis taken after 3 min, subsequent analyses were taken every 9 min. Catalyst regeneration was performed by flushing the reactor with N2 at temperature before switching the flow to 5% O2/N2. The system was held under 5% O2/N 2 at temperature until no CO2 was detected in the reactor effluent. The catalyst could be cycled through this process of reduction, reaction, and regeneration. The infrared spectra were obtained using a commercial FTIR spectrometer (Mattson Centauri). The studies were performed in transmission mode, with the catalyst in the form of a pressed disc, using an environmental cell where the temperature, pressure, and gas composition could all be controlled [7]. All spectra were recorded at 4 cm l resolution with the co-addition of 300 scans using an MCT detector. The catalyst used throughout this study was a 0.65% w/w Pt/alumina prepared by impregnation of ),-alumina with chloroplatinic acid. The dried sample was calcined in air at 823 K for 3 h.

3. RESULTS/DISCUSSION A catalyst was subjected to a series of 11 reaction cycles where the number of propane pulses was increased by one on each cycle; hence, in cycle 1, only one pulse of propane was passed over the catalyst, while cycle 10 had ten pulses of propane passed over the catalyst. Cycle 11 was a repeat of cycle 1 with only one pulse of propane passed over the catalyst. The data indicated that changes were occurring with respect to the amount of carbon deposited and the ability of the oxygen to remove it effectively. In Figure 1 it can be seen that the amount of carbon deposited increases with amount of propane passed over the catalyst; however, the amount of carbon removed by the oxygen treatment at 873 K, although increasing, increases at a slower rate. Also it can be seen that the O:Pt ratio decreases from 2 initially to 0.7 indicating incomplete reoxidation. Note that although the metal particle sinters between cycle 1 and cycle 2 the O:Pt ratio does not change, as it is measuring a bulk oxidation. Therefore the residual carbon species, which are not removed by the oxygen treatment at 873 K, are affecting the ability of the Pt crystallite to reoxidize. The effect is also seen in a change in the activity/selectivity profile of the catalyst with cycle number (Figure 2). It can be seen that the catalyst loses activity but gains in selectivity with each cycle, again reinforcing that the oxygen treatment at 873 K does not restore the catalyst to its original state. A fresh catalyst was reduced and subjected to pulses of [2-13C]C3H8 as the propane feed and the product gases analyzed by GCMS. Initially the catalyst produces solely methane as a gas

169 RATIO 10 9

9

9

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2

0

,

,

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,

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,

6 CYCLE NO.

_

~ ............. ,* , ,-F-.....

8

10

12

C:Pt DEP C:Pt BURN O:PT __m

o

.... $---

Figure 1. Ratios of (1) deposited carbon, (2) carbon removal, and (3) oxygen consumed per available Pt in pulse experiments.

2E+18 o

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t-O

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2

m CH4/10 D

3

4

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5 6 7 Cycle No. C2H6 m

8

9

C3H6 m

10

C3H8

Figure 2. Variation of Pulse 2 activity/selectivity with cycle number. phase product. In this situation the ratio of [13C]CH4"[12C]CH4 was found to be 1:2 indicating complete fracture of the propane to C-1 fragments and indicating a common surface intermediate. Hence the surface species has a C:H ratio of less than 1:2. When a labeled pulse was followed by a nonlabelled pulse an exchange process was observed (Table 1) such that 8.3% of the carbon- 13 that was deposited in the first pulse was removed by the second pulse. The typical time between pulses was 0.25 h. After the catalyst had been subjected to 12 pulses, with pulses 1, 10, and 11 labeled with carbon-13, aliquots of oxygen were passed over the catalyst. The CO and CO2 produced from each oxygen pulse was analyzed by mass spectrometry and a 12C'13C ratio determined. If full surface mixinghad occurred the 12C:13C ratio would be 11.7:1, if no mixing had occurred then the ratio could be either 89:1 or 2:1. The first oxygen pulse removed

170 9.7 ~tmol C and had a 12C:13C ratio of 11.3. As the final propane pulse only contained carbon-12 and deposited 15.5 ~tmol C, then, in the absence of surface mixing, the 12C:13Cratio should have been 89:1. Indeed the lowest 12C'13Cratio observed in the CO and CO2 was 10:1 and the highest was 22.4" 1 (Figure 3). These results indicate that there is mixing of the carbon on the surface. FTIR analysis of the surface deposit confirmed the absence of adsorbed C-3 species at temperatures above 673 K. Table 1 Exchange between surface carbon species Pulse Amount of Amount of C Amount of 13C C.]H8(a) deposited (b) deposited [2-13C]C3H8 9.2 12.6 4.2 C3H8 11.2 11.3 --C3H~ 11.2 10.7 --(a) units, lxmol propane (b) units, lxmol C

12C:13Cfor methane 2:1 33"1 89"1

Amount of 13C exchanged --0.4 0.0

When a catalyst, which was not fully reduced, was subjected to pulses of [2-13C]C3H8, it was found that methane, carbon monoxide and carbon dioxide were all produced. However when the isotope ratio of each of the components was examined it was found that the CO had a 12C:13C of 1:1, while the CO2 and the CH4 had a 12C:13C ratio of 2:1. This is a surprising result yet it indicates that the route to CO2 does not go through CO, and suggests that the totally fragmented propane yields either CH4 or CO2. To achieve a 1"1 ratio it is likely that the propane has split into a C-2 fragment and a C-1 fragment with the C-1 species being retained by the surface. If the C-1 fragment was to enter the gas phase, as any of the observed components, the isotope ratios could not be obtained. The C-2 fragment is then converted to CO. This behavior suggests that site responsible for this reaction sequence is different from the metallic site and that obtained during regeneration, when the 12C:13Cratio is the same for all the gas phase species at 2:1. To study the influence of hydrogen with respect to carbon laydown and to determine whether isotope exchange occurred, a catalyst was reduced in [2H]H2 and aliquots of a 1:1 [2H]C3H8/[1H]C3H8 passed over it at 873 K at 0.25 h intervals. The only gas phase product de-

O I-,< rY "-7 (.) ..

2O

15

10

d

1

2

3

4

5

6

7

8

9

OXYGEN PULSE NUMBER

Figure 3. Removal of surface residues such as carbon dioxide.

171 tected was methane; the results with respect to hydrogen balance and carbon laydown are shown in Figure 4. It can be seen that the [1H]hydrogen (protium) balance for each pulse is negative indicating loss of protium to the surface. However the [2H]hydrogen (deuterium) balance is positive for the first four pulses indicating that the catalyst has retained deuterium from the reduction, even though the catalyst was swept with helium for 0.5 h at 873 K after the end of the reduction. The effect of this retained deuterium was to reduce to zero the carbon laydown from the first pulse. Note that this effect is specific to a deuterium reduction, carbon laydown is detected from the first pulse of propane after reduction in protium [5]. Therefore, rather surprisingly, there is a kinetic isotope effect in the laydown of carbon in this system. In subsequent pulses, as the amount of free surface hydrogen decreased, the amount of carbon deposition increased. Analysis of the methane is shown in Table 2. The table also contains data from a catalyst that was reduced in hydrogen and was subjected to sequential pulses of [2H]C3H8 and [1H]C3H8. As can be seen the first pulse reveals the presence of retained hydrogen, however the methane produced from the second pulse has a higher level of d2 and d3 species than the methane from the first pulse. Given that the carbon-13 exchange experiments indicated that 8% of the deposit could be removed by a subsequent pulse, it is likely that the enhanced yield of d2 and d3 methane during the second pulse represents that proportion of the deposit from the initial pulse that can exchange. 80

o9

6O

O u

40

0~_ ~ 20 o ~_ 0 I -20

-40

I

I

1

2

I

I

3 4 PULSE NUMBER

.B AC PROTI LARABON NUM C E BALANCE .

I

5

"DEUTERIUM BALANCE I[

Figure 4. C/H balance. A catalyst and the alumina support were run through 100 cycles of reduction/reaction /regeneration in a continuous flow reactor, where the time on each segment was reduction 6 min, reaction 1 min, and regeneration 6 min. At the end of this sequence the oxidized catalyst and support were cooled and extracted with acetone. A yellow/brown deposit found just below the reaction zone in the reactors was also dissolved in acetone. The solutions were analyzed by GCMS and were found to contain a range of polycyclic aromatics. The principal polyaromatic produced in the reactor containing the Pt/A1203 catalyst was pyrene (C16H10), followed by methyl pyrene (CI7H12). The polyaromatics produced in the reactor containing the alumina gave

172 Table 2 Distribution of isotopic methanes Ca) Pulse % CD 4 % CDaH 1. C3Hs/C3D8 64 30 2. C3Hs/C3D8 22 41 3. C3Hs/C3D8 10 34 4. C3Hs/C3D8 7 29 5. C3Hs/C3D8 5 21 1. C3Ds(b) 2 12 2. C3HsCb) 2 15 (a) Catalyst reduced in deuterium at 873 K. (b) Catalyst reduced in protium at 873 K.

% CD2H2 5 34 43 43 36 42 45

% CDH,3 0 3 13 20 39 22 22

% CH 4 0 0 0 1 0 22 16

rise to a similar range of compounds, although the principal compound was perylene (C20H12), but the quantities produced were lower by an order of magnitude. Therefore although the support has a low level of activity, it is the metal that is the primary source of these compounds. It is clear from the molecular formula and the addition of C-1 fragments that these compounds are not formed from C-3 units but are built up by the addition of C-1 units. In agreement with the isotopic labeling experiments which indicated complete destruction of the C-3 unit. Therefore the reactive C-1 units identified above combine over a period of time (< 1 min under continuous flow conditions) to form highly unreactive polycyclic aromatics. It is likely that these are the species retained by the catalyst, even through the regeneration cycle, and hence limit the extent of reoxidation. Samples of the support and catalyst were subjected to propane (2020 Pa) at various temperatures in the FTIR environmental cell. By 673 K the principal bands observable are those due to carboxylate and aromatic species (Table 3). Bands due to C-H stretch were not observed, and the intensity of the bands obtained from the catalyst sample was greater than those from the alumina by approximately a factor of 2.5. Once again there is a similarity between the species present on the alumina and those on the catalyst but with the amounts being significantly higher on the catalyst. Table 3 Infra-red spectral assignments .Sample alumina 1660 1543 Pt/alumina 1558 Assignment carboxylate

Frequency cm -1 1500 1440 1500 1440

aromatic/ aromatic/ aromatic/ carboxylate carboxylate carboxylate

1325

1184

carbonate

carbonate

Although the polycyclic aromatics detected by GCMS would give strong C-H stretch bands in the gas phase [8], if they were adsorbed with the ring system parallel to the surface then the intensity may be below detection limits. However the FTIR spectra also revealed carbonate and carboxylate species on the alumina surface. These species will also contribute to the differential between the carbon retained by the catalyst and that removed under regeneration.

173 A catalyst was subjected to 2883 cycles of reduction/reaction/regeneration in a continuous flow identical to those detailed earlier. At the end of 500, 1000, 1500, and 2883 cycles the catalyst was reduced, exposed to 10 pulses of propane, and reoxidized. Each pulse was analyzed and the reaction products fitted against a model of the principal reactions. The results are shown in Table 4.

Table 4 Analysis of reaction products
C3H6+H2 1 3.48 10 13.34 1 4.08 10 10.71 1 6.91 10 15.42 1 12.56 10 15.07 6000 hl, Temp. 873 K.

% of Propane reactin~ to 3C+4H2 40.45 21.29 27.10 16.31 29.37 18.85 16.30 7.30

2CH4+C 8.60 0.00 1.69 0.00 4.56 0.00 1.72 0.00

% C2H4 reacting CHa+C2H4 17.22 13.72 7.55 5.40 12.37 8.82 9.51 8.77

to C2H6 100 60.54 100 27.10 91.44 37.89 33.67 12.63

Clearly the main carbon deposition reaction is C3H8 ---->3C + 4H2, while the hydrogenolysis reaction, C3H8 ---->2CH4 + C, is only of importance in the early stages of the catalyst life. It can also be observed that the formation of ethane is sensitive to the amount of hydrogen produced in the system. These results indicate that there are sites, which adsorb propane, and catalyze C-H and C-C bond breaking. However these sites are only slowly deactivated. Therefore the carbon deposited on them is either easily removed by the regeneration treatment, or the carbon may diffuse away from the site. 4. CONCLUSIONS The deactivation and regeneration of a Pt/alumina catalyst used for the dehydrogenation of propane has been studied. Initially the deposit is made up of identical C-1 fragments having a C:H ratio of 1:1 or 1:2, with the overall carbon deposition reaction being represented by C3H8 ---> 3C + 4H2. The carbon deposition reaction represented by C3H8 ---> 2CH4 + C is only of importance early in the reaction cycle. As well as the C-1 species the carbonaceous deposit also contains polycyclic aromatics, that are formed by ageing of the C-1 species. These compounds are formed principally on the metal but a small amount can be formed by the support. FTIR analysis of the deposit also indicates the presence of carbonate species. Approximately 8% of the carbon deposited can exchange with the gas phase after 0.25 h, no exchange is observed at later times. Therefore the exchangeof carbon is limited to the C-1 fragments. With increasing carbon deposition the O:Pt ratio obtained during regeneration decreases implying that some of the carbon is resistant to removal from the metal function by treatment in oxygen at 873 K. Removal of isotopically labeled carbonaceous deposit from the surface reveals that there is surface mixing of carbon deposited at different times. The hydrogen associated with the deposit can exchange with both the surface and gas phase in both the fresh and aged forms. With time on stream and with increasing reaction cycle the catalyst becomes increasingly selective to propene,

174 indicating that the carbonaceous deposit has a major role in defining the selectivity as well as the activity of these catalysts.

REFERENCES

1. 2. 3. 4. 5.

6. 7. 8.

G. Homady, F. Ferrel, and G. Mills, Adv. Petrol. Chem. Refin., 4 (1961) 451. Phillips Ltd., US Patent 4,536,196 (1985). UOP Ltd., US Patent 4,716,143 (1986). P. Biloen, F. M. Dautzenberg, and W. H. M. Sachtler, J. Catal., 50 (1970) 77. S.D. Jackson, J. Grenfell, I. M. Matheson and G. Webb, in "Catalyst Deactivation 1994", Studies in Surface Science and Catalysis Vol. 88, Eds. B. Delmon and G. F. Froment, p.297 (1994). S.D. Jackson, P. Leeming, and J. Grenfell, J. Catal., 150 (1994) 170. S Munro, Ph.D. Thesis, University of Aberdeen, (1996). J. Szczepanski and M. Vala, The Astrophysical Journal, 414 (1993) 646.

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Elsevier Science B.V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

175

A N o v e l M e c h a n i s m of Catalyst Deactivation in Liquid Phase Synthesis G a s - t o - D M E Reactions X. D. Peng*, B. A. Toseland and R. P. Underwood Air Products and Chemicals, Inc., 7201 Hamilton Boulevard, Allentown, PA 18195-1501 Commercial methanol synthesis catalysts and ),-alumina are stable when used separately in a slurry phase reactor for the methanol synthesis and the methanol dehydration reaction, respectively. However, when they are used as a physical mixture for our slurry-phase synthesis gas-to-DME process, both catalysts deactivate rapidly. Traditional causes of catalyst deactivation, including hydrothermal sintering, leaching, coking, and poisoning, were investigated and subsequently ruled out. Detrimental interaction between the two catalysts was identified as the cause of the rapid, simultaneous deactivation of both catalysts. Intimate solid-state contact between the two catalysts is necessary for this interaction to take place. Most likely, the interaction is due to the migration of Zn- and Cu-containing species from the methanol catalyst to ?-alumina. Although not discussed in this paper, a proprietary dual catalyst system with good stability has been developed. 1. INTRODUCTION In a Liquid Phase Di-Methyl Ether process (LPDME), synthesis gas (syngas) is converted into dimethyl ether (DME) in a single slurry phase reactor over a catalyst system. Both methanol synthesis and methanol dehydration function as a physical mixture of a methanol synthesis catalyst and a dehydration catalyst (dual catalyst system). Three reactions take place simultaneously in the system, namely: Methanol synthesis reaction:

CO + 2H2 CH3OH

(1)

Water gas shift reaction:

CO + H20 CO2 "t- H2

(2)

Methanol dehydration reaction:

2CH3OH CH3OCH3+ H20

(3)

where the methanol synthesis catalyst catalyzes reactions 1 and 2, and the dehydration catalyst catalyzes Reaction 3. The dehydration reaction drives the methanol synthesis reaction away from unfavorable equilibrium restrictions. Water decreases the rate of dehydration, and is formed in the dehydration reaction, but it is consumed by the water gas shift reaction. Furthermore, hydrogen formed by the water gas shift reaction makes the reaction environment more favorable for methanol synthesis. This synergy results in high syngas conversion per pass and low cost DME production. The second feature of this process is its efficient heat management due to the presence of an inert liquid medium and the good mixing in a slurry-bubble-column reactor. The heat generated from highly exothermic methanol synthesis reaction can be sufficiently dissipated to avoid the formation of hot spots, which are detrimental to the methanol catalyst. The efficient heat transfer makes it possible to

176 use CO-rich syngas directly from a coal gasifier [1]. The process can be incorporated in power plants using the integrated gasification combined cycle for peak shaving [2]. DME itself is a useful product as a potential alternative diesel [3] or cooking fuel [4], and as an interesting feedstock in the production of chemicals such as olefms and vinyl acetate [5]. A typical dual catalyst system studied in our lab consisted of a commercial, Cu-based methanol synthesis catalyst and T-alumina. While all the merits of this process were verified, the study showed that both catalysts suffered rapid deactivation under LPDME conditions. The problem seems common to many other dual catalyst systems containing different dehydration catalysts. This paper focuses on our investigation of the cause and the mechanism of the catalyst deactivation in this dual catalyst system. 2. E X P E R I M E N T A L The reactions were carried out in 300 cc autoclave reactors. Syngas was passed through a carbon trap before it entered the reactor to remove Fe and Ni carbonyls, both known poisons to methanol catalysts. The inlet and exit gases were analyzed by on-line gas chromatography. The BASF $3-86 Cu/Zn/A1/O-based methanol synthesis catalyst was used in all experiments. ),-alumina was prepared from Catapal B alumina (boehmite) by calcination in air in a muffle oven at 500 C for 4 hours. In a typical LPDME experiment, a powder mixture of the methanol catalyst and ),-alumina was loaded in the reactor along with a hydrocarbon oil (Drakeol 10). The typical weight of the oil was 120 grams, and the catalyst loading ranged from 10 to 30 grams. The standard methanol-to-dehydration catalyst ratio was 80:20 by weight. The methanol catalyst was activated, in the presence of ),-alumina, by a gas consisting of 2 mol% H2 in N2. The activation was carried out from room temperature to 240 C over a period of 23 hours. The system was then switched to the syngas feed and brought to the reaction conditions within an hour. The standard reaction conditions were 250 C, 750 psig, 6,000 GHSV. The reactor was stirred at 1,200 rpm. All operations were free of mass transfer limitations and the reactors operated as an ideal continuous stirred tank reactor. Catapal B ),-alumina doped with copper, zinc, or both copper and zinc were used for mechanistic studies. All three samples were prepared by incipient wetness impregnation, with Zn(NO3)2 and Cu(NO3)2 as precursors, followed by calcination in air at 500 C for 4 hours. 3. R E S U L T S AND DISCUSSION Figure 1 shows the catalyst activity as a function of time on stream for an LPDME run. The results from an LPMEOH run and a dehydration run are also shown, in which only the methanol catalyst or "t-alumina was used. All runs were carried out in a 300 cc slurry phase autoclave using powdered catalysts. The syngas used in both LPDME and LPMEOH run simulates that derived from a Shell-type coal gasifier. A 10% methanol in N2 feed gas was used for the dehydration run. For the LPMEOH and LPDME runs, the stability of the catalysts is expressed in terms of the apparent reaction rate constants which are normalized to the initial values are calculated from the rate expressions shown in the figure (where subscript m refers to the methanol synthesis reaction, d to the methanol dehydration reaction, and appr. stands for the approach of the reaction to equilibrium, respectively.) They serve as a good semi-quantitative kinetic tool in the domain of the reaction conditions used in our experiments.

177

For the dehydration run, the normalized DME productivity was used to describe the stability of the alumina. It is clear from the figure that both the methanol catalyst and the alumina are much less stable under the LPDME conditions than when they are used separately in methanol synthesis and methanol dehydration reactions. 300 cc autoclave, 250 C, 750 psig, 6,000 GHSV 30%H 2, 66%CO, 3%CO 2, 1%N 2 (Shell gas)

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Time on stream (hr)

Figure 1. Catalyst deactivation in different liquid phase processes. 3.1. Investigation of the Possible Causes of Catalyst Deactivation The differences between LPDME and LPMEOH were identified. First, the DME concentration is much higher in LPDME, typically 6 mol% vs. <0.1% for LPMEOH. DME is known to disintegrate the texture of a methanol catalyst [6]. It may also cause leaching of the active components out of the methanol catalyst, leading to methanol catalyst deactivation. Second, the water concentration in LPDME is generally higher than that in LPMEOH due to the dehydration reaction. Depending on syngas composition, it ranges from 0.06 mol% to 1.6% in LPDME compared to 0.03%-0.2% for LPMEOH. The higher water level may cause hydrothermal sintering of the methanol catalyst, leading to the more rapid deactivation of the methanol catalyst in LPDME. Third, coking by heavier oxygenates, such as higher alcohols, ethers and esters, and, possibly, some nonvolatile species, may deactivate 7-alumina. Fourth and finally, the co-presence of the two catalysts may be the problem. We investigated each one of these factors individually as the possible cause of catalyst deactivation in LPDME. The details of all these experiments have been given previously [7]. DME, water, and coking were ruled out as the cause of the more rapid deactivation of the catalysts in the LPDME process. The cause was revealed by the following two experiments.

178

3.l.a. The Catalyst Compatibility Experiment. This experiment was designed to check if the compatibility of the two catalysts could be the problem. A standard catalyst mixture was loaded into a 300 cc autoclave. Following a normal catalyst reduction (using 2% H2 in N2), this catalyst system was left under flowing reduction gas (50 sccm/min.) at reaction temperature (250 C) and pressure (750 psig) for 117 hours with normal stirring. After the holding period the activity of the catalyst system was measured using the simulated Shell syngas. It was found that the methanol rate constant dropped by 68% and the dehydration rate constant by 62%. Control experiments of holding the methanol catalyst only or a mixture of the methanol catalyst and a silica gel (Davison, grade 55, zero dehydration activity) in the reactor under the same conditions did not cause any deactivation of the methanol catalyst. Furthermore, this experiment avoided exposure of the catalysts in the holding period to syngas and LPDME reaction products. Therefore, the deactivation of the two catalysts in this specific experiment can only be attributed to the incompatibility of the two catalysts. Note that this experiment is not sufficient to pinpoint the incompatibility as the cause of catalyst deactivation under LPDME conditions, since the gas phase composition in LPDME is very much different from 2% H2 in N2. But it does show the likelihood of a detrimental interaction between the two catalysts.

3.l.b. The Experiment Using Robinson-Mahoney Basket Internals (R-M Experiment). This experiment was conducted by submerging Robinson-Mahoney basket internals in the mineral oil in a 300 cc autoclave. Loaded inside the basket was a mixture of the pellets of the methanol catalyst and "t-alumina (1 to 3.4 mm size). While the basket was kept stationary during the run, the oil was agitated to provide the mixing. The original goal of the experiment was to obtain individual methanol catalyst and alumina samples for analysis so that some insight into the deactivation mechanism could be obtained. The experiment, however, happened to provide a set-up that contains all the components, except the intimate contact among the micron-size powders of the two catalysts, in a normal LPDME run.

The surprising finding from this experiment is the lack of methanol catalyst deactivation after 504 hours on stream. The "f-alumina lost 38% of its initial activity, but only within the first 130 hours. The first thing we learned from this experiment is that all of the reactants or products can be ruled out, either individually or synergistically, as the cause of catalyst deactivation in the normal LPDME run, since the gas phase composition in this run was similar to that in a normal LPDME run. More generally, we can conclude that nothing in the gas and the liquid phase is the cause of the deactivation of the methanol catalyst under LPDME conditions. Secondly, since this experiment very much eliminates all other possible causes, the detrimental interaction between the two catalysts becomes the only conceivable possibility left. This cause has been suggested by the catalyst compatibility experiment. The lack of catalyst deactivation in the R-M run also goes along well with the hypothesis of the detrimental interaction. A major difference between this run and a normal LPDME run is the contact between the two catalysts. Instead of the point contact among millimeter-size pellets in the R-M run, more intimate contact between the two catalysts is available in a normal LPDME run. This is because mixing, collision and agglomeration can bring the particles of the two catalysts in contact, and the micron size of the catalyst

179 particles provides large contact surface area. The lack of catalyst deactivation in the R-M run can be explained by the lack of intimate solid-state contact between the two catalysts, assuming that the intimate contact is necessary for the detrimental interaction to take place. In summary, the results from all these phenomenal studies can be forged into one self-consistent explanation: the more rapid deactivation of the methanol catalyst and y-alumina in LPDME is caused by the detrimental interaction between the two catalysts, and the intimate, solid-state contact of the two catalysts is necessary for the interaction to take place. 3.2. Mechanistic Studies

The most probable mechanistic interpretation for the detrimental interaction between the two catalysts is inter-catalyst migration. Cu- and Zn-containing species may migrate from the methanol catalysts to the alumina, causing the loss of dehydration activity via. poisoning. The migration also results in the deactivation of the methanol catalyst because of the loss of its active components. The reverse process may also take place, i.e., migration of Al-containing species from the alumina to the methanol catalyst, to poison the active sites for methanol synthesis. Different forms of copper have been de'tected under the methanol synthesis conditions, including large metallic crystallites, small metallic clusters, and copper (e.g., Cu +) dispersed in ZnO [8]. The migration through surface diffusion of Cu, Cu + and ZnO is known in the literature [9-14]. All these species are very likely also present under LPDME conditions. Any one of them could serve as the sources of inter-catalyst migration, y-alumina, a widely used catalyst support, is known to have great dispersing capability for metals, metal oxides, and salts [ 10]. Under LPDME conditions, "t-alumina may serve as a chemical potential sink that attracts these mobile species, which otherwise would undergo much slower sintering processes under pure methanol synthesis conditions. In a word, we believe that it is the dispersion of mobile species onto ),-alumina in LPDME, verses copper or ZnO sintering in LPMEOH that results in the more rapid deactivation of the methanol catalyst under LPDME conditions. Poisoning of the dehydration sites on alumina by Cu- and Zn-containing migrating species is supported by LPDME experiments using Cu- or Zn-doped Catapal B )'-alumina. Alumina doped with 13.6 wt% of ZnO exhibited only 30% of its original activity, while doping alumina with 5 wt% of CuO dropped the activity by 35%, judging by the dehydration rate constant. The alumina particles in the spent catalyst mixtures were analyzed for Cu and Zn by SEM and EDS to further check the migration hypothesis. SEM enables us to pick out individual alumina or methanol catalyst particles from the spent powder mixtures of the two catalysts, while EDS provides elemental information of these particles. The resolution of SEM and EDS is around 1 micron, and the lower detection limit of EDS is 0.5 wt%. Figure 2 illustrates how a typical analysis was conducted. An SEM micrograph of an alumina particle, a cross section embedded in epoxy in this case, was generated first. Then the electron beam was parked at different spots in this cross-section and an EDS elemental analysis was performed. The sampling volume in this mode of analysis was an order of magnitude smaller than the dimension of alumina particles.

180

Fila~ure 2. The cross section of an alumina particle in a spent catalyst mixture. I

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181 surface where no physically attached methanol catalyst particles were visible. Cu and Zn were also detected in the cross-section of the alumina particles (Figure 2c). In contrast, no Cu and Zn were found in the outer surface and the cross-section of the alumina used in the R-M experiment. A traditional elemental analysis using chemical digestion and atomic absorption also showed no zinc and copper build-up in this sample. 1.1

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Figure 4. Zn-to-Cu EDS peak ratio of different samples. ( . ) Pure, freshly reduced methanol catalyst particles; ( 9) methanol catalyst particles in a spent catalyst mixture (714 hr); (A) methanol catalyst fines on the surface of the alumina particles in the spent mixture; ( A ) clean spots on the outer surface of the alumina particles in the spent mixture, solvent dispersed; (o) clean spots on the outer surface of the alumina particles in the spent mixture, dry; ( o ) cross section of the alumina particles in the spent mixture; ( + ) cross-section of the alumina particles in a freshly reduced catalyst mixture; ( x ) Cu- and Zn-doped y-alumina, as prepared. We also looked at the ratio of the Zn-to-Cu EDS peak area from different alumina samples as a way to distinguish the Cu and Zn by migration from those picked up from the methanol catalyst background. Figure 4 shows that the Zn-to-Cu ratio is very consistent (around 0.29) from different methanol catalyst samples, including a pure, freshly reduced methanol catalyst, the methanol catalyst particles in a spent mixture, and the methanol catalyst fines attached to the outer surface of alumina particles in the spent mixture. In contrast, the Zn-to-Cu ratio is very scattered from the alumina particles in the spent catalyst mixtures. More importantly, the scatter from the catalyst mixture that was 714 hours on stream is around 0.38, significantly greater than that from the methanol catalyst. While one could argue that this scatter is due to the poor signal-to-noise ratio of the EDS spectra (cf. Fig. 3c to e), the results from the Cu- and Zn-doped alumina sample indicates

182 that this may well be the inhomogeneous nature of deposited Zn and Cu on "~-alumina("x" in Fig. 4). Figure 4 shows that the ratio is very scattered, although the EDS spectrum from this sample has a good signal-to-noise ratio (Fig. 3b). These results suggest that the Cu and Zn detected from the spent alumina are likely to be from migration. No definitive conclusion can be made yet because we cannot firmly rule out the possibility that the detected copper and zinc are from the background methanol catalyst. Moreover, the poor signal-to-noise ratio does not allow us to look at the correlation of Cu and Zn accumulation in alumina to the loss of its activity. Efforts are being made to circumvent the background interference by analyzing the thin film prepared from spent alumina particles with TEM. Other mechanisms that could also explain the detrimental interaction between the two catalysts include: 1) the solid state reaction between ~,-alumina and ZnO from the methanol catalyst to form bulk mixed oxides, and 2) the formation of physical agglomerates from the fines of the two catalysts. These are under investigation. REFERENCES

1. X.D. Peng and B. A. Toseland, in: Proceedings of the Second China/USA Joint Chemical Engineering Conference (Beijing, China, 1997). 2. D.M. Brown, B. L. Bhatt, T. H. Hsiung, J. J. Lewnard and F. J. Waller, Catalysis Today, 8 (1991) 279. 3. A. M. Rouhi, C&E News, 29 May 1995, p. 37. 4. Z. Chen and Y. Niu, Coal Chemical Technology, 67 (1994) 1. 5. B.A. Toseland, R. P. Underwood and F. J. Waller, in: Proceedings of 1994 DOE Contracts' Review Conference, p. 307. 6. M. Kuczynski, Ph. D. Thesis, Twente University, Netherlands, 1986. 7. X.D. Peng, B. A. Toseland and R. P. Underwood, in: Proceedings of 1995 DOE Contracts' Review Conference, p. 371. 8. G. Ghiotti and F. Boccuzzi, Catal. Rev.- Sci. Eng., 29 (1987) 151. 9. J.T. Sun, M. Sahibzada and I. S. Metcalfe, in "The 1996 ICHEME Research Event/Second European Conference for Young Researchers", p. 466. 10. Y. C. Xie and Y. Q. Tang, Adv. Catal. 31 (1990). 11. G. L. Price, V. Kanazirev and D. F. Church, J. Phys. Chem. 99 (1995) 864. 12. Y. Kanai, T. Watanabe, T. Fujitani, M. Saito, J. Nakamura and T. Uchijima, Catalyst Letters, 27 (1994) 67. ~ 13. Y. Kanai, T. Watanabe, T. Fujitani, T. Uchijima and J. Nakamura, Catalysis Letters, 38 (1996) 157. 14. D. S. King and R. M. Nix, J. Catal., 160 (1996) 76.

9 Elsevier Science B.V. All rights reserved Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

183

Activity, Selectivity and C o k i n g o f Bimetallic N i - C o - S p i n e l Catalysts in Selective H y d r o g e n a t i o n Reactions. J.C. Rodriguez 1, C. Guimon 2, A.J. Marchi 3, A. Borgna 3 and A. Monz 6n 1. 1 Department of Chemical Engineering.Faculty of Science. University of Zaragoza. 50009 Zaragoza. Spain. e-mail: [email protected]. 2 L.P.C.M. URA-CNRS 474. Avenue de l'Universit6. 64000 Pau. France. 3 INCAPE. FIQ-UNL-CONICET. Santiago del Estero 2654. 3000 Santa Fe. Argentina. The influence of the calcination and reduction temperatures of a Ni-Co-ZnA1204 catalyst was correlated with its catalytic activity in the hydrogenation of acetylene. A well interdispersed Ni-Co catalyst supported in a ZnA1204 spinel-like structure was obtained by using a coprecipitation method. Cobalt apparently effects a dilution of Ni surface ensembles, increasing the selectivity to ethylene. The influence of the operating temperature on activity, coking rate and selectivity was analyzed using a deactivation kinetic model. 1. INTRODUCTION. NiO-A1203 oxides are, after a reduction step, the usual catalysts in reactions such as: steam reforming of methane or light hydrocarbons [ 1], production of synthesis gas [2], and selective hydrogenations [3,4]. Impregnation and coprecipitation are widely used in the preparation of these catalysts. The coprecipitation method typically produces a solid solution of mixed oxides with a spinel-like structure [5-7]. Moreover, the high temperature conditions of the steam reforming, syngas production and some hydrogenation processes requires the use of catalysts with good thermal stability such as spinels or other types of mixed oxides. Besides Ni +2, other transition metals including Zn 2+ or Co 2+, also can form aluminates with an inverse spinel structure, some at lower temperatures than Ni 2+ spinels (i.e. 773 K or lower) [7,8]. The final catalyst in some of these processes usually consists of a very well dispersed metal phase in a mixed spinel matrix. The attainment of a well-dispersed metal phase may be important in the case of reaction structure sensitivity. While acetylene hydrogenation is a structureinsensitive reaction, acetylene hydrogenolysis is a structure-sensitive reaction that occurs with the hydrogenation reaction to produce undesirable products like coke and methane. Thus, acetylene hydrogenation consists of a complex reaction network with apparent structure sensitivity [4]. The catalytic behavior of reduced NiO-AI203 mixed-oxide spinels for selective hydrogenation has been studied by our research group for several years [4-6]. These materials have been successfully modified with structural (Zn, Cr) or catalytic (Cu, Co) promoters in order to obtain higher activities and selectivities for the desired products and to decrease coking reactions [4]. This study focused on the effects of calcination and reduction temperatures of a bimetallic Ni-Co catalyst on composition and extent of reduction of its surface and bulk phases and on its catalytic activity in the hydrogenation of acetylene. Furthermore, the effects of operating temperature on activity, coking rate and selectivity were studied and analyzed by means of a kinetic model that simultaneously fits activity and deactivation rate data.

184 2. EXPERIMENTAL. The catalyst was prepared by coprecipitation at a constant pH of 7.2+0.2 and a temperature of 333 K. Decomposition and calcination of the dried precursor (hydrotalcite-like structure) was performed in a N2 atmosphere at 773 K or 873 K for 14 hours, giving a solid with a non-stoichiometric spinel-like structure of the following nominal composition: (NiO)0.5(CoO)0.5ZnA1204. XPS (X-ray Photoelectron Spectroscopy) spectra were recorded with a Surface Science Instrument (SSI) spectrometer, using A1Ktx radiation. The C 1s band at 248.9 eV was used as internal standard. The system was equipped with an in-situ chamber in which pretreatment (reduction at different temperatures) of samples was carried out before analysis. Temperature Programmed Reduction(TPR) experiments were carded out in a separate unit using a total flow rate of 20 cm3/mincontaining 5% H2 in N2, at a heating rate of 10 K/min. SEM (Scanning Electron Microscopy) micrographs of catalysts before and after reaction were obtained with a JEOL JSM 6400 microscope. Solids were covered by electrodeposition with a thin gold film to improve sample conductivity. Reaction tests were carried out in a thermogravimetric system, operated as a differential reactor. Prior to the kinetic tests, catalyst samples were reduced "in situ" for 3 hours at temperatures ranging from 673 K to 873 K. Hydrogenation of acetylene was performed under the following conditions: total flow rate = 700 cm'/min, mass of catalyst = 200 mg, pressure = 1 bar, feed composition (% vol.) (H2/C2H2/N2)= 45/15/40, temperature = 448-498 K. 3. RESULTS AND DISCUSSION. 3.1 Influence of Calcination and Reduction Temperatures on Catalytic Performance. Binding energies of the different metals contained in the catalyst after calcination at 773 or 873 K results are presented in Table 1. No fundamental differences were found at either temperature. The values of B.E. obtained for Ni and Co are those typically reported for NiO+NiA1204 and well-defined COA1204 structures, respectively [7,8]. Nevertheless, the results shown in Table 2 (first and fourth rows) clearly indicate that an increase in the calcination temperature preferentially favors the incorporation of Co +2 cations into the spinel lattice, increasing the relative amount of nickel in the catalyst surface. This can be related to the greater tendency of cobalt to form a well-defined bulk spinel phase [5].

Table 1 Bindinl~ energies of the metallic cations. Element T calc. = 773 K T calc. = 873 K B.E. (eV) (%) B.E. (eV) (%) A1 (2p) 74.40 100 74.40 100 856.03 62.50 Ni (2p3/2) 856.20 58.2 862.00 41.8 862.02 37.50 781.39 70.88 Co (2p3/2) 781.45 67.2 784.17 13.2 783.60 14.71 787.25 14.41 787.58 19.6 1022.37 100 Zn (2p~t2) 1022.48 100

After reduction at different temperatures, XPS and TPR results [4], showed equal percentages of reduction of Ni ~ and Co ~. The similarity of the reduction temperatures of Co and Ni, and their high degree of interdispersion in the original precursor of the catalyst apparently favors a strong synergy between metals. This is supported by a shift of-0.7 eV in the binding energy of Co ~ (778.1 eV), with respect to the value obtained for this element in a monometallic catalyst Co-Zn-A1 prepared for comparison purposes (778.8 eV). However, reaction tests revealed only minor differences in the performance of the samples of catalyst calcined or reduced at different temperatures (ethylene selectivities equal to 0.75-0.8, coke concentrations after 3 h of reaction equal to 7.5-9.5 mg/100 mg cat. (results not shown) and similar reactant conversions). In fact, XPS shows that, in spite of the increase in the

185 percentage of reduction of Co and Ni, as the temperature of reduction increases, sintering becomes more important (Table 2, first and second columns). The coupling of both processes apparently leads to similar quantities of available surface metal, and therefore to similar catalytic behavior. The only exception is the sample of catalyst calcined at 873 K and reduced at 673 K. The activity of this sample for the main reaction (ethane and ethylene production) is lower than that expected from the total amount of reduced metal in the surface (see Table 2, last column). Table 2 Not only the type of Surface atomic ratios of Ni and Co metals. Catalyst metal (cobalt or nickel), but calcined and reduced at several temperatures. its accessible fraction, Tcalc.= 773 K determine the catalytic Tred. NiT/A1 COT/A1 Ni~ ~ Ni~176 behavior. All the catalyst unreduced 0.213 0.153 0 samples, with the exception 673 K 0.133 0.106 1.51 0.066 of that reduced at 673 K, 773 K 0.080 0.054 1.79 0.075 have similar amounts of Tcalc.= 873 K metal and Ni~ ~ ratios in Tred. NiT/A1 COT/A1 Ni~ ~ Ni~176 the surface, and thus, exhibit similar activities for unreduced 0.187 0.076 0 the main and coking 673 K 0.126 0.078 0.32 0.011 reactions. The sample 773 K 0.101 0.058 1.92 0.084 reduced at 673 K, on the 873 K 0.083 0.047 1.92 0.082 contrary, is unusually enriched in cobalt leading to a near zero activity for the main reaction, but this 12 apparently does not affect 12 ' ' ' ' ' ' 1 ' ' ' ' ' ' m Tea =773 K, T ~ t = 6 7 3 K the coking reaction (7.5 mg 10 of coke/100 mg cat. after 3 10 9 Tc~=873K, T,=t=773K h of reaction) (results not o T cai=873K, Tract=873K R shown). This result, that is 8 fully analogous to the behavior observed with the )= monometallic catalyst (CoZn-A1), can be explained by 4 o considering the differences in the adsorption strength .~ Tqzr.=448 K 2 of reactant molecules on 2 {_ H2/C2H2=3/1 nickel and cobalt sites. On cobalt atoms, acetylene is 0 4000 8000 12000 0 4000 8000 12(~) '0 strongly adsorbed leading to Time (s) an almost stoichiometric reaction (this adsorption Figure 1. Influence of the reduction temperature on the could explain the non-zero ethane and ethylene yields. coke production when the Co-Zn-A1 catalyst is used). On the nickel atoms, reaction takes place, favored by the weaker reactant adsorption. Considering also that cobalt has a dilution effect on the Ni crystallites, the combination of both effects leads to facts leads to an increased selectivity to the desired product (ethylene) [9,10].

186 3.2 Influence of the Operating Temperature on Catalytic Performance. The influence of the operating temperature on the activity, selectivity and coking rate were studied for a sample of catalyst calcined and reduced at 773 K. An increase in the operating temperature from 428 K to 498 K led to a greater initial coking rate, a higher f'mal coke content, a lower initial conversion and a more selective catalyst (Figures 2, 3 and 4). In Figure 2, it is evident that an increase in temperature produces an increase in the coking rate, and thus, in the final coke concentration. Independent of temperature, coke content sharply increases during the initial stage of the reaction, and the coking rate progressively decreases with time, to reach a pseudo-steady state, in which coke grows at a con~ 15 stant rate. These two reaction O O regimes could be related to ,r o o o o two different kinds of coke on the catalyst surface. Thus, t.on the bare metallic surface, 8 coking progresses quickly, by ~ 5 hydrogenolysis of acetylene 1) 428 K 4) 488 K 8 molecules, and methane 2) 448 K 5) 498 K production (results not 3)468 K shown). During this stage, , 0100 ~ I , I ~ I , I , I 00 2 4000 60(X) 8000 1 0 0 0 0 12000 filaments are nucleated and "Sme (s) the metallic surface is covered with a layer of coke causing a Figure 2. Influence of the operating temperature on the pseudo equilibrium to be coke deposition. Line: experimental data ; Symbols: data reached that depends on the predicted by model. operating conditions (partial pressures of reactants and temperature). This fast deactivation also determines 14 the initial level of catalyst 12 activity. Thus, the higher the temperature, the lower the 0 0 10 initial activity of the catalyst A & (Figure 3) because of the v 8 -o higher initial coking rate ~) BB B ~ ~ m~ ~ ~ ~ F m (Figure 2). During the second 6 period of coking, that 9 428K #o 448K 4 constitutes most of the 9 468K reaction period, growing 488K 9 496K carbon filaments (see Figure 5) accumulate, while the , I , I , I , I , I 00 2000 4000 6000 8000 1 0 0 0 0 12000 modification of the metallic surface by acetylene ~me (s) hydrogenolysis and filamental nucleation is less Figure 3. Influence of the operating temperature on important. The negligible total gas yield (ethane + ethylene + methane). effect of filamentous coke on the activity of metallic 20

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187 particles which remain at the filament tip is well established (11). In fact, no important decrease in catalyst activity with reaction time is evident within the temperature interval studied (see Figure 3) (even, after 30 h of reaction at 448 K, catalyst activity remained constant). It can be concluded that the first stages of the reaction determine the state of the metallic surface, i.e. the percentage of bare sites for the main reaction, and hence, the overall performance of the catalyst. Two types of coke 1.0 . . . , 9, 9, , , , , 1.0 deposits (filamentous carbon and amorphous 0.9 0.9 k coke) were evident from 0.8 0.8 SEM. The process of filament nucleation gives 0.7 0.7 methane as a secondary reaction product. The increasing rate of encap0.5 - - ~, 0.5 sulation of the metallic 0.4 I~ 0.4 crystallites by carbon during this stage is re9 o~ sponsible for the negative 0.2 0.2 value observed for the 0.1 activation energy for the 010 4000 8000 12000 0.00 0.05 0.10 0.15 0.20 main reaction (Figure 3). On Time (s) Coke Content (mg/mg cat) the other hand, as mentioned previously, coke Figure 4. Influence of the operating temperature on the deposition causes an ethylene selectivity. incremental increase in ethylene selectivity (Figure 4) and a parallel decrease in the selectivity to ethane and methane. Thus, more than activity, coke affects product distribution with time. The mechanism developed by Thomson and Webb [10] to explain the hydrogenation of acetylene is presently commonly accepted with slight modifications [9]. It is proposed that ethane can be produced by two different routes, i.e. direct hydrogenation of acetylene on the metallic bare sites, via the organometallic-likeintermediate called ethylidine, or in successive hydrogenations. Ethylene is easily formed, even on the metallic sites covered by coke, via a hydrogen transfer mechanism from the coke layer to the second adsorption layer. Methane is preferentially formed on the bare metallic surface, and is probably a by-product of the nucleation of carbon filaments. Taking into account this mechanism, the incremental increase in selectivity to ethylene with reaction time could be a result of the progressive change of the catalyst surface. The presence of a well-dispersed cobalt-nickel active phase, in which reactants are adsorbed more strongly on cobalt than on nickel, probably leads to the abrupt increase in the selectivity to ethylene, i.e. the reaction product that requires smaller metal ensembles to be produced [9,10]. This fact explains the results shown in Figure 4, where it can be seen that the ethylene selectivity and coke levels are almost independent of the operating temperature. Thus, coke concentration determines the product distribution.

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3.3. Kinetic Modeling of Activity and Coke Coverage. A kinetic model based on other previously proposed models [12], was developed to simultaneously fit activity and coke coverage. From the above results, for the present catalystreaction system, coke apparently causes a progressive decrease in the hydrogenation capacity. Consequently, activity was calculated as the quotient between the rates of hydrogen consumption (to ethane, ethylene and methane) at a given time, and at zero time.

188 The model proposed assumes reversible formation of amorphous coke (Ccm) and simultaneous formation of whiskers (Ccw) on the catalytic surface. The amount of monolayer coke at infinite time (Ccmax) depends on the operating conditions (e.g. temperature) and is governed by values of the rate constant for

dCc m rcm -

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kso (S-1) "103 kRo (s "1) "102 kwo (s'l)'106

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5.242 _+0.38 8.428 +_0.89 5.29+ 0.268

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Eas(cal/mol) EaR(cal/mol) Eaw (cal/mol) Ccm~ (mg/mg cat)

Figure 5. SEM micrograph of coke formed on the catalyst surface after 3h reaction time.(X 4,300). Temp. Reaction = 448 K.

Parameter

8271 + 776 4634 + 923 4887 + 435 0.129 + 0.004 monolayer coke formation (ks) and that for coke removal (kR) (Eqn. 1). Coke removal takes place by reaction with hydrogen, generating other nonanalyzed reaction products (e.g. C4+) [9]. The growth rate of filamentous coke is assumed to be constant (Eqn. 2), according to literature data [11 ] and in view of the lack of any noticeable effect upon the activity of the catalyst. The total coking rate was predicted as the sum of Eqns. 1 and 2; Eqn. 3 relates catalyst activity to the mono-layer coke content. A global view

189 of the reaction-coking path-way is shown in Figure 6. The parameters shown in Table 3 were calculated using a multivariable, nonlinear regression fitting algo-rithm. The best results were obtained for h=2, m=l and s=2, and the dependence on temp-erature of ks, kR and kw, was expressed by the Arrhenius equation in its reparametrized form (Tm=468 K) (Eqn. 4). Consistent with experimental data (Figure 7), the model predicts a residual activity (as) for the hydrogenation reaction. The value of as is given by the ratio kR/(ks+kR) (12). Thus, the value obtained for EaR which is 1.8 times the value of Eas, explains the diminution of the residual activity with the observed temperature. On the other hand, the value of kw0 is nearly 1000 times lower that ks0, which also is, qualitatively consistent with the higher proportion of the amorphous coke observed by SEM (Figure 5) [ 11 ]. It is also interesting to note the low value of the activation energy calculated for the process of filamental growth (Eaw), since generally values within the range 3040 kcal/mol are reported [11]. Nevertheless, these values are usually obtained once the filament growing process has reached the steady state, after the initial period of induction for nucleation and formation of the filament. In this case, as a consequence of the parameterized calculation method used, the value of Eaw does not correspond directly to the temperature dependence of the rate of filament lengthening, but rather the difference between this rate dependence and the activation energies for filament nucleation and formation, and for encapsulation of metallic particles during the induction period.

Rlament formation

I

C=H. (g)

I

ICH,tg) Ii

Filament

growth

T

C,H= (g) - ~ C,H. la"ff) " - 14q9 > Filament ii -X~ ~TH=(g) I nucleation

I

Ij

adsorbed II layer of C=H. ,

~

iI

1.0

I

I

0.8

I I

1

reversible monolayer coke formation

Figure 6. Scheme of the reactioncoking pathway for the hydrogenation of acetylene.

|

!

9

|

!

9 428 K

o 448K o

4~K

9

9

A 488 K x 498 K

0.7

I

C,H, (g)~ I Polimerization , Deactivation .of the C=H layer I---> of the catalyst

|

0.9

I

encapsu ~1 at"ion I J

9

O

0.6

o o

0.5 0.4 0"30

&

X

x

A X

,

I

2000

i

I

4000

i

I

6000 9 me

i

I

8000

i

I

10000

,

12000

(s)

Figure 7. Comparison of the experimental activity data (symbols) and the data predicted by the model (continuous line)at different operating temperatures.

4. CONCLUSIONS. It is possible to obtain well-dispersed Ni-Co catalyst supported in a ZnAl204 nonstoichiometric inverse-spinel structure, by using an appropriate coprecipitation method. The close proximity of the reduction temperatures of Ni and Co produces a high degree of interdispersion and a synergetic increase in the activity of the active metals. Thus, catalytic

190 performance of the Ni-Co catalyst is better (higher ethylene selectivity and lower coke formation) than either of the Ni- or Co-based catalysts. The ratio of Ni~ ~ on the catalyst surface, after different calcination and reduction temperatures, determines catalyst activity and selectivity. Cobalt apparently effects a dilution of surface Ni ensembles, thus reducing the population of multiatomic nickel ensembles, leading to an increased selectivity to ethylene. Increasing operating temperature increases coking rate causing a lower initial conversion but a more selective catalyst. Within the temperature interval studied, the main reaction has a negative apparent activation energy which is a consequence of encapsulation of part of the metallic particles during the initial period of the reaction. During this stage, major changes occur on the active surface as a result of cracking and hydrogenolysis of the acetylene molecules which lead to higher methane production. Coke deposition is modeled as a reversible process. The activation energies calculated for the coke formation and coke removal reactions, are 8.3 kcal/mol and 4.6 kcal/mol respectively. These values explain the diminution of residual hydrogenation activity with increasing reaction temperature. Finally, a value of 4.9 kcal/mol is calculated for coke filament growth. This low value is explained by considering that it is the difference between the activation energy for the filament lengthening process and that corresponding to a combination of filament nucleation/formation, and encapsulation of metallic particles during the initial induction period.

Acknowledgments. The authors wish to acknowledge the financial support of DGICYT (Spain) for this work (Project PB94-0568).

References. J.R.H. Ross, M.C.F. Steel, and A. Zeini-Isfahani, Appl. CataL 52 (1978) 280. Y.G. Chen, and J. Ren, Catal. Lett. 29 (1994) 39. J.P. Jacobs, A. Maltha, J.G.H. Reintjes, J. Drimal, V. Ponec, and H.H. Brongersma, J. CataL 147 (1994) 294. 4. J.C. Rodriguez, E. Romeo, A. Monz6n, A., Borgna, and A.J. Marchi, Proc. of XV Iberoamerican Symposium on Catalysis, C6rdoba (Argentina), 1996, pp. 909-914. 5. F. Cavani, F. Trifir6, and A. Vaccari, Catal. Today, 11 (1991) 173. 6. J.A. Pefia, J. Herguido, C. Guimon, A. Monz6n, and J Santamaria, J. Catal. 159 (1996) 313. 7. G.R. Gavalas, C. Phichitkul, and G.E. Voecks, J. Catal., 88, (1984) 65. 8. G. Fomasari, S. Gusi, F. Trifir6, and A. Vaccari, Ind. Eng. Chem. Res., 26 (1987) 1500. 9. J. Margitfalvi, L. Guczi, and A.H. Weiss, J. Catal., 72 (1981) 85. 10. S.J. Thomson, and G. Webb, J. Chem. Soc. Chem. Comm. 526 (1976). 11. Figueiredo, J.L., Progress in Catalyst Deactivation. (J.L. Figueiredo, Ed.), NATO Adv. Stud. Inst. Ser.; Ser E: Appl. Sci., Vol. 54, 1981, pp 45-63. 12. J.C. Rodriguez, J.A. Pefia, A. Monz6n, R. Hughes, and K. Li, Chem. Eng. J., 58 (1995) 7. 1. 2. 3.

9

Elsevier Science B.V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

Stability and R e g e n e r a t i o n Dehydrogenation

of

191

Supported

PtSn

Catalysts

for

Propane

Cristina L. Padr6, Sergio R. de Miguel, Alberto A. Castro and Osvaldo A. Scelza. Instituto de Investigaciones en Catfilisis y Petroqufmica (INCAPE), Santiago del Estero 2654, (3000) Santa Fe, Argentina. FAX 54 42 571162. This paper reports the catalytic performance for propane dehydrogenation of Pt and PtSn catalysts on inert supports, studied by means of pulse and flow reaction techniques, in order to determine effects of support on catalytic behavior. K addition to Pt/A1203 and PtSn/A1203 significantly decreases catalyst deactivation, by lowering the amounts of coke deposited on the support and metal, while producing a coke with a lower degree of polymerization. Sn addition to Pt/A1203 improves the activity and selectivity to propylene and decreases deactivation. The addition of Sn to Pt/ZnAI204 not only enhances activity but also improves catalyst stability. The existence of strong interactions between Pt and Sn, with probable alloy formation, is suggested both on an acidic support like A1203, and on a pHneutral support like ZnA1204. Bimetallic PtSn/ZnA1204 catalysts appear to have the highest stability and reversibility after several reaction steps including the corresponding regeneration steps between them. 1. I N T R O D U C T I O N Dehydrogenation of light paraffins is an endothermic reaction that must be carried out at high temperatures in order to obtain high yields of olefins, due to thermodynamic barriers. The use of metallic supported catalysts for this reaction has increased in recent years. Thus, supported Pt catalysts modified by the addition of an inert metal, like Sn, Ge, Pb, In, etc., appear to be promising catalytic agents [ 1,2]. In these catalysts, the nature and behavior of the support are important topics to be considered. In fact, considering the high reaction temperatures and the highly exothermic reactions which take place during the regeneration steps, supports must exhibit both high thermal stability and high capacity to disperse the metallic phase. Moreover, the dehydrogenation of light paraffins can be accompanied by undesirable reactions such as formation of gases (by cracking and hydrogenolysis) and coke deposition. These undesirable reactions can be catalyzed by acidic supports and/or by the active metal. The modification of the structure of Pt by the addition of an inactive metal can modify the hydrogenolysis activity. Meanwhile, the cracking and polymerization reactions, catalyzed by acid sites, can be eliminated using neutral supports like Znml204spinels, or low acid supports like K-doped alumina. The catalytic performance in the propane dehydrogenation reaction of Pt and PtSn catalysts on inert supports (like K-doped alumina and ZnAlaO 4, and on an acidic AIaO3,) under conditions of high severity was studied by means of pulse and flow reaction techniques to determine the effect of the nature of the support on the catalytic behavior. The

192 stability of the different catalysts was followed by experiments of successive reactionregeneration cycles. Besides, TPR, XRD, H 2 chemisorption and test reaction experiments (cyclohexane dehydrogenation and cyclopentane hydrogenolysis) were also carried out. The coke deposition was followed by a pulse reaction technique and TPO experiments. 2. E X P E R I M E N T A L Two different supports were used: a commercial (~-A1203 (CK-300 from Cyanamid Ketjen, SBET-" 180 mEgl) and a ZnA1204 (SBET= 34 mEgl) prepared in the laboratory. The ZnAIEO 4 spinel was obtained by solid phase reaction between (?-AlEO3 and pure ZnO at 1173 K for 72 h. The solid was then cooled to room temperature and submitted for XRD analysis, which showed the characteristic lines of ZnA1EO4 and weak lines of ZnO. To eliminate the unreacted ZnO (which could produce PtZn alloys with a very low dehydrogenation capacity), the solid was repeatedly washed with portions of (NH4)ECO 3 solution (1 M). After washing, the solid was dried and a new XRD analysis showed the total absence of the ZnO phase. Pt (0.3 wt%) and PtSn (0.3-0.3 wt%) catalysts supported on A1EO3 and ZnA1EO4 were prepared by wet impregnation. Thus, supported Pt catalysts were obtained by impregnation of the support with an acetic solution of HEPtC16 (AcH concentration: 0.3 M, HEPtCI6 concentration: 0.011 M). After impregnation, samples were dried overnight at 393 K and then calcined in an air stream for 3 h at 773 K. Bimetallic PtSn catalysts supported on A1EO3 and ZnAI204 were prepared by impregnation of their precursors, Pt/A1203 and Pt/ZnA1EO4 obtained according to the aforementioned procedure, with a hydrochloric solution of SnC12 (HC1 concentration: 0.4 M, SnC1E concentration: 0.0176 M). After impregnation, the solids were dried overnight at 393 K, and finally calcined in an air stream for 3 h at 773 K. Pt (0.3 wt%)/A1EOa-K and PtSn (0.3-0.3 wt%)/A1EO3-K were obtained by impregnation of the monometallic or bimetallic catalyst with a KOH aqueous solution (0.092 M) to obtain a final K concentration in the catalysts of 0.50 wt%. Then, samples were dried overnight at 393 K, and finally calcined for 3 h at 773 K. Monometallic and bimetallic catalysts were activity-tested for propane dehydrogenation in a fixed-bed flow reactor. The sample, placed in a quartz reactor, was previously reduced for 3 h at 853 K under flowing H E. Then the reaction mixture, composed of H E, C3H8, and He (HE/Call 8molar ratio = 0.9, He/Call 8 molar ratio = 8.1), was fed to the reactor. Weight time was 4 x 10-3 g cat min m1-1.Four minutes after the reaction began, the first sample of reactor effluent was injected into the chromatographic FID system with a dimethylsulfolane on Chromosorb column. The total reaction time was 80 minutes. A pulse reaction technique was used to study: (1) the behavior of the catalysts in the first step of the reaction, (2) the initial carbon deposition and, (3) the effects of carbon deposition on conversion and selectivity. Experiments were carded out at 853 K by injecting pulses of pure propane (pulse volume: 0.50 cm 3 STP) to the catalytic bed, which was maintained under flowing He (30 ml min 1) between two successive pulses. The samples were previously reduced for 3h at 853 K under flowing H 2. The reaction products were analyzed in a FID chromatographic system with a packed column (Porapack Q). Temperature-programmed reduction (TPR)experiments were carried out in a conventional TCD system using 10 ml min -I of an H 2 (5 wt%)-N 2 gaseous mixture at a constant heating rate of 6 K min -1 from room temperature up to 1073 K. Temperatureprogrammed oxidation (TPO) profiles were obtained using a mixture of O E(6 %v/v)-N 2 and

193

a heating rate of 12 K min ~. Fresh samples were evaluated in cyclohexane (CH) dehydrogenation and cyclopentane (CP) hydrogenolysis reactions. CH dehydrogenation was performed at 573 K using an H2/CH molar ratio of 29. CP hydrogenolysis was carded out at 573 K using an H2/CP molar ratio of 25. Before activity testing, samples were reduced for 3h at 773 K under flowing H 2. To study catalyst stability, catalyst samples were subjected to five propane dehydrogenation reaction steps (6 h at 853K). After each reaction step, samples were submitted to a regeneration treatment with an 02 (5%v/v)-N 2 mixture (6 h at 773 K). 3. RESULTS AND DISCUSSION Table 1 shows both the initial propane conversion (X0) and the selectivity to propylene (So), measured 4 min after the reaction started, and the deactivation parameter (X/X 0 obtained in flow reaction experiments). This parameter is defined as the difference between the initial propane conversion (Xo) and the final conversion (Xe) measured at 80 min of the reaction time, referred to the initial propane conversion for the different catalysts. Table 1. Values of initial conversion (X0), initial selectivity to propylene parameter (X/X0) for the different mono and bimetallic catall;csts. Catall;cst X0 , % S0, % Pt/A1203 30.0 58.9 PtSn/A1203 45.0 90.2 Pt/A1203-K 31.0 55.0 PtSn/A1203-K 43.7 91.1 Pt/Znml204 23.0 95.0 PtSn/ZnA1204 38.7 94.0

(So) and deactivation (X/X0, % 77 67 48 38 44 16

When the Pt/A1203 catalyst is doped with K, slight modifications of both the initial conversion and the selectivity to propylene, and a significant decrease of the deactivation parameter, are observed (Table 1). Additionally, the amount of coke deposited on Pt/ml203K (1.26 wt%) is lower than that deposited on the undoped monometallic catalyst (1.95 wt%), as shown by TPO and pulse technique results (Figures 1 and 2a). When K is added to alumina, a significant decrease of the strong acid sites of alumina was found [3] due to the blocking effect on the tetrahedral A1+3 acid sites [4]. However, other important effects must be considered. First, a decrease of the Pt dispersion after the alkali metal addition (from 55 % for Pt/A1203 to 35% for K-doped catalyst) was detected. This increase in the metal particle size produces a decrease of the CH dehydrogenation/CP hydrogenolysis ratio, as shown in Table 2. Since that CP hydrogenolysis is a structure-sensitive reaction whereas the CH dehydrogenation is a structure-insensitive one, the increase of Pt particle size would produce a decrease of the dehydrogenation rate and a parallel increase of the amounts of Pt ensembles necessary for the hydrogenolysis reaction. Hence, K addition to Pt/A1203 produces two opposite effects: on the one hand, a poisoning of the acid sites of the support, and on the other, an increase of the Pt particle size. However, these effects would not justify the pronounced decrease of the deactivation of Pt/AI203-K. This fact suggests an additional effect of the K addition to Pt. Thus the TPR profile of Pt/A1203-K shows very important

194 differences with respect to the monometallic catalyst (Figure 3), which can be attributed t o a modification of the structure of the supported Pt by the alkali metal addition. It has been reported in the literature that Pt is electronically modified by K addition [5], which would favor the labilization of olefins, decreasing the fouling effect on the metallic function. This agrees with the TPO profile of Pt/A1203-K (Figure 1), which shows not only a decrease of the carbon deposited on the support (2nd. oxidation peak), but also an important decrease of the carbon deposited on the metal (lst. oxidation peak), with respect to the undoped catalyst. Moreover, the first peak in the TPO profile is shifted to lower temperature, which can be attributed to a lower polymerization degree of coke.

I

I 0 273

473

T I K 673

873273 [

473 T , K 673

875

,,o

u

i~s r

0

I

I 2 PULSE

(3.. I 3

l

4

NUMBER

I 5

0

2

,

PULSE

3

4

NUMBER

Figure 1. TPO profiles of samples after flow reaction experiments in the propane dehydrogenation Figure 2. a) Relative retention of carbon referred to that of Pt/A1203 a s a function of the pulse number of propane; b) propane conversion vs. pulse number.

195 Table 2. Initial reaction rates for CH dehydrogenation Sample ( Rc. ), mol/h g Pt Pt/ml203 111.8 Pt/AlaO.~-K 54.0

(Rcn) and CP hydrogenol~,sis (Rc,) (Rcp) mol/h g Pt (Rc.)/(Rcp). 13.13 8.51 10.43 5.17

The effect of Sn addition to Pt/ml203 is reflected in the increase of the initial activity and selectivity to propylene in flow reaction experiments accompanied by a decrease of the deactivation parameter (Table 1). The carbon retention obtained in the pulse experiments was also substantially lower for the bimetallic catalyst compared with Pt/A1203 (Figure 2a). The apparent increase of the initial activity for PtSn/A1203 observed in the flow experiments with respect to that of the monometallic one can be explained by taking into account the results obtained in pulse experiments (Figure 2b). In fact, the conversion in the first pulse is higher for Pt/A1203 than PtSn/A1203, while for the fifth pulse the slope of activity-pulse number curve is higher for the Pt/A1203 catalyst. This can be attributed to the higher deactivation rate for the monometallic catalyst. The higher selectivity to propylene for PtSn/A1203 compared to that of Pt]A1203, can be assigned to a modification of the nature of the metal function by Sn addition, such as shown by TPR and test reactions. The TPR profile of PtSn/A1203 (Figure 3) shows a single reduction peak at 523 K with a very small shoulder at 780 K. This profile is very different from the one corresponding to the sum of the TPR profiles of the monometallic catalysts. Moreover, the activation energy for CH dehydrogenation for the PtSn/A1203 catalyst (50.8 kcal mo1-1) is much higher than that on the Pt/A1203 samples (20.6 kcal mol~). These results suggest a strong interaction between both metals with probable formation of PtSn alloys in the reduced bimetallic catalyst, which agrees with other results reported in the literature [6,7]. K addition to bimetallic PtSn catalyst supported on A1203 does not modify the initial activity and selectivity during the flow reaction of propane dehydrogenation but produces a decrease of the deactivation parameter, as shown in Table 1. This effect is accompanied by a lower carbon retention of the PtSn/AI~Oa-K catalyst with respect to the undoped bimetallic sample during the pulse experiments (Figure 2a). Moreover, TPO results on the K-doped bimetallic catalyst shows a lower coke deposition on both the metal and the support with respect to Pt/A1EOa-K(Figure 1). In this catalyst, the poisoning of the acid sites of alumina would be due not only to K but also to Sn addition. The TPR profile of the bimetallic catalyst doped with K (Figure 3) shows a modifying effect of the alkali metal on the metallic phase. Furthermore, considering the results obtained in test reactions, it was found that K addition to PtSn/A1203 significantly decreases the activation energy for CH dehydrogenation relative to that of the undoped bimetallic catalyst (from 50.8 to 35.3 Kcal mol-1), which would suggest that K decreases the interaction between Pt and Sn. This modification in the interaction between both metallic components could be due either to a direct action of K on the metallic phase or to an indirect effect of the alkali metal addition to the support, which could change the metal-support and the metal-metal interactions, such as suggested in the literature [5,8]. The monometallic Pt/ZnA1204 catalyst shows an initial conversion lower than the Pt/A1203, but a much higher selectivity to propylene (95%) and a lower deactivation parameter (Table 1). This effect is also reflected in a lower coke deposition (1.95 wt% C for Pt]A1203and 0.94 wt% for Pt/ZnA1204), as is observed in TPO profiles (Figure 1). The different behavior of these two catalysts cannot be related to the metal dispersion, since the

196 Pt dispersion for Pt/ZnAl204 (57%) was similar to that of Pt/A1203(55%), according to the hydrogen chemisorption measurements. Therefore, the different performance of both monometallic catalysts can be attributed to a different structure of the metallic phases. Thus, Bosch et al [9] observed no systematic relative orientation between crystallographic planes of the support and those of the Pt particles supported on ZnA1204. In this sense, the TPR profile of Pt/ZnA1204 (Figure 3) is sharp and well defined, which could be due to the presence of a narrow population of metallic species. All this evidence could suggest that Pt is composed of small, homogeneous particles, with a high concentration of sites for propane dehydrogenation and a low concentration of hydrogenolytic sites, consistent with the cyclopentane hydrogenolysis results.

P

I-. (1.

''~

.,. . . . .

2"F3

..,.-"

473

%,

6?3

073

273

L, 473

T,K

1

G73

1

ffV3

"r,K

Figure 3. TPR profiles of different catalysts. The addition of Sn to the PIfZRA1204catalyst apparently promotes activity in the propane dehydrogenation reaction, while the selectivity to propylene remains practically unchanged, and the deactivation rate is markedly decreased (Table 1). However, coke contents were similar for both PI/ZnA1204 and PtSn/ZnAI204catalysts (0.94 and 0.86 wt%, respectively). To explain these results, TPO experiments were performed. The profile of the bimetallic sample (Figure 1) displays an oxidation peak at 560 K, and a shoulder at higher temperatures, probably due to the carbon deposited on the support. The first peak, corresponding to carbonaceous species deposited in the neighborhood of the metal, is shifted to higher temperatures with respect to the peak found in the monometallic catalyst, probably due to a different carbonaceous structure. Therefore, the lower extent of deactivation observed for the bimetallic catalyst could be due to the beneficial effect of Sn favoring deposition of the coke at the metal-support interface, without affecting the active sites of the metal, thus producing a coke of lower toxicity at a different location. Figure 3 shows TPR profiles of the PtSn/ZnA1204 catalyst and those corresponding to monometallic Pt/ZnA1204 and Sn/ZnA1204 catalysts. The profile of the bimetallic sample displays a sharp reduction peak coincident with the Pt reduction, with a small reduction zone at temperatures

197 where the Sn reduction in Srl]Znml204 takes place. These results would indicate that the metallic phase of this bimetallic catalyst is composed of different species, and that a significant fraction would correspond to intermetallic compounds or PtSn alloy particles (probably with a well-defined composition). Table 3 shows the evolutions of initial activity (X0) and initial selectivity (So) for the different propane dehydrogenation reaction steps (obtained from experiments of reactionregeneration cycles). Some differences are evident relative to the trends reported in Table 1. These differences can be attributed to a different residence time in both experiments. It can be observed in all cases that the initial activity (X0) decreases through the successive steps. The percentage decrease of the initial conversion between the first and the fifth step, with respect to the initial activity of the first reaction step, follows this sequence:

Pt/A1203 (76%) > Pt/A12Oa-K (47%) > PtSn/A12Oa-K (42%) > PtSrotZnA1204 (15%) Table 3. Values of initial conversion (X0) and initial selectivity (S0) during successive cycles of reaction-regeneration. R.S.: reaction step. CATALYST 1st R.S. 2nd R.S. 3rd R.S. 4th R.S. 5th R.S.

Pt/A1203 Pt/A1EOa-K PtSn/A1EOa-K PtSn/ZnAlzO4

X0

So

X0

So

X0

So

X0

So

X0

So

25 30 40 40

46 52 87 96

22 27 30 39

47 54 89 96

17 22 28 36

48 62 90 98

10 18 27 36

62 64 90 97

6 16 23 34

62 68 89 98

These results show (1) the highest reversibility of the regeneration process for the bimetallic catalyst supported on ZnA1204, (2) a higher extent of deactivation for PE]A1203 relative to Pt/A1EOa-K, and (3) improved stability due to Sn addition to Pt especially for the catalyst supported on ZnA1204, whose conversion is practically constant through the successive reaction steps (Table 3). With respect to the selectivity to propylene, the most pronounced changes correspond to Pt/A1203 and Pt/A1EO3-K catalysts. Selectivity to propylene remains practically constant through five cycles for PtSn/ZnA1204 and PtSn/A1EOa-K. However, selectivities to propylene are higher for the PtSn supported on ZnA1204 (96-98%)relative to those for PtSn supported on the K-doped alumina (89-90%). These results clearly show the important role of the ZnA1204 spinel support in promoting activity, selectivity and stability of the metallic phase. 4. C O N C L U S I O N S 1. K added to Pt/A1203 blocks acidic sites of the support. It also substantially decreases the deactivation of the catalyst, diminishing the amounts of carbon deposited not only on the support but also on the metal. The coke deposited on Pt displays a lower degree of polymerization. These effects are probably caused by an electronic modification of the metallic phase due to the K addition. 2. The significant modification of the metallic phase of Pt due to Sn addition, with probable alloy formation, was confirmed by TPR results and by the substantial changes of

198 the activation energy for propane dehydrogenation. Sn addition to Pt also enhances the selectivity to propene and decreases the catalyst deactivation rate. 3. The change from an acidic support (like A1203) to another with neutral or nonacidic characteristics (like ZnA1204)produces important modifications in the characteristics and catalytic performance of the metallic phase. The different behavior of Pt/ZnA1204 and Pt/A1203 cannot be attributed to an effect of particle size, but to changes in the structure of the metallic particles. In this way, the TPR profile of the Pt]ZnAI204catalyst would indicate the existence of a narrow population of surface species forming small particles of homogeneous characteristics, with a high concentration of sites for propane dehydrogenation and a low concentration of sites promoting hydrogenolysis. 4. Sn addition to Pt/ZnA1204 improves activity and decreases deactivation rate in the propane dehydrogenation reaction. This improvement in the stability is related to the lower toxicity of the coke deposited on the bimetallic catalyst located at the support-metal interface. ACKNOWLEDGMENTS We thank Secretarfa de Ciencia y Tecnologfa de la Universidad Nacional del Litoral (Program CAI+D) for the financial support. We also thank Dr. Carlos Querini for his contribution in TPO experiments. REFERENCES

1. L.C. Loc, N.A. Gaidai, S.L. Kiperman, H.S. Thoang, N.M. Podletnova and S.B. Kogan, Kinet. i Katal. 32(1991)72. 2. O.A. Barias, A. Holmen and E.A. Blekhan, Studies in Surface Science and Catalysis 88(1994)519. 3. A.A. Castro, Catal. Lett. 22(1993) 123. 4. S.R. de Miguel, O.A. Scelza, A.A. Castro and J. Soria, Topics in Catalysis 1(1994)87. 5. S.R. de Miguel, A.A. Castro, O.A. Scelza and J. Soda, Catal. Lett. 32(1995)281. 6. G.T. Baronetti, S.R. de Miguel, O.A. Scelza and A.A. Castro, Appl. Catal. 24(1986)109. 7. B.H. Davis, ACS Symposium Series N( 517, "Selectivity in Catalysis", 1993, Chap.8, p.ll0. 8. Y.H Park and G.L.Price, Ind. Eng. Chem. Res., 32(1992)469. 9. P. Bosch. M.A. Valenzuela, B. Zapata, D.R. Acosta, G. Aguilar-Rios, C. Maldonado and Y. Schifter, J. Molec. Catal. 93(1994)67.

~ Elsevier Science B.V. All rights reserved Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

199

D e a c t i v a t i o n and Shape Selectivity Effects in T o l u e n e N i t r a t i o n o v e r Zeolite Catalysts Jeffrey M. Smith, Haiyang Liu, and Daniel E. Resasco School of Chemical Engineering, University of Oklahoma, 100 E. Boyd St., Norman, OK 73019, USA

The nitration of toluene in the presence of solid acid catalysts (ZSM5, Mordenite, Beta, L, zeolites and MCM-41 and sulfated zirconia) has been studied in the vapor and liquid phases. The data indicate that the catalysts deactivate by two different modes, coke formation and pore plugging by adsorbed toluene.

1.

INTRODUCTION

Toluene nitration is an important chemical reaction producing commercially valuable intermediates such as nitrotoluene and dinitrotoluene. Current industrial nitration occurs in mixed acid, a concentrated solution of HNO 3 and H2SO 4, at 30-70 ~ in a CSTR. However, the rate of reaction quickly decreases with the dilution of sulfuric acid, as the reaction progresses due to the formation of water. Therefore, this process generates large amounts of spent sulfuric acid waste, which requires neutralization before disposal or re-concentration before being recycled (1). The replacement of sulfuric acid with a solid acid catalyst seems a good alternative considering the increasingly strict regulations on industrial waste streams and the cost of concentrating the spent acid. The mononitration products of toluene are the ortho, meta and para isomers. Among these, the para isomer has the greatest commercial value and sells for about three times as much as the ortho isomer (1). For this reason alone, it would be desirable to obtain the high para selectivity without great loss in activity. In the industrial nitration of toluene, the typical product distribution of ortho-meta-para isomers is roughly 58:4:38 and the ortho/para ratio is 0.76, close to the statistical value of 1.0. This process, therefore, offers little regioselective control of products. Its selectivity for the most valuable isomer, para-nitrotoluene, is only slightly better than that predicted by statistics for such an electrophilic substitution reaction. When the reaction is carried out on solid catalysts, significantly higher para selectivities have been obtained. Notably high selectivities have been obtained using n-propyl nitrate as the nitrating agent (2). Several heterogeneous catalysts have been investigated for the liquid and vapor phase nitration processes. An early attempt by McKee and Wilhelm (3) used silica gel as a solid catalyst for the nitration of benzene and toluene in the vapor phase. Various other heterogeneous catalysts systems have also been studied such as: sulfonated polyorganosiloxanes (4), acidic resins (5), modified clays (6), zeolites (7,8), sulfated zirconia

200 (9), supported sulfuric acids (10), and supported sulfonic acid (11). Zeolites have also been used as shape-selective catalysts in organic synthesis and electrophilic substitution of aromatics (12,13). Medium pore sized zeolite catalysts should result in higher para isomer selectivities due to the preferential diffusion of the para isomer through the channels of these catalysts. Furthermore, it is generally accepted that toluene nitration requires strong acid sites to take place (14). Therefore, acidic zeolites such as H-Mordenite, H-ZSM-5 and H-Y appear to be good candidates for this reaction(15). Bertea et al. (16) have found that zeolite catalysts exhibit good performance for the nitration of benzene in the gas phase with 65% nitric acid as the nitrating agent. The use of solid acid catalysts would eliminate waste disposal problems and allow for more advantageous control of product selectivities. However, rapid deactivation of these solid acid catalysts is a problem that significantly hinders the effective performance and selectivity of these catalysts. We have studied the performance of various solid acid catalysts for their activity/deactivation characteristics and also their shape selective effects. Specifically in the liquid phase system, unlike previous researchers, we have studied the activity/deactivation evolution with time on stream rather than rely on final product distribution only. This approach has allowed us to obtain unique data that clearly describe the deactivation pattern of these catalysts.

2.

EXPERIMENTAL

The protonic forms of the ZSM-5 (SiO2/A1203ratio = 24.3), Y-54 (SiO2/A1203ratio = 5.3), L, and Beta zeolites were prepared by treating the respective powder Na forms with NH4C1, then washing them in distilled water, and calcining them at 550 ~ H-Mordenite (LZ-M-8, SiO2/A1203 ratio -- 8), H-Y-85 and ZrO2-SO4 2- (9.9 wt. % SO4), referred to as SZ, were used as received. The MCM-41 was prepared at The University of Oklahoma. The H-Mordenite, Y and Beta zeolites were donated by UOP, as was the Na-ZSM-5 from Chemie Uetikon, the SZ from MEI, and the L zeolite from Tosoh Corp. Vapor phase experiments were performed in a fixed-bed continuous down-flow microreactor with a diameter of 12 mm at atmospheric pressure. Prior to reaction, the catalyst sample (0.10 g, except 0.15 g for SZ) was calcined at 500 ~ for 3 hours in a 30 cm3/min flow of air. The sample was then purged with 0.50 cma/min He for 0.5 hr at the reaction temperature. Reactant feed rates were 0.33 cma/min of 0.99 vo1% NO 2 /He and 0.17 cm3/min of He saturated with toluene by passing through a toluene reservoir kept at 0~ Reactions were performed at 150, 170 and 200 ~ and partial pressures of 2.3 torr toluene and 5-15 torr NO 2. The toluene space velocity (WHSV) was 0.148 for the SZ and 0.221 for all other catalysts. Toluene was used as the internal standard for the calibration of the mono-nitrotoluene products. Liquid phase experiments were performed with 5.0 or 10.0 g of catalyst pellets in a 300 ml Berty-type reactor under batch conditions at 66 to 100 ~ and 1.0 or 5.0 vol.% n-propyl nitrate. Prior to reaction, the catalyst sample was calcined at 500 ~ for 2 hrs. in 0.5 cm3/min of air. Then the catalyst was purged with He for 5 min before the addition of 250-280 ml of preheated toluene. Stirring at 300 rpm, the n-propyl nitrate was added via syringe injection, this point represents time zero. To follow the evolution of activity with reaction holding time, liquid samples of 1 ml were taken with a syringe and mixed with an internal standard of decane, diluted in hexane, as the calibration solution before analysis. Both vapor and liquid reaction

201

systems used a GC-MS (Hewlett-Packard, GCD) equipped with a 30 m x 0.25mm x 0.25mm cross-linked 5% phenyl methyl siloxane capillary column for sample analysis. Under the conditions described above, only mononitration products were observed. In a separate study, dinitrotoluenes were formed with high yield when excess amounts of fumic nitric acid were used as the nitrating agent. However, the addition of solid acid catalysts to this system resulted in negligible increases in dinitro-toluene production. It should also be noted that using nitric acid at lower concentrations resulted in greatly decreased amounts of dinitrotoluenes produced.

3.

R E S U L T S AND D I S C U S S I O N

As shown in Fig. 1, an induction period of 30-75 min was observed over all zeolite and MCM catalysts in the vapor phase runs. During this induction period, no toluene or products were observed in the effluent; toluene and nitrotoluene appeared only when the catalyst seemed to become saturated. Examination of the reactor effluent indicated that this adsorption only occurred when the two reactants were fed to the system simultaneously. When either toluene or NO 2 was fed separately, no adsorption was observed. For example, in one experiment, toluene was passed over the catalyst for 90 min. without noticeable adsorption. Then, as soon as the NO 2 flow was started, the toluene signal disappeared for about 50 min., after which both unreacted toluene and nitrotoluene products started to appear. It seems that the presence of NO 2 enhances the adsorption of the aromatic molecule which in turn is necessary for the catalytic process. However, this extensive adsorption may, simultaneously, cause pore plugging and catalyst deactivation. The induction period was much shorter on SZ, possibly because accumulation of hydrocarbons species occurs to a lesser extent on its relatively low surface area. It is interesting to note that SZ catalysts show long induction periods for the isomerization of n-butane, a different reaction system (17). In that case, the induction periods appear because a bimolecular step is necessary for the reaction. Accordingly, the presence of induction periods and the accumulation of hydrocarbon molecules in the nitration reaction may suggest steps involving more than one aromatic ring. The vapor phase studies also showed that the size of the pores has an important effect on the regioselectivity of the reaction. Fig. 2 illustrates the effect of pore size on product selectivity. The highest selectivity for para-nitrotoluene was observed over H-ZSM5 followed by Mordenite, L, and MCM. The kinetic diameters of the o-, m-, and p- isomers are 6.7, 6.7, and 5.25 A respectively. Experimental results shown in Fig. 2 verify that catalysts with channel diameters of less than 7 ~ preferentially form the para isomer. Based on these geometric arguments alone, the small amount of ortho isomer observed with the H-ZSM5 catalyst could be generated on acid sites located on the outer surface (2). A similar shape selective effect was observed in the liquid phase. Those catalysts with the smaller pore and channel openings were more selective for para-nitrotoluene. However, in the liquid phase, no induction period was observed. Rather, all catalysts exhibited significant deactivation throughout time on stream and after 5 hrs. little of the original activity remained. As shown in Fig. 3, the para selectivity was found to decrease with time on stream. This would indicate that deactivation occurs within the pore channels effectively reducing the preferential capacity of the catalyst to generate the para isomer. The decrease in para selectivity was not evident on Beta zeolite, which has larger pores and may allow for a more uniform production of

202 25

2.0E-07 r

O

E J

1.5E-07 -

20

!

o

1.0E-07

ZSM5 !

..~ 10

i,,,/x HMOR

r~

2

> 5.0E-08

5

..... MCM

O

SZ

'----O-

0

0.0E+00 50 100 Time, min.

i. 0

150

+

10 6O Pore Size (A)

Figure 1. Evolution of nitration activity conducted in the vapor phase flow reactor at 150 ~C illustrating the induction period observed for all the zeolite catalysts. Figure 2. Effect of pore size on the para-isomer selectivity in the vapor phase at 150 ~ ortho and para inside and outside the channels. Therefore, if pore plugging occurs, it would not affect the selectivity. Fig. 4 shows the evolution of total conversion as a function of holding time at different temperatures in the batch reactor. Here, the conversion is defined as the fraction of moles of the nitrating agent, n-propyl nitrate, converted into mono-nitrotoluene products. As in a typical batch operation, the conversion increased with time and reached a plateau. However, unlike many batch reactions, the very noticeable leveling off in conversion was not due to the shortage of reactants, but entirely due to catalyst deactivation. The unreacted toluene was always in great excess and injections of nitrating agent after reaching the plateau resulted in no further conversion.

20 0.8

ZSM5

15

> -

0.7 "-~ r.,t3

BETA --__--__11~_

=

[]

O

10

--__ --

0.6

>

SZ

g' 0.s

O

HMOR

5

0.4 0

100 200 Time, min.

300

0

100 200 Time, min

300

Figure 3. Evolution of the selectivity toward the para-isomer during the nitration of toluene at 93 ~ conducted in the liquid phase batch reactor in the presence of various solid acid catalysts. The selectivity is defined as the fraction of paranitrotoluene in the products. Figure 4. Conversion of the nitrating agent to mononitrotoluenes in the liquid phase at various temperatures in the presence of the H-ZSM5 catalyst. The solid lines are fitted power law curves.

203 The initial activity, measured from the slopes at time zero, increased with temperature with an activation energy of about 11 kcal/mol. To determine the effect of temperature on deactivation the activity data was fitted with a power law equation, indicated in Fig. 4 by the solid lines. The reaction rate data of Fig. 5 was obtained by taking the derivative of this equation with time and then normalizing it to a common reference, 20 min time on stream.

1 tt 0.8 g 1\\ I I 5 ]-k 0.6

66C . . . . . . 75C 94C

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N .

N

,...4

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Z

~ZSM5 . . . . . . MORD --D--- SZ BETA

0.8

Z

0.22

0.4 0.2

0 20

120 Time, min

220

320

20

80

140

200

Time, min

Figure 5. Effect of temperature on the deactivation of the H-ZSM5 catalyst during the nitration of toluene in the liquid phase. Figure 6. Deactivation patterns observed for various solid acid catalysts at 94 ~ These rate profiles show that higher temperatures lead to slightly faster deactivation of the catalyst. This is not unusual for any acidic catalyst which generally deactivates by formation of coke. However, the apparent activation energy calculated for the deactivation process from Fig. 5 (4.8 kcal/mol) is much lower than those typically observed on deactivation by coke. As discussed below, the deactivation of these catalysts may involve two modes, the adsorption of compounds resulting in pore plugging and the coverage of active sites due to the deposition of coke. It is conceivable that these two modes of deactivation have opposing temperature dependencies, the latter increasing with temperature, the former decreasing, resulting in a very low apparent activation energy. Fig. 6 shows the normalized rate profiles for reaction at 94 ~ in the presence of different catalysts. This plot was constructed from the concentration profiles of these catalysts similarly to the ones calculated above for H-ZSM5. It can be observed that H-ZSM5 and H-Mordenite exhibit a very similar deactivation pattern while sulfated zirconia is a little faster and beta zeolite significantly faster. Beta zeolite showed the highest initial activity but a much faster deactivation. At the same time, beta zeolite exhibited a poor selectivity to the para isomer. In order to analyze the role of sites located outside the zeolite structure, we attempted to selectively poison them using tetraethyl orthosilicate to silylate the surface. This method has been successfully used in previous work (18). In this case, silylation of H-ZSM5 resulted in greatly reduced activity but no improvement in para selectivity as compared to untreated HZSM5. This effect might indicate that, on H-ZSM5, silylation only results in pore blocking.

204

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,

' 6, : 66C

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e~

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:

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.

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E

e~

E9

o

2:

Z 0

1O0

200

300

Temperature, C

400

75

275

475

675

Temperature, C

Figure 7. Temperature Programmed Desorption profiles obtained by heating the spent H-ZSM5 catalyst in He flow and detecting the evolution of adsorbed species after reaction at various temperatures. Figure 8. Temperature Programmed Oxidation of the carbonaceous deposits left on the HZSM5 catalyst after reaction at various temperatures. We have also studied the effect of co-feeding a bulky basic molecule which should neutralize the external sites under reaction conditions but not penetrate the channels. Phenanthridine, a 3-ring aromatic base, was used as the neutralizing agent at a concentration of 4.86 mmoles/ml. This neutralization had tittle effect on the selectivity of H-ZSM5, but resulted in a 10% increase in para selectivity over Beta and Mordenite catalysts. We attribute this increased para selectivity to pore narrowing of these larger channels which results in preferential para isomer diffusion (18). The nature of deactivation of these catalysts was examined using Temperature Programmed Desorption (TPD), and Temperature Programmed Oxidation (TPO) experiments after the toluene nitration runs. As mentioned above, pore plugging by hydrocarbons and coke formation seem to be responsible for the rapid deactivation of these catalysts. The TPD and TPO experiments indicate that these two processes have opposing dependencies on temperature. As shown in Fig. 7, the TPD profiles obtained in the mass spectrometer using a 30 cc/min. He cartier at a heating rate of 4 ~ to 250 ~ indicated that the abundance of adsorbed molecules, which may plug the pores, decreased with increasing temperature of reaction. Mass spectrometric analysis indicated that toluene was the predominant species adsorbed, although lesser amounts of larger mass species were also detected. Conversely, TPO experiments, performed in 30 cc/min, of 5% 02 in He at a heating rate of 6 ~ to 800 ~ indicated that the abundance of deposited coke, as estimated from the evolution of CO 2, increased with increasing temperature of reaction (see Fig. 8). That is to say, as the temperature of reaction increases, molecular diffusion increases and fewer molecules are retained in the pores, however, coke formation increases. Lower temperatures produce less coke but more adsorbed molecules resulting in higher para-nitrotoluene selectivity over all catalysts tested. Deactivation is also a function of the catalyst used. Under similar reaction conditions, Beta zeolite deactivated more rapidly than did H-ZSM5. The Beta zeolite was found to contain more adsorbed molecules as well as larger amounts of coke. On the other hand, sulfated zirconia deactivated more rapidly than H-ZSM5 but formed much less coke. It has previously been

205 observed that sulfated zirconia can be deactivated with as little as 0.2 wt % C (17), indicating that a small fraction of the sulfate groups are active as acid sites. As shown in Table 1, a similar amount of coke was found on the deactivated sulfated zirconia used in this study. Table 1. Carbon deposition determined on the various catalysts by TPO after 3 hr reaction at 94 ~ in the liquid-phase batch reactor

Coke wt. %

4.

BETA 2.0

ZSM-5 1.2

Sulfated Zirconia 0.3

CONCLUSIONS

High selectivities to para-nitrotoluene can be obtained using acid zeolites at mild reaction conditions. However, the catalysts rapidly deactivate by two main mechanisms. At the low temperature range, the dominant form of deactivation is the plugging of the pores by the aromatic molecules. As the temperature increases, the adsorption of hydrocarbon species and pore plugging decreases, but the catalysts deactivate by coke formation.

ACKNOWLEDGMENTS This work was supported by NSF (CTS-9403199), and the international cooperative program NSF-CONICET (INT-9415590). We gratefully acknowledge NSF for a graduate traineeship for one of us (JMS).

REFERENCES 1. R.L. Adkins in "Kirk-Othmer Encyclopedia of Chemical Technology", 4th ed. (J.I Kroschwitz, M. Howe-Grant, Editors), Vol. 17, J. Wiley, New York, 1991 p. 153. 2. T.J Kwok and K. Jayasuriya, R. Damavarapu, and B.W Rodman, J. Org. Chem. 59 (1994) 4939. 3. M. McKee and R.H. Wilhelm, Ind. Eng. Chem. 28 (1936) 662. 4. S. Suzuki, K. Tahmori, and Y. Ono, Y., Chem. Lett., 747 (1986); ibid, J. Mol. Catal., 43 (1987) 41. 5. G.A. Olah, V. V. Krishnamurthy, and S. C. Narang, J. Org. Chem. 47 (1982) 596. 6. A. Cornelis, A. Gerstmans, and P. Laszlo, Chem. Lett., 1839 (1988). 7. A. Germain, T. Akouz, and F. Figueras, Appl. Catal. A: General, 136 (1996) 57.; and J. Catal. 147, (1994) 163. 8. D. Akolekar, G. Lemay, A. Sayari, and S. Kaliaguine, Res. Chem. Interm. 21 (1995) 7. 9. S.H. Nagi, E. A. Zubkoz, and V. G. Shubin, Izv. Aka. Nauk SSSR, Ser. Khim., 1650 (1990). 10. J. M. Riego, Z. Sedin, J. M. Zaldivar, N. C. Marziano, and C. Toratato, Tetra Lett., 37 (1996) 513. 11. E. Suzuki, K. Tohmori, and Y. Ono, Chem. Lett., 2273 (1987).

206 12. W. Holderich, M. Hesse, and Naumann, Agnew. Chem. Int. Ed. Engl., 27 (1993) 226. 13. P. B. Venuto, Microporous Materials, 2 (1994) 297. 14. K. Schofield, in "Aromatic Nitration", Cambridge University Press, Cambridge, 1980; G.A. Olah, R. Malhotra, and S. C. Narang, in "Nitration Methods and Mechanisms". Organic Nitro Chemistry Series, Verlag Chemie, Weinheim, 1989. 15. W. V. Holderich, and H. van Bekkum, Stud. Surf. Sci. Catal., 58 (1991) 677. 16. L. R. Bertea, H. W. Kouwenhoven, and R. Prins, Stud. Surf. Sci. Catal., 84 (1994) 1973. 17. W. E. Alvarez, H. Liu, D. E. Resasco, Appl. Catal. (in press). 18. J. Cejka, N. Zilkova, B. Wichterlova, G. Eder-Mirth, J. A. Lercher, Zeolites, 17 (1996) 265.

9

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Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

207

R e g e n e r a t i o n o f V P M e O Catalysts for n-Butane Oxidation by M e a n s o f M e c h a n o c h e m i c a l and Barothermal Treatments V. Zazhigalov a, A. Kharlamov a, J. Haber b, J. Stoch b, V. Yaremenko a, I. Bacherikova a, L. Bogutskaya a Ukrainian-Polish Laboratory of Catalysis: alnstitute of Physical Chemistry, National Academy of Sciences of Ukraine, pr. Nauki 31, Kyjiv, 252022 Ukraine blnstitute of Catalysis and Surface Chemistry, Polish Academy of Sciences, ul. Niezapominajek, Krakow, 30-239 Poland

Regeneration of VPBiO catalysts deactivated during commercial use by (i) oxidation to VOPO4, (ii)phosphorus removal, and (iii) loss of the Bi promoter was investigated. It is shown that a combination of methods, i.e., mechanochemicaland barothermal treatments with Bi203 or BiPO4 and H3PO4 or P205 are successful in regenerating used catalysts, increasing selectivity by 11-23% and restoring most of the original catalytic activity. Since introduction of the Bi promoter alone into the spent catalyst produces only a small increase in the selectivity (4-6%), the extent of regeneration can be substantially improved by simultaneous phosphorus introduction.

1. INTRODUCTION Vanadium-phosphorus oxide systems are well known to be efficient catalysts for the selective oxidation of C4-C5 hydrocarbons [1-3]. A (VO)2P207 phase having P/V = 1 is an active component of this system. However, commercial use of the system is often based on promoted catalysts having an excess of phosphorus [2,3]. It has been suggested that overstoichiometric phosphorus can either stabilize the V 4§ oxidation state [4] or lead to the distortion of the (VO)2P207 structure [5]. It is also established that the surplus phosphorus is mainly concentrated in a thin layer at the catalyst surface [6-9]. Its loss during catalyst use [6,9] leads to a decrease in catalyst selectivity for maleic anhydride (MA). One reason for this loss may be transformation of V 4§ to V 5§ at the surface to produce a VOPO4 phase, as suggested in [4]. Regeneration of deactivated VPO catalysts by means of treatments with phosphorus-organic compounds [ 10,11 ] increases the selectivity for MA by 7-11%, but the initial efficiency of the fresh catalyst cannot be reached. It has been reported that addition of metal promoters (e.g., Bi, Co, etc.) to the VPO catalyst can stabilize overstoichiometric phosphorus at the surface slowing down its depletion

208 [6,12,13]. It is also known [14] that these metal promot6rs are lost during catalyst use. Thus, it can be concluded that decreased catalytic performance can occur due to (i) oxidation of the surface to form VOPO4, (ii) loss of the phosphorus from the catalyst, and/or (iii) loss of metal promoters from the catalyst. In the present paper consideration is given to the solution of all three problems by use of mechanochemical and barothermal methods applied to catalyst regeneration.

2. EXPERIMENTAL Mechanochemical treatment was carried out in a planetary mill at 3,000 rpm in the presence or absence of a dispersant. Details of the treatment were described in [15]. For barothermal treatment a stainless steel autoclave equipped with internal jar made of Teflon or titanium alloy OT 4-0 was used. Catalyst was loaded into the jar and the modifying compound was placed into the space between jar and autoclave walls [16]. The VPBiO catalyst was prepared following [15] from V205, H3PO4 and Bi(NO3)3 5H20 in butanol. The precursor obtained after drying was extruded into rings of Dout= 5 mm, dint - 2.4 mm and L = 5-6 mm. Catalyst agingwas carried out in an industrial reactor (dint = 26 mm, 1 = 3.5 m). Testing of the fresh, spent and regenerated catalysts was performed in a laboratory reactor (d = 6 mm, lcat- 40 cm) using a reaction mixture of 1.7 vol.% butane in air under a load of 45 g C4H10/h-kgeat 9 Barothermal treatments were applied directly to catalyst extrudates. In the course of mechanochemical treatments, extrudates were physically milled, loaded into a planetary mill and extruded again after the treatment. Catalyst solids were studied by means of XRD, XPS and their specific surface areas (SSA) were measured similar to [ 13].

3. RESULTS AND DISCUSSION

3.1. Regeneration of oxidized VPO catalyst Layer-by-layer analysis of VPBiO samples after 600 h use during n-butane oxidation in an industrial reactor [17] showed that the lower layers (about 60 cm) were composed of 13VOPO4. In accordance with [18] the reason for this phenomenon is a hydrocarbon concentration gradient along the catalyst bed. Continuation of catalyst aging up to 2,000 h does not effect any additional change in catalyst composition along the bed. The catalyst composed of mainly ~-VOPO4 (D 1) has low-selectivity for MA in n-butane oxidation (Table 1). XPS data confirm the presence of V 5+ ions on the catalyst surface. Mechanochemical treatment of VPBiO sample (D 1) suspended in water and in the absence of any dispersant did not lead to important changes in phase composition. During the catalyst treatment in ethanol in the presence of hydrazine some visible changes occurred in its XRD pattern. Thus, for the sample treated for 40 min (D l-M40) the weak peaks at d - 0.386 and 0.313 nm attributed to the (VO)EP207 phase appear against a significant halo-effect. The

209 binding energy of V 2p electrons decreases due to an appearance of V 4+ ions on the surface. Continuation of the treatment up to 60 and 90 min (D1-M90) leads to diminution/disappearance of halo and increase of the above mentioned peaks intensities. Moreover, the remaining reflections attributed to (VO)2P207appear which is an evidence of complete transformation of the solid bulk into vanadyl pyrophosphate during the treatment. Table 1 Properties of a VPBiO catalyst aged in an industrial reactor: lower layers before (D 1) and after regeneration b~r mechanochemical (M) and barothermal (B) methods. Sample a F D1 D1-M40 D1-M90 D1-B400/28 F-B300/14

SSA, m2/~ 14.5 13.2 12.5 12.8 10.0 14.0

XPS, V 2p 517.5 518.4 517.7 517.6 517.6 517.5

Eb, eV b P 219 133.9 133.8 134.0 133.9 133.8 133.5

XHC, %

80 77 73 76 71 79

n-Butane oxidation ~ SMA,% W 10 s, mol/s*m 2 70 1.19 38 1.26 49 1.26 61 1.27 57 1.53 77 1.22

a F - fresh catalyst (after 100 h of work loaded out of the central part of the catalyst bed), F-B300/14 - fresh catalyst after 14 h of barothermal treatment at 300iC; b Eb O ls = 531.6-532.2 eV, Eb Bi 4f = 159.7-160.2; e T = 415~ W - the rate of n-butane oxidation The ultrafine powder obtained from the treated samples resisted pellet formation by extrusion. Thus a pellet-pressing method was used. Sample surface properties and their catalytic activities are listed in Table 1. It is evident from the data that a significant improvement in catalytic performance (for example, the selectivity increases by 23%) as compared with spent catalyst (D 1) can be achieved as a result of mechanochemistry, but it is not possible to regain the initial efficiency of the samples (F). Barothermal regeneration of the spent catalysts (D 1) was carded out in the presence of nbutanol. At temperatures lower than 350~ no changes in the state of the sample were observed. In the case of catalyst treatment at T > 400~ (VO)2P207 formation occurred and its most intensive reflections were detected by XRD. However, the specific surface area of the final solid was low (about 4 m2/g). The best improvement in catalytic performance was achieved after treatment of the used catalyst for 28 h at 400~ (sample D 1-B400/28, see Table 1). XRD of this sample showed weak peaks of the (VO)2P207phase only. V 4+ ions were detected on the sample surface (XPS data). We also attempted to activate additionally the fresh catalyst pellets by use of barothermal treatment. It was established that a treatment with butanol vapors under moderate conditions restores catalyst effectiveness for n-butane oxidation (F-B300/14, see Table 1).

3.2. Regeneration of used VPO catalyst by introduction of phosphorus Table 2 shows the properties of the catalyst after aging for 2000 h and taken from the central zone of the industrial unit (D2). It is seen that the catalyst has a lower P/V surface

210 ratio and reduced catalytic performance during n-butane oxidation compared with the initial catalyst (F). Barothermal regeneration of the spent sample D2 with orthophosphoric acid vapors (Bpa) is found to enrich sufficiently the phosphorus content at the surface (Table 2). In so doing, the catalytic properties are improved as compared with D2 sample and the selectivity towards maleic anhydride reaches that of the fresh catalyst. Experiments were conducted to increase the catalytic efficiencies of the sample regenerated from the oxidized state by means of barothermal treatment (D 1-B400/28, Table 1) and of the fresh sample (F). It was established that the barothermal treatment with phosphoric acid has a positive effect in both cases (Table 2). Barothermal treatment of the D2 catalyst with P205 vapors (Bpo) increases its selectivity by 9-15% (Table 2), but its activity is unchanged. After the barothermal treatment of the spent catalyst D2 with phosphorus vapors (Bph) one can observe a new unindexed line at 0.337 nm in XRD spectrum, and the P/V ratio at the surface is found to be sharply increased (Table 2). The treated sample selectivity increases slightly but the activity, on the contrary, decreases. The latter can be connected with a covering of the active surface with phosphorus. Thus, the barothermal treatment with phosphorus vapors is not a suitable method to regenerate VPO catalysts. Table 2 The properties of VPBiO catalyst loaded out of the industrial unit central zone after barothermal regeneration Sample a SSA, XPS, Eb, eV b n-Butane oxidatione m2/~ V 2p P 2p XHC,% SMA,% W*108 mol/s*m2 F 14.5 517.5 133.9 1.47 80 70 1.19 D2 10.2 517.2 133.7 1.21 71 59 1.53 D2-Bpa250/10 9.7 517.2 133.9 1.69 77 68 1.71 D2-Bpa300/6 9.2 517.3 134.0 1.76 75 71 1.76 D2-Bpa300/14 8.5 517.3 133.8 1.94 70 66 1.77 D 1-B400/28-Bpa 9.5 517.4 133.9 1.54 70 68 1.59 F-Bpa 13.0 517.2 133.9 1.89 79 75 1.31 D2-Bpo250/6 9.4 517.3 133.5 1.64 70 68 1.64 D2-Bpo300/6 9.0 517.2 133.5 1.96 68 74 1.63 D2-Bph300/6 8.3 517.5 133.7 2.38 59 65 1.53 a The figures after treatment index mean temperature and time of the treatment (h) b Eb O ls = 531.5-532.2 eV, Eb Bi 4f = 159.7-160.3 eV; r T = 415 ~

3.3. Introduction of promoters into VPO catalyst Surfaces of special films placed at the outlet of the reactor during VPMeO catalyst aging for an extended period of time[ 18] were analyzed by XPS. It was found that promoting additives, such as Bi, Co, La, were present on the films indicating the removal of promoters from the

211 catalyst composition. But no significant change in the Bi/V ratio in the catalyst bulk was observed when the reaction took place in the industrial unit (see D2 in Table 3). Mechanochemical treatment of deactivated sample D2 with BiPO4 (D2/BiPO4-M) leads to an improvement of its catalytic properties. However, it is not possible to reach the initial efficiency of the fresh catalyst by such a method. This could be due to phosphorus loss from the catalyst during reaction and its insufficient introduction in the form of BiPO4. Apparently, repromotion of the spent catalyst should be carried out together with phosphorus introduction. Successful promoter introduction into VPO catalyst was performed using mechanochemical treatment of the mixture VPO+Bi203 (BiPO4) for 10-20 min in ethanol (see Table 3). During the barothermal (vapor-phase) treatment of sample D2 with Bi203 or BiPO4, low vapor pressures of these metal compounds prevented effective addition to the catalyst. Indeed, at a treatment temperature of 500~ for 10 h a specific surface area decrease occurred and the activity of the sample decreased (32% conversion), although selectivity was unchanged. Nevertheless, it was found that a treatment at lower temperature (e.g., 400~ and 350~ see Table 3), had less of a negative effect on surface area and hydrocarbon conversion; moreover, MA selectivity increased slightly. It is worth noting that to effectively introduce the promoter into the catalyst by barothermal treatment, the use of easily-melted promoter compounds is recommended. Table 3 The properties of VPO catalyst after promoter introduction Sample SSA, XPS~Eb~eV a XPS n-Butane oxidation b m2/~ V 2p P 2p (P/V) s (Bi/V)s Xuc,% S~tA,% W*108 F 14.5 517.5 133.9 1.47 0.09 80 70 1.19 D2 10.2 517.2 133.7 1.21 0.08 71 59 1.53 D2/BiPO4-M 8.5 517.4 133.8 1.45 0.11 74 65 1.88 VPO 12.0 517.5 133.9 1.43 73 67 1.31 VPO/BizO3-M 12.6 517.6 134.0 2.02 0.02 83 72 1.42 VPO/BiPO4-M 13.3 517.2 133.6 1.96 0.03 88 74 1.43 D2/BiPOa-B400/10 7.8 517.3 133.7 1.37 0.10 67 63 1.55 D2/BiPO4-B350/20 9.0 517.4 133.9 1.41 0.10 72 61 1.72 aEb O ls = 531.6-532.2 eV, EbBi 4f = 159.7-160.2 eV; b T = 415~ W - mol/s*m2 In summary, regeneration of VPMeO catalysts, deactivated because of oxidation and/or phosphorus loss, can be successfully conducted using mechanochemical or barothermal treatments. By these methods it is possible to increase the selectivity of the spent catalyst by 9-23% which is a better result than that obtained by treatment with phosphorus-organic compounds.

212 REFERENCES

1. 2. 3. 4. 5. 6.

7. 8. 9. 10. 11. 12. 13. 14. 15. 16. 17. 18.

Vanadyl Pyrophosphate Catalysts. G.Centi-Ed., Catal. Today, 16 (1993) 1-147. G.J. Hutchin~ s, Appl. Catal., 72 (1991) 1. F. Cavani and F. Trifiro, Preparation of Catalysts VI. Stud. Surf. Sci. Catal., Elsevier, Amsterdam, 91 (1995) 1. G. Centi, F. Trifiro, J.R. Ebner and V.M. Franchetti, Chem. Rev., 88 (1988) 55. T. Shimoda, T. Okuhara and M. Misono, Catalyst (Jap.), 27 (1985) 431. V.A. Zazhigalov, V.M. Belousov, G.A. Komashko, A.I. Pyatnitskaya, Yu.N. Merkureva, A.L. Poznyakevich, J. Stoch and J. Haber, Proceed. 9th Int. Congr. Catal., Ottawa, Inst. Chem. of Canada, 4 (1988) 1546; 5 (1988) 493. A. Satsuma, A. Hattori, A Furuta, A. Miyamoto, T. Hattori and Y. Murakami, J. Phys. Chem., 92 (1988) 2275. F. Garbassi, J.C.J. Bart, R. Tassinari, G. Vlaic and P. Lagarde, J. catal., 98 (1986) 317. J. Haas, C. Plog, W. Maunz, K. Mittag, K.-D. Gollmer and B. Klopries, Proceed. 9th Int. Congr. Catal., Ottawa, Inst. Chem. of Canada, 4 (1988) 1632. M. Becker and J. Walden, Optimizing the yield of maleic anhydride catalyst, US Patent No 4 795 818 (1989). R.C. Edwards, Process for regenerating and stabilizing phosphorus-vanadium-oxygen complex catalysts, US Patent No 4 861 738 (1989). V.A. Zazhigalov, V.M. Belousov, A.I. Pyatnitskaya, N.D. Konovalova and G.I. Goldenberg, React. Kinet. Catal. Lea., 39 (1989) 311. V.A. Zazhigalov, J. Haber, J. Stoch, A.I. Pyatnitskaya, G.A. Komashko and V.M. Belousov, Appl. Catal. A, 96 (1993) 135. J. Stoch, I.V. Bacherikova and V.A. Zazhigalov, 1st Ukrainian-Polish Seminar Catal., Cracow, Inst. Catal. Surf. Chem., 1994, 11. V.A. Zazhigalov, J. Haber, J. Stoch, L.V. Bogutskaya and I.V. Bacherikova, Appl. Catal. A, 135 (1996) 155. V.A. Zazhigalov, I.V. Bacherikova, V.E. Yaremenko, I.M. Astrelin and J. Stoch, Teoret. and Experim. Chem., 31 (1995) 206. V.A. Zazhigalov, V.M. Belousov, G.A. Komashko, A.I. Pyatnitskaya, Yu.N. Merkureva, A.S. Shulman and T.A. Ivanova, Petrol. Chem. (Russ.), 29 (1989) 824. V.A. Zazhigalov, V.M. Belousov, A.I. Pyatnitsakaya, G.A. Komashko, Yu.N. Merkureva and J. Stoch, New Developm. Select. Oxid., Stud. Surf. Sci. Catal., Amsterdam, Elsevier, 63 (1990) 617.

9 Elsevier Science B.V. All rights reserved Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

213

S u l f u r T o l e r a n c e o f C u - a n d H - m o r d e n i t e Zeolite C a t a l y s t s for the R e d u c t i o n o f NO by Hydrocarbons Moon Hyeon Kim, In-Sik Nam* and Young Gul Kim Research Center for Catalytic Technology, Department of Chemical Engineering, School of Environmental Engineering, Pohang University of Science & Technology / Research Institute of Industrial Science & Technology, PO Box 125, Pohang 790-600, Korea.

Sulfur tolerances of Cu- and H-mordenite zeolite catalysts prepared by ion-exchange were examined in a fixed-bed flow-reactor system. Rates of reduction of NO over HM or CuHM with C2H4 and CuNZA with C3H 6 are decreased by SO2 included in the feed gas stream. Surface areas and sulfur contents of the deactivated catalysts, their TGA and TPSR patterns and observations by XPS and Raman suggest the formation of a sulfur species on the catalyst surface in the form of sulfate (SO42-) which causes the loss of NO removal activity of the catalysts. Data from Cu K-edge absorption spectra suggest sulfur electrostatically interacts with Cu ions on the catalyst surface. 1. INTRODUCTION Selective catalytic reduction (SCR) of NO x by hydrocarbons is under investigation as an alternative NO x removal technology. NO reduction by NH 3 is presently the commercial stateof-the-art technology available for reducing NO x from stationary sources and from the exhausts of lean-burn gasoline and diesel engines. A number of catalysts for the selective reduction of NO by hydrocarbons have been examined in previous studies [1-18]. Transition metal ionexchanged zeolite catalysts such as mordenite and ZSM-5 are the most effective SCR catalysts. Most flue gases from the combustion sources contain, in addition to NO x, H20 vapor and SO 2, which may cause SCR catalysts to lose activity. Therefore, sulfur and water tolerances of catalysts are needed for the successful commercial application of the catalytic process. Since a dramatic loss of NO removal activity of transition-metal ion-exchanged zeolite catalysts was observed for NO reduction by hydrocarbons at wet conditions [1, 4, 9, 11-13], several studies have focused on the cause of activity loss by H20 [14-18]. However, effects of SO 2 on SCR catalysts employing hydrocarbons as a reductant have received considerably less attention. Low levels of SO 2 in the exhaust gas stream are known to lower the catalytic activity of noble metal-supported catalysts, although this effect is considerably less compared to base metal catalysts [ 19]. Iwamoto et al. [20] observed a slight decrease of NO removal activity of the Cu-ZSM-5 for NO reduction by C3H 6 upon the addition of SO 2 to the feed gas stream. The conversion of NO at 300~ however, completely recovered to the initial state upon termination of the SO 2 feed. They suggested that the loss of the removal activity of the catalyst is probably due to the alteration of the copper ionic state on the catalyst surface but provided no evidence in support of this hypothesis. Mabilon and Durand [21] observed that when 100 ppm of SO 2 was admitted to Cumordenite, a moderate decrease in activity for reducing NO by C3H 8 occurred, while the oxidation of C3H 8 w a s significantly inhibited by SO 2. No recovery of its oxidation activity was observed after the termination of SO 2 from the reaction stream, although NO removal activity was recovered. Recently, Li and Armor[22] reported the effect of SO 2, with and without HE0, in the feed gas stream on the activities of Co-ZSM-5 and Co-ferrierite for the reduction of NO by C H 4. The sulfur tolerances of their catalysts without H20 were a strong

214 function of the catalyst type as well as the reaction temperature. They speculated that activity loss was primarily due to the sulfur poisoning of the cobalt cation on the catalyst surface. The influence of SO 2 on SCR of NO by hydrocarbons may be reversible [20, 21] or irreversible [22, 23]. Activities and selectivities of the main and/or the deactivation reactions might be negatively affected. However, most of the studies thus far were not sufficiently definitive to reveal the cause/effect relationships for SCR catalyst deactivation by SO 2. In the present study, sulfur tolerances of Cu- and H- exchanged mordenite catalysts for NO removal reaction were examined to elucidate the cause of the loss of NO removal activity of the catalysts. Deactivation by SO 2 of Cu2§ sites was also investigated by various catalyst characterization methods including TGA, TPSR, Raman and XAS.

2. EXPERIMENTAL 2.1. Catalyst Preparation A synthetic mordenite (Zeolon 900Na) was obtained from the PQ corporation. A natural zeolite mined from Youngil, Korea was also used in this study [9]. HM was prepared by exchanging Na § ions of Zeolon 900 Na with NH4§ ions in an aqueous 1.5 N NH4NO 3 solution followed by filtering, drying at 110~ for 10 h, and calcining at 500~ for 10 h in air. CuHM was prepared by soaking a sample of the ammonium ion-exchanged mordenite in a 1.0 N Cu(NO3) 2 solution at 93~ for 20 h and by following the same procedure for drying and calcining. The content of Cu ions in CuHM was determined by atomic absorption spectroscopic analysis. Table 1 Physicochemical properties of mordenite-type zeolite catalysts Catalyst Cation content (wt.%) Cu Na Ca Fe Mg K HM 0.37 CuHM 3.54 0.27 NZ 0.86 1.99 0.76 0.63 2.37 NZA 0.16 Tr 0.34 0.16 1.33 CuNZA 2.01 0.17 Tr 0.37 0.11 1.20

employed in this study Si/A1 BET Cu/A1 area(mZ/g) 5.29 449 5.02 433 0.24 3.87 130 8.65 8.89 210 0.22

A natural zeolite (NZ) containing mainly mordenite-type zeolite was treated with a 1.0 N HC1 solution at 93~ for 20 h in order to remove impurities contained in the ore and to stabilize the pore openings of the zeolite structure [9]. The CuNZA catalyst was prepared by further treatment of NZA (natural zeolite treated with acid) with a 1.5 N NH4NO 3 solution followed by drying and calcining at 500~ for 10 h and by repeatedly soaking in a 1.0 N Cu(NO3) 2 solution at 93~ for 40 h. The physicochemical properties of the catalysts prepared in this study are listed in Table 1.

2.2. Reaction System Sulfur tolerances of the zeolite SCR catalysts (20/50 mesh size) were determined in a fixedbed flow reactor system with a reaction mixture containing 500 ppm of NO, 1,000 ppm of C2H 4 or 2,000 ppm of C3H6, 4.2% of 02, 1,000 ppm SO 2 and He (balance). To observe the simultaneous effect of both SO 2 and H20 on the removal reaction, 7.3% H20 was also fed to the reaction system in addition to the reactants described above. H20 was injected to the feed gas stream by bubbling He into a water saturator with a small-pore frit immersed in deionized water. To avoid condensation of H20 vapor after the bubbler, reactor lines were heated to a temperature higher than the saturation temperature of the feed gas stream including H20. A gas flow rate of 300 cm3/min was employed for the present study, corresponding to a space velocity of

13,200 h1.

215

Reactants and products were analyzed by a gas chromatograph (Hewlett Packard 5890 Series II) equipped with a thermal conductivity detector. 13X molecular sieve and Porapak Q columns were used for the analyses of N 2, CO and 02 and of hydrocarbons, CO 2 and N20, respectively. The catalytic activities of the catalysts employed in this study for NO removal by hydrocarbons were evaluated in terms of NO conversion to N 2 as 2[N2]ouT/[NO]IN.

2.3. Catalyst Characterization BET surface areas of the SO 2 deactivated catalysts were measured with a Micromeritics Accusorb 2100E using liquid N 2 at 77 K after the catalysts had been pretreated in vacuo at 180~ for 10 h. Sulfur and carbon contents of the deactivated catalysts were determined by an oxidation method with a LECO SC-132 Sulfur Systems and CS-044 Carbon & Sulfur 781-000 Systems (LECO Co.). Changes in catalyst weight during exposure to SO 2 were measured by a SEIKO I 300 TGA system. TGA patterns were obtained within the temperature range from 25 to 800~ The flow rate of the high purity He (99.9999%) carrier gas was 40 cm3/min, and the temperature ramping rate was 10~ respectively. TPSR for the catalysts deactivated by SO 2 was made with a quadrupole mass spectrometer (MMPC-200D, VG Quadrupoles). Deactivated catalysts were placed in a quartz U-shaped reactor of 1/4" O.D. The reactor was surrounded by a cylindrical electric furnace which was controlled by a PID temperature controller with a K-type thermocouple. After the reactor was fully purged with He (99.9999%) at 50~ its temperature was ramped from 50 to 800~ at a heating rate of 10~ for identification by the mass spectrometer of compounds desorbing from the catalyst surface, e.g. H20, CO, CO 2, CH4, SO 2 and SO 3. XPS spectra for the SO2-deactivated catalysts were obtained using a Perkin-Elmer PHI 5400 XPS spectrometer using Mg K-ix radiation to observe the S 2p line. Charging effects were corrected for by using the carbon peak at 284.6 eV as a standard. Raman spectra for SO2-deactivated HM, CuHM and CuNZA were examined by a Spex 20 spectrometer equipped with a Spex Datamate computer. Zeolite samples were excited using the 514.5 nm line of a Spectra Physics 171 Ar ion laser powered with a Spectra Physics 265 exciter. To minimize local heating effects where the laser beam impinges on the catalyst surface, a rotating lens assembly was employed. XAS experiments were performed at Beam Line 7C of Photon Factory in National Laboratory for High Energy Physics (KEK-PF), Japan. The synchrotron radiation from the storage ring (2.5 GeV, 300 - 250 mA) was monochromatized by a channel-cut Si (111) crystal monochromator. The samples were ground into powder and then pressed into self-supporting wafers without any binder. Cu K-edge absorption spectra were collected for the coppercontaining catalysts such as CuHM and CuNZA at room temperature before and after the reaction for several hours under SO2-exposed condition, as well as for the references such as Cu(OH)2 and CuSO4.5H20. The Fourier transformations of EXAFS for the copper ions were carried out over the ranges of photoelectron wave vector, k, of 2.6 to 11.7 A-~. 3. RESULTS AND DISCUSSION 3.1. Effects of S O 2 o n NO Removal Activity of Mordenite-type Zeolite Catalysts Figure 1A shows activity versus time for the HM catalyst for NO reduction by C2H 4 with and without SO 2. About 60% NO conversion was maintained at 360~ for the SO2-free condition during 11 h of continuous operation. However, its NO removal activity gradually decreased about 10% within 8 h when 1,000 ppm of SO 2 was introduced to the reaction system. Conversion of the CuHM catalyst (58%) in the absence of SO 2 is constant with time (Figure 1B) while conversion decreases from 58 to 35% after 13 h of operation with SO 2. A more rapid loss of NO removal activity is evident for this catalyst relative to the copper-free catalyst.

216 80

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Figure l. Sulfur tolerance of the synthetic mordenite catalysts : HM (A) and CuHM (B). (O) without SO 2 ; (0) with SO 2. Reaction condition : NO 500 ppm, C2H 4 1,000 ppm, 0 2 4.2%, SO 2 1,000 ppm and T=360~ 100

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Reaction time (hr) Figure 2. Sulfur tolerance of CuNZA catalyst : (O) without SO 2 ; O ) with SO 2. Reaction condition : NO 500 ppm, C3H 6 2,000 ppm, O 2 4.2%, SO 2 1,000 ppm and T=400~

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Reaction time (min) Figure 3. Effect of H20 on NO removal activity of CuHM (@) and CuNZA LiD) catalysts. Reaction condition : N O 500 ppm, C2H4 1,000 (Q) or C3H 6 2,000 ppm (ll), 02 4.2%, 1-120 7.3% and T=360 (0) or 400~ (ll).

The effect of SO 2 on the CuNZA-C3H 6 SCR reaction system was also examined (Figure 2). In the absence of SO 2, a steady NO conversion of about 85 to 90% is observed for 10 h of reaction at 400~ NO conversion, however, is decreased more than 60% (to about 25%) when SO 2 is fed to the reaction system. Thus, this catalyst is very sensitive to catalyst deactivation by SO 2, even though its NO removal activity is higher than any other catalytic system examined in this study in the absence of SO 2, with or without H20 vapor present [ 16]. Data for the sulfur tolerance of NZA catalyst are not included in this paper since it exhibited

217

low NO removal activity within the range of reaction conditions covered in the present study, i.e., less than 20% NO conversion even for the SO2-free condition. The decreasing trend for oxidation of C2H 4 and C3H 6 in the presence of SO 2 was similar to that for NO conversion, regardless of the types of catalyst employed in this study. No activity loss by carbonaceous compounds deposited on the surface of the catalysts is anticipated during the course of reaction without SO 2. 100

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Figure 4. Effects of both SO 2 and H20 on the loss of NO removal activity of CuHM ( 9 ) and CuNZA ( B ) catalysts. 1,000 ppm of SO 2 was admitted in addition to the reaction conditions described in Figure 3.

3.2. Simultaneous Effects of H20 and SO 2 on NO Removal Activity o f Mordenite-type Zeolite Catalysts Water tolerances of the CuHM and CuNZA catalysts were examined for NO removal with and without SO 2. Figure 3 shows their water tolerances and reversibilities for NO reduction with 7.3% H20 as a function of reactor on-stream time. NO conversion of the CuHM catalyst immediately drops from 63 to 20% conversion within 5 min after H20 injection, while conversion decreases less than 10% for the CuNZA catalyst. The effect of H20 on NO removal activity is reversible, in agreement with previous work [ 16]. Figure 4 shows decreases in NO conversion efficiencies of CuHM and CuNZA catalysts due to the combined presence of SO 2 and H20 in the feed gas stream. For the CuHM catalyst, NO conversion decreases from 60 to 20% within 10 min after the simultaneous injection of SO 2 and H20 into the reaction stream. This effect is quite similar to the activity loss of the catalyst for NO reduction by C2H 4 with

Table 2 Physicochemical properties of the catalysts deactivated by S O 2 Catalyst

Sulfur content (wt. %)

BET surface area (m2/g)

HMa HMb HMc

0.00 1.10

449 419 384

CuHMa CuHMb CuHMc

0.00 1.86

433 418 207

CuNZAa CuNZAb CuNZAc

0.00 0.54

210 210 139

a fresh catalyst. b after reaction without SO 2. c after reaction with S O 2 of 1,000 ppm.

218

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H20 (Figure 3). However, the loss of NO removal activity of this catalyst by SO 2 alone was negligible during the initial 10 minutes of reaction (Figure 1B). This suggests that H20 plays a dominant role in the rapid loss of NO removal activity for Cumordenite while the deactivation by SO 2 is apparently negligible. In the presence of both SO 2 and H20, the NO conversion of CuNZA decreases to a steady state of about 30% within 2 h. This activity drop is similar to that observed with only SO 2 in the feed gas stream. Therefore, the apparent loss of the catalytic activity over the CuNZA catalyst is mainly due to the presence of SO 2. No synergistic effect of catalyst deactivation by both SO 2 and H20 is observed for either catalyst type. Although the simultaneous effect of SO 2 and H20 on the performance of HM was not examined in the present study, H20 vapor included in the feed gas stream may have dominated the activity loss since a drop in NO conversion from 67 to 20% was observed previously for this catalyst with 7.3% of H20 in the feed stream [ 16].

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Figure 5. TGA of the deactivatrA catalysts by SO2: Compared to the TPSR pattern of HM (A); CuHM (a); CuNZA (C). (a) fresh HM catalyst deactivated by SO 2, Cu catalyst; (b) after reaction with SO2; (c) after ions on the surface of the copper reaction without SO 2.

3.3. Characterization of Deactivated Catalysts by SO2 BET Surface Area and Sulfur Content. Surface areas of mordenite-type zeolite catalysts after SCR of NO with hydrocarbons without SO 2 were quite similar to those of the fresh catalysts, as listed in Table 2. This supports the hypothesis that little or no surface carbonaceous compounds formed on the catalyst surface. About 10% of the surface area of the HM catalyst was lost after exposure to SO 2, while more than 50% of the surface area of the CuHM catalyst was lost after

219

A

HM

exposure to SO 2. For the CuNZA catalyst, about 30% of its surface area is lost after exposure to a feed of SO 2. The correlation of a BET surface area loss with sulfur content of the SO Eexposed catalysts suggests that sulfur deposits poison or blocks the surface. A similar trend of decreasing activity and surface area with increasing sulfur content was also observed by Ham et al. [24] for the reduction of NO by NH 3 over a CuHM catalyst.

TGA and TPSR. TGA of S O 2 deactivated catalysts was conducted to w eobserve the decomposition of the deactivating agent on their surface. c Weight changes of three samples of HM CuHM catalyst (fresh, after reaction and after SO 2 exposure) with increasing temperature axe shown in Figure 5A. A continuous weight loss up to about 500~ was observed for the fresh catalyst, which is mainly due to the desorption of the : I I t I I ; adsorbed water on the catalyst surface. Its 50 250 450 650 8()0 weight loss pattern after the reduction of NO by hydrocarbons without SO 2 was Ramping temperature (~ similar to that of the fresh catalyst. However, the pattern for the catalyst Figure 6. Desorption of S O 2 during TPSR of the deactivated by SO 2 is different, i.e., turns deactivated catalysts, downward above about 470~ indicating that the decomposition of a sulfur compound on the catalyst surface may occur at temperatures higher than 470~ Figure 5B shows the TGA curves for the fresh and treated CuHM catalyst samples. Although the weight loss patterns for the catalyst without SO 2 were similar to those of HM catalyst, the curve for the deactivated CuHM catalyst drops more sharply from 420 to 680~ i.e., the weight loss for the CuHM catalyst is greater than for the HM catalyst. The weight loss of the deactivated CuNZA catalyst (Figure 5C) began by dropping more rapidly at about 400~ Thus, TGA for the deactivated catalysts strongly suggests the presence of a sulfur compound on the catalyst surface which can be decomposed at temperatures above 400-450~ From TPSR (Figure 6), desorption of SO 2 is observed at temperatures higher than 400~ regardless of catalyst type. A desorption peak maximum at 630~ is observed for the HM catalyst. For the CuHM catalyst, three SO 2 desorption peaks occur at 450, 600 and 680~ One SO 2 desorption peak is observed at about 600~ for CuNZA. From the TPSR data, it is evident that (1) NO sulfur-containing compounds on the surfaces of the previously SO2-exposed catalysts decompose to SO 2 at temperatures above about 500~ (2) there is a common peak for all 3 catalyst at 600-650~ that may involve a non-copper sulfur-containing species and (3) sulfur-containing species are present on the CuHM catalyst which decompose at higher temperatures than on HM and CuNZA catalysts indicating that they interact strongly with the surface and possibly with Cu ions on the catalyst surface. It should be mentioned that the amount of SO 3 desorbed from all the catalysts was relatively negligible.

lid ,m

m

220

Thus, BET surface area, sulfur content, TGA, and TPSR data strongly support the conclusion that sulfur species are formed on the catalyst surface of Cu- and H-mordenite catalyst during exposure to SO~. These species bind strongly to the surface of the catalysts but can be decomposed to SO 2 at temperatures above 500~ These species probably play a major role in the loss of the catalytic activity during the course of reaction in the presence of SO 2. XPS and Raman. S 2p XPS spectra for catalysts treated with SO 2 were obtained to identify the oxidation state of sulfur compounds on their surface. Figure 7 shows the S 2p bands for catalysts exposed to SO 2. An S 2p line at 169.4 eV is observed for the synthetic zeolite surface of these catalysts probably is a sulfate (SO42). It should be noted that the sulfate species on CuNZA was of lower binding energy than that on the synthetic zeolite catalysts. This suggests that the sulfate species on the CuNZA catalyst is less solid-like than that on the synthetic zeolite catalysts [30]. This is also supported by the TPSR spectra, where the sulfur compounds on the HM and CuHM catalysts exhibited higher decomposition temperatures than that on the CuNZA catalyst. Raman spectra for SO2-deactivated HM and CuM are shown in Figure 8 A shift band is evident at about 975 cm 1 for both catalysts after the reduction of NO with SO 2. When the fresh catalysts treated with a 0.2 N H2SO 4 solution, followed by drying at 300~ for 3 h, were employed as a reference, they also exhibited the Raman shift similar to the deactivated catalysts by SO 2. This indicates that the peak at 975 cm 1 is mainly due to a sulfate formed on the catalyst surface. Similar results were obtained for CuNZA.

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XAS (XANES and EXAFS). A Cu Kedge absorption peak at 8974 eV due to the

CuNZA 173

171

169

187

185

Binding energy (eV) Figure 7. XPS of the catalysts deteriorated by SO2: (a) after reaction without SO2; (b) after reaction with SO 2.

l s-3d transition observed for the SO2-exposed catalysts is characteristic of the Cu K-edge structure of divalent copper [31 ], regardless of the samples employed in this study. It indicates that the electronic structure of Cu ions in the catalysts is not significantly altered even after the reaction with SO 2. The ls-4pt~ (in-plane) transition [32] was observed for a cupric sulfate pentahydrate as a reference, as

221 well as for the catalysts deactivated by SO 2, but the ls,4Pn (out-0f-plane) transition [32, 33] was observed for the reference only. This suggests that the sulfate species formed during the course of reaction with SO 2 is not similar to a cupric sulfate pentahydrate which contains a copper ion surrounded by a tetragonally distorted oetahedron of oxygen atoms [34]. The Fourier transformed Cu K-edge k3-weighted EXAFS for CuHM catalyst exposed to the feed stream including SO 2 can be found in Figure 9A. Only a single major peak at 1.54 /~ corresponding to the Cu-O bond appeared for all the samples employed in this study; this result is in good agreement with previous EXAFS studies for CuY [35] and for cupric sulfate pentahydrate [36]. A similar Cu K-edge EXAFS spectrum exhibiting a catalysts, HM and CuHM. Sulfide and elemental sulfur exhibit the binding energies at 161.0 to 162.8 eV [25-27] and at 164.0 eV [25, 27, 28], respectively. Peaks at 167.8 to 171.0 eV are identified with sulfates [25, 29]. This indicates that a sulfur compound deposited on the prominent absorption of Cu-O bond at 1.51/k is also observed for the CuNZA catalyst, as shown in Figure 9B. Cu Kedge EXAFS spectra of the SO2-exposed catalysts, however, were similar to those of, but not the same as, reference copper compounds such as Cu(On)2 and C u S O 4 9 5H20. It should be noted that the sulfate species on the catalyst surface as identified by XPS and Raman does not exist in the form of a cupric sulfate pentahydrate.

(B)

(A)

(a)

d,

_=

I

200

500

!

I

I

800 4400 4400 1700 2000

Raman shift (crn "t)

200

"

500

I

I

I

I

800 1100 1400 1700 2000

Raman shift (cm "~)

Figure 8. Raman spectra of the synthetic mordenite catalysts deactivated by SO2: HM (A) and CuHM (B). (a) fresh catalyst; (b) after reaction with SO2; (c) fresh catalyst treated with a 0.2N nESO 4 solution.

222

Since Cu ions on the zeolite surface exist in an isolated environment, they may interact with the sulfate species on the catalysts deactivated by SO 2. The sulfate groups might partially surround the copper ions, as previously suggested by Choi et al. [37] and Hamada et al. [38]. The sulfate may also have some characteristics of a coordinate covalent bond, where Cu ions and sulfate species may act as a Lewis acid and base, respectively2 [39]. Ligands such as H20, NH 3, (C2Hs)3P, CO molecules and CI, CN, OH, NO2, and C204 " ions should at least contain a lone pair of electron to form a coordinate covalent bond between metal ions [40]. It should be noted that the sulfate catalyst species formed on the catalysts deactivated by SO 2 contains lone electron pairs on O atoms which surround S atom of SO 4 groups. Therefore, it is expected that the electrostatic interaction between Cu ions and sulfate species probably influences the local structure of Cu ions on the zeolite catalyst surface. 12

10 "il

12 ~

(A)

-,,

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8

'0 :~

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Figure 9. Cu K-edge EXAFS spectra for CuHM (A) and CuNZA (B) catalysts deteriorated by SO 2. ( - - - ) fresh catalyst " ( ---- ) after reaction without SO 2 ; ( ) after reaction with SO 2 ; ( m - - r e . ) Cu(OH)2" ( ~ " m ) CHSO4 t 5H20. 4. C O N C L U S I O N S a. H-mordenite and Cu-mordenite catalysts lose 20-80% of their activity over a period of a few hours during reaction in the presence of 100 ppm SO 2. Copper ion-exchanged catalysts such as CuHM and CuNZA are deactivated to a greater extent by SO 2 than the copper-free catalyst (HM). b. Effects of both SO: and/or H:O on the loss of their NO removal activity vary with catalyst composition. CuM is deactivated by H20; a natural CuNZA catalyst is more severely deactivated by SO 2. c. More sulfur is deposited on the CuHM catalyst than on the HM catalyst, consistent with the observed greater loss of BET surface area for CuHM. This is also consistent with the observation that the HM catalyst undergoes milder deactivation than the CuHM catalyst. d. TGA and TPSR data for the deactivated catalysts strongly suggest the formation of sulfur compound on the zeolite surface. This sulfur species is strongly bounded to the surface of the catalysts deactivated by SO 2 and is decomposed to SO 2 at temperatures higher than 500~ e. Based on XPS, Raman, SANES, and EXHFS data, the sulfur-containing species exists in the form of sulfate (SO42) but in a form different from that of cupric sulfate pentahydrate.

223 The somewhat altered local structure of Cu ions on the catalysts is probably due to the electrostatic interaction between Cu ions and SO42 on the zeolite catalysts. REFERENCES

.

6. .

8. 9. 10. 11. 12. 13. 14. 15. 16. 17. 18. 19. 20. 21. 22. 23. 24. 25. 26. 27. 28. 29. 30. 31. 32. 33. 34.

M. Iwamoto and N. Mizuno, J. Auto. Eng., 207 (1993) 23 and therein references. W. Held, A. Konig, T. Richter and L. Puppe, SAE paper No. 900496 (1990). T.J. Truex, R.A. Searles and D.C. Sun, Platinum Met. Rev., 36 (1992) 2. Y. Li and J.N. Armor, (a) Appl. Catal. B, 2 (1993) 239 ; (b) Appl. Catal. B, 3 (1993) L1 ; (c) J. Catal., 145 (1994) 1 ; (d) Appl. Catal. B, 1 (1992) L31. K. Yogo, M. Umeno, H. Watanabe and E. Kikuchi, Catal. Lett., 19 (1993) 131. M.C. Demicheli, L.C. Hoang, J.C. Menezo and J. Barbier, Appl. Catal. A, 97 (1993) Lll. T. Tabata, M. Kokitsu and O. Okada, Catal. Lett., 25 (1994) 393. X. Zhang, A.B. Waiters and M.A. Vannice, J. Catal., 146 (1994) 568. M.H. Kim, I.-S. Nam and Y.G. Kim, Appl. Catal. B, 6 (1995) 297. C.J. Loughran and D.E. Resasco, Appl. Catal. B, 5 (1995) 351. Y. Li, P.J. Battavio and J.N. Armor, J. Catal., 142 (1993) 561. R. Gopalakrishnan, P.R. Stafford, J.E. Davidson, W.C. Hecker and C.H. Bartholomew, Appl. Catal. B, 2 (1993) 165. K.A. Bethke, D. Alt and M.C. Kung, Catal. Lett., 25 (1994) 37. K.C.C. Kharas, H.J. Robota and D.J. Liu, Appl. Catal. B, 2 (1993) 225. T. Tabata, M. Kokitsu, O. Okada, T. Nakayama, T. Yasumatsu and H. Sakane, in B. Delmon and G. F. Froment (Eds.), Catalyst Deactivation 1994 (Studies in Surface Science and Catalysis, Vol. 88), Elsevier, Amsterdam, 1994, p. 409. M.H. Kim, I.-S. Nam and Y.G. Kim, (a) Stud. Surf. Sci. Catal., 105 (1997) 1493 ; (b) Appl. Catal. B, in press. R.A. Grinsted, H.W. Zhen, C. Montreuil, M.J. Rokosz and M. Shelef, Zeolites, 13 (1993) 602. J.Y. Yan, G.-D. Lei, W.M.H. Sachtler and H.H. Kung, J. Catal., 161 (1996) 43. T.C. Summers and K. Baron, J. Catal., 57 (1979) 380. M. Iwamoto, H. Yahiro, S. Shundo, Y. Yu-u and N. Mizuno, Appl. Catal., 69 (1991) L15. G. Mabilon and D. Durand, Catal. Today, 17 (1993) 285. Y. Li and J.N. Armor, Appl. Catal. B, 5 (1995) L257. F. Suganuma, M. Tabata, K. Miyamoto, M. Kawatsuki, T. Yoshinari, H. Tsuchida, H. Hamada, Y. Kintaichi, M. Sasaki and T. Ito, Shokubai, 35 (1993) 114. S.-W. Ham, H. Choi, I.-S. Nam and Y.G. Kim, Ind. Eng. Chem. Res., 34 (1995) 1616. J.F. Moulder, W.F. Stickle, P.E. Sobol and K.D. Bomben, in J. Chastain (Edr.), Handbook of X-ray Photoelectron Spectroscopy : A Reference Book of Standard Spectra for Identification and Interpretation of XPS Data, Perkin-Elmer Corp., Minnesota, 1992. V.I. Nefedov, Y.V. Salyn, P.M. Solozhenkin and G.Y. Pulatov, Surf. Interf. Anal., 2 (1980) 170. K. Ebitani, H. Konno, T. Tanaka and H. Hattori, J. Catal., 135 (1992) 60. T. Wang, A. Vazquez, A. Kato and L.D. Schmidt, J. Catal., 78 (1982) 306. E. Xue, K. Seshan, J.G. van Ommen and J.R.H. Ross, Appl. Catal. B, 2 (1993) 183. C.D. Wagner and J.A. Taylor, J. Electron Spectrosc. Relat. Phenom., 28 (1982) 211. W.P. Griffith, J. Chem. Soc. A, (1969) 1372. R.A. Bair and W.A. Goddard III, Phys. Rev. B, 22 (1980) 2767. N. Kosugi, T. Yokoyama, K. Asakura and H. Kuroda, Chem. Phys., 91 (1984) 249. T.A. Smith, J.E. Penner-Hahn, M.A. Berding, S. Doniach and K.O. Hodgson, J. Am. Chem. Soc., 107 (1985) 5945.

224 35. 36. 37. 38. 39. 40. 41.

G.E. Bacon and N.A. Curry, Proc. Roy. Soc. A, 266 (1962) 95. S. Tanabe and H. Matsumoto, (a) Chem. Lett., (1989) 539 ; (b) Bull. Chem. Soc. Jpn., 63 (1990) 192. R.W. Joyner, Chem. Phys. Lett., 72 (1980) 162. E.Y. Choi, I.-S. Nam, Y.G. Kim, J.S. Chung, J.S. Le~ and M. Nomura, J. Mol. Catal., 69 (1991) 247. H. Hamada, N. Matsubayashi, H. Shimada, Y. Kintaichi, T. Ito and A. Nishijima, Catal. Lett., 5 (1990) 189. J.E. Huheey, Inorganic Chemistry : Principles of Structure and Reactivity (3rd Edn.), Harper & Row, New York, 1983. F. Basolo and R.C. Johnson, Coordination Chemistry (2nd Edn.), Whitstable Litho, New York, 1986.

~ Elsevier Science B.V. All rights reserved Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

225

D e a c t i v a t i o n of C u - Z S M - 5 during Selective Catalytic R e d u c t i o n of N O b y P r o p a n e under Wet Conditions A. Martfnez, S.A. G6mez, and G.A. Fuentes Area de Ingenieria Quimica, Universidad A. Metropolitana-Iztapalapa A.P. 55-534, 09340 M6xico, D.F., MEXICO. E-mail: [email protected] The presence of H20 in the feed during NO reduction over Cu-ZSM-5 caused kinetic inhibition together with reversible and irreversible deactivation. These phenomena were sensitive to the operating conditions. Kinetic inhibition had small characterisUc times and was associated with competition for active sites between H20 and the reactants. Slow recovery of activity after H20 removal from the feed after long runs characterized reversible deactivation, probably caused by solid state rearrangements in the catalyst. Irreversible deactivation had a quasi-sigmoidal character, with a maximum rate occurring after more than 15 h on stream. Neither structural changes or dealumination of the zeolite were detected. Temperature programmed reduction showed that the state of Cu cations changed significantly during the diverse treatments. Bands associated with Cu 2+, Cu + and CuO were observed in different raUos, as well as a band at 593 that could not be associated with any particular species. In spite of those changes the state of Cu species did not correlate well with the activity of the catalysts. 1. INTRODUCTION Selective catalytic reduction of NOx with hydrocarbons (SCR-HC) under lean-burn engine conditions is an alternative for decreasing emissions that trigger photochemical smog formation in urban settings. Its application has been restricted by the lack of catalysts able to withstand operation under wet oxidizing atmospheres at high temperatures. So far, catalysts prepared by exchanging transition metal cations into ZSM-5 have been reported as having the best overall performance [ 1,2], but their stability under realistic operating conditions [3,4] is still questionable. Cu-ZSM-5 is one of the best catalysts for SCR-HC under dry conditions [5,6]. However, by using low H20 concentrations in the feed (2 2 %) it undergoes reversible deactivation [7,8,9]; above 10 % H20, deactivation becomes irreversible [10,11]. Among the different hypotheses about what causes deactivation, dealumination of the zeolite [10,12] and formation of inactive Cu species [ 11] seem the most plausible, yet the actual mechanism has not been fully identified. In this paper we report the kinetic behavior during NO conversion on Cu-ZSM-5 as a function of H20 concentration, as well as the characterization of fresh and spent catalysts using FT-IR, XRD, MAS-NMR and TPR. It is clear that the comprehension of the phenomena that occur under wet conditions is necessary in order to develop viable commercial catalysts. 2. EXPERIMENTAL

We prepared Cu-ZSM-5 catalysts by ion exchange of Na-ZSM-5 (Zeocat Pentasil PZ-2/54 Na, Si/A1 = 24.3) with an aqueous solution of Cu(CH3COO)2 (J.T. Baker). The samples were dried at 383 K before storage, and the Cu loading attained was 2.3 wt. %, equivalent to 6 1 % of ion exchange calculated as Cu(OH)+/A1. This catalyst was designated 2.3-Cu-ZSM-5. Catalytic

226 testing was done in a microreactor with on-line analysis by GC (Shimadzu 12A with Porapak Q column) and by a NO/NOx analyzer (Rosemount 951A). The standard reaction mixture was 1650 ppm of NO, 4.1% of 02, 3467 ppm of C3H8 and different amounts of H20 (0-10% v/v) introduced to the system through a saturator. The mixture balance was N 2 and the total flow was 150 cm3/min. Two space velocities were used: 180 000 and 22 000 h-1. Temperature programmed reduction (TPR) was performed in situ in a stream of 25 cm3/min of 10 % H2/Ar and a heating rate of 10 K/min. A thermal conductivity cell was used to follow H2 consumption. In the case of fresh samples, they were previously dried at 423 K under flow of N2 and cooled down to room temperature before running the TPR experiment. FT-IR spectra (Bruker IFS-66) were obtained by diffuse reflectance; the samples were dried beforehand. The catalysts were analyzed by XRD (Siemens D500) using a Cu K a radiation source filtered by Ni, and using 29Si and 27A1 MAS-NMR (Bruker ASX300WB) at spinning rates of 3.5 and 12 KHz respectively. 3. R E S U L T S AND D I S C U S S I O N

3.1 Reaction studies During temperature scanning 80 IE experiments with heating up to 923 K and cooling cycles at a space velocity of 0% of H20 70 [: 180,000 h -1, we found that NO reduction over 2.3-Cu-ZSM-5 was affected by H20 1.6% of H20 60 primarily by kinetic inhibition. Both NO conversion and the temperature of 5% of H20 = 50 0 maximum conversion were a function of H20 concentration (Figure 1). This ~ 40 ;> suggests that H20 was competing with the reactants for active sites. Our results were 0 ~ 30 in agreement with other authors [ 13,7,8], 0 who reported that the addition of ___2% of 2:20 H20 in the feed decreased NO conversion, but when H20 was eliminated from the 10 feed, the conversion returned to its original value. We also performed constant 0 ~ temperature experiments at 673 K and 200 300 400 500 600 700 800 900 1000 180,000 h -1 for up to 12 h on stream. In order to get a hold on the kinetic Temperature K dependence with respect to H20, deactivation trace can be approximated as Figure 1. Scanning temperature reactions during straight lines even for H20 concentrations NO reduction of 2.3-Cu-ZSM-5 catalysts with as high as 10% in the feed, as shown in different amounts of H20 in the feed. SV= 180,000 Figure 2. During experiments at low space h-1. velocity (22,000 h-l), no deactivation was detected up to 12 h on stream. We only observed an inverse first order kinetic dependence on H20 concentration (0 to 10% in the feed) between 673 and 923 K. Indicating that H20 acted only as a competitor against the reactants. At both levels of space velocity studied, NO conversions never reached 100% even under dry conditions. This suggests the presence of a selectivity limit given the competition between reactions of HC with NO and with 02 [ 11,9]. 9 ,...i r~

227 100

In order to force a significant degree of deactivation on the catalyst, we performed an experiment alternating dry and wet 80 operation and high and low temperature at a)" 180,000 h -1. The reaction was started at 673 K under dry conditions. After 12 h on stream the reaction was stopped and the =o 60 system was purged using N2. At this point, the temperature of the reactor was increased to 923 K under flow of N2. The o 40 Do reaction was then restarted, but in the 9 presence of 10% H 2 0 in the feed. After 9 25 h the reaction was stopped and the reactor 20 was cooled down to 673 K in N2. The activity was then probed again against the original reactant mixture. During this I I I I I I I experiment, the catalyst lost 60% of its 100 200 300 400 500 600 700 800 initial activity for NO reduction (Figure 3a) and 80% for C3H8 conversion (Figure Time on stream (min) 3b). This catalyst will be referred to as Figure 2. NO reduction vs. time on stream at 673 2.3-Cu-ZSM-5/T. K. a) 0%, b) 1.6%, c) 5% and d) 10% H20 in the A long term experiment was performed feed. SV=180,000 h -1 at 673 K with 14% of H20 in the feed at 180,000 h-1. The temperature was =

a)

b) ,oo

of.

~

50 40

80

923 K, 10% of H20 .~

60

30

4O

20

2O

10 / lli,

0

I,,,,

500

I , , , , I , , , , I , ,

1000

1500

2000

Z

0 Catalyst Before Deactivation

DeactivatedCatalyst

Time on Stream (min) F i g u r e 3. Effect of reaction at 923 K and 10% HzO in the feed. a) NO conversion vs. time on stream, b) C3H8 conversion at 673 K measured under dry conditions. SV=180,000 h ~ increased from room temperature to 673 K in the presence of a dry reaction mixture. Once the

228

temperature was stabilized, H20 was introduced to the feed. There was an immediate decrease in NO conversion (Figure 4), caused by the kinetic inhibition. The conversion then decreased slowly for up to 1100 minutes, but after about 1200 minutes on stream we observed a rather peculiar fast decay process that followed essentially a sigmoidal curve. This behavior could correspond to the activation of a solid-state process, perhaps related with a critical concentration of active sites. In the case of the long term test in 100 the presence of 14% of H20 in the feed, the catalyst lost its activity completely after 25 h on stream, but 80 when H20 was eliminated from the feed, the catalyst slowly recovered 60 about 30% of its initial activity for NO reduction. The loss of activity measured against a standard dry feed H20-off 40 > was 66% for NO reduction and 80% for C3H8 oxidation. This catalyst was referred as 2.3-Cu-ZSM-5/W. We 0 20have found that by decreasing the Z amount of H20 in the feed non-zero pseudo-steady state conversion could be reached. The percentage of NO reduction _20 i . . . . i . . . . I . . . . i . . . . I . . . . I . . . . activity lost for both 2.3-Cu-ZSM-5/T and 2.3-Cu-ZSM-5/W deactivated 0 500 1 0 0 0 1500 2000 2500 3000 catalysts was similar to the value Time on Stream (min) reported by Yan et al. [14]. For a deactivated catalyst they determined a Figure 4. Long term evaluation of NO reduction at loss of 60% in conversion during NO 673 K in a stream containing 14% H20. reduction. However, they made their SV=180,000 h-1 activity measurements under dry conditions after reaction with 10% of H20, whereas here we measured always under wet conditions. This is important because, as Figure 4 shows, there was a significant time lag (about 6 h) before the catalyst reached a stationary NO conversion when H20 was eliminated from the feed. It was noteworthy that both spent catalysts kept ca. 30% of their initial activity when measured under standard dry conditions, even though the distribution of Cu n+ species was totally different, as shown in the next section. We are analyzing the implications of this behavior.

,,,,,,,u.O-on

3.2 Structural characterization H2-TPR experiments showed the presence of two bands that have been assigned to Cu 2+ ---> Cu + (513 K) and Cu + ---> Cu ~ (623 K) reduction steps [15]. The Cu2+/Cu+ ratio for the deactivated catalyst identified as 2.3-Cu-ZSM-5/T was similar to that of the fresh catalyst, except for a new reduction band that appeared at 593 K on the H2-TPR pattern. The catalyst 2.3-Cu-ZSM-5/W had bands at ca. 553 K and 593 K (see Figure 5). The band at 553 K was attributed to CuO reduction, because CuO standards reduced at 573 K in one step. Yan et al. [14] found by TPR a band at 593 K, and suggested that it could correspond to Cu species interacting with extraframework AI produced during dealumination of the zeolite. However, preliminary characterization of catalysts operated at 22,000 h -1 and at temperatures as high as 923 K under dry and wet conditions (20% of H20 in the feed) via 29Si and 27A1 MAS-NMR, showed that there was no dealumination after reaction in presence of H20. X-ray diffraction patterns also proved that the crystallinity was barely modified during

229

a)

j--

reaction, and no other signals were detected besides those belonging to the zeolite. This was confirmed by FT-IR because the bands between 1250 and 400 cm-1 remained unchanged, indicating no dealumination of the zeolite [ 16].

We observed that the maximum temperature of the band assigned to Cu + reduction shifted to higher temperatures after reaction when H20 was present in the feed, even at concentrations as low as 1.6%. This shift could indicate that a fraction of Cu n+ was now located in a position that made more stable against reduction, as has been suggested by Matsumoto et al. [ 11], who claimed that a fraction of the Cu 2+ species could be located in cages different from those of the active species. Deactivation of Cu-ZSM-5 catalyst has been explained as resulting from ii II i , , I , , , I , , , I , ,,I,,,I, dealumination [14,12], CuO formation [ 17] and by migration of Cu n+ 320 400 480 560 640 720 800 880 species to inactive positions. Our Temperature K structural results indicated that dealumination was not significant and that both migration and CuO formation were possible depending Figure 5. TPR patterns of deactivated catalysts. on experimental conditions. a) 2.3-Cu-ZSM-5/T; b) 2.3-Cu-ZSM-5/W Interestingly it appears that some of these species could be spectators because their concentration did not track with activity. As an example, the Cu2+/Cu + ratio changed almost 40% after reaction under H20 concentrations below 10%, yet the activity remained essentially constant. An analysis of the TPR traces suggests that deactivation can be associated with the shift in the band corresponding to Cu + reduction. We are pursuing further this lead. ~

o ~..=(

b)

,

4. C O N C L U S I O N S The presence of H20 in the feed stream during NO reduction with C3H8 caused kinetic inhibition of the main reactions as well as the apparent activation of solid state phenomena possibly associated to Cu migration. As a result, both reversible and irreversible deactivation for NO and C3H8 conversions were observed. Catalyst deactivation followed a nearly linear trend at short times on stream (t < 12 h), but after long times we observed an unexpected sharp increase in deactivation. The drop in activity was a function of both temperature and H20 concentration, and a low steady state conversion was finally reached. It appears that the main reason for the deactivation of Cu-ZSM-5 is the mobility of Cu n+ in the presence of H20 rather than ZSM-5 dealumination.

230 ACKNOWLEDGMENTS

We acknowledge the financial support of Consejo Nacional de Ciencia y Tecnologia-Mexico (CONACYT, Project 400200-5-3420A), Instituto Mexicano del Petroleo (IMP, FIES No. 95F140-III), JICA (Infrastructure grant), and Universidad Autonoma Metropolitana-lztapalapa (UAM-I). REFERENCES

1. ,

3. ,

5. .

.

8. 9.

10. 11. 12. 13. 14. 15. 16. 17.

Y. Yuu, Y. Torikai, S. Sato, H. Hosose, H. Yahiro, N. Mizuno, and M. Iwamoto, Shokubai, 33 (1991) 6 I. S. Sato, Y. Yuu, N. Mizuno, and M. Iwamoto, Appl. Catal., 70 (1991)L 1. M.Iwamoto, H. Yahiro, S. Shundo, Y. Yu and N. Mizuno, Appl. Catal. B: Environ., 69 (1991) L15. E. Kikuchi, K. Yogo, S. Tanaka and M. Abe, Chem. Lett., (1992) 1063. H. Yahiro, H. Hirabayashi, H.K. Shin, N. Mizuno and M. Iwamoto, Trans. Mater. Res. Soc. Jpn., 18A (1994) 409. H.K. Shin, H. Hirabayashi, H. Yahiro, M. Watanabe and M. Iwamoto, Catal. Today, 26 (1995) 13. S. Scire and R. Burch., Appl. Catal. B: Environ., 3 (1994) 295. Y. Li, P. J. Battavio, and J.N. Armor, J. Catal., (1993) 561 S. A. G6mez, G.A. Fuentes y A. Martfnez, Actas del XV Simp. Iberoam. de Catdlisis, Vol. 2 (1996) 709. R.A. Grinsted, H.W. Jen, C.N. Montreuil, M.J. Rokosz and M. Shelef, Zeolites, 13 ( 1993) 602. S. Matsumoto, K. Yokata, H. Doi, M. Kimura, K. Sekizawa, and S. Kasahara, Catal. Today 22 (1994) 127. T. Tanabe, T. Iijima, J. Koiwai, K. Yokota and A. Isogai, Appl. Catal. B: Environ., 6 (1995) 145. M. Iwamoto, and N. Mizuno, J. Auto. Eng., 207 (1993) 23. J.Y. Yan, G.-D. Lei, W.M.H. Sachtler, and H.H., J. Kung, J. Catal., 161 (1996) 43. J. S~k~iny, J.L. D "Itri, and W.M.H. Sachtler, Catal. Lett., 16 (1992) 241. Van Boof and J.W. Roelofsen, Introduction to Zeolite Science and Practice, Elsevier, vol. 58 (1991) 241. K.C.C. Kharas, H.J. Robota, and D. Liu, Appl. Catal. B: Environ., 2 (1993) 225.

9 Elsevier Science B. V. All rights reserved Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

231

H - M o r d e n i t e D e a c t i v a t i o n during the S C R o f NOx. A d s o r p t i o n and D i f f u s i o n o f P r o b e M o l e c u l e s on Fresh and D e a c t i v a t e d Catalysts E. E. Mir6 a, L. Costa a, J.M. Dereppe b and J.O. Petunchi a a Instituto de Investigaciones en Cat~lisis y Petroquimica, INCAPE (FIQ, UNL-CONICET). Santiago del Estero 2829 - 3000 - Santa Fe - ARGENTINA. e-mail: [email protected] Laboratoire de Chimie-Physique, Universit6 Catholique de Louvain, P1 L. Pasteur 1, Bte. 3A, 1348 Louvain-la-Neuve, Belgium. e-mail: [email protected]

b

H-mordenites with various Si/A1 ratios (5.9 - 16.9) have proved to be active for the SCR of NO with CH4 in the 400-600~ temperature range. However, they suffered an irreversible deactivation after an incursion at 650~ for 1 h under reaction stream, due to a dealumination process. While acid dealumination only affects the free exchange of gaseous molecules between the main channels and the side-pockets (as seen by 129Xe NMR), the aluminum extracted from the lattice of the mordenite during the SCR of NOx at 650~ (without water vapor in the feed) also hinders the diffusive transport along the main channels. 1. INTRODUCTION Since Hamada et al. (1) reported that solids not containing transition metals catalyze the selective reduction of NO with hydrocarbons in the presence of O2, many research groups have studied this matter. Yogo et al. (2) reported that H-zeolites are able to reduce NO with CH4. They found that the selectivity for this reaction decreases in the order: H-ferrierite > H-mordenite > H-ZSM5 >> H-Y. We have previously (3,4) characterized in detail a series of dealuminated H-mordenites with Si/A1 r a t i o s - 5.9, 7.3, 11.0, and 16.9 through FTIR spectroscopy, 27A1-MAS NMR, XRD, and nitrogen adsorption. We have also found a correlation between the aluminum content of the samples and the catalytic activity for the selective reduction of NO with CH4 in the presence of oxygen excess. Despite the aluminum content, catalysts are partially deactivated on dry stream at temperatures of 6500C or higher. In order to gain insight into these deactivation phenomena, we have investigated the adsorption and diffusion of probe molecules on fresh samples with different Si/A1 ratios and deactivated on samples on stream during the SCR of NOx. While the 129Xe NMR of physisorbed Xenon was useful to study the exchange of gaseous molecules between the main

232 channels and the side-pockets, ethylene diffusion measurements gave us information about the possible obstruction inside the main channels of the mordenitic structure. TPD of ammonia and nitrogen adsorption were also performed to obtain supplementary information about these solids. 2. E X P E R I M E N T A L

2.1. Catalysts preparation and pretreatments. A commercial H-mordenite (Norton Zeolon 900 H) with a Si/A1 ratio of 5.9, determined by chemical analysis, was the starting material. The dealumination was performed with HNO3 at 90~ under different acid concentrations and leaching time (4). The following chemical Si/A1 ratios were obtained: 5.9, 7.3, 11.0 and 16.9. In all the samples (included the commercial H-Z) extra-lattice aluminum was detected by 27A1 MAS NMR (3). All samples were calcined in oxygen at 400~ From here onwards, catalysts will be designated as HMRF (where R is the Si/A1 ratio and F indicates fresh) and HMRD (where D indicates deactivated). The reaction (NO+CH4+O2) was carded out using 0.5 g of catalyst placed in a fixed-bed flow reactor. This was a 12 mm i.d. tubular quartz reactor with an internal thermowell. The typical reacting mixture consisted of 1000 ppm of NO, 1000 ppm of CH4, 10% of 02, balanced at 1 atm with He (GHSV = 6500 h-l). The catalysts were deactivated on stream during 1 h at 650~ More details of reaction experiments and apparatus are given elsewhere (3). 2.2. 129Xe NMR of physisorbed Xenon and 27A! MAS NMR. Prior to measurements, the samples were treated under vacuum (10-5 torr) at a heating rate of 2~ with three plateaus at 110oc (90 rain), 2100C (120 min) and 400oc (480 rain). Xe was adsorbed at 250C from a conventional apparatus which allowed us to measure the amount of Xe physisorbed at different equilibrium pressures. 129Xe NMR spectra were taken at a frequency of 82.91 mHz using a MSL-300 Brucker Spectrometer. Between 1000 and 5000 scans were performed in order to obtain a good signal-to-noise ratio. The chemical shift was measured between 10 and 600 mbar of Xe pressure, and it was referred to the NMR signal of Xe in the gas phase extrapolated at zero pressure. E7A1 MAS NMR spectra were taken at aluminum frequency of 78.15 MHz. Typically, 10,000 to 25,000 transients were acquired at a spinning frequency of 12 kHz, using a 4 mm diameter rotor with a double bearing system. 2.3. Ethylene diffusion measurements. The Zero-length Column Chromatography (ZCC) method was used (5). Between 10-50 mg of catalyst, 150 cm3/min He flow and a concentration of C 2H4 of 500 ppm were employed. Ethylene was adsorbed at different temperatures till the saturation of the sample was achieved. The hydrocarbon flow was then stopped and its concentration was measured during

233

desorption time. Details about the mathematical procedure to calculate diffusion coefficients can be seen in (5). 2.4. A m m o n i a TPD.

NH3 was adsorbed at 25 ~ till the saturation of the sample was achieved. It was kept at 100*C for 4 h in He stream. The TPD was conducted with a heating rate of 10*C/min between 100 and 700~ The solid was maintained at 7000C until recovering the base line. 3. RESULTS AND DISCUSSION Since in previous studies (3,4) we found that the NO conversion is related to the presence of AlIV sites of the zeolitic structure, it is possible to assume that the origin of the deactivation observed is, in part, due to the loss of these sites by dealumination. Together with the loss of active sites, the aluminum extracted from the network could block the channels of the mordenite, thus affecting the adsorption and diffusion properties. Zeolite dealumination consists in the removal of aluminum from a zeolite network. This process may be purposely done, under controlled conditions, or may occur during use through combined exposure to heat, water vapor and/or reactants and products, producing significant changes and modifications in catalytic properties. The extra-framework aluminum originated could remain, at least in part, inside the zeolite channels originating different kinds of aluminumresidues. Table 1 shows that both the dealumination process by reaction at 650~ (compare MH5.9F with MH5.9D) and the acidic dealumination decrease the volume available for the physisorption of N2 and Xe, the effect being even more marked in the case of Xe. The decrease of desorbed NH3 in TPD experiments is related to the decrease of the AlIV sites observed by 27A1 MAS NMR. 129Xe NMR of physisorbed Xenon, and ethylene diffusion measurements help us to better understand the deactivation phenomena. Table 1 Effect of deactivation and acid dealumination on the adsorption properties. Sample Vp(cm3/g)a 129Xe(1021/g)b NH3(1021/g)C AIIV(1021/g) MH5.9F 0.231 0.46 0.67 MH5.9D 0.205 0.37 0.19 MH7.3F 0.207 0.42 0.40 MH11F 0.203 0.28 0.38 MH 16.9F 0.215 0.30 0.34 a Micropore volume, b Measured at 100 mbar. c Desorbed between Calculated with the integrated AlIV intensity from 27A1 MAS NMR.

0.55 0.39 0.49 0.31 0.21 300 and 6000C.

d

234 3.1. 129Xe NMR of physisorbed Xenon. Fig. 1a shows the adsorption isotherms of 129Xe for the different dealuminated samples of H-mordenites. As observed in the graph, the adsorption capacity of the sample decreases as the dealumination degree increases. These results are in agreement with those obtained by Springuet-Huet and Fraissard (6) who compared two solids dealuminated by acid treatment and hydrotreatment, respectively, and the parent sample. Fig. lb. compares the adsorption isotherms of the H-mordenite catalyst with a Si/A1 ratio - 5.9, for fresh samples and deactivated samples on stream at 6500C. The lesser adsorption capacity of the catalyst which has been under reaction conditions indicates a structural change of the latter probably due to the combination of temperature with small amounts of water (approx. 500 ppm) formed as product. The behavior of the other catalysts under study is qualitatively similar. 10 (a)

(b)

HM 5.9 F

/•HMT.

HM5.9F

3F

8

f~6

HM11 F H~

e~

2

00

100

2O0

300

4O0

Pressure (mbar)

500

O0

~

I

100

,

I

200

,

I

300

Pressure (mbar)

,

4

Figure 1. Xenon adsorption isotherms at 25~ on (a) dealuminated samples, (b) fresh and used catalysts 129Xe NMR spectra of physisorbed Xenon for the different H-mordenites are shown in Fig. 2a. for a pressure of approximately 200 mbar and 250C. The spectrum corresponding to the Na-Mordenite sample is also shown as reference. The catalyst with Si/AI - 5.9 shows a single peak due to the well-known coalescence effect, indicating the Xenon free exchange between the main channels and the side-pockets. As the dealumination degree increases, the peak corresponding to the side-pockets starts appearing again, even though smaller if it is

235 compared to the Na-Mordenite. This is probably due to the fact that the residual extra-lattice aluminum makes the access of the Xenon atoms to the side-pockets difficult. These results, which indicate the presence of extra-lattice A1 (AlVI) even in the mordenites dealuminated by acid treatment, are in agreement with our previous results obtained with 27A1 MAS NMR (3). This suggests that the amorphous material formed by the extraction process of lattice alumina was not eliminated. This difficulty in eliminating the extra-lattice A1 altogether has also been observed in other studies (7,8). Klinowski et al. (7) argue that a possible reason could be the fact that octahedrally coordinated alumina created during the dealumination process are deposited with equal probabilities in the channels formed by eight- and twelve-member rings, respectively. During the acid washing, the extra-lattice species form (ai i -"1"35ppm (b) the AI(H20)63+, which is easily removable from the bigger channels but not from the smaller 250pp~ ] NaM 166ppm ones. Regarding the samples deactivated on stream, (3) we reported two changes if the 27A1 MAS NMR spectra are compared to those of the fresh samples (3). First, the AlVI/AlIV ratio is invariably increased and second, the signal corresponding to AlVI becomes slimmer in used solids, which is symptomatic of a more symmetrical environment of octahedral A1. These results show that catalysts are not only 1F 250 0 ppm dealuminated during reaction at high temperatures but also that there exists a change in the extra-lattice AI structure. However, the variation of tetrahedral alumina calculated between fresh and used samples is smaller than HM16.N,.,,,~,/~.: .~ . F expected if it is compared to the deactivation degree.

I

250

I

0

ppm

Fig. 2b. shows the spectra for the HM5.9F and HM5.9D samples. As stated above, for the fresh sample there is only one peak due to the coalescence phenomenon. However, for the used sample the peak corresponding to the side-pockets is also defined, and that of the main channels shifts to lower values of the chemical shift. This suggests that extra-lattice A1 is reaccommodated during the chemical reaction, thus partially obstructing the Xe free exchange between the main channels and the side-pockets. This restriction may be one of the causes for the loss of activity of the catalyst. Figure 2. 129XeNMR spectrum (a) dealuminated samples. (b) fresh and used samples.

236 Fig. 3a. shows the chemical shift (ppm) of the main peak for different adsorbed amounts of Xenon on fresh samples. The curves corresponding to catalysts with Si/AI 5.9 and 7.3 present a minimum at approximately 6 atoms of Xenon per gram, the form of the curve being qualitatively similar to that reported in (9) for H-Zeolon. However, for higher degrees of dealumination, side-pockets are obstructed, thus minimizing the coalescence effect. The interaction of Xenon atoms inside the main channels of H-mordenite with Si/A1 16.9 is similar to the one corresponding to Na-Mordenite, suggesting that their structure is not severely affected by acid leaching. Fig. 3b. shows chemical shifts with increasing amount of adsorbed Xenon, a shift to lower values being observed in all cases for the deactivated sample. 250

(a)

(b)

200

200

E ~"~

v

5.9 F

180

c-

Or) -

~

.~_150

E ID

160

tO

M-I 5.9 D MH 16.9 F

140

e<

100

NaM

o

'

5

'

~

'

~

Xa ,adsorbed (at. 102~

~

'

Io

Imo

'

~

'

~

'

$

'

)(e ,adsorbed (at. 102~

~

'

~o

Figure 3. Chemical shift of Xe adsorbed on (a) dealuminated samples, (b) fresh and deactivated catalysts. Comparing the results obtained with 27A1 MAS NMR to those obtained by 129Xe NMR, a coincidence between both techniques can be appreciated since both indicate that the extra-lattice A1 suffers a structural change during the chemical reaction, which leads to a partial blocking of the side-pockets thus contributing to the observed deactivation process.

3.2. Zero-Length Column (ZLC) Chromatography. The ZLC method recently developed by Eic and Ruthven (5) is an important new technique for measuring the intracrystalline diffusion in Zeolites since only a very small sample of

237 sorbent is required and relatively high diffusivities can be measured. An excellent review in this technique has been published by Rees (10). Figure 4 shows the experimental results obtained with this technique for the MH5.9F sample at three ~9 _ 25 ~ different temperatures. C indicates the measured concentration during 50 ~ 9 ethylene desorption at time t, and -3 Co indicates the inlet ethylene concentration. ~ Table 2 shows the results of ethylene diffusivity values together 7~oc 9 with NOx conversion for the different H-mordenites studied in this work. Note that while D/R2 & 9 decreases an order of magnitude in the deactivated sample (if compared -7 i 0 100 200 300 400 with the fresh one), it remains "time (sec.) approximately constant for the samples subjected to acidic Figure 4. Typical results of ZLC method. Ethylene dealumination. This result suggests desorption from MHF5.9 sample. that the aluminum extracted from the lattice of the mordenite during the SCR of NOx at 650~ hinders the diffusive transport of the mordenitic structure along the main channels. This effect could be in part responsible for the deactivation observed in the MH5.9D sample.

A~ ~

9

i,

I

i

I

,

I

,

Table 2 NO conversion, 129XeNMR and C 2 H 4 diffusion upon dealumination and deactivation. Sample NO x conversion (%) b Chemical shift (ppm) c D/R2(10-3/see) d MH5.9F 43 176 1.66 MH5.9D 30 162 0.17 MH7.3F 27 154 1.55 MH11F 25 123 1.90 MH16.9F 15 120 1.30 NaM a 122 a Sodium mordenite sample, b Measured at 500~ c Measured at 0.4 1021 atoms of Xe adsorbed/g, d Measured at 25~

238 4. CONCLUSIONS It has been shown that the H-mordenite catalytic activity during the SCR of NOx with methane in the presence of oxygen excess depends not only on the total amount of framework-aluminum but also on the distribution of amorphous residuals inside the mordenitic structure. 129Xe-NMR of physisorbed Xenon and diffusion measurements by ZLC Chromatography were useful tools in determining this distribution in acid dealuminated H-mordenites and in the same catalysts deactivated on-stream during the SCR of NOx. The framework-aluminum extracted during acid leaching is not fully washed under the conditions used in this work. Part of it remains inside the mordenitic side-pockets, partially obstructing the free exchange of gaseous molecules between the former and the main channels. The aluminum extracted from the lattice during the SCR of NOx at 650~ is located both in the side-pockets and in the main channels. These residuals obstruct the free exchange between the main channels and the side-pockets and hinder the diffusive transport along the main channels. Both effects, together with the loss of AlIV species, are responsible for the catalyst deactivation. ACKNOWLEDGMENTS

This work was performed under European Union Contract # C1 1* CT93-0090. Thanks are given to CONICET and UNL for their partial support, and to Elsa Grimaldi for the English edition. REFERENCES

.

6. 7.

10.

H. Hamada, Y. Kintaichi, M. Sasaki, T. Ito and M. Takata, AppL Catal., 64 (1990) L 1. K. Yogo, M. Umeno, H. Watanabe and E. Kikuchi, Catal.Lett., 19 (1993) 131. M. Lezcano, A. Ribotta, E. Mirr, E. Lombardo, J. Petunchi, C. Moreaux, and J.M. Dereppe, Stud. in Surf Sci. and Catal. 101B (1996) 971. A. Ribotta, M. Lezcano, E. Mir6, E. Lombardo and J. Petunchi, J. Dereppe and C. Moreaux. "Deactivation of Co and H-mordenite upon NO selective reduction with CH4". First World Conference in Environmental Catalysis, p255 Pisa, 1995. M. Eic and D.M. Ruthven, Zeolites 8 (1988) 40. M.A. Springuel-Huet and J.P. Fraissard, Zeolites 12 (1992) 841. J. Barras, J. Klinowsky and D.W. McComb, J.Chem.Soc.Faraday Trans., 90 (1994) 3719. L.D. Fernandez, P.E. Bartl, L. Fontes Monteiro, J. Gusmao da Solva, S. Cabral de Menezes and M.J.B. Cardozo, Zeolites 14 (1994) 533. T. Ito, L.C. De Menorval, E. Guerrier and J.P. Fraissard, J.P.Chem.Phys.Lett., 111 (1984) 271. L.V.C. Rees, Stud.in SurfSr and Catal., 84 (1994) 1133.

9

Elsevier Science B. V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

239

D e a c t i v a t i o n o f P t / A l u m i n a Catalysts for the H y d r o d e c h l o r i n a t i o n o f 1,1,1-Trichloroethane Kevin A. Frankel l, Ben W-L Jang 2, George W. Roberts 1, James J. Spivey 2. 1North Carolina State University, Raleigh, NC 27695 2Research Triangle Institute, Research Triangle Park, NC 27709 *Corresponding Author Vapor phase catalytic hydrodechlorination of 1,1,1-trichloroethane (TCA) has been studied using various supported platinum catalysts in a plug microflow reactor. The reactor was operated at temperatures ranging from 250 to 350~ a H2:TCA:He ratio of 10:1:89, a space velocity of 24 L/g cat-h, and atmospheric pressure. To study the deactivation process, tests were carried out by dividing the catalyst bed into three segments (inlet, middle, outlet) separated by glass wool plugs. Although the catalysts showed high initial activity, rapid deactivation was also observed. For example, when using a Pt/rl-alumina catalyst at 250~ essentially complete TCA conversion was observed initially; however, after 15 h TCA conversion had declined to < 25 percent. To understand the deactivation process, surface acidity and basicity, coke content, chlorine content, and platinum content were measured for both the fresh and the used catalysts. These measurements showed that up to 40 wt% coke formed on the supported platinum catalyst and that the acidity changed significantly during the reaction at 350~ 1. INTRODUCTION Catalytic hydrotreatment is widely used in the petroleum industry to remove sulfur, nitrogen, and oxygen from crude oil fractions. However, its use to treat chlorocarbons has not been widely reported despite the widespread use of these compounds in industrial and military operations, 1 and despite the negative environmental impact associated with most disposal options. Catalytic hydrotreatment has the potential to be a safe alternative for the treatment of chlorinated wastes and has advantages over oxidative destruction methods such as thermal incineration and catalytic oxidation. Some of these advantages include the ability to reuse the reaction products, 2'3 and minimal production of harmful byproducts, such as C12, COC12, or fragments of parent chlorocarbons. 2 1,1,1- Trichloroethane was chosen for this research because it is widely used in industry as a solvent 4'5 and is on the EPA Hazardous Air Pollutant list as a toxic air contaminant and ozone depleter. 6 The reaction for the complete hydrodechlorination of TCA to ethane is" C2H3C13 + 3 H2 -') C2H6 + 3 HC1.

(1)

240 Such hydrodechlorination reactions are generally exergonic (AGr <0) and exothermic (AHr <0). For example, AGr and A H r for Reaction (1) at 298 K are -243 and -220 kJ/mol, respectively. Much of the research dealing with hydrotreating of chlorocarbons has been on either chlorinated a r o m a t i c s 3'4'7'8 or chlorinated olefins 9-11 with very little reported on chloroalkanes. Although rapid deactivation has been observed, 4'8'~2 fundamental studies on the deactivation process have been limited. Deactivation has been ascribed to interactions between HC1 and the catalyst, 8'9 or to acid catalyzed oligomerization reactions. 4'12 Deactivation due to HC1 poisoning has been described as a rapid "self-poisoning" interaction between the HC1 product and the catalyst. 8'9 For example, the rate of hydrodechlorination of 1,2-dichloroethylene over a Rh/SiO2 catalyst at 240~ decreased by a factor of two within 50 min 9 and the chlorine content of the catalyst increased from 0.11 to 0.46 wt% after 600 min on stream, suggesting that HC1 poisoning caused the deactivation. Hydrogen treatment at 300~ restored 80 percent of the original activity and decreased the chlorine content from 0.46 to 0.17 wt%. However, in a similar study 8 of the hydrodechlorination of chlorobenzene with Rh/A1203 at 100~ hydrogen treatment was ineffective in restoring the catalyst activity. Chlorine could only be removed by an oxidizing treatment followed by a reduction step. 8 Coke formation has also been suggested as a major cause of deactivation. 4'12 Jang observed extensive coke deposition on Ni/ZSM-5 for the hydrodechlorination of TCA at 450~ 12 He suggested that coke was caused by oligimers formed on acid sites, and that HC1 poisoning was not a major factor. 12 This is consistent with the results of Creyghton et al., 4 who observed deactivation of Pt on various supports (alumina, BEA zeolites) for chlorobenzene hydrodechlorination. They studied the effect of the zeolite acidity on stability and activity. Replacing the Bronsted acid sites in Pt/BEA by sodium ions (i.e., decreasing catalyst acidity) increased the stability of the catalyst. Differential scanning calorimetry (DSC) of the various catalysts after 18 h of operation showed that coke deposition was the smallest on the least acidic catalyst. Creyghton et al. argued that HC1 poisoning was not the cause of deactivation because deactivation profiles of the various catalysts were different at nominally identical conditions, despite the fact that each fresh catalyst had the same amount of exposed platinum surface area (platinum loading multiplied by dispersion). 4 This study attempts to correlate the rate of deactivation with the physical and chemical characteristics of the catalyst. Catalyst properties studied included total acidity and basicity, BET surface area, coke content, and metal loading. The objective was to develop a better understanding of the primary cause of deactivation of the hydrodechlorination of chloroalkanes. Techniques are discussed later in detail. 2. EXPERIMENTAL Figure 1 shows a schematic diagram of the experimental equipment. The main components are the feed preparation chamber, reactor, and analytical system. Hydrogen and helium were stored in cylinders. The flow rates were controlled by Tylan mass flow controllers (MFCs). The H2 gas stream was passed through a temperature-controlled impinger filled with TCA. The impinger inlet was fritted to create an even distribution of the gas stream in the liquid. A tempering beaker

241 connected to a circulating bath controlled the temperature of the impinger. The impinger pressure was controlled by a metering valve downstream of the impinger. The saturated hydrogen stream from the impinger was mixed with He and/or H2 in a tee. The mixed feed stream then flowed through a fixed catalyst bed housed in a tubular furnace.

Hydrogen

Q

Hydrogen

G Helium Q Q

Hydrogen

Flow

Mass Controller

Impinger for HCl Analysis

Va~e

Vent

Heating/~ Jacket

~---h~'~

I I~1 ?pinger

~ Imlml I

Vent

Vent

Tubular Fumace

i

~:~m/MpSte Station /

I-~ Thermocouple | Pressure Gauge Figure 1. Process flow diagram for hydrodechlorination of 1,1,1 trichlorethane. The reactor was a vertically mounted tube (Monel, outside diameter 0.375 in., inside diameter 0.305 in.) in a single-zone, Lindberg clamshell furnace. The heating zone was 12 in. in length. The heat to the furnace was controlled by an Omega temperature readout/controller. The catalyst was placed in the middle of the furnace where the temperature was most uniform. The depth of the catalyst bed was about 1.6 cm. The bed was divided into three equal segments separated by glass wool plugs (inlet, middle, outlet). This made it possible to study the effects of the reaction on the catalyst with respect to axial position. Omega thermocouples (K-type, Chromel-Alumel 304, Stainless Steel Sheath) were used to measure inlet and outlet temperatures. The thermocouple tips were placed within the glass wool that supported the catalyst bed. Typically, the temperature difference across the catalyst bed was <5 ~ Cole-Palmer pressure gauges (0 to 15 psig) were used to measure the pressure drop across the bed. Typically the pressure drop was < 0.25 psi and the reactor pressure was 1 psig. All catalysts were supplied by Engelhard Corporation. Table 1 shows the physical characteristics of the fresh catalysts. The reactor was loaded with 0.5 g of catalyst. The mesh size was between 60 and 120 (particle size = 0.125 - 0.250 mm). All catalysts were pretreated in 10 percent H2 in He (200 cm3/min) before reaction as follows: reactor temperature was increased from 50 to 300~ at 5~ held for 8 h, then heated at 5~ to 450~ and held for 1 h. A Hewlett Packard gas chromatograph (GC) equipped with a flame-ionization detector (FID) and a thermal conductivity detector (TCD) was used to analyze the feed and product streams. A 10-port sampling valve permitted on-line sampling. The sample flow rate to the GC was 20

242 cm3/min. The sample was split before the GC. One sample was sent to a 50-m wide bore capillary column (GSQ) which separated the hydrocarbons and chlorocarbons. The second sample was sent to a dual column system. The first column was a 9-m packed column (HaysepR), which separated H2 and He from the rest of the stream. The second column was a 15-m packed column (Carboxen 1000), which separated H2 from He. The first column was backflushed to remove the hydrocarbons once the H2 and He passed through. The FID was used to identify and quantify the pure and chlorinated hydrocarbons contained in the sample. Using nitrogen as the carrier gas the TCD was used to quantify the He and H2 concentrations. Table 1 Fresh catalyst characterization

Acidity Otgmol/g of

Basicity Otmol/g of

cat)

cat)

Catalyst

Surface area (cm 2)

Metal loading (wt%)

Weak"

Strong b

Weak"

Strong b

Pt/rl-alumina

200

2.8

29

33

116

205

fir-alumina

102

--

38

17

52

95

Pt/ct-alumina

6

2.8

--0

--0

<15

--0

tx-alumina

5

--

---0

--0

< 15

---0

"Weak sites- Ta,< 250 *C. bStrong sites- Ta, > 250 *C. To determine the acidic and basic sites of the catalyst, probe gases were adsorbed on the catalysts and then desorbed via a temperature program using an Altamira AM-1 system. Ammonia (NH3) was used to determine acidic sites and carbon dioxide (CO2) was used for basic sites. The designation "weak" has been assigned to sites from which the probe molecule desorbed at <250~ and "strong" has been used for sites where desorption requires >250~ However, in studies with zeolite ZSM-5, it has been shown that physically adsorbed NH3 can require desorption temperatures as high as 600 ~ Therefore, it is not certain that the sites for which the NH3 desorbed below 250 ~ are acidic in the conventional sense. Table 2 shows the methodology used to determine the acidic and basic sites of the catalysts. Coke content was determined using a thermal gravimetric analyzer (TGA). A gas stream (180 ml/min) containing 2 percent 02 in N2 was passed over about 15 mg of catalyst. The initial temperature of the system was 35~ The temperature was ramped at 10~ until it reached 120~

It was then held at 120~ for 1 h to ensure that all of the water vapor was removed from

the catalyst. The temperature was then ramped to 550~ at 10~ and held at 550~ for 2 h. Coke content was determined by subtracting the final weight of the catalyst from the catalyst weight after the water was removed.

243 Table 2 Methodology used to determine acidic and basic sites

Treatment Step 1 Gas: H2 Starting Temperature: 50 ~ Temperature 2:600 ~ C Temperature 3: 35 ~C

Flow rate: 25 cm3/min Ramping rate: 15 ~ Hold time at temperature 2:60 min Hold time at temperature 3:5 min

Treatment Step 2 Gas: CO2 (basic sites) NH 3 (acidic sites) Temperature: 35 ~

Flow rate: 25 cm3/min Hold time at temperature: 60 min

Temperature-programmed Desorption Gas: He Starting temperature: 35 ~ Ramping rate: 5 ~ Hold time at final temperature;. 30 min

Flow rate: 25 cm3/min Hold time at temperature: 15 min Final temperature: 550 ~ C Cool to room temperature

Platinum and chlorine content were determined using inductive coupled plasma-mass spectrometry (ICP-MS). To determine the platinum content, the catalyst was first digested by three concentrated aqueous acids (H2SO4, HNO3, HC104). The actual platinum content was determined by adjusting the measured Pt content by the amount of coke deposited on the catalyst. Chlorine content was determined by extracting the chlorine from the used catalyst with deionized water and then analyzing the water for chlorine. Thus, the chlorine content reported here accounts only for water soluble chlorine, not chlorine that may be part of the coke. 3. RESULTS AND DISCUSSION Microreactor studies using the four catalysts shown in Table 1 were done at the following conditions: temperature = 2 5 0 - 350~ pressure - 1 atm, space velocity = 24 L/g of cat-h, H2:TCA:He - 10" 1:89. The Tl~i-alumina support was used as a comparison to the Pt/~-alumina. Although the two supports were not identical it has been assumed that the two supports act similarly. Figure 2 shows that the alumina supports have very different activities, and that the addition of Pt to t~-alumina has a more dramatic effect than addition to Tl~i-alumina. For example, Tl~-alumina was active for TCA conversion without Pt, while the t~-alumina was nearly inactive. In addition, essentially complete conversion was observed using the PtAI-alumina catalyst for 12 h, while the maximum conversion observed using the Pt&t-alumina was the initial conversion of 90 percent. Possible reasons for the lower activity of the s-alumina include much lower catalyst surface area, acidity, and basicity (see Table 1) for the t~-alumina than for the ~-alumina.

244 100

~,9 "\ ....ZX '\

- - 0 - - or-alumina

80

rn

~l&.alumina

Pt/q-alumina

\ \ \

"-=.ta

ro

o

Pt/ct-alumina

9.-Z~.-

\

60

.o

-I-

40

"B" B"B" -K, ==. . t S . . , 20

-11.11

_.___,_ J

'

I-

0

-

4

--

\

_._

I

I

I

8

12

16

"

I

20

24

Time on Stream (h) Temperature = 250 "C, Pressure = 1 aim, SV = 24 Ug cat-h, H2:TCA:He = 10:1:89

Figure 2. Effect of catalyst support on the hydrodechlorination of TCA. Figure 2 also shows that the addition of Pt increased the stability of the catalyst. For example, essentially complete conversion was achieved with the Pt/rl-alumina catalyst for 2 h longer than with the rl~5-alumina support, although there might be a difference between the activity of the 1"1;5alumina support and the rl-alumina support. In addition, when Pt was added to the a-alumina support, initial conversion increased from < 5 to 90 percent. Figure 3 shows the differences in hydrocarbon selectivity among the four catalysts. Products in the effluent were separated into two groups--pure hydrocarbons (HC) and chlorocarbons (CHC). Hydrocarbon selectivity is the percent of all products in the exit gas that are hydrocarbons on a carbon basis. Chlorocarbon selectivity is not shown and may be determined by subtracting the hydrocarbon selectivity from 100. Table 3 shows the products observed in the effluent for each individual experiment.

lO0

-~m_=~= ~-__= _ - _v

~. \

80

\ \

_.o

6O

_

40

-

0 -ID 9

i

"-0-II-

a-alumina Pt/ot-alumina --~-9 ~15-alumina

~

"~ \

Pt/q-alumina

._. m

o

z

\

\

\ ~.j~r~p--~P--~

20

9 -

~

0

4

8

12

I

16

,=h ,w

I

20

,IL ,w

24

Time on Stream (h) Temperature = 250 "C, Pressure =1 alto, SV = 24 Ug cat-h. H2:TCA:He = 10:1:89

Figure 3. Effect of catalyst on hydrocarbon selectivity for TCA hydrodechlorination.

245 Pt seems necessary for complete removal of the chlorine atoms from TCA. Neither a-alumina nor rl~5-alumina (without Pt) produced any pure hydrocarbons. However, when Pt was added to either support, ethane was observed in the effluent. In addition, ethylene was observed in the effluent when the Pthx-alumina catalyst was used, suggesting a weaker hydrogenation function for the Pt/(x-alumina catalyst compared to the Pt/rl~5-alumina catalyst. This weaker hydrogenation function may account for the more rapid deactivation since the unsaturated intermediates may polymerize to form coke. Table 3 Product Observed in Experiments at 250~

Experiment

Hydrocarbons observed

Chlorocarbons observed

co-alumina

None

DCE"

Pt/~-alumina

Ethane, ethylene

DCE', DCA '

riB-alumina

None

DCE

Pt/vl-alumina

Ethane, methane, butane

DCE, DCA

Note: Major compounds observed in effluent are in boldface. aConcentration of chlorocarbons observed in effluent was < 75 ppm DCE = 1,1-dichloroethylene, DCA = 1,1-dichloroethane Figure 3 also shows that the change in selectivity with respect to time was different for the two supported platinum catalysts. For Pt/TI-alumina, essentially complete conversion of TCA to ethane (Reaction 1) was observed for the first 10 h on stream. This was followed by a change in selectivity to primarily DCE, a partially dechlorinated product (Reaction 2) and an accompanying decrease in TCA conversion (Figure 2). C2 H3C13 ---) C2H2C12 + HC1.

(2)

When using Pt/ct-alumina, the product selectivity did not change with time (ethane was the primary product), although TCA conversion declined rapidly. The difference between the selectivities for the two Pt catalysts as they deactivate suggests that the Tl-alumina support may initiate the reaction. Specifically, HC1 elimination to form DCE (Reaction 2) appears to be the initial reaction that occurs primarily on the acidic sites of the Tl([i)-alumina support. Hydrogenation of this unsaturated chlorocarbon to DCA then occurs on the Pt. This is consistent with the observation that large concentrations of DCE were only observed on the two catalysts that had acidic sites (Tl~i-alumina, Pt/TI-alumina; Table 1). However, Pt has activity for the HC1 elimination step since the Pt/o~-alumina was active for the hydrodechlorination reaction while the or-alumina was not. In addition, because DCE was observed in the effluent when using the Pt/TIalumina before the conversion decreased, deactivation most likely occurred first by the loss of

246 hydrogenation activity (on the Pt) followed by a loss in HC1 elimination activity (on the acidic sites). Additional data were taken using the Pt/rl-alumina and riB-alumina catalyst at 350~ (space velocity = 24 L/g of cat-h, pressure = 1 atm, and a H2:TCA:He ratio = 10:1:89). Figure 4 compares the two catalysts. Clearly, by comparison with Figure 2, deactivation was slower at 350 ~ than at 250 ~ As at 250~ TCA conversion was complete for a longer period of time for the Pt/rl-alumina. The product distribution was similar to that observed at 250~ for both Pt/rl-alumina and rl~5-alumina. One major difference was that when the Pt/rl-alumina catalyst was used at 350~ the initial HC selectivity was primarily to methane instead of ethane. After 3 h, however, selectivity to methane was < 1 percent and the primary HC product was ethane. 100 ~'A.

(9 O

-->' .>

80

AA

Closed symbols - T C A Conversion

_.r r

Open symbols

- HC Selectivity

O o

9

Q

60

Q ~ o

~9 t...

Q

40

0

> 0

0

9

20 -

A

0

3% P t h l - a l u m i n a ~lS-alumina

20

0 0

40

60

Q

80

100

T i m e on s t r e a m (h)

Pressure = 1 a t m , S V = 24 LIg cat-h, H 2 : T C A : H e = 1 0 : 1 : 8 9 , W = 0.5 g, 3-bed system

Figure 4. Effect of Pt on TCA Conversion and Selectivity at 350 ~ Upon completion of this run at 350~ the segmented catalyst beds were unloaded and characterized by NH3 TPD, TGA, and ICP-MS. Table 4 shows that a large amount of coke (up to 40 wt%) was deposited on the catalyst and that coke deposition decreased slightly from the middle segment to the outlet. The large amount of coke deposited on the Catalyst suggests the coke was primarily being deposited on the support. Platinum and chlorine content also changed. There appeared to be a 14 percent loss in Pt after 96 h, although there was no significant trend with respect to distance. Loss in Pt was unexpected and there is no clear reason for it at this time. Repeat ICP-MS analysis of the fresh catalyst showed <1 percent variability, suggesting that this difference is significant and not due to reproducibility problems. The difference may be due to inaccuracy in the coke measurement, since the measured Pt concentration was corrected to a "fresh catalyst". Chlorine content appeared to increase with reactor distance. However, the measured chlorine content only included what was water extractable, and not what might be chemisorbed to the Pt, reacted with the alumina, or part of the coke.

247 Table 4. Change in catalyst characteristics of 3% Pt/rI-Alumina after 96 h of operation at 350~ Apparent Acidity (lamol/mg of cat) Coke content (wt%)

Pt content (wt%)

CI content (wt%)

Rate constant (L/g cat-min-atm) at 350 ~

Weak

Strong

Fresh

0

2.8

0.00

> 12

29

33

Inlet

40

2.4

0.26

0.22

39

15

Middle

40

2.3

0.33

0.32

40

17

Outlet

35

2.3

0.66

0.49

43

18

Bed segment

A psuedo-first order rate constant was determined for each deactivated bed segment. Before testing the catalyst, each segment was heated in pure He (180 cm3/min)at the reaction temperature (350~ for 1 h to drive off any water vapor. The conversion of TCA was then measured for 1.5 h on each individual segment. The rate constants were calculated assuming that the reaction was first order in TCA concentration and zero order in H2 concentration. Table 4 shows that there was a slight increase in activity with reactor distance. However, the rate constant for the fresh catalyst was at least two orders of magnitude greater than those of the coked catalyst and perhaps much greater than 12 L/g cat-min atm since essentially complete conversion was observed using 0.15 g of fresh catalyst. The reaction caused a large change in apparent surface acidity. Figure 5 shows in more detail the change in apparent acidity of the three catalyst segments and a comparison to the fresh catalyst. The increase in "weak" sites may simply be the result of increased area for physisorption due to coke deposition. Strong acid sites were lost as the catalyst deactivated, suggesting that coke may form on and cover the strong acid sites. Table 5 shows how the rlS-alumina changed during a 3-h experiment. The objective of this short run was to obtain a "snapshot" of what was occurring on the catalyst before complete deactivation occurred. Coke was rapidly deposited on the catalyst. Coke content increased slightly with reactor length, rather than decreasing, as for the Pt/TI-A1203. This coking profile suggests that the coke precursor may be DCE rather than TCA. It appears that water extractable chlorine is deposited rapidly on the catalyst. The amount of extractable chlorine deposited on the riB-alumina support by itself after 3 h is about the same as the amount deposited on the Pt/rlalumina catalyst after 96 h.

248 600 In

7 S

Temperature - 500

-

Middle

5 01

7

"

- 400

Fresh C a t a l y s t

(.)

- aoo ~

a

~2 I--

..... ".'.-,t

E

200 I--

I 0

100

-1

W e a k Acid Sites

-2

~-,,

!

0

20

40

Strong Acid Sites !

!

i

!

60

80

100

120

T i m e (min)

Microreactor Conditions: Temperature = 350 *C. P r e s s u r e = 1 atm, S V = 24 L/g cat.-h, H2:He:I 11 T C A = 10:1:89, T i m e on Stream = 96 h

Figure 5. Strong acid sites of 3% Pt/rl-alumina catalyst are lost due to reaction Table 5 Change in Catalyst Characteristics of rl~5-Alumina after 3 h of Operation at 350~ Apparent Acidity (pmol/mg of cat) Bed segment

Coke content (wt%)

CI content (wt%)

Weak

Strong

38

19

Fresh

0

0.02

Inlet

3.0

0.32

Middle

3.1

0.29

52

15

Outlet

3.5

0.28

55

34

Table 5 also shows that the change in apparent surface acidity for the deactivated catalysts is qualitatively different than that observed on the Pt/TI-alumina. Note, that the fresh rl~5-alumina catalyst has less apparent acidity than the fresh Pt/rl-alumina. Possible reasons include the differences in the supports and the addition of Pt. Figure 6 shows the NH3 TPD for the catalyst segments. As with the Pt/Tl-alumina, there was a large increase in weak apparent acidity. However, there appears to be an increase of strong acid sites at the outlet of the bed that was not observed in any of the segments of Pt/rl-alumina catalyst. One possibility is that strong acid sites are created by HC1 before they are covered by coke.

249 600 - 500

Middle

Temperature

O

5 m " ._~

a I--

400

"-"

300

E ID Q. E II) I---

4 3 .....-.--,.

2

200

100

0

!

i

i

I

!

i

20

40

60

80

100

120

Time

0

(rain)

Microreactor conditions" Temperature = 350 ~ Pressure = 1 atm SV = 24 L/g cat-h, H2:111 TCA:He = 10:1:89, Time on Stream = 3 h

Figure 6. Acidity of the qS-alumina support increased due to reaction. 4. D E A C T I V A T I O N M O D E L

The preceding results suggest deactivation model for the hydrodechlorination of TCA on the qS-alumina support by itself and on the platinum/q-alumina catalyst. The model is based on one originally proposed by Froment and Bischoff.13 The proposed model for the support by itself is a series reaction in which TCA is converted to DCE on the catalyst surface, which may then either desorb or form coke. This model is based on two observations: (1) DCE is the only product observed in the effluent and (2) coke content increased with reactor distance (Table 5), suggesting a series coke-forming mechanism. The proposed deactivation model for the Pt/rlSalumina catalyst is based on a series/parallel reaction. Here TCA is first converted either to DCE or an intermediate that is not observed in the effluent. This reaction intermediate then either completely and rapidly hydrodechlorinates to form ethane or reacts to form coke on the catalyst, via successive, rapid surface reactions which probably involve unsaturated intermediates. This series/parallel model for the supported Pt catalyst is based on the following observations: (1) the reaction intermediate (DCE) is observed before TCA conversion decreases (Figures 2 to 4) and (2) saturated compounds p e r se are unlikely coke precursors. From this model, coking profiles and a mathematical deactivation model are now being developed. 5. C O N C L U S I O N S

Deactivation of the 3% Pt/q-A1203 catalyst may be due to coke formed from the reaction intermediate DCE, probably on the strong acid sites of the alumina support. DCE forms coke rapidly on the qS-alumina without Pt. The function of the Pt may be (1) to hydrogenate DCE to DCA, permitting additional hydrodechlorination, and (2) to maintain a high ratio of saturates to unsaturates, thereby reducing coke formation. Even in the presence of Pt, the rate of coke

250 formation is significant. Deactivation might be reduced by increasing the rate of hydrogenation of DCE and minimizing the strong acid sites on the alumina. ACKNOWLEDGMENTS

The financial support provided by the Army Research Office and Project Officer Dr. R. Seiders under contract DAAHO4-95-K is gratefully acknowledged. The authors would also like to acknowledge Dr. J. Butt of Northwestern University for his assistance in the deactivation model and formulating the segmented bed experiments. REFERENCES

.

3. 4.

.

8. 9. 10. 11. 12. 13.

National Research Council. "Alternative Technologies for the Destruction of Chemical Agents and Munitions." 1992. Kales, T.N.; James, R.B. Environmental Progress, 1988, 7, 185-191. Gioia, F.; Famiglietti, V.; Murena, F. J. Haz. Mat., 1993, 33, 63-73. Creyghton, E.J.; Burger, M.; Jansen, J.; van Bekkum, H. Applied Catalysis A: General. 1995, 128, 275-278. Wolf, K.; Yazdani, A.; Yates, P. J. Air Waste Manage. Assoc. 1991, 41(8), 1055-1059. "1,1,1 Trichloroethane," Agency for Toxic Substances and Disease Registry Public Health Statement" [Web site: http:/atsdl.atsdr.cdc.gov:8080/ToxProfiles/phs9027.hmtl]. Hagh, B.; Allen, D. Ind. Eng. Chem. Res. 1994, 33, 2942-2945. Coq, B.; Ferrat, G.; Figueras, F. J. Cat., 1986, 101,434-445. Bozzelli, J.; Chen, Y. Chem. Eng. Comm., 1992, 1, 115-124. Weiss, A.H.; Krieger, K.A..J. Cat., 1966, 6, 167-185. Weiss, A.H. Vallinkski, S. J. Cat, 1982, 74, 136-143. Jang, B. "Conversion of Polychlorinated Hydrocarbons and Acetylene over a Nickel Modified Shape Selective Zeolite Catalyst," Ph.D. Thesis, University Of Texas, Arlington, 1992. Froment, G.; Bischoff, K. Chem. Engrg. Sci., 1961, 16, 189-193

Farneth, W. E. and Gorte, R. J. Chem. Rev., 1995, 95, 615-635.

9

Elsevier Science B. V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

251

U n d e r s t a n d i n g Claus Catalyst Deactivation Mechanisms: Optimization o f A l u m i n a U s i n g Physico-chemical Parameters Christophe N6dez a and Jean-Louis Ray b aRh6ne-Poulenc, Usine de Salindres, 30340 Salindres, France bprocatalyse, 212-216 avenue Paul-Doumer, 92500 Rueil-Malmaison, France This work is devoted to the comprehension of mechanisms involved in the deactivation of Claus catalysts. The principal cause of deactivation of alumina is sulphation of the surface. The optimization of the alumina (Na20 level content, macroporosity around 0.1 ~tm., etc.) reduces considerably the consequences of aging, given a high level of catalytic performance.

1. INTRODUCTION Many natural gas fields and gases coming from petroleum contain hydrogen sulphide. Moreover, in order to obtain cleaner gasolines for the environment, the use of hydrodesulphurization (Eq. 1) in modern refineries is increasing. These two factors could combine to triple the amount of sulfur compounds and hence H2S to be treated

[1]. RSH + H2 -~ RH + H2S

(1)

The catalytic transformation of hydrogen sulphide into elemental sulphur (Claus catalysis) is a major step in industrial pollution control. After a partial thermal oxidation of H2S into SO2 (Eq. 2), the Claus reaction (Eq. 3) and the hydrolysis reactions of the important secondary compounds such as COS (Eq. 4) and CS2 (Eq. 5) have to be catalytically achieved [2]. H2S + 3/2 02 <---) SO2 + H20 2H2S + SO2 +-~ 3/xSx + 2H20 COS + H20 ---) H2S + CO2 CS2 + 2 H 2 0 ~ 2H2S + CO2

(2) (3) (4) (5)

It is important to have efficient and stable catalysts, otherwise a decrease of activity produces an increase in the SO2 discharge. Furthermore stringent regulations require the reduction of gaseous emissions of Claus units [3]. In this work, we have studied the different aspects of the deactivation of Claus catalysts and reported our main observations.

252 2. E X P E R I M E N T A L

We evaluated the performance of different catalysts under two distinct experimental conditions: (1) under reactor conditions (R1), where the main objective of the catalyst is to hydrolyze COS and particularly CS2; and (2) under reactor conditions (R2), where the Claus reaction (Eq. 3) is more discriminating for a catalyst evaluation. Experimental conditions for the catalytic tests were highly representative of the conventional ones. The R1 conditions were: 320~ 6% H2S, 4% SO2, 1% CS2, 30% H20, N2, 10-2000 ppm 02, 2 or 3 s (TPN) of contact time. The R2 conditions were: 250~ 3% H2S, 1.5% SO2, 30% H20, N2, 500-2000 ppm 02, 2 or 3 s (TPN) of contact time. For the R2 conditions, the sulphur yield was calculated by assuming that the Claus conversion reaction approaches the thermodynamic equilibrium. After the elimination of water, product gases were analyzed by gas phase chromatography. Each experiment was designed to determine the equilibrium conversion.

3. RESULTS AND DISCUSSION 3.1.

Alumina

Most conventional Claus catalysts are alumina based. Many physico-chemical factors govern the performance of an alumina-based catalyst (Figure 1, Table 1). At equilibrium, the Aluminas X and Y reach respective sulphur yields of 73 and 71%, much lower than that of Procatalyse CR-3S (84.5%). The difference in performance between these catalysts does not depend on the contact time; i.e., for a 2 s contact time, at equilibrium, the respective sulphur yields of Alumina X, Alumina Y and CR-3S are 59, 58 and 72.5%. The deactivation suffered by a particular alumina is the consequence of several phenomena. Table 1 Physico-chemical characteristics of fresh aluminas Characteristics CR-3S Particle size (mm) 3.1-6.3 Na20 level (ppm) 2100 Surface(m2/g) 380 Total pore volume (ml/100g) 55 V750A (ml/100g) 19.6 V0. l~tm(ml/100g) 19.3 V lktm(ml/100g) 16.1 V1/V0.1 0.83

Alumina X 3.4-6.4 3235 380 56 20 19.5 10 0.51

Alumina Y 4-8 2025 359 52 12.2 12.0 8.0 0.67

253 We have recently shown [4-5] that an adjustment of macroporosity (between 0.1 and 1 ktm) of the used alumina beads permits considerable decrease in the diffusional constraints and thus allows attainment of distinctly improved catalytic performance. We have also shown the importance of the chemistry of the alumina surface: while a minimal level of Na20 content (> 1500 ppm) is needed for adequate activity, a level over 2500 ppm favors the sulphation of surface sites in the neighborhood of K atoms and effects a particularly fast chemical deactivation.

100 94

tD

\

88

................ ~

.......... .'.......... :..:.'....-...~..,...~ . . . . . . . . . . . . .

.,-.4

......................... ~ o

.=

82

. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .

~-..~..~. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .

~.. ..........

............................Alumina X ~ ~ s ~ 76 ........................................................................ ~-~----~...... . . . . . . CR-3S ---o-~-................. -o - Alumina Y 70 I t t t -

0

20

40 60 Reaction time (h)

80

Figure 1. Evolution of the sulphur yield (in R2 conditions) as a function of reaction time, at 500 ppm 02, for a 3 s contact time. During the first hours of the reaction, an adsorption equilibrium is established at the surface of the catalyst, causing a decrease in conversion. However, the most important chemical deactivation is related to the catalyst sulphation [6-11]. Several factors are responsible for sulphation, the most important of which seems to be linked to the presence of a relatively high level of alkalis (Na20 > 2500 ppm) [4,5,12] (compare CR3S and Alumina X). Sulfation is also clearly affected by the presence of hundred-ppm levels of oxygen in the stream (Figure 2). The mechanism of sulphate formation is still under debate. Recently, a kinetic model has been proposed for alumina sulphation [ 13]. Two different types of alumina surface sites could be involved in the elementary stages. The first one would be an oxygen atom on the surface that would favor SO2 adsorption and then sulphation. The second one would allow the dissociation of oxygen. After the formation of superficial sulphates on the surface of the alumina, a nucleation of core sulphates, via a slow process, causes formation of irreversible chemical species. Another important type of deactivation is due mainly to physical aspects of the catalysts at the high water content of the gases to be treated (ca. 30% vol.) linked with the temperature (200-350~ causing an irreversible hydrothermal ageing. Some other factors could play a role in the deactivation of a Claus alumina. For example, coke formation, due to the presence of hydrocarbon compounds in the feed,

254

could affect some sites at the surface. Deposition of sulphur in the micropores of the alumina could also lead to a decrease in the catalytic conversion. In this last case, a moderate elevation of temperature is sufficient to vaporize sulphur. In practice, it is the sum of all these different processes that leads to deactivation of Claus catalysts.

0

.................................................................................................................

-.9

Alumina X

84 .........................e.~ .....................................................-e .........CR-3S

~D

-. - - o - Alumina Y 78 ................................................................. 77:........~.................................

. ,...~

72 ........................6~....~ < - ~ i ............................................................ 66 60

I 0

I

500

1000

1~"~"~'~~'~'~ 1500

2000

Oxygen (ppm) Figure 2. Role of the oxygen level on sulphur yield (in R2 conditions), at near equilibrium, for a 3 s contact time. 3.2. P r o m o t e d alumina To delay deactivation, it is possible to promote the alumina by adding various promoters, which favor the Claus reaction (Eq. 3) and the hydrolysis reactions (Eqs. 4 and 5). However, it is essential to fully understand the effects of promoters on deactivation of the catalyst. For example, if promotion of an alumina by an alkaline earth element offers an advantage in Claus conversion during the first hours of reaction, a reverse effect occurs as the reaction proceeds further (Figure 3, Table 2). A good understanding of the surface chemistry explains these phenomena. Oxygen present at the levels shown (in general < 2000 ppm) causes sulphation of catalysts. An element such as calcium initially retains oxygen via intermediary formation of a sulphite, then a sulphate (Eq. 6). However, the calcium sulphate formed irreversibly is intrinsically less active than the alumina and is not regenerable. Thus, a Ca/alumina or a MgCa/alumina catalyst operates as a non-regenerable oxygen trap.

CaO + SO2 ---) CaSO3

CaSO3 + 1/2 02

---) CaSO4

(6)

The catalytic behavior of an iron-promoted alumina is clearly different. In the presence of the Claus gases, a catalytic cycle involves the conversion of oxygen impurities via transformation of iron sulfide into iron sulfate, which is easily regenerated in the stream to sulphur by reaction with H2S [14] (Figure 4). In summary, Fe/alumina

255

is a catalyst, whereas C a / a l u m i n a a n d C a M g / a l u m i n a are really only temporary traps, with regards to oxygen.

Table 2 Physico-chemical characteristics of promoted aluminas Characteristics

Ca/alumina

CaMg/alumina

Fe/alumina

Ti/alumina

Particle size ( m m ) N a 2 0 level (ppm)

3.4-6.4 3100

3.4-6.4 4920

3.1-6.3 1950

3.1-6.3 2000

CaO level (%) M g O level (%) Fe203 level (%)

13 0.17 -

5.8 1.5 -

6.6

-

TiO2 level (%) Surface(m2/g) Total pore v o l u m e a V0. ] wn (ml/100g)

256 40 14

313 49 13

288 45 14.9

5.25 272 56 15.7

V l~m (ml/100g)

9

3

10.4

11.3

V1/V0.1 aln ml/100g.

0.64

0.23

0.70

0.72

1 O0

~ j k ~ - = - ~ ~............................................................................................. ~ A

= "~

72

x

xx

%

...................................... ~ i ..........3 ...............................~................... " ~ , , , , "~ o x

~"

..........................................................

x x

o x~'~-~ ~ ~r

,, o

....................................... o

= 44 30

x

............................CaMg/alumlna ....... ,-~..~..~x :~, ...............-_-~-.........C ~ u ~ n a .................................~ i ; . . s ..... o~

- ~ - t CR-3S

I

t

J

20

40

60

80

Reaction time (h) Figure 3. Sulphur yield as a function of reaction time, under R2 conditions with 2000 p p m 0 2 , and 3 s contact time. P r o m o t i o n o f an alumina with TiO2 produces better performance (Figure 5) because titanium sulphates are not chemically stable in presence of H2S and titanium dioxide is very active in Claus conversion and hydrolysis [ 15].

256

100

~-~:

80

I,i

--

"~"

............................................................................................... * .... 9

"-

.................... ~

=~ 60 ~ =

L..I~

o

o

./...............

20

-

Ca/alumina

---,,-- Alumina

/ / ..................... ~ . .

!:i:

o40 ol) ~

.............. 1. .......................................................................

- -e-

Fe/alumina

............... :z~

0 " ~-

I

0

15

--

"

-I . . . . . . . .

I ~--zx-

30 45 Reaction time (h)

--

" - ~ o~ " - 0 }

1 ......

60

75

Figure 4. Oxygen consumption observed with different catalysts, under R2 conditions with 2000 ppm 02 and a 3 s contact time. 85

~- 70

F21Alumina X

.2

D CR-3S

>

55 Ca/alumina

o

~9

t'-,I r/3

I CaMg/alumina

40

Ti/alumina 25 2s

3s

Figure 5. CS2 conversion at near equilibrium, obtained with aluminas and promoted aluminas under R1 conditions, with 200 ppm 02, and contact times of 2 s and 3 s. In this last case, the choice of the alumina as carrier for TiO2 is crucial. We found that a great catalytic synergy links the active phase with the alumina. Precise control of the key physico-chemical parameters of the alumina (specific surface area, particle size, and more importantly macroporosity profile and surface Na20 content level) is crucial in the deactivation of the promoted catalyst. Macroporosity profile and the surface Na20 content level are the most important parameters for the alumina carrier in obtaining a high catalytic performance over a significant reaction time. For example, to avoid problematic diffusional constraints, a macroporosity around 700 A or even 0.1 lxm is not at all sufficient. It is necessary to choose a well-defined profile of the porosity

257

between 0.1 and 1 ~tm, for R1 conditions [4-5] as well as R2 conditions (Figure 6, Table 1).

85 ............~..........~..u..~n..a....X................................................................. /e CR-3S - e - S serie J9 / 82 ..............................................................................7 .../........................... x Alumina Y oi /

'x:l

/

79 ......................................................... / .........................

/f

~

0

...............................

J f

76 ............o " ~ ...........................................................................................

r.,r

73

..........................................................................................................-~..... •

70 10

i

i

i

i

J

12

14

16

18

20

V0. llam (ml/100g) Figure 6. Sulphur yield at near equilibrium, obtained with fresh aluminas, as a function of their volume at 0.1 lxm and their ratio V1/V0.1, under in R2 conditions, with 500 ppm 02, at a 3 s contact time (series S: all physico-chemical parameters similar to those of CR-3S, includes V1/V0.1 ratio, except V0.1~tm and V ll,tm values).

Sulphur yields

,005[

/

%f. TiO2/alumina

l

70h [r ~ ~y , d , l CS2 rolysis ! 1G 70 85 40/2 G

!

25

1G

40

Figure 7. Consequences of the aluminas' generation G on the catalytic performance, for a 3 s contact time, in R1 (under 200 ppm 02) and R2 (under 500 ppm 02) conditions. 4. C O N C L U S I O N S Carrier optimization allows for clearly improved performance (Figure 7) at equilibrium, and thus, higher sulphur yields. A notable reduction of HES discharge into the atmosphere is obtained, promoting better protection for the environment.

258 REFERENCES ~

2. 3. 4. 5. 6. 7. 8. 9. 10. 11. 12. 13. 14. 15.

Y. Barthel, C. Raimbault, P6trole et Techniques, 382 (1993) 40. Sulphur, 187 (1986) 1. D. Knott, Oil & Gas J., Aug. 1, 24 (1994); Sulphur, 228 (1993) 37. Ch. N6dez, and J.L. Ray, Catal. Today, 27 (1996) 49. Ch. N6dez, and J.L. Ray, Catal. Today, 29 (1996) 139. M. Graulier, and D. Papee, Energy Process., 1 (1974). Sulphur, 175 (1984) 34. Z.M. George, Can. J. Chem. Eng., 56 (1978) 711. C. Quet, J. Tellier, and R. Voirin, Stud. Surf. Sci. Catal., 6 (1980) 323. A. Pi6plu, O. Saur, J.C. Lavalley, O. Legendre, and Ch. N6dez, to be published. J.L. Ray, and O. Legendre, Preprints Sulphur'96, Vancouver, 20-23 October 1996. Z.M. George, M. Bensitel, M. Lion, O. Saur, and J.C. Lavalley, Appl. Catal., 43 (1988) 167. A. Pi6plu, O. Saur, J.C. Lavalley, M. Pijolat, and O. Legendre, J. Catal., 159 (1995) 394. J.L. Ray, O. Legendre, and Ch. N6dez, Preprints Sulphur'95, Abu Dhabi, 15-18 October 1995. T. Dupin, and R. Voirin, Hydrocarbon Process., 61 (l l) (1982) 189.

9

Elsevier Science B. V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

259

D e a c t i v a t i o n b y P o i s o n i n g and the I m p r o v e m e n t o f Three W a y C a t a l y s t s for Natural Gas-Fueled Engines Takeshi Tabata*, Hideo Kawashima and Kenji Baba Applied Research Center, R & D Department, Osaka Gas Co., Ltd. Torishima 6-19-9, Konohana-ku, Osaka 554, JAPAN Tel (+81-6)4664558, Fax (+81-6)4641805, E-mail: [email protected] A Pt-Rh three way catalyst used in natural gas-fueled engine systems for 21,000 h was investigated to analyze its deactivation mechanism. Specific deactivation characteristics, such as the decrease in the selectivity of NO reduction, which cannot be reproduced by heat treatment, were observed on the used catalyst. P and Ca were detected on the surface of the wash coat and Zn, Fe and trace amounts of Pb were observed in the wash coat layer. Of the model-poisoned catalysts a decrease in the selectivity of NO removal was observed only for the Pb-poisoned catalyst. All the characteristics of the deactivation on the used catalyst were reproduced on the model Pb-poisoned catalyst at high SV. Thus, it is concluded that the selectivity of NO reduction is decreased by chemical poisoning due to Pb on the used catalyst; moreover the absolute rates of NO reduction as well as for other reactions are reduced by a decrease in the effective surface area due to physical clogging in addition to a thermal effect. Furthermore, it was found that an improved catalyst, which can maintain a high Rh surface area even after thermal deactivation by separating high Pt/Rh ratio particles from low Pt/Rh ratio particles at a high Rh concentration, possesses high poison-resistance in the activity tests using model Pb-poisoned catalysts as well as field tested catalyst. Finally, its durability has been proven for more than 30,000 h. 1. INTRODUCTION Cogeneration systems using natural gas-fueled engines (gas engine), which can produce electricity and heat simultaneously, have become increasingly prevalent in Japan because of their high thermal efficiency. Simultaneously, strict regulations on the NOx emissions from such systems have been issued. A three way catalyst system, developed for the emission control of automobiles [1 ], is used to reduce NOx emission from gas engines [2-4]. However, the required performance of the catalysts for gas engine cogeneration systems is different from that for automobiles. The most difficult requirement is extremely high performance (NOx conversion > 98%) maintainable over several tens of thousands of hours of operation, which is more than one order of magnitude longer than for the normal automobile. Since the catalyst life directly influences the maintenance cost of cogeneration systems, it should be prolonged as much as possible. It has been reported that the major reason for the deactivation of three way catalysts for gas engines is the rapid decrease in methane oxidation activity, due to sintering of Pt by the thermal effect, causing the shift of the window (defined as the range of air/fuel ratio [A/F] at which NO and CO are simultaneously converted) of the catalyst [3, 4]. Therefore, a Pt-Rh catalyst with high Pt loading and with a high Pt/Rh ratio is used as a three way catalyst for gas engines in order to maintain the methane oxidation activity for a long time by maintaining

260 a high surface area of Pt. The durability requirement for the catalyst for gas engines has been established at 16,000 h. However, when we wish to prolong the catalyst life, it is necessary to pay attention not only to thermal deactivation but also to the deactivation by trace amounts of poisonous elements which are not problematic for automobiles due to the shorter life time required. For this purpose, a Pt-Rh three way catalyst used in gas engine systems for 21,000 h was investigated employing chemical analyses, EPMA, surface area measurement, catalytic activity tests as well as tests using model poisoned catalysts. From these investigations, the mechanism of deactivation was clarified. Furthermore, based on the new knowledge of the deactivation mechanism, a catalyst has been developed to overcome deactivation by poisoning. The results of the analyses of the deactivation mechanism and the development of a long life catalyst are reported in this paper.

2. EXPERIMENTAL A commercial Pt-Rh/CeO2/A1203 catalyst wash-coated on a cordierite honeycomb ("HL" catalyst supplied by Cataler Industrial; Pt - 3.5 g/l, Rh - 0.35 g/l; 300 cells/inch2) was used to analyze the deactivation mechanism. Improved catalysts were also prepared in washcoated honeycomb form by N. E. Chemcat. Model poisoned catalysts were mainly prepared by dipping a honeycomb catalyst sample heated in air at 1173 K for 5 h into 500 ml of an aqueous solution of metal nitrate for 1 min with stirring, followed by blowing off drops of water, drying overnight, and final calcination in air at 973 K for 5 h. Although this method may not accurately simulate the catalyst poisoning, it was employed because the source and chemical form of the elements are not clear and because such elements can be easily loaded throughout the wash coat layer uniformly with sufficient reproducibility. The chemical composition of the catalysts was quantified by x-ray fluorescence analysis. B. E. T. surface area was measured by the B. E. T. one point method using nitrogen gas, and the amount of CO adsorption was measured by the pulse method. These analyses were carried out for the ground catalysts containing cordierite. The deviation of the results due to the uniformity of the thickness of the wash coat layer was around 5%, when 5g of fresh catalyst sample was used for each. The distribution of elements was measured by line analysis using EPMA (Shimazu EPMA 8705). Catalytic activity was measured using a flow type reactor equipped with a stainless steel reactor tube (14 mm in inner diameter). Four ml of crushed catalyst sample (sifted into 1 - 2 mm) was placed in the reactor tube. A model exhaust gas, generated from standard mixed gases, was fed to the reactor tube after the addition of steam. The test gas compositions, analysis methods, and the calculation method of A/F (3,)is the same as [4,5]. For evaluating the selectivity of NO reduction, model gases containing 990 ppm NO, 2360 ppm O2 and 1000 - 6000 ppm CO were used. This confirmed that the results of the catalytic activity tests were unchanged for honeycomb catalysts without crushing at the same gaseous hourly space velocity (GHSV).

3. RESULTS AND DISCUSSION

3.1. Catalytic activity of the used catalyst The catalytic activities of the HL catalyst after heat treatment in an air stream at 1323 K

261 for 5 h is shown in Fig. la. The heat treatment at 1323 K severely reduces catalytic activity and corresponds to operation in an actual gas engine for more than 21 000 h. In Fig. 1a, both the conversions of 02 and NO start to decrease roughly at the same A/F. The conversion of methane forms a volcano shape and the peak is located in the window. The CO conversion is still high even on the rich side and exceeds H2 conversion due to its high activity of CO shift reaction. The conversions of CO and NO are almost 100% at the window. On the other hand, the results obtained for the catalyst used in actual gas engines for 21 000 h showed quite different characteristics (Fig. 1b). The conversion of CO decreases abruptly on the rich side, even though H2 conversion reaches almost 100%, and the conversion of CH4 on the rich side is 0%. These results mean that the activities of CO shift and steam reforming have been lost. On the other hand, the CH4 conversion peak is located at ~, = 1.0, at which the conversion of 02 starts to decrease, quite apart from the "assumed" window (the crossing point of the conversions of NO and CO). Furthermore, the conversion of NO is below 90% even on the rich side, and the window disappeared, while 02 conversion is almost 100%. (a)

~ - - ~ :-.%

100 ~ A

=

100 -" . . . . . . . .

(b)

80

80

60

60

---,~

-

40

40 o

o 20

20

O 0.985 [

,

I

,

I

~

I

0.990 0.995 1.000 Air/fuel ratio (~)

,

I

1.005

O

0.985

,,.

,=

~

0.990 0.995 1.000 Air/fuel ratio (;L)

-

1.005

Fig. 1. Activity of deactivated three-way catalysts for gas engines. (a) A catalyst deactivatedby heat treatment in air at 1323 K for 5 h. (b) A catalyst used in actual gas engine systems for 21000 h. Original catalyst: HL (Cataler Industrial, Pt = 3.5 g/l, Rh = 0.35 g/l). Reaction condition: T = 873 K, GHSV = 30000 h-~. Conversions: ( 9) NO, ( o ) CO, ( 9 CH4, ( zx) 05, ( 9) I-I2. To clarify the difference, the selectivity of NO reduction by CO versus CO oxidation by O2, considered to be the best index of the difference, was measured for fresh, heat-treated and used catalysts. The results are shown in Fig. 2. Here, the closer to the lower right comer the curve is, the lower the selectivity of NO reduction. The A/F's correspond to ~, = 1.015 at the origin and to ~,= 1.000 at the upper right comer. The CO conversion was 100% at every point between these two points. At the points near the stoichiometric A/F, at which the selectivity of NO reduction is most important, the curve for the catalyst heated at 1173 K or 1273 K is almost identical to that for the fresh catalyst, but the selectivity of NO reduction on the used catalyst is distinguishably lower. Thus, clearly a decrease in NO reduction selectivity, not explainable by thermal deactivation, occurs on the used catalyst.

262 100

A

80

3.2. Analysis of the used catalyst

The B. E. T. surface area of the used catalyst (15.1 m2/g) is similar to that of the catalyst .= heat-treated at 1173 K (15.5 mE/g: = 64% of the 0 value of a fresh catalyst). The heat treatment at " 40 1173 K corresponds to operation in a normal gas 0 0 engine for about ten thousand hours, and so, the 0 used catalyst was not deactivated as much as a z 2O catalyst thermally deteriorated by misfire. However, chemical analysis showed that P (0.38 wt.-%), Ca (0.42 wt.-%), Zn (0.95 wt.-%), Fe 20 40 60 ' 8'0 ' 1( )0 (0.41 wt.-%) and Pb (0.033 wt.-%) increased in 02 Conversion (%) the used catalyst compared to those in the fresh Fig. 2. The selectivity of NO reduction by CO on three- catalyst, while Pt, Rh, Ce, A1, Si and Mg were way catalysts. The original catalysts were the same as in unchanged and sulfur was below the detection Fig. 1. Catalyst: (a) flesh ( 9 (b) heat-treated in air at limit (0.01 wt.-%) in the used catalyst. 1123 K for 5 h (O), (c) treated in the same way at 1273 The distribution of the elements in the K ( 9 (d) the used catalyst same as in Fig. l b (o). direction across the monolith was analyzed by Reaction conditions: T = 773 K, GHSV = 30000 h-~, NO - 990 ppm, O2 = 2360 ppm, CO = 1000-6000 EPMA line analysis to identify the poisonous ppm. The conversion of CO was 100% except for the elements, because the change in the selectivity of NO reduction as shown in Fig. 2 is not thought highest point (CO = 6000 ppm; net rich condition). to be caused by physical poisoning such as the clogging of micropores but by chemical poisoning through the adsorption of poisonous compounds on an active metal distributed in the catalyst layer. The results are shown in Fig. 3. Rh is located in the catalyst (wash coat) layer, while P and Ca accumulate only on the surface; thus they do not chemically poison the noble metals. Both Fe and Zn exist partly in the catalyst layer, and Pb is distributed uniformly throughout the catalyst layer. e-

._o 60

3.3. Activity of model poisoned catalysts For Fe, Zn and Pb that may cause chemical poisoning, model poisoned catalysts were prepared by dipping catalysts into aqueous solutions of metal nitrates at various concentrations, and the catalyst carrying the nearest amount of each element was selected for the selectivity measurement. The uniform distribution of the loaded metals in the catalyst layer was confirmed by EPMA line analysis. While B. E. T. surface areas of the model poisoned catalysts differ little, the amount of CO adsorption decreases with the increase in the concentration of the poisonous metal, and it is noteworthy that the amount greatly decreases by loading a trace amount of Zn or Pb. The selectivity of NO reduction was estimated in the same way as in Fig. 2, and the results are shown in Fig. 4. NO reduction selectivity does not seem to decrease by loading Zn or Fe, while the selectivity considerably decreases on the Pb-loaded catalyst. In fact, for the Fe- or Zn-loaded catalyst, the characteristics of the dependence of the catalytic activity on A/F did not change from that of a heat-treated catalyst. However, as shown in Fig. 5a, the CH4 conversion peak was found away from the window on the model Pb-poisoned catalyst at GHSV - 30,000 h-1 and 873 K. Furthermore, at GHSV = 90,000 h-l, all the characteristics in Fig. 1b were reproduced qualitatively (Fig. 5b).

263

I

100 8O e-

.9 60 .=

t-

_=

I

o 0 7

i

I

20 40 60 80 02 Conversion (%)

Rh IG.as cha,

20

nnell

'. VVash coat' ,

0

,

,

,

Cordierite

I

,

,

,

0.1 Deplh (rnm)

Fig. 3. EPMA line analysis of the detected elements in the used catalyst. Catalyst: the same as in Fig. lb.

i

100

Fig. 4. The selectivity of NO reduction by CO on model-poisoned three-way catalysts. The original catalysts were the same as in Fig. 1. Catalyst: (a) only heat-treated in air at 1173 K for 5 h (ifl), (b) dipped in 5 wt.-% Zn aq. Zn(NO3)2 solution alter the heat treatment ( 9 (c) 0.3 wt.-% Fe aq. Fe(NO3)3 instead of Zn ( I ) , (d) 0.1 wt.-% Pb aq. Pb(NO3)2 instead of Zn ( 9 ), (e) the same "used catalyst" as in Fig. 1b (0). Reaction conditions are the same as in Fig. 2.

3.4. The deactivation mechanism Although the used catalyst was deactivated due to thermal degradation, deactivation due to poisoning is also evident. The most significant deactivation feature of the used catalyst is the decrease in the NO reduction activity, considered to be due to Rh which also contributes to steam reforming and CO shift reactions. Therefore, it is more likely that the activity of Rh is mainly decreased by Pb-poisoning, because only Pb reduces the selectivity of NO reduction of the catalyst. It is reported that Pb poisoning causes the decrease in the conversions of NO and hydrocarbon (HC) but does not affect CO conversion on Pt-Rh-Pd or Rh three way catalysts [6]. The fact that NO conversion decreases while CO conversion does not change implies a decrease in selectivity for NO reduction by CO, consistent with the present results. The present results also agree well with the reported result that the poisoning effect of Pb on HC oxidation was observed even from trace amounts [7]. On the other hand, the activity of the model Pb-poisoned catalyst (Fig. 5a) was generally a little higher than that of the used catalyst (Fig. 1b), although the Pb-poisoning effects are observed more strongly on the model Pb-poisoned catalyst than on the used catalyst (Fig. 4). It may be more likely that the catalytic activity for each reaction, such as CO or HC oxidation, NO reduction, CO shift, etc., is uniformly reduced by an additional factor on the used catalyst than on the model Pb-poisoned catalyst, because the activity of the model Pbpoisoned catalyst at high SV more closely resembled that of the used catalyst. As such an additional factor, the effective surface area of the used catalyst in actual operating conditions is considered to be lower than the measured B. E. T., because B. E. T. surface area was measured for ground catalyst although a considerable amount of accumulated

264 compounds consisting of Ca, P and Zn covers the catalyst layer as shown in Fig. 4. Thus, the deactivation observed on the used catalyst may be explained as follows. First, Pt is sintered and Rh is segregated on the surface of noble metal particles by the thermal effect [8, 9], and so, on only thermally deactivated catalysts, the reactions due to Rh such as NO reduction are not slowed, although methane oxidation due to Pt is considerably slowed [3, 4]. However, when Pb is adsorbed on Rh, the catalytic activity of Rh is suppressed and the reaction rate of NO reduction is decreased to the same order as the rate of CH4 oxidation. Further, the surface of the wash coat layer is covered with compounds consisting of Ca, P, Zn and Fe, and the effective surface area of the catalyst which the exhaust gas can reach decreases, causing a considerable decrease in NO conversion and the disappearance of the window, accompanied with decreases in CH4 conversion and CO conversion on the rich side. (a)

100

g 9

(b)

100

80 60-

80 60

.

4o

o

4o

o

2O 0

0.985

0.990 0.995 1.000 Air/fuel ratio (;L)

1.005

2O 0

0.985

0.990 0.995 1.000 Air/fuel ratio (L)

1.005

Fig. 5. Activity of the model-poisoned catalyst. Catalyst: the same as in Fig. 4 (d). Reaction conditions and symbols are the same as in Fig. 1 exceptfor GHSV; (a) 30000 h"l, (b) 90000 h-~.

3.5. Improvement of the catalyst for longer catalyst life It is difficult to determine the source of Pb because it is present only at extremely low concentrations and it is not easy to protect the catalyst from Pb-poisoning, because of the volatile nature of Pb compounds (e. g., Pb(CO)x). However, Fig. 5 teaches that the influence of Pb poisoning can be reduced if the effective surface area is kept high. To improve the catalyst itself, this means keeping high Rh dispersion even after thermal deactivation. Therefore, a high heat-resistant catalyst must also have high poison-resistance. However, a high surface area of Rh cannot be achieved by enriching Rh since it causes a decrease in CH4 oxidation activity [3, 4]. Apart from the study on poisoning, the authors have been developing a highly heatresistant three way catalyst for gas engines [5]. In this catalyst, high dispersions of both Pt and Rh can be maintained even after thermal treatment by separating high Pt/Rh ratio particles for CH4 oxidation from low Pt/Rh ratio particles for NOx reduction, because the Pt is hardly covered with Rh on high Ptha,h ratio particles, and the Rh is hardly involved in the Pt sintering on the low Pt/Rh ratio particles. The heat resistance of the improved catalyst is demonstrated [5].

265

For the present purpose, a high Rh surface area even after thermal deactivation is desirable. Therefore, one of the improved catalysts, in which a relatively high amount of Rh is loaded, "GEC-O 1" (Pt = 3 g/l, Rh = 0.6 g/l), was selected as a long life catalyst which would have high poison resistance. Figure 6 shows the activity of model Pb-poisoned catalysts. The characteristics of the activity of Pb-poisoned catalyst are clearly observed on the conventional catalyst, but hardly observed on GEC-O 1.

(a)

100 A

(b)

100

80

~, 80

6o

6o

40

40

20

20

0 0.985

0.990 0.995 1.000 Air/fuel ratio (~,)

0 0.985

1.005

0.990 0.995 1.000 Air/fuel ratio (k)

1.005

Fig. 6. Comparison of the activity of model-poisoned catalysts between the conventional and improved catalyst. Original catalysts: (a) the same as in Fig. 1 (conventional), (19) GEC-O 1 (improved; Pt = 3.0 g/l, Rh = 0.6 g/1. Both catalysts were heat-treated in air at 1323 K for 5 h followed by dipping in 0.1 wt.-% Pb aq. Pb(NO3)2

solution. Reaction conditions and symbols are the same as in Fig. 1. 100

0

.w > 0 X

O Z

95

|

0

20O00

40000

T i m e (h)

Fig. 7. Performance of the improved catalyst in commercial gas engine cogeneration systems. Catalyst: GEC-O1 (the same as in Fig. 6b). ( o, o, e ) Site A (1000 kw), ( 9 ) site B (460 kw), ( 9 site C (100 kw). Operation conditions: NO = 2500-3500 ppm, CH4 = 400-2000 ppm, CO = 2500-5000 ppm, T = 870-950 K, GHSV = ca. 30000 h-1. *Disproportionate flow due to the clogging of catalyst monolith by dust from engine oil. This catalyst has been introduced to customer sites, and its performance is shown in Fig 7. The catalyst showed sufficient NOx conversion over 30,000 hours of operation. One of the catalysts at site A was sampled after 27,000 h for analyses. NOx conversion on the

266 actual gas engine system was more than 98%. The B. E. T. surface area of the catalyst was 15.3 m2/g (59% of the value of fresh catalyst), and so the catalyst is considered to have been used under thermally normal conditions. The catalyst showed sufficient NOx reduction activity in the laboratory test, although it contained a large amount of contaminants: P (1.0wt.-%), Ca (0.79 wt.-%), Zn (2.7 wt.-%) and Pb (0.105 wt.-%). Since the concentration of Pb was 3 times higher than the highest value among all the samples used earlier in gas engine systems, it is thought that the operating conditions were considerably bad in terms of poisoning. Thus, the poison resistance of the improved catalyst GEC-O1 was well demonstrated and its durability has been proven for more than 30,000 h through this series of field tests.

4. CONCLUSION A Pt-Rh three way catalyst used in natural gas-fueled engine systems for 21,000 h showed specific deactivation characteristics, including a decrease in the selectivity of NO reduction, which can neither be reproduced by heat treatment nor explained by physical poisoning such as the blockage of micropores. Through chemical analyses, EPMA, and activity tests of the used catalyst and model-poisoned catalysts, it was found that the activities of Rh on the used catalyst were decreased by chemical poisoning due to Pb, causing a decrease in the NO reduction selectivity, and that the absolute rates of NO reduction and other reactions are considerably reduced by a decrease in the effective surface area of the catalyst due to accumulated compounds on the wash coat surface, in addition to thermal effects. An improved catalyst, which can maintain a high Rh surface area even after thermal deactivation by separating high Pt/Rh ratio particles from low Pt/Rh ratio particles was found to have high poison-resistance in activity tests using model Pb-poisoned catalysts as well as field tests, and its durability has been proven over 30,000 h. REFERENCES

1. K. C. Taylor, in J. R. Anderson and M. Boudart (Eds.), Catalysis, Springer-Verlag, Berlin, 1984, p.119. 2. J. Klimstra, SAE Technical Paper Series No. 872165 (1987). 3. T. Tabata, O. Okada, K. Baba and H. Fujita, Proc. 1st Jpn. EC Joint Workshop on the Frontiers of Catal. Sci. & Tech. for Alternative Energy & Global Environ. Protection, 1991, p. 306. 4. T. Tabata, K. Baba and O. Okada, Nihonkagakukaishi, 1995 (1995) 225. 5. T. Tabata, K. Baba, H. Kawashima, K. Kitade, T. Tanaka, M. Kokitsu, H. Ohtsuka and O. Okada, Stud. Surf. Sci. Catal., 92 (1995) 453. 6. D. R. Monroe, SAE Technical Paper Series No. 800859 (1980). 7. W. B. Williamson, J. Perry, H. S. Gandhi and J. L. Bomback, Appl. Catal., 15 (1985) 277. 8. B. J. Cooper, B. Harrison, E. Shutt and I. Lichtenstein, SAE Technical Paper Series No. 770367 (1977). 9. T. Yamada, J. Siera, H. Hirano, Y. Ikedo, B. E. Nieuwenhuys and K. Tanaka, Surf. Sci., 226 (1990) 1.

9

Elsevier Science B.V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

267

D e a c t i v a t i o n of M a n g a n e s e - C e r i u m O x i d e Catalysts during W e t O x i d a t i o n o f Phenol P. Oelker, A. Bernier, P. Ruiz, B. Delmon and P. Isnard a Laboratoire de Catalyse et Chimie des Mat6riaux Divis6s, Universit6 Catholique de Louvain Place Croix du Sud 2/17, B-1348 Louvain-la-Neuve, Belgium D6partement S6curit6/Environement, Rh6ne-Poulenc Industrialisation Avenue Jean Jaures, 69153 D6cines-Charpieux Cedex, France a

The catalytic deep oxidation of phenol in aqueous solution using manganese cerium oxide catalysts was studied by varying catalyst composition, pH of the reactant solution, partial pressure of 02, reaction temperature and possible intermediates of the reaction. A deactivation of the catalysts was observed. A leaching of cerium--and preferentially of magnesium--into the solution, and formation of carbonaceous species on the surface of the catalysts were observed. It is suggested that deactivation is due principally to the formation of a polymer on the surface of the catalyst during reaction. Modifications of the experimental conditions of the reaction (pH, temperature, pressure) or the composition of the oxides do not prevent this deactivation. Polymer formation seems to be well explained by a mechanism proposed in the literature attributed to a heterogeneous-homogeneous free-radical mechanism. Some lines for future research are proposed. 1.

INTRODUCTION

Catalytic wet air oxidation (WAO) is an attractive process for the purification of industrial waste waters containing a small amount of organic pollutants. It involves the total oxidation of dissolved organics to carbon dioxide and water (or at least a transformation into biodegradable molecules). Wet air oxidation thus appears as an alternative process to the thermal liquid-phase oxidation of organics, which has already found many industrial applications. It provides the advantages of milder working conditions, i.e. lower temperatures, lower pressures and lower corrosion rates. However in some cases, complete conversion of organic compounds to carbon dioxide and water cannot be easily achieved. This is the case, for example, in the oxidation of phenol, in which several undesired intermediate species that are much harder to destroy than the original compounds can be formed. Degradation of phenols is very important because they are highly toxic to aquatic life, have a high resistance to biodegradation and impart a strong unpleasant odor and taste to water even in very small concentrations. Contamination of water by phenol or substituted phenols is very often due to industrial activity, because these are commercially very important chemicals. Phenol, cresylic acids, and cresols are used for making phenol-formaldehyde resins and tricesyl phosphates. Phenol, alkylphenol, and polyphenols are important raw materials for a wide variety of organic compounds, dyes, pharmaceuticals, plastizers, antioxidants, etc. Phenols are also present in effluents from coke ovens, blast furnaces, and shale oil processing [1 ].

268 Several catalysts have been used to oxidize phenols and their substituted derivatives in water. A very common one is CuO-ZnO-A1203, which promotes the oxidation of phenol, chlorophenol and nitrophenol at low temperature (130~ and low oxygen partial pressure (3 bar) [2]. However, significant leaching of copper was observed during reaction, making the industrial use of copperbased catalysts unattractive. In addition to this, polymer formation on the surface of the catalysts was also observed. The same type of catalyst was used in the oxidation of p-coumaric acid, a compound representative of the polyphenolic fraction typically found in olive processing and wine distillery wastewaters. CuO-ZnO-A1203 was found to be effective although the leaching of metals was also found to occur. In this case it was found that the extent of leaching was strongly dependent on the operating conditions [3]. In recent years, manganese-cerium oxide with a well defined Mn/Ce ratio has been proposed as a promising oxide catalytic system for the oxidation of organics in aqueous solution [4]. High conversions of organics under moderate experimental conditions have been reported. However in spite of the good catalytic performance of this catalyst, the literature addressing details of the stability of Mn/Ce oxide catalysts during deep oxidation is scarce. It is clear that a detailed understanding of catalyst deactivation is absolutely necessary for successful industrial development of this process. This study focused on the deactivation of the Mn/Ce catalysts during reaction. The catalytic oxidation of phenol in aqueous solution to carbon dioxide, water and other side-products was selected as the test reaction. Catalysts were prepared from amorphous precursors using the citrate method and controlling the calcination temperature. Activity performance as a function of the time on stream was studied by simultaneously analyzing the conversion of phenol, the total organic carbon content of the catalyst, the cations eluted and the elemental composition of both cerium and manganese. Experimental conditions were widely varied. Fresh and used catalysts were also analyzed by BET surface area, X-Ray Diffraction and X-Ray Photoelectron Spectroscopy.

2. EXPERIMENTAL 2.1. Catalyst preparation A series of manganese-cerium oxide catalysts with different Mn/Ce atomic ratios was prepared by thermal decomposition of amorphous citrate precursors [5]. The samples were prepared by varying the ratio between the manganese and cerium salts. Citrate method: The preparation procedure consisted in the dissolution of the metal salts, namely Mn(NO3)2o4H20 and Ce(NO3)3o6H20, in distilled water, the complexation of the metallic cations with citric acid and the rapid concentration of the liquid by evaporation under vacuum. The viscous liquid was dried at 80~ and the amorphous precursor obtained was decomposed in air prior to calcination. Owing to the complexing property of the citrate anion, this procedure leads to the formation of finely dispersed two-phase systems or favours the formation of mixed oxide phases upon calcination, when these exist. Samples were calcined for 5 hours at 200, 300, 400 or 500~ The composition of the samples was expressed in weight per cent of Mn203 in the mixture, considering that the calcined samples are composed of Mn203 and CeO2 : wt.% Mn203 = weight of Mn203 / (weight of Mn203 + weight of Ce02)

(1)

Samples with mass ratios (Rm values)of 0, 0.20, 0.30, 0.335, 0.35, 0.45 and 1.0 were prepared.

2.2. Catalyst testing The concentration of phenol in water was 1.3 g/1. Tests were carried out at an oxygen pressure of 2.0 MPa in a temperature range of 100-140~ using a stainless steel batch autoclave reactor with an inner teflon jacket and equipped with a 600 rpm turbine (from Engineers Europe). At the end

269 of the experiment (30 or 120 min), the catalyst (2 or 4 g/l) was separated from the solution by filtration and the remaining organics were measured by means of a Total Organic Carbon (TOC) analyzer. The amount of cations eluted from the catalyst (Mn and Ce) was determined by Inductive Coupled Plasma-Mass Spectroscopy (ICP-MS). When the reaction was carried out for 120 min, liquid samples were withdrawn at regular time intervals and analyzed by HighPerformance Liquid Chromatography (HPLC) using a UV detector.

2.3. Characterization of the catalysts Surface areas (SBET) were measured with a Micromeritics Flowsorb II using the single point approximation of the BET equation for the adsorption of N2 at 77 K. Carbon deposits on the catalyst were quantified by elemental analysis involving combustion to CO2, absorption in a Ba(C104)2 solution and coulometric titration. XRD was perfoormed on a Kristalloflex Siemens DS000 diffractometer using the Kal,2 radiation of Cu (Z,=1.5418A) between 20 angles going from 2 to 90 ~ XPS analyses were performed with a SSX-100 model 206 X-ray photoelectron spectrometer from FISONS. The analysis chamber was operated under ultrahigh vacuum with a pressure close to 5x10 -9 Torr. The Cls, Mn2p, Ce2p, Ols, and Cls bands were swept successively. The binding energy (BE) values were calculated with respect to Cls (BE of C-C, H fixed at 284.8 eV). 3.

R E S U L T S AND DISCUSSION

.s

100

60 Ill

4o

00

I

20

40

60

8O

!

I

100 120 140 Time (min)

Figure 1. Wet air oxidation. Magnesium-cerium catalysts with different mass ratios as defined above. Conversion of phenol as a function of time. 20% MnO2; 9 30% MnO2; O 33.5% - Mn/Ce 1:1; 1 35% MnO2; A 45% MnO2 V 100% MnO2; & blank

270

3.1. Catalysts with different compositions The physico-chemical characterization of fresh Mn-Ce-O samples showed that they remained amorphous when decomposed at temperatures below 400~ X-ray diffraction bands for CeO2 and Mn203 were observed for calcined at 400 or 500~ The BET surface area of a sample with Rm=0.0 was 20 m2/g. For the other samples, BET surface areas were between 50 and 80 m2/g. The surface area of sample Mn/Ce= 1 was close to 50 m2/g. Surface areas increased after reaction, in some cases by about 100%. A weak crystallization to CeO2 and Mn203 and traces of Ce2(CO3)3 or MnCO3 were observed in samples after reaction. Only samples calcined at 500~ were used to measure catalytic activity. Figure 1 shows the conversion o f phenol in the presence of various Mn-Ce catalysts as a function of time. A deactivation was observed. At all compositions, a very high rate of phenol disappearance was observed during the first minutes of the reaction, but a plateau was rapidly reached. The consumption of phenol corresponds to a decrease of total organic carbon in the solution. Very few intermediate molecules (only traces) were identified in the solution. These were: benzoquinones, hydroquinone, catechol and some carboxylic acids. Acetic acid was not detected. The sample with Rm=0.335 was the most active. This sample was selected to study the deactivation process in more detail.

3.2. Catalysts with with Rm=0.335( Mn/Ce(atomic ratio)=l) A substantial leaching of the catalyst elements takes place during reaction; indeed, concentrations of 5 and 0.8 mg/1 of Mn and Ce respectively, were measured in the solution after reaction. A very high amount of carbon is detected on the used oxide. XPS analyses of the used oxide surface show a substantial increase of C on the surface of the catalyst after reaction (Figure 2).

Cis after catalytic test

..0_6 :b

Cls be fore catalytic test 290.9

I

I

288.9

286.9

I

~__

]L__

. . . .

I_

284.9 282.9 280.9 Binding energy (el/)

Figure 2. XPS analysis results. C 1s peak of Mn/Ce 1" 1 before and after the catalytic test (catalys 2g/l; phenol =l.3g/1; 2MPa of 02; temp.= 140~ reaction time = 30 min; pH=5.3). No XPS signal for Mn or Ce is detected after reaction. No new phases were detected by XRD. The carbon at the catalyst surface is probably present as a carbonaceous polymer, as has been reported previously for copper based catalysts [2]. At this stage of the investigation, we concentrated on polymer formation. As no new phase (neither another manganese or cerium oxide

271 nor a mixed Mn-Ce-O phase nor a manganese or cerium carbonate) was observed, the deactivation process has been assigned to either of the two following possibilities: (i) the elution of cations (principally Mn) and/or (ii) formation of carbonaceous polymer. Leaching of Mn +2 cations into solution could not be ruled out as playing a role in deactivation. But studies made in parallel showed that, even in the case in which cation concentration was lower, deactivation curves similar to those presented in Figure 1 were observed, suggesting the principal cause of deactivation to be polymer formation. The dependence of this process on experimental conditions was therefore studied in more detail. A catalytic test using an oxide that had already undergone a reaction, gave no phenol conversion. The catalyst was completely deactivated by the polymeric product that coated it. Submitting a used sample to a test without phenol in the aqueous medium with 2 MPa of oxygen did not destroy the polymer deposit. The polymeric product is strongly attached to the catalyst and is not soluble in conventional solvents (acetone, hexane). Additional tests were carried out to study effects of reaction and catalyst variables on the deactivation process. Results are presented in Tables 1-5. Thus the initial carbon present in the solution is transformed during the reaction principally into reaction intermediates, carbon dioxide and polymeric species. The rest remains as untransformed phenol. Elemental analysis of the catalysts quantifies the fraction of the initial carbon which has been transformed into polymers. The TOC analysis in the liquid phase permits calculation of the fraction of initial carbon which remains unreacted or has been transformed into reaction intermediates. Therefore, the initial carbon transformed into carbon dioxide was calculated as follows: CO2 (%)= [(ATOC- Elemental analysis) / (ppm C in initial solution)] x 100

(2)

Tables 1-5 present the effects of different reaction conditions and catalyst composition on the variation of TOC in solution, the percentage of initial carbon converted into polymers (quantity recovered on the catalyst divided by the quantity of carbon in the initial solution) and the percentage of initial carbon converted into CO2 (quantity of carbon reacted from the initial solution minus the quantity of carbon recovered on the catalyst, divided by the quantity of carbon in the initial solution), and the percentage of converted carbon recovered as polymers on the catalyst (quantity of carbon recovered on the catalyst divided by the quantity of carbon removed from the solution). Experiments were performed varying the pH (with NaOH or HC104) of the solution. One important objective of these measurements was to see if acidity favours polymerization. Effects of the oxygen partial pressure, the reaction temperature, the nature of the oxides, and hydroquinone, maleic acid and acetic acid, which have been proposed as intermediates in the mechanism of the reaction, were tested. Table 1 Influence of the pH in WAO of phenol, pH has been modified adding NaOH or HC104 polymer CO2 polymer from Conditions pH TOC from initial C from initial C converted C Mn-Ce oxide 2 g/l, 30 min, 140~ 2 MPa 02

8.5 7.5 6.5 5.3 2.5 1.2

(%)

(%)

(%)

(%)

72 72 77 79 80 18

55 58 63 66 68 18

17 14 14 13 12 3

77 80 82 84 85 >85

272 Table 2 Influence of the oxygen partial pressure in WAO of phenol polymer polymer from CO2 Conditions PO2 (MPa) TOC from initial C from initial C converted C (%) (%) (%) (%) Mn-Ce oxide 2 74 62 12 . 78 2g/1 4 87 71 16 78 140~ 6 86 72 14 80 2h

Table 3 Influence of the reaction temperature in WAO of phenol polymer polymer from C02 Conditions T (~ TOC from initial C from initial C converted C Mn-Ce oxide 2 g/1 2 MPa 02 2h

100 120 140

(%)

(%)

(%)

(%)

74 78 86

70 71 74

4 7 12

86 85 78

Table 4 Influence of the nature of the oxides in WAO of phenol polymer CO2 polymer from Conditions Oxide TOC from initial C from initial C converted C 4 g/l 140~ 2 MPa 02 2h

(%)

(%)

(%)

(%)

Mn203

93

77

16

82

Ce02

31

21

10

69

Table 5 Influence of the possible intermediates formed during WAO of phenol during reaction polymer CO2 polymer from Conditions Intermediate TOC from initial C from initial C converted C (1000 ppm C) (%) (%) (%) (%) Mn-Ce oxide 2 g/l, 140~ 2 MPa 02 30 min

hydroquinone maleic acid acetic acid

58 23 0

60

18

69

1

22

5

273 These tables show that, in nearly all these experiments, 75 to 85% of the converted phenol was transformed into non-oxidizable carbon. Carbon that disappears from the solution is found on the catalyst. Neither the pH of the reacting solution, nor 02 pressure nor reaction temperature have any influence on the extent of polymerization. Mn203 and CeO2 participate both to carbon deposit although less is deposited for CeO2. Aromatic molecules (phenol and hydroquinone) give large amounts of polymer. Only the non-aromatic molecule (maleic acid) does not give polymer. Surprisingly, acetic acid is not present in detectable amounts in our reaction. Without catalyst, phenol conversion is less than 3% at 140~ at a pressure of 2.0 MPa of oxygen. In addition, a test with 50 mg/1 of Mn +2 (quantity eluted when an experiment is run with 4 g/1 of Mn-Ce oxide) gives a phenol conversion of lower than 3%. This confirms that the formation of carbon deposit is heterogeneous. 3.3. M e c h a n i s m A plausible mechanism to explain the formation of polymers during reaction has been proposed by Pintar et al. [2] and Krajnc et al. [6]. This is based on the previous work of Sadana and Katzer [7], which suggested that heterogeneously catalyzed aqueous-phase phenol oxidation occurs by a free-radical mechanism which involves initiation on the catalyst surface by formation of a reaction intermediate, probably a hydroperoxide, homogeneous propagation and either predominantly homogeneous or heterogeneous termination, depending of the catalyst concentration. Catalysts would activate both reactants, phenol and oxygen, on different active sites at the catalyst surface. Oxygen adsorbed at the oxide surface is present mainly as superoxide 02- which decomposes further with the formation of O- ions which react with the phenol molecule. The corresponding radical undergoes protonation and splits off H202 to give the phenoxy radical. The authors [2,6] suggest that carbon deposition can be attributed to this heterogeneous-homogeneous free radical mechanism leading to polymer formation. They propose that the polymeric product mainly arises from two reactions taking place in the liquid phase: (i) the stepwise addition / polymerization of phenol to a glyoxal (C2 aldehyde) and (ii) polymerization of the C2 aldehyde. The phenoxy radical would be oxidized to acetic acid and/or C2-aldehyde and then oxidized to CO2 and water. Our results support the conclusion that intermediates (particularly hydroquinone) formed during reaction are probably promotors of polymer formations.

4.

CONCLUSION

An apparent high conversion of phenol is obtained by catalytic wet oxidation with Mn-Ce oxide catalysts. Phenol in solution reacts very rapidly over this catalyst but it is converted principally into a polymeric product. The catalyst deactivates completely during the first stages of the reaction. Two phenomena can explain deactivation: (i) leaching of the cations or (ii) polymer formation on the surface during reaction. Without excluding the role played by eluted cations or possibly by the formation of carbonates, it is proposed that deactivation of Mn-Ce oxide catalysts is principally due to polymer deposition. Polymer formation is practically independent of the composition of the catalysts, of the pH of the solution, of the oxygen partial pressure, of the temperature of reaction and of the composition of the oxides. Polymer formation can be attributed to a heterogeneoushomogeneous free radical mechanism.

5.

OUTLOOK

Mn-Ce oxide catalysts are very active in wet air oxidation of phenol. However, formation of polymer during reaction is a great drawback which limits its application. The same phenomenon has been observed with copper based catalysts. Further studies to control the processes by which

274 polymers are formed are crucial for further development or application of wet air oxidation of organic compounds. At this step of investigations five lines of research can be suggested for future studies: (i) identifying the nature and strength of the acid sites responsible of polymerization in order to modulate and to control their concentration and activity; (ii) applying solid state chemistry in order to isolate acid sites, for example by grafting the active components in a adequate solid matrix; (iii) using donors of oxygen spillover which have been effective to inhibit formation of carbonaceous polymers or precursors of coke during heterogeneous catalytic oxidation at high temperatures [8]; (iv) applying new chemical engineering concepts to control the liquid/solid ratio or more generally the way the liquid and oxygen contact the catalyst; (v) considering the possibility of purposely producing polymer [9], namely, designing a lowconcentration soluble polymer catalyst that would efficiently connect the phenols to polyphenols and be precipitated out with the polymer. REFERENCES

1. Vedprakash S. Mishra, Vijaykumar V. Mahajani and Jyeshtharaj B. Joshi, Ind. Eng. Chem. Res. 34 (1995) 2 2. Pintar A. and Levec J., J. Catal. 135 (1992) 345 3. Mantzavinos D., Hellenbrand R., Livingston A. G. and Metcalfe I. S. Applied catalysis B: Environmental, 7 (1996) 379. 4. Imamura S., Nakamura M., Kawabata N., Yoshida J. and Ishida S., Ind. Eng. Chem. Prod. Res. Dev., 25 (1986) 34 5. Courty P., Ajot H., Marcilly C. and Delmon B., Powder Technology 7 (1976) 21. 6. Krajnc M. and Levec J., The First European Congress on Chemical Engineering, Florence, Italy, May 4-7, 1997 7. Sadana A and Katzer J. R., J. Catal. 35 (1974) 140 8. L.T. Weng, L. Cadus, P. Ruiz and B. Delmon. Catalysis Today "Catalysts Deactivation" Vol. 11, N ~ 4, 22 January 1992 (J.J. Spivey, ed.), 455. 9. C.H. Bartholomew (private communication)

~ Elsevier Science B.V. All rights reserved Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

275

C o k e Deactivation o f Hydrotreating Catalysts: A Variable Site M o d e l F.E. Massoth, Department of Chemical and Fuels Engineering, University of Utah, Salt Lake City, UT

The decrease in activation energy with increasing coke content of a hydrotreating catalyst was modeled by a normal (Gaussian) variable-site activation energy function. Catalyst deactivation was correlated by site-selective deactivation for anthracene coke and site-preference deactivation for bitumen coke. 1. INTRODUCTION Hydrotreating catalysts, consisting of sulfided CoMo or NiMo supported on alumina, during processing undergo initial deactivation due to coke deposits from coking precursors in the feed. Coke forms from adsorption of polyaromatics and asphaltenes in the feed, followed by polymerization-condensation reactions. Model compound studies of coked catalysts showed that activity decreased with increase in coke content [1,2]. Further, activity losses for different reactions, e.g. hydrodesulfurization, hydrogenation, hydrodenitrogenation, otien do not show the same deactivation relationship with coke levels. Of significance, it has been recently found that the activation energy for a given reaction over a series of coked catalysts falls with increase in coke content [ 1]. The predominance of kinetic studies have assumed uniform sites on the catalyst surface. However, it has long been recognized that many catalyst surfaces exhibit non-uniform sites. Boudart and Djega-Mariadassou [3 ] have discussed the kinetics of non-uniform surfaces and conclude that "a non-uniform surface behaves catalytically ..like a uniform surface..", and that "rate equations are similar for a given mechanism on a uniform or non-uniform surface". This result justifies "the common practice of neglecting non-uniformity of catalytic surfaces in kinetic studies". However, it appears that uniform catalyst sites cannot adequately explain catalyst deactivation phenomena. The objective of the present study was to explain deactivation in terms of a model based on a variable activation energy site distribution on the catalyst. 2. THEORY

Since coke causes a decrease in catalytic activity for all reactions, it is obvious that it deactivates some catalytically active sites on the catalyst. If all the active sites have equal propensity for coverage of coke and are consequently completely deactivated towards reaction, then the activation energy should remain the same for the remaining unaffected sites, since activation energy does not depend on the number of sites present. This assumes that a deactivated site has no activity and that the activation energy of an uncovered site is not affected

276 i

by neighboring coke or coke-covered sites. In order to account for the lowering in activation energy with increasing coke content, we assume a distribution of catalytically active sites having a range of activities of different activation energies with the same intrinsic preexponential factor. The rationale for the latter assumption is that for a given reaction, the same mode of adsorption and change in entropy should occur over the reaction path. A normal (Gaussian) distribution has been chosen to describe the activation energy function. 2.1. Sulfided catalyst The normal (Gaussian) distribution of active sites, n(E), of activation energy E, is given by n (E) =

1 E X ~ - (E-E~ 2 OV/'2n 2o2

(1)

where Eo is the mean activation energy of all sites and sigma (o) is the distribution parameter. The total number of sites, n, and the mean activation energy, E.~ are given by EH

n=f n (E) dE

(2)

Er.

EH

E

m

En (E) dE f =EL

(3)

EH

f n(E) dE EL

where EL and E. are the lower and upper integration limits. The limits of integration strictly should be from 0 to .% but for practical considerations, it is useful to confine the limits to where the n(E) values become vanishingly small. For the sulfided catalyst, Em= Eo and the denominator term is unity. The individual rate constant for sites of activation energy E is k (E) = . (E) A oexp (-E/R 7') = . (E) e xp (-E/R T)

(4)

where Ao is the intrinsic preexponential factor. It is assumed that Ao is independent of E, as the same reaction takes place on each site. Since we will not be concerned with absolute rate constants but only relative rate constants, Ao is taken as unity. The global rate constant is then given by

k=fk(E) dE Ell

Er.

(5)

277

3.2. Coke deactivated catalysts For coked catalysts, deactivation is associated with poisoning of sites by coke precursors. Two cases are considered. (1) Site Selective Deactivation (SSD) - In this case, the coking precursor is exclusively adsorbed on the sites of highest activation energy first. Continued poisoning by coke then progressively deactivates sites of lower activation energy. That is, for E from EL to E', no sites are poisoned, while for E above E', all sites are poisoned. The degree of deactivation is given by the value of the deactivation variable E'. Thus, the number of active sites left on the coked catalyst, n', are E I

(6)

n/=[n (E) dE Er.

and the mean activation energy becomes E I

m /11

f En (E) dE

(7)

EL

1"1

I

The individual rate constant for sites of activation energy E is now k'(E) = n ( E ) e x p ( - E / R T )

, E < E'

(8)

The global rate constant and the relative activity of the coked catalyst is then given by E I

k'= f k'

(E) dE

(9)

E L

n=k~

(~0)

(2) Site Preference Deactivation (SPD) - In this case, the coking precursor adsorbs preferentially, but not exclusively, on the high energy sites. As poisoning proceeds with increasing coke level, more of the high activation energy sites are deactivated than the low activation energy sites. Thus, the requisite site preferential function, g(E), will be related to the value of E, and should have the following properties: (a) for no sites poisoned, g(E) = 1 (b) for all sites poisoned, g(E) = 0 (c) for intermediate site poisoning, g(E) should depend upon the difference between E and Eb since sites of activation energy EL are not poisoned (except at complete deactivation).

278 One such function, although others are possible, is g(E) = 1 - {(E - E , ) / ( E , - E , ) } t

(11)

where g is a site preference variable. This function fulfills the above criteria, i.e. for all sites poisoned: g = 0, g(E) = 0 for all E > EL for no sites poisoned: g -..% g(E) -. 1 for all E < EL In addition, the parameter g describes the particular variation in g(E) with E. Thus, for g < 1, g(E) curves concave with respect to E for g = 1, g(E) is linear in E for g > 1, g(E) curves convex with respect to E. Hence, the parameter g describes the various possible forms that g(E) can take relative to E. The site preference function only describes the propensity for adsorption of the coke precursor on a given site, but not the loss of sites due to coke. A deactivation function is now needed to describe the sites remaining at a particular degree of deactivation. The deactivation function, p(E), must meet the following criteria: (a) for no sites poisoned, p(E) = 1 (b) for all sites poisoned, p(E) = 0 (c) for intermediate poisoning, 0 < p(E) < 1. A form of deactivation that meets these criteria is p (E) = g (Efl'

(12)

where p is a site deactivation parameter, which is related to the degree of deactivation at a given E. Since g(E) is always less than unity (except at EL), then 0 < p < .0 meets the above criteria. Variations in p then relate to different degrees of deactivation of the coked catalyst. The number of active sites remaining at E and the mean activation energy now become n'(E) : , ( E ) p ( E )

(13)

EH

f En /(E) dE

E/= EL m

(14)

EH

f n (E) dE Er.

The variation in rate constant with E is now

k '(E) = n '(E) e x p (-E/R T)

(15)

279 and the global rate constant becomes EH

k'=f k' (E) d E

(16)

Er,

Finally, the relative activity for this case is given by (17)

R '= k~ 1"

30

0.8

0.8 -20

0.6-

'10.4-

4"

9

0.410

o.20

I'-"'

0

0.2"

I

9'

;

6 %C

8,

, 10

0 12

Fig. 1. Relative HYD activity R(HYD) and activation energy E, of coked catalysts vs. %C [ 1].

0 lO

30

Em, kcal/mol

Fig. 2. Correlation of R(HYD) for ANcoked catalysts by SSD and BIT-coked catalysts by SPD.

3. RESULTS Relative activity (R) and activation energy (E,) data for the hydrogenation (HYD) of naphthalene (NAP) vs. carbon content of the coked catalysts were taken from Chu et al. [ 1]. As seen in Fig. 1, both properties decreased with increase in coke level. Whereas catalyst deactivation was essentially unaffected by the coke precursor anthracene (AN) vs. bitumen (BIT), the decrease in activation energy was more pronounced for the AN-coked catalysts. From the data of the sulfided catalyst, the experimental activation energy, Eo, is known but not the distribution parameter sigma. It is assumed that deactivation by AN occurs by site selective deactivation (SSD). Various values of sigma were chosen and Eqs. (1) to (10) were solved to obtain a plot of R vs. E=. Integrations were carried out by summation of 30 points, with upper and lower limits ofEn = Eo + 50 and E L = Eo - 50. This ensures a cutoff of less than l xl 0"5% of the total sites. Use of this wide range of E values was necessary, as while n(E) becomes vanishingly small at low E's, the exp(-E/RT) term in Eq. (4) becomes very large. The E= values obtained from Eq. (7) were compared to the experimental data of Fig. 1 until the best fit was obtained. This was obtained for a sigma of 2.9, as shown by the left-hand curve in Fig.

280 2. This then provided the distribution parameter of the fresh catalyst for the HYD reaction. Figure 3 illustrates the relationship between k(E) and n(E) vs. E for the sulfided catalyst. It is seen that the lowest E sites, i.e. 18 to 22 kcal/mole, contain the highest activities, despite the relatively few sites in this region. On the other hand, values above about 25 kcal/mole represent a small fraction of the activity of the catalyst, despite having the majority of the sites. Thus, the catalytic activity resides predominantly in the region of low activation energy sites. It turns out that above about 7% carbon, the number of active sites becomes extremely small, despite there still being appreciable catalytic activity leg (40% of fresh activity). 0.2-

1 R(HYD)

0.8

o 0.6-

i

x~

n'(SSD]

,.=

0.1 o

m

0.05

0.4-

0.2"

(1

10

15

20

25 30 E, kcal/mol

35

Fig. 3. Distribution of sites and rate constants as a function of site activation energy.

40

0

......................"'-:...._ ........... 0

|

5

i

%C

10

15

Fig. 4. Fractions of sites let~ and relative activity vs. carbon content of coked catalysts.

The deactivation of the catalysts coked by BIT shows a much different relationship with respect to activation energy (Fig. 1). The experimental data are displaced to higher activation energies from those of the AN-coked catalysts. Since the distribution parameter of the fresh catalyst must be the same for the coked catalysts, is it obvious that deactivation by BIT cannot be by SSD. Hence, the BIT-coked catalysts are considered to deactivate by site preference deactivation (SPD). A value of the site preference parameter, g, was selected and a plot generated by varying p. Another g was selected and the process repeated until the best fit to the experimental BIT data of Fig. 1 was obtained. This occurred for a value of g of 2.7, and the fit is shown by the curve on the right-hand side of Fig. 2. Figure 4 illustrates the fractional loss in active sites with carbon content due to AN coke via SSD and BIT coke via SPD. Also shown is the relative activity curve derived from a power-law fit [2] of the data of Fig. 1. As can be seen, BIT coke deactivates considerably less active sites than does AN coke at the same carbon content, due to its lesser selectivity for adsorption on active sites. In both cases, appreciable residual catalytic activity exists when the fraction of remaining active sites becomes extremely small, in line with the high activity of the very low activation energy sites, as shown in Fig. 3. Data for the CNH reaction of indole over the same catalysts are displayed in Fig. 5 (CNH refers to the first C-N bond-breaking step forming o-ethyl aniline, e.g. the disappearance of

281 indole). Similar trends as for the HYD data are seen. Again, it was assumed that deactivation by AN followed the SSD model, while deactivation by BIT followed the SPD model. Figure 6 show fits of the experimental data with the deactivation models. The distribution parameter in this case was higher (3.3) than for the HYD case (2.9), while the g factor was slightly lower (2.4 vs. 2.7). 1-

30

0.8

f

0.8 -20

o.6 z

0.6-

o,li

-- D

-\

I

0.4-10

O.

0.2-

a _

./

0

2

4

6 %C

8

10

12

Fig. 5. Relative CNH activity R(CNH) and activation energy Ea of coked catalysts vs. %C [ 1].

10

30

Era, keal/mol

Fig. 6. Correlation of R(CNH) for ANcoked catalysts by SSD and BIT-coked catalysts by SPD.

4. DISCUSSION It is generally accepted that catalytic activity is associated with vacancies in the Mo-S slabs [4]. Reaction centers could involve a number of vacancies to accommodate the reacting molecule. Thus, it is not unreasonable to suppose that a distribution of centers of varying number and configuration of vacancies could exist. The normal energy-site distribution employed gave good correlations with the data. A gamma-function distribution, often employed in characterizing hydrocarbon fractions [6], also gave a satisfactory fit; however, one of the distribution coefficients was uncharacteristically large (56). For a large distribution coefficient, the gamma function approaches the normal distribution. Thus, the normal distribution seems to be the best choice for the energy-site distribution for the catalyst. The basis of the models is that all remaining sites in a partially deactivated catalyst retain their original activity, i.e. the activity of a site adjacent to a poisoned site is not affected, and that all poisoned sites have no activity. The results of the models show that catalyst deactivation can be satisfactorily correlated by a variable energy site model. The SSD model assumes that the coke-precursor species adsorbs initially on the highest energy sites. This is similar to adsorption of gases on heterogeneous adsorbents [5]. Continued adsorption of precursors occurs on progressively lower activation energy sites. At the same time, additional coke can grow on the adsorbed coke precursor by adsorption/polymerization reactions. It is interesting that the number of active sites let~ decrease sharply with only small amounts of coke present.

282 The SPD model relaxes the requirement for exclusive adsorption of the coke precursor on the highest energy sites. But, at the same time, adsorption is not uniform on all sites (else the activation energy of the coked catalysts would not change from that of the sulfided catalyst). Thus, an additional 'preference' parameter is introduced to account for this. Of course, we cannot be certain that deactivation by AN occurs by SSP. It could be due to weak SPP. It is not possible to distinguish this because sigma is not known a priori for the sulfided catalyst. But since sigma represents the distribution of the sulfided catalyst, it must be the same for deactivation of a given reaction by AN and BIT. Hence, it is certain that deactivation by BIT must follow SPP deactivation, in view of its displacement to higher E='s. A possible rationale for different models for AN- vs. BIT-coked catalysts is as follows: anthracene is relatively weakly adsorbed, and consequently requires high energy sites; whereas the coke precursors in BIT adsorb strongly, and consequently can utilize weaker sites, as well as strong sites. The latter may be due to the relatively high nitrogen content of the bitumen (1.1 wt.%), mostly in the form of N-heterocyclics. The latter are known to exhibit stronger adsorption than aromatics (AN). Although there is some scatter of the data for the CNH plot of Fig. 6, it is significant that the CNH data showed a poor correlation with the sigma value for HYD. Although sigma is a characteristic of the catalyst, it is also related to the active sites for aparticular reaction. Thus, sigma need not be the same if different sites are involved for the two reactions. Likewise, a different g factor (2.4) was obtained for the CNH data fit from that of the HYD data fit (2.7). Thus, it appears that the distribution of sites for HYD and for CNH are different, implying different sites for the two reactions. This is also manifested in the different extent of deactivation vs. coke content for the two reactions (Fig. 1 vs. Fig. 5). A similar conclusion was reached in a kinetic study involving indole and naphthalene [7]. This is not surprising as HYD represents the addition of hydrogen to an aromatic ring while CNH represents a C-N bond-breaking reaction. REFERENCES

1. K.-S. Chu, F.V. Hanson and F.E. Massoth, Fuel Processing Tech., 40 (1994) 79. 2. B.D. Muegge and F.E. Massoth, Fuel Processing Tech., 29 (1991) 19. 3. M. Boudart and G. Djega-Mariadassou, Kinetics of Heterogeneous Catalytic Reactions", Princeton University Press, Princeton, N.J., 1984. 4. H. Topsoe, B.S. Bjeme and F.E. Massoth in Catalysis Science and Technology, J.R. Anderson and M. Boudart (Editors), 11 (1996) 1. 5. S. Sircar, Ind. Eng. Chem. Res., 30 (1991) 1032. 6. C.H. Whitson, Soc. Petrol. Eng. J., (1983) 683. 7. F.E. Massoth, K. Balusami and J. Shabtai, J. Catal., 112 (1990) 256.

~ Elsevier Science B.V. All rights reserved Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

283

On the M e t a l D e p o s i t i o n Process during the H y d r o d e m e t a l l a t i o n of V a n a d y l Tetraphenylporphyrin J.P. Janssens ~, R.M. de Deugd, A.D. van Langeveld, S.T. Sie and J.A. Moulijn Delft University of Technology, Faculty of Chemical Technology and Materials Science, Julianalaan 136, 2628 BL Delft, The Netherlands, Tel. (31)-15-2784381, Fax. (31)-15-2784452 Present address: Unilever Research Laboratory Vlaardingen, Olivier van Noortlaan 120, 3133 AT Vlaardingen, The Netherlands, e-mail: [email protected]

This paper focuses on the metal deposition process during hydrodemetallisation (HDM) of vanadyl-tetraphenylporphyrin (VO-TPP) under industrial conditions. In catalyst pellets of a wide pore, low loaded molybdenum on silica, the vanadium deposition process was determined with EPMA and HREM. The effect of quinoline and H2S on the vanadium deposition profile is studied and an attempt is made to simulate the deposition profiles based on intrinsic reaction kinetics and percolation concepts. A reference vanadium deposition experiment is carried out in order to assess the influence of quinoline and H2S. Quinoline showed to decrease the rate of metal removal, the amount of vanadium deposited is lower as compared to the reference experiment. The shape of the vanadium deposition profiles is similar in both cases. A deposition maximum is observed in the centre of the pellet, indicating that the vanadium deposition process is not diffusion limited and that a sequential reaction mechanism applies for VO-TPP HDM. Low H2S partial pressure resulted in different vanadium deposition profiles as a function of the axial position in the reactor. At the inlet of the reactor, similar shaped profiles as the reference experiment were found, however, at the outlet of the reactor a shift towards M-shaped profiles was found indicating a diffusion limited vanadium deposition process. This shift in vanadium deposition profiles is explained by the build-up of the last intermediate resulting in a higher metal deposition rate. Simulations of the vanadium deposition profiles showed reasonable results at reactor inlet conditions. With HREM, vanadium deposits were found in a layered structure both in the surrounding of the active phase, causing active site poisoning, and as isolated vanadium sulfide clusters, likely causing active site generation. In our vanadium deposition experiments (up to 9 wt.% vanadium) no significant catalyst deactivation occurred, indicating that both effects are compensated.

284 1. INTRODUCTION Catalyst deactivation in hydrodemetallisation (HDM) is caused by the interaction of the metal deposits with the original active phase ('active site poisoning') and the loss of pore volume due to the obstruction of catalyst pores ('pore plugging') (1). However, metal deposits also have an auto-catalytic effect on the hydrodemetallisation reaction, thus 'active site generation' may occur in low active phase loaded or bare support catalyst systems. Modelling of the metal deposition process is of great importance in process design and operation, in order to give reliable predictions on catalyst life-time and activity. Detailed knowledge of reaction kinetics and restrictive diffusion are indispensable, description of the catalyst porous texture during metal deposition is necessary and often neglected in literature. The catalyst deactivation model is an integration of reaction kinetics, diffusion and the catalyst porous texture in the mass balances. The reactor model essentially determines the boundary conditions of the mass balances. In present work, a study is performed on metal deposition profiles in a wide-pore silica catalyst using vanadyl-tetraphenylporphyrin (VO-TPP) as model compound. The simplified HDM reaction mechanism of Vanadyl-TetraPhenylPorphyrin is depicted in Figure 1 (2). The reactant VO-TPP is reversibly hydrogenated to Vanadyl-TetraPhenylChlorin (VO-TPC), which can be reversibly hydrogenated to Vanadyl-TetraPhenylisoBacteriochlorin (VO-TPiB). The last step is a lumped ring cleavage step in which VO-TPiB is further hydrogenated and broken into ring fragments. Vanadium is removed and deposited on the catalyst surface. A detailed study on this reaction mechanism is presented elsewhere (3, 4).

~

I-I2

k,

(~

H2 k2

VO-TPP

H2

H2

VO-TPC

1(4

" '-'

'-'

VO-TPiB

ks ~

1_.t2S

VS , deposit + ringfragments

Figure 1. Hydrodemetallisation reaction mechanism ofvanadyl-tetraphenylporphyrin (VO-TPP).

A previously proposed state-of-the-art catalyst deactivation model, based on percolation concepts (5, 6), is proposed to tackle the problem of the changing catalyst porous texture (see Figure 2).

285

initialcatalystporoustexture

~

~

voidspace solidmaterial

l l =metaldeposition

catalyst porous texture after metal deposition process

Figure 2. Cubic tessellation model to describe the changing catalyst porous texture.

The effect of a nitrogen containing compound (7), i.e. quinoline, and the presence of H2S on the vanadium deposition is assessed. An attempt is made to predict the experimentally found vanadium deposition profiles by the catalyst deactivation model based on percolation concepts. Furthermore, the vanadium deposits are localised using HREM.

2. EXPERIMENTAL AND COMPUTATIONAL TECHNIQUES

2.1. Preparation and characterisation of the fresh catalyst For the vanadium deposition experiments a molybdenum on silica catalyst is prepared. A molybdenum catalyst is used since it is active for the hydrodemetallisation of VO-TPP and it can be used as a reference element for determining the radial vanadium deposition in catalyst pellets. The catalyst is prepared by high temperature batch adsorption (HTBA) (8) which ensures an uniform distribution of the active phase in the catalyst pellet. The characteristics of the prepared catalyst are summarized in Table 1. Table 1 Characteristics of Mo/SiO2 catalyst.

Mo/SiO2

surface area m2/g

pore volume ml/g

pore radius nm

Mo loading wt.%

52.5 + 1.3

0.69 + 0.04

27.0 + 1.7

1.79 + 0.23

286

2.2. HDM reaction kinetics The kinetic experiments were carried out with VO-TPP and o-xylene as solvent in a 250 ml batch autoclave. Approximately 80 mg of catalyst (pellet diameter: 63 -160 ~tm) was weighed and introduced in the reactor. The catalyst was sulfided in situ for two hours with dimethyldisulfide (623 K, 6.0 MPa H2). After sulfiding, VO-TPP dissolved in o-xylene (total reaction volume: 125 ml) was introduced. The reactor was filled to 0.34 MPa with a mixture of 15 mol % H2S in H2. Next H2 was added to a total pressure of 5.0 MPa and the reactor was heated to a temperature of 553 K. The HDM reaction was carried out at industrial conditions, 553 K and 9.0 MPa H2 pressure. At certain time intervals a liquid sample was taken from the autoclave and analysed ex-situ with UV-Vis spectroscopy to determine the porphyrin concentrations (9). 2.3. Vanadium deposition experiments 2.3.1. Microflow equipment The vanadium deposition experiments were performed in a so-called microflow reactor (9). A liquid feed, consisting of VO-TPP dissolved in an isomer mixture of xylene, is mixed with hydrogen and fed into a fixed-bed reactor at 9.0 MPa and 553 K. The reactor has a length of 35 cm and a internal diameter of 15 mm. Four round trays of 5 mm high were placed in this reactor with the catalyst pellets. The top side of each tray is open and the bottom side consists of a wire gauze. These trays enable the separation of the catalyst pellets during filling and emptying of the reactor and determine the effect of the axial position in the reactor on the vanadium deposition process. At the bottom of the reactor a sieve plate was placed to avoid particles flowing out of the reactor. The catalyst pellets with a diameter of 2.1 mm were to large for microflow applications (10). For this reason the catalyst beds was diluted with fine (0.1 mm) SiC pellets. Glass wool and the SiC on top of the catalyst beds are sufficient to provide a good radial distribution of the liquid in the reactor and complete wetting of the catalyst pellets. 2.3.2. Sulfiding and reaction conditions The catalyst was sulfided before starting the vanadium deposition experiment. The sulfiding has been carried out with 3 N1/h 15 mol % H2S in H 2 at 623 K for three hours at atmospheric pressure. Three different long duration vanadium deposition experiments, I, II and III, have been carried out. Experiment I was a test under reference conditions. Experiment II was carried out to determine the influence of quinoline on the vanadium deposition process. Experiment III was carried out to determine the effect of a low H2S partial pressure on the vanadium deposition profiles. Each experiment was split into two parts A and B. Experiment A is the first part of the vanadium deposition experiment. At the end of experiment A, the reactor was removed from the microflow equipment and catalyst pellets from every tray were collected and analysed. After this short stop, the reactor was mounted again in the microflow equipment, the catalyst sulfided and the experiment B started. Details on the reaction conditions of the three experiments are summarized in Table 2.

287 Table 2 Reaction conditions applied in vanadium deposition experiments. Experiment

IA

IB

II A

II B

IIIA

III B

Duration (h) Liquid feed rate (ml/h) Concentration VO-TPP (mol/m 3) Concentration DMDS (mol/m 3) Feed rate hydrogen (N1/h) Concentration quinoline (mol/m 3)

350 12.9 0.39

338 12.9 0.39

251 28.0 0.39

294 28.0 0.39

286 397 12.3 12.3 0.039 0.039

10.6

10.6

10.6

10.6

1.06

1.06

3.8 0

3.8 0

4.9 150

4.9 150

2.5 0

2.5 0

The major differences between experiment I and II are the liquid feed rate and the addition of quinoline. Approximately 150 mol/m 3 quinoline equivalent to 2000 wt.ppm nitrogen is added to the liquid feed in experiment II. The liquid feed rate was 12.9 ml/h for experiment I and 28.0 ml/h for experiment II. Dimethyldisulfide was used for the in-situ production of H2S, which is a reactant in the lumped ring cleavage step of the VO-TPP hydrodemetallisation reaction. As indicated in Table 2, experiment III has a ten times lower concentration DMDS in the feed, which resulted in a lower H2S partial pressure in the reactor. The liquid effluent was analysed during the experiments with UV-Vis spectroscopy.

2.3.3. EPMA and HREM analysis After each vanadium deposition experiment, two pellets from every axial position in the reactor were analysed with EPMA (Electronprobe microanalysis, Jeol JXA 733). Together with atomic absorption spectroscopy (AAS), in which the overall weight fraction of vanadium and molybdenum were determined, a quantitative analysis of the radial distribution of metals could be obtained. Spent catalysts were also investigated with a Philips EM 30-ST HREM equipped with EDX for element identification. 2.4. Computer simulations on vanadium deposition in catalyst pellets The model used to describe the metal deposition process during hydrodemetallisation includes a system of nonlinear parabolic partial differential equations (PDEs) in one space variable (5, 6). These equations were solved numerically with a CONVEX 3840 workstation. Subroutines to solve the set of PDEs are obtained from the NAG Fortran library (1988) (11). The simulations were performed in using the reaction rate constants obtained from the autoclave experiment. The bulk diffusion coefficients for reactant and intermediates were estimated

288 according to the Wilke-Chang correlation (12). A value of 3.0" 10-s m2/s was found, the restrictive factor for hindered diffusion determined by Satterfield (13) was used. The values characterising the porous texture of the catalyst are taken from Table 1. A cubic tessellation was used as percolation model to evaluate the change in catalyst porous texture.

3. RESULTS

3.1. Estimation of reaction kinetic parameters of VO-TPP hydrodemetallisation Figure 3 shows the concentration versus time profiles as obtained from autoclave experiments for VO-TPP hydrodemetallisation over the Mo/SiO2 catalyst. As can be observed, the concentration of reactant VO-TPP is decreasing, the first intermediate VO-TPC is passing through a maximum and approaching zero. The concentration of the second intermediate VOTPiB always remains very low. Kinetic parameters were obtained by evaluation of the concentration-time data according to a two-site (hydrogenation/ring cleavage site) LangrnuirHinshelwood model with competitive adsorption of hydrogen (2). Table 3 summarizes the values of the parameters of the reaction kinetics test, kl (first hydrogenation), k 3(second hydrogenation), and k5 (lumped ring cleavage) normalized on the total catalyst surface area. The solid lines in Figure 3 represents the fitted concentrations according to the reaction kinetics model.

0.25

0.20 o 0.15

0.10

0

~

"~

0.05

0

20

40

60

80

100

120

Time (ks)

Figure 3. Concentration-time profiles for VO-TPP HDM over a Mo/SiO2 catalyst. (m. VO-TPP, +: VO-TPC; *: VO-TPiB;--" fit according to two-site model (2))

289 Table 3 Values for the VO-TPP hydrodemetallisation reaction rate constants.

Reaction rate constants

kl m3/m2eat'S

k3 m3/m2eat'S

k5 m6/m2cat.mol-s

(1.7 + 0.1).10 -8

(6.6 + 1.1).10 .8

(4.6 + 2.4)-10 -7

3.2. Vanadium deposition experiments This section is divided into the three vanadium deposition experiments I (reference), II (addition of quinoline) and III (low H2S partial pressure). Quantitative vanadium deposition profiles in catalyst pellets are presented as a function of the axial position in the reactor. Experiment I A and B Figure 5a shows the vanadium deposition profiles in the Mo/SiO2 catalyst pellets as a function of the axial position in the reactor for experiments I A and B. As can be observed, for all pellets the vanadium deposition maximum is observed at the centre of the pellet. An increasing amount of deposited vanadium is found in axial direction in the reactor. Experiment 11A and B Figure 5b shows the vanadium deposition profiles as a function of the axial position in the reactor for experiments II A and B. As can be observed much lower amounts of vanadium are deposited on the catalyst pellets. As compared to experiments I A and B, deposition maxima are observed in the centre of the pellet. Experiment III A and B Figure 5c shows the vanadium deposition profiles as a function of the axial position in the reactor for experiments III A and B. In this experiment, similar reaction conditions were applied as in experiments I A and B, except a lower H2S partial pressure and a lower feed concentration VO-TPP was added. A shift in the shape of the vanadium deposition profile is observed in axial direction in the reactor. The first two trays show, similar to experiment I and II, a vanadium deposition maximum in the centre of the pellet. Tray 3 and 4 show a deposition maximum in the outer shell of the pellet and very little vanadium deposition in the centre. Again the deposited amount of vanadium are low as compared to experiment I.

UV-Vis analysis of the effluent liquid showed no significant change in conversion level of reactant and intermediates during experiment I, II and III.

290

tray 1

tray 1 4

tray 1

3 2

......... ........ ii,ii

9

,.,::; 9

B

I -

:::::::::::::::::::::: ~

IIII

......

All

B

r +..;

15

"

o

-.

"~176

....

"""''"'"a

i A

0 A

tray 2 2o

."

.

tray 2

4

tray 2

3" o,, . . . . . .

9~

"

B

9176 2~

"

-"

B

5

o

-

~.:::::::'~7~7"::::-

A ~

ii

B

|

-'-'a,.

tray 3 "~

2o

tray 3 tray 3 B .

o.

o o-

9 ...~

. ~ ~~ o ;_.~., t; .............

" 9

A

o.

.~ " 9176

B ~

-

9o -.. 9149

~

. .....

o...o"

......

_ tray 4

9

.o''

.....

-

A

A

-. tray 4

tray 4 ~ ~

-

.,.

.

9 B

o

2 o

.-

......

.-o..

o

o.

.

-~

- ' ~ 1 7 6 1 7 6 1. . .4. 9 " ' ' I -|~176

A o Outside

Centre

a.

~ .~., p

i

~ ~ 1 4 9 ,,

t

0 I

.. .. .. .. .. .. ..

. .."" ".....'.. -.,. . . ........ .

A

Outside Outside

Radial position (-)

B

B ~176

Centre

Outside

Outside

Centre

Outside

Radial position (-)

Radial position (-)

b.

c.

Figure 5. Vanadium deposition profiles in Mo/SiO2 catalyst pellet as a function of the axial position; a- reference experiment I A and B; b- experiment II A and B with quinoline addition and c- experiment III A and B with low H2S partial pressure.

3.3. Localisation of vanadium deposits o n M o / S i O 2 catalyst Two different morphologies of vanadium deposits on the M o / S i O with HREM and EDX; i) a phase non-visible for HREM, but present according to EDX, ii) a crystalline phase, a layered structure of vanadium sulfide.

2

catalyst could be identified

Figure 6 shows a typical view of the spent Mo/SiO2 catalyst with the layered structure of vanadium sulfide. According to EDX, vanadium sulfide is present on silica. The copper peaks in the EDX spectrum are due to the grid which supports the catalyst powder in the HREM. HR~M/EDX analysis at several point on the spent catalyst showed that vanadium was present as isolated clusters on the bare support and in the surrounding of the active phase.

291 Cu Si

v

Mo

Mo

Element

Figure 6. HREM image and EDX spectrum of spent catalyst.

4. DISCUSSION Vanadium deposition profiles in catalyst pellets have been determined by various researchers

(14-30). For this purpose porphyrinic model compounds and industrial feedstocks are used. The used catalyst are mainly conventional hydrotreating catalysts with narrow pores. Therefore, metal deposition profiles show mainly deposition in the outer shells of the catalyst pellets (M- or Ushaped profiles), indicating that the metal deposition process is diffusion limited. Present work shows the vanadium deposition process on a low molybdenum loaded silica catalyst with wide pores. Our results show, for the reference experiment I A and B, maximum vanadium deposition in the centre of the pellet. The diffusion rate of the reactant and intermediates into the catalyst pellet is high enough to reach the centre of the pellet. The deposition maximum in the centre of the pellet can be explained by the sequential reaction mechanism of VO-TPP hydrodemetallisation. At the outside of the pellet the concentration of reactant VO-TPP is high, however in the first tray almost no intermediates are present. VO-TPP has to diffuse in the pellet, gradually forming the intermediates VO-TPC and VO-TPiB and finally vanadium is deposited on the catalyst surface. In the subsequent trays in axial direction in the reactor, intermediates are present at the outside of the pellet due to hydrogenation of the reactant and the diffusion out of the catalyst pellet. The build-up of the last intermediate causes a higher rate in the vanadium deposition (proportional to the concentration) and therefore higher amounts of vanadium are found in the subsequent trays. The results of experiment II A and B show an identical course in the vanadium deposition profiles as the reference experiment I. The amounts of vanadium deposited are of course lower which is due to the competitive adsorption of quinoline and the higher liquid feed rate. The decrease in the rate of metal removal caused by quinoline was already found in previous work

(8).

292 Experiment IIIA and B, which were carried out under a low H2Spartial pressure and a lower feed concentration of VO-TPP showed a remarkable shift in the vanadium deposition profiles as a function of the axial position in the reactor. In the first two trays a similar vanadium deposition profile is found as in the reference experiment I. However, the shift towards M-shaped deposition profiles in tray 3 and 4 indicate a diffusion limited metal deposition process. This shift is due to the low H2S partial pressure which causes a build-up in the concentration of the last intermediate VO-TPiB. Apparently, in the last two trays the concentration of VO-TPiB becomes so high, that the metal deposition rate overrules the diffusion rate into the catalyst pellet which causes metal deposition in the outer shell. An attempt is made to model the experimental vanadium deposition profiles of the three experiments. Figure 7 shows the experimental and simulated vanadium deposition profiles for experiment I, II and III A and B for the first tray in the reactor. As can be observed identically shaped vanadium deposition profiles are obtained for both experiment and simulation. However, an increasing discrepancy was found in the axial direction, which is caused by uncertainties in the input parameters of the model, such as the boundary concentrations of reactant and intermediates in the subsequent trays and variation in reaction rate constants.

20t,o5

ayl I

.

tray 1

o Outside

Centre

Outside Outside

Cenlre

Radial position (-) Radial position (-)

a.

b.

Outside

Outside

Centre

Outside

Radial position (-)

c.

Figure 7. Comparison between experiment I, II and IIIA and B and model simulations for the first tray in the reactor (the solid lines represent the model simulations, a- Experiment I; bExperiment II and c- Experiment III). Vanadium deposits showed to have two different morphologies, i.e. disperse and crystalline. Vanadium could be present as layered structures (similar as MoS2) an are therefore difficult to identify. The location of vanadium, is either in the surrounding of the active phase or as isolated clusters on the support material. Vanadium deposits in the surroundings of the active phase cause active site poisoning due to the coverage of the active phase or the edges. Active site generation occurs by the formation of isolated vanadium phase on the support material which has catalytic activity for hydrodemetallisation. Both phenomena could explain the UV-Vis results of the liquid effluent. Apparently the nett effect is zero catalyst deactivation up to 9 wt.% vanadium deposition.

293 5. CONCLUSIONS Present work details on a study on the metal deposition process during hydrodemetallisation of vanadyl-tetraphenylporphyrin on a wide pore molybdenum on silica catalyst. The effect of quinoline and a low H2S partial pressure on the vanadium deposition profiles in catalyst pellets are determined in comparison with a reference experiment. The vanadium deposition process showed profiles with deposition maxima in the centre of the pellet indicating the absence of diffusion limitations and supporting a sequential reaction mechanism for VO-TPP HDM. Quinoline addition showed to have an decreasing effect on the rate of metal removal and showed similar shaped deposition profiles. The low H2S partial pressure caused a change of the vanadium deposition profiles into M-shaped profiles due to the build-up of the last intermediate and an increasing metal deposition rate. Simulations of the vanadium deposition process based on intrinsic kinetic data and percolation theory showed reasonable agreement at reactor inlet conditions. The location of vanadium deposits are both in the surrounding of the active phase, causing active site poisoning, and as isolated vanadium sulfide clusters, causing active site generation. In our vanadium deposition experiments (up to 9 wt.% vanadium) no significant catalyst deactivation occurred, indicating that both effects are compensated.

Acknowledgments Dr. P. Kooyman, Dr. H. Zandbergen, P. Colijn and E. Fakkeldij of the Materials Science Department of Delft University of Technology are gratefully acknowledged for the HREM and EPMA measurements. The European Community (contract number JOUF 0049) and Shell International Oil Products (SIOP, Amsterdam) are thanked for the financial support.

REFERENCES 1. Sie, S.T., in 'Catalyst deactivation', Delmon, B. and Froment, G.F. (Eds.), Elsevier Scientific Publishing Company, Amsterdam (1980) 545. 2. Bonnr, R.L.C., Steenderen, P. van, and Moulijn, J.A., Bull Soc. Chim. Belg., 100 (11-12) (1992) 877. 3. Janssens, J.P., Characterisation, testing and deactivation of sulfided catalysts in the hydrodemetallisation of vanadyl-tetraphenylporphyrin' Ph.D thesis, Delft University of Technology, The Netherlands, Chapter 7 (1996). 4. Janssens, J.P., Elst, G., Schrikkema, E.G., Langeveld, A.D. van, Sie, S.T, and Moulijn, J.A., RecL Trav. Chim. Pays-Bas, 115 (1996) 465. 5. Janssens, J.P., Characterisation, testing and deactivation of sulfided catalysts in the hydrodemetallisation of vanadyl-tetraphenylporphyrin' Ph.D thesis, Delft University of Technology, The Netherlands, Chapter 5 (1996). 6. Janssens, J.P., Langeveld, A.D. van, Sie, S.T., and Moulijn, J.A.,ACS Symp. Ser., 634 (1996) 238.

294 7. Janssens, J.P., Witte, G., Langeveld, A.D. van, Sie, S.T., and Moulijn, to be published in FUEL. 8. Bosch H., Ph.D thesis, University of Twente (1987). 9. Janssens, J.P., Characterisation, testing and deactivation of sulfided catalysts in the hydrodemetallisation of vanadyl-tetraphenylporphyrin' Ph.D thesis, Delft University of Technology, The Netherlands, Chapter 2 (1996). 10. Sie, S.T., Rev. Inst. Fr. P~troL, 46 (4) (1991) 501. 11. NAG Fortran library (1988). 12. Tsai, C.H., Massoth, F.E., Lee, S.Y., and Seader, J.D., Ind. Eng. Chem. Res., 30 (1991) 22. 13. Satterfield, C.N., Colton, C.K., and Pitcher, W.H., AIChE J., 19 (3) (1973) 628. 14. Sato, M., Takayama, N., Kurita, S., and Kwan, T., Nippon Kagaku Zasshi, 92 (1971) 834. 15. Oxenreiter, M.F., Frye, C.G., Hoekstra, G.B., and Sroka, J.M., Jpn. Petrol Inst., Nov. (1972) 30. 16. Dautzenberg, F.M., Klinken, J. van, Pronk, K.M.A., Sie, S.T., and Wijffels, J-B, A CS Symp. Ser., 65 (1978) 254. 17. Tamm, P.W., Hamsberger, H.F., and Bridge, A.G., Ind. Eng. Chem. Process Des. Dev., 20 (2) (1981 ) 262. 18. Galiasso, R., Blanco, R., Gonzalez, C., and Quinteros, N., FUEL, 62 (1983) 817. 19. Pazos, J.M., Gonzalez, J.C., and Salazar-Guillen, A.J., Ind. Eng. Chem. Process Des. Dev., 22(1983) 653. 20. Agrawal, R., and Wei, J., Ind. Eng. Chem. Process Des. Dev., 23 (1984) 505, 515. 21. Takeuchi, C., Asaoka, S., Nakata, S., and Shiroto, Y., Prepr. A CS Div. Petrol. Chem., 30 (1985) 96. 22. Ware, R.A., and Wei, J., J. Catal., 93 (1985) 122, 135. 23. Khang, S-J, and Mosby, J.F.,Ind. Eng. Chem. Process Des. Dev., 25 (2) (1986) 437. 24. Aldag, A.W., Prepr. ACS Div. Petrol. Chem., Denver Meeting, (1987) 443. 25. Sie, S.T., i2-Procestechnologie, 5 (1987) 37. 26. Pereira, C.J., Beeckman, J.W., Cheng, W-C, and Suarez, W., Ind. Eng. Chem. Res., 29 (1990) 520. 27. Toulhoat, H., Szymanski, R., and Plumail, J.C., Cat. Tod., 7 (1990) 531. 28. Simpson, H.D., in 'Catalyst deactivation', Bartholomew, C.H. and Butt, J.B. (Eds.), Elsevier Science Publishers, Amsterdam (1991) 265. 29. Hubaut, R., Dejonghe, S., Grimblot, J., and Bonnelle, J.P., React. Kinet. Catal. Lett., 51 (1) (1993) 9. 30. Fozard, P.R., McMillan, J.W., and Zeuthen, P.,J. Catal., 152 (1995) 103.

9

Elsevier Science B.V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

295

C o k e F o r m a t i o n in Fluid Catalytic C r a c k i n g M.A. den Hollander, M. Makkee*, and J.A. Moulijn Department of Chemical Engineering, Section Industrial Catalysis, Delft University of Technology, Julianalaan 136, 2628 BL, Delft, The Netherlands.

The catalysts used in Fluid Catalytic Cracking (FCC) are reversibly deactivated by the deposition of coke. Results obtained in a laboratory scale entrained flow reactor with a hydrowax feedstock show that coke formation mainly takes place within a time frame of milliseconds. In the same time interval conversions of 30-50% are found. After this initial coke formation, only at higher catalyst-to-oil ratios some additional coke formation was observed. In order to model the whole process properly, the coke deposition and catalyst deactivation have to be divided in an initial process (typically within 0.15 s) and a process at a larger time scale. When the initial effects were excluded from the modeling, the measured data could be described satisfactory with a constant catalytic activity.

1. INTRODUCTION Fluid Catalytic Cracking (FCC) is one of the main industrial catalytic processes, in which heavy hydrocarbons are converted into lighter hydrocarbons. The main products are gasoline and light cycle oil. The reactions are carried out in an entrained flow reactor, in which the catalyst is transported by the vaporous hydrocarbons. Typical contact times between . I Reactor feedstocks and catalyst are 2-10 s. Within this Products \ll time frame the catalyst is reversibly deactivated by coke deposition. Coke is an inevitable bySpent ]~l~ product of the cracking reactions. The heat that Ca~.l.ys.t..I.......~is generated by burning off the coke in the C er 1 Feed regenerator is used to generate the necessary heat for heating and vaporization of the feedstock and for the endothermic cracking /i reactions. The special characteristics of the FCC process are the entrained flow regime and // short contact times. These characteristics make it difficult to study coke formation in catalytic cracking on a laboratory or industrial scale. To Figure 1. Outline o f the microriser. *To Whom correspondence should be addressed

296 improve the modeling of FCC, a good description of coke formation and catalyst deactivation on a realistic time scale are essential. In this paper attention will be paid to coke formation in FCC as a function of residence time (x, between 0.15 and 5 s, based on outlet conditions), and catalyst-to-oil ratio (CTO, between 2 and 6), studied in a laboratory scale entrained flow reactor, the so-called microriser, as illustrated in figure 1. In previous work it has been shown that the reactor behaves for more than 98% as a plug flow reactor.

2. THEORY A commonly used, simplified model to describe FCC cracking kinetics is the five lump model according to Corella et al. [2], as depicted in figure 2. In this model different components are lumped together into pseudo-components according to their boiling point and only the different relevant reactions between these lumps are taken into account. The different lumps, as defined in table 1 are: Heavy Cycle Oil (HCO), Light Cycle Oil (LCO), gasoline, Coke, Y5

Table 1. Definition of lumps used in the five lump model

HCO, Yl

,.-

L e o , Y2

Gasoline, Y3

,.-

Gas, Y4

lump gas gasoline LCO HCO coke

composition/boiling point range H2, Cl-C4 C5- 494 K 494-643 K >643 K carbon deposit

Figure 2. Five lump reaction scheme [2]. gas, and coke. The weight fractions of the different lumping groups are represented by yl-ys, respectively. For a fundamental description of the cracking kinetics this model is not suitable, since properties of the feedstock and catalyst are not incorporated in the model, and the product quality of the different lumps is not taken into account. The catalyst and feedstock properties affect the values of the reaction rate constants and it is necessary to measure and evaluate the reaction rate constants for every new combination of catalyst and feedstock. Kinetic models based on elementary steps do not suffer from these disadvantages [3,4]. The present paper aims at establishing the coke profiles in combination with activity and, as a consequence, for the modeling that will be described the five lump model is satisfactory. The general equation for the reaction rate of reaction j is given by: rj = ~j kj Yini in which

rj: ~j" kj: Yi: ni:

(1)

reaction rate activity function reaction rate constant weight fraction of component i reaction order for component i

[kgi kgcatl S"1] [-] [kgi TM kgfeedn kgcatl S-l] [kgi kgfeed-l] [-]

297 The values used for the reaction order n are 2 for the reaction of HCO, 1.5 for the reaction of LCO, and 1 for the reaction of gasoline [5, 6]. The deactivation of the catalyst is caused by the deposition of coke, so the activity functions should be related to the amount of the coke deposited on the catalyst. As a first estimation the activity of the catalyst will not be based upon the amount of coke formed on the catalyst, but upon the residence time x of the catalyst in the reactor, represented by the following equation: d{~j d'l: - kd {~ in which

kd: m:

(2) deactivation rate order of deactivation

IS-l] [-]

Different orders of deactivation will be used in the modeling. A simplification will be that the activity function for all components is the same. Using the reaction scheme given in figure 2 and the reaction rates as described by equation 1, in combination with a plug flow reactor model, mass balances can be derived for the different lumped groups, resulting in: dYl - r CTO ( - k I - k e - k 3 - k 7 ) y2 d'c

(3)

dye dx - r

(4)

(kl y2 + ( - k 4 - k 5 - k z ) y ~ "5)

dy3 - r CTO (k 2 y~ + k 5 Y21.5 _ k6 y 3) dx

(5)

dy4 - q~CTO (k3 y21 + k 4 Y ~.5 + k 6 Y3 ) dx

(6)

dY5 d1: - r CTO (k 7 y2 + ks y 12.5)

(7)

in which

CTO: catalyst-to-oil ratio

[kgcat kgfeed-l]

3. E X P E R I M E N T A L

Cracking experiments have been performed using the microriser, an industrial equilibrium catalyst, and a commercially available hydrowax feedstock. This feedstock was selected for its low aromatics content, low Conradson Carbon Residue (CCR), and very low sulfur or metal concentrations. The characteristics of the feedstock are given in table 2. The microriser, as illustrated in figure 1, is a laboratory scale entrained flow reactor. The actual reactor consists of a bended tube (bend radius 0.09 m) with an inner diameter of

298 4.55 10 -3 m. The length of the reactor can be varied to obtain residence times of the oil and catalyst Density (288 K), kg m 3 850 based on outlet conditions between 0.15 and 5 s. The Conradson carbon, wt% 0.09 catalyst was transported by nitrogen from a feeding Basic nitrogen, ppm 12 device via a preheat oven to the reactor. At the oil Sulfur, wt% 0.01 injection point preheated feedstock was injected Vanadium, ppm perpendicular to the catalyst flow with a feeding rate Nickel, ppm of 0.067 g s l. Downstream of the reactor the catalyst Boiling point composition: was separated from the gas by a cyclone and stored Initial boiling point, K 607 in a hopper for post-run stripping. The gas was stepFinal boiling point, K 823 wise condensed and collected for analyses. The > 643 K, wt% 92.8 composition of the gaseous product was determined by gas chromatography; simulated distillation was Table 3. Reaction conditions used to determine the boiling point distribution of T(catalyst preheat), [K] 948 the liquid product, and the amount of carbon on the T(reactor), [K] 798 catalyst was determined using a LECO carbon T(stripper), [K] 798 analyzer. An overview of the analytical procedures is p(reactor), [ 105 Pa] 1.1 given in [5]. All experiments had mass balances of x(outlet conditions), [s] 0.15-5 98_+2 wt%. The catalyst-to-oil ratio has been varied CTO, [k~cat k~feed"1] 0-6 between 2 and 6 by adjusting the catalyst feed rate. To check the amount of thermal cracking, experiments have been performed without catalyst. The reaction conditions used are given in table 3. After each experiment the catalyst was regenerated in an external regenerator. Throughout the whole series of experiments the activity of the regenerated catalyst was the same. Table 2. Feedstock properties

The cracking experiments have been modeled according to the five lump model with different deactivation orders. The parameters of the model have been calculated using a Levenberg-Marquard minimization routine written in Fortran. Discrimination of models was based on: 1) the residuals, the difference between calculated and measured fractions of products, had to be independent of residence time; 2) the 95% reliability interval of the parameters; 3) minimization of the sum of squares residual, SSres.

4. RESULTS AND DISCUSSION

4.1. Experiments Figure 3 shows the yields of gas, gasoline, and LCO as a function of conversion, defined as the amount of HCO range hydrocarbons converted:

Conversion =

Yl , f e e d s t o c k

--

Yl , f e e d s t o c k

Y1

(8)

299

100 75gasolinea~A

"t:l

5 0

- -

~ v,.,.~

25--

~~o

0

I

Leo

I

I

25 50 75 100 Conversion [wt%]

0

All individual experiments that were performed are shown is this figure. It is remarkable that all measured values, obtained either at a high residence time with a low CTO, or at a low residence with a high CTO, can be represented by one simple function. The yield of gasoline increases linearly with conversion up to conversions higher than 90 wt%. At such high conversions in an industrial riser reactor, gasoline over-cracking is observed. The yield of LCO increases to maximal 17 wt%, and then decreases. It should be noted that the feedstock already contained 7.2 wt% hydrocarbons in the LCO boiling point range. The gas yield smoothly

Figure 3. Product yields as function of increases with conversion. The conversion and the coke yield as a conversion; tZT..'gas, A:gasoline, o:LCO. function of residence time with different CTOs are shown in figure 4a and 4b, respectively. The conversion in the absence of catalysts is a measure for the amount of thermal cracking. At a residence time of 4 s it is at maximum 6 wt%. The catalytic conversion after 0.15 s is remarkably high, over 30 wt%. The conversion increases proportionally with residence time up to a conversion of ca. 75 wt%, then levels off and reaches a maximum value of 90-95 wt% after 3-5 s. The coke yield (figure 4b) is hardly influenced by residence time, except for a minor increase at a CTO of 6. Clearly, almost all the coke has been generated within 0.15 s.

T =

O

100 t

80

~

"

-

-

O

~

1.5-

60

,0"

''!

"

a ~

.,,-q

-

20

r..)

0 a

O_O " O

40 g

2

--~o

1 - ~I [] . . . . . . .

i~1- -[]

0.5-

I- , - - x 0

2.5 5 Residence time [s]

b

0

2.5 5 Residence time [s]

Figure 4. Conversion (a) and coke yield (b) as a function of residence time and catalyst-to-oil ratio; l~.'CT0=2, A:CTO=4, o:CTO=6, x: no catalyst. The lines were drawn to guide the eye.

4.2. Modeling The data at short residence times (<0.15 s) show a quite different behavior from those obtained at long residence times (>0.15 s). It was decided to model the regions separately. Firstly, the part of the curves at residence times longer than 0.15 s is considered. The data measured at 0.15 s were used to define an imaginary feedstock composition. The almost

300 Table4. Modeling results of the data obtained between 0.15 s and 5 s. Reactions according to eq. 2-7

constant coke yields indicate that, after initial deactivation, the catalyst has a constant rest activity. no deactivation with deactivation (m=l) Additional coke formation is only k~ 8.610 "2"4" 4.510 -2 7.710 "2 "4" 4.310 -2 significant at high CTO. Two k2 1.810 -1_ 0.710 -1 1.610 ~ _+0.710 -l different models have been used: k3 5.810 2 + 6 . 6 1 0 -2 5.310 - 2 _ 7 . 1 1 0 -2 first a model without deactivation k4 0.0 -+ 2.4 10-I 0.0 -+ 2.9 10-I of the catalyst, ~(x)=l; the second ks 2.1 10-1 _ 2.3 10-l 1.8 10-1 -+ 2.8 101 model with a first order k6 6.010-8+ 1.110 -2 2.8 10 -9 _+ 1.4 10 -2 deactivation, ~(x)=exp(-lq 1:). k7 0.0 + 3.110 -2 0.0 4" 3.210 "2 The resulting reaction rate ks 7 . 7 1 0 4 _+ 6.610 -2 7.810 -4 _+7.110 -2 constants with their 95% reliability kd 2.510 -13 _+4.210 .5 interval are given in table 4. Based SSres O. 10 O. 11 on the criteria for discrimination that have been mentioned above, it is not possible to discriminate between the two models. The values of the calculated reaction rate constants are almost the same, and the deactivation constant k~ is negligibly small. From the data in table 4 it can be concluded that the deactivation of the catalyst with this hydrowax feedstock after 0.15 s is negligible. Because the model without deactivation describes the data with the same accuracy and less parameters, this model is selected for the optimal description of the data. Secondly, it was tried to model the complete data set with one deactivation function. A first order deactivation and a much higher order deactivation (m=6) were used to account for the initial deactivation and constant rest activity. The resulting reaction rate constants are given in table 5. Based on the relatively large inaccuracy of the reaction rate constants calculated with the highest deactivation order, this Table 5. Modeling results of the complete data set model was rejected. It can be with different deactivation orders m (eq. 2-7) concluded that the measured data are best described with a first order m= 1 m=6 deactivation. The reaction rate 1.5 10 -1_+0.510 -I 2.3 _+9.2 kl constants with the highest reliability 3.610 1 _ 0 . 4 1 0 -I 5.9 _+23 k2 are for the reactions of HCO to 6.4 10-2 _3.1 10-2 1.1 _+4.4 k3 LCO (kl), and gasoline (k2). The 3.5 10 -2_+8.010 .2 6.610 5-+3.5 k4 reaction rate constants that describe 2.210 -1-+1.210 l 3.8 _+16 k5 the formation of gas: k3, k4, and ks, 1.8 10-12___ 4.7 10 4 5.1 10 .2 _+ 36 10 .2 k6 have a relatively high inaccuracy. 1.2 102 _ + 2 . 4 1 0 .2 2.1 10-1_+8.210 -1 k7 This may be caused by the fact that 6.0 10-6 _+2.8 10-2 1.1 10-12_+2.6 10-4 k8 in the gas lump the catalytic product 1.3 10-1 _+0.4 10-1 18 _+75 kd LPG and the thermal product fuel 0.22 0.15 SSres gas are lumped together. Separation of the gas lump into two different lumps should improve the accuracy of the model. The parameter for gasoline over-cracking k6 is extremely low. This is not surprising, since the experimental data on which the model is based, showed a linear relation between conversion and gasoline yield.

301 A comparison between the models developed for x>0.15s (without deactivation) and for 80 x=0-5 s (first order deactivation) is given in figure 5. When only the experiments for x>0.15 s 60 ,, are taken into account, the data are reasonably "~ 40 described without deactivation. However, to =o 2o describe all data, a first order deactivation is necessary. Conversion has reached a maximum 0 value after about 3 s, so only the first part of the 0 2.5 5 reactor is effectively used. This is illustrated in residence time [s] figure 6, where HCO and gasoline yields versus Figure 5. Conversion as a function of reactor length are given. The data of figure 6d, residence time (CTO=4), experimental (CTO=10) have been calculated with the model data: A, modeling results; (--): >0.15 s, developed for x>0.15 s. From figure 6 it is clear ( ~ ) : 0-5 s. that only the first 5-10 m of the reactor are effectively used for conversion. 100

CTO=2

100

CTO=4

100 .-.'" 80 ~

80 6o

~

40

40

~=

2o

20

lh

0

0 0 a

0

5 10 15 20 25 Reactor length [m]

b

CTO=6

100

100 l ,--, 80

5 10 15 20 25 Reactor length [m] CTO=10, extrapolated

80

60

&

6o

40

~-

40

20

~

20

0 0 c

5 10 15 20 25 Reactor length [m]

0 d

5 10 15 20 25 Reactor length [m]

Figure 6. HCO (17) and gasoline (A) yields along the reactor length at different CTO. The data of figure a-c have been measured, whereas data in figure d have been extrapolated using the model without initial effects and with a constant catalytic activity. (symbols in figure d do not refer to measured data.

302 It is clear that modeling of the coke deposition is essential for a satisfactory description of FCC unit operation. It is not useful to describe this phenomena based on a simple power rate law. A more sophisticated model has to be derived in which the initial effects are accounted for. It is shown that coke deposition and catalyst activity have to be divided in an initial process (typically within 0.15 s) and a process at larger residence times. A simplified lumped kinetic model can be adequately used for this purpose, but a realistic coke formation model has to be developed.

5. CONCLUSIONS The conversion of a hydrowax feedstock due to thermal cracking at a residence time of 4 s was at maximum 6 wt%. The product yields obtained catalytically could all be described by the same function of conversion, regardless the reactor length or the CTO used. Up to a conversion of 95 wt% no gasoline over-cracking has been observed. Initial effects play an important role in FCC, the major part of the coke formation and almost 45% of the total conversion, have taken place within 0.15 s. At a residence time >0.15 s hardly any additional coke formation was observed. Translation of these results to effective use of reactor length showed that only the first 5-10 m of the reactor were used effectively for conversion. For the models evaluated in this work, the best model to describe all experiments was the five lump model with a first order deactivation, although it did not describe the first part of the reactor correctly, obviously due to an incorrect description of the initial effects. When the initial effects were excluded, a model with a constant activity described the data satisfactory. Therefore, coke deposition and catalyst deactivation have to be divided in an initial process (<0.15 s) and a process on a longer time scale.

Acknowledgment The authors are grateful to ir. A. Plesman and ir. J.M.A. van der Kamp for their contribution to the experimental work and the modeling, respectively. Dr. J.M.H. Dirkx of Shell International Oil Products is acknowledged for the stimulating discussions, Shell International Oil Products is acknowledged for the supply of feedstock and catalyst and for the financial support.

REFERENCES 1 M.P. Helmsing, M. Makkee, J.A. Moulijn, Chem. Engng. Sci., 51 (11) (1996) 3039. 2 J. Corella, E. Franc6s, Am. Chem. Soc., Symp. Series, 452 (1991) 144. 3 E. Vynckier, G.F. Froment, in G. Astarita, S.I. Sandler (Eds.), Kinetic and thermodynamic lumping of multicomponent mixtures, Elsevier, Amsterdam, 1991, 13 i. 4 F. Kapteijn, J.A. Moulijn, in G. Ertl, H. Kn6zinger, W. Weitkamp (Eds.), Handbook of heterogeneous catalysis, VCH, Weinheim, 1996, chapter 6.1. 5 M.P. Helmsing, PhD Thesis, Delft University of Technology, 1996. 6 T. Takatsuka, S. Sato, Y. Morimoto, H. Hashimoto, Int. Chem. Eng., 27 (1) (1987) 107.

9

Elsevier Science B. V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

303

M a t h e m a t i c a l M o d e l i n g o f Deactivation b y Coke F o r m a t i o n in the C r a c k i n g o f Gasoil H.S. Cerqueira a, E.C. Biscaia Jr. a and E.Falabella S.-Aguiarb'c'* PEQ/COPPE/UFRJ - Ilha do Fund~o, CT. Bloco G, CP. 68502; fax (55-21) 290-6626, CEP 21945-970 Rio de Janeiro, Brasil. E-mail: [email protected]. a

b CENPES/PETROBRAS - Divisao de Catalisadores, Ilha do Fundao, Quadra 7, fax (55-21) 598-6626, CEP 21949-900 Rio de Janeiro, Brasil. E-mail: [email protected] c EQ/UFRJ - Ilha do Fundao, CT. Bloco E, CP. 68542; Fax (55-21) 590-4991, CEP 21949900 Rio de Janeiro, Brasil. In the present article the profile of coke formation for different gasoil (aromatic, paraffinic, naphthenic and equalized) feedstocks in a fixed bed traditional MAT reactor is studied. In order to achieve this goal, a 12-lump model was used. In this model the feedstock is divided into 4 different lumps: paraffins, naphthenes, aromatic tings and aromatic substituent groups. Three different deactivation functions were tested. Results have shown that a one parameter exponential deactivation function is satisfactory. The mathematical solution was performed using a differential-algebraic approach. Two different initial conditions were tested. Although higher coke formation in the begining of the bed is encountered for all feedstocks, the aromatic feedstock is the one that yields more coke. Also, the injection time was varied, showing that the higher the catalyst-to-oil ratio, the higher the influence of the injection time on the coke yield.

1. INTRODUCTION Fluid Catalytic Cracking (FCC) is one of the most important process in oil refining. The evaluation of the catalysts in the laboratory scale is often carried out in a micro-reactor, the so called micro-activity test [1-3] (MAT). Coke formation plays an important role in the deactivation of FCC catalysts, which can be deactivated either permanently (loss of surface area, zeolite collapse, metals) or temporarily deactivated (coke). The approach proposed by FROMENT and BISCHOFF [4,5], that considers activity decay as a function of the coke on catalyst was used.

Corresponding author

304 2. MATHEMATICAL MODEL The kinetic model employed is a 12 lump model [6], based on the 10 lump model proposed by WEEKMAN et al. [7-9]. The reactional scheme could be seen on figure 1. A model for the fixed bed integral reactor was used, without feed gradients or temperature gradients inside the catalyst particle, at the interface particle/fluid and in the radial direction. The variation of temperature along the bed is low. Both axial dispersion of mass and axial gradients of temperature are disregarded. Considering that the reactor operates in a kinetic regime and chanelling is absent, the assumption of no mass axial dispersion is reasonable. The adsorption of reacting species is not considered because the controlling step in the paraffin cracking reaction is the formation of an intermediate carbocation. PH

NH

GAS2~ ' f

AH

CAH

COKE

Figure 1. Reactional scheme for the 12 lump model. LCO is given by the sum Of PL, NL, AL and CAL, as well as gasoil is the sum OfPH, NH, AN and CAn. With the previous assumptions the continuity equation for each componente may be written.

~

I 'ail IP'ai'Vzl +

0t

0z

=r.

1

(1)

Assuming uniform transversal area, uniform porosity and plug flow. ~a.

~ + G . 0t

~

1_

r.1

G=p-v z Assuming quasi-steady-state, the above equation may be simplified:

(2)

(3)

305

~a.

G'Oz !-r. 1

(4) Where,

G

W H S V . p~ 9L

(5)

~

The bed lenght in a dimensionless form looks like. z

(6)

Z=m

L The average molecular weight of the gaseous mixture is: 1 rlVlm

=

(7)

n-1

~-'~ ak k=l

The reaction rate may be defined as follows.

(8)

r i - ([:)(Co) 9s (Kij 9aj) 9P--e-~9P j=l

[~

Where: --(k,+~ +k~) o 0 o

.. "--

0

0

0

-(k. +k,. +k,) 0 0 0 -(k, +k, +k~ +k,o) 0 o o -(k,, + ~ )

0

0

o

o

o

0

o0-

0 0 o

0 o o

o

o

o

o

oo

o

o

o

o

o o

o

o

o

o

o o

v,.k,

o

o

o

-0~+k,)

0

o

o

o

o

oo

0 o

v,.k, o

0 v,.k,

0 o

0 o

-0q,+k,,) o

o

o

o

o

o o

-Oh,+k,,)

o

o

o

o o

0

0

V,o-lqo

v,, .lq,

0

0

o

-k,

o

o

o o

Z, .a~,

o

--(lOgo -I-kl1-1-k22 )

0

0 0

",o'k~ ~,.k~, ~.k~

v,, .lq,

o

0

0

0

0

0

0

0

o

0

0

0

v12-lq2

0

0

o

v,~.k~

+(Cc)

= deactivation

-(k~+k.) 0 0 ~.k~ 0 0 ~,.k~ 0 0

function

12

(i)(C~)" ~ (K~j-aj) 6~a i

"'OZ

P

WHSV-R. T

j=l 11

~ai i=l

(9)

306 The coke content on catalyst for a given axial position in the bed up to time tc is given by: C - a~3 (Z, t ) = WHSV. c

i

t~ 0 a 1 2 o

0Z

drl

(10)

Differentiating the expression above with respect to time, one obtains one more differencial equation (in time), that together with the continuity equations (differentials in Z), forms a system of partial-differential equations. C3a13 P 0 a12 . . . . . 0t~ R.T 0Z

(11)

The mathematical solution was performed by dividing the reactor into finite elements and applying a second order approximation in each element. This approximation is obtained using the point located at a distance corresponding to one third of the distance between two subsequent elements, Figure 2 illustrates the procedure. x

0

1/3

1

...1__1 Zi-1

Zi-2/3

...

Zi

Figure 2. One element of the reactor discretized. With previously published kinetic constants [10,11], presented in Table 1, the model was solved for two different initial conditions. The former one (type I) presumes the existence of a reactants and products profile at t=0, whereas the latter (type II) considers that the reactor is empty at t=0. The resultant differential-algebraic system of equations was solved by backward finite differences formula with variable step, implemented in the DASSL code [ 12,13]. The numerical convergence was assured by increasing the number of finite elements until no further modification in the model simulations was obtained. Hence, the number of elements was gradually increased from 20 to 100. Nevertheless, no significant diference was observed, i. e., all product yields were obtained with errors smaller than 10-5.

3. DEACTIVATION FUNCTION In order to solve the model, one must choose an appropriate deactivation function. Three different deactivation functions were tested [ 14-16]. B+I qkA = B + e x p ( A . C c )

B - exp(-ct.

Cc)

(12) (13)

307

{bc =

1

(14)

I+K.Cc Where: Cc = coke content on catalyst (wt %), A, B, K and a are constants to be fit.

With those functions, assuming the same deactivation for all reactions and an injection time of 90 s, several simulations using the 12 lump model were made. It allowed one to obtain values for the constants A, B, K and a that produce similar results regardless of the deactivation function used in the model. Figs. 3 and 4 show gasoline and coke yields. Analysing Figs. 3 and 4, one may note that quite similar results are obtained.Hereinafter, the one parameter exponential deactivation function will be adopted, employing different values for the parameter or, for heavy lumps cracking, LCO cracking and other cracking reactions (Table 2) [6]. Table 1 Kinetic constants used in the 12 lump model,. Temperature = 755 K. [(~ cat/cm3)lh l] Values Values kl = kphPl 8578 k9 = kAhgasl 34200 k17 = kAl~ k2 = kpm 17540 kl0 = kAhcAl 50000 k18 = kAlgasl k3 = kPhgasl 0 kll = kCAhCAI 5860 k19 --kCAlcoke k4 = k~Nl 22500 kl2 = kcAhcoko 14639 k20 = kGg,,l k5 = kNhgasl 14870 k13 = kpIG 8307 k21 = kGgas2 k6 = k~a 84570 k14 = kplgasl 7096 k22 = kGcoke k7 = kAhAl 19000 k15 = kNIG 66150 k23 = kgaslgas2 k8 = kAm 63000 k16 = kNlgasl 8180 k24 -- k~aslcoke 50 - -

Function

e~.xp(-150"c~)

Deactivation

~b=1/(1+13(X)*Cc)

4~==(22.3+1y(22.3+exp(218*ce)) ._=

1000

2340 0 0 2204 4880

10 - -

Deactivation

Q

Values 1850 3630

o-~

+

6--

9

30-

Function

~a=exp(-150*Cc) ~ b=1/(1+1300"Cc) { ~(22.3+1)/(22.3+exp(218*Cc))

Q

m o i/1 t~

0 0

' 0

0

I 10

4--

20

v~-tsv

30

40

50

Figure 3. Gasoline yield versus WHSV for different deactivation functions, tc=90 s.

' 0

I 10

'

I 20

'

v~sv

I 30

'

[ 40

Figure4. Coke yield versus WHSV for different deactivation functions, tc=90 s.

308 Table 2 Values for the exponential deactiyat!on function parameter (dimensionless) ~] 471.9 ~2 624.0 (~3 477.1

4. RESULTS AND DISCUSSION Four different feedstocks were used in the simulations. The first feed has the same paraffinic, naphthenic and aromatic contents, the others are richer (50% wt) in one of them. Figures. 5 and 6 display the coke profiles for the different feedstocks with an injection time of 90 s and WHSV=10. Coke yields for the equalized feed and different running times are presented in Figures 7 and 8. Results indicate that regardless of the feedstock used, the total coke yield is approximately the same for all WHSVs. With respect to the coke profile, however, differences concerning different feedstocks are quite clear (Figs. 5 and 6). The aromatic feedstock yields more coke than the other feedstocks along the entire bed length. Higher coke formation in the begining of the bed is encountered for all feedstocks. Also, the injection time was varied, showing that the higher the catalyst-to-oil ratio, the higher the influence of the injection time on the coke yield (Figs. 7 and 8).

1.30

Feedstock

Feedstock

--q.- ( ._ )

--..4.--(---) A

1.20

-~

~

A

paraffinic

='=

naphthenic

paraffinic naphthenic

m

A

aromatic

aromatic ,~,

"~m 1.10 0

1.10

1.00

g

g 1.00

0.90 ~1

0.90 0.00

0.20

0.40

0.60

0.80

Z

Figure 5. Coke profile for different feedstocks, initial condition type I.

1.00

0.80 ] 0.00

I 0.20

'

I

'

0.40

I 0.60

'

I

0.80

'

I 1.00

Z

different profile Figure 6. Coke for feedstocks, initial condition type II.

309 10

8

6

-

-

=-~

_

-

.-~

Injection Time

4--

)(

t=60~

1|

r

t=90S

/

.,-,., =~

Injection Time 6

=

== o

O O

)(

t=60s

1111

r

t=90s

i |

"~

=

~

=

4

2-2

_

0

' 0

I

'

10

I

'

20

I 30

'

I

'

40

~ V

Figure 7. Coke yield for different injection times, initial condition type I.

I 50

o

~

0

I 10

~

I 20

~

~ V

I 30

~

1

~

40

Figure 8. Coke yield for different injection times, initial condition type II.

5. C O N C L U S I O N S

The mathematical model of a MAT reactor, considering a 12 lump model, has been discretized using a finite element method in the direction of gas flow. The resulting system of differential-algebraic equations (DAEs) has been solved by an appropriate computer code (DASSL). The non existence of oscillation in the proposed discretization allied with the robustness of the computer code indicate that the association of these techniques seems to be adequate to solve the mathematical model of gasoil cracking. Other discretization methods, such as global polynomial approximation [ 17], could bring about oscillatory profiles through the bed length. Albeit several deactivation functions have been tried, the one parameter exponential function provides results which are very similar to those obtained with functions having more parameters. Therefore, this simpler function seems to be a convenient approach to the solution of the rather complicated problem of catalyst deactivation in gasoil cracking. Nomenclature al = tools of PH lump / g gas, a2 = tools of NH lump / g gas, a3 = tools Of AH lump / g gas, a4 = mols of CAH lump / g gas, a5 = tools of PL lump / g gas, a6 = tools of NL lump / g gas, a: = mols of AL lump / g gas, a8 = mols of CAL lump / g gas, a9 = tools of gasoline lump / g gas, alo = mols of gasl lump / g gas, all = mols of gas2 lump / g gas, a12 = g coke / g gas, a13 = coke content on catalyst,

cat/cm3)lh-1],

ki = kinetic constant [(g where, lbr the 12 lump model: kl = kl,hPl, 1(2 = kph~, k3 = kphgasl, k4 = kNhN1, k5 = kNhg~l, k6 = kNhG, k7 = kAltal, k8 = kAhG, k9 = kAhgasl, klo = k~c~a, kll = kc~c~a, kl2 = kc~coke, k13 = kplG, k14 = kplg~l, kl5 = kNlo, k16 = kNlgasl, k17 = kalG, kl8 = kAlgasl, k19 = kcAlcoke, k20 = kc,g~l, k:l = k c ~ 2 , k22 -- kG~oke, 1(23 = kg~slgas2, k24 = kg~slcoke, Ci = pi a = molar concentration of the i component, Cc = coke on catalyst [g coke / g cat], CTO = catalyst-to-oil ratio [g cat / g oil], G = mass velocity,

I 50

310

ki = velocity constant for the reaction i ~ j [cm3/(g cath)], L = bed lenght, n = number of lumps, PMm = average molecular weight, symbols of the 12 lump model [wt %]: PH = paraffm molecules in heavy fuel oil (>342 ~ NH = naphthene molecules in heavy fuel oil (>342 ~ AH = aromatic substituent groups in heavy fuel oil (>342 ~ C~a~= aromatic rings in heavy fuel oil (>342 ~ PL = paraffin molecules in light fuel oil (216-342 ~ NL -- naphthene molecules in light fuel oil (216-342 ~ AL = aromatic substituent groups in light fuel oil (216-342 ~ CAL= aromatic rings in light fuel oil (216-343 ~ G = gasoline (C5 - 212 ~ gasl = primary gaseous products, gas2 = secondary gaseous products, coke = coke. Tc, t~ = running time = injection time, Vz= axial velocity, WHSV = weight hourly space velocity [g feed/(g cath)].

Greek Symbols c~= exponential deactivation function coeficient, = porosity of bed,

~A, ~B, ~bc=deactivation functions, v~, v4, v7, v~0,v~ = ratio between the molecular weigths of the heavy fuel oil and the light fuel oil, v2, v6, vs = ratio between the molecular weigths of the heavy fuel oil and gasoline lumps, v3, vs, v9 = ratio between the molecular weigths of the heavy fuel oil and gasl lumps, vn =molecular weight of the heavy fuel oil lumps, Vl3, v~5,v~7 = ratio between the molecular weigths of the light fuel oil and gasoline lumps, v~4, v~6,v~s = ratio between the molecular weigths of the light fuel oil and gasl lumps, vl9 = molecular weigth of the light fuel oil, v20 = ratio between the molecular weigths of gasoline and gas l lumps, v21 -- ratio between the molecular weigths of gasoline and gas2 lumps, v22 = molecular weight of gasoline, v23 = ratio between the molecular weigths of the gasl and gas2 lumps, v24 = molecular weight of gasl lump, p = gas density [g/cma], pc = bed density [g cat/cm3].

REFERENCES 1. 2. 3. 4. 5. 6.

A S T M - D 3907, section 5 (1989) 733. G.D.L.Carter and G.McElhiney, Hydrocarbon Processing, september (1989) 63. J.B.Cropley, Chem. Engng. Progress, february (1990) 32. G.Froment and K.B.Bischoff, Chem. Eng. Sci., 16, 3 (1961) 189. G.Froment and K.B.Bischoff, Chem. Eng. Sci., 17, 4, (1962) 105. L.L.Oliveira, "Estimag~o de Par~maetros e A v a l i a ~ o de Modelos de Craqueamento Catalitico, M.Sc. Thesis; COPPE/UFRJ, Rio de Janeiro, RJ, Brazil, march (1987). 7. S.M.Jacob, B.Gross, S.E.Voltz and V.W.Weekman Jr., AIChE Journal, 22, 4 (1976) 701. 8. V.W.Weekman Jr., Lumps, Model and Kinetic in Practice, AIChE monograph series (1979). 9. B.Gross, S.M.Jacob, D.M.Nace and S.E. Voltz, US Patent 4 187 548 (1980). 10. L.L.Oliveira and E.C. Biscaia Jr., I. & E.C. Research, 28, No. 3 (1989) 264. 11. V.W.Weekman Jr. and D.M. Nace, AIChE Journal, 16, 3 (1970) 397. 12. K.E.Brenan, S.L. Campbell and L.R. Petzold, Numerical Solution of Initial Value Problems in Differential-Algebraic Equations, Elsevier Sc. Publishing Co, lnc, New York, (1989). 13. L.R.Petzold, DASSL Code, version 1989, Computing and Mathematics Research Division, Lawrence Livermore National Laboratory, L316, PO Box 808, Livermore, California, (1989). 14. M..Forissier and J.R.Bernard, Modeling the 'Micro-activity Test of FCC Catalysts to Compute Kinetic Parameters, AIChE Meeting, Houston, USA, april (1989). 15. M..Forissier and J.R.Bemard, Deactivation of Cracking Catalysts with Vacuum Gas Oil; Catalyst Deactivation, C.H. Bartholomew and J.B. Butt editores, Elsevier, Amsterdam (1991) 35 9. 16. J.Pitault, D.Nevicato, M.Forissier and J.Bernard, Chem. Eng. Sci., 49, 24a, (1994) 4249. 17. J.Viladsen and M.L. Michelsen, Solution of Differential Equation Models by Polynomial Approximation, Prentice Hall (1978).

9

Elsevier Science B.V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

311

Catalyst Decay by Simultaneous Sintering and Poisoning: Effects of Intraparticle and Interfacial Gradients M. Chocron a'b, N. Amadeo a, M. Laborde a

a Departamento de Ingenieria Quimica, Universidad de Buenos Aires, Pabell6n de Industrias, Ciudad Universitaria, (1428) Buenos Aires, Argentina. e-mail: [email protected] b U.A.Q., Comisi6n Nacional de Energia At6mica, Buenos Aires, Argentina.

The influence of intemal and interfacial diffusion on catalyst deactivation by simultaneous sintering and poisoning is examined. The study focuses on the copper catalyst used in the water gas shill reaction (WGSR). It is found that catalyst life increases when internal and extemal poison diffusional resistance increases. Temperature reduces the total average activity but this effect is partially neutralized by the diffusional effects undergone by the reactants inside the pellet.

1. INTRODUCTION Copper catalysts used in methanol synthesis and in the water gas shill reaction (WGSR) are irreversibly poisoned by small quantities of chlorine (in the feed) and by temperature (sintering). The effect of internal diffusion on catalyst decay by poisoning has been discussed in a previous paper [ 1]. The aim of this work is to analyze the effect of internal and interfacial diffusion on a catalyst pellet in the presence of an activity decay by simultaneous poisoning and sintering. The study has been applied to a copper based catalyst used in the WGSR: CO + 1-120 = COz + H2, for which a Langmuir-Hinshelwood type kinetics has been considered [2].

2. M A T H E M A T I C A L M O D E L For an isothermal, cylindrical catalyst particle with constant properties, the mass conservation equation for the reactant (CO) is:

De.co 9l d ( d Cr c ~ 1 - rcoa p (r.t)ar(r.t) rdr dr

(1)

with the following boundary conditions: at r = 0, dCco/dr- 0; at r = R, -Do~co(dCco/dr)s = kgco(CSco-Cbco). The kinetic expression for the main reaction is given by [2] :

312

kCco Cu2o (1- fl)

~co (l+XcoCco +I%oC,,~o +K~o~C~o~+x,,~C,,~)~

(2)

where 13 = Cc02 CH2 / ( Cco CH20 Keq ) for which the parameter values have been given elsewhere

[1].

The concentration of the i~ species is related to the CO concentration by [3] :

C~-[C, ~ k c~ k giaco

a~De~o

+ acoDe~ (Cco-Cco

(3)

The Wheeler equation [4] was used in order to evaluate the effective diffusion coefficients of each component and the poison inside the pellet [ 1]. 1

1

- - - - - + ~

1

Dey~D~DK~

(4)

Mass transfer coefficients in the film were estimated from the correlation given by Froment and Bischoffas a function of a modified Reynolds number, Re' = dp G / gg [5]:

k gipgSc 2/3 Jv -

a

(5)

For the multicomponent gas mixture, the effective binary diffusivity for each component that diffuses through the mixture ( D ~ ) was calculated following the treatment given by Taylor et al. [6], as well as the average physical properties of the gaseous phase [7]. Thermal deactivation of a copper catalyst has been experimentally studied [8] and the results were expressed by the generalized power law equation [9]:

da r (r, t) = _koM(r ) exp(- Ed / Rg T)(a r (r, t ) - ar ~176 )2 dt

(6)

where M(r) =(Cco + CUE) / (Cco2 + CUE)has been used to represent the feed ratio. For instance, if M > 1 the reaction atmosphere is reductive. Catalyst deactivation by poisoning (chlorine) was modeled by the following set of equations:

DefP r --~r r ---~r J = r , a , ( r , t )

(7)

with the following boundary conditions: at r = 0, dCp/dr = 0; at r = R, -D~(dCp/dr)s = kgp(C~-C~).

(8)

313

da p (r , t) ~=-k2(1-Xp)ap(r,t dt

)

(9)

Kinetic coefficients of eqns. (8) and (9) were evaluated from industrial reactor data [10]. The effectiveness factor of CO defined by Chocr6n et al. [ 1] has been u s e d v

IrcoapardV 0

(10)

r co apar ~vV where: v

f ap(r,,)a,(,,OdV (a'a')ov--~

(11)

V

3. RESULTS

The effect of the reaction temperature, poisoning and sintering rates as well as the Reynolds number on the effectiveness factor evolution have been investigated. In Figure 1, activity profiles affected by poisoning and activity profiles affected by both sintering and poisoning for three different kl values, T = 503 K and two months of operation are shown. It can be seen that the profiles are sharper when kl increases, consequently the average activity, as defined by eqn. (11), remains high. When both deactivation effects are superimposed, the

1.0

t

~.__

0.9

~ ~ J

O

K 1 =0.2

9

0.8

--

0.7

K1=2000

9

K1=8000

rn o l / I h m olllh rn o I I I h

Poisoning

P o i $ o n i n g + S in

te ring

0.6 0.5 0.4

l

0.3 0.2

L

. . . . . . . .

0.1 0.0

9

9-I .

.........

. . . . . . . .

. . . . . . .

_

9. . . . . . . .

1

0.1

9. . . . . . . . l

0.2

9. . . . . . . . i

0.3

9. . . . . . . . "J . . . . . . . . . J..-'.'.'.'.'.D'-'.'.'-'.'.'.'.lt .......... :.:.~.ll ................ iill i

0.4

i

0.5

I

0.6

,

i

0.7

,

i

0.8

,

,

0.9

1

1.0

r/R

Figure 1. Activity profiles during poisoning and simultaneous sintering and poisoning.

314 activity for each kl value diminishes but the profile shape is maintained. This is due to the fact that the deactivation by sintering is almost uniform along the pellet radius ( the effect of M on the activity profile is negligible ). Figure 2 shows the effect of activity decay on CO profiles during poisoning and simultaneous sintering and poisoning, at the same operating time, temperature and kl values of Figure 1. It can be appreciated that the higher the kl value is, the higher the consumption of CO near the pellet surface is. This behavior can be explained by the fact that the average activity remains higher as it was pointed out above. When both deactivation effects are considered, the CO concentration for each k~ value increases but holds, as it was depicted for the activity, the same profile shape. The behavior observed in Figures 1 and 2 can be explained by the fact that kl increases along with poison diffusional resistance.

/ ........ -T . . . . . . . . 1-. . . . . . . ~

.

~, .........

0.7 ~ 8 oo 8

O

9.........

r

....

.

--

9 9 9

Poisoning

r- ........ T-~/ -----_"__'"_r__'~Poionin__rs . g+Sintedng

........

9

........

*

. ...........

..,"

0.6 !-

K1=0.2 mol/Ih K1=2000 mol/Ih K1=8000 mol/Ih

. .......

!

...-;,i-...),!

I

..... 9

0.5 0.4 0.3 0.2

0.1,i .................................

.

0.0 ........................ 0.1 0.2

0.3

0 i 4 .... 0 : 5 ................................. 0.6 0.7 0.8

0.9

1.0

dR

Figure 2. Effect of deactivation on the CO profiles in the pellet. Same conditions as in Fig. 1. Figure 3 shows the effect of kl on the effectiveness factor for T = 503 K as a function of time. The first conclusion is that the effectiveness factor as defined in equation (10) tends to 1 at longer times whatever kl or the deactivation mechanism is. This behavior was found in a previous paper [ 1] and current results confirm that the given effectiveness factor definition is also valid for the present situation. When the effect of kl is analyzed considering only the poisoning deactivation, the following considerations can be mentioned: (i) at a given time, when kl increases, the effectiveness factor decreases; (ii) along the time, when kl increases, the effectiveness factor rises more slowly and for the highest k~ the initial value can be lower [1 ]; ('tii) when both, sintering and poisoning phenomena are considered, at a given time and kl value, the effectiveness factor increases with respect to the previous case, showing an abrupt increment at short times. This is attributed to the

315

i

l

10~

..0"

7 "~~

65

LL

4

I

I

I

......

v

v-

.--"

.'"

.~r"

~

.V " ' " " "

"9

3

7

"

.v"

i

0

i

......... 9

I

1

h

2

! i rin

I

3

4

I

5

6

i

7

8

Time(months)

Figure 3. Variation of the effectiveness factor as a function of time with kl as a parameter fact that the deactivation rate by sintering, at shorter times, is higher than that by poisoning. ( see equations 6 and 7 ). Figure 4 shows the usual representation of effectiveness factor vs. Thiele modulus modified in the present case by the square root of average activity at T = 503 K . This figure presents the evolution of both parameters with time. The shape remains approximately similar to that of a nondeactivated pellet, at least for the smaller k~ values. For the highest kl selected, 11 initially decreases. This fact can be explained by considering eqn. (10) and making a simultaneous analysis of the activity and CO profiles. The effect of Reynolds number on the activity and effectiveness factor is shown in Figure 5 for k] = 8000 mole / 1-h and T = 503 K. In this figure the effectiveness factor reduction defined r-~ + 9r- . . . . . . . . 10 or

T-, -+---T ~- ~-.~-.,.-r....

N ............ ~

~

'--

'

--

L

'i

6

9 9 9

. . . . .

"-~"

K I = 0 . 2 m olllh K 1 = 2 0 0 0 m olllh K 1 = 8 0 0 0 m olllh

~

....... ~ "

~'e

"m."+ilR

Poisoning Poisoning+Sintering

.........

1 0 -1 I L l

1 0 .2

,

a

,

L

2

. . . . . . . . . . . . .

3

4

I

,

,

1 0 "1 Thiele

*

,

. . . . . . . . . . . . .

2

3

4

I

. . . . . . . . . . . . . . . . .

100

m odulus*(av*aT)

2

a

4

1

101

. . . . . .

2

112

Figure 4. Effectiveness factor as a function of the modified Thiele modulus. T=503 K.

316 as ATI = [ ( rl ~176 - 11 ) / 1"1~176 ] * 100 and the increment in the average total activity defined as A (ap aT )av = {[ ( ap aT )~v - ( ap aT )av~176 ] / ( ap aT )av~176 } * 100 have been represented. It can be seen that the activity increases when the Reynolds number decreases. This fact can be explained by the higher diffussional resistance undergone by the poison. Consequently, the effectiveness factor diminishes with a Reynolds number at any time. Nevertheless, as it was expected, the effectiveness factor tends to 1 at longer times whatever the Reynolds number is. c o

=

!

22

o

:=

"~

~

20

'--

#

18!

~e

O [] 'IV .

14,

.

.

.

.

.

.

.

Ro'=611 Re'=397 R o'= 1 9 8

Effectiveness factor Total activity

.

:""

= o= 12~-

== > 1 0 L

U

9

"-.

..

Q'..

..

u

m

-~ ............: 1 1 :ii: . . . . . . i_:::::_.=

: ......

2

~

0

1

2

3

1t:: . . . . . .

4

5

6

7

8

Tim e(m onths)

Figure 5. Effect of the Reynolds number upon activity and effectiveness factor as a function of time. Figure 6 shows the effect of temperature on the effectiveness factor for the highest deactivation rate (k~ = 8000 mole/l-h ) and in absence of film resistance. It was assumed that the effect of temperature on the poisoning reaction rate is negligible compared to that on the sintering deactivation rate. Although the total average activity decreases when temperature increases it can be appreciated that the effectiveness factor decreases because the diffusional control regime in the pellet becomes more pronounced. At short times and high temperatures rl increases faster with time because deactivation by sintering is more marked under those conditions.

1~176i

....... 9

9

.., 99

'I ..... O ..... yD " "

m ........ o"

2

HI 10

-1

9

999

4

i

v

..

9

...o

.rim"

,v"

.

.

." . ..-

..

9

----o.--

.

.

n.-. . . . ~'-.-.

..'

"

.

.r

.....

T9.

i

,

i

-:-il

9

.

.i . ...-

" ...m

m ...... :

9" . . v ..

..

. .

. .

"

...o-"

.ll

. .o

99

:ll:.i./:

.., 9

.

.

.

T=453 T=503

K K

T=553

K

.

0

Tim 9 ( m onths)

Figure 6. Variation of the effectiveness factor with time as a function of temperature.

317 4. CONCLUSIONS The effect of internal and interfacial diffusion on catalyst decay by both, poisoning and sintering, was analyzed. Catalyst life increases when internal diffusional resistance for the poison increases ( high kl values ). Catalyst life also increases when interfacial diffusional resistance increases ( low Reynolds numbers ). Although this effect is less pronounced it might be taken into account when industrial reactor design and its operation are studied. Temperature reduces the total average activity but this effect is partially neutralized by the diffusional effects into the pellet. It was confirmed that the generalized definition of the effectiveness factor given in a previous paper [ 1] is valid for any type of kinetics and multiple deactivation mechanisms.

ACKNOWLEDGMENTS Support of this work provided by the Universidad de Buenos Aires and the Comision Nacional de Energia At6mica ( Argentina ) is gratefully acknowledged. Notation

%: activity identified with poisoning deactivation aT: activity identified with sintering deactivation aT~176 steady state activity by sintering. b: exponent defined in eqn. (6) Ci: concentration ofi component dp: catalyst pellet diameter. Dra: Knudsen diffusion coefficient of i component DM~: effective binary diffusion coefficient ofi component Dora: effective diffusion coefficient ofi component Ea: activation energy [eq.(6)] G: superficial mass flow velocity k : kinetic coeff', of main reaction. I~q: equilibrium constant of main reaction Ki : adsorption constant ofi component kl: kinetic coefficient of poisoning rate, [eqn.(8)] k2:kinetic coefficient of deactivation rate, [eqn.(9)] k#: mass transfer coefficient ofi component ko : pre-exponential factor in eqn. (6) M: parameter defined in eqn (6) r: pellet coordinate rco: reaction rate of WGSR rp: poisoning reaction rate R: pellet radius Rg: gas constant Sc: Schmidt number

T: temperature t: time V: pellet volume Xp: conversion of the poison Subscripts and superscripts av: average value s surface concentration b, o: bulk concentration oo. value calculated in absence of film resistance Greek Letters. or: Stoichiometric coefficient of i component 13: equilibrium term in equation (2) 11: effectiveness factor (I)co : Thiele modulus of CO pg: gas mixture density ).tg:gas mixture viscosity.

318 REFERENCES

1. M.Chocr6n, Ma. del C. Raffo Calder6n, N.E. Amadeo and M. Laborde, Chem. Eng. Sci., 51 (1996) 683. N.Amadeo and M. Laborde, Int. J. Hydrogen Energy, 20 (1995) 949. E.Petersen, Chemical Reaction Analysis. Prentice Hall Inc., NJ, 1965. A.Wheeler, Catalysis, Vol. II, Chap.2 (Edited by P.H. Emmett). Reinhold Pub. Corp., NY., 1965. G.Froment and K. Bischoff, Chemical Reactor Analysis and Design, J. Wiley & Sons, NY, 1979. R.Taylor and R. Krishna, Multicomponent Mass Transfer, J. Wiley & Sons, NY, 1993. T.Salmi and J. W~im~, Chem. Eng. Sci., 15 (1991) 715. M.L.Aparicio, M.Laborde and N.E.Amadeo, Informaci6n Tecnol6gic~ 7 (1996) 29. G.A.Fuentes, Appl. Catal., 15 (1985) 33. 10. J.Gonz/dez Velasco, M. Guti6rrez Ortiz, J. Gonz/dez Marcos, N. Amadeo, M. Laborde, M. Paz, Chem. Eng. Sci., 47 (1992) 1495.

9

Elsevier Science B.V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

319

A M o d e l for Catalyst D e a c t i v a t i o n in Industrial Catalytic R e f o r m i n g L.M. Rodrfguez OtaP, T. Viveros Garciab, M. SS.nchez Rubio c bArea de Ingenieria Qufmica, Depto. Ingenieria de Procesos e Hidraulica, UAM-Iztapalapa. Av. Michoacan y la Purisima, Col. Vicentina, Iztapalapa, Mrxico D.F. 09340. apemex Refinaci6n, Refinerfa Salamanca, Unidad de Evaluaci6n y Planeaci6n, Salamanca Gto., Mexico. cPemex Refinaci6n, Refinerfa Salina Cruz, Unidad de Evaluaci6n y Planeaci6n, Salina Cruz Oax, Mexico.

The deactivation of industrial catalytic reforming by carbon formation was analyzed. The study was based on the performance of three different industrial Platforming units using Pt-Sn/A1203 and Pt-Re/A1203 catalysts to produce high octane gasoline. It was determined that the difference in the operating temperature and the theoretical temperature needed to obtain a defined RONC can be used as a measure of the degree of deactivation. A model was obtained from a statistical analysis of parameters such as operating temperature and pressure, PONA analysis, residence time, H2/HC ratio. The model is capable to predict the deactivation degree in terms of the operation parameters and properties of the feed.

1. INTRODUCTION Naphta catalytic reforming (CR) is one of the basic processes in the petroleum industry. Catalytic reforming is an important process for the production of aromatic hydrocarbons, as well as being employed to enhance octane number (RONC) in gasolines and an important hydrogen producer [ 1]. CR processes presently employ platinum-based catalysts, either monometallic, bimetallic or polymetallic, depending upon the type of process: semiregenerative or continuous catalytic regeneration (CCR) [2]. Platinum based catalysts are employed in semiregenerative and continuous regeneration processes. Pt-Re, Pt-Sn and Pt-Ir are the bimetallic catalysts most frequently used in the Platforming and Aromizing processes. These processes, Platforming and Aromizing, can be used in both modes semiregenerative and CCR, with both catalysts Pt-Re and Pt-Sn. Main reactions in CR processes are: dehydrogenation of cyclohexane and alkylcylohexanes, cyclization of alkanes, isomerization of n-parafines, alkylcyclopentanes and alkylaromatics, and hydrocracking. Secondary reactions are: the demethylation and cracking of cyclic compounds,

320 and alkylation-dealkylation of aromatics. The polymerization of hydrocarbons on the catalytic surface is beleived to be the source of carbon formation which in turn deactivate the catalyst. The main deactivating phenomenon of CR catalysts is carbon deposition, but the presence of metallic compounds, sulphur compounds and an inadecuate H20/C1 ratio produce additional deactivation. Carbon formation is a phenomenon influenced by the operating conditions, feed composition, the nature of the second metal (promoter) and the quality of the product to be obtained [3]. Since carbon deposition is the main deactivating source, CR catalysts are regenerated after reaching a certain level of carbon, normally monitored by the increase in the operating temperature to obtain a certain product quality. In this work we attempt to evaluate the deactivation of industrial catalytic reforming catalysts due to carbon formation. The work is based on a statistical analysis of the results of three different industrial sitesusing Pt-Re and Pt-Sn catalysts in either semiregenerative or continuous catalytic regeneration Platforming processes to produce high octane gasoline. The model considers the influence of the operation parameters (temperature, pressure, residence time, and hydrogen/hydrocarbon ratio) the nature of the feed, the quality of the product and the type of catalyst.

2. ANALYSIS The study was based on the evaluation of industrial catalytic reforming units of the Platforming process. These units operate with Pt-Re/A1203 and Pt-Sn/A1103 catalysts producing high octane gasoline. The analysis is performed on data obtained from 7 operating cycles with Pt-Sn CR catalysts and 5 cycles with Pt-Re CR catalysts. A cycle is known as the operation from startup with fresh or regenerated catalyst until the operation was considered to be working with deactivated catalyst. The operation proceeded then to regeneration. Typical conditions for semiregenerative catalytic reforming units are given below in table 1. In this table feed PONA analysis, H2/HC molar ratio and operation conditions are given. Table 1. Typical operation conditions and feed characteristics

Catalyst Catalyst weight, lbs Number of reactors Operating pressure, kg/cm 2 Operating temperature, ~ H2/HC molar ratio Feed flow, B lls/day Feed composition, vol% parafines naftenics aromatics

Pt-Sn 77 000 4 13 479 7.5 14 000 70 20 10

321

Plant performance is initially visualized by the results of PONA analysis. The composition of feed and reformate are given in figures 1 & 2. The data shown correspond to cycle No.5 of a PtSn catalyst, and were taken from startup to 80 Blls/lbcat which represents around 365 days of operation. It can be seen that the feed is rich in parafines, with medium concentration of naftenics and low in aromatics. The reforming process yields an increase in the aromatics concentration and a considerable decrease in naftenics. It is also clear that the feed concentration varies during the cycle due to variations in the feed quality.

,,,J' b,,we

1

Par~eiNas A N~TENICS Q AROMATICS

t

.-e

A P~,FINF.S A NAFTENICS AROMATICS

I

I

I

I

I

BLLS/I.,B CAT.

I

I

80

Fig. 1. PONA analysis of feed. Cycle No.5 Pt-Sn catalyst CR unit.

O~

i

i

I

I

I

I

I

80

Fig. 2. PONA analysis of reformate. Cycle No.5 Pt-Sn catalyst CR unit.

From the analysis of the initial data (t---~O), polynomial equations were obtained for the prediction of reformate yield, the temperature needed to obtain a certain octane number (RONC) and hydrogen production in terms of the feed composition, and operating conditions. For example, the theoretical temperature is given by the following equation obtained for Pt-Sn catalysts: T = 449.16 - 0.4007 (N + 3.5 A) + FLHSV + 10(0.01679 RONC + 0.3032) where 9 FLHSV = (log LHSV + 4.1819 x 10-4 ) / 0.0183247 and for the reformate yield"

322

C5+ = FC 1 + FC2(RONC - 80) + FC3(RONC - 80) 2 + (X1 + X2RONC)(9TME -126.67) + 5/9(FX 1 + FX2 + FX3) where: FX1 = X9 + RONC(X4 (5TME/9 - 228) 2 + X6 FX2 = (RONC X8 + X7) (N + 2A) FX3 = (2112.67 - H) ( - 0.7 4124 + (- 1.6227x 10 -4 RONC + log 1.046)( - 0.426383 RONC(0.9639)(N + 2A))) FC1, FC2, FC3 are constants that depend upon the PONA. TME is the ASTM 50% volume distillate temperature. X 1, X2 ..... X9 constants function of RONC. H : operating pressure. A: vol % of aromatics in the feed. N: vol % of naftenics in the feed. From these equations it is posible to predict the temperature required to obtain a certain octane number throughout a cycle. This is shown in Fig. 3 for cycle No. 5. This graph shows the behaviour of the operation temperature and the theoretical temperature for the cycle. 90

0

OPERATION TEMPERATURE 70

I rr~or~rlCaLr ~ r t r ~

t

j

80

,I

611

50-

40-

311-

h

20-

10-

0-

-10 0

BI.,I.,S/I.,B CAT.

80

Fig. 3. Theoretical and operation temperatures in cycle No.5. Pt-Sn catalyst.

0

1~

2b

3b

4b

A IL il 0

CY~,E No. CYCLENo. CYCLENo. CYCLENo.

0 II

CYCLENo. 6 CYCLENo. 7

;0

6"0

1 :2 3 5

7"0

80

BLL$/I.~ CAT.

Fig. 4. Pt-Sn catalyst deactivation measured as increase in temperature. 6 cycles.

Clearly as time on stream increases in order to maintain a specified RONC, the operation temperature should be increased due to the deactivation of the catalyst. The shape of the theoretical temperature plot also shows a normal industrial operation policy. To maintain a specified RONC the reactor temperature is increased, but if a certain level of temperature considered high enough is attained, the base RONC is allowed to decrease, so that the

323

theoretical temperature to obtain this level of reforming decreases. Fig. 3 shows that after aproximately 40 Blls/lb cat the theoretical temperature decreases slowly until the end of the cycle. This figure suggests that the difference between the actual and theoretical temperatures is a measure of the loss in activity of the catalyst. Besides, industrially, temperature is the parameter normally used to control the operation to achieve an established RONC. The deactivation measured as a temperature increase, AT, was obtained for all the cycles for both types of catalysts. The result for the Pt-Sn catalyst is shown in Fig. 4. In this figure the deactivation increases with time on stream. The shape of the different plots is similar and the increase in temperature could reach 90~ (cycle No. 7). The deactivation rate depends on the characteristis of the feed and the operating conditions. It can be shown that [4] the deactivation rate is a function of the H2/HC molar ratio, The feed composition, the final boiling point of the feed (FBP), octane number of the reformate and operation pressure. In order to generalize the results in fig. 4 and to simulate the deactivation, the data obtained is standardized to the deactivation that could be obtained for certain reference conditions. This yields the relative deactivation ATr. In our case the chosen conditions were: RONC = 93; (N + A) vol% = 50; FBP = 182 ~ pressure = 31.7 kg/cm2; H2/HC molar ratio = 7. Two equations were obtained to represent the relative deactivation, these are: ATr = 0.67723 + 10 (133108 - 026346 (H2/HC))

and ATr = - 0.068728 + 10(3.27447-

1.24617 log n)

The relative deactivation for all the cycles using Pt-Sn catalysts are given in fig. 5. From these data the following can be derived: a) For 5 cycles no deactivation is innitially observed up to 10 Blls/lbcat (approx. 45 days). b) Two sets of deactivation rates are observed, one with 4 cycles and another with two cycles, having a faster deactivation. The later behaviour was correlated with the presence of sulphur compounds in the feed in one case, and an excess of chlorine on the catalyst for the second. The increase in the acidity due to the chlorine excess favored the formation of carbon, and sulphur is known to poison metallic catalysts, and to accelerate carbon deposition. The set of 4 cycles showing the same behaviour was considered to have operated normally, so that the main deactivation was basically due to carbon deposition [5]. As a result an expression was obtained to correlate the relative deactivation of the 4 cycles, for an industrial semiregenerative CR unit operating with Pt-Sn/A1203 catalyst. The resulting equations are the following: from 0 to 60 B lls/lbcat ATI = -4.58397 + 10 (0.01475 Blls/lbcat

+ 0.57527)

and from 60 to 80 B lls/lbcat: AT2 = 0.497134 Blls/lbcat + 2.5616 x 10.3

324 50

40

i

~ ~o

!,o

CYCLE No. 1 i

CVCU~No. 2

| CYCLENo. 3

0

0 CZCU~No. s

0 CYCtZNo. 6 -10

| CYCLE No. 7

o

1~

I

20

3~

4~

BI,I,SS..B CAT.

i 50

i 60

i 70

Fig. 5. Standard deactivation of Pt-Sn catalyst for 6 cycles. The curve obtained from these equations is shown in fig. 6 together with that obtained for semiregenerative Pt-Re/A1203 CR catalysts. It is clear the difference in deactivation behaviour for both catalysts. Pt-Sn catalysts show a monotonic decrease in activity, whereas Pt-Re catalysts have a fast initial deactivation but then it stabilizes out from around 11 to 45 Blls/lbcat (aproximately 50 to 200 days) followed by a slow deactivation process. The consequence of these different performances suggest that for semiregenerative processes a Pt-Re catalyst is better than a Pt-Sn catalyst, but if a continuous regeneration unit is considered then a Pt-Sn catalyst should be preferred. The equations obtained are being used to predict the catalyst performance, and as a tool to decide when the regeneration should be done. Also early detection of catalyst malfunctioning and evaluation of the quality of the catalyst regeneration can be done.

325 Deactivation(~ 50

Pt-Sn/AI203 X Pt-Re/AI203 [

45 40 35 30

20 x .-X"

~.xx,

.._x.~.........~, ~

0

~

0

-'

50

x

~

•

.....~ .,~

~ ...........x ...........X . . . . . . . . x

•

X

X

.... X "

'

,

i

i

i

100

150

200

250

300

i

350

Operation time (days)

Fig. 6. Generalized standard deactivation for Pt-Sn (.) and Pt-Re (x) catalysts.

3. CONCLUSIONS The deactivation of catalytic reforming catalysts was studied in Platforming units. The deactivation was measured by the increase in the operation temperature necessary to obtain a specified performance in processes using Pt-Re and Pt-Sn catalysts. The statistical analysis produced an equation capable to predict the deactivation degree in terms of the operation parameters and properties of the feed.

REFERENCES 1. Wuithier, P., E1 petroleo, refino y tratamiento qufmico. Ed. CECSA, S.A., Madrid (1991) 2. C.M. Tsang, J.B. Butt, W.M.H. Sachtler, Symposium on Advances in Naphta Reforming, New Orleans, August (1987). 3. D.M. Little, Catalytic Reforming, Penwell Books (1985). 4. Manual de Reformaci6n Catalitica de Naftas, Subdireccion de Transformaci6n Industrial, Ed. IMP, Mexico (1985). 5. L.M. Rodrfguez Otal, M.Sc, Thesis, UNAM, Mexico (1989).

This Page Intentionally Left Blank

9

Elsevier Science B.V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

327

Effects o f the Metal-Metal Interactions on the Stability o f Pt-Re/A1203-C1 R e f o r m i n g Catalysts J. Barbier a, P. Marecot a and C.L. Pieck b aUniversit6 de Poitiers, U.M.R.C.N.R.S. 6503 Laboratoire de Catalyse en Chimie Organique 40, Avenue du Recteur Pineau 86022 Poitiers Cedex, France. blnstituto de Investigationes en Catalisis y Petroquimica I.N.C.A.P.E. Santiago del Estero 2654 3 000 Santa Fe, Argentina.

Bimetallic Pt-Re/AI203 catalysts were prepared by catalytic reduction (CR) of a Re salt on a monometallic Pt/AI203 catalyst using hydrogen as the reducing agent. The degree of interaction between Pt and Re was indirectly measured by the cyclopentane hydrogenolysis reaction. This interaction is enhanced for catalysts prepared by the CR method. A further calcination can destroy that strong interaction. The most beneficial effects of Re on the stability of Pt (lower coke deposition, lower sulfur coverage, lower toxicities of coke and sulfur) are obtained when Pt and Re are in close interaction.

1. INTRODUCTION The bimetallic Pt-Re/A1203-C1 catalyst is the most widely used in naptha reforming. The addition of Re strongly improves the stability of the traditional monometallic Pt catalyst. Such improvement is explained by a double effect of rhenium: stabilization of the metallic phase on the support and higher resistance to deactivation by coke deposition [ 1-8]. Nevertheless the role and the nature of the interaction between Pt and Re are the subject of many controversies [9-14]. The main difficulty in the preparation of bimetallic catalysts is to bring the two metals into close interaction. Classical coimpregnation or successive impregnations often prove to be unsatisfactory and new techniques are being used. In this work, we investigated the preparation of bimetallic Pt-Re catalysts by surface redox reaction between hydrogen adsorbed on a "parent" monometallic platinum/alumina catalyst and the cation of the second metal (Re) according to the following reaction : nH~s+M n+ --~ M~s+nH +

(1)

328 where Hads is the hydrogen adsorbed on platinum, M n+ is the cation of the second metal (rhenium) and Ma~ is the second metal (rhenium) adsorbed. The objective of this work was to compare the catalytic properties of Pt-Re/A1203-CI catalysts prepared either by the surface redox reaction or by classical coimpregnation. According to Augustine and Sachtler [15-16], the interactions between platinum and rhenium were probed by using hydrogenolysis of cyclopentane as a test reaction. Finally the effect of Pt-Re interactions on sulfur adsorption and on coke deposition on the different Pt-Re/A1203-CI catalysts was investigated. The toxicity of these poisons was evaluated with regards to a pure metallic reaction, in this case dehydrogenation of cyclohexane.

2. EXPERIMENTAL

2.1. Catalyst preparation The support was a cubic T-alumina (Rh6ne Poulenc) with a specific surface area of 246 m2/g, a particle density of 1.11 g/cm3, a total pore volume of 0.60 cm3/g and an apparent density of 0.65 g/era3. The support was ground and then sieved in order to retain particles with sizes between 0.25 and 0.10 mm. A monometallic platinum/alumina catalyst was prepared by impregnation with H2PtC16 dissolved in 0.2 M HC1. After drying at 120~ overnight, the catalyst was calcined in flowing air at 450~ for 4 h and then reduced in flowing hydrogen at 500~ for 8 h. The Pt/A1203 catalyst thus obtained had 0.60 wt% Pt, 1.10 wt% CI and 56% metallic dispersion. In order to obtain catalysts with different metallic dispersions, a portion of the batch was sintered by means of heat treatment at 600~ under oxidizing atmosphere (1% 02 in nitrogen) for different times (2 and 10 h). Bimetallic platinum-rhenium catalysts were prepared either by coimpregnation or by catalytic reduction (redox reaction).

2.1.1. Coimpregnation A given amount of alumina support was immersed in a 0.20 M HC1 solution. After addition of Pt and Re precursors (H2PtCI6 and NH4ReO4), water was evaporated in a sand bath at 70~ Then catalysts were dried at 120~ overnight. Finally the catalysts were activated by calcination in air at 450~ for 4 h, followed by reduction in pure hydrogen for 8 h at 500~

2.1.2. Catalytic reduction This technique was developed in order to favor the deposition of the second metal just on top of the pre-reduced supported metal. The monometallic catalyst was modified by the surface reaction between chemisorbed hydrogen and the ion of the second metal; the schematically involved reaction, in the case of Pt-Re catalysts, is as follows:

329 ReO 4 + 8H § + 7e-

r

Re ~ + 4 H 2 0

(2)

7(Pt-H

r

Pt+H++e )

(3)

Re O 4 + 7Pt - H + H §

r

Re~

(4)

The addition of the second metal can be done in an inert atmosphere (e.g. nitrogen) or in the presence of hydrogen. In the first case, the reaction is limited to the number of hydrogen atoms irreversibly chemisorbed over the surface of Pt, and the maximum amount of the second metal that can be incorporated depends on its valence and the Pt accessibility. In our case, a maximum ratio Re/Pt of 1/7 would be obtained. If the reduction of the second metal is done in the presence of hydrogen (catalytic reduction method), its adsorption on Pt is produced simultaneously with its consumption by Re reduction. Therefore, with this procedure higher Re/Pt ratios can be obtained. The technique of catalyst preparation by the catalytic reduction method is as follows: a given amount of pre-reduced Pt/A1203 monometallic catalyst was placed in a 0.20 M HCI solution, which was purged with a nitrogen stream to eliminate dissolved oxygen. Then hydrogen was bubbled for two hours, and finally a solution containing the rhenium salt precursor, previously purged by bubbling nitrogen, was added without interrupting the hydrogen bubbling. After one hour, the catalyst was washed with distilled water, filtered and dried at 120~ overnight. Two activation procedures were used: direct reduction (Hz, 8 h, 500~ and calcination (air, 4 h, 450~ followed by reduction (H~, 8 h, 500~

2.2. Metal accessibility Metal accessibility was determined by the H2-O2 titration method in a static volumetric equipment. Hydrogen preadsorbed at 500~ was titrated at 25~ by oxygen. Isotherms have been determined in the 0-50 torr range. Extrapolation to the original values was used to calculate the number of exposed atoms. 2.3. Adsorption of hydrogen sulfide Hydrogen sulfide chemisorption was carried out in a continuous flow reactor at normal pressure. The catalyst sample was reduced again for 2 h under hydrogen flow at 500~ Then the hydrogen sulfide/hydrogen mixture (50 ppm of hydrogen sulfide) was introduced. A frontal analysis with a photo-ionization detector allowed the amount of sulfur adsorbed by the catalyst (Stot) to be measured. After the catalyst was saturated with sulfur, pure hydrogen was passed through the sample for 10 h, which led to the desorption of some sulfur (Sre~). The difference between Stot and Srev allows the determination of the amount of irreversible sulfur adsorbed at 500~ 2.4. Test reactions Catalysts were evaluated by means of different test reactions: cyclohexane dehydrogenation, cyclopentane hydrogenolysis and coking reaction with cyclopentane at high pressure (15 bar). The reaction conditions are summarized in Table 1. Before performing the test reactions, presulfided catalysts were reduced at 500~ for 8 h and non-sulfided catalysts for 2 h. Coke

330 deposition was analyzed by temperature-programmed oxidation (TPO) following the technique described by Barbier et al. [ 17].

Table 1 Reaction conditions (CH: c~r CP: c,yclopentane) Reaction CP CH conditions Hydrogenolysis Dehydrogenation Temperature (~ H 2 flOWrate (cm3/min) CP flow rate (cm3/h) CH flow rate (cm3/h) Mass of catal),st (ms)

290 80 2 100

270 100 2 20

CP coking High pressure 450 120 4 600

The toxicity of coke for the metallic function was defined by its fouling effect on the reaction of cyclohexane dehydrogenation. For catalysts coked at high pressure, cyclohexane dehydrogenation was performed ex-situ with an initial reduction time of 8 h at 500~ Previous experience has shown that the reduction time (2 or 8 h) does not affect the activity of precoked catalysts.

3. RESULTS AND DISCUSSION 3.1. Hydrogenolysis of cyclopentane

Table 2 shows conversion values obtained during cyclopentane hydrogenolysis on the two series of Pt-Re bimetallic catalysts (CI: Coimpregnation, CR: Catalytic reduction). The different catalysts were activated by direct reduction (R) or by calcination followed by reduction (C). Table 2 Conversion of CH dehydrogenation and CP hydrogenolysis and total metal dispersion of catalysts prepared by different methods (%Re)/ (%Re+%Pt)

CH conv. (%)

CP conv. (%)

Total metal dispersion

Preparation method*

0.50 0.50 0.50

14.2 12.2 15.0

3.8 14.0 4.2

60 44 58

CI CR (R) CR (C)

0.67 0.67 0.67

14.8 9.4 13.0

9.0 26.9 10.6

55 42 53

CI CR (R) CR (C)

* R: Catalysts activated by direct reduction; C: catalysts activated by calcination followed by reduction.

331 In the case of catalysts prepared by catalytic reduction (redox technique) the conversion is higher when catalysts were activated by direct reduction than when activation was done using calcination followed by reduction (C). In that last case, activity for cyclopentane hydrogenolysis is comparable to that obtained on Pt-Re catalysts prepared by coimpregnation. Hydrogenolysis of cyclopentane is a demanding reaction [18], which needs large ensembles of Pt metal atoms. Ensembles containing both Pt and Re have higher hydrogenolytic activity than ensembles of the separate metals; accordingly, hydrogenolysis of CP can be used as an indirect measure of Pt-Re interaction [15-16]. This interaction was studied by Augustine and Sachtler [910] using Pt, Re and Pt-Re catalysts supported on SiO2 and A1203. They found a synergetic effect which was ascribed to the fact that the heat of adsorption on the metal surface is decisive in the formation of the complex substrate-metal that will undergo C-C bond scission. An ensemble containing only Pt atoms would have lower heat of adsorption than that of pure Re, while an intermediate value would be observed in the case of bimetallic particles containing Pt and Re. It can be argued that the pure Re catalyst has lower activity for hydrogenolysis due to the inability of Re to dissociate H2. When Pt is present it is possible that hydrogen spillover onto Re could produce a higher hydrogenolysis activity even when Pt an Re are not alloyed. From Table 2, it can be concluded that the interaction between Pt and Re increases according to the following sequence: coimpregnation < catalytic reduction (C) < catalytic reduction (R). So calcination greatly modifies the interaction between the metals in catalysts prepared by catalytic reduction.

3.2. Adsorption of hydrogen sulfide Different catalysts were sulfided at 500~ by the hydrogen sulfide/hydrogen mixture and then treated at the same temperature under pure hydrogen for 10 h. After such a treatment, only sulfur strongly bound to the metallic surface remains adsorbed on the catalyst since sulfur adsorption on the alumina support is wholly reversible under hydrogen at 500~ [ 19]. The sulfur coverages (0s) calculated from the amounts of irreversible sulfur are reported in Table 3.

Table 3 Sulfur coveral~e and cyclohexane (CH) dehydrogenation X = wt%Re/(wt%Re + wt%Pt) Catalyst wt%C1 Accessibility 0s CH dehydrogenation a/ao x 100 0.6 Pt/A1203 1.10 56 0.31 77 0.6 Pt/A1203 1.10 20 0.41 0.6 Pt/A1203 1.10 12 0.51 0.6 Re/A1203 1.10 36 1.31 PtRe/A12OaX-0.67 CI 1.10 55 0.94 14 PtRe/AlEOaX=0.67 CR (R) 1.10 42 0.80 38 PtRe/AlEOaX-0.67 CR (C) 1.10 53 0.92 26

332 For monometallic Pt catalysts, sulfur coverage increases with particle size while the sulfur adsorption capacity is higher for the monometaUic Re catalyst. This last result is in accordance with previous work, the adsorption of an electron acceptor like sulfur being enhanced on metals of low electronic affinities [20]. Indeed, electronic affinities are 2.12 eV and 0.16 eV for platinum and rhenium, respectively. With regard to the effect of platinum particle size, the increasing 0s results from the electron deficient character of small metal crystallites deposited on an acidic support [20, 21]. For bimetallic catalysts, the sulfur coverage (0s) is lower on catalysts exhibiting the highest PtRe interaction determined by cyclopentane hydrogenolysis (Table 2). The effect of irreversible sulfur on catalyst activity was investigated in the course of cyclohexane dehydrogenation (Table 3). The comparison of relative activities (activity of sulfided catalyst/activity of fresh catalyst) shows that dehydrogenation is more strongly inhibited by sulfur adsorption on bimetallic catalysts. However, the higher the Pt-Re interaction, the higher the remaining activity of sulfided bimetallic catalysts. 3.3. Coke deposition The effect of sulfur and Pt-Re interaction on coke deposition were studied at high pressure (15 bar) in the course of cyclopentane reaction at 450~ for 16 h (Table 4). The catalyst prepared by catalytic reduction and activated by direct reduction (close Pt-Re interaction), deposits less coke than catalysts activated by calcination followed by reduction (weak Pt-Re interaction).

Table 4 Coking of non sulfided and sulfided catalysts (16 h under c;r Catalyst (X) %C

flow at 15 bar) CH Dehydrogenation a/ao x 100

Non Sulfided catalysts 0.00 0.50 CR (R) 0.50 CR (C) 0.50 CI

2.11 0.35 0.65 0.60

1.1 36 22 25

2.47 0.12 0.14 0.30

0.7 18 11 10

Sulfided catalysts 0.00 0.50 CR (R) 0.50 CR (C) 0.50 CI

Catalyst presulfidation induces a strong decrease of the amount of coke deposited on bimetallic Pt-Re catalysts, the effect being more pronounced on catalysts prepared by catalytic reduction (Table 4). However, coke deposits are less toxic for cyclohexane dehydrogenation on catalysts activated by direct reduction, i.e. on catalysts where Pt-Re interaction is high.

333 Such higher residual activity could be in accordance with the model proposed by Sachtler, which assumes that S is preferentially adsorbed on Re sites in bimetallic Pt-Re particles [20]. Indeed, as the Pt-Re interaction is the highest, sulfur ought to divide the Pt-Re surface into very small Pt ensembles. Consequently, the reorganization of the carbonaceous overlayer into pseudo-graphitic entities, which are detrimental to the metallic function, could be impeded.

4. CONCLUSIONS Bimetallic platinum-rhenium catalysts can be prepared in aqueous acid medium, under hydrogen flow, by a redox reaction between hydrogen activated on a parent platinum-alumina catalyst and the perrhenate ion Re O ~. The activation treatments affect the metal-metal interaction which is higher when the catalysts are reduced after Re deposition. Such interactions are destroyed by calcination under oxidizing conditions. Bimetallic Pt-Re catalysts prepared by catalytic reduction and activated by reduction (close interaction) are less sensitive to sulfur adsorption than catalysts prepared by the classical coimpregnafion technique. The deposition of coke and its toxicity for the metallic function decrease as the Pt-Re interaction increases. Finally, sulfided Pt-Re catalysts prepared by catalytic reduction are less sensitive to coke deposition. Furthermore, coke deposited on that catalyst is less toxic for the metallic function. In conclusion, if the addition of Re strongly improves the stability of the classical reforming Pt catalyst, the beneficial effect is improved when the two metals are in close interaction in the same bimetallic particles.

Acknowledgements The research was carried out within the framework of an official International Scientific Collaboration (P.I.C.S.) between France and Argentina with the economical support of the C.N.R.S. (ECOTECH program) and the CONICET.

REFERENCES

~

7.

H.E. Klusksdhal, US Patent 3415737 (1968). J.M. Parera and J.M. Beltramini. J. Catal., 112 (1988) 357. J. Biswas, P.G. Gray and D.D. Do, Appl. Catal., 32 (1987) 249. J. Barbier, Appl. Catal., 23 (1986) 225. J. Margitfalvi, S. G6b616s, E. Kaysser, M. Hegediis, F. Nagy and Koltai, React. Kinet. Catal. Lett., 24 (1984) 315. D.M. Little, Catalytic Reforming, Permwell, OK, 1985. B.C. Gates, JR. Katzer and G.C.A. Schuit, Chemistry of Catalytic Processes,

334

.

9. 10. 11. 12. 13. 14. 15. 16. 17. 18. 19. 20. 21. 22.

Mc Graw-Hill, New-York, 1979. R.J. Bertolacini and R.J. Pellet, Stud. Surf. Sci. Catal., 6(1980) 73. M.F.L. Johnson and V.M. Leroy, J. Catal., 3__5(1974) 434. S.M. Augustine and W.M.H. Sachtler, J. Catal., 116 (1989) 84. N.S. Nacheff, L.S. Kraus, M. Ichikawa, B.M. Hoffman, J.B. Butt and W.M.H. Sachtler, J. Catal., 106 (1987), 263. N. Wagstaff and R. Prins, J. Catal., 59 (1979) 434. B.D. Me Nicol, J. Catal., 4...66(1977) 438. A.N. Webb, J. Catal., 3..29(1975) 485. S.M. Augustine and W.M.H. Sachtler, J. Catal., 106 (1987) 417. S.M. Augustine and W.M.H. Sachtler, J. Phys.Chem., 91 (1987) 5953. J. Barbier, E.J. Churin, J.M. Parera and J. Rivi+re, React. Kinet. Catal. Lett. 2_29 (1985) 323. R. Maurel, G. Leclercq and J. Barbier, J. Catal., 3..7.7(1975) 324. C. Apesteguia and J. Barbier, J. Catal., 7..88(1982) 353. V.K. Shum, J.B. Butt and W.M.H. Sachtler, J. Catal., 96 (1985) 371. J. Barbier, P. Mar6cot, L. Tifouti, M. Gu6nin and R. Frety, Appl. Catal. 19 (1985) 375. J.C. Vedrine, M. Dufaux, C. Naccache and B. Imelik J. Chem. Soc. Faraday Trans., 7__4 (1978) 440.

9 Elsevier Science B.V. All rights reserved Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

335

T e m p e r a t u r e P r o g r a m m e d O x i d a t i o n of D e a c t i v a t e d Pt/Nb205 Catalysts D.A.G. Aranda a, J.C. Alfonso b, R. Fretyc and M. Schn'lal a'd "NUCAT-Universidade Federal R. Janeiro - Cx. Post. 68502-CEP 21945-970-RJ-RJ BRASIL e-mail: donat o@peq, coppe, ufrj. br bInstituto de QuimicaJUniversidade Federal do Rio de Janeiro, Brasil CInstitute Recherches sur la Catalyse and LACE/CNRS, Villeurbanne, France dEscola de Quimica/Universidade Federal do Rio de Janeiro, Brasil

This work presents a detailed study of coke deposition on deactivated Pt/Nb205, Pt/AI203 and Pt-Sn/Nb205 catalysts during n-heptane dehydrogenation. The spent catalysts were characterized by means of temperature-programmed oxidation analyzed by mass spectrometry (TPO/MS) and by differential scanning calorimetry (TPO/DSC). The low density of acidic sites and the SMSI effect yielded a more hydrogenated coke on Pt/Nb205 catalyst than on Pt/Al203. Addition of tin promotes a decrease of the total coke which is more hydrogenated than the coke on Pt/Nb205 samples. Analysis of soluble coke suggests a relationship between coke precursors and metallic surface structure. 1.INTRODUCTION Nafla reforming and alkane dehydrogenation processes are directly connected on platinum alumina based catalysts. The stability and selectivity requirements for industrial purposes induced the addition of a second metal like Sn, Ge or Re. In fact, these promoters coupled to other acidity controllers added on the support enhanced the catalytic efficiency. [1-3]. Recently, Pt/Nb205 catalysts have been investigated on dehydrogenation of alkanes. These systems have presented advantages on selectivity towards olefins when compared to Pt/A1203 or even Pt-Sn/A1203 catalysts [4-6]. The promoting mechanism is related to both the SMSI (Strong Metal Support Interaction) effect and the low acidity of the support, which produce a sharp decrease in hydrogenolysis and aromatization, respectively [5]. Despite of good selectivity, these catalysts suffer fast deactivation due to the highly dehydrogenated by-products, which cause blockage of the metallic sites and metal/support interfaces, thus decrease the catalytic activity. Coke deactivation on Pt/AI203 catalysts have been studied intensively in the literature. Previous works have focused on: the kinetics of catalyst deactivation [7]; the influence of additives on coke formation [8]; the coke deposition on different morphologic surfaces [9]; the structure [10] and chemical composition of coke [11]. Deactivation by coke deposition on niobia supported catalysts, or even on other reducible supports which promote SMSI effect has not been studied.

336 This work presents a detailed study of coke deposition on Pt/Nb205 catalysts which were deactivated with n-heptane dehydrogenation. Temperature programmed oxidation (TPO) allowed to the identification of different coke oxidation zones which are related to the distribution of carbonaceous species on the support and on the metallic sites. Chemical analysis of soluble coke was performed allowing to a better understanding of the deactivation processes on niobia supported catalysts. 2. EXPERIMENTAL The Nb205 support (BET area = 65 mZ/g) was obtained by calcination of niobic acid (CBMM, HY 340lAD 929) in air at 773 K for 2 h. A commercial ~'-A1203 from Harshaw (AI 3996, BET area: 200 m2/g) was employed as support to Pt/AI203 reference catalyst. The monometallic and the bimetallic catalysts were prepared by incipient wetness technique, following a procedure described elsewhere [4]. The platinum and tin content were 1% wt/wt on mono and bimetallic catalysts. The deactivation was obtained by n-heptane dehydrogenation. Prior to the catalytic tests, the samples were dried with N2 in situ at 423 K and then reduced with H2 at a heating rate of 10 K/rain up to 773 K. The reaction was performed in a flow microreactor at atmospheric pressure and at 773 K. The molar ratio was H2/n-C7H16 = 16 and the space velocity WHSV=2-7 h "1. After 4 h under reaction, the catalysts were purged with N2 flow for 30 min and cooled to room temperature. The surface area and pore volume of the fresh and deactivated catalysts were analyzed through BET (ASAP 2001) and the elemental analysis of carbon through a LECO Analyser. The deactivated Pt/Nb205 and Pt/A1203 catalysts (12 g) were submitted to extraction of soluble coke in a soxhlet apparatus with toluene for 24 hours. The extracts were analyzed by gas chromatography-mass spectrometry (GC/MS). The experimental procedure is given elsewhere [ 12]. Temperature programmed oxidation (TPO/MS) was performed with 50 mg of the deactivated samples in a microreactor with 1% O2/He, flowing at 30 ml/min (heating rate: 5 K/min). The reactor effluent was monitored by a Balzers QMS200 quadrupole mass spectrometer. Carbon monoxide was not detected in the effluent gas. Thus, the profiles of O2 consumed and CO2 produced represent the complete oxidation. The deactivated catalysts were also burned in a calorimeter under compressed air flowing at 40 ml/min. (Thermal Analysis Station TAS 100 - Rigaku - TG 8110). Analysis were performed by differential scanning calorimetry (TPO/DSC) and combustion heats were obtained by integration of the profiles and using appropriate calibration samples. 3. RESULTS 3.1 Textural properties and coke amount A visual observation shows a deep and uniform dark color of the deactivated catalysts alter 4 hours on stream. Some of their textural properties are given in Table 1.

337 Table 1 Textural properties

Specific area BET (mZ/g) 3 Pore Volume (cm/8)

Fresh Catalysts Pt/NbzO5 Pt/Al203 58 200 0.131 0.700

Deactivated Catalysts Pt/Nb205 Pt/A1203 56 195 0.128 0.650

Table 2 presents the results (%wt/wt) of soluble and insoluble coke as well as the total amount of coke on the different catalysts. Table 2 Data relative to the amount of coke on deactivated catalysts Type of Coke %wt/wt Pt/Nb205 Pt/A1203 Total coke 3,7 3,1 Insoluble coke 1,9 2,2 Soluble coke 1,8 0,9

Pt-Sru'Nb205 1,2 -

3.2 Chemical composition and identification of the compounds The chemical characterization of the soluble coke was obtained from mass spectral interpretation and from the analysis of key ion mass fragmentograms [12], as well by identification of homologous series and by injection of standards of respective components. Table 3 presents the identified compounds in the extracts. Table 3 shows the major presence of aromatic structures in both extracts and a large distribution of compounds. Noteworthy is the presence of 2-rings aromatic compounds (alkylnaphthalenes and byphenyls), which were obtained from the coke on the Pt/Nb205 samples, whereas the 3-4 condensed rings aromatic structures on Pt/AIzO3 catalysts. 3.3 TPO/MS Figure 1 displays the profiles of 02 uptake and CO2 formation during the TPO of deactivated Pt/Nb2Os, Pt/AI203 and Pt-Sn~b205 catalysts. A physical mixture of the deactivated Pt/AI203 catalysts (25 mg) and Nb205 (25 mg) was also analyzed by TPO searching for the activity of niobium oxide on coke oxidation. The 02 and CO2 profiles are very similar. This suggests that 02 is only consumed for the formation of CO2. Thus, the 02 uptake for the reoxidation of the partially reduced niobia and the metallic particles is negligible. The coke burn occurs at two oxidation zones on the niobia catalysts: the first one, of limited contribution, at low temperatures between 528 and 553 K, and the second zone at higher temperatures (623-663 K). The Pt/Nb205 catalyst indicates higher temperature ranges for the oxidation zones compared to the bimetallic Pt-Sn/Nb205. The deactivated Pt/Al203 catalyst exhibits two oxidation peaks at 538 and 707 K. However, the addition of niobia promotes a slight decrease of the oxidation peaks at 523 and 692 K, respectively.

338 Table 3 Classes of organic compounds present in the extracts Classes of Hydrocarbons n-alkanes (%) CH3-(CH2)n-CH3 n = 10,11 ..... 29

Pt/Nb205 Pt/AI203 (%wtlwt)" (%wtlwt) ~ 5.2 3.3 3.0

Branch. alkanes C15-C26

(%)

C15-C26

1.3

1 ring (%)

C2-Cs-alkylbenzenes(MW 162-190)

2.5

2 rings (%)

Co-C4-alkylnaphthalenes (MW 128-184) biphenyl, methyl and dimethylbiphenyi (MW 154182) phenanthrene and anthracene (MW 178) C~-C3-alkyl-phenanthrene./anthrac. (MW 192-220) methyl-phenylnaphthalenes (MW 218-260) trimethyl-triphenyls (MW 272) fluorene, methyl and dimethylfluorene (MW 165-193)

27.5 38.5

6.3 4.3

1.5 7.0 1.5

18.0 23.0 4.3

Mono-olefins

(%)

Aromatic compounds:

3 rings (%)

4-5 rings (%)

6-7 rings (%)

triphenylene (MW 270) methyl-phenylphenanthrene (MW 268) bi-naphtyl (MW 254) pyrene and fluorantene (MW 202) methylpyrene and fluorantene (MW 216) benzopyrene / methylbenzopyrene (MW 252-266) C21Hl5 (MW 267) C20Hll (MW 251) chrysene (MW 228)

C28H24 (MW 360) C2sH20 (MW 356) C28H!9 (MW 3 55) C27H18 (MW 342) a %wt/wt of soluble coke *MW = molecular weight

-

15.0

1.5

1.3

2.6 1.4

1.2

1.2

2.3 3.5 -

4.5 3.6

4.3 1.0 1.7 2.1

0.5 -

0.5 1.5

1.0 0.5 2.1

339

02

02 /1 f

'--,\,

/,," ",,\

Pt/Nb205

Pt/A1203 < v

/ ~ .

.

303

.

.

.

400

.

~',,, CO2

.

.

830

.

703

300

403 5O0 O30 7O3 8O0 T(K)

T(K)

02

O2

~]//-""

Pt/AI 203 + Nb205 <

Pt-Sn/Nb205

v

~CO I

,

I

,

I

,

I

i

I

i

I

33O 40O .930 83O 7OO 80O T(K)

~0'

I

400

'

I

500

' ~ 0

T(K)

.

2 I

700

.

I

830

'

Figure 1 - TPO/MS profiles of deactivated catalysts Table 4 presents the CO2/O2 ratio for the oxidation of the niobia catalysts. This ratio decreases in the bimetallic Pt-Sn/Nb205 catalyst forming less CO2, which suggests a more hydrogenated coke. Moreover, the compounds of the soluble coke, already shown in Table 3, indicate that coke was more hydrogenated on Pt/Nb205 than on Pt/Al203 catalyst.

340 Table 4 CO2/02 ratios during TPO Catalysts Pt/Nb205 Pt-Sn/Nb205

CO2/O2 (area ratios) 1,05 0,90

3.4 TPO/DSC Table 5 presents the heat of combustion of dispersed coke on the deactivated catalysts. The Pt-Sn/Nb205 catalyst exhibits the highest exothermicity during the coke burn, which is related to higher H/C ratios in coke. Similarly, the Pt/Nb205 catalyst evidences the presence of lighter hydrocarbons in coke than the Pt/AI203 catalyst. Lower heat of combustion was observed in the insoluble coke compared to the total coke. This result suggests a correlation between the soluble coke and the highest H/C ratios. Table 5 Heat of Combustion of Coke Catalysts Pt/A1203 (total coke) Pt/A1203 (insoluble coke) Pt/NbzO5 (total coke) Pt/NbzO5 (insoluble coke) Pt-Sn/Nb205 (total coke)

Heat of Combustion (Kcal/gcoke) 9,1 8,0 10,0 9,8 10,5

4. DISCUSSION The presence of different oxidation steps during coke burn can be related to the acidity of the support and the structure of the metallic phase. Barbier [13] has shown that the first oxidation zone at low temperatures corresponds to the carbonaceous materials on the metallic phase, whereas at high temperatures of oxidation are attributed to the coke on the acidic sites of the support. In our case, the total acidity varies in a wide range after calcination at 773 K. In fact, the NH3 uptake on alumina was 750 ktmol/g catalyst, whereas on calcined niobia it was only 112 l.tmol/g catalyst [5]. Moreover, according to Pittman and Bell [14], the acidic sites on a similar niobia support are basically weak Lewis sites. On the other side, coke burn in the neighboring of the metallic sites can be also associated to the promoting effect of platinum and the platinum-niobia interface. The TPO profile of the Pt/AI203 catalyst is very similar to those presented in the literature. However, the Pt/Nb205 system displays a different behavior, particularly lowering the temperature peak at the high temperature oxidation zone. Augustine et al [15] suggested that the polymerization activity of hydrocarbon structures during the coke formation was attributed to Lewis acidity. As niobia support is less acidic than alumina, the polymerization on niobia is also less effective, thus leading to a lighter and less complex coke formation than on the alumina supported samples. Moreover, the heat of combustion of the total coke and of the insoluble coke on Pt/Nb205 are very similar (Table 5). If insoluble coke indicates more polymerized species in coke, it seems reasonable that polymerization reactions on niobia supported catalysts lead to a lower

341 decrease of the H/C atomic ratio than on the alumina-supported catalysts. In the latter case, polymerization with formation of condensed aromatic rings leads to a greater loss of hydrogen when compared to formation of byphenyl structures on niobia. The H/C atomic ratio in the carbonaceous deposits is 20% higher on Pt/Nb205 than on Pt/Al203. In fact, the TPO profiles and DSC results exhibit very distinct oxidation zones on Pt/Al203 in contrast to the Pt/Nb2Os.catalyst The strengths of the Lewis acidic sites are important for draining the coke precursors to the support which promote more polymerized coke than the metallic sites. Acidic sites at the niobia surface were not able to attract these carbonaceous species. Thus, the metallic sites were rapidly saturated which cause fast deactivation on Pt/Nb205 catalyst. The structure of the metallic phase is also relevant. Niobia supported catalysts present partial reduction of the support in presence of the metal (under these reduction temperature conditions). Previous results agreed also with the decoration model, where niobia reduced species migrate onto the platinum surface [16]. According to the soluble coke composition results (Table 3), the byphenyl-type structures prevail on Pt/Nb2Os, whereas condensed aromatic rings prevail on the Pt/AI203 system. Thus, it suggests directly that the decorated platinum particles influences the growing of the carbonaceous structure on Pt/Nb205 catalyst. Despite of all the differences of chemical nature and location of coke on Pt/AI203 and Pt/Nb205 catalysts, the oxidation activity of niobium oxide itself cannot be discarded. In fact, the TPO results of the physical mixture of deactivated Pt/A1203 and fresh niobium oxide exhibit a slight decrease of the oxidation peak in the high temperature oxidation zone. On the other side, the addition of a second metal like Re, Ir or Sn, which increases the stability of Pt/AI203 catalyst, has been studied extensively in the literature. In summary, the effect of Sn is ascribed to: i) a reduction of the number of Lewis sites [17]; ii) an increase of selective hydrogenation of coke precursors (dienes) [18]; iii) a suppression of hydrogenolysis reactions [19]. Moreover, Sn can promote alkane oligomerization or even Diels-Alder type reactions [20]. The TPO profile was drastically changed after tin addition to the Pt/Nb205 catalyst. The heat of combustion (Table 5) and CO2/O2 ratio results (Table 4) suggest that coke formation on Pt-Sn/Nb205 is even lighter and more easily oxidized than the coke on Pt/Nb2Os. This effect cannot be attributed to the reduction of the number of Lewis sites, since it would inhibit the migration of carbonaceous species onto the support and thus a faster deactivation of the catalyst. On the contrary, Pt-Sn/Nb205 is more stable than the monometallic and the results of the ammonia chemisorption did not show any increase in the acidity [5]. The effect of tin addition can be explained as described before [18,19], which are related to the ligand and geometric effects in the Pt-Sn interaction. Again, the oxidation activity of the bimetallic system cannot be discarded, since Pt/SnO2 catalyst is very active in the CO oxidation [21 ]. This latter assumption needs further experiments. 5. CONCLUSIONS The lower density of acid sites and the structure of metallic sites under SMSI state lead to a type of coke on Pt/Nb205 catalyst which is more hydrogenated and lighter than the coke on Pt/AI203. Coke on Pt/Nb205 is more easily oxidized; however, the oxidation character of niobia cannot be discarded.

342 Weak acidic sites on niobia surface were not able to induce high coke polymerization on the support. Thus, the carbonaceous sub-products blocked the metallic sites and promoted fast deactivation with n-heptane dehydrogenation. In addition, it is proposed that the decorated platinum particles influences directly the growing of the carbonaceous structure. The addition of tin leads to a decrease of coke content on the Pt/NbzO5 catalyst, which was lighter than the coke on the monometallic system. In this case, the effect was explained by Pt-Sn interaction. ACKNOWLEDGEMENTS We are grateful to Ricardo Silva Aderne (NUCAT/COPPE/UFRJ) for TPO/DSC analysis. We thank also CNPq for financial support and CBMM for providing the niobic acid. Donato A.G. Aranda is grateful to FAPERJ (Funda~:~o de Amparo ~. Pesquisa do Estado do Rio de Janeiro) for the financial support. REFERENCES 1. Dautzenberg, F.M.; Helle, J.N.; Biloen, P. and Sachtler, W.M.H., J. Catal., 63 (1980) 119. 2. Passos, F.B.; Vannice, M.A. and Schmal, M., J. Catal., 160 (1996) 106. 3. Beltramini, J. and Trimm, D.L.; Appi. Catal., 32 (1987) 71. 4. Aranda, D.A.G.; Noronha, F.B.; Schmal, M.; Passos, F.B., Appl.Cat., 100 (1993) 77. 5. Aranda, D.A.G.; Noronha, F.B.;Passos, F.B.;Schmal, M., Cat.Today 16 (1993) 397. 6. Aranda, D.A.G.; Ramos, A.L.D.; Passos, F.B. and Schmal, M., Cat. Today 28 (1996) 119. 7. Barbier, J.; Marecot, P.; Martin, N.; Elassal, L. and Maurel, R., Catal. Deactiv. (1980) 53. 8. Margitfalvi, J.; Szedlacsek, P.; Heged, M. and Nag~y, F., Appl. Catal., 15 (1985) 69. 9. Espinat, D.; Freund, E.; Dexpert, H. and Martino, G., J. Catal., 126 (1990) 496. 10. Gallezot, P.; Leclercq, C.; Barbier, J. and Marecot, P., J. Catal., 116 (1989) 164. 11. Wolf, E.E. and Alfami, F.; Catal. Rev. Sci.Eng., 24 (1982) 329. 12. Afonso, J.C.; Schmal, M.; Cardoso, J.N. and Frety, R., Ind. Eng. Chem. Res., 30 (1991) 2133. 13. Barbier, J. Appl. Catal., 23 (1986) 225. 14. Pittman, R.M.; Bell, A.T. Cat. Lett. 24 (1994) 1. 15. Augustine, S.M.; Alameddin, G.N. and Sachtler, W.M.H., J. Catal., 115 (1989) 217. 16. Haller, G.L. and Resasco, R.E. Adv.Catal., 36 (1989) 173. 17. Sarkany, A.; Lieske, H.; Szil/~gyi, T.; T6th, L., Proc. 8th. Int.Cong.Catal. Berlin; V 2, p. 613. 18. Butch, R. and Mitchell, J.A., Appl. Catal., 6 (1983) 121. 19. Volter, J. and Kurschner, Appl Catal., 8 (1983)167. 20. Afonso, J.C.; Schmal, M. and Frety, R., Fuel Proc.Tech., 41 (1994) 13. 21. Schryer, D.R.; Upchurch, B.T.; Van Norman, J. D.; Brown, K.G. and Schryer, J.; J. Catal., 122 (1990) 193.

9

Elsevier Science B.V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

343

Effects o f Sulfidation o f M o Nitride and C o M o Nitride Catalysts on T h i o p h e n e H D S Son-Ki Ihm a, Do-Woan Kim and Dong-Keun Lee b Dept. of Chem. Eng., Korea Advanced Institute of Science and Technology

Mo and CoMo nitrides were prepared by temperature-programmed reduction of the corresponding oxides with flowing ammonia; effects of sulfidation of these catalysts on their thiophene hydrodesulfurization (HDS) activities were investigated. The HDS activity of fresh CoMo nitride is higher than that of Mo2N. The properties of nitrides vary significantly with sulfidation temperature. After sulfiding Mo2N transforms into MoS2 and the CoMo nitride into MoS2 and Co9S8. This indicates that these nitride catalysts are not resistant to sulfur under severe sulfidation conditions. The transformation, on the other hand, results in a synergistic activity of HDS among the three phases present.

1. INTRODUCTION High surface area molybdenum nitride (MozN) and alumina-supported molybdenum nitrides have recently received a great deal of attention due to their high activities comparable to those of commercial alumina supported catalysts [1-5]. Especially, Mo2N is promising for HDS and HDN because it is reactive toward direct removal of sulfur and nitrogen with less consumption of hydrogen [1,2]. Catalytic activities in hydrodesulfurization reactions have been reported for a number of new bimetallic sulfide catalysts [6,7]. In addition new bimetallic nitrides such as C03Mo3N, CuWN2, FeWN2, Ni3Mo3N and V-Mo-O-N phase have been synthesized [8-12]. In our previous study, it was found that CoMo nitride has higher specific HDS activity than Mo2N [8] and that the addition of Co to Mo nitride improves its HDS conversion. It is anticipated that sulfur compounds might affect the properties of metal nitrides during HDS reaction. However, there has been relatively little investigation of the stabilities of monometallic and bimetallic nitrides during HDS reaction. In this study, Mo2N and CoMo nitrides were prepared from their respective oxides by means of temperature-programmed reduction (TPR) with ammonia, and their reactivities toward thiophene HDS were investigated. Effects of sulfidation of these metal nitrides on thiophene HDS were also examined. X-ray diffraction, elemental analysis and scanning electron microscopy were also carried out to study the structure and compositional changes after sulfidation. aTo whom the correspondence should be addressed bpresent address: Dept. of Chem. Eng., Res. Inst. of Environ. Prot., GyeongsangNational University.

344 2. EXPERIMENTAL 2.1 Preparation of Mo and CoMo nitrides Mo2N and Co3Mo3N were prepared by reacting the corresponding oxides, MoO3 and CoMoO4, with ammonia [8]. CoMoO4 was synthesized by evaporating an aqueous solution of Co(NO3)2" HE0 and (NH4)6Mo7024" 4H20 followed by calcination in air at 500~ for 4 h. Nitriding of the oxides was conducted in a quartz reactor following the procedure by Choi et al. [5]. One gram of the oxide was nitrided by TPR with ammonia. The reduction temperature was rapidly increased from room temperature to 350~ over 30 min, then to 450~ over 150 min, from 450~ to 700~ over 75 min, and finally held at 700~ for 60 min. The freshlyprepared nitrides were passivated by exposure to a 1% O2/He gas mixture at room temperature over 60 min. The ammonia flow rate was fixed at 160 ml/min. 2.2 Sulfidation of Mo and CoMo nitrides Sulfidation of the nitrides was likewise conducted in a quartz reactor. One gram of nitride was sulfided in flowing 10% HES/H2 (60 ml/min)while heating linearly at a rate of 10~ from room temperature to the desired sulfiding temperature under helium flow and then holding at the sulfiding temperature for 5 hours. The sulfiding temperatures were 400, 500, 600 and 700~ respectively. MoS2 and Co9S8 were also prepared for comparison by sulfiding one gram each of MoO3 and Co304 with 10% H2S/H2 mixture (60 ml/min) at 500~ for 10 h respectively. 2.3 Thiophene Hydrodesulfurization Thiophene HDS was carded out over 20 mg of nitrided catalyst at 400~ in a stainless steel microreactor operated at 1.5 Mpa. The flow rates of hydrogen and liquid thiophene at room temperature were 150 ml/min and 0.035 ml/minrespectively. The mole ratio of hydrogen to thiophene flow was 15. Before starting the reaction, each catalyst was sulfided in-situ at the desired sulfiding temperature for 2 h in flowing 10% HES/I-I2 (30 ml/min). Steady state was achieved after 5 h of run time. Reaction products were analyzed by gas chromatography equipped with TCD and a column packed with OV-101. 2.4 Characterization Crystalline structures of oxides and nitrides were analyzed by X-ray diffraction (XRD). Diffraction patterns were collected using a Rigaku DMAX-II diffractometer with CuKo~ radiation. Structural and compositional changes of the CoMo nitrides were observed by SEM (Philips, 535M) and elemental analysis (Elementar Analysensystem. GmbH Vario EL). Surface areas of the samples were measured by N2 adsorption at 196~

345 3. RESULTS AND DISCUSSION 3.1 F o r m a t i o n of C o M o Nitride Phase

Mo2N and a new CoMo bimetallic nitride were obtained by nitriding MoO3 and CoMoO4, respectively. The XRD patterns of the starting oxides and the resulting nitrides are shown in Figure 1. Mo nitride, typically Mo2N, has been known to have high surface area (> 200 m2/g). The reaction between MoO3 and flowing NH3 was topotactic as Mo nitrides had a well-defined crystallographic orientation relative to the MoO3 underlayer [3]. Also a new CoMo nitride phase appears upon reduction of CoMoO4 with ammonia (Figure 1). Elemental analysis shows that the Mo/N ratio for Mo nitride is 1.3, while both Co/N and Mo/N ratios for the CoMo nitride are 2.5. In order to elucidate the chemical composition of the new CoMo bimetallic nitride, Rietveld analysis[11, 13] was conducted on this catalyst based on the assumption that the new XRD peaks of the CoMo bimetallic nitride are representative of the Co3Mo3N phase. The calculated peaks from Rietveld analysis coincided well with the observed XRD pattern. Thus, the new nitride phase is postulated to have a composition of Co3Mo3N.

9

I

M~

:i ._zr

CoMo04

CoMo Nitride

30

I 40

I 50

60

2 Theta

Figure 1. Changes of XRD patterns for MoO3 and CoMoO4 through temperature-programmed reduction with ammonia.

346 Surface areas and catalytic activities of the freshly-prepared Mo nitride and CoMo nitrides for thiophene HDS are summarized in Table 1. The HDS conversion activity of the CoMo nitride (even with a factor of 10 lower surface ar,'a) is comparable to that of Mo2N. However the specific activity (based on surface area) is four times higher for the CoMo nitride phase. Table 1 Catal~ctic activities of Mo2N and C03Mo3N. Nitride Surfacearea, (m2/g)

HDS conversiona, (%)

Mo2N 115 19.5 CoMo nitride (Co3Mo3 N) 19 13.4 a determined after 5 h of run time b corresponding to mole of reacted thiophene/surface area of catalyst

Specific activity b, (~mol/m2 min) 37 155

3.2 Effects of Sulfidation on Mo and CoMo Nitrides

Effects of sulfidation of nitrides on their morphologies were observed after treating them with a 10% H2S/H2 mixture at various sulfiding temperatures. Changes in XRD patterns of Mo2N with the sulfiding temperatures are shown in Figure 2. Markel et al., [4] reported that there was no evidence for MoS2 formation from Mo2N sulfided at 400~ In this study, the characteristic XRD pattern for Mo2N was observed up to sulfiding temperatures of 500~ but disappeared at a sulfiding temperature higher than 600~ At higher sulfiding temperatures the formation of MoS2 was unambiguous. Thus, these results indicate that Mo2N was transformed into MoS2 under severe sulfiding condition. Figure 3 shows the XRD patterns of CoMo nitrides at different sulfiding temperatures. No significant changes are observed at 400~ except for the appearance of Co9Ss. As the sulfiding temperature increases to 500~ the CoMo nitride is transformed gradually into MoS2 and Co9S8. The CoMo nitride phase almost disappears at 700~ Thus it appears that the CoMo nitride is also unstable under H2S atmosphere. Morphological transformation from Mo2N into MoS2 could be observed from SEM photographs. Figure 4 shows the SEM photographs of fresh Mo2N and the same catalyst after sulfiding at 700~ In Photograph A plate-like shapes representative of Mo2N are clearly seen. After sulfiding at 700~ a band-like morphology appears. Surface areasand HDS activities of Mo2N vary significantly with sulfidation temperature (Table 2). As Mo2N is sulfided at 400~ its surface area decreases dramatically from 115 m2/g to 75 m2/g without any phase change. At higher temperatures HDS conversion and surface area of the Mo2N decrease significantly with sulfiding temperature due to transformation into MoS2. The decrease in HDS conversion at lower sulfiding temperature must be due principally to a loss in Mo2N surface area, since its specific activity remains almost the same after sulfidation at 400-500~ However, at higher sulfiding temperatures (600 and 700~ activities based on surface area decrease with the degree of sulfidation, since the specific activity of MoS2 is apparently much lower than that of Mo2N.

347 9Co3Mo3N I |

9 ~I~ 2

,~ MoS 2 9C%S 0

ii

9 M~

9

9

400Oc,r

400~

9

-C

4~

0

500 C-Kdf'Klation

500~

w C

c e

9

9

9 9

t

T

"

I

0

9

,,j". 9

I

t

I

II

t

20

,

L

30

,

I,

t

40

50

..l

60

.

9

,.

i

70

80

20

30

9

I

I

40

50

2 Theta

Figure 2. XRD patterns of sulfided Mo2N.

0

700 C-~f~lation

. . . .

I

I

60

70

.,.

80

2 Theta

Figure 3. XRD patterns of sulfided CoMo nitride.

Table 2 Catalytic activ!t!es and surface ~eas of MozN with increasin~ temperature of sulfidation. Sulfidingtemp. Phase b Surface area, HDS conv c, Specific activity, (~ (m2/g) (%) (~tmol/m 2 min) . . . . . . . . . . . . . . . . . Fresh a 400 500 600 700

Mo2N Mo2N Mo2N+ MoS2 MoS2+ Mo2N MoS2

115 75 55 34 34

19.5 13.2 7.8 2.5 2.1

37 38 31 16 13

catalyst not previously sulfided with 10% HzS/H2. b listed in order of the height of the predominant peak for that phase in the XRD pattern. c determined after 5 h of run time.

a

348

(A) Fresh Mo2N

(B) MoS2 sulfided at 700~

Figure 4. SEM Photographs of (A) fresh Mo2N and (B) MoS2 obtained by sulfiding Mo2N at 700~

Table 3 summarizes changes in the catalytic properties of Co3Mo3N with increasing temperature of sulfidation. Sulfiding at 400~ of Co3Mo3N causes no significant changes in its specific activity and structure, although there is a decrease in surface area. At higher sulfiding temperatures (500-700~ it can be seen that this bimetallic nitride is converted first to Co9S8 and finally to MoS2 and Co9S8 (especially at 700~

Table 3 Catalytic activities and surface areas of CoMo nitride with increasing temperature of sulfidation. Sulfiding Phase b Surfacearea HDS conve., Specific activity, temp. (~ (m2/g) (%) (~mol/m2 min)

Fresha

C03Mo3N 19 13.4 155 400 Co3Mo3N+Co9S8 11 3.1 160 500 Co3Mo3N+Co9Ss+MoS2 13 19.5 327 600 Co3Mo3N+Co9S8 + MoS2 18 13.5 164 700 Co9S8+MoS2 20 10.3 112 a catalyst not previously sulfided with 10% H2S/H2. b listed in order of the height of the predominant peak for that phase in the XRD pattern. c determined atter 5 h of run time.

349 It should be noted that the specific activity of the catalyst sulfided at 500~ is very high (327 lxmol/m2 min). This appears to be due to a synergistic effect among the three phases present, i.e. Co3Mo3N, MoS2 and Co9S8. The specific activity of each phase is given in Table 4 for comparison. The synergy in activity is obvious for sulfidation of Co and Mo nitrides at 700~ and of Co, Mo, and CoMo nitrides at 500~ respectively.

Table 4 Comparison of catal~ctic activities of MoS2, C0988,C09S8+MOS2and Co3Mo3N. Phase Surfacearea HDS conversion 9 a Specific activity b (mE/g) (%) (lamol/m2 min) 3.4 MoS2 (from MOO3) 22 1.4 Co988 (from Co304) 9 13.4 Co3Mo3N (fresh) 19 10.3 Co9S8+MoS2(700~ c 20 19.5 Co3Mo3N+Co9Ss+MoS 2 (500~ c 13 a determined after 5 h of run time e lxmoles of reacted thiophene per surface area of catalysts per min b sulfidation temperature

34 35 155 112 327

4. CONCLUSIONS MoaN and CoMo nitrides were prepared by temperature-programmed reduction of the corresponding MoO3 and CoMoO4 in flowing ammonia. These catalyst were activity-tested in the thiophene HDS reaction, and their resistances to sulfidation were investigated. The bimetallic nitride phase is proposed to be C03Mo3N based on results from elemental and Rietveld analyses. Specific HDS activity of the CoMo nitride is higher than that for Mo2N. The properties of nitrides vary significantly with sulfidation temperature. During severe, hightemperature conditions of sulfidation Mo2N transforms into MoS2, and Co3Mo3N into MoS2 and Co9Ss. HDS conversion, surface area and specific activity of MOEN decrease significantly with sulfiding temperature due to transformation to MoS2. However, in the case of Co3Mo3N, a synergistically higher activity is observed for moderate temperatures of sulfiding relative to MoS2 and C09S8. Neither Mo2N nor C03Mo3N is resistant to sulfiding under high-concentration, hightemperature sulfiding conditions (i.e. 10% H2S/H2 and 500-700~ Thus, reaction over MOEN and CoaMoaN under typical process conditions, i.e., with sulfur-containing molecules such as thiophene, could transform these nitrides into sulfides. The transformation, however, could under moderate reaction conditions be beneficial to their catalytic performance due to possible synergistic effects among the nitride and sulfide phases present.

350 REFERENCES

1. 2. 3. 4. 5. 6. 7. 8. 9. 10. 11. 12. 13. 14.

Schlatter, J.C., S.T. Oyama, J.E. Metcalfe, III and J.M. Lambert, Jr, Ind. Eng.Chem.Res., (1988) 1648. Nagai, M., T. Miyao and T. Tuboi, Catal. Lett., 18, (1993) 9. Volpe, L., S.T. Oyama and M. Boudart in G. Poncelet, P. Grange and P.A. Jacobs(Editors), Preparation of Catalysts III, Elsevier, Amsterdam, 1983, p. 147. Markel, E.J., and J.W. Van Zee, J. Catal., 126, (1990) 643. Choi, J.G., J.R. Brenner, C.W. Coiling, B.G. Demczyk, J.L. Dunning and L.T. Thompson, Catal. Today, 15, (1992) 201. Moon, Y.H., and S.K. Ihm, Catal. Lett., 22, (1993) 205. McCarty, K.F., J.W. Anderegg and G.L. Schrader, J. Catal., 93, (1985) 375. Kim, D.W., S.K. Ihm and D.K. Lee, Catal. Lett., 43, (1997) 91. Zachwieja, U., and H. Jacobs, Eur. J. Solid State Inorg. Chem., 28 (1991) 1055. Bem, D.S., and H.-C. zur Loye, J. Solid State Chem., 104 (1993) 467. Bem, D.S., C.P. Gibson and H.-C. zur Loye, Chem. Mater., 5 (1993) 397. Yu, C.C., and S.T. Oyama, J. Solid State Chem., 116 (1995) 205. Izumi, F., "A Software Package for the Rietveld Analysis of X-ray and Neutron Diffraction Patterns", KEK Report, National Laboratory for High Energy Physics, Tsukuba

9

Elsevier Science B.V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

C a t a l y s t D e a c t i v a t i o n b y Metals and C o k e D u r i n g H y d r o d e m e t a l l a t i o n M. NOfiez, E. G6mez, C. Suarez and E. Carmona ECOPETROL-ICP (Colombian Petroleum Institute). Autopista Bucaramanga - Piedecuesta Km 7. A.A. 4185 Bucaramanga, Colombia. E-Mail [email protected]

The methods for the study of metals and coke poisoning of hydrotreatment catalysts normally require long and expensive tests in pilot plants. The Colombian Petroleum Institute. ICP, has developed a methodology to age catalyst samples in a commercial reactor for their subsequent analysis and study in short pilot plant testing. Because the conditions inside the reactor change spatially the study of this catalyst samples provides good information about the relationship between the different variables involved in catalyst design, the quality of the feedstock and the operating temperature.

1. INTRODUCTION Previous investigations have demonstrated that the hydrotreatment catalyst loses its activity with the poisoning caused by metal and coke deposits [ 1,2]. Metal deposition occurs slowly during long periods covering the internal surface of the pores, this is speeded up when these deposits plug the mouth of the pores and block out the access to the inside [3,4]. Coke poisoning occurs through different physical and chemical mechanisms, the latter ones are closely related to temperature, so that during the first working hours of the catalyst there is a strong deactivation caused by the active surface dirtying. This deactivation subsequently continues but at a lower rate as the result of a series of temperature stimulated chemical reactions of the coke. [5,6]. Commercial reactors present a temperature profile due to the heat generated by exothermic reactions which originate many different conditions along the reactor, its control is therefore vital for the good operation of a fixed bed reactor system.

2. EXPERIMENTAL PHASE This research was carried out in a system of two trickle-bed reactor for hydrotreatment of deasphalted vacuum bottoms with a capacity of 22.000 b/d at 1.500 psi and a maximum temperature of 425 ~ The quality of the inflow and outflow streams are given in Table 1. At the start of the run the catalyst has a high activity and can obtain the quality of the product at a relatively low temperature, but as the catalyst loses activity is necessary to increase the inlet temperature in order to maintain the quality of the products.

352 Since the reactions are exothermic a temperature profile is generated, being higher at the discharge. When the maximum temperature design is reached, it is necessary to change the catalyst since it cannot maintain the quality of the products. Each run lasts approximately 8 months. Table 1 Quality of the inflow and outflow streams in the studied system Inflow S Ni V API CCR

1.3 % 11.5 ppm 12.5 ppm 16.5 4.2

Outflow 0.3% 1 ppm 1.5 ppm 18.5 2.2

The two reactors are disposed in series. The first reactor is loaded with a big pore diameter demetallation catalyst designed to withstand high metal poisoning; the second reactor is loaded with a smaller pore size desulfurization catalyst which gives a larger reaction area and therefore, a larger hydrotreatment capacity. The reactor's optimization lead to the development of sampling methods which in turn ended in the implementation of a system by which a set of previously identified baskets containing samples of catalysts, are introduced into different spots during the loading process of the reactor. The hydrodesulfurization and hydrodemetallation reactions are considered as the most important reactions for the objectives of the plant. The first one generates H2S which comes out of the reactor with the gaseous stream while the latter ones originate metallic deposits that accumulate inside the catalyst. These deposits which act as a poison constitute a measurement of the accumulated work of the catalyst, their analysis and pilot plant tests allow to obtain information which contributes to a better understanding of the poisoning and aging mechanisms of these systems. Table 2 illustrates some characteristics of the catalysts introduced into the metallic baskets. Catalysts A, C, D and E that have big pore diameter are designed to withstand higher metal poisoning. Catalyst B by virtue of its smaller pore diameter withstands less metal poisoning but has a larger internal area and therefore is more active. Catalysts A and B are the same used in the industrial reactor. The first reactor is loaded with catalyst A and the second reactor with catalyst B. 2.1. Parallel aging of two catalysts Baskets were filled placing inside samples of Catalyst A and B. The baskets were stored in different locations throughout the two reactors so that they would be exposed to different operational conditions regarding to the temperature and feedstock quality. As the charge moves along in,,'de the reactors there is a decrease in its metal and sulfur contents and an

353 Table 2 Some characteristics of the catalysts studied. Catalyst A B C D E

Dp (/~) 190 90 140 120 120

DP (mm)

Ni(%)

2.2 1.5 1.8 1.4 0.9

1.2 1.3 1.8 1.0 0.7

Mo(%) 7.5 9.5 9.0 3.8 2.0

Dp (A): Pore diameter in Angstroms. DP (mm): Equivalent diameter in mm.

increase in the temperature as the result of the exothermic reactions. Samples recovered at the end of the 8 months run were analyzed to determine the metals content as shown in Figure 1. It can be observed in this figure that the points located around the region of high metal retention show a higher ratio of metals retained by the Catalyst B over the catalyst A. As seen in Equation 1 [7] the diffusion is favored by larger pore diameter and smaller molecule size.

De - Dm~ {OCM[l - (rmol/rM )] 4 nt- g m [ 1 - Cr.ollr.)] 4) T

(1)

where: Dmol

= Normal diffusivity coefficient.

c M = Macroporosity given as volume of macropores per volume of particles. em = Microporosity given as volume of micropores per volume of particles. r r~o l

rM rm

= Tortuosity. = Mean radius of molecule. = Mean radius ofmacropore. = Mean radius ofmicropore.

At the inlet of the reactor there is a mixture of molecules probably with different sizes so that the reactions are more selective towards the smaller ones. As the reaction continues there is a concentration of the larger molecules left unreacted which makes that the catalyst with the big pore diameter increase its performance if compared to the smaller pore size catalyst in the bottom of the second reactor. Total coke deposits in the aged samples, determined after 24 hours of Soxhlet extraction with toluene were approximately 8%, showing no clear tendencies that could relate these deposits to the catalyst nature or to the temperature. This fact does not exclude the possibility that differences in the quality of the deposits could be found using other techniques [5].

354 2.2. A g i n g of Catalysts in T w o Cycles The samples of catalysts aged during one complete run stored at the inlet of the first reactor were aged in a second run in order to determine if they still continue retaining metals. Table 3 illustrates the results of this experience.

16 o~ 14>., o

12108

o9

6

_J

<

4

m

2 0 0

I

I

I

I

5

10

15

2O

25

METALS (%)A Catalyst Figure 1. Relationship between metals retained by two catalysts with different texture

Table 3 Metals retained by different catalysts at the end of the first and the second cycle Catalyst

First cycle (Ni+V)

Second cycle (Ni+V)

A B C D E

25.2 16.8 19.8 30.6 31.2

43.2 32.4 36.4 54.8 57.4

In every case it is observed that the catalysts still preserve their metal retention capacity which means that the pores have not been blocked by plugging. The fact that these catalysts continue retaining metals under the plant normal operational conditions suggest that the limitation in the performance of the plant at the end of the run is mainly originated by the nature of the coke deposits.

355 2.3. Pilot Plant Tests

In order to know the catalysts activity over different stages of aging some tests were carried out on a fixed bed pilot plant. The test were carried out at industrial operational conditions related with feedstock quality, LHSV and pressure at different temperatures. Hereafter the tests made showing the information obtained in each one of them are described briefly: -The quality of the product at the outlet of the first reactor at the start of the run at different temperatures was determined with the flesh demetallation catalyst and the spatial velocity of the first industrial reactor (1F). -The quality of the product at the outlet of the first reactor at the end of the run and at different temperatures was determined using the aged demetallation catalyst with the spatial velocity of the first industrial reactor (1A) -The quality of the product at the outlet of the second reactor at the start of the run and at different temperatures was determined using a bed composed of the flesh demetallation and desulphuration catalysts with the spatial velocity of the two reactors (1F2F) -The contribution of the catalyst in the second reactor at the end of the run and at different temperatures was determined using a bed composed of an unused demetallation catalyst and an aged desulfurization catalyst (1F2A).

3. ANALYSIS OF THE RESULTS Figures 1 and 2 show the metals and sulfur contents of the products obtained in the pilot plant. Curve 1A corresponds to a bed simulating only the first reactor with aged demetallation catalyst; 1F corresponds to the first reactor with flesh catalyst, 1F2F corresponds to the first and second reactor with flesh catalyst, 1F2A corresponds to the first reactor with flesh catalyst and the second one with aged desulfurization catalyst. Comparing curves 1A and 1 F, i.e. the quality of the product at the outlet of the first reactor and at the end of the cycle, we observe that the quality of the product can be maintained by increasing the temperature without exceeding the design temperature (425~ The hydrotreatment levels obtained at 350~ with the flesh catalyst are obtained at approximately 385~ with the used desulfurization catalyst and at 390~ for the demetallation catalyst, this indicates that the catalyst has not yet collapsed due to plugging of its pores by metallic deposits, confirming the results obtained by aging the samples of this catalyst in a second cycle. Comparing the curves obtained with the flesh demetallation catalyst (1F) in this same catalyst plus the aged bed of the second reactor (1F2A), we can observe that the quality improvement obtained by the addition of the second catalyst is very little and is not representative if compared against the one obtained with flesh beds (1F2F). This confirms the fact that the catalyst in the second reactor is completely exhausted.

356 12 10 13_ v

03 ._1

<: I-iii

-,--. 1A +IF

0 320

340

360

380

400

1F2A •

1F2F

TEMPERATURE (~

Figure 2. Metal contents in the outflow obtained in pilot plant for different combinations of catalytic beds with aged and fresh catalysts.

1.2 1 -~

--'~-- 1A

0.8

--.i9 1F

0.6

1F2A

0.4

-M 9 - 1F2F

0.2 0 320

340

360

380

400

TEMPERATURE (*C)

Figure 3. Sulfur Contents in the outflow obtained in pilot plant for different combinations of catalytic beds with aged and fresh catalysts. Table 4 shows the metals contents for catalyst B samples on different locations inside the catalytic system. Since metal levels in the second reactor are too low if compared with the real capacity to assimilate metals shown in the first reactor, then it is concluded that the exhaustion is produced by coke deposition either in quantity or quality. In this case these deposits are

357 related to higher temperature on the second reactor and not subjected to the charge quality since the feedstock has been improved as it goes inside the reactors.

Table 4 Metals retained by samples of catalyst B aged at different locations inside the industrial reactors Locations inside reactors

Ni+V(%)

Upper part of the first reactor Middle part of the first reactor Lower part of the first reactor Upper part of the second reactor Middle part of the second reactor Lower part of the second reactor

15.8 15.3 12.4 9.8 6.1 1.5

4. CONCLUSIONS In the system studied, the duration of the run is limited by coke poisoning which is related to the high temperature of the second reactor. There was not found a clear relationship between temperature and extent of the deposits, through the techniques used; so it seems that quality more than quantity of the coke deposited determine the catalyst activity lost which must be associated to the temperature under which the reactions are carried out. Apparently there is not a clear relationship between the catalyst nature and the coke poisoning susceptibility but the tendency to retain and withstand higher metal poisoning of some catalysts over others is clear. Catalysts with pore size near to 120 A present the maximum metals retention for this type of feeds, but a better definition of the texture needs more information and is the reason of further research at ICP.

REFERENCES 1. 2. 3. 4. 5. 6.

S. Kobayashi, Ind. Eng. Chem. Res., 26 (1987) 2241. S. Kobayashi, Ind. Eng. Chem. Res., 26 (1987) 2245. A. Bum-Jong, AICHE J., 30 (1984) 739. E. Newson, Ind. Eng. Chem. Process Des. Develop., 14 (1975) 27. N.O. Egiebor, Appl. Catal., 55 (1989) 81. J. Bartholdg, Symp. Resid Upgrading, Div. Petroleum Chem. Inc., 205th National Meeting, American Chem. Soc., Denver, CO, March 28- April 2, 1993,386. 7. C. Pereira, Ind. Eng. Chem. Res., 29 (1990) 512.

This Page Intentionally Left Blank

9

Elsevier Science B.V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

359

Deactivation of Pt-Sn/A1203 Catalysts by Coking: Influence of the Preparation Method. G. Corro, P. Marecot, J. Barbier Universit6 de Poitiers, URA CNRS 350, Laboratoire de Catalyse en Chimie Organique, 40 avenue du Recteur Pineau, 86022 Poitiers Crdex, France

Bimetallic Pt-Sn/A1203 catalysts were prepared either by coimpregnation (CI) or successive impregnations (SI)of platinum and tin. In the SI method, the tin salt solution is introduced on a prereduced parent Pt/A1203 catalyst under hydrogen bubbling. SI catalysts are less sensitive than CI catalysts to deactivation by carbon deposition for isobutane dehydrogenation and coking with cyclopentane. A more effective interaction between the two metals is responsible for the higher stability of the SI catalysts.

1. INTRODUCTION Deactivation of catalysts by carbonaceous deposits is an important technological problem in the oil reforming and petrochemical industry. Bimetallic catalysts based on platinum and tin, generally supported on y--alumina, are widely used in the dehydrogenation of alkanes and in the reforming of hydrocarbons [1-3]. The main purpose in adding tin to a platinum catalyst is to increase the selectivity and the stability towards coke formation. It has been shown that addition of tin inhibits the formation of highly dehydrogenated hydrocarbon species required for coking, isomerization and hydrogenolysis reactions [4,5]. The beneficial effect of tin results from the formation of platinum tin alloys and in the decrease in platinum ensemble size required for coking and hydrogenolysis reactions [5,6]. The attenuation of these reactions enhances the selectivity of the dehydrogenation reactions that can proceed over small ensembles of surface platinum atoms. However, platinum-tin interactions depend strongly on the preparation and activation procedures of bimetallic catalysts. The main techniques used in the preparation of these catalysts are the impregnation of platinum derivatives on alumina-supported tin oxide, the impregnation of tin derivatives on alumina-supported platinum and the coimpregnation of metallic precursors on the support. In the present work, bimetallic catalysts were prepared by successive impregnations of platinum and tin precursors (chloroplatinic acid, stannic chloride) with an intermediary reduction step. Moreover, the tin precursor is introduced under hydrogen bubbling in the hope of reducing part of the tin chlorides on platinum particles. A second series of Pt-Sn/A1203 catalysts

360 was prepared by the classical method of coimpregnation. Stabilities and selectivities of the two series of catalysts were studied for isobutane dehydrogenation. The various catalysts were also characterized by hydrogen-oxygen titration and by their activities for different reactions: cyclopentane hydrogenolysis, cyclohexane dehydrogenation and coking reaction with cyclopentane.

2. EXPERIMENTAL 2.1. Catalyst preparation The support (T-alumina, surface area 246 m2/g) was ground and then sieved to retain particles with sizes between 0.25-0.10 mm. A monometallic platinum/alumina catalyst was prepared by impregnation using chloroplatinic acid in an acidic medium (0.2M HC1). After drying at 120~ overnight, the catalyst was calcined in flowing air for 4 h at 450~ and reduced in flowing hydrogen for 8 h at 500~ Bimetallic catalysts were obtained by coimpregnation of the metal precursor salts (H2PtCIt, SnCI4) on the support. After drying at 120~ overnight, the catalysts were finally activated by reduction in pure hydrogen for 8 h at 500~ A series of bimetallic catalysts was also prepared by successive impregnations (SI) of the metal precursors, with an intermediary reduction step. Thus, a sample of monometallic Pt catalyst previously activated by calcination-reduction was immersed in an acidic medium (HCI) and submitted to hydrogen bubbling. After 1 h, an acidic solution containing the stannic chloride, previously degassed under a nitrogen stream, was added to the hydrogen bubbling mixture. Catalysts were then filtered, dried overnight in air at 120~ and reduced under hydrogen for 8 h at 500~ A blank Pt/A1203 catalyst was prepared following the same procedure without addition of the tin salt. 2.2. Metal accessibility measurements Platinum accessibility was obtained by the H2-O2 titration method at room temperature in a static volumetric apparatus. Isotherms were obtained in the 0-50 torr range. Extrapolation to the origin value was used to calculate the number of exposed atoms. 2.3. Catalytic tests All samples were reduced again at 500~ under hydrogen for two hours before the catalytic tests. Turnover frequencies (TOF) of the catalysts were calculated from initial rate data. A fixed bed tubular continuous axial flow reactor was used for all reactions. Cyclohexane dehydrogenation. The reaction conditions were as follows: temperature: 270~ 0.025 g of catalyst; 0.97 bar of hydrogen and 0.03 bar of cyclohexane. Cyclopentane hydrogenolysis. Catalysts were tested with cyclopentane hydrogenolysis under the following conditions: 0.9 bar hydrogen and 0.1 bar of cyclopentane at 290~ lsobutane dehydrogenation. The amount of catalyst loaded into the reactor was 0.025 g. The reactant mixture consisted of 75 torr hydrogen, 25 torr isobutane and a total pressure of 760 torr. Coking with cyclopentane. This reaction was carried out under 0.9 bar hydrogen and 0.1 bar cyclopentane at 500~ for one hour. Coke deposition was analyzed by temperature-programmed oxidation (TPO) following the technique described earlier [6].

361 3. RESULTS AND DISCUSSION Table 1 summarizes the various catalysts prepared in this study. The tin loading on bimetallic catalysts prepared by successive impregnations was varied by using different concentrations of the tin solution introduced on the parent catalyst. Moreover, in Table 2, comparison of catalysts 11 and 14 which were obtained under hydrogen and nitrogen streams respectively, shows that hydrogen slightly improves the deposition of tin on the parent catalyst. Thus, before catalyst activation, most of the tin deposited would be localized on the support and only a small part would be reduced on the platinum surface by hydrogen. This large deposition of tin on the support is in agreement with previous studies showing the formation of surface complexes during the impregnation of stannic chloride on alumina [7]. Table 1 Summary of the different catal~csts studied Catalyst

Pt loading (wt%)

Sn loading (wt%)

1 2

0.6 0.6

. . ---

3 4 5 6

0.6 0.6 0.6 0.6

0.5 1.0 2.5 5.0

coimpregnation coimpregnation coimpregnation coimpregnation

7 8 9 10 11 12 13 14

0.6 0.6 0.6 0.6 0.6 0.6 0.6 0.6

0.1 0.2 0.3 0.8 1.5 1.7 5.0 1.2

SI (H2) SI (H2) SI (H2) SI (HE) SI (HE) SI (HE) SI (HE) SI (N2)

.

.

Preparation procedure a .

. blank

a: SI = Successiveimpregnations Table 2 Amount of Sn deposited as a function of Sn concentration in the solution Catalyst [Sn] solution (g 1"i) wt% Sn catalyst Atmosphere 7 8 9 10 11 12 13 14 ,,,

0.18 0.36 0.73 1.46 2.80 4.50 22.50 2.80

0.1 0.2 0.3 0.8 1.5 1.7 5.0 1.2

H2 H2 H2

H2 HE HE HE

N2

362 Figure 1 shows that the accessibility of platinum decreases as a function of the Sn/Pt atomic ratio whatever the method used for catalyst preparation. This evolution arises from the poisoning effect of tin interacting with the platinum surface, as has been reported earlier [4]. In this figure, one can see that tin poisoning is stronger on catalysts prepared by successive impregnations under the hydrogen than on those prepared by coimpregnation. Likewise, for the same Sn/Pt atomic ratio, the platinum accessibility is less affected for catalysts prepared by successive impregnations under nitrogen (catalyst 14). This result shows the prominent role of hydrogen in the interaction between platinum and tin in the course of the preparation of bimetallic catalysts by the SI method. The different catalysts prepared in this study were characterized by their catalytic activities for cyclohexane and cyclopentane hydrogenolysis which, in reaction on platinum, are generally considered to be structure-insensitive and structure-sensitive, respectively. ..

14

e~0 oo

10 ~

6 -

N

4

9

0 0

5

Sn/Pt

10

15

Fig. 1. Accessible Pt atoms x 1018/g cat. as a function of S~Pt atomic ratio in catalysts prepared by: m coimpregnation; 9 SI; A Pt:AlzO3; X SI without Hz. In Figure 2, TOF for cyclohexane dehydrogenation is shown as a function of accessible platinum atoms for the two series of Pt-Sn catalysts. If cyclohexane dehydrogenation were a structure insensitive reaction on platinum-tin catalysts, one would expect a constant TOF independent of platinum surface accessibility. However, activity was not determined as a function of Pt surface concentration for this reaction, although TOF was dependent on both Sn/Pt atomic ratio and preparation procedure. In general, the TOF values for cyclohexane dehydrogenation are higher on the SI catalysts than on the coimpregnated ones for the same Sn/Pt ratio. When tin content is higher, a negligible catalytic activity is detected even if the catalyst is still active for hydrogen chemisorption. Results obtained from cyclopentane hydrogenolysis are reported in Figure 3. In this figure, cyclopentane conversion decreases drastically as the S l ~ t atomic ratio increases. Moreover, a complete suppression of the catalytic activity is observed on SI platinum-tin catalysts while this effect is not found on catalysts prepared by coimpregnation. It is interesting to note that catalyst 14 (SI without hydrogen) presents a similar behavior to that observed on coimpregnated PtSrgA1203 catalysts. These results show that the method of successive impregnations under hydrogen would be more appropriate to divide large platinum ensembles which are responsible for hydrogenolysis reactions.

363 10000

8OO0 r~

6000

x,

[.., 2OOO

0 0

2

4

6

8

10

ace. Pt atoms x 10"18 gcat"1

12

14

Fig. 2. Turnover frequencies for cyclohexane dehydrogenation (benzene production) on PtSrl/AI203 catalysts as a function of accessible platinum atoms x 10-18/g cat. Catalysts prepared by: II coimpregnation; 9 SI; A PVA1203 blank;

A

A v

1

1

=o 3 ~ r~

o r

v

0

A v

,~a "v~

I

5

10

v

15

Sn/Pt Fig. 3. Cyclopentane conversion in n-pentane as a function of Sn/Pt atomic ratio. Pt-Sn/AI203 prepared by: 1 coimpregnation; 9 SI; &, Pt/AI203; X SI without hydrogen. The stability and the resistance to coke deposition of the different catalysts were studied using the reactions of isobutane dehydrogenation and coking with cyclopentane at 500~ Results in Figure 4 show that turnover frequencies for production of isobutene increase with Sn/Pt atomic ratio. Moreover, catalysts prepared by coimpregnation exhibit lower TOF values than those prepared by the SI method. Figures 5 and 6 present percent conversions for isomerization and hydrogenolysis reactions during isobutane dehydrogenation at 500~ as a function of Sn/Pt atomic ratio. These figures show for SI catalysts a drastic suppression of the transformation rate for these two secondary reactions as the tin content increases. For coimpregnated catalysts, the total suppression of the catalytic activity for these reactions was not found, even for high Sn loading.

364

~2-

9

r~

I

J

0 0

I

I

5

10

15

Sn/Pt Fig. 4. Isobutane dehydrogenation. Turnover frequencies for isobutene production over PtSn/AI203 catalysts as a function of Sn/Pt atomic ratio. Catalysts prepared by: m coimpregnation; 9 SI; APt/A1203 blank; X SI without hydrogen. 0.4 0

.~

0.3

0.2 0.1

9

0

v v

.v ,A

2

~

4

A v

I

6

!

Sn/Pt

!

8

10

!

12

A v

.

14

Fig. 5. Isobutane isomerization. % conversion to n-butane as a function of Sn/Pt atomic ratio. Catalysts prepared by: m coimpregnation; 9 SI; A Pt/A1203 blank; X SI without H2. 3 v------r---~ ,-..,~ 2 O

"

0

5

Sn/Pt

10

15

Fig. 6. Isobutane hydrogenolysis % products as a function of Sn/Pt atomic ratio. prepared by: m coimpregnation; 9 SI; A Pt/AI203 blank; X SI without H2.

Catalysts

365 Isobutane dehydrogenation activity was studied over a period of 24 h on selected catalysts. Figure 7 shows the evolution of percent conversion in isobutene at 500~ as a function of reaction time. In this figure, one can see that catalyst 12 prepared by SI method (1.7 wt% Sn) gives a higher conversion than catalyst 5 (2.5 wt% Sn) prepared by coimpregnation. After 24 h on stream, the quantitative determination of coke deposits performed by TPO shows that carbon deposition decreases following the sequence Pt/A1203 > coimpregnated Pt-Sn > SI Pt-Sn. Thus, the highest activity of the SI catalyst for isobutane dehydrogenation results from the lowest coke deposition on the catalyst. :30

O

2O

9 aam%

O r~

o~

Emmlmmmmm m m S

10

A==

0 0

~

400

600

~

1000

1~

1400

1600

Time (min) Fig. 7. Evolution of % conversion in isobutene as a function of time of reaction. II Pt-Sn/A1203 wt% 0.6-2.5 prepared by coimpregnation; 4 Pt-Sn/A1203 wt% 0.6-1.7 prepared by SI; A Pt/A1203 blank. The deactivation of the three same catalysts by carbon deposits was also studied in the course of the coking reaction with cyclopentane. Contrary to the results obtained for the isobutane dehydrogenation reaction, the carbon content of catalysts submitted to the cyclopentane reaction at 500~ for 1 h decreases following the sequence: SI Pt-Sn > coimpregnated Pt-Sn > Pt/A1203 (Table 3). The analysis of the TPO curves shows that the behavior of the SI catalyst arises Table 3 Carbon deposition Catalyst Pt-Sn/A1203 1.7 wt% Sn Pt-Sn/A1203 2.5 wt% Sn Pt/A12Os

Prepared by

%C a

%C b

SI

0.21

6.17

coimpreg. blank

0.51 0.90

5.47 4.77

a isobutanedehydrogenation bcoking with cyclopcntane

366 mainly from an increase of the amount of carbon oxidized at high temperature, i.e. an increase of the amount of coke deposited on the support [8]. This result is in accordance with the highest concentrations of cyclopentene and cyclopentadiene observed on the SI catalyst during the reaction. Indeed, coke on the support is produced on the acid sites by polymerization of dehydrogenated intermediates generated by the metallic function [9].

4. CONCLUSIONS In conclusion, bimetallic Pt-Sn/alumina catalysts prepared by successive impregnations with an intermediary reduction step and introduction of the tin salt (SnCI4) under hydrogen are less sensitive to coke deactivation than catalysts prepared by coimpregnation. This behavior probably results from a more effective interaction between the two metals, leading to smaller platinum ensembles, as evidenced by the low hydrogenolysis activity. However, the amount of coke deposited on the whole catalyst depends on the nature of the feed and therefore on the nature of the dehydrogenated species which are more or less active precursors for coke deposition on the support.

REFERENCES ~

2. 3. 4. 5. 6. .

8.

9.

S.J. Miller, U.S. Patent 4727216 (1986). T. Imai and C.W. Hung, U.S. Patent 4,430,517 (1983). Compagnie Fran~aise de Raffinage, French Patent 2031984. R.D. Cortright and J.A. Dumesic, J. Catal., 148(1994) 771. R.D. Cortright and J.A. Dumesic, J. Catal., 157 (1995)576. J. Barbier, E. Churin, J.M. Parera and J. Riviere, React. Kinet. Catal. Lett. 29 (1985)323. Y.X. Li, Y.F. Zhang and K.J. Klabunde, Langmuir, 4 (1988) 385. J. Barbier, P. Marecot, N. Martin, L. Elassal and R. Maurel, Catalysts Deactivation (B. Delmon and G.F. Froment, eds.) Elsevier, Amsterdam, 1980, 53. J. Barbier, Appl. Catal. 23 (1986) 225.

9

Elsevier Science B.V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

N e w D e v e l o p m e n t s in F C C Catalyst Deactivation by Metals: Metals M o b i l i t y and the V a n a d i u m M o b i l i t y Index (VMI) Lori T. Boock a'*, Joanne Deady b, Tow Foon Limc and George Yaluris ~, W.R. Grace & Co.-Conn., a 7500 Grace Drive, Columbia, MD 21044 b 10 East Baltimore Street, Baltimore, MD 21202 c W.R. Grace (Singapore, Pte Ltd.) 501 Orchard Rd, Singapore 0923

We have developed a laboratory procedure using FCC Ecat which quantifies the mobility of vanadium, relative to nickel, in a particular FCC unit. The procedure involves the separation of the Ecat into different density fractions, and subsequent measurement of the nickel and vanadium on each fraction. The results of these measurements are combined into what is referred to as a "vanadium mobility index" or VMI, which is a number between one and about 15. A VMI of less than five represents low mobility, whereas a VMI greater than eight represents high mobility. In this paper we present data on the VMI of a number of different FCC units, representing all types of catalysts and operation. We have found that units with high VMI tend to have more vanadium deactivation of their catalysts; however, vanadium trapping technologies perform better in these same units. We will present results which show that the commercial performance of a Davison vanadium trap is well correlated with the unit VMI. We have also examined what specific factors affect the mobility of metals in an FCC unit. Our work on VMI has shown large variation in vanadium mobility; however, we have not been able to clearly correlate changes in VMI with specific unit operating conditions, such as full vs. partial burn, regenerator configuration, catalyst addition rate, unit type or carbon on catalyst. In order to take a more rigorous approach to understanding the effect of regenerator conditions on vanadium mobility, we have undertaken a laboratory study to evaluate the effect many of these conditions have on metal mobility. We have designed a bench scale regenerator [ 1], which simulates the essential elements of operation of an FCC unit regenerator, and can be used to study metal mobility. Using this apparatus, we can control the important operating parameters of the regenerator, such as temperature, contact time, oxygen content and steam partial pressure, and evaluate their effect on vanadium mobility. In this paper we discuss our recent results on the effects of these parameters on the mobility of vanadium on Ecat and lab deactivated catalysts.

368 I. INTRODUCTION Nickel and vanadium are deposited on FCC catalysts from the feedstock. These metals do not remain in the state and location they are deposited, but typically change oxidation states and move throughout catalyst particles and from particle to particle during regeneration and reaction steps in the FCC unit. It is believed that metals mobility is highest when exposed to FCC regenerator conditions, where air and steam are present. Once flesh FCC catalyst has been exposed to feedstock containing Ni and V contaminants and cycled through the FCC unit a number of times, it is referred to as equilibrium catalyst or Ecat. The metals on the Ecat are now less active for catalyst destruction and dehydrogenation than when they were initially deposited. However, they still have some activity for unfavorable reactions and this activity varies with the metal level, age of the catalysts particles, conditions they were exposed to, and other factors. Thus, FCC Ecat is not a single entity, but is a conglomeration of catalyst particles at different stages of deactivation, all with different activities and selectivities. Vanadium, in particular, is a metal which has large effects on FCC activity and selectivities. Not only does vanadium migrate to and destroy the zeolite, but also under certain conditions, it has high dehydrogenation activity [2]. Recent work has shown that vanadium is most mobile in its fully oxidized state, the state in which it is most destructive towards the zeolite [3], and that there are many factors which affect vanadium mobility. Nickel, which also has dehydrogenation activity, is much less mobile than vanadium and does not attack and destroy the zeolite. Thus, from a refiners' perspective, it would be valuable to know how mobile the vanadium is in their FCC unit. In this paper, we use commercial Ecat and laboratory experiments on flesh catalyst to help answer the question of how mobile is vanadium and what are the factors that affect this mobility.

2. EXPERIMENTAL The procedure we have developed to determine the VMI of a particular FCC unit involves obtaining approximately 100 grams of an FCC Ecat. This Ecat is first calcined in a muffle furnace at 593~ for 2 hours to remove any residual carbon. After cooling, the Ecat is separated into eight fractions by a sink/float density separation. The density separation is performed by mixing the catalyst sample with tetrabromoethane (TBE) which has a density of 2.96 g/cc. This material is heavier than most FCC catalyst (excluding additive materials) so the catalyst sample will float in the TBE. Tetrochloroethane (TCE), which has a density of 1.58, is then added to the mixture to lower the density of the liquid and cause some of the denser catalyst particles to sink. The amount of added TCE is controlled to obtain the desired amount of sink and float fractions. The specific procedure for obtaining an eight fraction sink float involves placing 50g of catalyst in an 8 ounce glass jar with a lid. Approximately 60 g of TBE is added to the jar, which is shaken well to mix. After mixing, enough TCE is added to sink approximately half of the catalyst sample. This procedure is done with two 50 g samples and the sink and float from each sample are combined. After adding the TCE, the catalyst/liquid mixture is centrifuged for 10-15 minutes at 2000 rpm. Alter centrifuging, the jar is carefully removed and it is visually determined if an approximately 50/50 split has been obtained. If the desired split has not been

369 obtained, TBE or TCE is added to increase the sink or float fraction, respectively, and the sample is re-centrifuged. The sink fraction will be tightly packed in the bottom of the container, whereas the float fraction will be floating in or on top of the liquid. A piece of Whatman #2 filter paper is placed in a Buchner funnel, with the vacuum on, and the float portion of the sample is decanted into the funnel. Usually the float fraction can be liited while decanting the liquid. Care is taken not to disturb the sink fraction. The float fraction is rinsed twice with Hexane, then placed in an evaporating dish to dry in a hood overnight. The same procedure is then followed with the sink fraction. Both fractions are then heat treated in a muffle furnace for 3 hours at 538~ The above procedure is repeated with each fraction until a total of eight approximately equal fractions have been obtained. The skeletal density of each fraction can be calculated based on the amount of TBE and TCE which was used to float or sink the fraction. Each fraction is submitted for ICP analysis to determine the Ni and V on each fraction and this information is used to calculate a VMI, as described below. The same sink/float procedure can be used to remove additives, such as Davison's RV4+ vanadium trap from an Ecat sample; however, since most FCC additives are much denser than FCC catalyst, this separation can usually be performed using only TBE. To study the effects of the regenerator operating parameters we used the lab-scale system we recently developed for the study of the chemistry of the FCC unit regenerator [ 1]. This apparatus does not attempt to simulate all of the operation characteristics of refinery-scale regenerators, rather it is a one-catalyst pass system that operates as close as possible to the operation conditions of refinery regenerators and captures the chemical phenomena occurring in a true regenerator. It features a fluidized bed reactor, and the ability to control the catalyst residence time from 5 rain. to more than 2 h. It also allows us to control the steam and excess oxygen in the reactor, as well as the reactor temperature. We use nitrogen to achieve fluidization of the catalyst in the reactor (600 - 1100 sccm), while steam, oxygen or other gases are added in the reactor feed as needed. The catalyst flows continuously through the reactor using a custom-designed catalyst flow system that can accurately control the catalyst flow rate. Depending on the catalyst flow rate, the system can hold enough catalyst to operate for 8 - 24h. After the catalyst passes through the bed it is collected in a sealed container. A Horiba oxygen analyzer (MPA-510) is used to measure the oxygen in the reactor effluent, and through a feedback loop, a PID controller controls the amount of air or oxygen needed to maintain the desired excess oxygen in the effluent. All other reactor gas products are analyzed by an On-Line Technologies 2002 FTIR Multigas analyzer which can measure CO2, CO, NO, N 2 0 , NO2, SO2, H20 , HCN and other FTIR-active gases. Detailed mixing experiments have shown that the catalyst mixing patterns in the reactor are very close to those of a CSTR. We have studied V and Ni transport in two catalyst blends. The first blend contained 75 wt.% ECAT with 2400 ppm V and 2400 ppm Ni blended with 25 wt.% Davison RV4+ metals trap [4].The second blend contained 75 wt.% of a commercial lab-deactivated catalyst impregnated with 4900 ppm V by Davison's CPS deactivation procedure [2], mixed with 25 wt.% lab-deactivated Davison metals trap (RV4+). These catalyst blends were tested at various reactor temperatures (675-750~ catalyst average residence times in the reactor (14 120 rain), amounts of excess oxygen (0 to 3%), and steam levels in the reactor feed (0 to 30%). Samples of 20 - 40 g of catalyst were collected and separated by the sink-float method.

370 Subsequently both the "float" and "sink" fractions were submitted for elemental analysis by ICP.

3. RESULTS 3.1. V M I

As described above, an Ecat sample is separated by density into 8 fractions and the Ni and V content of each fraction is measured. The heavier fractions are higher in metals and represent older catalyst. The VMI for that particular Ecat, and FCC unit, is calculated by plotting the Ni/V of each fraction divided by the Ni/V of the base Ecat versus the density of the fraction. The slope of that line is a measurement of the relative mobility of V to Ni. Since Ni is assumed to be very immobile, the VMI can be taken as a measure of the vanadium mobility of a specific unit at a specific time. Figure 1 shows an example of a VMI plot for two units, one with high VMI and one with low VMI. We have calculated the VMI for many commercial FCC units using various regenerator types, operating conditions and catalyst. The results are summarized in Table 1. Based on the results in Table 1, there is no obvious controlling factor which leads to a high or low VMI. However, a few observations can be made which suggest some operating conditions which may affect VMI.

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Example of a Low VMI and a High VMI

In general, there are more partial combustion, high carbon on regenerated catalyst (CRC) units with lower VMI than higher VMI. Two stage, partial/full combustion regenerators seem to span the full range of VMIs, so clearly there are factors other than regenerator type and

371 operation which affect vanadium mobility. Metal level and unit type do not have any clear effect on unit VMI. Catalyst replacement rate, or more precisely, the percentage of the unit inventory which is replaced every day, does seem to effect VMI, with the units with a higher replacement rate having higher VMIs. This is also observed when examining the three VMI measurements on Refinery # 2, with the primary difference being that when the high VMI was measured, the unit was experiencing loss problems and more fresh catalyst was being added to the unit. This results suggests that fresher metals on catalyst have higher mobility and the older metals may have been passivated. In general, catalyst type does not appear to effect VMI, since units using identical catalysts may have very different VMIs. However, in certain cases, the catalyst may have an affect. Two examples are shown in Table 1, refinery # 19 and 20. Unit 19 switched from a base catalyst to a catalyst with a vanadium trap, and the VMI dropped dramatically. This result is not unexpected, since a vanadium trap, when operating properly, is expected to immobilize vanadium so it cannot destroy zeolite and produce coke and gas. Unit #20 was a unit starting up. Initially the VMI was low, but the inventory was largely start-up Ecat and the VMI being measured was likely a combination of the Unit #20 and the unit where the Ecat was from. However, the VMI increased as the Ecat lett the unit and the higher VMI is likely more representative of the actually unit V mobility.

3.2. Vanadium traps and V PUF The sink/float procedure can also be used to measure the performance of a separate particle vanadium trap, such as Davison's RV4+ trap. This performance measurement is characterized by a vanadium pick-up factor or V PUF. The V PUF is defined as the vanadium on the additive fraction divided by the vanadium on the catalyst fraction, and for an additive present in the unit at 4-6%, a V PUF of 5 or higher is considered to be good performance [4]. We have been able to show, both theoretically and experimentally, that the VMI of a particular unit corresponds with the V PUF. Figure 2 shows a plot of V PUF versus VMI for a number of commercial units where RV4+ was used. It is interesting to note that for some of the units, the V PUF was predicted based on the VMI before the RV4+ was added to the inventory, and the actual V PUF matched the predicted result. This implies that if a FCC unit is considering using a vanadium trap, that a VMI measurement can predict whether or not the trap would work well and therefore be economically effective. 3.3. Regenerator Test Unit (RTU) Our measurements of metals transport in the lab-scale regenerator unit show that for the Ecat no significant transport occurs at the conditions of our experiments, even when there is 30% steam in the reactor gases. No measurable metal transport was observed after steaming the same Ecat/metal trap sample with 100% steam at 732~ for 2 h. It appears that Ecat metals do not transport in any significant amounts within the time frame of our experiments, regardless of the operating conditions. This is consistent with the VMI results that show that older metals are much less mobile. Table 2 shows that in the absence of steam in the reactor, and at a catalyst residence time of 36 min (reactor temperature 715~ 1.1% oxygen at the reactor effluent) no vanadium transport is observed for the lab-deactivated

372

Table 1" A s u m m a r y o f Indices. Refinery Regenerator 1 full burn 2 full burn 3 partial burn 1(repeat) full burn 4 full burn 5 full bum 6 2-stage P/F 7 full burn 8 9 partial burn 10 full burn 11 partial bum 12 full burn 13 partial bum 14 2 stage P/F 15 partial burn 2 (repeat) full burn 2 (repeat) full burn 16 partial burn 17 partial burn 18 full burn 19 partial burn 20 2 stage P/F 21 full burn 22 partial burn 23 2 stage P/F 24 full burn 25 partial burn 26 full burn 27 full burn 20 (repeat) 2 stage P/F 20 (repeat) 2 stage P/F 28 partial burn 29 full burn 30 partial burn 31 partial burn 19 (repeat) partial burn 32 full burn 33 partial burn 34 2 stage P/F 35 partial burn 36 partial burn 37 2-stage P/F 38 partial burn

c o m m e r c i a l F C C o p e r a t i n g conditions and V a n a d i u m M o b i l i t y Ni/V 1100/1100 1300/3300 2220/2280 800/1400 3000/3700 434/1610 1400/3300 600/2100 400/3000 200/1400 1200/3000 2070/4570 1200/3100 1400/4890 3000/3800 508/1700 1400/3400 1800/3200 1630/2880 1770/3600 3141/3410 2400/3300 4800/4900 1700/2700 2300/3600 1900/6500 2600/4800 2700/6700 2000/2000 1300/2600 5200/5900 5200/5900 600/3600 500/1900 2700/5000 2200/4200 3100/2900 360/1900 1400/3300 3000/6000 2000/4000 730/1400 1600/6000 1000/2000

Catalyst Resoc Spectra Octacat Resoc Ramcat Spectra Octacat XP Octacat XP Orion Advance Super Resoc Spectra Vision Orion Spectra Spectra Orion Residcat Resoc Advance Orion LC Spectra Harmorex Octacat Harmorex Horizon Orion XP Orion LC Orion LC Advance XP Advance GO Resoc Octacat Vektor Vision Orion Octacat Orion Orion

VMI 12 12 12 9.7 9.1 8.7 8.7 8.4 7.9 7.5 7 6.9 6.7 6.6 6.5 5.8 5.7 5.7 5.3 5.1 5.1 5.1 5.9 4.8 4.5 3.9 3.8 3.7 3.4 3.4 3.3 3.2 3.2 3 2.9 2.8 2.7 2.5 2.2 2 1.8 1.4 1.4 1.2

CRC 0.02 0.04 0.11 0.02 0.03 0.03 0.04 0.02 0.05 0.12 0.03 0.1 0.06 0.02 0.15 0.03 0.03 0.14 0.07 0.04 0.1 0.02 0.11 0.16 0.06 0.03 0.2 0.05 0.06 0.02 0.01 0.24 0.05 0.39 0.21 0.11 0.03 0.31 0.19 0.07 0.03 0.15

Inventory 110 132 544 110 220 88 200 240 380 150 85 140 496 198 386 350 132 132 165 827 364 193 882 77 95 70 500 325 882 882 320 140 154 260 193 500 450 600 270 50 165 320

T/D 5 6 20 5.5 6.5 3.5 11 5 4 3 2.5

%/Day 4.55 4.55 3.68 5.00 2.95 3.98 5.50 2.08 1.05 2.00 2.94

8

1.61

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3.89 1.71 1.52 1.06

8 6.6 9 33 2.2 3 1.5 8 10 12 2.5 33 22 5 2 6 8 9 7 6 8 3.5 1 1.5 10

0.97 1.81 4.66 3.74 2.86 3.16 2.14 2.00 3.69 3.74 2.49 1.56 1.43 3.90 3.08 4.66 1.40 1.33 1.33 1.30 2.00 0.91 3.13

373

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2

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4

~

6

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VMI

Figure 2.

Performance of a commercial V Trap: VMI versus V PUF

Table 2. Vanadium transport to CPS deactivated metals trap from a commercial FCC catalyst, blended at 75 wt.% catalyst and 25 wt.% trap. The FCC catalyst contains 4900 ppm V. All values reportedare in ppm V.

Sample

Total V

FCC Catalyst Metals Trap 75 wt% Catalyst- 25 wt.% Trap Catalyst Blend Steamed 2 h/1005 K 0% water in gas feed 5% water 15% water 27% water 27% water

4890 80 3690 3420

Float Fraction (Catalyst)

Sink Fraction (Trap)

4360 4620 4550 4470 4630 4520

740 60 280 330 320 340

%'~/ Transported

4.5 o.4 2.o 2.3 2.4 2.0

catalyst, either. However, in the presence of steam, some metal transport can be observed from this lab-deactivated catalyst, where the metals tend to be more active and probably not as tied down as in the actual Ecat. The amounts of V transported at 5% and 30% steam in the reactor are virtually the same. Thus, the amount of V transported does not appear to be correlated to the concentration of steam in the reactor gas. When the same catalyst blend was steamed in a fluidized bed steamer for 2 h at 100% steam and at 732~ more than a two fold increase of V transport was measured. We believe that the larger amount of vanadium on the metal trap is due to the longer contact time and the higher reactor temperature rather than the increased steam concentration. These results are in agreement with our understanding that

374 steam, longer contact times, and the state of metals on the catalyst contribute greatly to the mobility of metals [2, 3].

REFERENCES

1. G.Yaluris, A.W. Peters and X. Zhao, Impact of the Clean Air Act on Fuels Production and Use, 212th ACS National Meeting, Orlando, FL, August 25-29, 1996, Vol 41(3), ACS, Washington, DC, 1996. 2. L.T.Boock, T. F. Petti, and J. A. Rudesill, Ch. 12, Cyclic Propylene Steaming of FCC Catalysts, in Deactivation and Testing of Hydrocarbon-Processing Catalysts, P. O'Connor, T. Takatsuka and G. L. Woolery, (Eds.), ACS, Washington, DC, 1996. 3. R.F.Wormsbecher, W.-C. Cheng, G. Kim and R.H. Harding, Ch. 21, Vanadium Mobility in Fluid Catalytic Cracking, in Deactivation and Testing of Hydrocarbon-Processing Catalysts, P. O'Connor, T. Takatsuka and G. L. Woolery (Eds.), ACS, Washington, DC, 1996. 4. T.J.Dougan, U. Alkemede, L. T. Boock and B. Lakhanpal, 1994 NPRA Annual Meeting, San Antonio, TX, Paper AM94-46.

9 Elsevier Science B. V. All rights reserved Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

375

Stability o f an F C C C a t a l y s t M a t r i x for P r o c e s s i n g Gas Oil w i t h R e s i d P. Gamero M., C. Maldonado M., J. C. Moreno M., O. Guzman M., E. Mojica M., R. Gonzalez S. Instituto Mexicano del Petroleo. Eje Central Lazaro Cardenas 152. P. O. Box 14-805, Mexico City, Mex. 07730

Silico-aluminates prepared from kaolinite were investigated as potential FCC matrix additives for use in upgrading of heavy fractions of gas oil and residue. These materials have lower deactivation rates compared to commercial FCC catalyst matrices. By inclusion of these materials in FCC catalysts, surface area and area retention of the matrix are increased. In consequence of their high activities and s~bilities, these matrix additives have promise for developing more efficient FCC catalysts for cracking of bottoms to high quality gasoline and highly olefinic LPG.

1. INTRODUCTION FCC catalysts are typically formulated with a Y-zeolite in a silica-alumina matrix. Catalytic properties are strongly dependent on the nature and composition of the zeolite and the matrix. An important role of the matrix is to provide acid sites for primary cracking of large molecules to molecules of intermediate size which can be subsequently cracked by the internal sites of the zeolite to desirable products [ 1]. The matrix is formed from a combination of active silica-aluminas, aluminas, binders and clays. The clay is commonly kaolin, which contributes little to the surface area and activity of the catalyst. The binder is used to improve mechanical properties required to retain the FCC catalyst in the industrial unit. Active alumina and silico-aluminates facilitate the production of high-value products from barrel bottoms. Although alumina or silica-alumina is an altemative to increasing meso- and macroporosity, under regenerator conditions they tend to progressively lose surface area, porosity and active sites [2]. It is, therefore, important to include high-stability materials. Under FCC regeneration conditions, the zeolite is quickly deactivated and, at the same time, its unit cell size is modified. The commercial FCC unit has a catalyst inventory made up of a mixture of particles of different ages. The newer particles contribute most of the activity to convert the gas oil. A part of the catalyst inventory has a low zeolite activity, and, notwithstanding, it has a substantial activity in the matrix, because the matrix is more stable than the zeolite. The oldest fraction is responsible for providing the activity to crack the heavy fraction of the feed. It is imperative to pay special attention to the design of the matrix and specifically to its pore size distribution, surface area, activity and stability. As alternative matrix additives for FCC catalysts, several silico-aluminates with controlled structural properties were prepared and characterized. These silico-aluminates have the potential to crack high molecular weight hydrocarbons without a loss of selectivity. In this

376 paper, experimental measurements of the structural, deactivation and activity properties of these new experimental silico-aluminates obtained from kaolinite are reported and discussed.

2. EXPERIMENTAL A series of controlled-texture silico-aluminates were obtained by heat and chemical treatments applied to kaolinite. The main variables of the study were calcination temperature of the kaolinite, the leaching agent and its concentration, reaction temperature and time. The procedure for preparing silico-alurninates of controlled pore-structure is discussed by way of example: Fresh kaolinite clay or clay calcined at 600 to 1200~ was subjected to a milling process. The resulting material was reacted with 0.2 to 1.6 g of sodium hydroxide per gram of kaolin under reflux conditions and then dried in a Niro Atomizer Mobile Minor spray dryer under the following conditions: inlet temperature, 350~ outlet temperature, 140~ air pressure, 0.8 Bar; feed rate, 120 ml/min; and solids concentration in the feedstock, 30 wt. %. The micro-spheroidal product was found to have an average particle size of 50 to 60 microns. The main control parameters in the characterization of experimental and reference materials were textural, structural and catalytic properties, the last involving the ability to crack modified gas oil with 5 wt.% residue. Textural properties of fresh and deactivated samples were determined using an ASAP 2000 analyzer (Micromeritics). To determine total surface area, the BET equation was applied to adsorption uptakes obtained in the relative pressure interval, P/Po, of 0.01 to 0.05. Zeolite and matrix surface areas were determined according to the ASTM D-4365 method, using P/Po pressures of 0.01 to 0.60. Mesopore volume and pore-size distribution were determined according to the ASTM D-4222 method at P/Po relative pressures of 0.01 to 0.99. Chemical structure was analyzed using a SIEMENS D-500 X-Ray diffractometer with a secondary beam mono-chromator and software for the optimization of the equipment and data processing. Methods to determine the cell parameter and crystallinity in zeolites were ASTM D-3942-90 and ASTM D-3906-85 respectively. Catalytic activity of synthesized and reference materials was evaluated using the ASTM D-3907 method. Gas and liquid streams were obtained as products of the reaction. The gaseous fraction was analyzed by on-line gas chromatography using a Hewlett Packard, Model 5890, Series II chromatograph equipped with two 13 X molecular mesh columns, two Sebaconitrile columns and one Porapak column. The liquid fraction was analyzed in a Hewlett Packard, Model 5890, Series II chromatograph, using an UCW-98 phase column, to obtain simulated distillation information. In the determination of coke content a stream of air was passed through the hot catalyst (650~ and combustion products were analyzed by IR.

3. RESULTS AND DISCUSSION Matrix area of commercial catalyst before and after steaming was determined to define their potential for promoting high molecular weight hydrocarbon cracking. Results are summarized in Table 1 for catalyst in which zeolite was added either in situ or by integration during preparation of the catalyst spheres.

377 Table 1. Matrix surface areas for commercial catalysts before and after deactivation at 788 ~ 4 h, 100 % Steam D E FA In Situ a In Situ a Integrated b Integrated b Integrated b Areas after h~cdrothermal treatment Total, m2/g 166 / 172 161 182 188 277 Zeolite, m2/g 121 152 74 135 77,/100 Matrix, m2/g 40 108 36 142 Steamed catalyst area retention Total, % 71 73 65 73 73 64 Zeolite, % 64 67 68 70 76 72 75 81 55 Matrix, % 60 80 57 a Process in which the zeolite is crystallized in situ in a natural clay-derived matrix. b Zeolite is prepared first and subsequently incorporated into a matrix. Catalyst Method

A In Situ a

Surface areas of zeolites prepared by in-situ crystallization are greater than in the catalysts obtained by integration of components. Zeolite area retention is also greater for the integrated catalysts. This difference in stability may be attributed to compositional variables, such as zeolite type and rare earth content. The in situ catalysts, after having been hydro-thermally deactivated, retain 79% of the matrix area as an average, while the integrated catalysts retains on the average 57%. Of the total area, matrix area in the in situ catalysts and the integration catalysts comprises 49% and 35 % respectively. Pore size distributions of the commercial catalysts are shown in Figure 1. For comparison purposes, a catalyst specific for heavy feed cracking (Integrated FA) was included. The data in Fig. 1 show that these catalysts including the Integrated FA catalyst have meso- and macropore fraction which are relatively low. Accordingly, high molecular weight hydrocarbons have a low probability of being cracked on these materials within the time available for the reaction, which would lead to significant yields of heavy cycle oil and coke. Nevertheless, it was reasoned that cracking of heavy hydrocarbons on these commercial catalysts, might in principle, be increased by the integration of a matrix component having controlled-textural properties and activity. Experimental silico-aluminates obtained through thermal and chemical treatments of kaolinite are an alternative to matrix additives for promoting high molecular weight hydrocarbon cracking. Table 2 shows the results obtained from the base material and kaolinite-derived silicoaluminates having promise in the formulation of FCC matrices for primary cracking of heavy hydrocarbons. The starting material was a kaolinite with low porosity, low area and, consequently, low heavy hydrocarbon cracking capability. As a result of the treatments to which kaolinite was subjected, its textural properties were substantially modified, increasing surface area up to ten-fold and pore volume up to four-fold. For processing of the heavy fractions of gas oil and resid, in addition to matrix area, pore volume distribution is important. The catalyst must have a large pore volume to accommodate more liquid hydrocarbons and more coke than conventional catalyst matrices. The optimal pore volume is from 2 to 6 times the size of the molecule to be cracked [3]. Macromolecules

378

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10

............. , ....... , ..... i

. . . .

100 DIAMETER PORE, A ~

i...i..i..i..i 1

1000

Figure 1. Pore-size distributions of commercial catalysts.

Table 2. Controlled Pore Distribution of experimental silico-aluminates

Kaolinite Carb- 11 a Carb- 14b Carb- 15 c Filttq d

Surface area, m2/g

Pore volume, cc/g

Pore diameter, A

18 163 186 192 125

0.13 0.50 0.51 0.54 0.35

245 89 70 77 94

a

Obtained with 0.8 g NaOH / gelay at 85~

2 h.

b

Same as Carb-11, with 1.6 g NaOH. c Same as Carb-11, 4 h . d

Same as Carb-14, with 1.6 g NaOH/gmullite at 95~

6 h.

are converted in the matrix into small entities that can penetrate the zeolite for its selective conversion to products that make up gasolines. However, too much matrix activity can lead to non-selective cracking; therefore, the catalyst must have an optimal zeolite/matrix activity ratio to maximize bottoms cracking with a minimal loss of gasoline yield [4, 5]. Figure 2 shows pore-size distributiona of silico-aluminates having promise for FCC matrix application. For comparison purposes, the initial kaolinite is included. Modification o f the kaolinite increases porosity, and, consequently, pore volume and surface area are increased. A bi-modal pore distribution is created, generating large fractions of pores in the 20 - 70 A and

379 1.5 1.0 0.5 ~0.0

i

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O

~ ~1.0 m0.5

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Figure 2. Pore-size distributions of experimental matrix materials. 200 - 1000 A regions. It is mainly the pores in the latter region that provide acid sites for primary cracking of heavy hydrocarbons, thereby generating molecules of intermediate size for subsequent cracking in the 20 to 70 A region of the matrix and in the zeolite structure. Hydrothermal stability is a critically important property of the constituents of FCC catalysts; accordingly prototype and reference materials were subjected to a hydrothermal deactivation treatment at a temperature of 788~ 4 and 8 hours and 100% steam. Results are summarized in Table 3. Synthetic silico-aluminate (MX-0994) and alumina/kaolin (MM0894) base matrices, which are used in the preparation of commercial FCC catalysts designed to crack heavy feedstocks, were used as references. Data in Table 3 show that after high temperature treatment the alumina/kaolinbase matrix retains, 32 % of its original surface area, while the silico-aluminate of reference retains 58 %. The experimental prototypes treated under the same conditions, in addition to having a considerably greater area, after being deactivated, appear to be more stable, since they retain 78% on the average of their original surface area. An additional advantage is that hydrothermal treatment of experimental silico-aluminates does not modify substantially their pore distribution (Figures 2 and 3). Although the reference silico-aluminate, does not contain macropores, it retains its mesopores, while the alumina-based matrix, undergoes area reduction accompanied by a loss of porosity. The final test that determines if experimental silico-aluminates can be applied in the formulation of FCC catalysts is the evaluation of their catalytic behavior. Materials deactivated for 4 and 8 hours at 788~ in 100% steam were evaluated in a microreactor at 520~ with a cat/oil ratio of 5, WHSV of 15.82 h -1 using a modified gas oil feedstock with 5 wt. % of atmospheric residue. The material of reference in this study was the kaolinite with no heat nor chemical treatments coded as Filsttq, which typically is used in the formulation of FCC catalysts. Two experimental silico-aluminates, CARB-11 and Filttq, which are the result of thermal and chemical treatments of the Filsttq kaolinite, were evaluated.

380 Table 3. Stability of experimental silico-aluminates. CARB- 11 CARB-14

Catalyst Total area Pore volume Pore Diameter Total area Pore volume Pore diameter Area retention Total area Pore volume Pore diameter Area retention

1.2 0.8 0.4 0.0

o

m2/g cc/g A Material

156 149 0.41 0.41 106 A 108 % 96 80 Material treated @ 788~

117 101 0.40 0.20 139 69 59 81 8 h, 100 % steam

cc/g

cc/g

A %

FILTTQ

163 186 125 200 0.51 0.50 0.23 0.35 89 70 94 treated @ 788~ 4 h, 100 % steam

m2/g

m2/g

MX-0994

131 0.39 86 80

145 0.43 78 78

115 0.21 42 58

94 0.36 151 75

MM-0894 103 0.11

34 0.09 106 33

33 0.09 71 32

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............i.............. i....i...i...i..i.'~. MX0994 8 H ' " "~''~ ~ ...... -..

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..................... ii ii!!

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PORE DIAMETER, A ~ Figure 3. Pore-size distributions of steamed experimental silico-aluminates.

381 According to the results summarized in Table 4, a consequence of the kaolinite modification is a considerable change in the catalytic properties, i.e., an increase in the conversion rate of at least 40 units and a reduction in heavy cycle oil of 75 wt. %, with respect to the initial material. It is important to note that the amount of LPG was increased by 500 % as well as its olefinicity, measured as (Total olefins/LPG) x 100. The resulting average olefinicity of 81% is higher than that of a typical FCC catalysts. If the concentration of olefins is high in LPG, it can be expected that the gasoline fraction will be high as well and, consequently, its octane number. On other hand, selectivities to coke and dry gas are similar to those commonly observed in FCC catalysts, and it may be expected that when these materials are included in an FCC catalyst matrix, these parameters will not be unfavorably modified.

Table 4. Catalytic evaluation of experimental silico-aluminates.

Dry gas LPG Isobutane Iso-C4Total C 4C3 = + C4 = Gasoline LCO HCO Coke Conversion

FILSTTQ CARB11-4H CARB11-8H FILTTQ-4H FILTTQ-8H 1.6 1.4 2.0 1.7 0.6 12.4 9.6 10.0 10.9 2.1 1.0 1.7 1.1 0.4 1.0 2.1 1.9 2.3 2.0 0.2 4.8 5.6 5.0 0.7 5.6 9.7 8.0 8.1 1.4 9.1 32.5 29.9 36.7 36.5 6.2 24.9 25.3 26.1 26.0 15.5 23.1 29.8 18.3 20.9 74.3 5.4 4.0 4.9 1.4 5.9 10.3 55.5 53.1 52.0 44.9

4. CONCLUSIONS Modified silico-aluminate materials were investigated as candidates for inclusion in matricies of FCC catalysts for cracking heavy hydrocarbons. Through thermal and chemical treatments of kaolinite is possible to gain at least 40 conversion units with respect to the base material, as well as a reduction of 75 wt.% of heavy cycle oil and a 500 % increase in higholefinic LPG yield. Due their enhanced textural properties, hydrothermal-stability, and catalytic properties, the experimental silico-aluminates are promising alternative candidates for application as matrix additives to FCC catalysts for the efficient upgrading of bottoms into higher quality gasoline and LPG with high olefinicity.

382 REFERENCES:

Humphries., J. R. Wilcox, Oil & Gas Journal pg. 45, Feb. 6 (1989). A. F. Sweezey., Catalyst Report., Engelhard Corp., (1996). VIP. O" Connor., Ketjen Catalyst Symposium Scheveningen., Netherlands 1986. M. Bourgogne., Catalytic Cracking of Residues Feedstocks at Total Flanders Refinery, Akzo Nobel Catalyst Symposium 1994, p 121. 5. Ritter., Grace Davison., Catalyst Trends and New Materials for FCC., Primer Foro de Avances en la Industria de la Refinaci6n., M6xico D.F. Agosto 28, 1995.

1. 2. 3. 4.

9

Elsevier Science B.V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

383

K i n e t i c s o f C o k e C o m b u s t i o n during T e m p e r a t u r e - P r o g r a m m e d O x i d a t i o n o f D e a c t i v a t e d C r a c k i n g Catalysts Cam Le Minh, Chaoen Li and Trevor C. Brown Department of Chemistry, University of New England, Armidale, NSW, Australia

The kinetics of catalytic-coke combustion has been investigated by analysing evolved CO, CO2 and H20 during the temperature-programmed oxidation (TPO) of coked cracking catalysts. Structure in the TPO spectra is affected by variations in the rate-determining oxidation step. An assumed oxidation mechanism forms the basis for calculating rate parameters for carbon oxide evolution from industrial spent, cyclohexene coked and 1-hexene coked cracking catalysts. Optimised activation energies allow a comparison of "soft" and "hard" coke combustion reactivity. Measured temperature-dependent H/C ratios indicate that low hydrogen content ("hard") coke, forms from adsorbed hydrocarbon molecules ("soft" coke) via dehydration or dehydrogenation.

1. I N T R O D U C T I O N The technique of temperature-programmed oxidation (TPO) combined with either evolvedgas [ 1-4] or gravimetric analysis [5, 6] is often employed to investigate catalytic coke. Details of the proximity of coke to metal additives or the support and the nature of coke are often inferred from the shape of resultant TPO spectra [4, 6, 7]. However, recently the temperaturedependence of the combustion mechanism has been shown to also affect the spectral shape, and so TPO peaks may not always be a direct indicator of location or type [8]. Therefore, for a complete interpretation of TPO spectra a fundamental understanding of the kinetics of coke combustion is required. Mechanisms of coke combustion are uncertain, although some investigations have been reported. In early work [9-11 ] intracrystalline diffusion processes, under certain conditions, were shown to affect combustion rates. Under the conditions employed in our work diffusion is unlikely to be affecting the apparent oxidation kinetics [ 12]. Other catalyst and coke properties may, however, be important. When simulating evolved CO2 during TPO Querini and Fung [3] considered variations in coke reaction order due to either different coke geometries or the extent of conversion or coke particle-size distribution, and variations in rate coefficient parameters due to different coke properties. Several groups have also shown that hydrogen-rich coke is highly reactive [9, 13-15]. Extensive studies on amorphous carbon and coal char combustion have demonstrated that surface-oxides are important precursors to CO and CO2 formation [ 16-19]. In order to gain details of the rate determining steps involved in catalytic-coke combustion during TPO we have continuously monitored the rate of CO, CO2 and H20 evolution from a coked cracking catalyst. Those who do monitor evolved gases during TPO of catalytic coke generally record CO2 with CO being converted to CO2 prior to analysis and in most cases H20 is not routinely measured.

384 2. E X P E R I M E N T A L

The apparatus constructed for controlled coke deposition and to continuously monitor CO CO2 and H20 evolution during subsequent oxidation has been previously described in detail [8]. Briefly, 500 mg of cracking catalysts are coked, after calcination to 800 ~ by means of a controlled injection of 1 mL of 1-hexene (Aldrich), 1-octene (Aldrich) or cyclohexene (Aldrich) at a set temperature in the range 200 - 600~ The deactivated substrate is then removed and divided into 100 mg batches for TPO and surface area analysis. Spent catalysts collected from the industrial catalytic cracking at Ampol, Queensland, Australia were investigated without further treatment. Water vapour is monitored by means of a Model 95 A Super Dry Hygrometer (Alpha Moisture Systems). This records the ambient dew point over the range 0 to -80 ~ which corresponds to vapour pressures of 6000 to 0.5 ppm. The hygrometer is calibrated by decomposing accurately weighed quantities of Ca(OH)2 (>98% BDH). Evolved carbon oxides are measured, following conversion to methane over a ruthenium catalyst, via dual flame ionisation detectors (FIDs). One FID records the combined CO plus CO2 gas, while the other monitors only CO by removing CO2 over ascarite. The difference between the two FID signals is evolved CO2. The FID's are calibrated using standard gases supplied by BOC (119 + 2 ppm CO and 519 + 10 ppm CO2 both in N2). All signals are connected to a Strawberry Trees data acquisition system and personal computer. Evolved gas partial pressures are converted to moles / m 2 catalyst / ~ by assuming CO and CO2 velocities are equivalent to the N2 velocity which is stabilised by Bronkhurst mass flow controllers. For H20 the pressured is adjusted to the evolution rate by taking into account the slower velocity of H20 molecules through the system [20]. BET surface areas were measured using equipment constructed and calibrated in our laboratory. For our experiments TPO involves exposing the coked substrate to 0.939% 02 in N2 at a temperature (typically 200 ~ where reaction is minimised then heating to 1000~ in a controlled linear manner. The heating rate is set at either 5 or 10 ~ min -1.

3. C O K E OXIDATION M E C H A N I S M Figure 1 illustrates the evolved-gas signal for a spent cracking catalyst exposed to 0.939% 02 and heated from 200 ~ to 1000 ~ at 10 ~ min -1. A single peak is observed for CO, three peaks or shoulders for CO2 and three broad peaks or shoulders for H20 evolution are apparent. Overall molar hydrogen to carbon ratio is 1.7:1. By comparison evolved gases from the TPO of the cracking catalyst coked at 400 ~ with 1-octene (Figure 2) shows a large water peak at ca. 430 ~ which is accompanied by CO and CO2 peaks; in addition to those observed for the spent TPO. Overall molar H/C ratio for coking the cracking catalyst with 1-octene at 400 ~ is 2.2:1. The overall H/C ratio decreases with increasing 1-octene coking temperature, ie., 2.6:1 (200 ~ 2.5:1 (300 ~ 2.2:1 (400 ~ 1.6:1 (500 ~ and 1.0:1 (600 ~ Temperaturedependent H to C ratios provide more information as to hydrocarbon properties of the coke. This is shown in Figure 3 up to 650 ~ for 1-octene coking at 200, 400, 500 and 600 ~ as well as for the spent cracking catalyst. Ratios at temperatures greater than 650 ~ and less than 800 ~ were small. Non-systematic fluctuations were observed in the temperature dependent ratio beyond 800 ~ This is due to the carbon signals being very small, which introduces large errors. The temperature-dependent ratios curve for the spent catalyst lies between the 500 and 600 ~ 1-octene coking curves. Industrial cracking is generally carried out at ca. 550 ~

385

0.014 0.012 o

0.01 0.008 P

0.006 0

0.004

,// :a. 0.002 0 ~umn~uimim--- r 200 300 400

f

2

~

,

N

P'

N~,

1 500

r 600

,- 700

800

900

1000

Temperature (~ Figure 1. Evolved H20 ( - - ) , CO( 9 ) and CO2( o ) during TPO (10 ~ n'fin -1) in 0.939 % 02 of a spent cracking catalyst.

0.04 0.035 o 0.03 - 0.025 8 0.02 o 0.015

/ / !

\

\ \

/ ( /

O

fil

"--'X

0.01 0.005 0 I 200

I

300

400

500

600

700

800

900

1000

Temperature (~ Figure 2. Evolved H20 ( - - ) , CO ( 9) and CO2 ( o ) during TPO (5 ~ min -1) in 0.939 % 02 of a cracking catalyst with 1-octene at 400 ~ Of particular note in Figure 3 are H to C ratios as high as 10 and the decrease in the maximum of these ratios as the coking temperature is increased. Low temperature and high hydrogen content suggests that the first oxidation step is removal of hydrogen from adsorbed hydrocarbons; H:C for alkenes is 2:1 and for alkanes is 2n+2:n, where n is the number of carbon atoms in the molecule. Molecules with ratios greater than 4:1 are not possible. Large but decreasing H/C ratios may be explained by low hydrogen content ("hard") coke forming from "soft" coke either by combustion or pyrolysis, ie., n

-CnH2n + ~O2 ~ -CnH2n A

-Cn + nH20

-Cn + nil2

(1) (2)

386 where -CnH2n and Cn denote "soft" alkene and "hard" coke respectively. Series of dehydrogenation reactions have been suggested for coke deposition at metal sites [21 ], but coke formation on acidic oxides is generally considered to occur via carbonium ions [22]. The role of chemisorbed oxides in the oxidation step (reaction (1)) is unknown. Because of this and expected complications due to readsorption of water onto the catalyst surface the rate parameters for H20 evolution during TPO have not been calculated. 10

200 ~

~" 6 ~

4 2 _

0 250

300

350

400

450

500

550

600

llllllllf[{

650

Temperature (~ Figure 3. Temperature-dependent molar hydrogen to carbon ratios for cracking catalysts coked with 1-octene at 300, 400, 500 and 600 ~ and for the industrial spent cracking catalyst. Evolution of carbon oxides during the TPO of the "hard coke" and of charcoal has been reported to be strongly dependent on the oxidation mechanism [8]. This is based on an observed temperature and heating rate-dependence of CO2 evolution during TPO. The proposed mechanism is: -Cf + 02 -'-'> -C(O2) -C(O2) +-Cf ---> -C(O) + CO -c(o) ---> c o -c(o) -Cf + 02 --) C02 -C(02) .-) C02

(3) (4) (5) (6) (7)

where -Cf denotes a free carbon site available for 02 chemisorption, -C(O) and -C(02) are dissociated and undissociated surface oxides respectively.

4. KINETIC MODEL Although in all recorded TPO spectra of "hard" coke CO evolution is represented by a single Gaussian curve, reaction steps (4) and (5) suggest two peaks should be observed. The contribution of step (5), however, is expected to be small as such unimolecular desorption processes have been shown to continue, in the absence of oxygen, at temperatures well beyond those required for complete conversion during TPO [16]. Despite this, a small peak,

387 corresponding to reaction (5), may still underlie the large peak associated with reaction (4). The rate of CO evolution is represented by dpco iv dt = k4[-C(02)] [-Cf]

(8)

dpco v dt = k5[-C(O)]

(9)

where Pco iv and Pco v refer to the pressures of carbon monoxide evolved from reaction (4) and reaction (5), respectively, while the subscript for each rate constant correspond to the reaction number. Square brackets around adsorbed oxides or deposited carbon represent coverage. By assuming a steady-state concentration of-C(O2) and that k-3 >> k4 reaction (8) becomes: dpco iv dt = kiVpo2[-Cf]2

(10)

where k iv - k3k4 and k_3 is the rate constant for the reverse of reaction (3). k_3 Three peaks are observed in evolved CO2 during TPO. The first (P1) is associated with evolution from coke located near combustion-promoting metal sites or unpolymerized coke with a low C:H ratio. This may be represented by reaction steps (3) and (7) and hence the rate, assuming a steady-state concentration of-C(O2) and k_3 >> k7, is dpc02 vii dt = kviipo2[-Cf]

(11)

where k vii - k3k7 The second peak (P2) in the evolved CO2 spectra represents polymerised -

k-3"

coke silica/alumina sites and also evolves via steps (3) and (7) albeit with a different rate. The rate expression, therefore, is identical to equation (11). For the third peak (P3) all remaining -C(O), formed via step (4), are removed by reaction with 02 (step (6)). It is unknown whether an adsorbed 02 intermediate is involved in this step, but the rate expression may most simply be written dpco2 vi dt = k6 Po2[-Cf][-C(O)]

(12)

Free carbon sites may be non-existent and so [Cf] is replaced by [-C(O)], hence second order with respect to this stable oxide. In the model the concentration of dissociated oxides are taken as the concentration of [-C(O)] resulting from step (4). Thus the difference between the area under the curve representing P3 minus the area of the CO curve indicates either the extent of step (5) if the CO area is greater than the P3 area or the availability of-Cf for step (6) if the CO area is less than the P3 area. In the present model the pre-exponential factors and activation energies are optimised by non-linear regression using the Solver module on Microsoft Excel. The amounts of coke associated with CO2 evolution at P1 and P2 are also optimised, while a mass balance determines the coke quantity associated with CO at P5 and CO2 at P3. The oxygen partial

388 pressure is assumed always to be in excess. Figures 4 and 5 illustrate the modelled CO2 and CO peaks for the industrial spent catalyst. Optimised parameters for the spent, cyclohexene coked and 1-hexene coked cracking catalyst are listed in Table 1. Errors in all optimised parameters have been estimated by systematically adjusting the parameters until the simulation differs from the experimental data by the predicted experimental errors. The peaks in Table 1 correspond to the following elementary steps: CO2 formation via steps (3) and (7) at metal sites or of "soft" coke, CO2 formation via steps (3) and (6) of "hard" coke, CO2 formation via steps (3) and (7) of "hard" coke, CO formation via steps (3) and (4) of "hard" coke, CO formation via step (5) of "hard" coke.

P1; P2; P3; P4; PS;

1.6 1.4

o

~

1.2 1 0.8 0.6 0.4 0.2 0 200

~llllll

P2 ~ a I][

300

400

500

600

I

I

700

800

[] 900

Temperature (~ Figure 4. Experimental CO2 ( O ) evolution and simulated peaks during TPO (10 ~ min -1) in 0.939 % 02 of a spent cracking catalyst.

o

1.5 1

O ::I.

0.5

200

300

400

500

600

700

800

900

Temperature (~ Figure 5. Experimental CO( 9) evolution and simulated peaks during TPO (10 ~ min -1) in 0.939 % 02 of a spent cracking catalyst.

389 All listed TPO spectra show three CO2 peaks (P1, P2 and P3), while two CO peaks (P4 and P5) are calculated for the spent and cyclohexene coked catalysts, but only P4 is apparent for the 1-hexene coked catalyst. When present, P5 only represents 5 - 10% of the total carbon associated with CO formation and the error in the calculated rate parameters for this peak is large. Care is also required in the interpretation of P3 rate parameters because this is associated with the last 30 - 50 % of the deposited carbon. For this residual coke the structure may have collapsed and a broad distribution of rate parameters are likely. Also at higher heating rates (10 ~ this peak becomes very large and oxygen depletion may become rate controlling. Table 1 Rate parameters for carbon-oxide evolution during TPO of an industrial spent cracking catalyst and of a fresh cracking catalyst coked with 1-hexene at 500 ~ Coked Evolved A factor Activation Energy Carbon per peak* (~tmole g- 1 C- 1) Catalyst Oxides (kJ mol- 1) ~tmole / g-1 Spent (1.6 + 1.2)• 102 60+2 C02 P1 44+3 (1.1 + 0.3) • 106 130+2 P2 42+6 (9.7 + 1.0) • 103 111+2 P3 267 + 25 (3.4 + 0.4) • 104 113+2 CO P4 278 + 25 (5 + 4) • 104 102+ 4 14 + 10 P5 (7.3 + 5.5) • 103 Cyclo75+3 C02 P1 131+9 (2.1 + 0.6) x 107 hexene 137+3 P2 187 + 27 (4.4 + 0.5) • 107 164 + 5 P3 822 + 90 (1.4 + 0.2) x 105 CO P4 110+2 922 _+83 (2.7 + 0.3) • 101 20+2 P5 101 _+77 1-hexene (1.9 + 1.4) x 103 65+3 CO2 Pl 83+6 (6.4 + 1.7) • 106 128+2 P2 132 + 19 (7.9 + 0.8) x 106 149+5 P3 532 + 50 (2.0 + 0.2) x 105 CO P4 111 _ 2 465 + 4 2 * Heating rate of 10 ~ min -1. While A-factors and activation energies are independent of the conditions, the carbon associated with each peak is dependent on the heating rate. The most significant rate parameters are those calculated for P1, P2 and P4. For all coked substrates, the corresponding A-factors and activation energies for these peaks are a similar magnitude. Any differences in A-factors between equivalent peaks for different substrates can be associated with differences in the amount of carbon per peak. Hence reflecting the propensity for coke formation. Similar activation energies for corresponding P2 and P4 indicate the oxidation reactivity of the final "hard" coke is independent of the initial hydrocarbon source of the coke. One important difference in activation energies, however, is apparent for P1, which has been associated with "soft" coke probably adjacent to metal sites, where the activation energy of the cyclohexene-coked catalyst is ca. 15 kJ mo1-1 higher than for spent and 1-hexene coked catalysts. This difference may be related to the dynamics of hydrocarbon adsorption. For linear alkenes only one end of the chain may be attached to the catalyst surface, while for small cyclic alkenes, which have a more compact structure, the entire molecule may be adsorbed onto or electronically affected by the catalyst surface. That is, adsorbed cyclohexene which may lie

390 flat on the surface, requires a higher activation energy for oxidation than does adsorbed 1hexene. A comparison of activation energies between P1 and P2 for the same coked substrates, show that the energy barrier for the former is half the barrier for the latter. These barriers provide a comparison of the oxidation reactivities of soft and hard coke. The higher A-factors for P2 when compared to P1 correspond to the compensation affect, where for the same reaction-type, larger activation energies result in larger A-factors [23].

5. CONCLUSIONS An analysis of the rate of CO, CO2 and H20 evolution during TPO of industrial and laboratory coked cracking catalysts has provided information on the mechanism and energetics of coke combustion. The mechanism has been deduced from previously reported studies on amorphous carbon oxidation [8], while rate parameters have been calculated from non-linear regression simulations of the TPO spectra. The rate of water vapour formation has not been analysed due to re-adsorption expected to affect the apparent kinetics. "Soft" and "hard" coke have been identified in the spectra, and oxidation activation energies of each compared. A further outcome of this work is the proposal that coke deposition on cracking catalysts proceeds from "soft" to "hard" coke via a series of dehydrogenation or dehydration steps. REFERENCES

1. 2. 3. 4. 5. 6. 7. 8. 9. 10. 11. 12. 13. 14. 15. 16. 17. 18. 19. 20. 21. 22. 23.

G.C. Bond, C.R. Dias, and M.F. Portela, J. Catal., 156 (1995) 295. J. Bartholdy, P. Zeuthen, and F.E. Massoth, Appl. Catal. A, 129 (1995) 33. C.A. Querini and S.C. Fung, Appl. Catal. A, 117 (1994) 53. S.M. Augustine, G.N. Alameddin, and W.M.H. Sachtler, J.. Catal., 115 (1989) 217. M. Larsson, M. Hulten, E.A. Blekkan, and B. Anderson, J. Catal., 164 (1996) 44. C.L. Pieck, R.J. Verderone, E.L. Jablonski, and J.M. Parera, Appl. Catal., 55 (1989) 1. K. Moljord, P. Magnoux, and M. Guisnet, Appl. Catal. A, 121 (1995) 245. C. Le-Minh, R.A. Jones, I.E. Craven, and T.C. Brown, Energy & Fuels, 11 (1997) 463. R.G. Haldeman and M.C. Botty, J. Phys. Chem., 63 (1959) 489. P.B. Weisz and R.B. Goodwin, J. Catal., 2 (1963) 397. P.B. Weisz and R.B. Goodwin, J. Catal., 6 (1966) 227. N.Y. Chen, T.F. Degnan, and C.M. Smith, Molecular Transport and Reaction in Zeolites. Design and Application of Shape Selective Catalysts, New York, VCH Publishers, 1994. F.E. Massoth, I&EC Process Design and Development, 6 (1967) 200. F.E. Massoth and P.G. Menon, I&EC Process Design and Development, 8 (1969) 383. P. Magnoux and M. Guisnet, Appl. Catal., 38 (1988) 341. A.E. Lear, T.C. Brown, and B.S. Haynes, in Twenty-Third Symposium (International) on Combustion. 1991,, The Combustion Institute, Pittsburgh, 1191. Z. Du, A.F. Sarofim, and J.P. Longwell, Energy & Fuels, 5 (1991) 214. S. Ahmed, M.H. Back, and J.M. Roscoe, Comb. and Flame, 70 (1987) 1. B.G. Tucker and M.F.R. Mulcahy, Trans. Faraday Soc., 65 (1969) 274. P.A. Redhead, Vacuum, 12 (1962) 203. D.L. Trimm, Appl. Catal., 5 (1983) 263. J. Scherzer and A.J. Gruia, Hydrocracking Science and Technology, New York, Marcel Dekker, 1996. A.K. Galwey, Adv. Catal., 26 (1977) 247.

9

Elsevier Science B.V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

391

Activity, Selectivity and D e a c t i v a t i o n o f Y Zeolite: A C o m p r o m i s e in Fluid Catalytic C r a c k i n g Catalysts

C. Maldonado M., P. Gamero M., F. Hern/mdez B., J.A. Montoya F., J. Navarrete, J.A. Toledo A., A. Vfizquez R. and A. Vargas Instituto Mexicano del Petr61eo, Eje Central Lfizaro C/trdenas 152. Apartado Postal 14-805 07730 M6xico D. F.

The hydrothermal deactivation of Y zeolite containing 0, 4, 7 and 12 wt.% of REO and its effects on catalytic activity, stability and selectivity were investigated. The Y zeolites were hydrothermally deactivated at 788~ in three consecutive cycles of two hours each. The fresh and deactivated zeolites were characterized by measuring Unit Cell Size (UCS) and surface area. The acidic properties were measured by the Temperature Programmed Desorption (TPD) of ammonia and IR-pyridine desorption. In order to correlate structural, textural and acid properties with catalytic behavior, the zeolites were evaluated in the conversion of cyclohexane. The Hydrogen Transfer Index (HTI) measured as a ratio of paraffins to olefins is a parameter of the selectivity. It was found that the REO was incorporated into zeolite structure up to high concentrations modifying to some extent XRD deflection, the acidic properties and the HTI ratio. After deactivation, the acidity and HTI were diminished and the Lewis/Bronsted acid ratio was modified. HTI decreased as REO concentration increased.

1. I N T R O D U C T I O N Zeolite Y has excellent activity for promoting the cracking of hydrocarbons. However its stability is not high enough to resist deactivation at FCC regenerator conditions (1). Rare Earth Oxide (REO) is applied to stabilize the zeolite structure in terms of Unit Cell Size (UCS), surface area and acidity to produce a catalyst with properties meeting the requirements of an FCC unit. Production of gasoline, octane, octane barrel, olefins, etc., requires specific zeolite and REO contents (the latter ranging from 0 to 2.4 % wt) for each application (2). However it is well known that the Y-REO-zeolite is deactivated by dealumination to obtain an FCC equilibrium catalyst, and this process is followed by a decrease of the cell parameter, although dealumination

392 is more difficult for Y-REO-zeolites than for Y-zeolites (2). Zeolite deactivation is made to adjust the acidity properties and consequently to increase the rate of cracking relative to hydrogen transfer increasing the FCC gasoline octane (3). However, when dealumination of a Y-zeolite is carried out until to the limits of a FCC equilibrium catalyst (cell parameter ao 24.35 A) there is about 15% loss of crystallinity which is an undesirable process but it is inevitable (4). The conversion of cyclohexene has been proved as good test reaction to evaluate the hydrogen transfer properties of the Y-zeolite (3). The aim of the present work was to study the behavior of the Y-zeolite with variable content of REO at the first steps of deactivation (dealumination) in order to avoid the problem of zeolite amorphisation or loss of crystallinity.

2. EXPERIMENTAL

Sample preparation The Y-zeolite was synthesized by conventional route. The Na-Y-zeolite, named ZYOF, was exchanged with NH4OH to obtain the acid form and after this it was used to prepare samples with 4, 7 and 12 % wt of rare earth by exchanging with rare earth nitrate solution. The YREO-zeolites were calcined in air flow at 500~ during two hours, this samples will be identified as ZY4F, ZY7F and ZY12F. The hydrothermal deactivation procedure of the zeolites consisted in saturating the zeolite with water and put quickly inside of a muffle which has been previously preheated at 788~ After soaking for 2 hours at this temperature, the samples were pulled out of the muffle and airquenched. This procedure was repeated three times. The name of the deactivated samples will be ZYOD, ZY4D, ZY7D and ZY12D.

Characterization The powder X-ray diffraction pattems were measured in a D-500 SIEMENS diffractometer with a graphite secondary beam monochromator and CuKt~2 contribution was eliminated by the DIFFRAC/AT software to obtain a monochromatic CuKtx~. The Unit Cell Size (UCS)was measured following the ASTM D-3942-90 procedure. The Surface areas were measured by nitrogen adsorption at 75 K on a Micromeritics Accusorb 2100 E equipment using the ASTM method D-3663-78. Temperature Programmed Desorption (TPD) of ammonia and pyridine adsorption by Infrared Spectroscopy (IR) were used to characterize the acidity of the zeolites. For IR-Pyridine the spectra were recorded each 100~ and the characteristic bands of Lewis and Br6nsted acid sites (1444 cm ~ and 1540 cm l, respectively) were integrated in order to obtain the total acid sites.

Catalytic activity The catalytic activity of the obtained zeolite was carried out measuring the conversion of cyclohexene at micro reactor level, using 0.10 g of catalyst, at constant temperature(250~ and different flow rate (contact time). To analyze the gas effluent an on-line Gas

393 Chromatograph was used. The Hydrogen Transfer Index (HTI) measured as ratio of paraffins and olefins was a parameter of selectivity.

3. RESULTS AND DISCUSION 3.1 Structure, cell parameter and texture The incorporation of REO into the Y-Zeolite structure modified both the intensity and position of the reflections in the X-ray pattern. This modifications depended on the amount of REO. The loss of intensity is more drastically observed for (220), (311) and (620) reflections which practically disappeared for 12% wt of REO (Figure 1). This behavior is more clearly shown by the comparison between the summation of the integrated intensity of the most intense peaks (XRD-II) for each sample and taken the 0% wt of REO (ZY0 fresh) as reference material (% XRD intensity - 100 x (XRD-II problem / XRD-II of ZY0)) (Table 1). It is important to mention that traditionally this ratio has been related with the crystallinity of the Y-zeolite but the original crystal structure is not been modified with REO addition, instead of that the REO is only accommodating into the supercage occupying the $4 and $5 sites before calcination and migrating to $2 sites (sodalite cage) as zeolite undergoes calcination, as it has been recently founded (5). In our case the REO content is higher in comparison with the conventionally FCC catalyst provoking that almost all the sodalite cages are filled modifying the intensity of some reflections. The crystallinity ratio only applies for cases when a given YREO-Zeolite is deactivated and the crystalline structure is broken gradually. From the table 1, it is observed that the % XRD intensity diminish from 100% to about 47% when the REO goes from 0% to 12% wt. The same behavior is observed for deactivated samples. It is also found that the deactivated samples did not lose % XRD intensity respect of the fresh samples (crystallinity), in comparison with Y-REO-zeolites deactivated at levels of equilibrium FCC catalyst where % XRD intensity decreased 15 % (4). In the samples ZY12 fresh and deactivated it is detected the presence of La203 as segregated phase, which suggests that 12% wt of La203 concentration overflows the adsorption capacity of the zeolite. For fresh samples the micropore area increased lightly with content of REO but it decreased for the 12% wt of REO (table 1). The micropore volume suffered a small increment for 4 and 7% wt REO content but decreasing for 12% wt of REO. In contrast, for deactivated samples both micropore area and micropore volume remained constant for ZYOD, ZY4D and ZY7D but diminishing in sample ZY12D. For all deactivated samples the micropore area and micropore volume decreased respect to the fresh samples.

394

r v

'r 03 v o

=.,,

w

c

I

10

5

20

15

25

9

I

.

30

9

35

20

Figure 1.- XRD patterns of fresh Y-Zeolites exchanged with different concentration of REO. A) ZY0, B) ZY4, B) ZY7 and C) ZY12.

Table 1 .- % XRD intensity and texture characteristics of fresh and deactivated samples. SAMPLES

FRESH

SAMPLES

% XRD M. Area M. Volume Intensity (m2/gr) (cc/gr) ZYOF

100

560

0.278

ZY4F

76

572

0.286

ZY7F

56

588

ZY12F

47

554

DEACTIVATED

% XRD Intensity ZY01)

M. Area M. Volume (m2/gr) (cc/gr)

98

515

0.256

ZY4I)

75

514

0.254

0.281

ZY7D

54

512

0.253

0.274

ZYI21)

42

502

0.249

M. Area, M. Volume: Micropore Area and Micropore Volume.

395 In figure 2, it is related the cell parameter with the REO content of fresh and deactivated YZeolite. It is shown that the cell parameter expanded linearly with amount of REO, this behavior is well known as the Vegar's law (6) which is interpreted as an evidence that the REO was incorporated into sodalite cage of Y-zeolite. This behavior was the same for both fresh and deactivated materials but in case of the deactivated samples the cell parameter shrank relative to the corresponding fresh samples. It is important to notice that the shrinkage is lower as the REO increases. This suggested that at similar deactivation condition the REO content prevented the dealumination. However, it is important to mention that dealumination of REO-Y-Zeolite does not necessarily mean a decrease in the unit cell parameters. Figure 2. UCS vs. REO 24.70 24.65 24.60 ~ 24.55 . ~ 24.50 24.45 24.40

I

I

3

6 RB:), % wt.

m DEACTWATEDl I 9

12

3.2 Acidic properties The total acidity of fresh and deactivated materials measured by TPR-NH3 is presented in figure 3. As it was expected the total acidity was lower for deactivated samples for either composition. For 4% wt REO the total acidity, in both fresh and deactivated samples increases but beyond this REO concentration the total acidity diminished. This fact seemed not to be according with general postulation where the total acidity and HTI increases linearly with the expansion of unit cell parameter. However this assertion is found when a low REO concentration is used typically < 2.4% wt. and the cell parameter is between 24.23 A - 24.45 A (2). Figure 3. Total acidity vs. REO 7o0 6o0

_= o

,,~ 500

_,1

o 400 300

IO FRESH

m DF_ACTWATED) I

4

(

RED, % w t .

8

12

396 The Lewis/Bronsted acidity ratio (L/B) as a function of REO concentration for both fresh and deactivated materials is presented in figure 4. Additionally to the decreasing of the total number of acid sites with the REO content, the L/B ratio is also changing. From figure 4, it is observed that in fresh samples the number of Lewis acid sites decreases as the REO amount increases but an opposite tendency is observed in case of deactivated samples. This results suggest that in all fresh samples the REO is strongly interacting with the -Si-O-Si framework increasing the Bronsted acidity but as the samples are hydrothermally treated the REO could migrate to the supercages decreasing the REO-framework zeolite interaction and REO behave as an isolated oxide with the typical Lewis acidity. Therefore it reasonable to assume that for REO-Y-Zeolites the Bronsted acid sites generation is directly related to the amount of REO cations occupying $2 sites. Fig.4. Lewis/Bronsted rat/o vs REO 3.50 3.00 I-

2.50 2.00

ta 1.50 .j 1.00

9

0.50 0.00 0

4

RED, % wt:

8

12

3.3 Catalytic activity The cyclohexene conversion to paraffins and olefins which is well known as the Hydrogen Transfer Index (HTI) ratio was used to evaluate the catalytic activity of the fresh and deactivated materials at constant conversion (30 mol %) varying the contact time and at 250~ The HTI results were related with the REO content, cell parameter and total acidity and are plotted in figures 5, 6 and 7, respectively. Fig.5 Hydrogen transfer index v~REO 3.5 3.0 2.5 I----2"0 -r 1.5

~

1.0 0.5

,

d

~,"

[,FRESH ] /B DEACTIVATED/

0.0 0

I 4

REO, %wt.

I 8

12

In figure 5, the HTI for fresh samples presented a maximum in the case of sample ZY4F and beyond 4 %wt the HTI decreases. In contrast for deactivated samples the HTI decreases as the

397 REO increases. This behavior suggested that the REO incorporation inhibit the HTI reaction. It is important to notice that in case of zeolite with 0% wt of REO the HTI decreases as the zeolite is deactivated this behavior agrees with the mechanism of I-iTI reaction proposed by Pine (2) in which the lowering of the cell parameter by dealumination favors the olefins formation while decreasing the HTI reaction due to the lowering of site density. However, this model does not consider the expansion of the unit cell size by the variable high REO incorporation. In our case, when the HTI was plotted as function of the cell parameter size (see figure 6), it is observed that as the unit cell parameter was reduced the HTI increases toward the HTI of zeolite free of REO for both fresh and deactivated samples. This results could be explained in terms of the Pine model in which the homogeneity of the HT (Hydrogen Transfer Reaction) site density decreases as a consequence of their blocking by the REO. Our results seemed to be complementary to that of low REO concentrations. Fig. 6 Hydrogen transfer indexvs UCS 3.5 3.0 2.5 --2.0 I-r 1.5

FRESH

1.0

DEACTIVATED

0.5 0.0 24.45

24.50

i

24 55 24.60 UCS, A

i

24 65

24.70

The figure 7 shows the HTI dependence with the total acidity. It is observed that the HTI increases as the total acidity is risen for fresh and deactivated zeolites, this is according with the cell parameters and REO content results. Additionally, as total acidity of fresh samples was higher than their corresponding deactivated samples then the HTI of fresh samples was also higher. However if we compare fresh and deactivated samples showing the same value of total acidity it would be expected similar activity in HTI, but from the figure 7 it is observed that it is not true. A possible explanation of that fact is taking account the Pine model where in spite of the samples have the same total acidity the HT site density is different in both cases due the distinct REO concentration. Fig. 7 Hydrogen tranfer indexvs Total Acidity 3.0 F-r

A

2.0 1.0 0.0 300.00 TOTAL ACIDITY, m m ol Nlq31g

398 4. CONCLUSIONS Deactivation for Y-REO-zeolite with different high REO concentrations was studied. The REO can be incorporated into the Y-zeolite around 12% wt and it can be followed by the cell parameter modification. For both fresh and deactivated samples, below 4% wt of REO the total acidity, and the HTI were increased but beyond this REO concentration the total acidity and the HTI decreased as the REO amount increased. The cell parameter and the total acidity diminished for all samples after the deactivation. This behavior was observed when the HTI was analyzed as a function of cell parameter, REO content and total acidity. It is clear that the control of REO content in FCC catalyst is one of the most important task for high octane gasoline production.

REFERENCES 1 .- D. J. Rawlence et al, Appl. Cat. 43, 213 (1988). 2.- L.A. Pine et al, Journal of Catalysis, 85, 466 (1984). 3.- W-C. Cheng and K. Rajagopalan, Journal of Catalysis, 119, 354-358 (1989). 4.- T.H. Fleisch, B.L. Meyres, G.J.Ray, J.B. Hall and C.L. Marshall, Journal of Catalysis, 99, 117-125(1986). 5.- J.G. Nery, Y.P. Mascarenhas, et al, Zeolites 18: 44-49, 1997 6.- Klug and Alexander, John Wiley & Sons, Second Edition, P. 562, 1974.

~

Elsevier Science B. V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

399

M / H - M F I C a t a l y s t D e a c t i v a t i o n and R e a c t i o n Kinetics for n - O c t a n e Transformation Papa J.*, Santos F., Becerra M.J., Giannetto G., Yanez F. and Garcia L. Escuela de Ing. Qufmica, Fac. de Ingenierfa, Universidad Central de Venezuela, Apartado 47268, Los Chaguaramos 104 l-A, Caracas, E-mail: jpapa~eacciun.ve

n-Octane transformation over an M/H-MFI zeolite-based catalyst was studied using a fixed bed tubular reactor, operated at 280~ and 200 psig under hydrogen atmosphere and at various WHSV. In order to simulate the amount of observed reaction products, an operative scheme of reactions and their kinetics expressions were proposed. An important inhibition effect by reactants was observed and, depending on the presence of inhibitors and/or poisons, so was catalyst deactivation that in certain cases, became almost total in a relatively short time.

I. INTRODUCTION New environmental regulations that apply to emissions of internal combustion engines, are imposing more severe formulations to gasoline each day. These new regulations demand for the development of alternate technologies to comply with them [1 ]. In developed countries the use of lead in gasoline has been forbidden, and the allowed amount of aromatics, oleflns and sulfur containing hydrocarbons has also been sensibly reduced. Hence, it is important and necessary to develop new catalytic processes to enhance the gasoline octane number and at the same time keep the content of those compounds at a low level. The octane number of a non branched paraffin is low when compared with a branched one [2]. Hence, for raw gasoline cuts rich in linear paraffm, that number can be increased developing catalytic processes that are able to crack hydrocarbons with large carbon chains, and to transform C5 and C6 ones to branched isomers. Zeolite based bifimctional catalysts that are able to selectively activate the formation of branched paraffin have been the subject of active research since 1960. In 1970 Shell introduced an isomerization process using a dealuminated mordenite impregnated with a noble metal as the catalyst [3]. The zeolite geometric selectivity has also been the basic property used to develop processes "like the Selectoforming and the M-forming using erionite and ZSM-5 respectively, or the Dewaxing with mordenite or ZSM-5, all in the protonic form and impregnated with platinum or nickel [4, 5, 6] The development of new catalytic processes for the selective hydroisomerization of linear paraffin, as an alternative process toward gasoline with a higher octane number, require the understanding of their catalysts bifimctional activity and selectivity. They have to be highly selective toward hydroisomerization keeping the cracking toward light gases to a minimum.

400 Zeolite-based catalysts in their acidic form are active in isomerization and cracking but, without the metallic function are not very active and selective, and deactivate very fast [7, 8]. In this work we report the kinetic behavior of a bifimctional catalyst based on a M/H-MFI zeolite with unique shape selectivity and active for enhancing gasoline octane number. As a test reaction the n-octane transformation was used. The catalyst showed to be subject to inhibition as well as deactivation effects other than the effects of coke formation. To test these effects, the presence of inhibitors and poisons in the feedstock, was also studied [9].

2. EXPERIMENTAL

The catalyst, 50% M/H-ZSM-5 and 50% alumina, was supplied by INTEVEP S.A. as 1/16 inch diameter cylinders. Its catalytic properties (activity, yield and selectivity) was studied in an isothermal fixed bed reactor operated under hydrogen atmosphere keeping the ratio of H/n-C8 always equal to 42.06. Deactivation tests were done under hydrogen atmosphere and at the following reaction conditions: WWH = 2.87 h ~, T = 280~ P = 200 psig., until a stable activity level or a complete deactivation was reached. For pure n-octane transformation over a stabilized catalyst, the change in products distribution with contact time, from 0.2 to 2 h, and with reaction temperatures, from 200 to 470~ was also studied [9]. Experimental results obtained at a WWH = 2.87 h l, and temperatures between 200 and 470~ showed that with temperatures below 280~ the reaction rate is under chemical control and that the activation energy is around 20 Kcal/mole, which is characteristic for acid catalysis. The undesirable cracking toward C1, and C2 hydrocarbons increases with temperature, but at 280 is still low, while the conversion gets high enough as to be interesting. Results showed that the product molar distribution did not change significantly within the 200-280~ range, and that aromatics are present only as traces, which indicates that the cyclization reactions are still not activated. This observation explains why the studied catalyst was able to increase the gasoline octane number without the formation of aromatics. Results for pure n-octane and for n-octane with different amounts of benzene (2%, 5% and 10%), pyridine (50 ppm, 250 ppm and 1000 ppm) and carbon disulfide (1%), are shown in Figure 1 as a function of time on run. Benzene shows minor effects on the deactivation level, whereas carbon disulfide acts as a strong inhibitor, and pyridine, when present in amounts over 250 ppm, behaves as a poison. Although benzene does not change the level of stable activity obtained with pure n-octane, the selectivity toward products is changed in a measurable amount. With a catalyst stabilized using pure n-octane, the degree of conversion and the selectivity was studied as a function of contact time. The selectivity showed to be nearly independent of conversion and the ratio of branched to non-branched hydrocarbons was always lower than one except for C7 compounds. The isomers distribution was found to be the typical of medium pore zeolites. Experimental results, lumped by the number of carbon in the chain, are shown in Figure 2.

3. DEACTIVATION KINETICS The nature of the catalyst deactivation, until an inactive state or a stable one is obtained,

401

varies from case to case. With pure n-octane and for a reaction temperature of 280~ the deactivation is mainly due to the formation of hydrocarbons with large and mainly linear carbon chains, and of small quantities of alkylbenzene and alkylnaphtalene, having boiling points up to 400~ [ 10]. They are strongly retained within the zeolite pores mainly due to their low volatility, and in minor amounts at the pores intersection, due to their sizes. A stable state is obtained when the production rate becomes equal to the removal rate, mainly by cracking.

= ~J

0.6 0.5 0.4

0.6 t 0.5 0.4

!i

::tt -"-'- 2% C6H6 o.1

0

200

400

600

0.50"6t

i 200

i 400

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0.60.5. . . .m. . . . . . . . . . . . .

0.4

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0.5

t

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04-

9r

0.2 t 0/

Or 0

200

01 0.2-

5% C6H6 I

300

o.1-

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~

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0

10% 06H6

I

I

200

400

600

0.3

I

I

~t

1% CS2, ~176t

0 0.3 ]

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100

200

,

300

250 ppm C5H5N

......

ol

0.3]T

0.25

0.15

0.15

0.05

0.05

0/ 0

100

200

300

400

0

0;

0

50 :m C25N I

,

,

100

,

200

300

I

400

1000ppm C5H5N ......

,

100

--

,

200

Time on run (min) Figure 1.- n-Octane conversion as a function of time on stream.

=--300

402 The first step of paraffin conversion over H-ZSM-5 catalysts showed to be of first order [11, 12], giving intermediates compounds that oligomerize between them and crack very fast. With n-octane, dimeric molecules and some of those formed by oligomerization give rise to the main compounds of the coke formed within the analyzed temperature range. These high molecular weight hydrocarbons are not detected and as a consequence it is not possible to establish a step by step reaction kinetics for their formation with our data. Considering the fact that intermediate compounds react very fast, and that dimeric ones are not detectable, it seems reasonable to assume that the formation of coke precursors is the controlling step, and that this step is linear with respect to the main reactant.

80 ) I I L

25

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Of 0

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0

,

t

50

1O0

150

, 0

1 0

~

i

O!

100

50

" 1O0

150

Contact time [min] Figure 2 Reaction product evolution with contact time Under the fixed reaction conditions, neither CS2 nor Pyridine react with n-octane or its reaction products, but when added to the reactor feed stream they adsorb on the catalyst active sites taking them away from the main reaction in an amount that depend with the time on run. Their effect adds to the coke effect on the catalyst activity and in order to take them into account, the kinetic expression is multiplied by a time dependent empirical factor

403

r = x(1) CC,H.'

x(2) + exp(- x(3) * t) 1 + x(2)

where x(1) is the rate constant, x(2) and x(3) two adjustable parameters, and "t" the time on run. This empirical factor mimics processes that reach a stable steady state as it is our case. For an isothermal plug flow packed bed reactor the following system of differential equations can be obtained d(C i 9V) J = ~=~aj~ .Rj .pc with i = 1, 2, 3 ..... C (1) dz j

-

which can be transformed into ni=n~+m~

j_-~

aj~.Rj,

\ mc J

with i = 1 , 2 , 3 .... ,C

(2)

This system was solved using a seventh order Runge Kutta-Vemer method, with adjustable step size and error control, which showed to be stable and fast. The regression was done using a non-linear regression software, M3D [ 11 ], with the following objective function

(3) n= 1

n i ,mod elo

Simulated data are shown in Figure 1 with solid lines and the parameter values that best fit experimental data in Table 1. The specific reaction rate constant x(1) has the same value for all cases and this is consistent with the fact that each deactivation run was done starting with fresh catalyst. On the other hand, parameter x(2) reflects some type of equilibrium state reached for free active catalytic sites. It takes very small values for a poison as pyridine, intermediate values for an inhibitor like CS2, and values close to one for pure n-octane and for n-octane mixed with the smallest amount of benzene. In larger amounts, benzene seems to act as coke inhibitor, which is consistent with the fact that it does not change activity in a significant manner but it does change the distribution of reaction products. Table 1 Parameter values for the deactivation experiment with pure n-octane in the ofinhibitors or poisons (WWH = 2.87 h -1, T = 280~ P = 200 psig) PYRIDINE ppm BENZENE % C8H18 50 250 1000 2 5 10 X(1) 0.0427 0.0427 0.0427 0.0427 0.0427 0.0427 0.0427 X(2) 0.346 0.0602 0.0155 1.06 1.30 1.35 1.06 X(3) 0.0295 0.0366 0.0407 0.0175 0.0234 0.0229 0.0138

presence CS2 % 1 0.0427 0.558 0.0265

4. R E A C T I O N K I N E T I C S Catalysts stabilized with a mixture of n-octane and hydrogen were used in experimental studies of product evolution with contact time. Experimental results, lumped by their carbon number, are shown in Figure 2.

404

-03

C7H16

C6H14

C4H10

C4H10

C8H18

C7H16

C3H8

1 .._ C5H12

-

4

-

C6H14

Figure 3 Selected reaction pathway to simulate reaction product distribution.

Since isomer distribution for each carbon number group does not vary significantly within the conversion range 0% to 40%, it was possible to simplify the model by restricting the simulation to those groups. Several reaction schemes were studied in order to select the most. appropriate scheme. Whenever possible, the selection was made by the Fisher test at a confidence level of 95%. The selected reaction scheme is shown in Figure 3. Experimental data was modeled assuming Langmuir-Hinshelwood and Rideal reaction mechanisms [11 ]. The set of reaction rate expressions corresponding to the reaction scheme in Figure 3, was selected using the Fisher test with a confidence level of 95% and it is shown in Table 2. A single site controlling step was assumed for each reaction in the scheme. Simulated data are shown in Figure 2 with solid lines. Table 2 Selected reaction rate expressions for the reaction scheme shown in Figure 3 rl

X(I ) * C8H18 I + X(13)* (C6H14 + C7H16 + C8H18) X(2)*C8H18-X(8)*C4H10 ^ 2 I + X(13)*(C6H14 + C7H16 + C8H18) X(3) * C4H 10 * C 3 H 8 - X(9) * C7H16 I + X(13)*(C6H14 + C7H16 + C8H18) X(4)* C3H8 ^ 2 - X(10)*C6H14 I + X(13) * (C6H14 + C7H16 + C8H 18) X(5) * C8H18 * C5H12 - X(I I ) * C6H14 * C7H16 I + X(13) * (C6H14 + C7H16 + C8H18) X(6) * C8H18 * C3H8 - X(12)* C6H14 * C5H12 I + X(13) * (C6H 14 + C7HI 6 + C8H18)

1'7

X(7) * C8H18 * C3H8 I + X(13) * (C6H14 + C7H16 + C8H 18)

405 Table 3 Kinetics and adsorption parameters for rate expressions on Table 2.

Parameter X[1] X[2] X[3] X[4] X[5] X[6] X[7]

Value 0.0115 0.00800 0.110 0.319 0.185 0.320 0.0196

Parameter X[8] X[9] X[lO] X[ll] X[12] X[13] -

Value 0.0521 0.0502 0.0203 1.93 1.18 192 -

5. CONCLUSIONS The kinetic model reproduces satisfactorily experimental results. Deactivation experiments seems to indicate that the mechanism of deactivation changes with the nature of the contaminant used. When a strong poison for active acidic sites like pyridine is used, the catalyst gets totally deactivated when its concentration is over 250 ppm. In this case, the deactivation is faster than with CS2, but not as fast as an acid base reaction should be. The behavior can be explained assuming that the pyridine reaction with acidic sites is a diffusion controlled phenomenon enhanced by its molecular size, which is very near to the zeolite pore size. The presence of a mixed mechanism of deactivation and inhibition is also evident.

NOTATION C -total number of compounds. C i -concentration of compound i (mole/m3) J -total number of reactions. m c -catalyst mass (kg) ni, n ~ -molar flux of compound i (mole/s). Rj -reaction rate for reaction j (mole/kg.s) otji -stoichiometric coefficient of compound i for reaction j. Pc

-catalyst apparent density (kg/rn3).

BIBLIOGRAPHY

1. 2. 3. 4. 5.

Kouwenhoven H.W., Advances in Chemical Series, N. 121, p. 529 (1973). Physical Constants of Hydrocarbon and Non-Hydrocarbon Compounds, ASTM Data Series D54B, 2nd.Edition. Kouwenhoven H.W and Van Zijll W.C., "Shell's Hydro-Isomerization Process", Chem. Eng. Prog., 67, p.65 (1971). Csicsery S.M., "Zeolite Chemistry and Catalysis", A.C.S. Monograph 171, J. Rabo De., Washington D.C., p. 680 (1976) Chen N.Y., Gorring R.L., Ireland H.R. And Stein T.R., Oil Gas J., 75, p.23 (1977).

406 6. 7.

Weitkamp J., Jacobs P.A. and Martens J.A., Appl. Catal., 8, p. 123 (1983). Sachtler W.M.H.and Zhang Z., "Zeolite Supported Catalysis", Advances in Catalysis, Acad. Press Inc., 39, p. 129 (1993) 8. Kouwenhoven H.W, Advances in Chemical Series, 121, p. 529 (1973) 9. Becerra M., "Transformaci6n de n-Octano sobre un catalizador M/H-MFI", Trabajo de Grado, Universidad Central de Venezuela, Octubre de 1994. 10. Rodrigues M.G.F., Magnoux P., Guisnet M. and Choudary V.R., Proceedings of the XV Simposio Iberoamericano de Cat~ilisis, C6rdoba, Argentina, V. 1,p. 91 (1996) 11. Papa J., Santos F., Le6n G. and Giannetto G., Proceedings of the 5th World Congress of Chemical Engineering, Vol I, p. 191 (1996). 12. Santos F., "Modelado cin6tieo de la transformaei6n eatalitiea de n-oetano sobre catalizadores a base de zeolitas", Trabajo de Grado, Universidad Central de Venezuela, Junio de 1996.

~

Elsevier Science B.V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

407

I s o b u t a n e A l k y l a t i o n with C40lefins: L o w T e m p e r a t u r e R e g e n e r a t i o n o f Solid A c i d Catalysts with O z o n e C.A. Querini, E. Roa, C. L.Pieck, J.M. Parera Instituto de Investigaciones en Cat/disis y Petroquimica (INCAPE) Sgo del Estero 2654 - (3000) Santa Fe - ARGENTINA

The regeneration of Y-zeolite catalysts used in isobutane alkylation with C4 olefins was studied. The coke formed on these catalysts during this reaction needs temperatures higher than 500~ to be burnt off with air. Ozone was used in this study to eliminate most of the coke at a much lower temperature. After a treatment at 125~ with ozone, the small amount of coke remaining on the catalyst can be removed with air at 250~ The ozone not only eliminates coke from the catalyst, but also modifies its burning characteristics as measured by Temperature Programmed Oxidation, shifting the peak to lower temperatures. This allows a combined treatment with ozone at 125~ followed by air at 250~ to restore the activity and stability of Y-zeolite catalysts for isobutane alkylation.

1. INTRODUCTION The alkylation of isobutane with C4 olefins using solid acid catalysts has become a growing research field during recent years. The main reason is that the currently used processes in industry have the HF or H2SO4 acids as catalysts, both of them being very difficult of being handled or disposed of; they also present severe problems for the environment. However, in spite of an important research effort carried out both by industrial (1-4) and academic (5) laboratories, it has been very difficult to solve the major problem that the solid acid catalysts present, which is the fast deactivation due to the coke deposition. Zeolites (6,7), heteropolyacids (8,9), sulfated zirconia (10,11), and other materials (12) have already been explored. All of these materials deactivate in a rather short time, ranged in the order of minutes to hours, and therefore any process involving solid acid catalysts for isobutane alkylation would require frequent regenerations. Most of the research reported in the literature was oriented towards analyses of product distribution, influence of operating conditions and reaction mechanisms. Recently, spectroscopic characterization of coke deposits formed on Y-zeolite catalysts was reported (13,14), and Stocker et al presented regeneration studies of the EMT catalyst, by heating in air up to 600~ (15). The alkylation reaction is carried out at low temperatures, typically lower than 80~ The coke deposit formed under these conditions has an aliphatic structure (13-15). The reason why this coke requires very high temperature to be fully eliminated is that when temperature is

408 increased above reaction temperature, in order to burn the coke off, the structure of coke changes from the originally aliphatic structure to an aromatic one. Therefore, as temperature is increased, the temperature needed to burn the coke shifts towards higher values (16). Temperatures higher than 500~ are needed to eliminate the coke. If a small amount of coke is left on the catalyst by carrying on partial regeneration at 500~ the initial catalyst activity is not restored. Even though this activity could be restored by treatments at 600~ (15), this would represent a major drawback for the process, taking into account that the reaction is carried out at temperatures lower than 100~ In this work, the regeneration with ozone of Y-zeolite catalysts, exchanged with lanthanum, is studied. The objective is to find a low temperature regeneration procedure for the solid acid catalysts used in the isobutane alkylation reaction.

2. EXPERIMENTAL 2.1 Catalysts Y-Zeolite (UOP, Y-54), with a Si/A1 ratio 5.3. The sodium form of the zeolite was exchanged with (NO3)3La 0.58 M, following two different procedures: - exchange with (NO3)3La in reflux, during 4 hours, drying at 100~ calcining at 550~ during 4 hours, and then a second ionic exchange step, using the same conditions as in the first one. The exchange level was 91%. This sample is labeled LCL-Y. - exchange with (NO3)3La in reflux, during 4 hours, drying at 100~ calcining at 550~ and then a second ionic exchange with NO3NH4, during 4 hours at reflux, drying and then calcining at 550~ The exchange level was 68% in the first step, and 2 1 % in the second exchange. After the second exchange, lanthanum was not detected in the final solution. This sample is labeled LCH-Y. 2.2 Activity test The alkylation reaction of isobutane with a mixture of C4 linear olefins was carried out in liquid phase at temperatures between 25 and 80~ and at 30 Kg/cm2, in a fixed-bed reactor. The space velocity was WHSV = 1 h1 referred to the olefins. The isobutane is premixed with the olefins. The molar ratio used in this study was 15. The C4 olefins fraction contains 38% 1-butene, 22% trans-butene, 14% cis-2-butene and 26 % isobutene. In order to analyze the products coming out of the reactor, a ten-loop valve was used to collect the sample to be analyzed after the run. Products are analyzed by GC, using a 100 m squalane column. Prior to the reaction, catalysts were pretreated in-situ, heating up to 250~ in an air stream. In selected runs, the catalyst was unloaded separated in fractions, in order to obtain information about the coke profile along the reactor, after the reaction or a~er a regeneration. 2.3 Coke Characterization Coke was characterized by Temperature-Programmed-Oxidation (TPO). These experiments were carried out using a modified unit. The CO2 produced during the coke burning is converted to CH4, in a methanator reactor. A H2 stream is fed to this reactor, which is loaded with a Ni catalyst, in order to quantitatively convert CO2 into CH4. This compound is then continously monitored by a flame ionization detector (FID). With this configuration the sensitivity and resolution of the classical TPO technique is greatly improved. Typically, 10 mg

409 of coked catalyst is loaded into a quartz cell, heating at 12~ using 5%O2/N2 as carrier gas. Additional details of the technique can be found elsewhere (17). 2.4 Regeneration Regeneration experiments were carried out both with discharged samples and in-situ. Air and ozone were used to remove coke deposits. The ozone was generated by flowing air between two electrodes at high electrical potential, using an equipment similar to that described by My and Sahghal (18). The stream coming out of this equipment contains 1% ozone approximately. 3. RESULTS AND DISCUSSION Figure 1 shows TPO results of LCH-Y and LCL-Y catalysts after a reaction time of 40 minutes at 80~ The catalysts contain 13.1 and 11.2 % of coke respectively, and display similar TPO profiles. The second peak in these TPO profiles corresponds to the coke that changed its structure as a consequence of the increase in the temperature (16). At temperatures higher than 600~ the coke is still burning (Fig. 1). In this case the analysis is dynamic, and therefore it could be expected that by heating up to a temperature lower than 600~ and holding it for some time, a large fraction of coke could be removed. Figure 2 A and B show results of regeneration experiments, using ozone and air. The labels in each curve indicate the regeneration temperature, and the %Coke remaining on the catalysts after the regeneration. All the samples were regenerated in the ozone containing stream, except the one with the AIR label in Figure 2 A. In this case the catalyst was regenerated by heating from room 20 temperature up to 500~ at 3.7~ and holding the final temperature for 33 minutes. After this regeneration treatment the catalyst recovered LCL-Y o activity, but the stability is much poorer > ~t than in the first cycle, as has been ,~ 10 t $ ~ reported in (16). This occurs in spite of z ~ LCH-Y t the catalyst having only 0.25%C after o '_~. 1 ~ I~9 the regeneration. Probably this occurs .I ar;', j ~" " because the coke that was not ] eliminated from the catalyst remains blocking the stronger acid sites, which .t'.: ~, -0 are required for the hydrogen transfer 0 200 400 600 800 reaction, a necessary step in the TEMPERATURE, *C alkylation reaction. It seems that if coke is not fully eliminated during the Figure 1-TPO profiles of coked LCL-Y and LCH-Y regeneration, it selectively deactivates catalysts, the acid sites involved in the alkylation reaction, and therefore a more efficient regeneration procedure is required. In Figure 2 A and B the TPO profiles for the LCL-Y and LCH-Y catalysts after an ozone treatment are shown. Most of the coke can be eliminated by this treatment at all the i

|

410 temperatures between 80~ and 180~ The amount of coke left on the catalyst after the ozone treatment at 125~ is similar to the amount left after an air treatment at 500~ However, an important difference can be observed in Figure 2A. This is that the main peak in the TPO profile is shifted towards lower temperatures. This shift is because during the regeneration with ozone, partially oxidized groups are generated on the coke. These groups

2

B

2

LCH-Y

180 ~

~, o >_.. -~1 "~ z

o >

a

u.

0.55%C

~z o

1.1.,

0 O0

200 400 TEMPERATURE, ~

600

~ 0

k

~ ~ 1 0.13%C 5 0'~-q,~ 200 400 TEMPERATURE, ~

600

Figure 2: (A) TPO profiles of regenerated LCL-Y and LCH-Y catalysts, regeneration in ozone 4 hours; in air 33 min.; (B) TPO profiles of LCH-Y catalyst, regenerated with ozone. are then more reactive and either they are more easily burnt with air or decompose at lower temperatures. Therefore the main peak in the TPO appears at rather low temperatures. The ozone generates atomic oxygen both in the gas phase and on the catalyst by decomposition, even at low temperatures (19), and this highly reactive species attack the coke at low temperatures. The double bonds are more reactive with the atomic oxygen, and therefore are the first bonds to be oxidized (20). In the examples shown in Figure 2A and B, with regeneration temperatures between 80 and 150~ the temperature of the main peak of theTPO profiles is between 250~ and 300~ approximately, depending upon the regeneration temperature. Since the pretreatment for the alkylation reaction is carried out at 250~ it can be expected that the combined treatment of ozone at 125~ followed by air at 250~ could remove most of the coke. Additionally, it has been found that after 03 treatment, oxygen remains adsorbed on the alumina surface (21), being possible that this adsorbed species plays a role in coke gasification when heating up to temperatures higher than the temperature used during the ozone treatment. Figure 2A and B also shows that as regeneration temperature in ozone increases up to 150~ the maximum in the TPO profiles shifts to lower values. This indicates that the coke remaining on the catalyst after the ozone treatment becomes more easily burnt as the pretreatment temperature in ozone increases, which is due to a higher degree of oxidation under this condition. The regeneration at 180~ is less effective than the regeneration at lower temperatures, as indicated in the TPO of Figure 2B. In this case, the TPO also shows

411

that the effectivity of ozone starts to decrease above 150~ The high temperature peak in this TPO clearly shows that there is an optimum temperature to eliminate coke with ozone. When the temperature is too high, in this case above 150~ ozone decomposes and oxygen atoms recombine very fast, with the consequent decrease in coke gasification rate. The effect of temperature during the ozone regeneration is shown in Figure 3. The amount of coke remaining on the LCH-Y catalyst, aiter the ozone treatment during 4 hours, is plotted as a function of the regeneration temperature. As the 0.a temperature is increased, the amount of coke decreases, but between 100~ and 160~ the variation is very small. 0.6 On supported metal catalysts, it has been found that the effectiveness of tu the ozone treatment presents a O maximum around 125~ Above this o 0.4, temperature, the amount of coke remaining on the catalyst a~er the m~.~ ~ ozone treatment increases, probably 0.2 ~ due to a very fast ozone 9 decomposition in the presence of platinum (22). Even with the empty 0 reactor, the amount of ozone sharply 80 100 120 140 160 180 decays above a given temperature TEMPERATURE, ~ (22), which depends on experimental conditions such as residence time, Figure 3- %Coke as a function of regeneration reactor material, etc. In our case, the temperature. Regeneration in 03, 4 hours, optimum temperature is between 130160~ At lower temperatures, the rate of reaction between ozone and the coke is low. At higher temperatures, the recombination of atomic oxygen atoms prevails over the reaction with coke. To verify if the ozone regeneration in-situ is as efficient as with the samples unloaded from the reactor, an experiment was carried out, with a reaction cycle and a regeneration in ozone during 4 hours at 125~ After this, the catalyst was unloaded and separated in fractions. Figure 4 A shows the TPO profiles for these samples. Samples taken from the top of the reactor have very low coke content, in the order of 0.2-0.3%, which is similar to the amount found on the sample regenerated out of the reactor. However, samples taken from the bottom of the reactor have higher coke content, and the TPO profile indicates that this coke needs high temperature to be gasified. This means that ozone was consumed on the top of the reactor and during most of the regeneration experiment the bottom of the reactor was not in contact with ozone. A part of this ozone is consumed by the coke, and another part decomposes either between the ozone generator or inside the reactor. Figure 4 B shows how the coke content changes along the reactor aider this regeneration, clearly indicating that the bottom of the reactor was indeed in less contact with the ozone. Nevertheless, on the bottom of the reactor the amount of coke was diminished from the 13.1% up to 2.94%. This profile could be improved, decreasing the amount of coke all along the bed, by optimizing the regeneration conditions, such as temperature and flow rate.

412

u~ A

4

o >

3

z 0 u~ ca

2

.J

Ii

1

0

0

200 400 TEMPERATURE, ~

600

0

0

20

40

60 %BED

Figure 4: A) TPO spectra of LCH-Y catalyst, regenerated with ozone at 125~ %coke as a function of the position along the reactor.

80

100

4 hours.; B)

Therefore, to assure that the full catalyst is in contact with ozone, a 16-hour regeneration experiment was carried out, followed by a second reaction cycle. Figure 5 shows the results of this experiment. The trimethylpentanes are produced during 30 minutes approximately, and after that time, dimethylhexenes start to appear. This change is due to the coke deposition that leads to the loss of hydrogen transfer capacity. The hydrogen transfer activity is an essential step for the alkylation reactions. When the acid sites involved in this reaction deactivate by coke, the catalyst loses its alkylation activity but it is still active for oligomerization reactions and therefore, the amount of dimethylhexenes in the C5+ fraction becomes dominant. Therefore, to assure that the full catalyst is in contact with ozone, a 16-hour regeneration experiment was carried out, followed by a second reaction cycle. Figure 5 shows the results of this experiment. The trimethylpentanes are produced during 30 minutes approximately, and after that time, dimethylhexenes start to appear. This change is due to the coke deposition that leads to the loss of hydrogen transfer capacity. The hydrogen transfer activity is an essential step for the alkylation reactions. When the acid sites involved in this reaction deactivate by coke, the catalyst loses its alkylation activity but it is still active for oligomerization reactions and therefore, the amount of dimethylhexenes in the C5+ fraction becomes dominant. After the 16-hour regeneration with ozone the catalyst recovered the initial activity and stability, with a second cycle essentially identical to the first one. Even though after the 4 hour ozone treatment a small amount of coke is left on the catalyst, the temperature of the maximum in the TPO spectrum of this coke is 250~176 approximately. As previously mentioned, the pretreatment after the regeneration is carried out at 250~ using air. This treatment is necessary in order to eliminate the water produced during the regeneration. When the catalyst regenerated in ozone at 125~ during 4 hours, which contains 0.29%C, is treated in air heating up to 250~ and holding this temperature during 2 hours, the amount of coke remaining on the catalyst is 0.13%C. The TPO maximum temperature of this coke (spectrum not shown) shifts from 300~ (see Fig. 2) up to 350~ This indicates that the

413 partial oxidized groups left on the coke after the ozone treatment, slowly decompose at 250~ and therefore the coke remaining on the catalyst after the air treatment requires higher temperatures to be removed. The ozone treatment carried out in this experiment was 16 hours long, in order to assure a complete regeneration. An optimization of the regeneration conditions, i.e. temperature and flow rate, could lead to a substantial reduction of this time. However, it is still comparable to that reported in (15), where a total of 15 hours was used in the regeneration in air, but with a final temperature of 600~ The catalysts active for the

A

9

B

i"

I

t* s

1

iiB.---mmiimlimlllml

~

O

"

0.5 '5 o

0 0

10

20 TIME, min

30

40

-.=

- - 0

10

20

30

TIME, min

Figure 5:

Activity results for LCH-Y catalyst, (A) Fresh, (B) after regeneration in ozone during 16 hours at 125~ Reaction conditions: 80~ 30 Kg/cm2, WHSV = lh ]. isobutane alkylation should display an acid strength distribution not quite different from the zeolites catalysts. Therefore, it can be expected that any catalyst that could be developed will require very high temperatures to be regenerated in air. In some cases, such as the heteropolyacids, the catalyst structure collapses at a temperature much lower than the one needed to burn the coke in air. Under these circunstances, low temperature regeneration procedures must be used in order to preserve the catalyst integrity. Even with the zeolites, when heating up to 600~ conditions must be carefully controlled, since under the regeneration conditions in air, the structure could be destroyed. After the second cycle, the catalyst was unloaded, separated in fractions and analyzed by TPO. The coke content is slightly higher on the top of the bed; 16% on the top, and 14.3% on the bottom. The reason why after this second cycle the amount of coke is higher than after a normal cycle is unclear.

4. CONCLUSIONS The zeolite catalysts deactivate very fast during the isobutane alkylation with C4 olefins due to coke deposition. This coke requires very high temperature and long times to be fully eliminated in air. However, the regeneration in ozone can be carried out at low temperatures,

414 in the range of 100-150~ being possible under theses conditions to modify both the amount of the coke and the nature of the coke that remains on the catalyst after the ozone treatment. This is revealed by the TPO analyses, that show that after the ozone regeneration the coke has a maximum in the TPO spectra around 250-300~ Therefore, if a small amount of coke is left on the catalyst after the ozone treatment, some of it is further eliminated during the pretreatment of the catalyst before the next reaction cycle. However, if regeneration in ozone is long enough all the coke is eliminated, and the catalyst recovers the activity and the stability. The low temperature regeneration procedure using ozone could be an option to be considered for a process of isobutane alkylation with solid acid catalysts.

ACKNOWLEDGMENTS

The authors greatly acknowledge the financial aid of Universidad Nacional del Litoral and thank JICA (Japan International Cooperation Agency) for the support to this project.

REFERENCES

1- US Patent 5, 414,18 7. 2- US Patent 5,157,197 3- EP Patent 650394 4- US Patent 5, 22 O,095 5- A. Corma, A. Martinez, Catal. Rev.-Sci. Eng.35(4)(1993)483 6- A. Corma, A. Martinez, C. Martinez; J. Catal. 146(1994)185 7- F. Cardona, N.S. Gnep, M. Guisnet, G. Szabo, P. Nascimento, Appl. Catal. A: General 128(1995)243 8- T. Okuhara, M. Yamashita, K. Na, M. Misono; Chem. Letters (1994) 1451 9- N. Essayem, S. Kieger, G. Coudurier, J. Vedrine, St. Surf So. and Catalysis 101(1996)591 10- C. Guo, S. Yao, J. Cao, Z. Qian; Appl. Catal. A: General 107(1994)229 11- A. Corma, M.I. Juan-Rajadell, J.L6pez-Nieto, A. Martinez, C. Martinez; Appl. Catal A: General 111(1994)175 12- US Patent 4,180,695 13- C. Flego, I. Kiricsi, W.O. Parker, M.G. Clerici; Appl. Catal. A: General 124 (1995)107 14- C. Flego, L. Galasso, I. Kiricsi, M.G. Clerici; St. Surf So. and Catalysis 88(1994)585 15- T. Rorvik, H. Mostad, O.H. Ellestad, M. Stocker, Appl. Catal. A: General 137(1996)235 16- C.A. Querini, E. Roa; Appl. Catalysis A: Genera# submitted 17- S.C. Fung, C.A. Querini, J. Catal. 138(1992)240 18- N.L. My, P.N. Sahghal; The Chem. Eng. Journal, 40(1989)15 19- J.I. Steinfeld, J.S. Francisco, W.L. Hase, Chemical Kinetics and Dynamics, Prentice Hall, Englewood Cliffs, New Jersey, (1989) 105 20-Kirk, Olhmer, Encyclopedia of Chemical Technology, 3rd Ed., 16(1981)683 21- A. Klimovskii, A.Bavin, V.Tkalich, A.Lisachenko, React. Kinet. Catal. Lett. 23(1983)95. 22- C.L. Pieck, E.L. Jablonsky, J.M. Parera, St. Surf So. and Catal. 88(1994)289

~

Elsevier Science B.V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

415

Influence o f Chloride during Coke Burning of a N a p h t h a R e f o r m i n g Catalyst C o k e d in a C o m m e r c i a l Cycle M.C. Rangel', M.N.M. Barbosa ~, C.L. Pieckb and N.S. Figolib aInstituto de Quimica, Universidade Federal da Bahia, Federaq~o, 40 170-280 Salvador, Bahia, Brazil blNCAPE, Instituto de Investigaciones en Catfilisis y Petroquimica, Santiago del Estero 2654, 3000 Santa Fe, Argentina

The effect of chloride addition during coke burning-off on the regeneration of the catalytic functions of a Pt-Re-S/AI203-C1 catalyst was studied. Samples were collected in an industrial unit during the fourth operation cycle, in which naphtha cuts of C6-Cs and C7-C9 were processed. It was found that regeneration with the highest chloride concentration produces the highest activity of the acid function. The activity of the metallic function is easily recovered and is not influenced so strongly by chloride. The described regeneration procedure assures the recovery of both catalyst functions in only one step operation, saving time and preserving the catalyst. However, large amounts of chloride are required to regenerate all particles, since the efficiency of the process depends on the diffusional limitations of chloride in the outermost layers of the pellets.

1. INTRODUCTION The demand for enhancing the octane number of gasoline, as well as for providing highvalue aromatic hydrocarbons (such as benzene, toluene and xylenes) is largely met by naphtha reforming. It is also an important source of hydrogen for hydrotreating processes. The chemistry of reforming involves several reactions which are promoted by a bifimctional catalyst, in which the metallic function (platinum) catalyses hydrogenation-dehydrogenation and hydrogenolysis reactions and the acid function (chloride-modified alumina) has enough acidity to promote isomerization and cracking. Both functions are required for isomerization and dehydrocyclization. Some undesirable reactions also occur, such as hydrocracking and coke formation. In commercial catalysts, a second metal was introduced allowing operation at higher temperature and lower pressure, since it improves selectivity to aromatics and prevents coke formation [1]. Many catalysts such as Pt-Re/AI203, Pt-Sn/AI203, Pt-Ir/Al203 and PtGe/A1203 have been proposed, but the rhenium-based one is by far the most used in petroleum industry worldwide. In spite of its high stability and selectivity, this catalyst deactivates with time, mainly due to coke deposits. To compensate this decay of activity, the operation temperature is increased up to a maximum value recommended by the technology of the process; at that time the catalyst must be submitted to a regeneration process to restore some

416 of its original properties. By this procedure, coke is burned and chlorine added and the acid and metallic functions are restored. Coke deposition affects both catalyst functions. It has been accepted that coke formation starts on metal sites and their surroundings and continues on the support [ 1, 2]; it was stated that coke on the metal is more hydrogenated and is eliminated first during the regeneration process, while the coke on the support requires longer times and higher temperatures to be removed [3- 5]. A lot of work has been carried out dealing with the effect of coke on activity and selectivity as well as with coke burning [2-7]. However, there are comparatively few papers about regeneration, despite its economic importance. Because of the reforming conditions, most of the catalyst damage occurs during regeneration, which demands for optimizing this process. The steps commonly involved in catalyst regeneration are: elimination of coke by controlled burning, restoration of the acid and metallic function by chlorination, reduction with hydrogen and passivation. However, some technologies recommend the addition of chloride during coke burning. Almost all the works, in the open literature, deal with coke burning followed by oxychlorination [4, 8-10]; as the coke burning leads to the major injuries of the catalyst, it is important to investigate what happens to the acid and the metallic functions when chloride is added during this step. In this work, the effect of chloride addition during coke burning on the regeneration of the catalytic function of a Pt-Re/A1203 catalyst, was studied. It also intends to evaluate the possibility of using a smaller amount of chloride so that the environmental damage in discarding the chloride compounds may be reduced.

2. EXPERIMENTAL

A Pt-Re-S/A1203 (0.22%Pt and 0.44% Re) commercial catalyst was used. Samples were coked in an industrial unit during the fourth operation cycle of the catalyst, in which naphtha cuts of C6-C8 and C7-C9 were processed. They were collected after one (sample 1), ten (sample 10), seventeen (sample 17) and twenty-one (sample 21) operation months. Carbon contents of each sample was measure using a LECO equipment. In the regeneration experiments, about 10 g of catalyst were heated under nitrogen up to 400~ then, the nitrogen stream was stopped and a 5% oxygen in nitrogen stream was introduced and maintained during 6.5 h at 400~ This oxidation stream passed through two saturators containing water and dichlorometane (the last one could be kept at different temperatures). Three types of regeneration experiments were performed: i) without chloride (F); ii) using 0.03% chloride (M), which is the concentration used in the commercial plant and iii) using 0.01% chloride (I). After this, the chloride content was analyzed by the modified Volhard-Charpenter method [11] and the solids were characterized by means of temperature programmed reduction using a Ohkura TP 2002 S equipment. The activity of each catalytic function was determined by means of toluene hydrogenation and n-pentane isomerization, reactions closely related to the activity of the metallic and acid functions, respectively. Some samples were not homogeneous but contained pellets of different colors. In these cases the brown and black particles were analyzed by temperature programmed oxidation using a specially designed equipment [12]. The toluene hydrogenation reaction, was carded out with 0.15 g of the catalyst sieved to 30-80 mesh using a glass tubular reactor operating at atmospheric pressure, 40~ WHSV 1.5

417 and hydrogen/toluene molar ratio 50. For n-pentane isomerization, the conditions were 500~ WHSV 4.5 h-1 and hydrogen/n-pentane molar ratio 6. Before the tests, the catalysts were reduced in hydrogen at 500~ for 2 h.

h "1

3. RESULTS AND DISCUSSION

The samples showed quite different colors as a function of chloride concentration after regeneration. The catalysts treated with the maximum chlorine concentration (0.03%) contained white pellets whereas those regenerated with the intermediate amount (0.01%) were made of white, brown and black particles. On the other hand, the samples heated without chloride were homogeneous and black. All dark particles (brown and black) did not have any coke as indicated by TPO analysis and had a low amount of chloride. It means that the diffusion in the outermost layers of the pellet determines the efficiency of the process. Because of this, all TPR experiments, as well as the activity tests were carried out using only white pellets, in the case of the samples regenerated with the intermediate chlorine concentration. Table 1 shows the results of the catalytic test reactions as well as carbon and chlorine contents. It can be noted that samples regenerated with a stream containing 0.01% and 0.03% chloride show almost the same amount of chloride. In both cases these results correspond to white pellets and we can thus conclude that low amounts of chloride (e.g. 0.01%) are enough to produce particles with concentrations close to that in fresh catalysts. However in this case, some dark particles with a concentration as low as that in samples burnt without chloride are also obtained. It confirms that, in the regeneration experiments performed with the intermediate value of chloride, there is a diffusional limitation in the outermost layers of the pellets. From these data we can see that, for samples collected atter one month of commercial operation (sample 1), the acid function is completely restored during coke burningchlorination, whereas the metallic function achieves higher values of activity, when compared with the fresh catalyst. Atter ten months of operation (sample 10), regeneration using the highest chloride concentration produces the complete recovery of the acid function. For samples collected aiter seventeen months (sample 17), regeneration of the acid function strongly depends on the chloride concentration and this influence is lower for the metallic function, which achieves activities even higher than the one of the fresh catalyst for M type regeneration. ARer twenty-one months of commercial operation (sample 21) almost the same activity is obtained, when chloride is present during burning. These results show that, in general, regeneration with the highest chloride concentration produces the highest regeneration of the acid activity. The activity of the metallic function is easily recovered and is not influenced so strongly by chloride amount. It means that small amounts of chloride are enough to promote the formation of the surface complexes [PtO2] and [ptIVOxCly]. As pointed out [13, 14] they are responsible for a well-dispersed metallic platinttm In some cases, the activity of the metallic function is even higher than that of the fresh catalyst, perhaps because some sulfur is lost from the metallic function during regeneration, thus producing an increase in activity [4]. It can also be noted that the regeneration of the samples is related not only to the chloride concentration but also to the nature of coke which varies during the run as a consequence of the different steps of its formation and accumulation. Therefore, the described regeneration procedure assures the recovery of both catalyst functions in a single step operation, saving time and preserving the catalyst.

418 Table 1 Coke content before regeneration and activity of the acid and metallic functions after coke burning with three chloride concentrations Carbon content Chloride content n-pentane Toluene Catalyst sample (%) after regeneration conversion conversion Fresh

-

1.6

(%)

29 28 16 14 29 19 7 25 8 4 20 22 13

3.2 3.5 2.4 2.0 2.5 2.6 2.3 4.8 2.7 1.7 3.3 3.9 2.8

1.01 0.20 1.02

10M

10I 10F

(%)

1.02

1M

lI 1F

(%)

1.10

1.00 0.17 17M 0.99 17I 10.9 0.70 17F 0.17 21M 0.99 21I 15.8 0.98 21F 0.21 M = regenerated with 0.03% (maximum value) I = with 0.01% (intermediate value) F = l~ee of chloride 6.0

The TPR curves show similar profiles for samples regenerated under the same conditions. Figure 1 illustrates this behavior for the catalysts collected after twenty-one months of operation. In all cases, there is a peak around 200~ assigned to platinum reduction as well as to the reduction of rhenium in the platinum surroundings [13]. At higher temperatures, different profiles are noted: in the case of samples regenerated in the presence of chloride, there is a peak around 600~ attributed to the platinum-rhenium alloy formation, whereas the other curve show a double peak associated to rhenium reduction as well as to alloy formation. Therefore, in the absence of chloride part of the rhenium remains segregated. As reported earlier [15], this alloy is responsible for the catalytic performance of the bimetallic system not only at atmospheric pressure but also at the elevated pressure typically used in industrial reforming. This can explain why the metallic sites regenerated without chloride are less active towards toluene conversion.

4. CONCLUSIONS The addition of chloride during coke burning provides an efficient method of regenerating the Pt-Re-S/AI2OrC1 catalyst, since it promotes the recovery of both acid and metallic functiom. Use of the highest chloride concentration produces the highest regeneration of the acid function. The activity of the metallic function is easily recovered and does not depend strongly on the chloride concentration. The presence of chloride favors the platinum-rhenium alloy formation improving the activity of the metallic sites. However, the regeneration procedure described is controlled by the diffusional limitations in the outermost layers of the pellet so that large amounts of chloride are needed to regenerate all the particles. We can thus

419 I

'

I

'

I

'

I

'

I

'

600

.,o

400 ...~

i: I 9 0;

o.

,: I"

#:

200

.-

. 9

:

9

"...'...:... . . . . ..

!,

,,,

-.%

.-

~.

'~

o

c,.)

I

l

0

I

200

Y~.~..

.6.

,: e-

-

',~

%.

i

I

400

l

I

600

J

I

800

l

1000

Temperature, ~ Figure 1. TPR curves of samples collected twenty-one months of operation regenerated without chloride ( m Sample 21 F) and with 0.01% (--- Sample 21 I) and 0.03% chloride (... Sample 21 M)

conclude that the large amounts of chloride used in industrial plants are probably needed because of diffusional features rather than because of chemical restrictions to the recovery of the sites.

ACKNOWLEDGEMENT The authors thank the financial support from the RHAE/CNPq Program to this work.

REFERENCES

1. J.M. Parera and J.N. Beltramini, J. Catal., 112 (1988) 357. 2. J. Barbier, in B. Delmon and G.F. Froment (eds.), Studies in Surface Science and Catalysis, vol. 34, Catalyst Deactivation 1987, Elsevier, Amsterdam, 1987. 3. J.M. Parera, N.S. Figoli and E.M. Traffano, J. Catal., 70 (1983) 484. 4. C.L. Pieck, E.L. Jablonski and J.M. Parera, Appl. Catal., 70 (1991) 19.

420 5. C.L. Pieck, E.L. Jablonski, R.J. Verderone and J.M. Parera, Appl. Catal., 56 (1989) 1. 6. D. Espinat, E. Freund, H. Dexpert and G. Martino, J. Catal., 1126 (1990) 496. 7. W.A. Groten, B.W. Wojciechowski and B.K. Hunter, J.Catal., 138 (1992) 343. 8. C.L. Pieck, E.L. Jablonski and J.M. Parera~ Appl. Catal., 62 (1990) 47. 9. J. Barbier, D. Bahloul and P. Marecot, Catal. Lett., 8 (1991) 327. 10. A. Parmaliava, F. Frusteri, A. Mezzapica and N. Giordano, J. Catal., 111 (1988) 235. 11. N. S. Figoli, M.R. Sad, J.N. Beltramini, E. L. Jablonski and J.M. Parera, Ind. Eng. Chem~ Prod. Res. Dev., 19 (1990) 545. 12. S.C. Fung and C.A. Querini, J. Catal., 138 (1992) 240. 13. H. Lieske, G. Lietz, H. Spindler and J. VSlter, J. Catal., 81 (1983) 8. 14. G. Lietz, H. Lieske,H.Spindler, W.Hanke and J.V~lter, J. Catal., 81 (1983) 17. 15. K. Shum, J.B. Butt and W.M.H. Sachtler, J. Catal., 196 (1985) 371.

9 Elsevier Science B.V. All rights reserved Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

421

Deactivation o f P t - A u / A 1 2 0 3 Catalysts Prepared by Surface R e d o x Reaction: Effects o f Sulfur and Coke Deposition G.Espinosa a, G. Del Angel b, J. Barbier c, P. Marecot c and I. Schitter a aGcia, de Catalizadores, IMP, Eje Central No.152, 07730 M6xico, D.F., M6xico. bDepto de Quimica, UAMI, Apdo. postal 55-534, 09340 M6xico, D.F., M6xico c Laboratoire de Catalyse en Chimie Organique, URA CNRS 350 Facult6 des Sciences, 40 avenue du Recteur Pineau, 86022 Poitiers, France.

The effect of the coke and sulfur deposition on Pt-Au/AI203 catalysts for the hydrogenolysis of methylcyclopentane was studied. A higher resistance to the deactivation by coke was observed on the catalysts with medium and high content of gold (0.097, 0.35 wt%), with respect to the low content one (0.013 wt%) and to the Pt reference catalyst (1.0 wt%). The selectivity pattern was modified at high content of Au, being the methylcyclopentene yield increased. The sulfured Pt-Au catalysts showed lower activity and deactivation constant than the non sulfured ones. An important change in the selectivity pattern was also observed: the formation of methylcyclopentene was favoured with the content of Au.

1. INTRODUCTION In recent years numerous studies have been concerned to supported bimetallic or multimetallic catalysts, since their applications on the hydrocarbons conversion are largely known. One expect superior properties to those obtained with the monometallic compositions, i.e. higher selectivity and resistance to deactivation. The carrier and the preparation method of the multimetallic catalysts are reported as the most important variables for the final catalytic properties. The multimetallic catalysts have been prepared mainly by two methods: i) by coimpregnation of the support with a solution containing the metallic precursors, and ii) the successive impregnation method, in which the metal impregnation order can be modified. Depending on the preparation method used, separated metal particles, homogeneous alloys, segregated alloys or all them can be obtained during bimetallic preparation. Obviously a controlled preparation of bimetallic catalysts is needed in order to understand its role upon activity, selectivity and deactivation resistance. On this way, the controlled formation of surface bimetallic particles has been reported in catalysts prepared by the redox method [ 1]. In the present work, in order to define the role of the redox method in the surface properties of the Pt-Au alumina supported catalysts, we report the preparation, characterization and catalytic properties of a set of bimetallic catalysts with different gold content. The catalysts were evaluated using methylcyclopentane (MCP) hydrogenolysis as the test reaction.

422 2. EXPERIMENTAL

Monometallic catalyst. A low area alumina (SCS69, 60 m2/g) was used as the support. A 1 wt % Pt/AI203 was prepared by the impregnation technique using an aqueous solution of H2PtC16followed by drying overnight at 393 K. This solid was calcined in flowing air at 723 K for 4 h and reduced in flowing hydrogen at 773 K. The dispersion of this catalyst was 60 % determined by the hydrogen adsorption volumetric method. One sample of this catalyst was treated at 823 K for 3 h under oxygen-nitrogen flow (02 1% - N2 99 %) and then reduced in flowing hydrogen at 773 K for 6 h, as a result the dispersion decreased up to 44%. Bimetallic catalysts. The catalysts were prepared by the redox method [ 1] which consists in the reduction of AuCI4 by the preadsorbed hydrogen on the platinum surface. A slurry of the monometallic catalyst was put in contact with a solution of HCI (pH=l). The slurry was purged with flowing nitrogen for 30 min. Afterwards the hydrogen was introduced by bubbling (1 h) and a further nitrogen treatment was done in order to eliminate the dissolved hydrogen and the reversible adsorbed one. The AuCI4 solution with the desired concentration of gold was put in contact with the monometallic slurry. After the redox reaction the sample was filtered, dried overnight at 393 K, and finally reduced in flowing hydrogen at 773 K for 6 h. The nominal composition of the catalysts was 0.06, 0.12, and 0.48 wt % gold for the 0.5M, 1M and 4M respectively Reference catalyst. This catalyst was treated under identical preparation conditions used for the redox catalysts, except for the reaction with the AuC14. Such treatment produced an additional sinterization of the Pt as it has been reported [2]. The poisoning of the catalysts was carried out by impregnation with a (NH4)2804 solution and further reduction under hydrogen at 773 K. Catalytic activity. Methylcyclopentane (Aldrich, 98 % purity) conversion was carried out at atmospheric pressure and 623 K of temperature in a fixed bed continuous flow reactor. The conversion was kept lower than 15 %. The saturation temperature of the reactant was 273 K. The reaction products were analyzed on-line using a Perkin-Elmer 8410 gas chromatograph, equipped with a FID detector and a 50 m fused silica capillary column with methyl-5%phenyl-silicon coating. Turnover frequency (TOF), is defined as the number of molecules of methylcyclopentane converted per site per second (obtained at time zero of reaction); selectivity is defined as the molar ratio of product formed to methylcyclopentane reacted. The absence of diffusion limitations was verified by using different gas flow rates and weight of the catalyst. 3. RESULTS AND DISCUSSION In Table 1 the gold contents deposited on the Pt surface are reported. This method of preparation allows to obtain bimetallic catalysts in which the two metals are in close interaction. It can be seen the addition of a very low load of gold decreases the accessibility of Pt in a significant way. The Pt accessibility decreases with gold content. To avoid the chlorine role in catalytic properties, the chloride contents are also reported in Table 1, and was almost the same for all bimetallic catalysts.

423 Table 1 Pt-Au/A1203 bimetallic catalysts prepared by the redox method.

Catalysta

wt % Au

Reference 0.5 M 1M 4M

0.013 0.097 0.350

Accessibility %

34.0 17.1 14.9 6.5

wt % CIb

0.54 0.42 0.77

a) Pt 1%; b) Determined by X-ray fluorescence.

It is important to note that the results in activity and selectivity were obtained at low conversion (< 15%) to avoid mass and heat diffusion effects. The results then are not a function of the conversion degree. In Table 2 we report the activity values (TOF) for the hydrogenolysis of methylcyclopentane. It can be observed that the gold content on the Pt-Au catalysts has an important effect on the activity, with respect to the reference catalyst. A synergetic effect can be observed, the catalyst with the lower content of gold is slightly more active than the reference Pt/A1203 catalyst. Synergetic effects have been reported for several bimetallic catalysts and reactions. The synergetic effect on alloyed catalysts has been explained in terms of electronic or ligand effects [3,4]. Coke formation on the metallic surface can occur during the hydrocarbon conversion. The Pt and Pt-Au catalysts deactivates steadily at 623 K, and the catalytic decay followed the second order hyperbolic relation [5]. In order to make a more quantitative analysis of the deactivation

Table 2 Activity and deactivation constants of Pt/Al203 and Pt-Au/A1203 catalysts for the hydrogenolysis of methylcyclopentane at 623 K.

Catalyst

khxl 0 6 (moUs g)

kexl 0 4 (s "l)

Reference 0.5M 1M 4M

1.03 2.98 0.6 0.27

2.0 2.5 2.7 3.0

k'dxl 010 (kekh)

2.1 7.4 1.62 0.81

TOFxl 0 3 (s -1)

59 342 79 81

424 deactivation process, the model of Levenspiel [6] was applied to obtain the rate constant of deactivation in similar way to that reported in the literature [7,8]. Taking into account that the reaction order relative to the MCP is zero, and that hyperbolic decay correspond to a second order, application of the Levenspiel model led to the following equation ICe= (kd/kh)(Cb)p (Fb/w) where ke is the experimental rate constant of deactivation obtained from the hyperbolic law, kd is the deactivation rate constant, kh is the rate constant for the hydrogenolysis reaction, Cb is the methylcyclopentane concentration (mol/1), Fb is the feed rate of methylcyclopentane (mol/ s), w is the catalyst weight (g) and p is the order of deactivation relative to methylcyclopentane. Considering that the feed rate, the concentration of methylcyclopentane and the catalyst weight were kept constant in all experiments, one can write" k'd = kd(Cb)p (Fb/w) and therefore: k'd = ke(kh) The values of the experimental slope (ke), the initial rate of reaction (kh), and the deactivation constant (k'd) are reported in Table 2. The deactivation constant k'd is highest for the catalyst with the lowest Au content. The ke experimental rate constant of deactivation has the same order for all catalysts, the highest k'd (kekh) value corresponds to the catalysts showing the higher reaction rate. At higher Au contents the deactivation k'd constant decreases significantly. The 1M and 4M catalysts show a higher resistance to self poisoning, the k'd values are lower than those observed for the reference monometallic Pt/A1203 catalyst. Then it seems that for the bimetallic catalysts the coke formed during the reaction decreases as the Au content of the catalyst increases [9]. It has been pointed out that the coke formation on the metallic surface occurs mainly on the fiat planes (large particles) [10]. If we assume that Au at high coverage is deposited on the dense planes of the platinum particles, we can propose that Au acts as a diluent of the platinum ensembles, and hence that coke formation is inhibited. The level of coke and its nature are under study, and the results will be reported in a subsequent paper. The selectivity for the hydrogenolysis of methylcyclopentane (MCP) on the Pt-Au catalysts is reported in Table 3. The hydrogenolysis of MCP yields mainly 2Methylpentane (2MP) and 3Methylpentane (3MP). The ratio 2MP/3MP varies from 2.1 to 2.7, it means that the statistical rupture of the bisecundary C-C bonds (2MP/3MP=2) is almost maintained. The addition of Au to Pt does not modifies this ratio. The formation of nHexane (nil) decreases at low Au content and is not modified even at high contents of Au. Benzene (Bz) is formed at very low proportion. The selectivity towards methylcyclopentene (MCP =) is favored with the addition of gold, the 4M catalyst presents the higher yield of this product. It is well known that the cyclo-olef'ms are precursors of coke [ 11 ], then we would expect a higher deactivation of Pt-Au catalysts, however, the deactivation constant is smaller. Such results suggest that Au

425 Table 3 Selectivity pattern for the hydrogenolysis of methylcyclopentane on Pt-Au/A1203 catalysts at 623 K.

Catalyst

2MP

3MP

nH

MCP--

Bz

Reference 0.5M 1M 4M

50.1 53.1 57.7 42.8

21.7 24.7 21.4 20.1

20.0 12.4 11.3 11.4

2.4 7.6 9.3 24.4

5.7 2.1 1.2

3MP/2MP

2.3 2.2 2.7 2.1

hinders coke formation favoring the desorption of MCP--, and inhibits the catalysts deactivation. The deactivation constant for the sulfured catalysts are reported in Table 4. It can be observed that the specific activity (kh) of the sulfured catalysts is lower than that shown by the non sulfured catalysts (Table 1). However, the deactivation constants k'o of the sulfured catalysts are lower, implying that the bimetallic Pt-Au are the less affected by sulfur poisoning. The addition of sulfur produces a modification in the selectivity pattern for all catalysts. An important increase in the selectivity towards the MCP- is observed. The formation of the principal products obtained on the reference Pt catalysts, 2MP and 3MP decreases. The ratio 2MP/3MP decreases as the gold content increases on the sulfured Pt-Au catalyst. It is well known that cyclopentenes are coke precursors [4], because of their fast polymerization over the metal particles without desorption [ 11 ]. Therefore, it can be assumed that sulfur restrain the polymerization of the dehydrogenated cyclocompounds by weakening the double bond-metal interaction and favoring the desorption of the MCP--. Similar effects

Table 4 Deactivation constants of the sulfured Pt/A1203 and Pt-Au/AI203 catalysts for the hydrogenolysis of methylcyclopentane at 623 K.

Catalyst

Reference 0.5M 1M 4M

khxl 0 6 (mol/s g)

kexl 0 4 (s "l)

k'dxl 010

0.369 0.15 0.13 0.05

3.9 1.2 0.3 0.0

1.44. 0.18 0.04

(kekh)

426 Table 5 Selectivity pattern for the hydrogenolysis of methylcyclopentane on sulfured Pt-Au/A1203 catalysts at 623 K.

Catalyst

2MP

3MP

nH

MCP--

Reference 0.5M 1M 4M

62.0 52.0 34.4 21.0

20.6 18.2 20.1 0.0

10.1 2.2 6.1 1.7

7.3 27.5 39.5 77.3

2MP/3MP

3.0 2.9 1.7 -

can be observed during the Pt clusters dilution produced by gold. However, if we compare the gold effect with that produced by sulfur on the reference catalysts (Tables 3 and 5), we can observe that the sulfur effect is more important than the gold effect. When both effects, gold dilution effect and electron sulfur effect act at the same time, the cyclolefm desorption is magnified and the deactivation of the catalysts is diminished. These results are of great importance since to the best of our knowledge, no detection of large amounts of cycloolefins have been reported before. Certainly the redox preparation method used in the present work had an important role in the obtained results. The controlled preparation of bimetallic catalysts let us obtain a bimetallic surface in which the interaction between the metals is magnified.

ACKNOWLEDGMENT This work was supported by the "Program Cooperative of Post-graduation" (PCP) FranceM6xico.

REFERENCES

1. J.C. Menezo, M.F. Denanot, S. Peyrovi and J. Barbier, Appl. Catal., 15 (1985) 353. 2. J. Barbier, P. Marecot, G. Del Angel, P. Bosch, J.B. Boiteaux, B. Didillon, J.M. Dominguez, I. Schiller, G. Espinosa. Appl. Catal. A, 116 (1994) 179 3. S.H. Inami and H. Wise, J. Catal., 26 (1972) 92. 4. G.C.Bond and D.E. Webster, Platinum Met. Rev., 9 (1965) 12 5. J.E. Germain and R. Maurel, C.R. Acad. Sci. Paris Ser. C., 247 (1958) 1854. 6. O. Levenspiel, J. Catal., 25 (1972) 265. 7. S. Fuentes and F. Figueras, J. Catal., 54 (1978) 397. 8. M. Viniegra, V. Arroyo and R.Gomez, Appl. Catal., 44 (1988) 1 9. J.W.A. Sachtler ad G.A. Somorjai, J. Catal., 89 (1984) 35. 10. A.J. Den Hartog, M. Deng, F. Jongerius and V. Ponec, J. Mol. Catal., 60 (1990) 99. 11. J. Barbier, Appl. Catal., 23 (225)1986.

9 Elsevier Science B. V. All rights reserved Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

427

Nature o f Coke F o r m e d from Tri-isopropylbenzene Over U S Y Zeolites with Different Rare Earth Content C. A. Henriques a, E. FalabeUa S. Aguiarb, M. L. Murta Valler S. Varelad and J. L F. Monteirod a Instituto de Quimica, UERJ, Rua S~o Francisco Xavier, 524, CEP: 20559-900, Rio de Janeiro, ILl, Brasil b DICAT/GEAPRO - CENPES/PETROBRAS, Cidade Universit6da, Quadra 7, CEP" 21949900, Rio de Janeiro, RJ, Brasil cIndependent Consultant d NUCAT - Nficleo de Cat~lise, COPPE/UFRJ, CP68502, CEP: 21.945-970, Rio de Janeiro, RJ, Brasil

The nature of the coke formed on Y zeolites with different rare earth content during the cracking of triisopropylbenzene was studied. The results showed that increasing rare earth contents decreases both total and insoluble coke build-up. On the other hand, the main components of the soluble fraction were polyaromatics compounds with general formula CnH2n-12, C H2n-22 and CnH2n-32 for all samples. TPO/DSC profiles suggested that rare earth elements present a catalytic effect on promoting coke oxidation.

1. INTRODUCTION Y zeolites exchanged with rare earth cations are widely used as the active component for cracking catalysts in petroleum industry. Such cations improve the catalytic activity and the gasoline yield, lowering the gas production and the coke formation. They also promote the thermal and hydrothermal stability of the catalyst. The use of heavy feedstocks in FCC processes significantly increased the interest for studies about the chemistry of the cracking of voluminous molecules. However, there are not many works reported on the use of large molecules in model reactions. This approach has been used by Sousa-Aguiar and Murta Valle [1,2] who studied the use of 1,3,5-triisopropylbenzene (TIPB) as a tool to evaluate the accessibility in zeolites. They observed [2] different deactivation rates due to coke deposition for USY and CREY (calcined rare earth) zeolites. On the other hand, Henriques and Monteiro [3-6] have been giving special attention to the characterization of the coke responsible for the deactivation of mordenites using small hydrocarbon molecules, using the methodology developed by Guisnet and co-workers [7-12]. Thus, this study was developed aiming at the chemical identification of the coke formed from 1,3,5-triisopropylbenzene on USY and CREY zeolites with different rare earth contents, searching for a better comprehension of the way deactivation by coking takes place.

428 2. E X P E R I M E N T A L

Coke formation from TIPB (Fluka 92075) was investigated on three samples of Y zeolites. A NaY zeolite synthesized according to Ferreira [12] was the parent sample. It was ion-exchanged with a NH4C1 solution and hydrothermally treated with steam at 923K (USY). Alternatively, it was either exchanged with NH4C1 and next with a REC13 solution (CREY-1) or only with the RECI3 solution (CREY-2). Both samples were steam calcined at 923K. Zeolite composition was determined by X ray fluorescence, in a XRF Phillips PW 1407 spectrometer. The framework silica to alumina ratio (SAR) was determined by FTIR in the structure vibration region by means of the frequency shitt of a characteristic band around 800 crn~ with a Nicolet 60SXR spectrometer. The textural characteristics of the zeolites, such as BET specific area, micropore volume (t-plot) and mesopore area (BJH), were evaluated by physisorption of Nz at 77K in a Micromeritics ASAP 2400 and their acid sites density by TPD of N H3 (adsorption performed at 423K with a 4.0% Nn3/ne mixture and TPD performed at 20K/min up to 823K). The reaction was carded out in a glass fixed bed gas phase reactor at 698K and 1 atm. The reaction products were analyzed by gas chromatography using an HP PONA column for the liquid phase and an Ai203/KCL PLOT column for the gas phase. For those runs aiming at coke characterization, WHSV was adjusted to allow a significant coke formation at isoconversion. The experimental method for recovering the coke from the samples and for determining the composition of the methylene chloride soluble fraction was reported elsewhere [7]. For TGA/DSC analysis the coked samples were treated at a 10K/min heating rate in a N2 stream up to 703K and kept at this temperature for 30 min. They were next burnt in a stream of N2 + 0 2 (10% 02) at the same heating rate up to 1023K in a Rigaku Thermoanalyser TS-1000. The TGA/DSC data allowed the calculation of the total coke content and the inspection of the TPO profiles gave information on coke reactivity.

3. RESULTS AND DISCUSSION Table 1 shows the physico-chemical characteristics of the samples. The results clearly indicate that all zeolites have the same global SAR and approximately the same sodium content. However, the SAR of the framework is much lower for the high rare earth content zeolite, showing that such a high content prevented dealumination. This fact is confirmed by the higher microporous volume of this sample as well as the lower mesoporous specific area. As expected, the samples containing rare earth elements have a lower acid site density. The results of the catalytic tests showed that ortho, meta and para-diisopropyl benzene, isopropylbenzene and benzene were the main products in the liquid effluents, propene being the main gaseous product. These results are in accordance with the reaction scheme proposed by Murta Valle et al [2]. Table 2 shows both total and insoluble coke contents after 100 min t.o.s, along with the respective extraction yields (Y%). The greater the rare earth cations content, the lower the coke build-up, depicting the negative effect of rare earth on coke formation. This fact may be attributed to a reduction in both acid site density and mesoporosity. A decrease in acid site density hinders bi and polimolecular reactions involved in coking while the absence of a

429 significant secondary mesoporous system limits the room available for voluminous coke precursor molecules. The latter characteristic resuks in a particularly significant reduction in insoluble coke content with decreasing mesoporosity. For all samples, coke is mainly CHaC12soluble and the extraction yields decrease with increasing coke contents, as previously reported [3,6,8-10].

Table 1 Physico-Chemical Characteristics of the Samples USY

CREY- 1

CREY-2

%RE203

0.0

2.83

12.3

%Na20

3.9

3.5

4.6

%SIO2

72.71

71.1

63.2

%A1203

23.37

22.6

19.7

SAR global

5.3

5.4

5.4

SAR framework (FTIR)

10.5

9.70

5.72

Total Acidity (Ixmol NH3/g)

650

420

310

V~c~o (cm3/g) (1)

0.262

0.273

0.275

S r~o (ma/g) (2)

33.9

27.1

20.2

S BET(m2/g)

623

649

645

(1) Micropore Volume ---}t-plot 3 to 5A ; (2) Mesopore Area ---}t-plot 6 to 9A

Table 2 Coke Contents (% C and % C ins) and Extraction Yields (Y %) Sample

%C

% C ins.

Y (%)

USY

15

4.5

70

CREY-1

14

2.8

80

CREY-2

10.6

1.6

85

The soluble coke components were classified in famih'es of general formula C~LI2r~z and Table 3 shows the most probable structural formula of the main components. Analysis of the CH2Cla-soluble fraction either by CG/MS or MS with direct introduction of the samples showed that their composition was not significantly affected by the different rare earth contents, since the same families were identified over the three samples. Famih'es 1 (CnH2n-12 alkylnaphtalenes), 3 (CnH2n-22 - alkylpyrenes) and 6 (CnH2n-32 - alkylindenepyrenes) were the main ones, followed by famih'es 4 (CnH2n-26 - alkylcyclopentapyrenes), 5 (CnH2n-28 indeneanthracenes / indenephenantrenes) and 8 (CnH2n-38).

430 Families 1, 2, 3 and 4 result from reactions involving alkylaromatic and olefmic compounds. So, the alkylation of alkylaromatics by olefins (propene or propene oligomers) followed by cyclization and then aromatization through hydrogen transfer reactions may explain the formation of those families [9,10]. Family 6 results from the nucleophilic attack of a benzilic carbocation by a pyrene molecule followed by dehydrogenation and family 7 by alkylation of family 6 with olefins and then by cyclization and aromatization [9]. On the other hand, families 2, 5 and 8 were formed from bimolecular reactions involving benzilic-type earbocations and diarylmethane intermediates [3,8]. Families 3, 4, 6 and 7 were also identified by Guisnet and co-workers [7-11 ] in the coke formed from toluene, n-heptane, propene and cyclohexene on USY zeolites at 723K. The coke molecules, whose dimensions [8,9] are close to that of zeolite cavities, remain trapped in the porous system due to their low volatility and/or by steric blockage inside the porous structure. Temperature programmed oxidation of coke deposits obtained after 100min t.o.s, are shown in Figure 1. It can be seen that samples USY and CREY-1 presented similar TPO/DSC 566 profiles. The oxidation of coke was somewhat easier for sample CREY-2 and this behavior can be tentatively associated to either the reduced coke content (particularly insoluble coke) of this sample and/or to its higher rare CREY-2 :3 earth content.

EY-1

"

400

~

'

500

'

600

'

700

USY

'

800

Temperature (K)

Figure

1. T P O

profiles for coked samples

Moljord et al [12] have shown, for protonic Y zeolites, that density of acid sites is the most important factor in determining the rate of coke oxidation and that the larger the number of A1 atoms or protonic acid sites per unit cell, the easier the coke combustion. This was not observed in the present work, since coke formed on USY (high acid sites density) was more difficult to burn than that formed on CREY-2 (low acid sites density). Although the role of rare earth cations in coke combustion has yet to be further explored, this observation suggests that these cations present a catalytic effect on promoting coke oxidation.

431

Table 3 Main Components of the CHzCIz - Soluble Coke Fraction Family

Molecular Formula

Structural Formula

[~~

Cx 2~x~6

CnH2n-12

-Cx

[~~

~

M.W.: 156- 212

CnH2n_18

Cx or

x=l

M.W.: 192

X CnH2n_22 O~x~4

M.W. 202-258

CnH2n-26 2_x_4

~

Cx

M.W.: 245-282

~

O~x~4

X

M.W.: 266-322

( ~ ~ C x

CnH2n-32

or 0~x~5

M.W.: 276-346

-cx

CnH2n-36

0<x<5

~

~Cx

M.W.: 300-370

Cx

or 0<x <6

CnH2n-28

M.W.: 354-424

C.H2.-3s

432 4. CONCLUSION The results of chemical analysis (GC-MS and MS) indicated that, regardless of the zeolite composition, the same families of components were present in the CH2Cl2-soluble fraction of the coke formed over Y zeolites with different rare earth content. Alkylnaphtalenes, alkylpyrenes and alkylindenepyrenes have been identified as its main components. Reactions of alkylation of alkylaromatics with propene, followed by cyclization and aromatization by hydrogen transfer are responsible for the formation of such compounds. From the evaluation of TPO/DSC profiles, a catalytic effect of rare earth elements on promoting coke oxidation could be suggested.

ACKNOWLEDGMENTS

The authors thank Ricardo da Silva Ademe (NUCAT/COPPE/UFRJ) for the analysis by TGA/DSC and Jos~ Orlando de Jesus Santos and Carla Maria Salerno Polato for their aid. C. A. Henriques expresses her gratitude to Fabrica Carioca de Catalisadores S.A for its financial support.

REFERENCES

1. E. Falabella S. Aguiar, M.L. Murta Valle, M. P. Silva and D. F. Silva, Zeolites, 15, 620, (1995). 2. M.L. Murta Valle, E. Falabella S. Aguiar, C. P. Costa, F.B.S. Castro and M. E. Espinho, Actas XV Simp. Iberoan~ Catal., V-I, p. 305, September, 1996, Cordoba, Argentina. 3. C. A. Henriques, J.L.F. Monteiro, A. M. P. Bentes Jr, P. Magnoux and M. Guisnet; to be published. 4. C. A. Henriques, J. C. Afonso, P. Magnoux, M. Guisnet and J. L F. Monteiro, 1 lth Int. Zeol. Conf., Book of Abstracts, pp 140, August, 1996, Seoul, Korea. 5. C. A. Henriques, J. C. Afonso, J. L. F. Monteiro; Actas XV Simp. Iberoam. Catal., V-II, p. 1213, September, 1996, Cordoba, Argentina. 6. C. A. Henriques, D.Sc. Thesis, PEQ/COPPE/UFRJ; Rio de Janeiro, Brazil, 1994. 7. P. Magnoux, P. Roger, C. Cannaff, V. Fouche, N. S. Gnep and M. Guisnet, Studies in Surface Science and Catalysis, 34, (1987) 317. 8. M. Guisnet and P. Magnoux, Appl. Catal., 54 (1989) 1. 9. P. Magnoux, F. Machado and M. Guisnet, in: L. Guczi et al (Editors), New Frontiers in Catalysis, Proc. 10th International Congress on Catalysis, Budapest, Part A, (1993), p.435. 10. P. Magnoux, C. Canaff, F. Machado and M. Guisnet, J. Catal.,134, (1992), 286. 11. K. Moljord, P. Magnoux and M. Guisnet, Appl. Catal. A, 122 (1995) 21. 12. K. Moljord, P. Magnoux and M. Guisnet, Appl. Catal. A, 121 (1995) 245. 13. J. M. Ferreira, M. Sc. Thesis, COPPE/UFRJ, Rio de Janeiro, Brasil, 1991.

9

Elsevier Science B.V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

433

Differential Effect of Coke Burning with O x y g e n or Ozone on Pt-Re Interaction in Pt-Re/A1203

C. L. Pieck and J. M. Parera Instituto de Investigaciones en Cathlisis y Petroquimica - INCAPE- (FIQ-UNL, CONICET) Santiago del Estero 2654, 3000 Santa Fe (Argentina)

Pt-Re/A1203 catalysts were prepared by different methods (coimpregnation, successive impregnations and catalytic reduction). The degree of Pt-Re interaction on the catalysts was indirectly determined by the hydrogenolysis of cyclopentane. The deposition of coke was performed at 500~ and atmospheric pressure. The modifications in the metallic phase when eliminating coke by combustion with oxygen at high temperature (450~ and with ozone at low temperature (125~ were studied. The regeneration with ozone enabled the elimination of coke without modifying the metal phase of the catalyst, keeping the degree of interaction between Pt and Re and maintaining Pt dispersion nearly constant. The regeneration with oxygen produced the elimination of Pt-Re interaction and an increase in the dispersion of Pt.

I. INTRODUCTION The deactivation of naphtha reforming catalysts is mainly due to coke deposition. Bimetallic catalysts have been developed to minimize the loss of activity, Pt-Re/AI203 being the most widely used formulation. In previous articles (1-3) we reported that for test reactions carded out at atmospheric pressure, the coke deposition was lower at higher degrees of Pt-Re interaction. Catalysts prepared by catalytic reduction and activated by direct reduction exhibited a higher degree of Pt-Re interaction than those prepared using the techniques of coimpregnation or successive impregnation. Pt-Re interaction is destroyed by calcination at 450~ If calcination is carried out at temperatures lower than 300~ the amount of hydroxyl groups remaining on the alumina surface is enough to let the migration of Re species proceed. The metallic phase interaction can therefore be reconstructed during Pt and Re correduction (4-7). For this reason, Pt-Re interaction could be maintained by carrying out catalyst regeneration at lower temperatures. Moreover, segregation of Pt and Re oxides should be lower due to an incomplete oxidation of metallic particles at such temperatures. Current regeneration procedures used in reforming plants burn coke at temperatures higher than 500~

434 so Pt-Re interaction disappears completely. Coke can be eliminated at lower temperatures using oxidants like ozone or nitric oxide that generate the much more reactive atomic oxygen. In this work the modifications that occurred in the metallic phase, namely in the Pt-Re interaction, during coke elimination with oxygen or ozone were studied using catalysts prepared by classical (eoimpregnation and successive impregnation) and recently developed (surface reaction or catalytic reduction) methods.

2. EXPERIMENTAL

Catalysts: All catalysts prepared contained 0.3% Pt, 0.6% Re and 1.0% CI. The precursor metallic salts used were H2PtCI6 and NH4ReO4. A ),-Al203 (180 m2/g) was used as support. Monometallic catalysts with the same composition were prepared as reference materials. Bimetallic catalysts were prepared according to the following methods: Coimpregnation: The support was previously treated with a 0.20M HCI aqueous solution in a beaker. Then the solutions of metals salts were added and the beaker content was slowly heated with simultaneous stirring in a sand bath until a dry powder was obtained. This material was oven-dried overnight at 120~ and finally activated by calcination (4h in air at 450~ and reduction (8h in hydrogen at 500~ For some preparations the calcination step was omitted. Successive impregnations: A monometallic Pt catalyst was first prepared and activated by calcination (4h in air at 450~ and reduction (8h in hydrogen at 500~ Then the solution of Re salt was added. The impregnated sample was first dried in a sand bath at 70~ and then in an oven at 120~ overnight. The catalysts prepared in this way were activated by the mentioned procedure of calcination/reduction or by direct reduction (8h in hydrogen at 500~ Catalytic reduction: A prereduced sample of Pt/Al203 monometallic catalyst was covered with a 0.20 M HCI aqueous solution and the solution was purged with a nitrogen stream to remove dissolved oxygen. Then the solution was reduced by hydrogen bubbling for 2 h and the aqueous solution containing the Re salt (previously purged with nitrogen) was added, letting the system react for 1 h with a continuous hydrogen flow. The solid was washed with distilled water and filtered. After oven-drying at 120~ overnight, the catalyst was activated using any of the procedures described above, namely direct reduction or calcination and reduction. Test reactions: Catalyst characterization was performed using test reactions of cyclopentane (CP) hydrogenolysis at 290~ cyclohexane (CH) dehydrogenation at 270~ and coking with CP at 500~ All catalysts were reduced with hydrogen at 500~ before the test reactions runs. Catalyst regeneration: Catalysts coked with CP at 500~ were regenerated with oxygen (2.5% in N2) at 450~ or with ozone (1% in air) at 125~ for 10 h. Ozone was produced in an equipment similar to that described by My and Sahghal (8). Temperature-programmed reduction : TPR diagrams were recorded in an Ohkura TP2002S. All samples were previously oxidized with oxygen (2% in argon) at 280~ At lower oxidation temperatures hydrogen uptake signals were low due to an incomplete oxidation of the metallic particles. On the other hand, higher oxidation temperatures would favor the segregation of metal oxides. The reducing gas was hydrogen (4.8 % in argon) and the temperature ramp was 10~

435 3. RESULTS AND DISCUSSION Figure 1 shows CP conversion values for fresh and regenerated samples of catalysts prepared by different procedures. It can be observed that catalysts activated by reduction have higher hydrogenolysis activity than those activated by calcination-reduction. According to Augustine and Sachtler (6), Pt-Re ensembles have higher hydrogenolysis activity than Pt or Re alone so CP hydrogenolysis can be considered as an indirect measurement of Pt-Re interaction. It can be concluded from Figure 1 that Pt-Re interaction in fresh catalysts is always higher for those directly reduced (R) than for the catalysts activated by calcination-reduction (C-R). For the only reduced (R) catalysts the degree of Pt-Re interaction increases according to the sequence: Coimpregnation<Successive impregnation
25

20

I--'lFresh 1~21Coke bumoffwith 03

15

/ C o k e bumoffwith 02

0 A, CR

B, CR A,R

C, CR B,R

Pt/AI203, CR C,R

Figure 1: Activity for cyclopentane hydrogenolysis. Catalysts: A: catalytic reduction; B: coimpregnation; C: successive impregnation; CR: calcined and reduced; R: only reduced catalyst. The destruction of the Pt-Re interaction by calcination is attributed to the formation of Pt and Re oxychlorided species strongly bounded to A1203. Those species cannot migrate during reduction so the correduction of Pt and Re is impossible (7). It can also be observed in Figure 1 that the catalysts regenerated with ozone have activities for cyclopentane hydrogenolysis (Pt-Re interaction) comparable to those of flesh catalysts. On the other hand, the catalysts regenerated with oxygen suffer a greater drop in the hydrogenolysis activity. Differences in activity between flesh and regenerated samples are small for catalysts activated by calcination-reduction. This can be explained taking into account that Pt-Re interaction was already partially destroyed during calcination. The degree of Pt-Re interaction in catalysts prepared and activated by different procedures was also measured using TPR analysis. As it can be seen in Figure 2, Pt oxide is reduced at lower temperatures than Re oxide when catalysts are monometallic, which agrees

436 800

1000

Catalytic Reduction (R)

-4 ., 8 0 0

:= 600

(6

...,.

Q.

E

o0~

600

e~

E

tO O rG) 4 0 0 01 0 t_

O 200 "O

"0

r9 2 0 0

0

400

rO O r

PUAI20

100

200

300

400

"r"

3

500

600

700

Temperature, ~

Figure 2: TPR of monometallic Catalyst

1~o ' 2~o ' 3;o' ,;o' 5~o' e~o '~o Temperature, ~

Figure 3: TPR of bimetallic catalyst

with current literature (7). Figure 3 shows the TPR diagram for Pt-Re/Al203 catalysts prepared by catalytic reduction (activated by reduction only) and coimpregnation (activated by calcinationreduction). The catalyst prepared by catalytic reduction exhibits a broad reduction peak centered about 330~ which is attributed to the correduction of Pt and Re oxides due to the catalytic effect of Pt (5,9). The catalyst prepared by eoimpregnation has a sharp peak at 130~ due to the reduction of segregated Pt oxide which is higher than the observed for the catalyst prepared by catalytic reduction, and a second broad peak around 350~ which includes the reduction of Pt oxide (starting a little below 200~ and the correduction of Pt and Re oxides at higher temperatures. A third broad peak can also be observed around 620~ which corresponds to the second reduction peak of Re/AI203 shown in Figure 2, due to segregated Re oxide. This peak is more important than the observed for the catalyst obtained by catalytic reduction due to a higher Re oxide segregation. Figure 4 shows conversion values for cyclohexane dehydrogenation corresponding to fresh and regenerated with oxygen or ozone catalysts. Cyclohexane dehydrogenation activity is directly proportional to the amount of surface Pt atoms (10), Re being almost inactive for this reaction. It can be seen that fresh catalysts activated by calcination and reduction (CR) have higher activity than those activated by reduction (R) independently of their preparation method. These results could indicate that the destruction of Pt-Re interaction during calcination increases the amount of surface Pt atoms. Comparing activity values for fresh and ozone regenerated catalysts shown in Figure 4, it can be seen that the regeneration with ozone does not significantly modify the metallic phase in agreement with data from Figure 1. Conversely, when catalysts are regenerated with oxygen at 450~ the metallic phase suffers a considerable modification: the degree of Pt-Re interaction diminishes (Figure 1) and the amount of exposed Pt atoms increases (Figure 4).

437

30

-

20

-

l0

-

o r~

O o

x

N rj

i

7 o

i A, CR

B, C R A,R

Figure

4:

C, C R B,R

Pt/A 1203, CR C,R

Activity for cyclohexane dehydrogenation. References as Figure 1.

Relative values of the increase in dehydrogenating activity after regeneration with oxygen are higher for only reduced catalysts. In catalysts calcined and reduced there exists a previous partial destruction of Pt-Re interaction by calcination, so, they are less affected by regeneration with oxygen. Values of total metallic dispersion [Pt+Re] (not shown) and dehydrogenating activity for different catalysts exhibit the same trend. The monometallic Pt catalyst has high dispersion and suffers sintering during regeneration, this effect being less important when regeneration is accomplished with ozone at lower temperatures. 1200 1200 1000

~: Signal/10 03 .

..I/\[':.,,03 H 2

lOOO c:" o

,Numinatreated / with 0 3 /

:r~ 800

E

600

'-

~

400

~"

,Numina treated with 0 2 ,

i

,

i

,

i

,

i

,

i

,

1O0 200 300 400 500 600 700 Temperature,

oC

Figure 5: TPR of alumina treated with ozone and oxygen.

i :

i

"

i

"

400

"1- 2130

1200

600

o

0 {~

800

0

-

i

: \

i

;

:J

i~

~

02

......... "*"-:'.,../;," ............. "........

;

1002130300400500600700 Temperature,

~

Figure 6: TPR of catalyst prepared by catalytic reduction regenerated with ozone and oxygen.

438 It is impossible to measure with TPR analysis the interaction degree of Pt and Re after regeneration with ozone, due to the formation of surface oxidized compounds that have a high hydrogen uptaking capacity. In Figure 5 TPR diagrams are shown for alumina treated with oxygen at 450~ and with ozone at 125~ The former does not exhibit any appreciable hydrogen uptake, but the sample treated with ozone has two reduction peaks at 230~ and 440~ Figure 6 shows the TPR diagrams of a catalyst prepared by catalytic reduction and activated by direct reduction after its regeneration with oxygen or ozone. It also includes a diagram obtained for an ozone regenerated catalyst a~er treatment with hydrogen (2h, 300~ The catalyst regenerated with ozone exhibits only one reduction peak corresponding to the first peak of alumina alone. It may be supposed that Pt has a catalytic action making it possible the reduction of the surface oxidized compounds at low temperatures. It can also be observed that the hydrogen uptake corresponding to the reduction of oxidized compounds in the metallic phase for the catalyst regenerated with oxygen ' is much lower than the corresponding to oxidized compounds in the metal and support for the catalyst regenerated with ozone. In order to eliminate oxidized surface compounds generated during ozone treatments, the samples regenerated with ozone were heated at 300~ in flowing hydrogen. Such compounds were reduced as expected, but the TPR diagram recorded a~er these treatments were different from the previous ones. This behavior could be possibly attributed to the exothermicity of oxidized compounds reduction.

4. CONCLUSIONS It can be concluded that the interaction Pt-Re depends on the catalyst preparation method, the catalytic reduction method being the one that leads to a better interaction degree. The calcination step in the preparation procedure is harmful for the interaction. Regarding coke burn-off, ozone at 125~ eliminates coke producing a low alteration in the metallic phase. When the elimination is carried out by oxygen at 450~ a greater modification in the metallic phase is produced: the Pt-Re interaction decreases and the fraction of exposed Pt is increased.

REFERENCES 1. C.L. Pieck, Thesis Doctoral, Poitiers University, France 2. C.L. Pieck, P. Marecot, J. M. Parera and J. Barbier, Actas XIV Simp. Iberoam. Catalysis, Concepci6n, Chile, 1994, p. 263 3. C.L. Pieck, P. Marecot, C. A. Querini, J. M. Parera and J. Barbier, Appl. Catal. A. General 133 (1995)281 4. R. Mieville, J. Catal. 87(1984)437 5. B.H. Isaac and E. E. Petersen, J. Catal. 85(1984)8 6. S.M. Augustine and W.H. Sachtler, J. Catal. 116(1989) 184 7. P. Malet, G. Munuera and A. Caballero, J. Catal. 115(1989)567 8. N.L. My and P. N. Sahghal, Chem. Eng. Journal 40(1989)15 9. L. Chen, Y. Li, J. Zang, X. Luo and S. Cheng, J. Catal. 145(1994)132 10. V. Krishnasamy and K. Balasubramanian, J. Catal. 90(1984)351

~ Elsevier Science B.V. All rights reserved Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

439

H y d r o d e s u l f u r i z a t i o n of Dibenzothiophene on a Nitrided Supported M o l y b d e n a A l u m i n a Catalyst Masatoshi Nagai*, Yosuke Goto, Hiroshi Sasuga, and Shinzo Omi Department of Advanced Materials, Graduate School of Bio-applications and Systems Engineering, Tokyo University of Agriculture and Technology, 2-24 Nakamachi, Koganei, Tokyo 184, Japan

The activity and selectivity of 12.5% Mo/A1203 nitrided at various temperatures for the hydrodesulfurization (HDS) of dibenzothiophene and the effect of re-treatment of NH3 on dibenzothiophene HDS were studied. The nitrided catalyst was significantly more active toward the scission of the C-S bond from dibenzothiophene with hydrogenation of dibenzothiophene. The sulfur species accumulated on the surface of the nitrided Mo/A1203 catalysts by replacement of nitrogen species after reaction was analyzed by XPS measurement. The formation of molybdenum sulfide during the HDS dibenzothiophene led to a decrease in the activity of the nitrided catalyst, which approached that of the sulfided catalyst.

1. INTRODUCTION Interest has increased in exploring the catalytic properties of molybdenum nitrides. Recently, Markel and Van Zee [ 1] studied HDS of thiophene on an unsupported Mo2N catalyst and reported that the Mo nitride had higher activity for hydrogenation than that for the C-S bond scission of thiophene. For the HDS of dibenzothiophene, however, the nitrided Mo/A1203 catalyst was more active than the catalyst sulfided at 623 K and was extremely selective for the C-S bond breakage of dibenzothiophene to produce biphenyl [2]. Thus, the selectivity of nitrided catalysts for the C-S hydrogenolysis to hydrogenation was not yet understood. Furthermore, although the nitrided Mo/A1203 catalysts was reported to be very active for HDS, the sulfiding process of the nitrided catalyst during the HDS reaction and the nitriding effects of the catalyst for the catalyst activity were not fully understood. Therefore, in this study, the effects of re-treatment with NH3 on the activity and selectivity of the Mo/A1203 catalyst for dibenzothiophene HDS were determined. The sulfur on the nitrided catalysts after reaction was analyzed by XPS with the purpose of determining the cause of the deactivation of the nitrided catalysts.

2. EXPERIMENTAL A 12.5% MoO3/A1203 catalyst was prepared by using a mixture of-alumina and ammonium paramolybdate, and calcined in air at 823 K for 3 h (Nikki Chemicals Co.). Nitriding of the

440 12.5% M o O 3 / A 1 2 0 3 was carried out by temperature-programmed reaction with NH3 [3,4]. The 12.5% M o O a / A 1 2 0 3 precursor was treated in flowing N H 3 at 4 liters/h at a rate of 1 K/min from 573 K to 773 K and held at 773 K for 3 h (Low-Temperature-Nitriding; LTN catalyst), from 573 K to 973 K and held at 973 K for 3 h (MTN catalyst), and from 573 K to 1173 K and held at 1173 K for 3 h (HTN catalyst). The flesh Mo/AI203 was presulfided in flowing 10% H2S/H2 (15 liters/h) at 573 K (LTS catalyst) and 773 K (HTS catalyst) for 3 h for comparison. Surface compositions of the nitrided catalysts were measured using a Shimadzu ESCA 850 spectrometer with monochromatic M g ~ exciting radiation (8 kV, 30 mA). For the activity measurement, the nitrided and granular catalysts (2.0 g, 0.84-1.19 mm) were passivated in 1% O2/He stream after nitridation. The HDS of dibenzothiophene was carried out using a stainless steel microreactor in a high pressure flow system. The liquid feed, consisting of 0.25wt% dibenzothiophene in xylene, was introduced into the reactor at 20 ml/h with a H2 flow of 6 1/h. The HDS rate for dibenzothiophene HDS on the Mo/A1203 catalysts was calculated based on the rate for disappearance of dibenzothiophene.

3. RESULTS AND DISCUSSION 3.1. HDS of dibenzothiophene The activities of nitrided catalysts for dibenzothiophene HDS versus time are shown in Fig. 1. The LTN catalyst was 1.5 and 1.6 times more active than MTN and HTN 1.0 -za ,

I

'

I

'

I

'

I

'

I

0.9 0.8 0.7

9~

o r,.)

.

~----O_

O-----dD-O--

. ~

0.5 0.4

-

~i'~

0.3

i -

9~I_-_

"

0.2

0. I 0.0 0

'

I

2

,

I

4

,

I

6

,

I

8

,

I

10

Time on stream [h] Figure 1. The HDS of dibenzothiophene on the ( O ) LTN, (O)MTN, (A)HTN, (II)HTS, and ([3)LTS catalysts as a function of time on stream at 573 K at 10.1 MPa total pressure. catalysts, respectively. In comparison the HTS catalyst was substantially less active for HDS. The HDS activity of the Mo/AIzO3 catalysts decreased with increasing nitriding temperature.

441 The activity of the LTN catalyst decreased with the time on stream and reached steady state in 3h. The major product was biphenyl with a small amount of cyclohexylbenzene in dibenzothiophene HDS on the nitrided catalysts. In the scheme of the HDS of dibenzothiophene on the sulfided [2] and nitrided [5] Mo/A1203 catalysts, sulfur-removal reactions as well as hydrogenation are present, as shown in Fig. 2. The reaction takes place through either the direct C-S hydrogenolysis of dibenzothiophene to biphenyl before hydrogenation or via hydrogenation of dibenzothiophene followed by sulfur-removal. In addition, a previous study [5] showed that biphenyl is not hydrogenated to form cyclohexylbenzene in the presence of dibenzothiophene, since dibenzothiophene adsorbed stronger than biphenyl. The hydrogenation of dibenzothiophene is a rate-determining step in the reaction at 573 K. Therefore, the molar ratio represents the selectivity of the nitrided catalyst for the C-S bond breakage in dibenzothiophene HDS. The molar ratio of biphenyl to cyclohexylbenzene is shown in Fig. 3. The formation of biphenyl was greater for the nitrided catalysts than for the sulfided catalyst. As a result, the nitrided catalyst was significantly more active toward the C-S bond fission without hydrogenation in dibenzothiophene. 1,2,3,4,4a, ga-Hexahydrodibenzothiophene

3,4-Tetrahydro~enzothiophene

Cyclohexytbenzene

Dtbenzothiophene

Bipheny[

Figure 2. The mechanism of the HDS of dibenzothiophene

3.2. The effect of NH3 re-nitridation

The effect ofNH3 re-nitridation on dibenzothiophene HDS is shown in Fig. 4. The reaction was interrupted in 10 h after the reaction started. After the reaction feed was stopped and hydrogen gas was changed to ammonia gas, the catalyst was heated to 773 K in flowing NH3, and kept at 773 K for 3 h. After the NH3 treatment, the formation of biphenyl and the conversion of dibenzothiophene increased by 17% and 75%, respectively, but the formation of cyclohexylbenzene decreased by 24%, compared with the reactions before the NH3 retreatment. However, the HDS activity of the catalyst rapidly decreased in 2 h to less than the activity before the NH3 treatment. The initially high activity and selectivity of the nitrided catalyst in the HDS reaction gradually decreased and approached that of the sulfided catalysts, as shown in Figs. 1 and 3. The XPS analysis showed that the S/A1 ratio of the LTN catalyst without NH3 treatment was 2.68 but NH3 treatment lowered the S/A1 to 2.3. These results suggest that the re-nitriding of the catalyst removed sulfur from the catalyst surface and regenerated the active molybdenum species near only the surface of the catalyst.

442

NH3 fl.ow

0.15

""

i

I

20

& o.lo-

I;/

"r~

I I

i

0

0

O.O5 -

I I

0 0

2

4

6

8

10

f~L Ii 9

I

I

I

0

I

2

3

T i m e on Stream [h]

Time on Stream [hi

Figure 4. Effect of re-nitridation of the catalyst with N H 3 o n the HDS of dibenzothiophene at 573 K.

Figure 3. The selectivity of the (O)LTN, (O)MTN, (A)HTN, (11) HTS, and ([]) LTS catalysts for the C-S bond breakage of dibenzothiophene.

N

/

N

%

[]

[]

//

Mo

/

Mo

\

3.3. Catalyst

/

N

\ H0s

O~dibenzothiophene; A biphenyl; A, cyclohexylbenzene.

S

\\

~

\

S

/

\

\

/

Mo

/

I

// Mo

N

S

NH3 % ~

\

[]

[]

N

/

Mo

/

// Mo

+ Ntt3S

\

deactivation

The relationship between carbon deposition and the decrease in activity was determined through the ratio of XPS C ls/A1 2p peak area. In Table 1, the C/AI ratios of the LTN and HTN catalysts were greater than the others and that of the MTN catalysts was the smallest. This indicates that the drop in activity of the catalysts during reaction was not related to carbon deposition on the surface of the catalyst. To determine if the activity of the nitrided

443 catalyst was diminished by sulfur poisoning, the S 2p XPS spectra for the nitrided catalysts before and after the reaction were evaluated. The S 2p XPS binding energy for the nitrided catalysts aider the reaction was low, at 162.8 eV (S/A1 ratio, 0.0268), as shown in Fig. 5. For the sulfided catalyst, the S 2p XPS was high, at 162.5 eV (S/A1 ratio, 0.171). The 773 Knitrided catalyst after sulfiding at 573 K in 10% H2S/H2 also has a peak centered at 162.7 eV (S/A1 ratio, 0.0552) with a shoulder peak of 163.8 eV. Although the nitrided catalysts had low intensity of S 2p XPS binding energy compared with that of the sulfided catalysts, the peak for the catalysts at 162.5 eV was attributed to $2[6], whereas the a shoulder peak at 163.8 eV for the MTN catalyst could be due to sulfur, So (164.7 eV [6]; 163.9 eV [7]). In Table 1, however, the difference between the N/Al atomic ratio of the nitrided catalysts before and after the reaction linearly decreased as the S/Al ratio of the catalysts aider reaction increased. From the XPS data, therefore, the surface of the spent catalyst was covered with sulfur ions of MoS2 during the HDS reaction. The formation of molybdenum sulfide during the reaction led to a quick decrease in the activity and selectivity of the nitrided catalyst, which approached those of the sulfided catalyst. The decrease in activity and selectivity of the nitrided catalyst was attributed to the structural changes associated with sulfur deposition by replacement of nitrogen atoms and adsorbed nitrogen species with sulfur atoms. Moreover, alumina in the catalyst was not nitrided under these pretreatment conditions, since the A1 l s XPS spectra were indistinguishable for the nitrided Mo/AlzO3 catalyst before and aider the reaction. S 2p lrz-3rz

S2 162.8 !

(a)

.~

.'4

f

o,-,4

162. 5

,.

172

......... i...., ....

16/.

156

BindingEnergy (eV) Figure 5. The S 2p XPS spectra for (a) the 773 K-nitrided catalyst a~er the reaction, (b) the nitrided catalyst was sulfided at 573 K a~er nitriding at 573 K, and (c) the 573 K-sulfided catalyst before reaction.

444 Table 1 Surface area and surface compositions of the nitrided 12.5% MoO3/m1203 catalyst at reaction temperature of 573 K

Catalyst LTN

Surface Area a) (m2/g) 226

N/Mo C 1s/A12p Ratiob) 0.24 2.5

XPS a) SIAlc) x 10z 2.68 (1.80) ~ 1.99 1.09

N/A1d) x 102 3.58 (1.55) ~ 2.15 1.76

MTN 195 0.46 1.7 HTN 138 0.35 2.8 a) Evacuated at 473 K and at 1 Pa b) Nitrogen content of N was measured by oxygen burning method and the Mo loading was 12.5% in the sample measured by atomic absorption spectroscopy. The catalyst was evacuated at 973 K. c) XPS S2p: 164 eV, after reaction at 573 K. d) XPS N/AI before reaction minus XPS N/A1 after reaction e) reaction temperature at 533 K.

4. CONCLUSIONS (1) The LTN catalyst was 1.5 and 1.6 times more active than MTN and HTN catalysts, respectively. The HDS activity of nitrided Mo/A1203 catalysts decreased with increasing nitriding temperature. (2) Nitriding enhanced the removal of sulfur atom in dibenzothiophene to form biphenyl. The nitrided catalyst was significantly more active toward sulfur removal without hydrogenation of dibenzothiophene. (3) Although the nitrided catalyst exhibited high selectivity for the C-S bond scission of dibenzothiophene, the selectivity of the catalyst gradually decreased and approached that of the sulfided catalysts. (4) Re-nitriding enhanced the HDS activity of the nitrided catalyst during the reaction through interruption, but the recovered activity of the catalyst rapidly decreased becoming less than the activity before the NH3 treatment. (5) The nitrogen in Mo nitride was removed and replaced by sulfur.

REFERENCES

1. 2. 3. 4.

E.J.Markel, and J. Van Zee, J. Catal., 126 (1990) 643. M.Nagai, T. Miyao, and T. Tsuboi, Catal. Lett., 18 (1993) 9. L.Volpe, and M. Boudart, J. Phys. Chem., 90 (1986) 4874. L.Volpe, S.T. Oyama, and M. Boudart, Preparation of Catalysts III, G. Poncelet, P. Grange, E A. Jacobs (eds.), Elsevier, Amsterdam (1983), p. 147. 5. M.Nagai, T. Sato, and A. Aiba, J. Catal., 97 (1986) 52.

445 6. J.R.Brown, and M. Ternan, Ind. Eng. Chem. Prod. Res. Dev., 23 (1984) 557. 7. C.D.Wagner, W.M.Riggs, L.E.Davis, J.F.Moulder, G.E.Mullenberg (eds.), Handbook of Xray Photoelectron Spectroscopy, Perkin-Elmer Corp., (1979).

This Page Intentionally Left Blank

9 Science B.V. All rights reserved Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

447

Industrial E v a l u a t i o n o f Selective H y d r o g e n a t i o n Catalyst P o i s o n i n g B. Didillon a, j. Cosyns a, C. Cameron a, D. Uzio a, p. Sarrazin b, j.p. Boitiaux b

a IFP, 1&4 avenue de Bois-Pr6au, 92852 Rueil Malmaison cedex, France b Procatalyse, 212-216 avenue Paul Doumer, 92500 Rueil Malmaison, France This paper presents a survey of the principal selective hydrogenation catalyst contaminants and their sources. The particular special cases of Hg and As are discussed. More recently silicon poisoning has become important. This subject will also be addressed. For each contaminant type, the IFP solution to suppress the undesirable effects will be detailed on the basis of results obtained in industrial plants.

1. INTRODUCTION Selective hydrogenation catalysts are susceptible to contact with a wide variety of contaminants (S, Hg, As, Si, Na, Fe, etc.) during their lifetime. Their influence on catalytic performance depends on the type of reaction, the operating conditions, the active phase and the contaminant itself [ 1-3]. The effects range from a simple reversible inhibition to an irreversible poisoning and consequently the necessity to change the catalyst. In the first section, we present the industrial background of selective hydrogenation processes and the variety of contaminants that catalysts may encounter. The nature of the effects (reversible or not) on catalyst activity will be discussed on the basis of industrial experience. In the following sections, we focus on the different purification technologies developed by IFP for various types of contaminant. 2. THE INDUSTRIAL SITUATION Unsaturated compounds (acetylenes, diolefins and monoolefins)which are not present in natural petroleum and coal sources, are produced by conversion processes such as steam cracking, fluid catalytic cracking, coking or visbreaking. All these processes produce a mixture of mono and polyunsaturated hydrocarbons which cannot be directly utilized in petrochemical processes. Indeed, polyunsaturated compounds have a strong ability to form polymers (green oil for example in the case of the C2 cut) which have negative effects on catalysts and process units in terms of deactivation or pressure drop. Selective hydrogenation processes are often required to reach the specification required by refining and petrochemical industries. Moreover, the hydrogenation feedstocks may contain a wide variety of impurities, in organic or mineral form (see Table 1). The utilization of ever more diversified feedstocks containing metal impurities can lead to an increase in operating problems. Catalytic processes such as selective hydrogenation, employed downstream of the steam cracker separation train, may suffer from reduced cycle lengths and lifetimes due to poisoning.

448

Table 1. Some typical hydrogenation feedstock contaminants from ref [21. Contaminant Normal s o u r c e Concentrationrange Mercury(Hg) Naphtha 1-2000 ppb steam cracker C3 Natural gas condensate Arsenic FCC off-gas 1-150 ppb Arsine (ASH3) Refinerygas 1-4 ppb steam cracker C3 60 ppb Natural gas condensate 65 ppb Carbonyl sulfide (COS) FCC off gas < 5 ppm ethane 0.1 - 1000 ppm LPG FCC C3 Hydrogen sulfide Refinerygas trace to % (H2S) Naphtha Gas oils FCC off gas LPG C2 and C3 mixtures 0.1 - 0.3 % vol Carbon monoxide steam cracker H2 gas FCC offO.5 - 1% vol gas C3 mixtures

Oxygen

Refinerygas FCC C3, C4, C5

0 - 50 ppm

Refinerygas

2 - 1000 ppm

Consequences AI equipment corrosion, safety, Catalyst poisoning Catalyst poisoning Safety

Environment Catalyst poisoning Corrosion Product specification Environment Catalyst poison Corrosion Productspecification

Front End C2 reactor stability Product specification PSA or methanator stability Environment Catalyst poison Corrosion Product specification Environment Catalyst poison Equipment fouling NOx and ~um formation

3. D E S C R I P T I O N O F T H E P O I S O N I N G P H E N O M E N A In addition to "normal" deactivation due to operating severity or to the unstable characteristics of the processed cuts, impurities brought by feedstocks are sometimes so poisonous that the catalytic transformation is no longer possible. It is very important to distinguish between "inhibitors" and "poisons". When inhibitors are present, catalytic activity is reduced to a lower level of performance but the initial activity can be recovered when the inhibitor disappears from the feedstock (Figure 1). In the case of poison, the activity decrease may be so important that the activity is nil when the catalyst has accumulated a given quantity of poison. Moreover when the quality of the feedstock is restored, the activity of the catalyst is not recovered. Among these poisons, one can distinguish further between reversible poisons and lethal poisons. For reversible poisons, it is possible to recover at least part and sometimes all of the catalytic activity by carrying

449 out oxidative catalytic regeneration. However, with lethal poisons, no solution is available except changing the catalyst.

Activity NO IMPURITY

IMPURITY

NO IMPURITY INHIBITOR

POISON

r

Time

Figure 1. Effect of inhibitors and poisons on catalytic activity

4. IFP EVALUATION OF CATALYST POISONING During our most recent 20 years of experience in selective hydrogenation, IFP has analyzed thousands of used catalyst samples covering all types of problem directly linked with catalyst poisoning. This technical assistance activity has led to an important database providing reliable correlations with catalyst activity and selectivity and to new lab scaled tools and methods used for routine catalyst evaluation procedures. This know-how is particularly effective for delivering rapid and reliable diagnostics on catalyst properties (including residual activity and selectivity). As shown in Figure 2, all the IFP selective hydrogenation process catalysts are represented in this database with a marked predominance for gasoline hydrogenation catalysts. Various contaminants have been reported for selective hydrogenation catalysts. Figure 3 indicates the type of poisoning encountered on Pd-alumina catalysts. These results correspond to more than 3000 analyses performed in IFP's R&D laboratories. Among these elements, Hg and As are the more frequently analyzed, but more recently, poisoning with Na, Si and Fe have increased significantly. The origin of sodium is generally attributed to presence of sea water in the feedstocks or related to a caustic wash step to remove sulfur compounds (HES or mercaptans), and may reach high levels (thousands of ppm). Silicon poisoning may be related to the use of additives (anticoking, antifoaming or anticorrosion) in processes upstream of selective hydrogenation units. A classic example of the application of IFP expertise performed on catalyst samples taken from industrial units is illustrated in the case of silicon poisoning. Two types of Si contamination have been observed: - the first class involves Si species having weak interaction with the active metal phase. This generally corresponds to large crystals of SiO2 retained in the pores of the support and comes from inert balls present above and below the catalyst bed in industrial reactors. This type of Si contaminant has little effect on the activity of the catalyst. Thus catalytic activity remains stable even with very high levels of SiO2,

450

Frequency %

4O 35

30

25

20

15

10

5

-

0

-

.

.

.

.

.

.

.

.

S

~;

Figure 2. Frequency of IFP analysis versus the type of catalyst.

18

Frequency

(%)

16 14 12 10

-r

<

z

ft..

~:

N

o

Figure 3. Type of comaminant encountered on a Pd based selective hydrogenation catalyst.

451

- the second class is represented by organosilicon species which are adsorbed and decomposed (hydrogenolysed) on the metal surface. These molecules are much more poisonous than the previous class and have a detrimental effect on activity, which cannot be recovered by oxidative regeneration) Plotting the results of analyses performed on used Pd-A1203 catalysts, a correlation has been found between residual hydrogenation activity and the amount of Si present on the sample (see Figure 4). This correlation facilitates estimating the remaining life of a catalyst when its level of Si poisoning is known. The deactivation curve obtained shows two distinct zones: a first zone exhibiting a large change in slope and a second zone where the slope change is limited (almost horizontal). These two regions depend very much on the specific metallic structure of the Pd catalyst. With the LD265 catalyst, the remaining activity level (in the low slope region) is 30 %. After the regeneration procedure, the catalyst does not recover its initial activity, most likely due to the formation of an SiO2 overlayer and/or reconstruction of Pd particles with Si. This reoxidation step probably does favor the migration of Si to the support to form SiO2 and the partial access of the Pd surface area, as recently reported in literature [4].

1 00

Remaning tivity 9 (%)

80

LD265 (PROCATALYSE)

60 40

Ibm

Ii

i

20 Si concentration (ppm)

Figure 4. Influence of Si concentration on the remaining activity for the LD265 catalyst. When dealing with uncontrolled imported feedstocks, one way to partially overcome the effects of Si poisoning is to increase the remaining activity. While the best solution is to avoid Si poisoning, consideration of realistic operating constraints for the petrochemical industry has led us to fous on improving activity, stability and resistance of Pd catalysts. Iron is also a frequently analyzed contaminant. Corrosion or colloidal iron species (dust) may be the cause of this type of poisoning. Pd based catalysts are sensitive to this contaminant because 1Industrial catalysts used for long periods in an industrial plant are in general best regenerated in an oxidative atmosphere under controlled conditions representative to those found in industrial units before laboratory characterization.

452 of their ability to form alloys. The characteristics (particle size, geometric and electronic structure) of the active phase are modified to a large extent, and finally activity and selectivity are degraded. It is worth noting that this type of poisoning is irreversible. Nevertheless, iron contamination may have a much less damaging effect, as in the case of Si poisoning, when large dust particles in feedstocks or originating from corrosion problems are trapped by the porous network of the catalyst support. The interactions with Pd are then weak and little deactivation is observed. The reddish iron oxide on the outside of catalyst spheres may be removed by a simple mechanical rubbing. Recent work in the literature mentions these two types of Pd-Fe association [5].

5. A R S E N I C A N D M E R C U R Y I F P P U R I F I C A T I O N T E C H N O L O G I E S

5.1. Arsenic poisoning Arsenic compounds are transformed in the FCC operation into relatively reactive arsine (ASH3), the boiling point of which falls within the C3 cut boiling range. As a consequence, the FCC C3 cut very often contains significant amounts of arsine, a poison for the selective hydrogenation catalysts. Indeed, the utilization of the propylene contained in this cut requires the selective hydrogenation of the remaining traces of MAPD (methylacetylene and propadiene). Severe deactivation can be observed but the original performance of Pd-based catalysts is readily recovered by calcination in air. This oxidation step breaks the Pd-As bonds and the As203 formed is trapped on the alumina support in a non-poisonous catalytic form. In order to prevent As contamination, a lead oxide on alumina adsorbent can be added to the fir: reactor which permits a good removal of ASH3. Figure 5 compares the cycle length of tl~ hydrogenation catalyst with and without lead oxide guard bed. A s = 150 p p b

As = 0 ppb

A s = 150 p p b

FCC C3

A s = < 10 p p b

Fee C3 STEAMCRACKER C3

STEAMCRACKE c3 Palladium Catalyst

Palladium Catalyst

C y c l e length relative

Cycle length

41-

4

20

100

Figure 5. IFP technology for As removal in FCC C 3 feedstock.

453 Arsenic found in non olefinic liquid hydrocarbons is present as organoarsenic species and thus must be converted to a more reactive form to be efficiently trapped. This can be achieved using the single step process. Feedstock treatment with a metal sulfide containing catalyst adsorbent, under the appropriate reaction conditions permits the total conversion and trapping of arsenic compounds as shown in the following mechanism:

/R1 R2 " " A s R3 R2

~As I Ni

.R1

H

\

Ni

H

\

R2 ""As ,R1 m I Ni

m Ni

Ni

~As I Ni

+

R3

sR1 +

R2

/H

/H

5.2. Mercury poisoning The removal of all mercury species from liquid hydrocarbon feedstocks (raw condensates, crude oil, condensate cuts) is difficult since the majority of the mercury in these feeds is in organometallic form. Most of the time the presence of mercury is associated with the injection of natural gas condensates into the steam cracker. These heavy condensates can contain very high levels of mercury compounds and these mercury species are spread over all the cuts according the following distribution: 0-5 % C2, 30-35 % C3, 55-65% C4, 5-10% others cuts. If a mercury polluted stream contacts a palladium catalyst, severe deactivation occurs. This phenomenon is partially reversible as a thermal, oxidative regeneration permits recovery of some of the catalytic activity. Moreover and in contrast to arsenic, mercury is removed from the solid during this thermal treatment. This means that the gas effluent is polluted by mercury. Another important damaging consequence of mercury is the corrosion of aluminum containing materials (heat exchangers). To avoid all of these effects, IFP proposes a two step process for arsenic and mercury removal, upstream of the steam cracker. The impure feedstock is mixed with hydrogen then treated under process conditions. The feed enters the first reactor where both hydrogenolysis reactions and arsenic capture occur using a metal sulfide based catalyst (CMG841 - Procatalyse). The arsenic free product containing metallic mercury is then cooled down and treated in the second reactor for mercury trapping on a different metal sulfide (CMG273 - Procatalyse).

6. CONCLUSIONS Deactivation of hydrogenation catalysts is a problem which is faced daily by industrial operators. It involves complex phenomena at the surface and in the bulk of the metallic active phase or the carrier. The increased utilization in terms of sophisticated high performance catalysts for refning and petrochemical applications coupled with the use of more diversified feedstocks requires a working knowledge of inhibitors and poisons which may be present in specific feedstocks and how they affect catalyst performance.

454 REFERENCES

1. J.P. Boitiaux, J. Cosyns and F. Vema, Studies in Surface Science and Catalysis, 34, 105 (1987). 2. D.R. McPhaul, "Ethylene Plant Feedstock Contaminants Treatment", AIChE Spring National Meeting (1995). 3. J.A. Reid, D.R. McPhaul, Hydrocarbon Processing, 45-53 July 1996. 4. G.V. Smith, F. Notheisz; S. Tjandra; M. Musoiu; T. Wiltowski; M. Bartok;, Journal of Catalysis (161), 441-450, 1996. 5. M.M. Bhasin, "Importance of Surface Science and Fundamental Studies in Heterogeneous Catalysis"/4 th North American Meeting, Snowbird - Utah, June 11-15, 1995.

9

Elsevier Science B.V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

455

The M e c h a n i s m o f Metal Poisoning by Cyclic Deactivation in Fluid Cracking Catalysts F. Hernfindez, R. Garcia de Le6n, E. M6gica, J.C. Moreno, R. Gonz~ilez, and E. Garciafigueroa Instituto Mexicano del Petrrleo, STI, Eje Central L~aro C~denas 252, PO Box 14-805, 07730 Mrxico D.F., Mrxico, e-mail" [email protected] A commercial REUSY fluid cracking catalyst was subjected to a cyclic deactivation (CD) procedure in a fluid cracking pilot plant using a metal-spiked VGO feed. The properties of the catalyst were compared to those of equilibrium catalysts obtained from an industrial FCC unit and a sample metallated by a CPS deactivation method. The catalysts were characterized by surface area measurements, scanning electron microscopy, electron dispersive spectrometry and microactivity test. Metals distributions on catalyst particles and aging rates were found to be the main differences between catalysts treated at the two different scales. Oxidation of metals has negative effects upon coking and hydrogen production relative to those produced in the reduced state. 1. INTRODUCTION Understanding catalyst deactivation, fouling and aging mechanisms is of paramount importance for developing better catalysts. This is particularly true for catalysts used in the fluid catalytic cracking (FCC) process. Metals present in the hydrocarbon feeds in commercial FCC units deposit continuously on the catalyst. Nickel and vanadium are the primary contaminant metals of a typical fluid cracking feedstock. Although a number of factors are actually involved in the catalyst deactivation it has been demonstrated that these two metals are very frequently the origin of activity and selectivity changes during operation. The main deleterious effects of nickel and vanadium poisoning are high coke and hydrogen yields and loss of liquid product. Extensive literature [1-4] related to this subject is available. The quality of feedstocks processed in FCC units continues to decline. Therefore more units will reach operating constraints due to larger amounts of coke precursors and metal contaminants. While special process and unit design modifications are developed to handle the processing of these heavier feedstocks highly metal tolerant catalysts are required. Laboratory cracking catalyst deactivation protocols are used in order to simulate long term commercial catalyst deactivation in an accelerated way. Different methods reported to date have been reviewed [5]. It has been observed [6-7, 10] that the oxidation state influences greatly the catalytic effect and aging rates of metals deposited on cracking catalysts. Therefore current laboratory procedures for studying metal deactivation in FCC catalysts include cyclic reduction and oxidation treatments. Such methods are based mostly on a pore volume metal impregnation based on the Mitchell method [8]. Although very simple to handle, these procedures tend to produce a rather homogenous distribution of metals throughout the catalyst particle which differs to what is actually observed in industrial equilibrium catalysts [5].

456 The aim of this work was to further understanding of how effect of metals relate to the procedure by which they are deposited in the catalyst. We focused our examination on a deactivation procedure based on the so-called cyclic deactivation method [5,9]. The cyclic deactivation method has proven to be one of the best approaches for simulating properties of a fluid cracking equilibrium catalyst. Basically it involves metal poisoning during cracking in a cyclic, integrated plant using a natural high-metal or metal-spiked feedstock. It is hoped this information will lead to better design methodologies for testing fluid cracking catalysts at the laboratory scale. 2. EXPERIMENTAL

2.1. Catalysts Some of the properties of catalysts used in this study are listed in Table 1. The base catalyst was a commercial REUSY catalyst currently used in an industrial unit operating in a partial combustion mode at 964 K with the same feedstock used in this study. This catalyst contains 1.5% Re203 and 35 % A1203 . The zeolite to matrix surface area ratio (fresh catalyst) is ca. 2.7 thus indicating a moderate matrix activity. Two samples of equilibrium catalyst coded ECAT1 and ECAT2 differing in their zeolite surface area and vanadium content were used as references. ECAT2 was separated into four fractions by means of density differences using organic solvents. Thus ECAT2F1 represents as much as 45 wt % of the most aged fraction of the equilibrium catalyst. A sample referred to as CPS was prepared by Grace-Davison. The modification of the base catalyst was done by incipient wetness impregnation with nickel and vanadium napthenates plus 30 cycles of reduction (propylene)-oxidation treatment. In order to probe the influence of the oxidation state of metals upon product selectivity samples were subjected to oxidation at 973 K or H2 reduction at 923 K treatments for 4 h prior to MAT tests. These samples were labeled O and R respectively.

2.2. Deactivation procedure Prior to metals deposition the fresh catalyst was deactivated at atmospheric pressure for 4 h at 1061 K and 100% steam in a fluidized bed. This sample was identified as HTD. Cyclic deactivation (CD) experiments were performed in a FCC pilot unit containing a 5.6 m folder riser, disengager, stripper and regenerator. 2300 g of catalyst were continuously circulated in the unit. The temperature of the riser was kept at 793 K. The regenerator operated isothermally at 943 K with an air flow of 367 1 h ~ thus assuring complete coke bum operation. The residence time of the catalyst in the regenerator was around 30 min. The feed was a Vacuum Gas Oil (VGO) of 23.4 ~ 11.9 KUOP, 0.5 l%wt Carbon RB, 2.0%wt S, 0.16 ppmw Ni, 1.2 ppmw V and 722 K for a 50 %vol distillate. Two reference runs were done with this feed using HTD and ECAT1 catalysts. The CD run was conducted using the HTD catalyst and the charge stock spiked with organometallics of light molecular weight provided by OM Group Inc. as Nickel and HEX-CEM CURRX to concentrations of 105 ppmw Ni and 460 ppmw V. The contaminated VGO was fed at a rate of 600 g h "1 under continuous stirring. The pilot plant was operated at a catalyst to oil (CTO) ratio of 6 and a WHSV of 2 1 h "1 thus enabling the catalyst to be circulated through the plant for approximately 50 cycles within a 24 h period in order to achieve target metal concentrations (Sample 24CD). This was followed by a second 24 h period for catalyst aging (sample 48CD). The Dry Gas and GLP product distribution were determined by means of a Hewlett-Packard Gas Refinery System GC analysis. The liquid portion was analyzed by a Hewlett-Packard GC operating under a simulated distillation method setting the gasoline cut to 494 K. Mass balance calculations (within 97 wt%) were performed each 3 h. Conversion was calculated on a weight basis from the difference

457 between the weight of VGO fed to the plant and the recovered fractions of Light Cyclic Oil (LCO) and Heavy Cyclic Oil (HCO). No fresh catalyst was added during these runs. The metal content in the catalyst increased linearly during 24 h operation. The catalyst obtained at the end of the run (samples 24CD and 48CD) contained Ni and V concentrations very close to the target values of 700 and 2900 ppmw. A V/Ni ratio of 4 is typically found in equilibrium catalysts in Mexico. The metal content produced on the catalyst by the VGO itself was less than 50 ppmw. Table 1. Catalyst properties Sample FRESH HTD HTD (10 h) ECAT1 ECAT2 ECAT2F 1 CPS 24CD 48CD

Surface area (m2/g) Zeolite Matrix 165 62 127 34 118 33 111 36 106 36 99 29 129 36 118 37 117 36

U.C. size (nm) 2.457 2.432 2.430 2.429 2.429 2.429 2.430 2.431 2.431

Metal (ppmw) Ni V . . . . . . . . . . . . 460 2183 503 2760 666 2804 1030 2170 680 2930 680 2930

2.3. Characterization and testing Surface areas of catalysts were determined by N2 adsorption using an ASAP 2000 analyzer from Micromeritics. Matrix and zeolite surface areas were calculated by the t-plot method accordingly to the ASTM-D-4365 standard test [11 ]. Zeolite unit cell size (UCS) was determined by X-Ray diffraction using a SIEMENS D-500 automated analyzer according to the ASTM-D3942-80 standard [ 11 ]. Catalyst samples were also characterized by Scanning Electron Microscopy (SEM) and Energy Dispersive Spectroscopy (EDS) using a JEOL 35CF electronic microscope equipped with a Kevex Microanalyst 7500, a Si/Li detector and Be windows. For approximate measurements of concentrations and distribution of metals in the catalysts V and Ni signals were referenced to the Si signal. Overall concentrations of metals in catalysts were determined in acid digested samples in a Perkin-Elmer 2380 Atomic Absorption (AA) spectrometer. Upon deactivation and metallation values of zeolite UCS and matrix surface area (Table 1) are similar to those for ECAT except for the most aged fraction ECAT2F1. It appears that zeolite surface area reduction is essentially due to steaming. Upon metallation it undergoes further reduction similar to that obtained after 10 h steaming. The microactivity test (MAT) based on the ASTM-D-3907 [1 l] standard was used to determine activity and product selectivity of catalysts. MAT runs were performed in a Xytel automated equipment with 4.0 g of catalyst using the same VGO as in the pilot plant runs. Unless otherwise specified catalyst samples were previously calcined at 853 K for three hours. Operating conditions were 793 K, CTO ratio of 4, 75 s injection time and WHSV of 15.7 h "1. Product analysis and conversion and selectivity calculations were done as in the pilot plant. The relative error of data was 5%. We analyzed coke burn products "in-situ" by IR analysis using an HORIBA VIA-510 analyzer. Product distribution was expressed in terms of product yield/ activity ratio as defined in Table 2 as currently used for interpreting MAT numbers in equilibrium catalysts.

458 3. RESULTS AND DISCUSSION Figure 1 presents the deactivation curves of catalysts obtained in the FCC pilot plant runs. Under constant operating conditions the pilot plant conversion decreases steadily. In the absence of metals, the deactivation rate (samples HTD and ECAT2) is lower i.e., a decrease of only 2-3 wt% is observed after 24 h. Deactivation is faster in the metal-spiked feed, i.e., conversion decreases 74 wt % to 66 wt% for sample CD within 24 h. During the second 24 h aging period (24CD to 48CD) conversion decreases only slightly (1.5 wt%). As metallation proceeds the product selectivity changes. The coke factor (defined in Table 2) in sample CD increases from 1.5 at 8 hours to 2.1 at 24 hours operation. The gas factor changes from 0.6 to 1.1 for the same period of time. By comparison, factors for ECAT and HTD (1.4 and 1.5) and gas factors (1.3 and 0.6) are almost constant. The gasoline selectivity varies from 67 to 75% for sample CD, 62 to 64% for sample HTD and remains almost constant 69-70% for sample ECAT2.

77

o'--9, 75t-

.o

e0

71 -

I,,-

r

69-

to

0

CD ( V G O + M E T A L ~

67-

65 0

4

8

t

I

12

16

I

20

24

28

Time (h) Figure 1. Conversion vs. time on stream in pilot plant. The decrease in conversion upon metal addition is concomitant with decreases in zeolite surface area (Table 1) and MAT activity (Table 2) for the catalysts. It should be noted that an additional 24 h aging period (24CD vs 48CD) had no further effect on these parameters. While samples 24CD and CPS have similar or lower MAT activities compared to ECAT samples the zeolite surface areas of the former are higher. Therefore it can be concluded that modifications in samples 24CD, 48CD and CPS involves catalyst matrix properties rather than zeolite properties. This is consistent with the observations of Pompe et al. [2]. Figures 2a to 2c show SEM images for samples ECAT2F1, 48CD and CPS. The appearances of particles of samples ECAT1, ECAT2 and HTD are very similar to ECAT2FI's. The wide age distributions and metal contents of catalyst particles in samples ECAT1 and ECAT2 prevented an accurate interpretation of the EDS analysis. In contrast, since sample ECAT2F1 consisted of particles of similar age and metal content the analysis by EDS was more accurate and representative. High concentrations of V on the surface of particles were observed while the Ni signal was too low for accurate analysis. It should be noted that the surfaces of some particles contained in samples 48CD (similarly to 24CD) appeared to be modified somewhat. In general,

459

(a)

(b)

(c)

Figure 2. SEM images for samples: (a) ECAT2F1; (b) 48CD; and (c) CPS.

460 analysis by EDS indicated higher concentrations of V and Ni on particle surfaces. Referenced to silica, average V and Ni concentrations were 10 and 50 fold higher than concentrations obtained by AA while analyses performed on these modified particles revealed 30 and 200 fold higher concentrations. Based on these results we interpret the appearance of such particles as due to Ni and V accumulation on, and extensive sintering of, the surface. Table 2 MAT (CTO=4 and 793 K) catalyst data Sample Activity a H2/CH4 Gas Coke GLP Selectivity c (%) factorb factorc olefinicitv d GLP e gasoline e bottoms t HTD 2.60 0.04 0.87 2.05 58.5 21.4 68.1 67.5 ECAT1 2.37 0.05 0.72 1.72 54.8 21.5 70.3 63.8 ECAT1R 2.41 0.08 0.58 1.38 64.8 19.8 73.5 64.3 ECAT2R 1.93 0.14 0.57 1.39 71.2 18.1 76.5 61.7 ECAT20 2.01 0.10 0.65 1.51 68.7 19.8 73.7 61.2 ECAT2F1R 1.53 0.17 0.70 1.84 76.3 15.4 78.7 55.3 ECAT2F10 1.51 0.16 0.74 1.87 74.0 17.6 75.8 57.4 CPSR 2.09 0.16 0.60 1.73 71.7 17.3 75.4 64.4 CPSO 1.97 0.34 0.92 3.73 70.9 17.7 68.6 63.9 24CD 2.24 0.22 0.74 2.56 64.2 18.7 70.7 65.3 48CD 2.24 0.18 0.79 2.77 64.1 19.4 69.0 65.8 48CDR 2.33 0.10 0.73 2.10 67.4 18.6 72.0 67.1 48CDO 2.12 0.19 0.87 2.84 65.1 17.8 70.6 65.9 (a) [conversion/(100-conversion)]; (b) [dry gas yield/activity]; (c) [coke yield/activity]; (d) [ (GLP olefins yield/GLP yield)* 100]; (e) [(product yield/conversion)* 100]; (f) [LCO yield/(LCO yield + HCO yield)]. Finally, the V concentrations on particle surfaces of sample CPS were only slightly higher than those determined by AA while no Ni was detected. Thus, we conclude that metals in the particles of sample CPS are distributed throughout the particles and more homogeneously than in CD samples. These results are roughly in line with what has been reported [5] regarding metal distribution in equilibrium catalysts and catalysts prepared by cyclic deactivation and impregnation. Metal poisoning of catalysts is reflected directly in the MAT data (Table 2). Significant differences are evident depending on the pretreatment atmosphere. In line with previously reported data for Ni [6] and V [7] the oxidative treatment of catalysts leads to comparatively higher dehydrogenation and hydrocracking activity. The lab-metallated catalysts are especially sensitive to these changes. The H2 pretreatment of metallated samples leads to product selectivity similar to that of equilibrium catalysts. By contrast, higher yields of coke and hydrogen are observed following pretreating in oxygen of 48CD and CPS. Under these conditions sample CPS produces the highest yields of coke and hydrogen. Those results indicate that lab-metallated samples contain metals of comparatively higher activity than in the equilibrium catalysts. The metals in samples ECAT are less active after longer aging in the FCC unit. Therefore we conclude that the differences observed in MAT data between oxidized and reduced catalysts reflect differences in the age of metals on the catalyst surface. The distribution and the effect of metals in CD samples can be explained by a coupling of the time basis for metal deposition and the kinetics of metal aging. Especially the requirement for Vanadium is longer compared to that of nickel as the former needs somewhat special conditions to produce mobile reactive species [4]. Simply put, the ratio of the agingtime to the metal addition

461 rate is much smaller for a lab sample compared to that of an equilibrium catalyst in an industrial unit. Metal porphyrins normally found in FCC feedstocks are comparatively more difficult to crack and condense in a reducing atmosphere [2] than lighter metal species. The nature of organometallics used in this work facilitates more rapid deposit and cracking on the catalyst. By comparison, the industrial operation involves a lower rate of metal deposition over the catalyst requiring a longer time before the metal concentration becomes significant or the vanadium redistributes throughout the catalyst particles. As metallation proceeds in a cyclic deactivation operation metals concentrate upon the external surface of the catalyst particles. Under these conditions V and Ni aging is difficult to achieve. Regeneration under conditions of higher temperature and steam partial pressure (more favorable for V mobility) might improve aging and should be recommended as variables for further study. The results of the mechanism described above are (1) that V and Ni concentrate on the surfaces of catalyst particles (as shown by EDS), (2) a high degree of metal sintering occurs on the catalysts particles (as observed by SEM) and (3) active metals are very sensitive to oxidationreduction treatments. We interpret the lower MAT catalytic activity (w.r.t. the zeolite surface area) as due to a high coke-makingtendency of the catalyst leadingto pore mouth plugging and increased intraparticle diffusional constraints. It has been shown that Ni is more active as it deposits near the zeolite component [ 10] or as it is more dispersed [4]. These two factors would help to explain the high gas, hydrogen and coke factors obtained with the CPS sample, especially in the oxidized state. Similarly, the relatively lower MAT activity of this sample could be explained either by an extensive matrix site poisoning or a large increase in the coke make. Although this sample has higher Ni content, the large differences in MAT product selectivity found between the oxidized and the reduced samples lead us to conclude, similarly to that for the CD samples, that metal aging was not readily accomplished. 4. CONCLUSIONS The properties of a commercial FCC catalyst modified with V and Ni by cyclic deactivation (CD) and impregnation-based laboratory procedures were compared to the properties of equilibrium catalysts tested in a commercial FCC unit. Principle differences were the degree of metal aging and concentration profiles in the catalyst particles. The CD procedure produced properties in the catalyst that were comparatively closer to those exhibited by the equilibrium catalysts. Differences in MAT numbers between oxidized and reduced catalysts are proposed to be related to the extent of metal aging. Metals were comparatively more active in the metallated samples than in the equilibrium catalysts. Although this was less evident for the catalyst deactivated by CD, it was observed that the metal aging procedure failed to reproduce completely the state of metals found in equilibrium catalysts. Therefore the aging process of deposited metals still remains a key factor for simulating more closely the properties of FCC equilibrium catalysts. REFERENCES 1. 2. 3.

A.W. Chester, Ind. Eng. Chem. Res., 26 (1987) 833. R. Pompe, S. Jaras and N-G.Vannerberg, Appl. Catal. 13 (1984) 171. M.L. Occelli, Catal. Rev. Sci. Eng., 33 No. 3&4, (1991) 241.

462 4. 5. 6. 7. 8. 9. 10. 11.

R.H. Nielsen and P.K. Doolin ; J.S. Magee and M.M. Mitchell (eds.) Fluid Catalytic Cracking : Science and Technology, Studies in Surface Science and Catalysis 76, Elsevier Sc. Pub. B.V., Amsterdam, 1993 P. O'Connor, A.C. Powels and H.N. Wijingaards, Proc. Int. Symp. on Deactivation & Testing of Hydrocarbons Conversion Catalysts, Div. of Pet. Chem., Inc., 210 th Meet., ACS, Chicago, IL, August (1995) 392. G.L.Woolery, M.D. Farnos and A.R.Quinones, Proc. Symp. Adv. FCC Conv. Cat., Div. of Pet. Chem., Inc., 211 th Meet., ACS, 42 No. 2, New Orleans, LA, March (1996) 403. J.A. Rudesill and A.W. Peters, Proc. Symp. Adv. FCC Conv. Cat. Div. of Pet. Chem., Inc., 211 th Meet., ACS, 42 No. 2, New Orleans, LA, March (1996) 407. B.R. Mitchell, Ind. Eng. Chem. Prod. Res. Dev. 19 (1980) 209. R.N. Cimbalo, R.L. Foster and S.J. Wachtel, Oil & Gas J., 70 No. 20 (1972) 112. P.F. Schubert and C.A. Altomare, M.L. Occelli (ed.), Fluid Catalytic Cracking: Role in Modem Refinery, ACS, Washington, 1988. ASTM Annual Book of ASTM standards, ASTM 1916 Race Street, Phil., PA., section 05 Petroleum Products and Lubricants (III) ; Catalysts ; V.5.03, 1991.

9 Elsevier Science B.V. All rights reserved Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

463

A F M a n d X P S Studies o f T h i o p h e n e and 1-Butanethiol D e a c t i v a t i o n o f Pd/A120 3 M o d e l C a t a l y s t s D u r i n g 1,3-Butadiene H y d r o g e n a t i o n K-H Lee, R. Catani a, R. Miglio a, and E.E. Wolf. Chemical Engineering Department, University of Notre Dame, Notre Dame, IN, 46556, USA. a On leave from SNAMPROGETTI, San Donato Milanese, ITALY. AFM and XPS data for model Pd supported catalysts show that the morphology of these catalysts is greatly affected by reduction in hydrogen and by addition of sulfur compounds to the surface before reaction. It was found that reduction of the catalyst before thiophene addition decreases the S/Pd ratio at which complete deactivation occurs. It was found that both the activation energy and the pre-exponential factor are quasi linear functions of sulfur coverage. Kinetic results after 1-butanethiol addition reveal a decreasing activation energy and pre-exponential parameter with increasing S/Pd ratio. Sulfur adsorption not only reduces the number of active sites but also weakens the adsorption strength of the remaining sites. It is demonstrated that physico-chemical information from AFM & XPS combined with kinetic data enables prediction of conversion at other conditions based on the S/Pd ratio. 1. INTRODUCTION Pd is one of the most active and selective catalysts for the partial hydrogenation of alkynes and alkadienes both in liquid or in gas phase. Poisoning of Pd catalysts by sulfur impurities is an important problem in many catalytic processes [ 1]. In addition to site blocking, sulfur adsorbed on noble metals sometimes forms new surface sites which improve reaction selectivity or reduce catalyst deactivation due to coking [2]. Sulfided compounds are toxic for catalytic surfaces due to a doublet of free electrons in their valence orbitals [3]. These unsaturated sulfur molecules form surface bonds through their n electrons with transitional metals surfaces and block active sites. The toxicity of the sulfided compounds varies depending on the nature of the active metals, the support, and reaction environment [4-6]. Sulfur poisoning not only affects the activity, but can sometimes alter catalyst selectivity [7]. While the literature of sulfur poisoning is very extensive, there are only a few studies on supported catalysts in which effects of sulfur on surface structure have been accounted for [ 8,9]. Following previous studies along the same lines [ 10-13], the objective of this work is to determine a correlation between catalytic activity, morphology, and the sulfur content. To this effect, we combine activity studies using a probe reaction with atomic force microscopy (AFM) and X-ray photoelectron spectroscopy (XPS) studies on model thin film catalysts. Two sulfur precursors were studied: thiophene and 1-butanethiol. The probe reaction studied is the hydrogenation of 1,3-butadiene because of its high reactivity on low-surface area Pd model catalysts [ 11-12], and because its wellstudied mechanism [ 14-15].

2. EXPERIMENTAL 2.1. Thin film preparation Thin film catalysts were prepared by vacuum evaporation of a Pd wire onto a single crystal of sapphire A1203 (Union Carbide Corp.). The A1203 substrate was cut into sixteen pieces (1.1 by 0.75 cm each) which were placed in a 30 ~ tilted platform in a vacuum evaporator (Denton DV-502) to allow uniform deposition of Pd.

464

2.2. Activity measurements A piece of film catalyst was placed in a quartz tube reactor (0.8 cm ID) and pretreated as specified below. Time on stream activity studies were carried out at 100~ or other temperatures for various periods of times. Selectivity versus conversion measurements were carried out by varying the temperature. Kinetic studies were conducted at lower temperatures to achieve differential reaction conditions (<15% conversion). The reactive mixture consisted of 1,3-butadiene (Matheson Gas Products), hydrogen, and nitrogen. The flow rates were as follows: butadiene 0.6 cc/min. (4 torr), hydrogen 75 cc/min, and nitrogen 36.5 cc/min. A high hydrogen-to-butadiene ratio (H2/HC=125) was used to reduce carbon deposition and minimize catalyst deactivation due to this factor. The reaction products (1-butene, n-butane, cis-2-butene and trans-2-butene) were separated in a gas chromatograph (GC). The GC was equipped with a 7 ft chromatographic column packed with 0.19% picric acid/graphpac packing (Alltech Associates, Inc.) and a FID detector.

2.3. Atomic Force Microscopy The AFM used in this study was a Nanoscope II system (Digital Instrument Inc.) operated at ambient conditions. Two AFM scan heads (700 nm range and 8000 nm range) with silicon nitride probes were employed. The AFM was operated in the height imaging mode and at low scan frequency (< 2 Hz). AFM analysis was conducted ex-situ on the fresh film, after H2 pretreatment, after S addition, and after the catalytic reaction. Four areas from each piece of catalysts were imaged before and after each reaction. The areas scanned varied from 100 x 100 nm 2 to 4000 x 4000 nm 2.

2.4. X-ray Photoelectron Spectroscopy The XPS analyses were performed in a Kratos 800 spectrometer using monochromatic Mg K s radiation (hv = 1,253.6 eV). The electron binding energy scale was calibrated by assigning 284.6 eV to the C l s peak position. The S/Pd surface ratio was calculated from the S 2p and Pd 3d signals. The samples were exposed to air while being transferred from the reactor to the XPS chamber. Pitchon et al. [15] have shown via XPS that Pd/SiO2 catalysts exposed to ambient-conditions exhibited no measurable surface oxidation, an effect confirmed by our measurements.

2.5. Sulfur poisoning Two sulfur containing precursors were studied: thiophene and butanethiol. Attempts to add the S containing precursor from nES , and from a gas phase stream bubbled through a solution containing the S precursor did not result in reproducible results due to the low surface area of the catalysts. The procedure adopted was to add directly liquid droplets of a cyclohexane-thiophene or cyclohexane-1butanethiol solution to the Pd films. Addition from the liquid phase yielded satisfactory reproducibility both in terms of the amount of S adsorbed by the films, as well as the activity after sulfur poisoning. Solutions with various concentrations of precursors were used to vary the amount of sulfur added to each sample. One cc of the solution was added via a syringe to the Film catalysts (fresh and pre-reduced). The films were then placed on a hot plate at 60~ to evaporate the solvent. 3. R E S U L T S

AND D I S C U S S I O N

3.1. Effect of reduction temperature on unpoisoned catalysts We first studied the effect of the reduction temperature (200-500~ on the unpoisoned films. AFM micrographs of the ct-A1203 single crystal substrate reduced at 500~ showed that the surface of the substrate appears close to atomically smooth. The Pd film catalysts were then imaged by AFM after hydrogen reduction. The initially smooth films sintered and broke into islands of varying size and height, even when the reduction temperature was as low as 200~ The films treated at 300450~ showed isolated islands of about 200 nm in diameter and 10 nm height. The AFM results obtained on the unpoisoned catalysts agree fairly well with previous results showing that exposure to

465

air during transfer to the microscope did not affect the microstructure, and that major morphological transformations were induced by the reaction environment [ 11-12]. Conversions obtained on films (-0.3-0.2 cm 2 ) reduced at various temperatures are summarized in Table 1. At low reaction temperatures, conversion decreases with increasing reduction temperature and then levels off at about 24% for reduction temperatures < 450~ At the higher temperatures the conversion is insensitive to the surface structure because of mass-transfer control of the reaction rate. Table 1 Effect of reduction temperature on conversion and kinetic parameters Temp of Conversion at Pd area Ea reduction ~ 30~ 45~ 60~ 100~ (cm 2) (Kcal/mol) 200 300 400 500

11.4 4.5 2.9 <1

20.7 8.1 9.0 1.5

24.4 23.8 24.4 12.2

24.8 24.8 24.8 13.0

0.3 0.21 0.19 0.19

17.2 15.8 14.1 21.1

A

(cm/hr) 1.07 3.62 2.51 3.89

E+ 11 E+9 E+10 E+ 12

Kinetic parameters obtained at differential conversion (< 15%) at different reduction temperatures are also summarized in Table 1. The activation energy and pre-exponential factor decrease and then increase with increasing reduction temperature up to 500~ Similar results involving activity loss with increasing reduction temperature have been reported during cyclohexane dehydrogenation and methane combustion on Pd catalysts [ 16,17]. These reports indicate that increasing the temperature of reduction up to 450~ decreases the reaction activation energy. The AFM results show that the reduction temperature affects the surface morphology thus changing the type of active sites. In addition, the reduction temperature could affect H2 solubility in Pd, which in turn could affect reaction kinetics and the rate of poisoning.

3.2. Effects of thiophene addition Thiophene was first added to the surface of fresh unreduced catalysts and then to the surface of reduced catalysts. Fresh unreduced catalysts. The S/Pd ratio remaining on the surface of unreduced films, as determined by XPS, is rather small when compared to the S/Pd ratio originally fed onto the catalysts. When the S/Pd ratio in the original solution was 10n, the following S/Pd ratio was obtained by XPS: n S/Pd

1 0.11

2 0.13

3 0.17

4 0.28

5 0.55

Even when pure thiophene was added to one film catalyst (n = 5), the surface S/Pd ratio is 0.5 5. The surface S/Pd ratio increased in a logarithmic relation with the sulfur concentration in the liquid added to the fresh catalysts. Thus, roughly, of 10" sulfur atoms per Pd atom added only about n/10 remains on the surface. It has been suggested [ 18] that thiophene adsorbs in a planar configuration with the ring facing the surface or in a perpendicular form with S strongly bonded to the surface. The planar form adsorbs reversibly at low temperature and activity is recovered when thiophene is removed from the feed. The second form is irreversibly adsorbed via hydrogenolysis of thiophene with evolution of n-butane leaving an inactive metal-sulfur species. In our case, we speculate that that only strongly chemisorbed thiophene molecules remain on the surface. All unreduced, poisoned catalysts were subjected to reaction at 100~ hence during this period, adsorbed thiophene can hydrocrack leading to the formation of adsorbed sulfur as shown by the binding energy of the S 2p, line which corresponds to elemental sulfur.

466 Conversion versus time-on-stream (TOS) data for various S/Pd ratios added to fresh catalysts (unreduced, 0.82 cm 2, 100~ are presented in Table 2. Initial conversion decreases with S/Pd ratio but only changes slightly from 1 to 5 hours of TOS. When the conversion at 100~ and 1 h of TOS is plotted versus S/Pd ratio for the poisoned, unreduced catalyst, it yields an almost linear decrease of 1,3-butadiene conversion versus S/Pd ratio (not shown). Selectivities for 1-butene and 2-transbutene increase as S/Pd ratio increases to 0.1 but change only slightly with further S addition. Table 2 Effect of thiophene addition S/Pd: 0 0.11 TOS (min) 10 65 52 30 64 51 60 63 50 120 61 49 lnA

Ea

25.2 20.0

18.4 15.1

0.13 0.17 Butadiene conversion.(%) 51 42 50 44 49 44 48 42 Kinetic parameters 17.1 14.6 14.2 12.4

0.28

0.55

27 30 31 30

0 0 0 0

7.6 7.5

However, butane selectivity decreases dramatically from 71% to 3% with sulfur addition, in agreement with previous literature [1]. The natural log of the pre-exponential factors (In A) and activation energies (Ea [kcal/mol])for the various S/Pd ratios are also summarized in Table 2. Activation energy and the logarithm of the pre-exponential factor decrease linearly with S/Pd ratio. AFM images of the poisoned, unreduced catalysts were not well resolved. Addition of S to the fresh catalyst did not alter significantly the morphology with respect to the fresh surface, showing a relatively unbroken film surface. However after 5 hours of reaction, the surface of the continuous film broke, exhibiting large rifts and terraces. Pre-reduced catalysts. The previous experiments were repeated again but catalysts were reduced before poisoning. Reduction of the f'llm at various temperatures before sulfur deposition decreased very dramatically the rate of reaction compared to fresh unreduced films. Total deactivation is attained at much lower levels of surface sulfur poisoning than in the case of the unreduced catalysts. For a S/Pd ratio of 0.18, conversion decrease from 60 (fresh unreduced) to 7% when the reduction temperature is as low as 200~ For a reduction temperature of 300~ only a 1% conversion is measured, and no conversion is detected when the reduction temperature is 400~ The AFM images of these catalysts show that the surface breaks up in islands of varying sizes. As the reduction temperature increases, the sizes of these islands decrease, but their heights increase.

3.3. Effects of 1-butanethiol poisoning All the catalysts were pre-reduced in H2 at 250~ for 2 hours before adding 1-butanethiol. In this case the log of the 1-butanethiol loaded also correlated linearly with the S/Pd ratio (1 cc 1butanethiol yields a Pd/S=0.47 ). Fig. la shows that 1,3-butadiene conversion (at 100~ decreases with increasing Pd/S ratio and that deactivation with time on stream is relatively slow. The conversion after 2 hours of TOS versus the Pd/S ratio is shown in Fig. lb. Figure 2 shows a series of AFM images of the Pd/t~-A1203films after various pretreatments and degrees of poisoning and reaction. The fresh Pd/tx-A1203 (not shown) is a rather featureless surface. After the catalyst undergoes 2 h of HE treatment at 250~ the surface contained particles of Pd of about 16 nm thick and 100x75 nm 2 size on a background of exposed A1203 substrate (Fig. 2a). After 2 h of reaction the surface morphology did not change much except that particle thickness increases.

467

1~176 I

60

80

v

~

0

S/Pd=O. 11

40

20 . = 1~5 S~/ P, d ,=17 O~ . ~ . ~ . m

~-~

. .

.

.

.

.

.

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.

.

'-

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~:

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~

0.2

Time (min)

-

,

o.;

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Surface S/Pd ratio

Figure 1.1,3-Butadiene conversion at (100~

(a) vs. time on stream and (b) vs. S/Pd ratio

a

b ....,

~,~,~

:i:ii!!~'

0

eO0

c

i!ii!?~

~0

600

800

d 8

i 0

500

1000

1~5a0

0

~0

1000

~500

Fig. 2. AFM micrographs of Pd/A1203 model catalysts after: (a) hydrogen reduction, (b) cyclohexane addition, hydrogen reduction and reaction, (c) hydrogen reduction and reaction, S/Pd=O.12, (d) hydrogen reduction and reaction, S/Pd - 0.33.

468 Images obtained after cyclohexane addition, reduction and reaction, show islands interlinked with web-like connectivity (Fig. 2b). Figure 2c of a film with a S/Pd molar ratio of 0.11, shows a surface similar to that of Figure 2b. Further addition of S (S/Pd= 0.15), causes agglomeration of Pd grains, and the interconnectivity of the Pd particles is greatly reduced leading to the formation of bigger islands. At S/Pd= 0.33, Pd particles slowly grow to around 125 to 150 nm in size and a thickness of around 30 nm. With pure 1-butanethiol addition, Pd grains remain large but the substrate is less visible, probably also due to sulfur coverage on both the Pd film and substrate. As in the case of thiophene, the limiting value of 0.47 for the S/Pd ratio is related to how sulfur is adsorbed on the surface. STM images at UHV shows that. on hcp single crystals such as Ru(0001) and Re(0001), at S/Pd < 1/3, the p(2x2) and (~t'~ x ~f3 ) R30 ~ structures coexist on the surface. Between S/Pd = 0.45 and 0.50, sulfur trimers have slowly converted to tetramers. Another surface with a (2,~/3 x 2~/3 ) R30 ~ structure of sulfur hexagon ring is also observed at saturation coverage of sulfur of S/Pd = 0.5 [19]. Oudar [20] reported a saturation sulfur coverage at 0.85 or even higher. Although others [21 ] have reported that sulfur adsorbs as S 2-, $22, and S 6§ at different reaction conditions, the XPS S 2p peaks cannot be convincingly deconvoluted to address the existence of such oxidation states. The morphological transformation seen by AFM is due to the combined effects of thermal and reducing gas treatments. It was found that cyclohexane significantly alters the morphology of the fresh films, and this effect is also observed but to a lesser extent on poisoned catalysts. The wetting effect from cyclohexane addition diminishes at higher solution-phase concentrations of 1-butanethiol. The lower surface tension of cyclohexane compared to 1-butanethiol on Pd probably lowers the interfacial tension between Pd and the substrate and thus cyclohexane wets Pdhx-A1203 better than a 1-butanethiol/cyclohexane mixture. The increasing size of the particles with sulfur is believed to be associated with sulfur adsorption on the surface of the Pd particles. Vasquez and coworkers [22] concluded that the expansion of the particles was the result of shell formation, probably forming a surface sulfide phase. At the maximum sulfur coverage of 0.47, at which complete deactivation occurs, the particles appear to be more compactly arranged on the surface. Arrhenius parameters were also obtained for each fresh and poisoned catalyst. Table 3 shows that the activation energy and pre-exponential factor decrease slightly and almost linearly with S/Pd ratio. Turnover frequencies (TOFs) measured at 300K are also displayed in Table 3. Table 3 Kinetic parameters and turnover frequencies of l~3-butadiene S/Pd 0. 0.11 0.15 Ea 11.7 9.7 10.1 lnA 18.9 16.4 15.3 TON 51.3 19.6 18.0

hydrogenation 0.17 8.1 12.6 17.3

at various S/Pd ratios. 0.31 9.1 13.2 9.1

The kinetic results can be summarized by separating the decrease in the number of active sites a from S blocking obtained from the S/Pd ratio and from the AFM micrographs, from the effect of S on the activation energy Ea and the frequency factor kp. Assuming that a

Pdsr- 7

(1)

ao

Ea - Eo-~(~d )

(2)

469

kp

exp(X, S

kp,o =

P-d )

(3~ S

RC4H6 then RC4H6,o

=

Pdsr - ~

e ee e Rr

(4)

From S/Pd at zero conversion one obtains a value of T= 1.21. From AFM estimates of P d s r , the total metal area exposed and the S/Pd ratio, the value of [3 was found to be 2. ~, was obtained by fitting Eq. 4 to the normalized rate data and found to be -19.5. The correlation between experimental data obtained from the above kinetic results and the conversion (another set of experiments) at the various S/Pd ratios is shown as a solid line in Fig. lb. Selectivities of various poisoned and flesh catalysts were also measured at various S/Pd ratios. The cis and trans-2-butene selectivities at different conversion levels and S/Pd ratios are constant throughout the conversion range up to 75%. The 1-butene selectivity slowly decreases from 60% to about 5% throughout the conversion range whereas n-butane selectivity increases with increasing conversion. The trend of the selectivity changes with S/Pd ratio of these two products is almost a mirror image of each other. These results show that the selectivity does not change with S addition, indicating that the reaction pathway is not affected by sulfur. The Arrhenius parameters listed in Table 3 offer additional insight into the role of sulfur adsorption in the reaction. L'Argentiere [23] suggested that noble metals become more electron deficient with the formation of covalent S-metal bonds. Following this line of reasoning we speculate that the lower activation energy at low sulfur coverage indicates that 1,3-butadiene molecules adsorb on the surface metal-sulfide and form a metal-butadiene complex having a weak interaction with adsorbed sulfur. Unfortunately, the deconvolution of Pd 3d peak is not accurate enough to identify the Pd-S oxidation states. Changes in the pre-exponential parameters also indicate that the activity of surface sites is reduced significantly with increasing sulfur coverage, but increases slightly at higher S/Pd ratio of 0.31. Changes in the pre-exponential factor, A, are more drastic compared to those in activation energy. Also from the correlation in Equation 4, it follows that the number of active sites, a/ao, dominates the decreasing trend of the activity whereas activation energy is a less significant factor. The selectivity results show that the adsorption of 1-butanethiol does not affect the reaction pathway. The decreasing and mirror-increasing selectivity trends of 1-butene and n-butane, respectively, show that the two products are in series with 1-butene as the intermediate. On the other hand, cis- and trans-2-butenes follow their own reaction pathways. From this observation, it is concluded that sulfur adsorption will block sites and reduce the Pd bond strength but it does not alter the characteristics of the reactive complex. The unaltered selectivity trend of cis-2-butene is also an indication of a structurally insensitive reaction. Equation (4) shows that physical information of the catalyst such as surface structure and composition obtained by various techniques can be related to the catalyst reactivity. At S/Pd-0.47, even though almost half of the Pd atoms on the surface are still available, there are no active sites left. This implies that hydrogenation of butadiene molecules requires two conjugate Pd atoms. The second-order dependence of the normalized active sites, a/ao, also indicates that adsorption and surface dissociation are both the controlling steps of the reaction, kp/kp, o in equation [4] indicates that another decaying function is necessary to take into account the change in activity. Presumably carbon left over from the decomposition of the carbon precursor is responsible for this effect. This is likely due to an exponentially decaying sticking coefficient which implies that S not only blocks sites but also reduces the bonding strength of the remaining active sites. 4. C O N C L U S I O N S AFM and XPS data for model Pd supported catalysts show that the morphology of these catalysts are greatly affected by reduction in hydrogen and by addition of sulfur compounds to the surface

470 before reaction. Reduction of the catalyst before thiophene addition decreases the S/Pd ratio at which complete deactivation occurs. The effects of pre-reduction on sulfur uptake and activity are attributed to both a sintering effect, that decreases the number of active sites, and hydrogen solubility on Pd, which can affect the S-Pd bonds. Both the activation energy and the natural logarithm of the preexponential factor are linear functions of sulfur coverage. 1-Butanethiol addition decreases catalytic activity but not selectivity in 1,3-butadiene hydrogenation. Both activation energy and pre-exponential factor decrease with increasing S/Pd ratio. Sulfur adsorption not only reduces the number of active sites but also weakened the adsorption strength of the remaining sites. It has been demonstrated that physico-chemical information from AFM & XPS combined with kinetic data enables prediction of conversion at other conditions based on the S/Pd ratio. 5. ACKNOWLEDGMENTS The financial support of RC and RM during their stays at Notre Dame by SNAMPROGETTI, and the support of the National Science Foundation Grant CTS 92 15339 are gratefully acknowledged. REFERENCES

1. J. Oudar, Catal. Rev.- Sci. Eng., 11(2) (1980) 171. 2. 1261126]. L.J. Hoyos, M.Reimet & H. Praliaud, J. Chem. Soc. Faraday Trans., 88 (22), (1992) 3367. 3. E.B. Maxted and A.G. Walker, J. Chem. Soc., (1940) 252. 4. P. Gallezot, J. Datka, J. Massadier, M. Primet, and B. Imelik, Proc. 6th Int. Congr. Catalysis, (1976) A11, London,. 5. H.P. Bonzel, and R. Ku, J. Chem. Phys., 58, (1973) 4617. 6. M. Boudart, A. Aldag, J.E. Benson, V.A. Dougharty, and C.G. Harkins, J. Catal., 6, (1966) 92. 7. N.J. Kirkpatrick, Adv. Catal. 3 (1951) 329. 8. C. Lee and L. D. Schmidt, J. Catal., 101 (1986) 123. 9. T. Wang and L. D. Schmidt, J. Catal., 78 (1982) 306. 10. K. L. Yeung and E. E. Wolf, J. Catal., 143 (1993) 409. 11. K. L. Yeung, K.H. Lee and E.E. Wolf, J. Catal., 156 (1995) 120. 12. K.H. Lee and E. E. Wolf, Catal. Lett., 26 (1994) 297. 13. B. Tardy, C. Noupa, C. Leclercq, J. C. Bertolini, A. Houreau, M. Treilleux, J. P. Faure and G. Nihoul, J. Catal., 129 (1991) 1. 14. C.M. Pradier and Y. Berthier, J. Catal., 129 (1991) 356. 15. V. Pitchon, M. Guenin and H. Praliaud, Appl. Catal., 63 (1990) 333. 16. L. J. Hoyos, H. Praliaud and M. Primet, J. Chem. Soc. Faraday Trans., 88(22) (1992) 3367. 17. L. J. Hoyos, H. Praliaud and M. Primet, App. Cat., 98 (1993) 120. 18. K. Ahmed, D. Chadwick and L. S. Kershembaum, in Catalyst Deactivation, 1987, p. 513, B. Delmon and G. F. Froment (Edits), Elsevier, Amsterdam. 19. B. Marchon, D. F. Ogletree, M.Salmeron, and W. Siekhaus, J. Vac. Sci. Technol. A, 6 (1988) 531. 20. J. Oudar, in "Deactivation and Poisoning of Catalysts", (1985) 51, J. Oudar and H. Wise (Edits), M. Dekker, New York. 21. A. Arcoya, A. Cortes, J.L.G. Fierro and X.L.Seoane, in "Catalyst Deactivation" (1991) 557, C.H. Bartholomew & J.B. Butt (Edits.). Elsevier, Amsterdam). 22. T.H. Fleisch, R.F. Hicks, A.T. Bell, J. Catal., 87 (1984) 398.

9

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Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

471

S u l f u r P o i s o n i n g o f N i c k e l - b a s e d Hot Gas C l e a n i n g Catalysts in Synthetic Gasification Gas J. Hepola and P. Simell VTT Energy, Energy Production Technologies, PO Box 1601 FIN-02044 Finland

The sulfur distribution and content of nickel catalyst beds were analyzed to account for poisoning effects of sulfur on the activities of catalysts which decompose tar, ammonia and methane in synthetic gasification gas. The desorption behavior of chemisorbed sulfur from the bed materials was monitored by temperature programmed hydrogenation (TPH). Sulfur adsorbs on nickel catalysts in different chemical states depending on the process conditions. At > 900~ the sulfur adsorbs on the catalyst forming an irreversible monolayer on catalyst surfaces, while at < 900~ adsorbed sulfur, probably composed of polysulfides (multilayer sulfur), is desorbed from the catalyst in a sulfur-free hydrogen-containing atmosphere. However, a monolayer of sulfur still remains on the catalyst after desorption. The enhanced effect of high pressure on sulfur-poisoning of nickel catalysts can be attributed to an increased amount of adsorbed sulfur, probably in the form of polysulfides. In addition, it was established that bulk nickel sulfide is active in decomposing ammonia in hightemperature gasification gas-cleaning conditions. The activity for decomposing methane is not affected by bulk nickel sulfide formation, but that of toluene is decreased.

1. I N T R O D U C T I O N Nickel-based catalysts have proven to be very efficient in decomposing tar, ammonia and methane in biomass gasification gas mixture at about 900~ [1-5]. In biomass gasification gas the H2S content is normally below 100 ppmv. When peat is used as feedstock the HaS content of gas may increase to 500 ppmv. However in coal gasification the content of sulfur in the gas can be much higher. It could be expected that tar, ammonia and methane compete, at least partly, for the same active nickel sites on the catalyst surfaces during the desirable H20/CO 2 reforming of hydrocarbons and ammonia decomposition reactions. However, especially under pressurized conditions the effectiveness of the catalyst is decreased due to sulfur adsorption on the nickel surfaces. In principle, to mitigate sulfur poisoning of nickel catalysts in the tar and ammonia decomposition process, the catalytic process should be operated at > 900~ On the other hand, we observed that ammonia conversion is enhanced by high H2S concentrations in the gas. The ability of nickel catalysts to decompose tar (toluene), ammonia and methane in synthetic gasification gas mixtures with or without H2S is reported and discussed in two complementary studies [6,7]. The results are based on experiments carried out in laboratoryscale fixed-bed tube reactors at 800-1000~ under 1, 5 and 20 bar total pressure. In this paper we present the results obtained in those studies when analyzing spent catalysts from sulfur-poisoning experiments to account for the different deactivating effects of hydrogen sulfide at different test conditions. The desorption of chemisorbed sulfur from bed materials was monitored by temperature-programmed hydrogenation (TPH) under argor~ydrogen atmosphere.

472

2. EXPERIMENTAL Atmospheric and pressurized sulfur poisoning tests were carded out in fixed-bed quartz reactors using synthetic gas mixtures. The atmospheric apparatus in [8] and the pressurized apparatus in [9] are described in detail. Small modifications in the testing facilities were made before the tests, mainly in the gas feeding (H2S/N2) systems and in the dimensions of the quartz reactors. In pressurized tests the H2S/N 2 mixture was injected directly to the inlet of the quartz reactor where it was mixed with the other gases before the catalyst bed. In atmospheric tests the H2S/N2 gas was mixed with the heated main gas stream in a tefloncoated mixing chamber before feeding to the quartz reactor. A high-purity calibration gas mixture of HES/N2 (+3%) was used in the experiments. Catalyst beds were completely saturated with sulfur during poisoning tests. In each experiment the catalyst bed was exposed to the desired H2S concentration for 4 - 6 hours to stabilize the catalyst performance. After the experiments, the catalyst beds were cooled to room temperature under dry nitrogen flow before opening the reactor vessel. In pressurized tests the pressure let down was not started until 20-30 min after nitrogen flushing to be sure that all remaining reactive gas components had been removed. Sulfur desorption from the poisoned nickel catalysts was studied by temperature programmed hydrogenation (TPH). The poisoned catalyst beds were powdered to homogenize them. Approximately 0.5 g of powdered catalyst was placed in an atmospheric quartz tube reactor which had a thermocouple shield. The reaction gas was argon/hydrogen (70% Ar/30% H2). The content of hydrogen in the gas was about the same as in the sulfurpoisoning tests. Gases leaving the catalyst bed were continuously analyzed by withdrawing a small fraction of the reactor effluent into a GAM 415 mass spectrometer. Changes in the intensity of the sulfur compound (mass numbers: 32, 33, 34, fragment ions: S § H2S+) peaks were monitored during the desorption tests. The sample was indirectly heated by an electrically heated oven surrounding the reactor. The reactor was heated from room temperature up to 970~ at a rate of 20~ The heating rate was controlled by a programmable temperature controller. The catalyst bed temperature was monitored continuously with a shielded thermocouple. To determine nickel surface areas of fresh catalysts, hydrogen chemisorption measurements were performed with the atmospheric TPH apparatus described above. Nickel surface areas of the catalysts were also measured by static volumetric hydrogen chemisorption (Coulter Omnisorp 100 cx). Before these measurements, samples were reduced in situ in a pure hydrogen atmosphere by raising the temperature up to 900~ (20~ The samples were then cooled in a helium flow and the measurements were performed at 30~ The catalyst materials and the analytical methods are described in [4,6,7].

3. RESULTS AND DISCUSSION 3.1. Effect of high hydrogen sulfide gas concentration In our pressurized tests [6,7] we found that when the H2S content of the gas was as high as 500 ppm the conversion of ammonia started to increase compared to lower sulfur levels. We also carded out additional experiments to see the effects of increasing H2S to a level where bulk sulfide formation is expected to form based on thermodynamics. Fig. 1 shows the effect of the high H2S content of the synthetic gasification gas on the conversion of toluene, methane and ammonia at 900~ 20 bar SV=15000 1/h. The experiments were performed by keeping the catalyst at each of the H2S levels for 4 - 6 hours to stabilize the catalyst performance. At high sulfur levels (> 500 ppmv) the stabilization of the catalyst

473 performance was slower than at lower sulfur levels. An interesting behavior was seen when the concentration of H2S was increased in the gas: the conversion of ammonia increased to nearly the same level as without H2S in the gas. On the other hand, the conversion of toluene decreased as the H2S content of the gas increased. However, the conversion of methane was not affected by high sulfur levels. The sulfur content and state of the catalyst beds were analyzed after the tests. An X-ray diffraction analysis revealed that bulk nickel sulfide was indeed formed at high sulfur levels (1000-2000 ppmv). It can be concluded that bulk nickel sulfide is, in addition to metal nickel, active in decomposing ammonia in the high-temperature gasification gas cleaning conditions. Moreover, the activity for ammonia decomposition of the catalyst increases even before bulk nickel sulfide formation, indicating that the change of activity caused by adsorbed sulfur species is not sudden. The observation is in agreement with information in two patents [ 10,11], according to which metal sulfides (iron, nickel, cobalt, molybdenum, vanadium, thorium) are active in decomposing ammonia compounds in high-temperature gas mixtures. In addition, according to Jobic et al. [ 12] and Zeuthen et al. [ 13], the dissociative adsorption of ammonia occurs on sulfide catalysts and the reaction with ammonia results in replacement of some catalyst sulfur. The indication that dissociative adsorption of ammonia is possible on sulfide catalysts supports the fact that dissociative adsorption can probably also occur on sulfide nickel in the conditions of the present study.

120

Toluene

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.

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1,000 1,500 H 2S (ppm)

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2,500

Figure 1. Effect of H2S on the conversion of toluene, ammonia and methane, Catalyst A1, T= 900 ~ P= 20 bar, SV= 15000 1/h. Synthetic gasification gas.

474 An observation of interest in the case of methane is that conversion is insensitive to bulk nickel sulfide formation on the catalyst. It was suggested by Rostrup Nielsen and Alstrup [14,15] that the ensembles of free nickel atoms available at high coverages of sulfur are sufficient for the conversion of adsorbed methane with steam or carbon dioxide. On the basis of the present study even the bulk nickel sulfide appears to have sufficient active sites for methane decomposition. When the sulfur content of the gas was low (< 500 ppmv) the conversion of toluene was fairly insensitive to sulfur. Rostrup Nielsen [14,15] and Ng and Martin [16] suggested that hydrogen from adsorbed H2S probably interacts with unsaturated hydrocarbons to give an hydrogenated adspecies. The decrease of toluene conversion when the sulfur content of the gas was high (< 500 ppmv) can be explained by the large molecular size of toluene/benzene compared to methane. Therefore, the tar decomposition could need more active nickel sites than methane. During increased sulfur adsorption and finally bulk nickel sulfide formation there are apparently not enough active nickel sites available on the catalyst for tar decomposition.

3.2. Sulfur and carbon contents of catalyst beds Table 1 shows total sulfur and carbon contents of Ni catalysts after poisoning in a fixed bed. Sulfur content varied with temperature, pressure and H2S concentration. In addition, the content of sulfur was affected by the catalyst type. The homogeneous catalyst with a calcium aluminate carder (C) adsorbed more sulfur under the same conditions. This is probably due, in part, to the calcium content of the catalyst [ 17] in contrast to Catalyst (A) with an A1203 carrier. The sulfur and carbon contents of selected spent catalyst beds (900~ 20 bar) were analyzed as a function of bed length (Table 1). The content of these components of the bed proved to be higher at the inlet than in the middle and bottom of the bed. Due to endothermic reactions, i.e., steam reforming of hydrocarbons and ammonia decomposition, the bed temperature decreased considerably (max. about 100~ at 20 bar) at the top of the catalyst bed, but after poisoning the temperature increased again [6]. This demonstrates, in addition to a strong non-uniform adsorption of sulfur on metals, higher sulfur and carbon contents at the inlet of the bed compared with those of the other parts of the bed. Probably due to higher temperatures the contents of sulfur and carbon were lower at the inlet of the bed when the sulfur content in the gas was higher (Table 1). The interaction of adsorbed sulfur and carbon formation could also have influenced the carbon content. The higher ammonia conversions obtained at higher sulfur levels in the gas may have been partly due to the temperature distribution. However, this cannot be the main reason as there was no temperature gradient in the bed during atmospheric experiments. Nevertheless, the uneven distribution of the poison as a function of location in the fixed bed complicated fundamental interpretation of experimental data. According to ESCA and SIMS analysis sulfur on the catalyst surfaces was, on the basis of binding energies (161-162 eV for sulfides) in the form of sulfides. No elemental sulfur was detected by ESCA measurements when the analyzed sample was cooled to liquid nitrogen temperature before analysis, to avoid the loss of easily volatized sulfur from the sample. On the other hand, when the content of sulfur in the gas was sufficiently high (1000-2000 ppmv) for bulk nickel sulfide (NiaS2) formation (ascertained by X-ray diffraction analysis), the distribution of sulfur in the catalyst bed was fairly even in different parts of the bed (Table 1) and also within catalyst particles. According to SEM/EDS analysis the content of sulfur on the surface ranged from 4.0 to 7.5 wt% and in the middle of particles from 5.3 to 9.3 wt%. However, nickel forms a liquid sulfide product at temperatures above 635 ~ [ 18, 19]. According to the SEM analysis of the fresh and spent catalyst particles, no liquid formation on the catalysts was detected.

475

Table 1 Sulfur and carbon content of nickel catalyst samples after poisoning tests in fixed bed reactors. Sulfur behavior in TPH experiments. Catalyst Process conditions Sulfur and carbon content of catalyst beds TPH-experiments T P H2S S, total S/C, top S/C, middle S/C, bottom S desorption ~ bar ppmv wt% wt% wt% wt% (S after the test) A2 800 1 500 0.06 yes, (S= 0.02 wt-%) A2 800 1 2 000 3.1 yes A2 900 1 500 0.02 no A2 800 20 200 0.18 yes (S---0.02wt-%) A1 900 5 500 0.05 no (S= 0.043 wt-%) A1 900 20 150 0.21 0.55/0.51 0.046/<0.1 0.044/<0.1 A1 900 20 500 0.12 0.23/0.24 0.086/0.18 0.054/0.12 A1 900 20 500 0.11 yes A1 900 20 1 000 3.2 3.7/0.39 3.4/<0.1 2.4/<0.1 A1 900 20 2 000 3.3 2.8/7.1 3.5/<0.1 3.5/<0.1 yes C 800 5 500 0.26 C 900 5 500 0.12 no C 800 20 500 0.42 yes C 900 20 125 0.07 C 950 20 500 0.07 C 970 20 500 0.05 AlzO~ 900 20 500 <0.02 no

3.3. Sulfur desorption from poisoned catalysts To better understand the behavior of sulfur chemisorption on the spent nickel catalysts at the applied conditions, TPH experiments were performed to determine how sulfur desorbs from the poisoned catalysts treated differently in fixed-bed poisoning tests. Fig. 2 shows the desorption of sulfur (mass number 34) from catalysts poisoned in different conditions as a function of temperature under Ar/H 2 atmosphere. Qualitative results of the desorption tests are also listed in Table 1. TPH tests with pure alumina (alpha) indicated that sulfur was not adsorbed on this material during fixed-bed poisoning tests, although sulfur adsorbed on nickel catalysts supported on alumina, thus indicating it adsorbs on the surface of nickel only. Tests with a pure alumina (alpha) bed also indicated that hydrogen was not adsorbed on it at the conditions for nickel surface area measurements by hydrogen. It can be seen from the Fig. 2 that sulfur desorption from the catalyst beds (Catalysts A & C) which had been poisoned at 800-900~ under 1 and 20 bar pressure begins when the temperature of the bed is above 400~ the most part being desorbed rapidly between 500700~ On the other hand, when the H2S concentration of the bed is sufficiently high for bulk nickel sulfide formation, the desorption of sulfur in the atmospheric tests occurs at the same temperature as the catalyst which has been treated in fixed-bed poisoning tests. However, in the 20 bar tests the desorption of sulfur begins at about 650~ This temperature is lower than the bulk nickel sulfide formation temperature of about 900~ but higher than the above mentioned desorption temperature of adsorbed sulfur species. Therefore, some adsorbed sulfur, in addition to bulk nickel sulfide, may have been present on the catalyst. Fig. 2 shows that the desorption of sulfur in the case of bulk sulfide occurs more slowly than in the case of chemisorbed sulfur.

476

AI/900 *C/20 bar H25=500 ppm

C/800 0C/20 bar H2S--.q00ppm ~D

A1/900 *C/20 bar H2S=2000 ppm

A2/800 *C/1 bar H2$=2000 ppm

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300

400

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600 700 Catalyst temperature ~

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800

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900

1,000

Figure 2. Temperature-programmed hydrogenation of sulfur from poisoned nickel catalysts (A 1, A2 and C). Heating rate 20~ gas atmosphere Ar/H v Some other results of other TPH experiments can be seen in Table 1. In atmosphericpressure tests at 900~ with 500 ppm H/S in the gas phase, sulfur was not desorbed from Catalyst A 1. The same phenomenon was noticed in the tests performed at 900~ under 5 bar pressure with Catalysts A2 and C. In addition, when the sulfur content of the catalyst beds was analyzed after TPH experiments, it was observed that only a small amount of sulfur was present on the catalyst. This observation indicates that sulfur adsorption is not completely reversible, but that part of the adsorbed sulfur remains on the catalyst. The effect of this phenomenon was also observed when a catalyst was regenerated by removal of H/S to the gas mixture in fixed-bed poisoning tests. The catalyst activity did not reach the original level (with no H2S in the gas) especially in ammonia decomposition. The analysis of the sulfur content of the bed showed that a small amount of sulfur was still present on the catalyst. The composition of the saturation layer (two-dimensional sulfide) has been determined by different methods [20, 21]. The sulfur content is close to 445 ~tgS/m 2 Ni. This result, which does not vary significantly from face to face, corresponds to about 0.5 sulfur atoms per nickel atom. It has been confirmed [20] that sulfur capacity correlates with hydrogen capacity, i.e., H/Ni = 1.0 and S/Ni = 0.4-0.5 depending on the surface plane. This method was applied for Ni/AI/O 3 catalysts in the present study to determine the correlation between the saturation layer and the measured sulfur content of the catalysts. A planar surface area of 6.5 x 10 -9. nm 2 per nickel atom, based on the average of the areas for the (100), (110), and (111) planes [20-22] was used in the calculations. However, according to [20] this method cannot be applied for catalysts containing oxides that react with sulfur, e.g., Ca or Ba. In addition, the H2/Ni system is complicated due to the existence of different hydrogen states where part of the hydrogen has been suggested to be present in subsurface sites [20]. This may explain why the hydrogen uptakes were found to increase with pressure up to 10 MPa [23]. In the present study it would have been useful to perform hydrogen adsorption measurements at the same pressure as the poisoning tests to evaluate the correlation between

477

sulfur and hydrogen adsorption, if any. However, such a high-pressure experimental apparatus was not available. Chemisorption uptakes of H 2 at 298-303 K for alumina supported nickel catalysts were measured. The corresponding nickel surface areas were calculated and subsequently sulfur contents at saturation were determined. The average values were of the same magnitude as the amounts of sulfur that were not desorbed from the catalysts during the TPH treatments. Hence, it may be concluded that the saturation layer of sulfur remains on the catalyst even after regeneration in hydrogen atmosphere. Evidence of the formation of multilayer or subsurface sulfides under conditions where bulk sulfides are not expected has been reported [24-28]. For both gold [26] and platinum [27] it was concluded that in aqueous acid medium, sulfur species formed from hydrogen sulfide, which are likely to be composed mainly of polysulphide and/or elemental sulfur layers on the metal sulfide. Buckley et al. [26] concluded that the multilayer sulfur deposit does not have the properties of bulk elemental sulfur, suggesting substantial interaction occurs between the multilayer sulfur and the underlying metal sulfide and substrate metal. Two models account for the characteristics of the deposit. The sulfur could grow as an extension of the initial metal sulfide lattice with the possible movement of some metal atoms from the substrate into the sulfur layer. Alternatively, the sulfur layer could develop as a conglomerate of chains, from the underlying metal sulfide by the formation of polysulphide bonds with surface sulfur atoms. In the present study the observed high sulfur content of the catalyst beds at < 900~ especially under pressurized conditions, can most probably be attributed to the formation of polysulfides on the adsorbed metal sulfide. These sulfur species were formed quickly on the catalyst surfaces depending on the temperature and pressure applied. In addition, these sulfur species were desorbed at > 400~ which is much lower than the decomposition temperature of bulk nickel sulfide. When the temperature was raised to > 900~ in poisoning tests the formation of the multilayer polysulfides was significantly decreased due to severe reaction conditions, and consequently the content of sulfur on the catalyst approached the monolayer coverage. This again resulted in increased catalytic activity. At high pressure the formation of polysulfides is enhanced probably due to higher rates of surface diffusion of sulfur on catalyst surfaces. High pressure is known to increase the mobility of the adsorbed species on the catalyst surfaces [29]. The possible enhanced sulfur allotropes formation under pressurized conditions could also have affected the sulfur adsorption at the inlet of the catalyst bed. ACKNOWLEDGMENTS This work was funded by VTT Research Programme on Chemical Reaction Mechanisms through the project "Fundamentals of Catalytic Gas Cleaning" and by the EU-JOULE project "Catalytic Upgrading of gas from Biofuels and Implementation of Electricity Production" (JOR 3-CT95-0053). The financial support of the Finnish industry is also gratefully acknowledged.

478 REFERENCES .

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4.

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o1997 Elsevier Science B.V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

479

E f f e c t s o f H 2S on B i f u n c t i o n a l Catalysts c. Flego 1, L. Galasso 1, S. Vidotto l, G. Faraci 2 l Eniricerche S.p.A. - Via Maritano, 26 - S. Donato Milanese (MI) - 1-20097 Italy 2Euron S.p.A. - Via Maritano, 26 - S. Donato Milanese (MI) - 1-20097 Italy Poisoning of noble metal catalysts by sulfur-containing organic compounds is a wellknown phenomenon. These compounds are frequently present in hydrotreating feedstocks and cause a selective decrease in the metal hydrogenation function of noblemetal-based hydrotreating catalysts. The effects of H2S, considered as a model sulfiding compound, were studied on a supported platinum catalyst. Low coverages of H2S are found to cause a selective poisoning of noble metal particles, which with increasing coverage becomes unable to chemisorb H2 at room temperature. At higher coverages H2S spills over to the support, modifying its acidic properties. Eventually new weak acid sites are formed. 1. INTRODUCTION Organic compounds containing heteroatoms (S, N) are present in natural hydrocarbon sources and industrial feedstocks. The adsorption of sulfur, organic sulfides and sulfurcontaining derivatives is therefore a common and serious cause of deactivation of bifunctional catalysts. Poisoning of metallic catalysts by sulfur has been extensively studied and involves a number of consecutive steps including reversible molecular adsorption of H2S, its dissociation, reorientation or reconstruction of the metal surface, formation of a layer of polysulfides, and a decrease in the metal capacity for hydrogen chemisorption [1-6]. H2S adsorbs strongly at high coverages of the metallic surface (irreversibly in the range of S/Pt = 0.4-1), without any sulfur dissolution into the bulk [7]; the strong covalent character of the S-Pt bond leads to a long lifetime of the adsorbed species [8]. The extent of the metallic functionality decreases with increasing S content. In spite of the large number of papers devoted to sulfur poisoning, the effects of sulfur on both metallic and acidic functions of a catalyst were not previously investigated at high sulfur impurity concentrations. 2. EXPERIMENTAL 2.1. Catalyst. The catalyst studied in this work was a platinum (0.205% wt) supported on an amorphous mesoporous silica-alumina, MSA (SIO2/A1203=100) extruded with 50% wt alumina. The synthesis of the catalyst and the metal deposition procedure were described in detail [9]. A used catalyst, prepared as the former and containing 0.186% wt of Pt, was studied after the hydroisomerization of an n-paraffin feed containing 10 ppm S. 2.2. Reactants. All gases (SIAD) were high purity grade (99.9995%) and flowed through an oxygen converter (Supelpure-O Trap, SUPELCO), which removed the oxygen to less than 5 ppm. Pyridine (MERCK) was reagent grade (95%), distilled and purified by cryo-treatments before use. 2.3. Sulfiding procedure. A fresh batch of catalyst was heated from 20 to 300~ and kept at 300~ for 10 m in N2 flow (around 15 ml/min); then a mixture 10% H2S/H2 was introduced, at varying durations (from 0.5 to 120 minutes) and flow rates (from 3 to 135 ml/min) in order to modify the amount of H2S contacting the catalyst. The catalyst was then cleaned at 300~ for 30 min and cooled to 20~ in N2 flow.

480 2.4. H2 chemisorption. Metal dispersion was determined by H2 chemisorption performed with a pulse flow method (PulseChemisorb 2700, MICROMERITICS). The samples (0.3-1.0 g) were placed in a Pyrex reactor and heated from 20 to 300~ (heating rate of 10~ in N2 flow (15 ml/min), treated at 300~ for 2 h in H2 flow (35 ml/min), kept under a stream of N2 for 1 hr to clean the surface and eventually cooled to 20~ in the same atmosphere. The chemisorption experiments were performed at 20+/I ~ Successive pulses of 86 ml of H2 were sent to the catalyst in a constant stream of N2 (15 ml/min); the time interval between successive pulses was 90 s. The total amount of adsorbed hydrogen was calculated from the difference between the saturation peak area and the area of the peak before saturation. From this amount metal dispersion parameters were calculated [ 10]: (1) the percent of platinum present on the surface with respect to the total amount in the catalyst (Pt*/Pt, %), (2) the catalyst surface covered by metal particles (Pt area, m 2 Pt/g cat) and (3) the average diameter of the Pt particles on the catalyst surface, using a spherical model for the aggregates (d, A). 2.5. TEM (Transmission Electron Microscopy) analysis. This analysis was done on a Philips 420T microscope (120kV, maximum resolution 5A) equipped with an EDAX PV9900 EDS. The catalysts were ground to a powder, embedded in epoxy resin and then microtomed with a diamond knife to obtain sections about 300A thick. Images were taken at 100 kV. Diameters of about 100 isometric-shaped Pt crystallites were measured for each sample. 2.6. FT-IR spectroscopy. Self-supporting wafers (4-8 mg/cm 2) were analyzed by FT-IR (Fourier Transform Infra-Red) spectroscopy (mod. 2000, PERKIN-ELMER). All treatments were performed in situ in a pyrex cell with KBr windows and spectra were registered at 20~ The following experiments were performed: (i) acidity determination by pyridine adsorption-desorption method [11]. After evacuation (300~ lh, 10 -3 mbar) the catalyst was contacted by 10 mbar pyridine vapor at 200~ for lh; desorptions were performed for lh at 200 and 300~ in dynamic vacuum; (ii) hydroxyl evolution after different "in situ" treatments (reduction, sulfidation, H2 adsorption). The wafer was heated at 300~ in a dynamic vacuum, contacted by a known amount of H2S/H2 mixture for 1 hr at the same temperature, reduced for 2 h at 300~ in H2 flow and cleaned in N2 flow at the same temperature. After this pretreatment, H2 chemisorption was performed at different temperatures (20-100~ followed by pyridine adsorption-desorption at 200~ 2.7. TG-DSC (ThermoGravimetry-Differential Scanning Calorimetry) analysis. Thermogravimetric-calorimetric analysis was performed in flowing air (40 ml/min) from 20 to 800~ at a heating rate of 10~ by a TG-DSC 111 apparatus (SETARAM). 3. RESULTS

3.1. Used catalyst Used catalysts were tested at 300~ for about 1000 h in the hydroisomerization of nparaffin with feed containing traces of sulfur compounds (10 ppm S). The reaction was stopped and the catalyst was divided in two parts; one was directly characterized, the other was cleaned as follows. The catalyst was washed with n-heptane (reflux at 100~ and dried by a mild calcination at 300~ for 16 hours in flowing air, in order to remove the physisorbed organic compounds without changing the physico-chemical characteristics of the catalyst. Thermogravimetric analysis of the fresh, used and washed batches of catalyst produced the results reported in Table 1. The initial weight loss was due to the atmospheric moisture, the other two TG signals were caused by release or combustion of organic compounds and dehydroxylation of some OH groups on the surface of the catalyst. The used catalyst contained a large amount (about 32% wt) of organic compounds (reactants and products of the hydroisomerization reaction). These compounds were almost completely removed by n-heptane, as no more organic compounds were detected (no C was found by chemical analysis) and the

481 thermogravimetric and calorimetric data were comparable with those of the flesh catalyst. Table 1. Weight loss (WL~ %) and calorimetric heat (kJ/g) evolution after air treatment. Sample fresh used washed total WL (%) 14.53 46.57 14.67 (20-200~ WL (%) 9.77 4.98 9.42 (230-340~ WL (%) 2.29 13.72 1.98 (340-455~ WL (%) 18.54 (455-610~ WL (%) 2.57 7.20 3.52 endo Heat (kJ/g) 3.26 6.58 3.67 exo Heat (kJ/g) 6.56 15.51 8.36 Table 2. Metal dispersion parameters and some Sample Pt*/Pt Pt area (%) (m 2 Pt/g cat) flesh 100 0.65 used 4 0.03 washed 44 0.28

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Figure 1. TEM micrographs of fresh (A) and used (B) catalyst and metal particle distribution Other physico-chemical characteristics measured for these samples included platinum dispersion parameters (see Table 2 and Figure 1), acidity distribution and hydroxyl concentration (see Table 3). The H2 chemisorption capacity of the washed sample decreased slightly in comparison with the fresh one, is consistent with poisoning or sintering. However, TEM analysis revealed no appreciable difference in the average crystallite dimension and crystallite size distribution of fresh and used samples. The S content (by chemical analysis) was very different in fresh and used samples, while that of the washed sample was comparable with that of the used one. The acidity distribution was obtained by pyridine adsorption-desorption (Table 3). The presence in the IR spectrum of a band at 1455 cm -l was assigned to the pyridine coordinatively bonded to Lewis acid sites, while a band at 1545 c m l was attributed to the pyridinium ion; using the known values of the extinction coefficient of the two IR signals [12] it was possible to determine the Lewis and Bronsted acid sites densities. The number of acid sites observed in the used catalyst was about half of that present in

482 the fresh one, while the washed sample showed a small decrease in the Lewis acid sites density and a large increase (more than twice) in the Bronsted acid sites density. This was confirmed by the increase in the intensity of the IR hydroxyl signals in the washed sample. In fact the IR band at 3595 cm l was attributed to vicinal SiOH or to acidic SiOHAI groups and that at 3740 cm -~ was attributed to isolated acidic SiOH groups. The presence of paraffinic organic species was confirmed by the high intensity of the IR signals in the 3000-2800 cm"range in the used catalyst. Table 3. Acid sites and IR signals distribution Sample Lewis Bronsted sites b 3740 cm- 3595 c m organic IR band sites a (mmol/g) t 1 (a.u./g) (mmol/g) (a.u./g) (a.u./g) fresh 257 14 1.393 9.140 0.382 used 135 7 0.783 6.299 1.313 washed 220 30 1.602 12.232 0.351 a from the intensity of the 1455 cm -~ IR band; b from the intensity of the 1545 cm ~ IR band 3.2. The influence of sulfiding on the fresh catalyst The total amount of H2S contacting the fresh catalyst is reported as theoretical S/Pt molar ratio, assuming 100% adsorption of the S atoms introduced as H2S/H2 mixture on surface Pt atoms of the catalyst, in spite of the fact that sulfur was also distributed throughout the support. The dispersion of Pt was good (Pt*/Pt=-82) with an average particle dimension of 12A. The metallic phase was very stable toward sintering in a reducing atmosphere up to 600~ although after 400 hours the average dimension of the crystallites was increased from 12 to 41A (Figure 2). The H2 chemisorption uptakes at increasing sulfur coverage (S/Pt=0 - 100) of the metal are shown in Figure 3. At higher S/Pt ratios the percentage of platinum particles able to interact with the probe decreased to 40% of the original value of dispersion at S/Pt=92. When the catalyst was contacted with larger quantities of H2S, the H2 chemisorption experiments showed a reverse trend: i.e. increasing the S/Pt ratio, the amount of H2 adsorbed at room temperature increased (Figure 4). No appreciable changes in the metal dispersion and crystallite size distribution between fresh and sulfided samples were observed by TEM analysis (Figure 5). This is consistent with a selective poisoning of the metallic sites, unable to chemisorb hydrogen. The behavior observed for high sulfiding coverages is not easily explained. In order to verify the influence of physisorption on these experiments, H2 adsorption uptakes were measured at different temperatures at increasing S/Pt molar ratio. A maximum in adsorption uptake was observed at 100~ for both S/Pt=0 and S/Pt=2 samples, while hydrogen adsorbed on the S/Pt=-383 sample was lower than for the other samples independent of temperature (Figure 6). The influence of the degree of sulfiding on hydroxyl groups of the support was evaluated by FT-IR spectroscopy (Table 4). Upon increasing the S/Pt ratio of the samples, a new band at around 2530 cm" (due to stretching vibration o f - S H groups) appeared and grew and the vibrations in the 1700-1555 cm" increased in intensity up to a ratio S/Pt = 139. In this IR region the bending vibration of water increased in intensity together with contributions of other groups related to the presence of H2S [13, 14]. At the same time broadening of the IR band at 3595 c m " increased, due to hydrogen bonding of H2S with the surface -OH groups. After H2 chemisorption at 21 and 100~ a shift to lower frequencies and a broadening of the IR band due to acidic -OH and to -SH groups was observed in all samples. This was consistent with the bonding of gas phase hydrogen with the hydroxyl and thionyl groups of the surface [ 15, 16].

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484 of the 1545 cm -1 one. This suggests that two acid sites of different elemental composition were present on the highly sulfided catalysts.

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'r ( ' c ) Figure 6. H2 chemisorption at different temperatures Table 4. Intensity of the IR bands at 3740 cm 1 (A), at 3595 cm 1 (B), at 2523 cm 1 (C) and in the range 1700-1555 cm -l (D) and acid site density in samples with increasing amount of HzS. S/Pt Treatment A B C D Lewis a sites Bronsted sites b (a.u./g) (a.u./g) (a.u./g) (a.u./g) (~nol/g) (btmol/g) 0 reduced 1.431 14.413 0 0.425 H2 chemi21~ 1.295 15.352 0.017 0.450 H2 chemiat 100~ 0.925 21.133 0.137 0.819 47 13 2 sulfided 1.545 10.202 0.041 0.639 H2 chemi 21~ 1.477 16.718 0.258 2.560 H2 chemiat 100~ 1.647 12.809 0.221 1.143 98 15 139 sulfided 1.717 14.976 0.114 1.259 H2 chemi 21~ 1.289 21.390 0.317 3.566 H2 chemiat 100~ 1.607 17.428 0.185 1.844 134 39 910 sulfided 1.445 15.030 0.061 0.472 H2 chemi21~ 1.198 23.488 0.137 1.340 H zchemiat 100~ 1.175 22.058 0.108 0.618 129 52 a from the intensity of the 1455 cm -I IR band; b from the intensity of the 1545 cm ~ IR band 4. DISCUSSION In the used catalyst two different kinds of deposits are present: organic compounds (around 32% wt) and adsorbed S species (10 lamol S/g). The former (due to reversible adsorption of reactants and products of the hydroisomerization reaction) are easily washed by a mild treatment, the latter remaining almost constant (7 Ixmol S/g) due to an irreversible interaction. The main changes observed in the catalyst after reaction were (1) the decrease in Pt availability to chemisorb H2, without a change in the dimensions of the metal crystallites, and (2) the formation of a larger number of Bronsted acid sites, confirmed by the higher intensity of the IR bands due to acidic and vicinal -OH groups. Irreversible sulfur poisoning causes the well-known inhibition of the metal hydrogenation function. In most reactions, acid site concentration is typically decreased by both selective poisoning of the acid sites and/or an aspecific pore blockage/surface coverage [16]. Therefore the increase of acid site density is an unusual side reaction of

485 sulfur exposure; thus the presence of S in the feed could catalyze different reactions involving both metallic and acidic function of the catalyst. The lowering of Pt dispersion in this catalyst is not caused by sintering of the metal particles during reaction because the average crystallite size is stable up to 600~ in a reducing atmosphere (similar to that of reaction). Selective treatments with the model probe molecule (H2S) were performed on the fresh catalysts, in order to follow the behavior of metal and acid sites after contacting with increasing amounts of S. The reduction of sulfur-treated samples does not affect the chemical state of sulfur [3] and was applied to clean the surface before H2 chemisorption experiments. At low H2S coverages (theoretical S/Pt molar ratio=2-100) the ability of Pt to chemisorb H2 decreases with increasing sulfur loading, while the metal distribution (both average Pt diameter and number of particles with a defmed diameter) does not change before and after sulfidation. This confirms a poisoning due to the strong bonding of S to Pt particles. The contact with larger amounts of H2S (S/Pt=-100-4500) results in a reverse behavior: i.e. the higher the S/Pt ratio, the larger the H2 amount adsorbed on the catalyst (without changes in Pt distribution). At the same time the formation of-SH and -OH groups is observed by IR spectroscopy, the bands at 2523 cm -1 and 3595 cm -l increasing in intensity with the H2S interaction. It is reported [17] that "f-A1203 itself is an effective H2S sorbent, although to a smaller extent than bifunctional catalysts, dissociating the molecule and to produce -SH groups with Bronsted character [ 18]. The formation of new acid sites with increasing H2S loading is confirmed by pyridine adsorption; i.e., the number of both Bronsted and Lewis acid sites increases at increasing H2S coverages. After H2 chemisorption a broadening and an increase in intensity of the bands in the 3700-3200 and 2700-2300 cm l IR regions are observed, due to the hydrogen bond complex between thionyl and hydroxyl groups and the probe. Metals have a promoting effect on the spillover process, though a direct reaction occurs between H2 and surface acidic groups causing a heterolitic dissociation of the molecules and eventually an irreversible adsorption [19, 20]. Spillover is affected by a large H2S content. Temperature influences the amount of H2 adsorption and the energy desorption profile during spillover [21]; at higher temperatures H2 adsorption increases because of the diffusion of atomic hydrogen onto the surface of the support, while desorption occurs at temperature up to 120~ [22]. In samples with higher metal availability the amount of H2 adsorbed at 100~ is higher than at 21~ On the contrary, in highly sulfided samples less surface Pt is available for facilitating spillover, while new weak -SH groups are able to bond H2 in a similar way at both temperatures. 5. CONCLUSION Treatment with low amounts of H2S decreases the availability of metal to adsorb H 2 in a bifunctional catalyst e.g., that used in hydroisomerization. Upon contacting with higher amounts of H2S the metal availability is still lowered but the sulfiding agent spills over to the support, creating new weak acid sites. Decreased availability of the Pt surface and an increase in the acid site density result from the presence of sulfur in the feed. The increase in the ratio of acid sites (detected by pyridine at 200~ to Pt metal (present on the surface and still able to interact with H2) could cause a change in the catalyst performance, e.g. increasing the cracking activity with respect to the isomerization. ACKNOWLEDGMENT The authors gratefully acknowledge Mr. B. Stocchi for helpful TEM contribution.

486 REFERENCES

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10. 11. 12. 13. 14. 15.

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9

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Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

487

P e r f o r m a n c e o f Ni/A1203 Pellets P o i s o n e d b y T h i o p h e n e D. Rusic a and S. Zrncevic b aFaculty of Technology, University of Split, Croatia bFaculty of Chemical Engineering and Technology, University of Zagreb, Croatia

A quantitative assessment of the intemal mass transport effect on a catalytic system involving chemical poisoning is based on solving a set of coupled partial differential equations describing the reactant and product concentrations and activity profile within the catalyst particle. It is convenient to describe the results by a pellet effectiveness factor for the main reaction which depends upon time, intraparticle position, diffusion and reaction parameters. Based on kinetics of gas phase benzene hydrogenation to cyclohexane, catalyst poisoning and pore diffusion time-dependent effectiveness behavior of a Ni/AI203 catalyst pellet were simulated. Deactivation data were collected in an isothermal fixed-bed reactor at hydrogen partial pressure of 99.82 kPa, benzene partial pressure of 7.55 kPa, thiophene partial pressure of 0.032 kPa and at reaction temperatures ranging from 403 to 473 K. The time-dependent effectiveness factor obtained as a function of Thiele modulus was compared with experimental observations for a commercial catalyst pellet. A degree of correlation between theoretical prediction and experimental results was generally satisfactory when benzene hydrogenation was represented by a power-law kinetic equation. By contrast, the fit obtained with the Hougen-Watson kinetic model for benzene hydrogenation is poor due to changes in adsorption constants with extent of poisoning.

1. INTRODUCTION In industrial heterogeneous catalytic processes catalysts inevitable lose activity over a period of time. Quantitative study of the activity-time relation is important in determining the optimum reactor design and operation. The first step in treating the problem is to analyze the behavior of a single catalyst pellet. The most important industrial catalysts are porous pellets involving diffusional transport of reactants to sites within particles before being converted. Because transfer resistances affect the main reaction rate it can be assumed that the rate of deactivation also depends on the extent of mass transfer resistance. The effect of intraparticle diffusion on a catalytic process with accompanying catalyst deactivation has been presented in a number of papers. Following general analyses of Wheeler [ 1], one of the earliest treatments of poisoning was given by Masamune and Smith [2]. They solved the mass conservation equations numerically for an isothermal catalyst pellet with firstorder reaction and deactivation. The analysis of Masamune and Smith was extended to a complex first-order reaction system by Murakami et al. [3] and to LangmuirHinshelwood kinetics by Chu [4], who introduced the concept of a relative effectiveness factor. Hegedus [5] studied the general case of combined external and internal mass transfer resistances and the effect of pellet geometry. The effect of impurity poisoning on steady state and transient behaviors of a diffusionally influenced exothermic catalytic

488 reaction was studied by Lee et al. [6]. Krishnaswamy and Kittrel [7] developed analytical procedures for nth order irreversible reactions and defined a deactivation effectiveness factor using the concept of activity as defined by Levenspiel [8]. Grzesik et al. [9] applied a time on stream theory to solve a fixed-bed reactor model considering both catalyst decay and intraparticle diffusion resistance. More recently Chocron et al. [10] defmed a general equation for the effectiveness factor which includes catalyst decay. The present research is an experimental and modeling study of the effect of impurity poisoning on the behavior of a diffusion-affected catalytic reaction. Benzene hydrogenation to cyclohexane over a Ni/A1203 catalyst poisoned by thiophene was used as a model reaction.

2. EXPERIMENTAL The reactor system and general experimental procedures were similar to those described by Zmcevic et al. [11] and are discussed in detail in our former paper [12]. Deactivation measurements were carried out in an isothermal fixed-bed reactor with a hydrogen partial pressure of 99.82 kPa, benzene partial pressure of 7.55 kPa, thiophene partial pressure of 0.032 kPa and reaction temperatures from 403 to 473 K. The size of commercial cylindrical catalyst pellets supplied by BASF was 5 x 5 mm (21% Ni on alumina). The catalyst was activated by reduction with hydrogen at 743 K for 10 h.

3. RESULTS AND DISCUSSION The main reaction, i.e. benzene hydrogenation occurring inside porous Ni-catalyst pellets, is accompanied by a poisoning reaction in which thiophene in the feed stream reacts irreversibly with the catalytic active sites. An analysis was made assuming isothermal behavior so that the heat balance within the pellet could be neglected [13]. Since the rate of catalyst poisoning is sufficiently small and diffusion of the reactant sufficiently fast, the steady-state continuity equation is a good approximation at all times. This approach has been widely used in the problem of poisoning [ 14-16]. Under these conditions the following mass balance describes the behavior of a catalyst pellet

d 2C a 1 dC a -~ ~ - r DB dr 2 r dr

B (t = 0) a(t)

with boundary conditions: X=I X= 1 where YB=

YB=I dYB/dX CB/CBo

t=O t>0

and X = r/R

(1)

489 The kinetic expression for the benzene hydrogenation unaffected by pore diffusion my be represented by a power-law kinetic model [12] in

n

rB = k C H C B

(2)

or Hougen-Watson kinetic model [ 17] k C B K s 4CH K H rB

(3)

~+C B KB+dCHK"~

with parameter values given in Table 1.

Table 1. Kinetic parameters for benzene hydrogenation. Equation 2 T, K 105k, n m mol kg-ls -1 403 5.63 0.24 0.35 423 6.66 0.35 0.32 448 7.58 0.41 0.30 473 11.10 0.39 0.29

Equation 3 102k, 103KB, mol kgls "1 Pa -1 0.30 3.46 0.57 3.52 0.78 4.53 1.58 3.81

106KH, Pa ~ 1.25 1.35 1.58 1.29

The catalyst activity a(t) in equation (1) is related to the activity of unpoisoned catalyst and is given by the kinetic equation [12] da dt - kh ad

(4)

where ka' = k d CT. The kinetic parameters for catalyst deactivation by thiophene determined in separate experiments are given in Table 2. Table 2. Kinetic parameters for catalyst deactivation. Tr K 105 kay S-1 d 403 6.22 1.34 423 5.41 1.36 448 5.52 1.50 473 4.55 1.64

Values of the kinetic and adsorption parameters in the above tables were extracted from a standard nonlinear regression technique employed to select the best fits of the experimental data. The regression methods used were based on a minimization of the squares of the difference between experimentally observed and predicted results.

490 A differential equation simulating the deactivation process within the catalyst pellet was solved numerically using the orthogonal collocation method [18]. Parameters employed in the simulation are given in Tables 1 and 2. The effective diffusivity, DB, of benzene within the particle was measured to be 5 x 10-3 cm 2 s1 via an helium-argon counter-diffusion experiment. It was assumed that the effective diffusivity was uniform throughout the catalyst pellet and remained unaffected by poisoning. Typical results of simulation are shown in Fig. 1, which compares the change of benzene profiles inside the catalyst pellet with time due to thiophene poisoning at 403 K.

~x-7..~-&~--

__

_

-

-

!\\

0.8

.

.

" ----.

0.6

0.4_ 1-

o.2-

0

2,2'- lOmin

~

_ 3,3'- 50min - 4,4' - 100 min 0.0

'

0.0

,

t.

0.2

~

~

~ ",2,,

. ' '.t ".'

0.4

~

0.6

'

i

0.8

'

1.0

r/R Figure 1. Variation of benzene concentration within the catalyst particle during poisoning (--- empirical model, eq. 2; ~mechanistic model, eq. 3). The concentration profile of benzene is, as expected, maximal at the surface of the catalyst pellet. With increasing time of operation benzene concentration becomes more uniform through the pellet giving a nearly flat concentration profile when the time on stream is about 100 min. The profiles depart more slowly from the initial profile (at t-0) when the empirical kinetic model for benzene hydrogenation is used. The overall effectiveness factor of a catalyst pellet can be characterized by the ratio of the observed reaction rate to the rate in the absence of poisoning or external mass transfer resistance. It is expressed in the form of a power-law kinetic model for benzene hydrogenation as

]'~E -" (% +

1)a(t)

1 N=3

W 4 L~-~-4 ~_~ w i y ~ n

J (X i) + 1

and for Hougen-Watson kinetic model as

(5)

491

rim = (;L + 1) a(t) (K2 + KsCs~ W4

Wi YB(xi) -I- 1 i=l (K2 + Kso YB(Xi)) 2

(6)

where K 2 = k 4Cn K . K B

1.00 ~.~

1

0.80-

9

2 m 3 A 4 # 5 v 6 ~t

0.60-

0.40-

5 20 40 60 80

mln mln m|n mln

min 1O0 rnln

Id

t~

0.20-

0.10-

1

2

'"

t 3

i 4

i 5

I" 6

i 7

i i 8910

i

Figure 2. Effectiveness factor versus Thiele modulus at various times of poisoning (powerlaw kinetic model for benzene hydrogenation).

Figures 2 and 3 show the calculated time dependent effectiveness factors TIE and rim respectively, as a function of Thiele modulus ~. The figures also show points obtained by experimental observation. The experimental values for the effectiveness factor for benzene hydrogenation were obtained by dividing the experimental reaction rate corresponding to a given particle size, with a true hydrogenation rate calculated from Equations (2) or (3). In general, there is a satisfactory degree of correlation between theoretical prediction and the experimental results when the kinetics of benzene hydrogenation are represented by the power law kinetic model (see Fig. 2). The results shown in Fig. 3 indicate that the Hougen-Watson kinetic model for benzene hydrogenation provides a poor fit of the data. Apparently, the experimental effectiveness factors are about 20% lower than the computed values. The effect of mass

492 transport on the rate of benzene hydrogenation is thus somewhat higher than predicted by the model. 1.00 ....................................... .

2 = 0.60

2

3 4 5 6

0.40

g..

20 rnin

,t

40 60 80 100

9

v Nt

rain

rain rain rain

0.20

0.10-

1

............

I

2

. . . .

i

3

.......

i

4

--i

5

i--i

6

-r

i ....i

78910

'

-

Figure 3. Effectiveness factor versus Thiele modulus at various times of poisoning (HougenWatson kinetic model for benzene hydrogenation). The origin of such behavior may derive from experimental uncertainty in determining kinetic parameters due to low conversions at high poison levels. Another reason might be the effects of poisoning on the individual rate constants for the adsorption and surface reaction step in a heterogeneous catalytic reaction. Smith et al. [19, 20] studied the effects of catalyst poisoning on adsorption and surface reaction rates for the hydrogenation of t~-methyl styrene on Pd/A1203. According to them poisoning did not affect the rate constant per site for surface reaction, i.e. the activity of the remaining active site. That was confirmed by the constant activation energy regardless of the extent of poisoning, the same result shown in our previous paper [12]. This suggests that poisoning is due to geometric rather than electronic effects. The adsorption rate constant and adsorption equilibrium constant per unit mass of a catalyst decreased with poisoning and that might explain the poor fit of the results obtained with the use of Hougen-Watson kinetic model. 4. CONCLUSIONS The problem of decreased catalyst activity due to irreversible thiophene adsorption was solved numerically using an orthogonal collocation method with three internal points. The numerical results were compared with experimental data obtained by

493 measuring concentration changes due to thiophene poisoning in benzene hydrogenation over Ni/A1203 catalyst pellets. Benzene hydrogenation was represented by power-law and Hougen-Watson kinetic models. A satisfactory degree of correlation between theoretical prediction and experimental results is found when benzene hydrogenation is represented by the power-law kinetic model. The data fit obtained when benzene hydrogenation is represented by Hougen-Watson kinetic model is poor; the results are complicated by changes in equilibrium adsorption constants with extent of poisoning. 5. NOTATION a

CB CH CT

d DB k kd' KB KH m n

r R t

T W

n

catalyst activity benzene concentration, mol dm "3 hydrogen concentration, mol dm "3 thiophene concentration, mol dm -3 deactivation order effective diffusion coefficient for benzene, cm 2 sl reaction rate constant, mol kg 1 s1 deactivation rate constant, s-i adsorption constant for benzene, Pa1 adsorption constant for hydrogen, P a -1 rate order for hydrogen rate order for benzene radial coordinate, cm radius of catalyst pellet, cm time, min temperature, K weight value geometric factor Thiele modulus effectiveness factor

6.ACKNOWLEDGMENT The financial support of this work by the Croatian Ministry of Science and Technology is kindly acknowledged. REFERENCES

1. 2.

A. Wheeler, Catalysis, Vol.2, Reinhold, 1955. S. Masamune and J.M. Smith, Performance of fouled catalyst pellets, A.I. Ch.E.J.,

3.

Y. Murakami, T. Kobayashi, T. Hattori and M. Masuda, Effect of intraparticle diffusion on catalyst fouling, Ind.Eng.Chem.Fundam., 7(1968)599. C. Chu, Effect of adsorption on the fouling of catalyst pellets, Ind.Eng.Chem.

12(1966)384. 4.

Fundam., 7(1968)509.

494 5.

L.L. Hegedus, On the poisoning of porous catalysts by an impurity in the feed,

lnd.Eng. Chem. Fundam., 13(1974)190. 6.

J.W. Lee, J.B. Butt and D.M. Downing, Kinetic, transport, and deactivation rate interactions on steady state and transient response in heterogeneous catalysis,

A.1.Ch.E.J. 24(1978)212. 7. 8. 9.

S. Krishnaswamy and J.R. Kittrell, Diffusional influences on deactivation rates,

A.I. Ch.E.J., 27(1981)120. O. Levenspiel, Chemical Reactor Engineering, 2nd Ed., Wiley, New York, 1972. M. Grzesik, J. Skrzypek and B. Wojciechowski, Modelling of intraparticle diffusion affected by the time-on-stream catalyst decay, Chem.Eng. Sci.,

47(1992)2805. 10. 11. 12.

M. Chocron, M.C. Raffo Calderon, N. Amadeo and M. Laborde, Effect on intraparticle diffusion on catalyst decay, Chem.Eng.Sci. 51(1996)683. S. Zmcevic, Z. Gomzi and E. Kotur, Thiophene poisoning of Ni-SiO2-A1203 in benzene hydrogenation. Deactivation Kinetics, Ind.Eng. Chem.Res. 29(1990) 774. S. Zmcevic and Z. Gomzi, Catalyst poisoning in the benzene hydrogenation,

Chem.Eng. Sci., 38(1983)1351. 13. 14.

S. Zmcevic, The intraparticle deactivation of Ni-SiO2-A1203 by thiophene in benzene hydrogenation, Chem.Eng.Sci., 39(1984)1245. C.E. Megiris and J.B. Butt, Effect of poisoning on the dinamic of fixed bed reactors. I. Isothermal in a cyclic policy of operation, Ind.Eng.Chem.Res.,

29(1990)1065. 15.

R. Christoph and M. Baems, Modelling of an adiabatic, catalytic fixed bed reactor with catalyst deactivation and pore-diffusional effects for the methanation of CO,

Ber.Bunsnges.Phys. Chem., 90(1986)981. 16.

A. Baiker, D. Epple and A. Wokaun, Behaviour of a pilot plant fixed bed reactor during catalyst deactivation, Chem.Eng,Sci., 41(1986) 779. 17. S. Zmcevic and D. Rusic, Verification of the kinetic model for benzene hydrogenation by poisoning experiment, Chem.Eng.Sci., 43(1988) 763. 18. J. Vilandsen and M.L. Micheelsen, Solution of differential equation models by polynomial approximation, Prentice-Hall, Engelwood Cliffs, N.Y., 1978. 19. L.A. Arrua, B.J. McCoy and J.M. Smith, Effect of catalyst poisoning on adsorption and surface reaction rates in liquid-phase hydrogenation,

lnd.Eng. Chem.Res., 29(1990)1050. 20. S.Y. Chen, B.J. McCoy and J.M. Smith, Dynamic studies of catalyst poisoning: Effect on adsorption and surface reaction rates for hydrogenation of o~-methyl styrene by Pd/A1203, A.I. Ch.E.J., 32(1986)2056.

9

Elsevier Science B. V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

495

Study o f the Simultaneous Deactivation b y Coke and Sulfur o f N a p h t h a R e f o r m i n g Catalysts using a Bifunctional Test Reaction A. Borgna, T.F. Garetto and C.R. Apesteguia Instituto de Investigaciones en Cat~ilisis y Petroquimica-INCAPE-, (UNL-CONICET) Santiago del Estero 2654, (3000) Santa Fe, Argentina 1. INTRODUCTION Under industrial conditions, sulfur poisoning of Pt-based naphtha reforming catalysts takes place in the presence of a simultaneous deactivation by coke formation. Very few papers have been published on the modeling of simultaneous deactivation processes [ 1,2], probably because the occurrence of two different deactivation mechanisms complicates both the analysis of experimental results and the determination of the intrinsic deactivation kinetics. In previous work [1,3], we have undertaken fundamental studies regarding sulfur poisoning of naphtha reforming catalysts in presence of coking. Specifically, we have characterized the (C + S) deactivation of Pt-based catalysts by employing two monofunctional metallic reactions: benzene hydrogenation, a well known structure insensitive reaction on platinum, and cyclopentane hydrogenolysis, which is a structure sensitive reaction over the same metal. Nevertheless, the main reactions in naphtha reforming proceed via a bifunctional mechanism involving both metallic and acid sites. In this work, we have extended the above studies to employ a bifunctional reaction, the reforming of n-hexane, for studying the simultaneous (C + S) deactivation of Pt-Re(Ge,Sn)/A1203 naphtha reforming catalysts. 2. EXPERIMENTAL A monometallic Pt catalyst was made by impregnation at 303 K of a high-purity y-AI203 powder (Cyanamid Ketjen CK300) with an aqueous solution of chloroplatinic acid and HC1. After impregnation, the sample was dried for 12 h at 393 K and heated in an air stream to 773 K. Next the chlorine content was regulated using a gaseous mixture of HC1, water and air. Finally, the sample was purged with N2 and reduced in flowing H2 for 4 h at 773 K. A bimetallic Pt-Sn catalyst was prepared by coimpregnation at 303 K of alumina CK300 with an aqueous solution of CI2Sn, H2PtCI6 and HC1. After drying overnight at 393 K the sample was calcined in air for 5 h at 773 K, purged in N2 and reduced in flowing H2 at the same temperature for 4 h. A bimetallic Pt-Ge catalyst was prepared by a similar method. The support was impregnated with a solution containing H2PtCI, GeCI4, and HCI. After successively drying overnight at 393 K and calcining in air for 5 h at 773 K the sample was reduced in H2 for 4 h at 773 K. Pt-Re/A1203 was obtained from commercial sources (Cyanamid Ketjen CK433).

496 Table l Catalytic tests without sulfur poisoning. Initial activities and selectivities. T = 693 K, P = 1 atm, H2/n-hexane = 8.8, WHSV = 5 h"~

Catalyst

SQ(%)~

SO(%)c,

Pt 7.7 3.5 Pt-Re 8.7 3.8 Pt-G-e 3.1 0.4 Pt-Sn 3.1 0.4 (a): ro= [mol n-hexane/h .g cat.]

S~_c ,o (%) 56.9 58.9 86.2 85.2

SMcpO(%)

SB.o (%)

16.2 15.8 6.9 7.4

roxl02(.)

15.6 12.7 3.4 3.8

2.2 2.1 1.8 1.7

Catalytic testing was carried out at 1 atm in a fixed-bed flow reactor. The reaction was performed at T = 693 K, P ~ = 0.114 atm and WHSV = 5 h"~. Sulfur poisoning experiments were performed by doping the feed with thiophene in concentrations between 0 and 25 ppm of S. The gaseous reaction effluent was analyzed by gas chromatography with a flame ionization detector The following nomenclature was used for analyzing the reaction products: gases (G), paraffins with 5 carbon atoms (C5), n-hexane isomers (i-C6), methylcyclopentane (MCP), and benzene (Bz). Selectivities (S) were calculated as: Si(%) = [conversion to product i / n-hexane conversion] x 100.

3. RESULTS

3.1. Catalytic tests without sulfur poisoning The initial catalyst activity and selectivity were measured by using thiophene-free n-hexane as reactant and by extrapolating the coke deactivation curves to zero. Results are given in Table 1. In all the cases, i-C6 isomers were the main products of the reaction. Pt/Al203 and PtRe/AI203 were more active than Pt-Ge(Sn)/Al203. The Pt/Al203 sample displayed the highest Catalyst: Pt-Re

] .01

Pt

] .0| '

0.9 .. 0.8 tr

~

Pt-Ge ~

0.8

~

0.7

"~0.6

0.6

0.4

"

ar (25.2 ppm S) 0.2!

I

I

I

I

I

30

60

90

120

150

Time (min) Fig. 1" Coke (aC), sulfur (a s) and simultaneous (C + S) deactivation (aT).

0.5

0%

;'

10

1;

20

2'5

IS], ppm

Fig. 2: Catalyst thiotolerance as a function of feed sulfur concentration.

497 selectivity to benzene whereas bimetallic Pt-Re was more active for the formation of gases. o o Pt-Ge(Sn)/AI203 samples exhibited lower S~, SMCP, and SBz values as compared to Pt or PtRe catalysts and, as a consequence, were more selective for the formation of i-C6 isomers. Deactivation by coking was studied from the curves representing the activity decay as a function of time (Fig. 1). The activity for n-hexane reforming without sulfur poisoning is defined as ac = rtc / roc. where roc and rtc the n-hexane conversion rates without sulfur poisoning at zero time and at time t, respectively. Similarly, the activity for the formation of the product i is defined as aic = ritC/rioc. The residual activity values at 150 min in the activity decay curves were taken as deactivation parameters to establish the relative coking sensitivities of the catalysts investigated and are designated here as a~c'ss. Results are shown in Table 2. From the values of the last column it is inferred that the n-hexane conversion was less inhibited by coking on Pt-Sn(Ge)/AI203 than on Pt/AI203 or Pt-Re/AI203. Table 2 Catalytic tests without sulfur poisoning. Deactivation by coking and residual activities c.ss aC' SS _c.ss _c.ss Catalyst a c'ss ac, i-c, aMcp aBz Pt Pt-Re Pt-Ge Pt-Sn

0.70 0.60 0.76 0.92

0.62 0.58 0.73 0.91

0.87 0.90 0.87 0.90

0.63 0.71 0.89 0.95

0.51 0.54 0.67 0.82

aC'SS 0.76 0.77 0.84 0.91

3.2. Catalytic testing with sulfur poisoning Sulfur poisoning curves were obtained by doping the feed with thiophene. Thiotolerance (TT), which is defined as the residual activity value in the poisoning curves (TT = a s'ss), was used for determining the relative sensitivity of the catalysts to sulfur poisoning. The catalyst thiotolerance for the n-hexane reforming reaction was calculated as TT = 1 + a T'ss - a c'ss by assuming that the overall deactivation was a simple sum of each individual deactivation rate (Fig. 1). a c'ss and a T'ss are the residual activities in the activity decay caused by coking alone and by (C + S) deactivation, respectively. Similarly, the catalyst thiotolerance for the formation of product i was calculated as TTi = 1 + a~~s - ac'ss . Fig. 2 shows the catalyst thiotolerance for the n-hexane reforming reaction as a fimction of the sulfur concentration in the feed whereas Table 3 summarizes the TTi values obtained when the reactant was doped with 15.5 ppm S. From Fig. 2 it is inferred that the catalyst thiotolerance for n-hexane conversion follows the older Pt -_-Pt-Ge > Pt-Re > Pt-Sn. On the other hand, Table 3 shows that the TTc,, TTBz and Table 3 Catalytic test with ~

poisoning. Catalyst thiotolerance for 15.5 ppm S

Catalyst

TT~

TTc,

T~_c,

TTM~

TTB~

Pt Pt-Re Pt-Ge Pt-Sn

0.74 0.63 0.94 0.53

0.52 0.47 0.71 0.48

1.13 0.64 0.90 0.62

0.83 0.67 0.98 0.86

0.77 0.56 0.90 0.53

498 TT~ values decrease in the order Pt-Ge > Pt > Pt-ge > Pt-Sn whereas in the case of T1~_c, the trend is Pt > Pt-Ge > Pt-Re -- Pt-Sn. 4. DISCUSSION

Fig. 3 shows a simplified scheme of the bifunctional paths proposed by several authors [4,5] for interpreting the n-hexane reforming reaction. The reaction network includes: (a) the isomerization and the dehydrocyclization of n-C6 to i-Ca and MCP, respectively, through a bifunctional mechanism involving metal- and acid-catalyzed steps; Cl-Cs Paraffins / l "~,Hydrocrackingand r (b) the hydrogenolysis and \ Hydrogenolysis hydrocracking of n-C6 and i-C6 i-hexane n-hexane n,i-hexane to C I-C5 paraffins, (c) the aromatization of n-C6 to Bz c through the ring expansion i-hexene 0 ' n-hexene ~ r 0 -Skeletal Ring k._J m reaction of methylcyclopentene tIsom,ization Closure T/ O .0 Ring (MCPe); (d) dehydrogenation of I J, MCPe to methylcyclopentadiene (MCPde) producing coke t precursors. This reaction network may Coke @ Precursors be employed for interpreting the P values of initial activities and Reactions on acid sites Fig. 3" Simplified reaction network for n-hexane reforming selectivities obtained here for different naphtha reforming catalysts (Table 1). Under our operating conditions, i-C6 isomers were the main products of nhexane reforming over all the catalysts. However, the initial selectivity to i-C6 was substantially = 58 %). The higher higher on Pt-Ge(Sn) ( S~i-c, = 86 %) than on Pt or Pt-Re catalysts (Si_c, o

]1

Tl

/

TI

Tl

11

selectivity to i-C6 exhibited by the Pt-Ge(Sn) catalysts was actually caused by a diminution in the formation rate of Bz and C1-C5 paraffins. This result is in line with previous studies which showed that the hydrogenation and dehydrogenation activity of Pt is severely depressed in PtGe(Sn) catalysts, particularly at low reaction pressures [6]. Regarding the Pt and Pt-Re catalysts, Pt is better hydrogenation metal than Re and this explains the higher selectivity to Bz exhibited by the Pt/A1203 sample as compared to Pt-Re/A1203. On the other hand, formation of light hydrocarbons was higher on the Pt-Re catalyst because Re is known to promote hydrogenolysis reactions. Pt/A1203 and Pt-Re/AI203 were both more affected by coke deactivation than PtGe(Sn)/A1203 (Table 2, cohmm 7). Because the Pt-Ge(Sn) samples displayed a low dehydrogenation activity, both the transformation of MCPe to MCPde and the further MCPde dehydrogenation to form coke precursors are inhibited over these catalysts. Pt-Sn/A1203 was the most resistant catalyst to coke deactivation. Beltramini and Trimm [7] have also found the reduced coke formation on Pt-Sn catalysts and postulated that the ability to decrease coke formation may result from enhanced gasification of coke precursors by tin. In all the cases, the n-hexane isomerization to i-C6 was only slightly affected by coking (Table 2, column 4). Isomerization of n-paraffins may occur over the metallic fraction via a bond-shift or via a cyclic

499 mechanism. It can also take place by a bifunctional metal-acid mechanism where the metal dehydrogenates the n-paraffin and the acid center isomerizes the olefin. Nevertheless, under reforming conditions the monofunctional mechanism is negligible compared to the bifunctional mechanism which is controlled by the acidic function [8]. On the other hand, several studies on coke formation over metal supported catalysts agree in that coke precursors are initially formed on the metal surface and then, through a slow diffusion mechanism, are condensed and accumulated on the support [9]. Thus, we can expect that under our experimental conditions coke formation deactivates mainly the metallic function. This interpretation explains why the acid-controlled formation of i-C6 isomers is practically not deactivated by r Catalyst thiotolerance for the n-hexane conversion followed the order Pt -_-Pt-Ge > Pt-Re > PtSn (Fig. 2). However, the thiotolerance trend is different when the individual reactions involved in the n-hexane reforming mechanism are analyzed separately. In fact, TTc,, TTBz and TT~ values decreased in the order Pt-Ge > Pt > Pt-Re >- Pt-Sn whereas in the case of TT,_c, the trend was Pt > Pt-Ge > Pt-Re _-- Pt-Sn (Table 4). These results may be explained by considering the nature of the sulfur adsorption on the catalysts. The adsorption of sulfur on Pt-based supported catalysts takes place on both the metal and the support [10]. Part of the adsorbed sulfur is resistant to H2 treatment at 773 K (Si, irreversibly held sulfur) and is located on the metal. The other part of the adsorbed sulfur is e"hminated in H2 at 773 K and is probably located on both metal and support sites. When the partial pressure of the mlfiw compound in the gas phase is low (as in the present work) only the irreversible sulfur on the metal is retained by the catalyst. Thus, the catalyst thiotolerance for a metallic reaction is related to the irreversibly held sulfur on the catalyst: the lower the Si amount, the higher the thiotolerance [6]. With the exception of the Pt-Sn catalyst, the catalyst thiotolerance trend determined here for the metal-controlled formation of Bz, C5 and gases was similar to that found for cyclohexane dehydrogenation, a monofimctional metallic reaction on Pt [6]. The Pt-Ge catalyst exhibited the highest values of TTc,, TTBz, TTMcP and TT~. This superior performance is attributed to the formation of Pt-Ge clusters upon H2 reduction at 773 K that decreases the electronic density of platinum and thereby weakening the strength of the S-Pt bond [11]. Germanium inhibits the adsorption of irreversible sulfur on Pt and, as a result, the Pt thiotolerance in the Pt-GeJAI203 catalyst increases. In the case of Pt-Re/AI203 the adsorption of sulfur on Re is thermodynamically favored compared to Pt and preferentially forms surface Re sulfide. A number of papers have confirmed that Re adsorbs more sulfur than Pt and that sulfur is more strongly bonded to Re than Pt [12]. This is consistent with the very low thiotolerance exhibited here by the Pt-Re catalyst for the metal-controlled reactions. Pt-Sn/AI203 displayed a thiotolerance for the formation ofBz, MCP, C5 and gases similar to that of Pt-Re/Al203. However, in contrast to Pt-Re/Al203, sulfur does not adsorb on tin and the amount of Si on Pt-Sn/AI203 is sin~ar to that on Pt/AI203 [13]. The low thiotolerance exhibited by the Pt-Sn catalyst can not therefore been explained only by the characteristics of the sulfur adsorption on the metal fraction. Probably, the sensitivity to gdfur poisoning was enhanced because the formation of coke was suppressed over Pt-Sn/Al203 (Table 2). Coke formation may protect the metal against sulfur poisoning [ 13] and this effect will be negligible on the Pt-Sn catalyst. The catalyst thiotolerance trend found for the isomerization of n-hexane to i-C6 (Table 3, column 4), an acid-controlled reaction, was different than that stated above for the metalcontrolled reactions. It has to be noted here that the acidity of all the catalysts was similar since they contained similar chlorine level (wt% C1 = 0.7-0.8) over the same support. Sulfur practically did not affect the isomerization activity of the Pt-Ge catalyst and this is attributable

500 to the inhibition by Ge of the S adsorption on the Pt active sites. The i-C6 formation was similarly inhibited on Pt-Re and Pt-Sn catalysts but for different reasons. Sulfur is strongly bonded to Re and the amount of Si per metal surface atom is higher on Pt-Re/A1203 than on Pt/AI203. Thus, the metallic function and the hydrogenation activity are drastically poisoned in the Pt-Re catalyst as the sulfur feed concentration is increased. As a consequence, the i-C6 formation which is controlled by the acid sites on the flesh Pt-Re catalyst becomes controlled by the metallic function upon sulfur poisoning because the number of accessible metal sites is drastically depleted. On Pt-Sn/A1203 the amount of Si is lower than on Pt-Re/A1203. However, because the hydrogenation activity of the fresh Pt-Sn catalyst is low, a small amount of sulfur on Pt is enough for suppressing its dehydrogenation activity. Thus, isomerization on PtSllfA1203 becomes controlled by the metallic function at a lower S coverage as compared to PtRe/AI203. Finally, the isomerization activity on the fresh Pt/AI203 catalyst increased upon addition of thiophene-doped n-hexane. Formation of i-C6 isomers on Pt/A1203 remains controlled by the acid function, at least up to the feed S concentration used in this work, because the high hydrogenation activity exhibited by the fresh catalyst compensates the partial blocking of the metal function by sulfur. Because sulfur is not adsorbed by the support, we can exclude the possibility that the catalyst acidity will increase upon S addition. Probably, the increase of the isomerization activity is due to an "indirect control" of the metallic function on i-C6 production [ 14]. In fact, by partially poisoning the metallic function, sulfur decreases both the hydrocracking of C6 paraffins and the transformation of olefins to MCP, thereby increasing the abundance of intermediates available for isomerization. REFERENCES

1. C.R. Apesteguia, T.F. Garetto and A. Borgna, in C.H. Bartholomew and J.B. Butt (eds), Catalyst Deactivation 1991, Elsevier, Amsterdam, 1991, p. 399. 2. J. Corella and A. Monz6n, Ind. Eng. Chem. Res., 27 (1988) 369. 3. T.F. Garetto, A. Borgna, A. Monz6n y C.R. Apesteguia, Proc. 10th Iberoamerican Symposium on Catalysis, Segovia, Espafia, Vol II, 1992, p. 1257. 4. G.B. Matin and G.F. Froment, Chem. Eng. Sc., 37 (1982) 759. 5. J.M. Parera, J.N. Beltramini, C.A. Querini, E.E. Martinelli, E.J Chutin, P. E. Aloe and N.S. Figoli, J. Catal., 99 (1986) 39. 6. A. Borgna, T.F. Garetto, A. Monz6n and C.R. Apesteguia, J. Catal., 146 (1994) 69. 7. J. Beltramini and D. Trimm, Appl. Catal., 31 (1987) 113. 8. J.H. Sinfelt, in J.R. Anderson and M. Boudart (eds.), Catalysis, Science and Technology, Springler-Verlag, Berlin, 1981, Vol. 1, p. 257. 9. J.M. Parera and J.N. Beltramini, J. Catal., 112 (1988) 357. 10. C.R. Apesteguia, T.F. Garetto, C.E. Brema and J.M. Parera, Appl. Catal., 10 (1984) 291. 11. T.F. Garetto, A. Borgna and C.R. Apesteguia, in J.W. Hightower, W.N. Delgass, E. Iglesia and A.T. Bell (eds.), Studies in Surface Science and Catalysis, Vol. 101, 1996, p. 1155. 12. C.R. Apesteguia and J. Barbier, J. Catal., 76 (1982) 352. 13. C.L. Pieck, P. Marecot and J. Barbier, Appl. Catal. A: General, 145 (1996) 323. 14. T.F. Garetto, A. Borgna and C.R. Apesteguia, in B. Delmon and G.F. Froment (eds), Catalyst Deactivation 1994, Elsevier Science B.V., 1994, p. 369. 15. C.A. Querini, N.S. Figoli and J.M. Parera, Appl. Catal., 53 (1989) 53.

9

Elsevier Science B.V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

501

D e a c t i v a t i o n and Attrition o f Iron Catalysts in Synthesis G a s Nancy B. Jackson al , Abhaya K. Datye b, Linda Mansker b, R. J. O'Brien c and Burtron H. Davis c aSandia National Laboratories, PO Box 5800, MS 0710, Albuquerque, NM 87185, United States bUniversity of New Mexico, Dept. of Chemical and Nuclear Engineering, Albuquerque, N M 87131, United States CCenter for Applied Energy Research, University of Kentucky, 3572 Iron Works Pike, Lexington, Pike, Lexington, KY, 40511, United States

A number of different iron Fischer-Tropsch catalysts were characterized with an emphasis on the study of morphological changes of the iron and iron carbide phases as well as the growth and morphology of carbon species during reaction. The effects of catalyst composition, reaction temperature, and pretreatment were investigated. Temperature programmed surface reaction of H2 with a number of used catalysts revealed the formation of more than one carbide under certain circumstances. This paper considers deactivation and catalyst attrition for catalysts reacted in a slurry phase autoclave reactor. Several similar catalysts were tested for activity for long periods of time (1000-3000 h), and it was found that when operated properly, these iron Fischer-Tropsch catalysts can be run in slurry reactors with only modest deactivation during the time periods tested. 1. INTRODUCTION Fischer-Tropsch (FT) synthesis has been used and studied for over 50 years. Several major eras of study include World War II, post-1950 in South Africa, and during the "oil crisis" of the 1970s. Each era has revealed, with improved technology, more understanding of the chemical processes of FT synthesis. In addition, improved reactor technology has brought about an increased demand in FT catalyst performance. Recently, interest has been focused on the slurryphase bubble column reactor which is an especially efficient reactor for the highly exothermic, wax-producing Fischer-Tropsch reaction. The bubble column reactor brings about different issues to catalyst deactivation than a fixed bed reactor, since a bubble column gives a more uniform catalyst dispersal and a more uniform reactant, product, and potential poisons (sulfur primarily) distribution throughout the bed, than a fixed bed reactor. Besides the traditional catalyst deactivation, an equally serious problem is the attrition of the catalyst. The bubble column 1Person to whom correspondence should be addressed.

502 reactor works best with catalyst particles on the order of 20-100 lxm. If catalyst attrition during reaction causes the particles to break down into sub-micron dimensions, the liquid in the bubble column will become very viscous and the bubble column will become difficult to operate. More importantly, even if a small portion of the catalyst attrits to <1-5 lxm size particles, the separation of the wax and catalyst becomes problematic. 1.1. Deactivation

The various factors that can contribute to deactivation of iron Fischer-Tropsch (FT) catalysts include transformation of the active phase into an inactive constituent, poisoning by carbonaceous species and heteroatoms, and loss of active phase surface area. Progress in elucidating the causes of deactivation is hampered by the inability to conclusively identify the active phase in iron FT catalysts. In recent work involving doubly promoted, unsupported iron catalysts, the sequence of phase transformations shown in Figure 1 that take the catalyst from its as-prepared hematite phase to iron carbide [1,2] was postulated. Hematite

L

..__._ .

Deactivation

Activation

AoEv~onin H2, CO, or syngas

(=)Calcination ~

Magneld~

~

(all amolphou,~CHx

~~H) 2 acl~,amlnio

/syngas .

qD: 10

-

% . o

o~-Fe

~ graphitization

carbide

1 9/ syngas / Iron

'~1t~1t~ /

(elReoxidati .

.

.

.

.

Figure 1. Summary of phase transformations. As shown in Figure l a, after calcination the catalyst is present in the form of hematite, txFe203, and is rapidly transformed pseudomorphically to magnetite. The overall shape and internal texture of the hematite crystallites are preserved. This process occurs readily in H2, CO, or in the syngas mixture (H2/CO = 0.7). Further reduction in H2 can lead to formation of o~-Fe (Figure lb). Upon activation in CO or syngas, transformation into iron carbide occurs (Figure 1c). The iron carbide particles grow as small nodules which nucleate from the magnetite phase. After use in a syngas environment, the carbide nodules always exhibit a surface layer that appears as a halo around the particle (Figure ld). The thickness of this surface layer does not seem to increase with reaction time on stream [ 1]. Furthermore, the surface layer can be readily

503 hydrogenated at reaction temperature in H2, hence we suspect the surface layer is composed of CHx species and may actually represent the pool of carbonaceous species on a working F T catalyst [3]. No such halo or surface layer is seen on the surface of the magnetite particles. A TEM image of a catalyst that exhibits such a surface layer is shown in Figure 2. The catalyst is identical to that used in a previous study [1]. This catalyst shows the carbide particles that have nucleated from the magnetite after catalyst activation in syngas at 270~ for 2 h followed by 45 h of reaction at atmospheric pressure and Figure 2. HRTEM of FT catalyst from fixed bed reactor H2/CO = 0.7. As shown following 45 h in H2/CO=0.7 at 630 Torr later, very similar carbide particles with a "halo" surface layer were observed also in a catalyst removed from a FT slurry reactor operated at 13 atm, 270~ under conditions that are fairly representative of industrial FT bubble column reactors. Previous work described in the literature [1,2,4] suggests that iron carbide is a necessary constituent of an active FT catalyst and hence it is reasonable to believe that the carbide surface area should correlate with catalyst activity. The actual surface area of the carbide phase on a FT catalyst is difficult to measure, because the carbide surface is covered with CHx which is reactive in the presence of H2 [ 1,2]. This makes the accuracy of HE chemisorption, one potential method, uncertain. Since CHx covers all the carbide particles and since it is unclear whether H2 would measure CHx surface area, iron carbide surface area, or react with the CHx, the accuracy of HE chemisorption is uncertain. In addition, the extreme air sensitivity of the catalyst, and the inherent wax build up on any catalyst that has been reacted in synthesis gas, makes potential sorption surface-measurement difficult since it requires removing wax from the pores of the catalyst in an oxygen free environment before a meaningful sorption experiment. But if the carbide particles all grow from the magnetite and are similar in particle size, then the total amount of carbide as seen by a bulk technique such as x-ray diffraction may also be used as a measure of the active phase. Such an approach was used in a previous study by Duvenhage, et al. in 1994 [5]. These authors performed a series of deactivation studies on catalysts which were first used in an industrial fixed bed reactor at SASOL, then removed, under nitrogen, from sections along the bed, and transferred to microreactors for further study. Their results [5] showed that the activity lost at the front of a fixed bed is caused by S poisoning. However, downstream into the reactor, they observed a clear correlation between the loss in activity and an increase in magnetite concentration. Further, they observed that in the region of highest activity, the catalyst contained

504 significant amounts of the carbide phase and hardly any magnetite, further confirming that the carbide is probably the active phase or associated with the active phase for FT synthesis. 1.2. Attrition

Previous work has indicated that attrition in iron FT catalysts is caused not just by a physical process, but that two "chemical" factors contribute to FT iron catalyst attrition. One is the catalyst transformation from magnetite to carbide which causes the crystals of magnetite to nucleate into crystallites of the carbide phase [1,2]. The other cause of attrition is the deposition of carbon on the carbide surfaces which causes the carbide crystallites to further separate from one another [ 1,2]. Considering the discussion above, and the role carbide and the carbon coveting the carbide surface plays in the activity of the catalyst, attrition resistance and activity may be at odds against one another. In order to understand deactivation and the attrition of iron FT catalysts, it is essential to understand the factors which enhance or deter the formation of iron carbides, carbon (CHx), and graphitic carbon during the course of the FT reaction. One point this paper addresses is the effect of copper and potassium on the formation of carbides and carbon deposition on catalysts during reaction by temperature programmed surface reaction. In addition, the effect of reaction temperature on the carbon and carbide formation for several Fe/Cu catalysts is also examined. Copper and potassium are common promoters that have been added to FT catalysts since the 1930's. Copper is added to make the Fe304 easier to reduce. Although it appears iron carbide is the active phase and not tx-Fe, catalysts that are made from highly dispersed iron which cannot be reduced to o~-Fe have been reported to be ineffective FT catalysts [6]. Potassium is added as a promoter, and it has a significant effect on the activity and selectivity of the catalyst. Adding increasing amounts of potassium increases the selectivity towards longer chain hydrocarbons, meaning more economically-attractive wax is made and less low-value C1-C4 products. Addition of small amounts of potassium increases the overall activity of the catalyst, but at higher potassium loadings the overall conversion decreases [7]. This paper also investigates the catalyst microstructure of iron catalysts reacted in slurry phase in an autoclave reactor. Catalyst activity was recorded as a function of time and catalyst samples were withdrawn at several times during a run and at the end of the run. The catalyst microstructure was studied via x-ray diffraction and transmission electron microscopy to shed insight into the nature of deactivation processes in FT catalysts. Based on our previous work, we can schematically suggest two pathways (see Figures le and If) that could lead to catalyst deactivation. The two pathways are oxidation of the carbide to magnetite or site blockage by inactive carbon species (e.g., graphitic carbon). A third pathway, the loss of active phase surface area via sintering is also possible but not shown in this figure. We have not considered the effect of S poisoning which has been adequately documented [5, 8-11 ] as a cause of catalyst poisoning. 2. EXPERIMENTAL The iron catalysts used in the temperature programmed adsorption experiments were synthesized using iron nitrate precipitated with a base. In the case of the copper/iron catalysts the base used was Na2CO3 and for the iron/potassium catalysts, NaOH was used to precipitate

505 the oxide. The catalysts were repeatedly washed to remove the sodium present prior to calcination. No sodium was detected after the final wash using X-ray fluorescence spectroscopy. The potassium was added by incipient wetness using K2CO3 solution. The catalysts were calcined in air at 300~ and were characterized using N2 adsorption with BET analysis. The temperature programmed surface reaction (TPSR) apparatus had a 1/4 inch stainless steel tube catalyst bed. The catalyst was placed in the TPSR reactor, and before TPSR, the FT reaction was run in the same reactor. Therefore, the catalysts did not need to be passivated following reaction in synthesis gas and prior to TPSR. The catalyst was loaded into the reactor and the effluent of the reaction was directed towards the vent with a cold trap in line to collect wax produced by the reaction. The catalyst was pretreated in H2 for two h. Previous studies showed that this pretreatment leads to the formation of magnetite and does not reduce the catalyst to a metallic state [1]. The catalyst was then reacted in synthesis gas (H2/CO-0.7) at 16 atm and 215~ for 24 h. In addition, several of the Fe/Cu catalysts were reacted at 270~ to gauge the effect of reaction temperature on carbon and carbide formation. Following reaction, the catalyst was cooled to 180-200~ and was flushed with He for anywhere from 4 h to 2 days until no more hydrocarbons could be detected in the helium stream by an FID detector. The reduction gas, 10% H2/90% He, was introduced after the catalysts had been cooled to 100~ The temperature was raised at 5~ until it reached 270~ where it remained for 1 h. Next the temperature was raised at 2~ until it reached 700~ In the case of the copper-containing catalysts this capability was limited to an upper temperature of 550~ The catalyst was held at the high temperature for 30 minutes or until baseline was reached. An FID detector was used to measure the methane produced from hydrogen reacting with the catalyst carbon. The use of an FID detector, versus a TCD, allowed us to differentiate between the desorption of water and the desorption of a carbon species as well as realizing high sensitivity to the product. The catalysts reacted in the autoclave reactor were examined by x-ray diffraction and transmission electron microscopy (TEM). Studies were performed in a l-liter autoclave reactor under high pressures and typical process conditions. Samples were withdrawn from the reactor at various times for analysis and were obtained via a dip tube in the CSTR. The RJO-189 series of Cu, Si, K-containing catalysts are based on 100 Fe/4.4 Si/2.6 Cu/1.0 K oxides (atomic % ratios to iron), activated with H2 at 220~ at 120 SLPH for 25 h, and then used for FT synthesis. LGX-171 and-175 (promoted with Si and K) were prepared from 100 Fe/4.4 Si/1.0 K oxides (atomic % ratios to Fe). LGX-171 was pretreated in syngas (H2:CO = 0.7) at 1 atm, at 270~ for 24 h in wax, and then reacted in syngas for 3164 h. LGX-175 was pretreated in CO at 1 atm at 270~ for 24 h in wax and the FT synthesis ran for 1160 h at 270~ The last sample, LGX-195, was prepared from a UCI C-73 ammonia synthesis catalyst (fused Fe/K/Si), reduced in H2, passivated, and used 192 h in synthesis gas at 270~ For XRD, all samples were analyzed as-received, in the wax matrix, using a SCINTAG PAD-V diffracted-beam monochromator instrument, slit sizes of 2 and 4 (incident beam), and 1 and 0.3 (diffracted beam) with a Cu source. Sample size was 0.6 cm 3. The LGX samples were smearmounted and fast-scanned with 0.05 ~ chopper increment, 0.5 ~ per minute. The RJO-189 series samples were pack-mounted and step-scanned with 0.02 ~ chopper increment, 0.12 ~ per minute. Step-scanned samples were passed through a Reitveld Refinement routine, DBWS=9411 to

506 deconvolute the known phases: Fe304 and {x-Fe. TEM analysis was performed using a JEOL 2010 microscope operated at 200 keV. For TEM, the wax was removed by flowing helium through a tube filled with the waxy catalyst and collecting the wax in a downstream trap. During this process, the temperature was ramped gradually up to 400~ to enable removal of all the volatile components. The sample could then be examined in the TEM without any further processing. The samples from the fixed bed reactor that were examined by TEM were prepared by flowing syngas (H2/CO = 0.7) at ambient pressure (630 torr in Albuquerque) over a catalyst bed of 1 g. Other experimental conditions are identical to those described in Ref. 1. After reaction, the fixed bed samples examined by TEM were flushed with helium at reaction temperature, cooled to room temperature and a mixture of 1% 02 in He was slowly pulsed over the gas to make sure that there were no noticeable exotherms. Further details on precautions necessary to ensure controlled passivation are described elsewhere [12]. 3. RESULTS AND DISCUSSION 3.1. Catalyst surface area and carbon content 3.1.1. BET surface areas Two series of iron catalysts were synthesized: one with varying amounts of copper and one with varying amounts of potassium. The catalysts and their BET surface areas are shown in Table 1. The microscope showed the Cu/Fe series of catalysts to have very large particles (80 nm), which, since the potassium and copper iron catalysts were prepared by different chemists, probably accounts for, in part, the significant differences in surface area between the two batches of catalyst. 3.1.2. Carbonaceous species by TPSR Earlier work showed that there are three temperature regions in which three general types of carbon react off the catalyst surface in a flowing hydrogen/helium mixture [3]. This was determined by stopping the H2 reaction process at various temperatures, carefully passivating the sample and studying the catalyst morphologies and surface carbon species in HRTEM. A very reactive surface carbon reacts at 270~ This surface carbon is very sensitive to passivation. A used catalyst that has been passivated prior to TPSR contains an order of magnitude less amorphous carbon than a catalyst reacted under similar conditions but not exposed to oxygen prior to TPSR. This carbon is speculated to be CHx as mentioned above. ~The second region of reaction ranges from 385~ to 510~ and represents iron carbide reduction. The variation in the temperature of carbide reaction indicates different iron carbides may be present. However, our HRTEM does not distinguish between iron carbide phases. Graphitic carbon, the third carbon type, reacts above 550~ [3]. Shown in Figures 3 and 4 are the results of temperature programmed surface reaction on an iron, 1% Cu/Fe, and 2.5% Cu/Fe catalysts following 24 h of reaction in synthesis gas (H2/CO=0.7). The TSPR behavior of the two copper-containing catalysts was quite similar. The major difference was between the copper-containing and the iron-only catalyst. Less CHx carbon is formed on the iron-only catalyst, and the substantial amount of graphitic carbon formed on the iron-only catalyst is considerably less reactive. Also, the more copper in the catalyst, the more

507 Table 1 Fe/K and Fe/Cu catalyst surface area following calcining

easily its carbide (the carbide with desorption centered at -400~ is reduced. A second carbide peak is wt% K wt% Cu BET m2/~ observed for the copper-promoted 0 0 100 iron catalysts. When the TPSR of the 0.2 0 98.6 unpromoted (0% Cu) catalyst was 1.0 0 132 repeated with maximum temperature 2.0 0 148 increased to 700~ an additional 3.0 0 233 large peak was found at 570~ 0 1 30.8 indicating graphitic carbon (see Fig 4). 0 2.5 86.4 Reaction at higher temperature is known to increase attrition in iron FT catalysts [2]. A comparison of 2.0E+06 I TPSR experiments done on 1% O l%Cu -~ ~ 1.5E+06 ~ ~ o 2.5%Cu Cu/Fe catalyst following reaction at ~1 - --Fe203 215~ and 270~ is shown in Figure 0 5. Note that the catalyst reacted at 0 higher temperature has significant .~ ~ 5.0E+05 graphitic carbon at 550~ that is not 0.0E+00 seen on the catalyst reacted at lower 100 270 295 402 509 550 550 temperatures. In addition, the catalyst treated at higher Temperature *C temperature contains primarily the Figure 3. Temperature programmed surface reaction of 3 more difficult to reduce carbide catalysts following FT synthesis for 24 h at 215~ and 16 (reaction peak centered at ~525~ atm. TPSR reduction carried out to 550~ Thus, the increased reaction temperature has a significant effect on the carbon species grown on the catalyst. Similar shifts in peaks take 2.E+06 "~ | Fe203 place when 2.5% Cu/Fe catalyst is ~ 1.E+06 (samecatalyst compared at reaction temperatures of '~ 215~ and 270~ (Fig. 5). ~a ~ 5.E+05 Four iron catalysts promoted with 9~ff varying amounts of potassium were O 0.E+00 I I t,tm reacted at 215~ in synthesis gas for o 109 270 331 450 568 685 700 Temperature ~ 24 h and then tested using TPSR. The results of the experiments are shown Figure 4. Temperature programmed surface reaction of an in Figure 6, which tabulates the Fe-only catalyst following FT synthesis for 24 h at 215~ quantities of amorphous CHx carbon, and 16 atm. TPSR carried out to 700~ carbidic, and graphitic carbon found on each catalyst. The most graphite formed on the unpromoted iron

;

#X A

508

2.E+06 o 9-

1% Cu Reacted at 215C 1% Cu Reacted at 270C

2.E+06

1 .E+06

--

5.E+05

0 .E+00

100

270

295

402 509 550 Teml~rature C

550

Figure 5. Temperature programmed surface reaction of 1% Cu/Fe catalyst following reaction at 215~ and 270~ catalyst. Similar results were found by Eliason and Bartholomew [ 13]. Carbide content increased with increasing K content up to 1% and after 1% K promotion was reached the amount of carbide remained almost constant. The amorphous carbon was lowest on the unpromoted catalyst, but also decreased with increasing g

~ lO00

~ ~ 800 .~ ~ 600

~ C H x carbon Carbides IIGraph

~ II

n

n

] [

~ 400 ~ 200 '~ ~ ~. 0 0

0.2 1 2 wt% Potassium in Iron

3

Figure 6. Relative amount of amorphous, carbidic, and graphitic carbon on unpromoted and potassium-promoted iron FT catalysts after 24 h on stream at a reaction temperature of 215~

potassium promotion with a minimum at 2% K. Dry [7] published a study on the effect of potassium on activity and hardwax (boiling point > 500~ selectivity and found that activity increased with increasing K promotion to a maximum where activity then fell off. For attrition concerns, it is interesting to note that the addition of low amounts of

potassium (which would be expected to increase the activity) decreases the amount of graphite formed on the catalyst versus the unpromoted catalyst. Taking information from Reference 7 on the relative activity of potassium-doped iron and plotting it versus the amount of graphitic carbon formed during the first 24 h of reaction shows that the most active K-promoted iron catalysts contained the least amount of graphitic carbon production (Figure 7).

509

3.2. Deactivation

Shown in Figure 8a is an activity-time profile for Run RJO-189 [4]. The data indicate a stable activity profile, with only modest deactivation over 3000 hours. An XRD powder pattern of four samples from that run is presented in Figure 9a. We were not able to analyze a sample after o = 160 the initial activation in H2 via XRD, and this = ~ 12o treatment has been shown to reduce the catalyst N u 80 to ct-Fe. Mossbauer analysis indicated that the ~ 40 sample was part magnetite and part oc-Fe, but that 0 was after exposure to 02. The XRD patterns ~ 20 40 60 obtained as a function of time on stream show that the catalyst appears to consist mostly of Relative activity (Ref. 7) magnetite, with little iron carbide. This analysis is Figure 7. Relative activity versus relative based on examination of the ratios of the 100% amount of graphitic carbon on Fe/K catalysts peak of magnetite at 2.53 A and next most intense with varying amount of potassium, peak at 2.098 A. Since the iron carbide has a prominent peak also around 2.098 A, a crude estimate of fraction carbide can be obtained by comparing the experimental peak ratio with that from a pure sample of magnetite, as described previously [ 1]. The characteristic "fingerprint" region for identification of iron phases lies between 30 ~ and 45 ~ 20. In that region, peaks can be observed at about 2.532 A (Fe304 100% intensity peak), at 2.172 A which corresponds to FexC (x, because nearly all the known carbides have a peak at this position), at 2.111 A (ascribed to Z- or e' carbide), and at 2.098 A, (a secondary Fe304 peak). It should be noted that the labeled d-spacings do not correspond exactly to the numbers cited above. Significant displacement errors were observed on analysis of these samples manifested by systematic shifts in 20 position, the amount being function of the phase. Since the peaks at 2.111 A and 2.098 A overlap considerably, the peak at 2.172 A is chosen as the one to be compared with the Fe304 peak at 2.532 A as a measure of relative 100 phase abundances of magnetite to carbide. In this set of samples that 80 ratio is high, at 3.69, 4.00, 3.35, "i 60 and 3.94 for 122, 888, 1796, and 3547 hours, respectively. It o RJO-189 CO conv o 40 should be emphasized that these r~ ratios are very rough estimates of ~ r~ 20 relative abundance. Also, the anomaly at 1796 hours can be I I I 4000 explained by further examining 1000 2000 3000 Figure 9a. Note that there is a T i m e o n s t r e a m (h) significant amount of t~-Fe phase Figure 8a. Activity plot for Run RJO-189. present in the sample, which

"~

"2

510 100

could also account for the presence of higher amounts of ( ~ : o Oo 0~ 80 _:Q:b ( ~ : b ( ~ o o o ~ o O o~ carbide than in the earlier samples, c~P o since we observe that when 60 carbide content is high, o~-Fe is o LGX-175 CO conv also present. As proof that the o= 40 peak observed is o~-Fe, we 0r~ 20 analyzed the sample, using a Reitveld analysis routine to strip t t I t I the magnetite and c~-Fe phases 400 600 800 1000 1200 (Figure 9b). The lack of negative 0 200 Time on stream (h) peaks in the residual translates to a high goodness of fit. Figure 8b. Activity Plot for Run LGX-175. The high magnetite-carbide ratios may appear, at first glance, 100 to be inconsistent with our 80 assertion that significant amounts O 0 o of carbide are required for high 0 11 60 0 activity. However, we suspect > 0 I o LGX-171 CO Conv that exposure of samples to the 0 o= 40 ambient air during transfer to the 0r~ 20 XRD apparatus may have led to formation of the magnetite during I i t I I sample collection and hence the relative amounts of carbide and 1000 2000 3000 Time on stream (h) magnetite do not represent the ratio actually present in the Figure 8c. Activity Plot for Run LGX- 171. reactor. We have also found that it is possible, over time, for some of the reactive iron phase to 100 oxidize, even embedded in wax, at ~ 80 room temperature. This is shown o in Figure 9c, where XRD powder "~ 60 o LGX-195 CO conv patterns are shown for the sample o freshly withdrawn at 1796 hours o 40 r~ 0 0 0 0 on stream recorded 9 months O 20 apart. The o~-Fe peak at around 2.027 A has disappeared after 9 I 0 I I months of sitting in air in the wax. 200 The o~-Fe peak was seen only in 150 0 50 100 Time on stream (h) the sample withdrawn at 1796 hours into this run. It is highly Figure 8d. Activity Plot for Run LGX-195. unlikely that this particular .Jl

,m

0I

511

Urdversity of I~rltuc W F~J0189 series i ................................

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Figure 9a. XRD of catalyst RJO-189 Series (100 Fe/4.4 Si/2.6 Cu/1.0K)

260

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. 50

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Figure 9b. Rietveld Deconvolution of magnetite, a-iron in Sample RJO-189M Jg-Month repeatability Study: RJO189M,~ 180 ..~ m

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Figure 9c. Comparison of two runs, 9 months apart, from sample RJO-189M

512 sample would be the only one showing the o~-Fe to be present. What is more likely is that this particular sample was not oxidized to the same extent as the others due to subtle differences in the sample withdrawal procedure. Since the catalyst was reduced to o~-Fe during initial activation, the o~-Fe peak seen in the XRD pattern may arise from an iron phase that never transformed into carbide after exposure to syngas or it may be possible that there are certain equilibrium amounts of o~-Fe and iron carbide that exists in the reaction mixture. Further study is necessary before the morphology of the et-Fe phase seen in the reaction mixture can be ascertained. From Figure 9a it can be concluded that there is little change in the composition (relative amounts of carbide and iron oxide phases) with time within the autoclave reactor catalysts. A similar conclusion was reached in a previous study of catalyst composition with time, where the sample withdrawal procedure involved the use of an argon inert gas purge to protect the sample from further oxidation [ 14]. The study in ref. 14 found the iron catalyst sample contained two or more carbide phases with little or no magnetite. In the data presented here in Figure 9a, there is more magnetite found than in ref. 14, however the amount of magnetite does not change with time. In addition, there does not appear to be any significant growth in crystallite size, as evidenced by the absence of change in the sharpness of the diffraction peaks, with time on stream. Hence, the cause for the modest deactivation seen here deserves further study. We plan to study in future work the nature of carbonaceous (particularly graphitic) species deposited on the catalysts. Conversion versus time data are shown for Runs LGX-171,175, and 195 in Figures 8b, c, and d. The catalysts used in Runs 171 and 175 were nearly identical, the only difference being the pretreatment method. The catalysts for Runs LGX-171 and 175 were similar in composition to the catalyst used for Run RIO-189, except they did not contain copper, which facilitates reduction of Fe304 to ~x-Fe. Despite the differences in activation treatment, the conversion versus time on stream is very similar for these three Runs. When we compare the performance with time in Runs RJO-189, LGX-171 and -175, we can conclude that activation treatment (H2, CO, or syngas) has very little effect on the long term activity of these catalysts. For Runs RJO-189, LGX-171 and 175, only one sample was available from each test, which was taken at the end of the run. LGX-175 is the catalyst with the highest CO conversion activity, varying from a high of about 88% to a low of 80%. Analysis by XRD of the sample from Run LGX-175 (Figure 10a), which had been kept in FT wax, shows a number of peaks in the fingerprint region, with a Fe3Oa:FexC ratio of 1.96. HRTEM micrographs of this catalyst (Figure 11 a and b) show that the morphology is indeed very similar to that depicted schematically in Figure 1. The wax was stripped off for TEM analysis as described above. However, not all of the wax could be volatilized, with the result that a wax residue covers the entire sample. Inside the wax residue are small particles covered with a surface layer that looks very similar to that shown in Figure 2. It is indeed remarkable that despite the differences in the method of preparation and initial composition, the primary particles in the catalysts seen in Figures 2 and 1 l a look so much alike. The carbide particles (identified by their lattice fringes), are surrounded with a halo of amorphous carbonaceous material which must differ in composition from the wax residue since there is sufficient contrast to distinguish them. In agreement with the XRD results from Run RIO-189 (time on stream 1796 hours) an o~-Fe peak is observed in the XRD pattern in

513

ISarq~ LG~-175, B-ID I '120 100 8O

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Figure 10b. XRD of catalyst sample LGX-171, in wax, from end of run. leo

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514

wax residue Fe carbide

9

L G X- 17 5 wax stripped

'~

.......... ~!~!~"= i

Figure 11. HRTEM micrographs of catalysts reacted in the autoclave reactor in synthesis gas. (a) LGX-175; (b) LGX-175; (c) LGX-195; (d) LGX-195. Fig 10a. The presence of a rich spectrum of carbide peaks is further evidence that oxidation during sample removal did not proceed to a great extent. Data from Run LGX-171 indicate intermediate activity (Figure 8c), with a sharp drop in activity near the end of the run. The XRD of the LGX-171 sample taken near the end of the run shows a fingerprint region with less detail, and a Fe3Oa:FexC ratio of 2.38. This indicates the presence of 21% more magnetite at the end of the run than in the LGX-175 sample. The composition of the catalyst does not completely explain the lower activity. Indeed, other runs (notably the ones in Figs 8a and b) continued for longer times without such a sudden loss of

515 activity. Unfortunately, HRTEM and TPSR analyses of the catalyst at the end of the run are not available to examine for clues to the sudden deactivation. Finally, for sake of completeness, we present a run with a fused Fe catalyst containing no promoters (Cu or K). This is an ammonia synthesis catalyst (C-73 from United Catalysts) and is known to be difficult to reduce. In the F-T synthesis, this catalyst (Run LGX-195) was the least active sample. From Figure 8d, one can observe that conversion was never higher than about 41%, and declined to a low of 35%. The Fe3Oa:FexC ratio is 6.332, which is consistent with the hypothesis that for a higher magnetite to carbide ratio, catalyst activity is lower. This is also consistent with the aforementioned findings of Duvenhage, et al., 1994 [5], and Jager, et al. 1995 [ 14]. HRTEM micrographs of this material (Figures 1l c and d) show hollow spheres composed of magnetite. As seen in the higher magnification image, the magnetite particles show no evidence of a surface carbon overlayer or halo surrounding the particle. The TEM image did not reveal any particles similar to those seen in Figs l a or lb. The low degree of reduction and of carbide formation is consistent with the low activity of this sample. 4. CONCLUSIONS 1. Catalyst composition does not change appreciably with time on stream for the runs performed above. 2. In this set of runs, sample pretreatment seems to have little effect on catalyst activity. All of these pretreatments (H2, cO or syngas) gave comparable long term performance. 3. An t~-Fe peak is often seen ~ in the active FT catalyst (Figures 9a and 1l a and Ref. 14). Surprisingly, this peak is seen in the catalyst reduced in H2 (RJO-189, Figure 9a) as well as the catalyst reduced in CO (LGX-175, Figure l la) at the end of the run. Therefore it is possible that there is a thermodynamic equilibrium amount of t~-Fe phase under these reaction conditions. 4. Operated properly, F-T catalysts in slurry reactors undergo only modest deactivation. The modest deactivation seen in these runs may be due to sintering or conversion of active carbons or carbides to less active forms, but further TEM and XRD will be needed to verify particle size growth and possible phase changes. Catalysts that are known to deactivate more quickly (1% CufFe reacted at 270~ Fig. 5, or Fe reacted at 215~ in Fig. 3) have carbides which react with H2 at higher temperatures than catalysts that are known to deactivate more slowly (1% Cu~e reacted at 215~ Fig. 5 and Fig. 3) 5. While we cannot conclusively say what the active phase is, it is clear that the presence of the carbide is essential. A carbon halo is also observed on the surface of carbide particles in an active catalyst. This surface film does not grow with time on stream, as evidenced by observation of comparable film thickness in catalyst recovered after almost 3000 hours on stream. The film probably represents the working pool of CHx species. A 'halo' is seen in the HRTEM because the carbonaceous species stay on the surface as the catalyst is cooled to room temperature when the reaction is quenched. 6. The increase in graphite formation correlates with a decrease in relative activity when looking at a series of Fe/K catalysts.

516 7. In order to understand the nature of the working catalyst, it is important to protect the sample against oxidation; hence removal under an inert blanket and protection of the sample from further oxidation is necessary. ACKNOWLEDGMENTS

The authors thank Mark Miller for help with XRD analysis and numerous helpful discussions. This work was supported at the University of New Mexico by U.S. DOE contract DOE-FG-2295PC-95210, National Science Foundation NSF HRD-93-53208, and the Microbeam Analysis Laboratory in the UNM Earth and Planetary Sciences Dept. where TEM was performed. At the Center for Applied Energy Research, this work was supported by U.S. DOE contract number DE-AC22-94PC94055 and the Commonwealth of Kentucky. The authors thank Mark Harrington and Ron Sandoval for their help regarding the TPSR experiments. At Sandia, this work was supported by the United States Department of Energy under contract DE-AC04-94AL850000. Sandia is a multiprogram laboratory operated by Sandia Corporation, a Lockheed Martin Company, for the United States Department of Energy. REFERENCES

1. Shroff, M. D., Kalakkad, D. S., Coulter, K.E., Kohler, S. D., Harrington, M. S., Jackson, N. B., Sault, A. G., and Datye, A. K., J. Catal., 156 (1995) 185. 2. Kalakkad, D. S., Shroff, M. D., Kohler, S., Jackson, N., and Datye, A. K., AppL Catal. A., 133 (1995) 335. 3. Shroff, M. D., Datye, A. K., and Jackson, N. B., in preparation. 4. O'Brien, R. J., Xu, L., Spicer, R. L., Bao, S., Milburn, D. R., and Davis, B. H., Catalysis Today, in press. 5. Duvenhage, D., Espinoza, R., and Coville, N., in Catalyst Deactivation 1994, eds. B. Delmon and G.F. Froment, Studies in Surface Science, Elsevier, 1994, p. 351. 6. Anderson, R. B., The Fischer-Tropsch Synthesis, Academic Press, Inc. Orlando, 1984, p. 147. 7. Dry, M.E., in "Catalysis: Science and Technology" (J. R. Anderson and M. Boudart, eds.), Vol. 1, Chapter 4, Springer-Verlag, Berlin and New York, 1981. 8. Chaffee, A. L., Campbell, I., and Valentine, N., Appl. Catal., 47 (1989) 253. 9. Madon, R. J., and Taylor, W. G., Hydrocarbon Synthesis (1979) 98. 10. Bartholomew, C. H., and Bowman, R. M., Appl. Catal., 15 (1989) 59. 11. Stenger, H. G., and Satterfield, C. N., I&EC Proc. Des. & Develop., 24 (1985) 415. 12. Shroff, M.D. and Datye, A.K., Catalysis Letters, 37 (1996) 101. 13. Eliason, S.A. and Bartholomew, C.H., in Catalyst Deactivation 1997, eds. C.H. Bartholomew and G.A. Fuentes, Studies in Surface Science, Elsevier, in press. 14. Jager, B., Espinoza, R.L., Catalysis Today, 23 (1995) 17. 15. Mansker, L.D., Bukur, D. and Datye, A.K., in preparation.

~ Elsevier Science B.V. All rights reserved Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

T e m p e r a t u r e - P r o g r a m m e d Reaction Study of C a r b o n T r a n s f o r m a t i o n s F i s c h e r - T r o p s c h Catalysts During Steady-State Synthesis

517

on I r o n

S.A. Eliason and C.H. Bartholomew BYU Catalysis Laboratory, Department of Chemical Engineering, Brigham Young University, Provo, UT 84602 Temperature-programmed surface reaction (TPSR) in hydrogen was used to differentiate up to seven different carbon types on iron Fischer-Tropsch (FT) catalysts. Both K-free and K-promoted catalysts were pretreated in 1:1 H2/CO syngas for 6-18 hours to achieve well-developed carbon surface coverages. Reaction temperatures of 215 and 245~ were selected to provide conditions of steady-state activity (no apparent deactivation) and relatively rapid deactivation. Possible carbon forms in the catalyst include (l&2) atomic carbon on planar and edge sites, (3) H~igg carbide, (4) epsilon-carbide, (5) amorphous carbon and (6) moderately- and (7) well-ordered graphitic carbons. Deactivation of iron catalysts in FT synthesis under the conditions of this study is associated with (1) transformation of high activity H~igg-carbide to less active epsilon-carbide and (2) formation of a highly ordered graphite on the surface of the catalyst. 1. I N T R O D U C T I O N One of the principal modes of catalyst deactivation for Fischer-Tropsch (FT) catalysts is loss of catalytic surface area due to the accumulation of carbonaceous species on metal/metal carbide surfaces and in the pores of the catalyst and/or formation of inactive carbide phases [ 1-3]. These carbon-containing species are probably products of the condensation/polymerization of atomic carbon or CHx reaction intermediates, formed during reaction by CO dissociation and subsequent partial hydrogenation of the atomic carbon [2]. Temperature-programmed surface reaction (TPSR) with hydrogen is a well-established, definitive technique for identifying and quantifying different carbon and metal carbide species present on the surfaces of supported metals and metal carbides [2,4]. It was used, for example, by McCarty et al. [4] to identify 6 different carbon types and bulk Ni carbide in a Ni/A1203 catalyst exposed to ethylene under different high temperature conditions. While a few previous TPSR studies were conducted to identify carbonaceous species during FT synthesis on Fe and Ru [5-7], none were conducted at sufficiently long reaction times for steady-state surface coverages to have been realized. This work involved the study by TPSR in H 2 of carbon and carbide species formed in iron catalysts during FT synthesis for up to 18 hours under steady-state conditions or up to 6 hours under severe, deactivating reaction conditions. The effects of reaction temperature and potassium oxide promotion on the distribution of carbon types were also determined.

2. EXPERIMENTAL 2.1 Catalyst preparation A singly-promoted, unsupported iron catalyst was prepared by decomposing a melt of Fe(NO3)3 ~ 9 H20 with 1 wt.% Al(NO3)3 ~ 9 H20 at 120~ until dry followed by calcination at

518 200~ for 12 hours. The solid was crushed with mortar and pestle to a powder. It was found that 1-2 wt.% of alumina helps maintain the structure and surface area of the fresh catalyst [8]. A doubly-promoted Fe/K/A1203 catalyst was prepared by adding dissolved KNO3 to reduced/passivated 99% Fe/1% A1203 followed by the same drying and calcining procedure as for the singly-promoted catalyst. The K/Fe molar ratio for the K-promoted catalyst was 0.013 (corresponding to about 0.9 wt.% K). Catalyst samples for the TPSR study were taken from batches prepared previously for use in high-pressure, deactivation studies using monolithsupported catalysts in a Berty reactor [8]. As such, they had been physically mixed with 30% by wt. Conoco Dispal alumina to help bind the catalyst to the monolith walls. Catalysts were reduced in flowing hydrogen at 400~ for about 24 hours.

2 . 2 Apparatus A 9 mm-diameter quartz cell with fritted disk was used as the reaction chamber for reduction, pretreatment and TPSR experiments. An external heater surrounding the 1-atm cell was controlled by an Omega temperature controller. Gases were purified before metering through Brooks mass flow controllers. Evolution of methane during the TPSR experiments was measured by a UTI 100C quadrupole mass spectrometer with comput-erized multiple peak selection. Static hydrogen chemisorption measurements were conducted in a standard volumetric glass adsorption apparatus.

2 . 3 Procedure A fresh catalyst sample (about 100 mg) was placed in the cell which was in turn connected to the TPSR system. Reduction, pretreatment and TPSR were performed consecutively without removing the catalyst from the cell or allowing air to enter the system. After reduction in hydrogen at 400~ for 24 hours, the catalyst was allowed to cool to the reaction temperature (either 215 or 245~ in flowing hydrogen. Syngas (H2/CO= 1) flow was then initiated and maintained for either 18 hours (215~ or 6 hours (245~ At the end of the reaction period, the syngas was replaced by helium and the catalyst was allowed to cool to room temperature (RT). For TPSR, hydrogen at 25 cc/min was flowed over the sample while the temperature was raised at 35~ from RT to 800~ Hydrogen chemisorption uptakes were measured by first evacuating a reduced sample at 400~ for 1 hour at 10-4 torr followed by cooling to 100~ Hydrogen was then admitted and the isotherm was measured at 100~ Hydrogen uptakes for reduced Fe/l% A1203 and Fe/0.9% K/l% A1203 catalysts were 70 and 50 l.tmoles/gcat respectively. 3. RESULTS Samples of both K-promoted and K-free catalysts were pretreated in situ under reaction conditions in a 1:1 H2:CO synthesis gas at either (1) 215~ for 18 hours or (2) 245~ for 6 hours before TPSR with hydrogen. Based on previous CSTR reaction studies [8], these two reaction regimes were chosen to be representative of (1) steady-state reaction under mild conditions without deactivation and (2) severe reaction conditions involving significant deactivation. Figure 1 shows plots of methane evolution rate for each of the four samples. Rates are expressed as the number of l.tmoles of methane (i.e., I.tmoles of carbon reacted) per gram of catalyst per second. Several peaks are evident in each spectrum. After reaction at the lower pretreatment temperature, the TPSR spectrum for the K-free catalyst (narrow, solid line) shows distinct peaks at around 390~ and 580~ with a definite shoulder on the left of the high-tempera-

519 5.0

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Figure 1. TPSR spectra of K-free and K-promoted catalysts after pretreatment in synthesis gas (H2/CO = 1) at 215~ for 18 hours or at 245~ for 6 hours. ture peak (about 510~ There is also a small peak centered around 300~ For the same catalyst pretreated at 245~ (wide, solid line), peak shifts and changes in peak intensity are observed. The peak at 390~ doubles in size while the shoulder at 510~ shifts to a slightly lower temperature and decreases in intensity, making the peak more distinct. Additionally, the peak at 575~ shifts to a slightly higher temperature (594~ and appears as a shoulder to a new, high-temperature peak with a maximum rate at 640~ Spectra for the K-promoted catalysts are shown as dashed lines in Fig. 1. The spectrum for the promoted catalyst pretreated at 215~ reveals peak locations similar to the deactivated, K-free catalyst plus a new HT peak at about 660~ An increase in pretreatment temperature for the K-promoted catalyst causes the high-temperature peaks (> 550~ to shift to still higher temperatures. Deconvolution and analysis of individual peak positions and areas were performed for each spectrum using a modified M6ssbauer curve-fitting program employing a variable metric minimization (VMM) technique. Such sophisticated fitting techniques have not been previously used much for TPSR peak analysis, partly because, as in the case of McCarty et al. [4], peaks were quite clearly delineated with little overlap and in part because there is little published information to guide such use. For example, Bianchi and Gass [6] did not deconvolute the two overlapping peaks in their TPSR spectra containing a total of three peaks. For the analysis of multiple overlapping TPSR peaks an assumption must be made about peak shape. From the TPSR data of McCarty et al. [4], the peaks appear to be Gaussian, although some asymmetry is evident; in fact, VMM analysis (assuming a Gaussian distribution) of their two high-temperature TPSR peaks after ethylene exposure at 1073 K produced reasonably close fits. Thus, TPSR spectra from this study were fitted with several (4 to 8) Gaussian lines, and the goodness of fit (based on the chi-squared value) was used as a criterion for choosing for each spectrum the optimum number of lines with

520 positions common to the other spectra. This approach facilitated the selection of statistically significant peaks in terms of location and intensity. Figure 2 shows TPSR spectra for the Fe and Fe/K catalysts at the two different reaction temperatures (215 and 245~ individual Gaussian peak contributions for different carbon or carbide species are shown in spectrum. The heavy dashed line is the sum of the individual peaks,

Unprom, 2~ .

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100

200

300

400

500

600

700

800

Temperature

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Figure 2. TPSR spectra showing individual peak contributions from the various carbon species. directly overlaying the measured spectrum in each case. The optimized analysis for the first spectrum yielded six peaks and the analyses for the last three yielded seven peaks. The observation of such a large number of TPSR peaks, while somewhat unexpected, is nevertheless consistent

521 with observations of 7-8 distinctly different carbon and carbide types in previous studies of carbonaceous deposits on Fe, Ni, and Ru catalysts following CO/H2 or CO treatments [4-17]. Table 1 lists temperatures associated with each H2-TPSR peak maximum for Fe and Fe/K catalysts previously reacted in CO/H2 at 215 or 245-8~ Peak areas are summarized in Table 2 in terms of monolayer equivalents, based on the hydrogen chemisorption uptakes of fresh, reduced K-free (70 ~tmoles/g) and K-promoted (50 l.tmole/g) catalysts, while the corresponding fractional compositions are listed in Table 3. Monolayer equivalent totals calculated from individual peaks are within 4% of the measured carbon amounts. TPSR peaks (1 through 7) of progressively higher peak temperature are designated as t~l, ct2, 71, 72, ~, ~1' and ~2; these forms are progressively less reactive with H 2. Peak 1, designated t~l, shifts toward higher peak temperature with increasing reaction temperature and through addition of K, as evident by a scan down the 2nd column in Table 1. The amount of this carbon also increases for the same catalysts and treatments, from 0.08 to 0.75 monolayer equivalents (see Column 2 of Table 2). The position and area under Peak 2 (ct2) remain essentially constant with treatment temperature and K content, coverage only varying from 0.48 to 0.64 monolayers and peak temperatures spanning only 18~ Table 1. TPSR peak temperatures for various carbon phases on spent Fe Peak Temperature (~ Catalyst / Conditions 0~1 ~2 71 72 ~ K-free / 215~ 282 392 525 575 594 K-free / 248~ 355 398 499 594 631 K-promoted / 215~ 355 392 491 588 637 K-promoted / 245~ 393 410 502 623 676

and Fe/K catalysts.

~1

~2

605 646 659 701

654 672 713

Table 2. TPSR peak areas for various carbon phases on spent Fe and Fe/K catalysts. Peak Areas (monolayer C equivalents) Catalyst / Conditions ~ a~2 ]rl 72 ~ 81 ~52 Total K-free / 215~ 0.08 0.50 3.31 0.96 0.44 0.25 5.53 K-free/248~ 0.34 0.60 1.70 3.35 0.96 0.49 0.25 7.69 K-promoted/215~ 0.51 0.64 1.79 2.24 1.81 0.81 0.40 8.20 K-promoted/245~ 0.75 0.48 0.98 2.88 1.49 0.96 0.38 7.92 Table 3. Fractional compositions for various carbon phases on spent Fe and Fe/K Peak Areas (Fractional Composition) Catalyst / Conditions c~ ~2 71 ~2 ~ ~1 K-free / 215~ 0.01 0.09 0.60 0.17 0.08 0.05 K-free / 248~ 0.04 0.08 0.22 0.44 0.12 0.06 K-promoted/215~ 0.06 0.08 0.22 0.27 0.22 0.10 K-promoted / 245~ 0.10 0.06 0.12 0.36 0.19 0.12

catalysts.

~2 0.03 0.05 0.05

The most dramatic changes, however, occur with Peaks 3-7. For example, the relative fraction of the third peak (71), which accounts for 60% of the carbon in the K-free catalyst after reaction at 215~ decreases going down Column 4 of Table 3 to only 22% of the total carbon on the same

522 catalyst after reaction at 248~ and decreases further to 12% for the K-promoted catalyst pretreated at 245~ This decrease in fractional coverage occurs at the same time that the fractions of the remaining four peaks ('t2, 13, ~51, ~52) are increasing. It is also noteworthy that the total surface coverage of the K-promoted catalyst reacted at 215~ is significantly higher than for the K-free catalyst tested under the same conditions.

4. DISCUSSION Analysis of the temperature-programmed surface reaction spectra in this study reveals 6 to 7 different carbon species on the spent catalysts (Tables 1 and 2 and Fig. 2), more than reported in any one previous study of carbon species deposited by CO or CO/H E reactions on iron. Since the analysis was based on computerized fitting of the data with Gaussian lines, it is appropriate to address the validity of assuming a Gaussian line shape for the different carbon forms. This assumption is probably reasonably valid for surface carbons (e.g. t~ and [~ carbons) but more approximate in the case of bulk phases (e.g. carbides), since TPSR involves diffusion of C atoms from the bulk to the surface until the bulk is exhausted--resulting in a pronounced high-temperature cutoff [18]. Furthermore, above 700~ methane formation from carbon and hydrogen on metal surfaces is reversible [ 18], possibly leading to a tailing of the reaction peak at higher temperatures. Such effects, if present, would affect the uncertainty in the reported peak positions. Nevertheless, it was possible to make peak assignments to various carbon species consistent with data from this and previous studies as discussed below. Designation of the first TPSR peak at about 300~ as Ca, an atomic surface carbon, is consistent with the work of McCarty et al. [4] and Bianchi et al. [6,7]. In this study the temperature at peak maximum for Ca and its surface coverage are observed to increase as reaction pretreatment temperature increases and with the addition of potassium. Absolute rates of hydrocarbon production (or conversion of CO or H2) apparently correlate with Ca coverage, i.e., rate increases with increasing temperature and increasing K content, in spite of the greater extent of deactivation taking place at higher temperatures [8]. Based on this work steady-state coverages of Co~ during CO hydrogenation on Fe range from 0.08 to 0.75 monolayer equivalentsmhigher than 0.05 ML observed for single crystal Ni [ 16] but consistent with monolayer coverages observed for CO hydrogenation on iron foils [9,10]. Addition of potassium to iron catalysts has been shown to strengthen the metal-C bond for adsorbed CO and the rate of CO dissociation [3,19]. Thus, a shift to higher TPSR temperature and Co~ coverage with addition of potassium is consistent with an increased C-Fe bond strength. The second peak (oh), centered at 400~ is unique in that its temperature at maximum rate does not shift significantly with increasing reaction temperature or addition of K (i.e., peak temperature shifts only 18~ compared with 37-111~ for the other peaks). It is also the only peak that maintains a relatively constant coverage (only 25% variation between extremes) while coverages of other peaks increase or decrease by at least a factor of three. McCarty et al. [4] in their study of carbonaceous species on alumina-supported nickel observed peaks similar in location to Peaks 1 and 2. Both low-temperature peaks, designated as ~1 and t~2 species, were attributed to chemisorbed carbon, possibly on different sites (e.g. terrace versus steps or ledges). The first of the two peaks was populated at "less than" one monolayer while the second was "up to" one monolayer. In view of the close correspondence between this study on iron and the previous study on Ni, it is logical to assign Peaks 1 and 2 of this study to atomic surface carbon and designate them ~1 and ix2.

523 Based on data from this study, deactivation is not expected on a K-free iron catalyst during FT synthesis at 215~ However, as conditions become more extreme, e.g. temperature increases or K is added to iron thereby increasing the surface coverage of carbon, a decrease in activity with time due to the accumulation of less active carbons and/or carbides is expected to become more pronounced. However, not all forms of carbon cause loss of catalyst activity [2]. During methanation or steam-reforming on Ni catalysts loss of activity can be attributed to (1) the accumulation on the nickel surface of films consisting of amorphous carbon, graphitic carbon, and/or coke and/or (2) the conversion of Ni to Ni carbide [2]. Since the surface chemistry of Fe is similar to that of Ni, the similar deactivation of Fe surfaces by carbide formation and/or graphitic films is a distinct possibility. However, one must be cautious in making such comparisons, since some carbides of iron (e.g. X and e') are apparently active phases for CO hydrogenation [ 11 ], while in Ni catalysts only the metal is active [2]. It would be tempting to assign Peaks 3 and 4 to amorphous C a carbon or coke, since these species of intermediate reactivity can be formed on metal surfaces by polymerization of CH x and C a [2,18]. However, in the case of the Ni-carbon system, Ni carbide has a higher reactivity (lower TPSR peak temperature) than amorphous C a. If the order of reactivities for the carbide and the amorphous carbon film formed during CO hydrogenation on Fe is similar to that for Ni, Peaks 3 and 4 might be assigned to iron carbides. Mtssbauer data from Niemantsverdriet et al. [11] support this hypothesis. Indeed, their in situ spectra after 24 hours of reaction under conditions similar to those of this study indicate that x-Fe5C2 and E'-Fe2.2C account for 65 and 35% of the iron, respectively. Areas of Peaks 3 and 4 for the K-free Fe catalyst after reaction at 215~ are present in toughly the same ratio (3.31 and 0.96 monolayer equivalents); moreover, these two peaks account for 77% of the carbon on the catalyst. Given that these catalysts have dispersions around 1%, it is expected that iron in the bulk of the metal crystallites would provide a significant storage volume for carbon as carbide (i.e., up to about 50 layers). Accordingly, Peaks 3 and 4 (at about 500 and 600~ are assigned to iron carbides. Before finalizing these assignments for carbides, it would be prudent to determine if the trends observed in the peak areas are consistent with what is expected for carbide formation and transformation. In the case of the K-free catalyst, a decrease in Peak 3 and an increase in Peak 4 are observed as temperature is increased. Enrichment of the bulk with carbon would account for the shift from the less carbon-rich Hiigg carbide (x-FesC2) to the more carbon-rich e-carbide (e'-Fe2.2C). This is consistent with the observed increase in total carbon content of the catalysts with increasing temperature. The observed changes in the areas of Peaks 3 and 4 for the potassium-promoted catalyst samples are similar to those for the K-free samplesmnamely, the area of Peak 3 decreases while that of Peak 4 increases with increasing temperature. Thus, evidence for assigning Peaks 3 and 4 to two iron carbides, while not conclusive, is at least consistent with previous work and with trends observed from the data presented in this paper. Shifts in peak temperature for both peaks are toward lower temperatures at lower coverages. Peak 5 can be assigned on the basis of previous work with Ni and Ru catalysts to amorphous coke or C a. The formation of a hydrogen-containing carbon species during CO hydrogenation has been verified on Ru [18] and iron catalysts [13]. While Dwyer observed no "polymethylene," as he calls it, on K-free iron after reaction for 8 hours at 7 atm, H2/CO = 3 and 540 K, identical treatment of a catalyst promoted with potassium produced a multilayer coverage of C a. Accordingly, the area of Peak 5 should increase with K addition. Indeed, upon the addition of K

524 the area of Peak 5 (after reaction at 215~ increases fourfold to nearly two monolayer equivalents. Exposure of a Ni/A1203 catalyst to ethylene at high temperatures produces a similar carbonaceous material of up to four monolayers [4]. Thus, coverages exceeding one monolayer for K-promoted catalysts are consistent with the assignment of Peak 5 to C~. The assignment of Peaks 6 and 7 (~51 and 82) to layered graphitic carbons is reasonably unambiguous. The fractional contribution of both peaks increases with increasing reaction temperature and hence extent of deactivation [Table 2, Figure 2; 8,14,15]. In fact, Jackson et al. [ 15] have correlated decreasing FT synthesis activity of preciptated iron catalysts with increasing content of graphitic carbon. Evidence of graphitic carbons has been provided by NMR for ruthenium catalysts [18] and recently by transmission electron microscopy (TEM) for iron FT catalysts in the form of films surrounding iron carbide particles [ 14,15]. Consistent with the high peak temperatures observed in H2-TPSR for these species in this and another recent study [ 15], graphitic carbon, with its highly ordered structure, would be expected to be more highly resistant to reaction with hydrogen than all other carbon species. The graphite referred to here cannot, however, be pure, crystalline graphite because the extremely low reactivity of highly crystalline graphite would require much higher temperatures than could be reached in our TPSR experiments. Accordingly, it is appropriate to differentiate graphitic carbon from crystalline graphite. Crystalline graphite occurs naturally in two forms: alpha (hexagonal) and beta (rhombohedral) [20]. Interconversion between the two forms can occur for given treatments. Graphite densities can vary from specific gravities of 1.9 to 2.3. Therefore, graphite represents a range of carbon species that vary in density, degree and type of crystallinity and reactivity. Size of the planar graphite phases as well as the extent of the three-dimensional structure may also affect reactivity. Based on their different but nevertheless low reactivities, Peaks 6 and 7 (81 and ~52) can be assigned to films of "moderately-oriented" and "more highly-ordered" graphitic carbons respectively; however, neither would have the high crystallinity and high density of threedimensional, single-crystal graphite. Are there other possible forms of carbon or carbides that account for the data of this study? While filamentous carbons have been observed on iron foils, oxidized iron foils (FeO) and unsupported Fe203 catalysts following exposure to CO or hydrocarbons at high temperature (i.e., > 450-500~ [12], they would probably not be formed under Fischer-Tropsch conditions (200300~ In fact, Datye et al. [14,15] find no evidence of filamentous carbon formation on precipitated iron catalysts tested for FT synthesis. Hence, it is not likely that filamentous carbon was present under the conditions of this study. While the high temperature delta peaks might be assigned to a carbide of high stability (hence low reactivity), there was no evidence from Mrssbauer spectra of the catalyst of this study that such carbides were formed under the conditions of this study leading to deactivation. Comparison of the data in Tables 1-3 for catalysts tested at 215 (with no loss of activity) and 248~ (with significant loss of activity) respectively indicates that deactivation of the iron catalysts in this study is qualitatively associated with (1) the transformation of high-activity Z-carbide to the less active e'-carbide and (2) the formation of graphitic carbons ([i1 and ~52) in the form of films or layers. While we attempted as part of this study to find a quantitative relationship between different carbon forms and the extent of deactivation of monolithic iron catalysts during FT synthesis at 10 atm in a Berty reactor, we were unsuccessful. The major problem was exposure of the spent

525 catalyst to air in the process of transferring it from the reactor to the TPSR cell. In the resulting TPSR spectrum no ~1 peak could be identified, while the intensities of Peaks 2 and 3 (c~2 and Yl) were significantly reduced relative to those observed for samples treated in situ under identical reaction conditions. The higher temperature (i.e., less reactive) carbons were not noticeably affected. 5. CONCLUSIONS a. Seven peaks have been identified from TPSR spectra of iron catalysts after PT synthesis at 215250~ 1 atm, and H2/CO = 1. The peaks are assigned to two atomic surface carbon species, ~1 and ~2, two Fe carbide phases, Y1 and Y2, corresponding to Z - and e'-carbides, a hydrogen-containing coke, 13, and graphitic carbons, 81 and 82. b. Catalyst activity for FT synthesis on iron catalysts can be correlated with the coverage of atomic, surface carbon (Otl) species. c. Deactivation of the iron catalysts under the conditions of this study is qualitatively associated with (1) the transformation of high-activity z-carbide to the less active e'-carbide and (2) the formation of graphitic carbons (81 and 82) in the form of films or layers.

REFERENCES

1.

M.E. Dry, The Fischer-Tropsch Synthesis, in J.R. Anderson and M. Boudart (Eds.), Catalysis: Science and Technology, Vol. 1, Springer-Verlag, Heidelberg, 1981, pp. 159-255. C.H. Bartholomew, "Carbon Deposition in Steam Reforming and Methanation," Catal. Rev.Sci. Eng. 24 (1982) 67-112. H. Arakawa and A.T. Bell, "Effects of Potassium Promotion on the Activity and Selectivity of Iron Fischer-Tropsch Catalysts," Ind. Eng. Chem., Proc. Des. Dev., 22 (1983) 97-103. J.G. McCarty, P.Y. Hou, D. Sheridan, and H. Wise, "Reactivity of Surface Carbon on Nickel Catalysts: Temperature-Programmed Surface Reaction with Hydrogen and Water," in Coke Formation on Metal Surfaces, eds. L.G. Albright and R.T.K. Baker, American Chemical Society, Washington D.C., 1982, p. 253. R.M. Bowman and C.H. Bartholomew, "Deactivation by Carbon of Ru]AI203 During CO Hydrogenation," Appl. Catal. 7 (1983) 179-187. D. Bianchi and J.L. Gass, "Hydrogenation of Carbonaceous Adsorbed Species on an Iron/Alumina Catalyst," J. Catal. 123 (1990) 310. H. Ahlafi, C.O. Bennett, and D. Bianchi, "Isothermal Hydrogenation of Carbonaceous Adsorbed Species on an Iron Catalyst," J. Catal. 133 (1992) 83. S. A. Eliason, Deactivation by Carbon of Unpromoted and Potassium-Promoted Iron FischerTropsch Catalysts, Ph.D. Dissertation, Brigham Young University, April, 1994. D.J. Dwyer and G.A. Somorjai, "Hydrogenation of CO and CO 2 over Iron Foils: Correlations of Rate, Product Distribution, and Surface Composition," J. Catal. 52 (1978) 291. 10. H.J. Krebs and H.P. Bonzel, "A Model Study of the Hydrogenation of CO over Polycrystalline Iron," Surf. Sci. 88 (1979) 269. 11. J.W. Niemantsverdriet, A.M. van der Kraan, W.L. van Dijk, and H.S. van der Baan, "Behavior of Metallic Iron Catalysts During Fischer-Tropsch Synthesis Studied with M6ssbauer Spectroscopy, X-ray Diffraction, Carbon Content Determination, and Reaction Kinetic Measurements," J. Phys. Chem. 84 (1980) 3363. .

,

.

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.

.

.

.

526 12. R.T.K. Baker, D.J.C. Yates, and J.A. Dumesic, "Filamentous Carbon Formation over Iron Surfaces," in Coke Formation on Metal Surfaces, in Coke Formation on Metal Surfaces, eds. L.G. Albright and R.T.K. Baker, American Chemical Society, Washington D.C., 1982, p. 1. 13. D.J. Dwyer, "Iron Fischer-Tropsch Catalysts: Surface Synthesis at High Pressure," Prep. ACS Div. Pet. Chem. 29 (1984) 715. 14. M.D. Shroff, D.S. Kalakkad, K.E. Coulter, S.D. Kohler, M.S. Harrington, N.B. Jackson, A.G. Sault, and A.K. Datye, "Activation of Precipitated Iron Fischer-Tropsch Catalysts," J. Catal. 156 (1995) 185. 15. N.B. Jackson, A.K. Datye, L. Mansker, R. J. O'Brien and B.H. Davis, "Deactivation and Attrition of Iron Catalysts in Synthesis Gas," in Catalyst Deactivation 1997, eds. C. H. Bartholomew and G.A. Fuentes, Elsevier, 1997. 16. D.W. Goodman, R.D. Kelley, T.E. Madey and J.M. White, "Measurement of Carbide Buildup and Removal Kinetics on Ni(100)," J. Catal. 64 (1980) 479. 17. J.G.McCarty, private communication, May 1997 18. T.M. Duncan, P. Winslow, and A.T. Bell, "The Characterization of Carbonaceous Species on Ruthenium Catalysts with 13C Nuclear Magnetic Resonance Spectroscopy," J. Catal. 93 (1985) 1. 19. J.L. Rankin and C.H. Bartholomew, "Effects of Potassium and Calcination Pretreat-ment on Adsorption and Chemical/Physical Properties of Fe/SiO2," J. Catal. 100, (1986) 533-540. 20. R.C. Weast, ed. CRC Handbook of Chemistry and Physics. CRC Press, Boca Raton, 1980.

~ Elsevier Science B. V. All rights reserved Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

527

D e a c t i v a t i o n o f Iron-based Catalysts for Slurry Phase F i s c h e r - T r o p s c h Synthesis Ajoy P. Raje, Robert J. O'Brien, Liguang Xu and Burtron H. Davis Center for Applied Energy Research, University of Kentucky 3572 Iron Works Pike, Lexington, KY 40511, USA.

Deactivation rates and aged catalyst properties have been investigated as a function of time on stream for iron-based Fischer-Tropsch catalysts in the presence/absence of potassium and/or silicon. There is a synergism in activity maintenance with the addition of both potassium and silicon to an iron catalyst. The addition of silicon appears to stabilize the surface area of the catalyst. Catalysts containing only iron or added silicon with or without potassium consist mainly of iron oxide at the end of the run. However, iron carbides are the dominant phase of the iron catalyst with added potassium alone. Catalyst surface areas increase slightly during synthesis. The bulk phase of the catalyst does not correlate to the catalyst activity. The partial pressure of water in the reactor is lower for potassium-containing catalysts and is not a reliable predictor of catalyst deactivation rate.

1. INTRODUCTION The Fischer-Tropsch Synthesis (FTS) converts synthesis gas (a mixture of CO and I-I_,)to hydrocarbons. Iron-based catalysts lose activity with time on stream (TOS). The rate of deactivation is dependent on the presence/absence of promoters such as potassium and/or binders such as silica [ 1.2]. Several possible causes of catalyst deactivation have been postulated [3 ]: (i) Sintering. (ii) Carbon deposition, and, (iii) Phase transformations. With respect to phase transformations, there is considerable disagreement whether the active phase for the FTS is iron oxide or carbide [4,5]. In addition, certain reactor conditions, such as a high partial pressure of water, are known to cause a decline in activity [6]. There were two major objectives of this study. Firstly. the effects of the addition of potassium and/or silicon on catalyst deactivation rates and changes in catalyst properties with TOS were investigated. Secondly, the possible causes of catalyst deactivation were examined by following aged catalyst properties and reactor conditions as a function of TOS for each catalyst. The FTS was carried out in a continuous-flow stirred slurry reactor to ensure uniformity in catalyst aging and reactor conditions throughout the reactor. The aged catalyst properties examined as a function of TOS were total surface areas, carbon deposits and phase transformations.

528 2. EXPERIMENTAL Four precipitated iron-based catalysts were used. The first catalyst consisted of only iron. The other catalysts contained either added potassium, added silicon or both. The catalysts were designated in terms of the atomic ratios as: 100Fe, 100Fe/3.6Si, 100Fe/0.71K and 100Fe/3.6Si/0.Tl K. The catalysts were prepared by continuous precipitation from iron (Ill) nitrate and concentrated ammonium hydroxide. For silica-containing catalysts, a colloidal suspension of tetraethyl ortho silicate was mixed with the iron nitrate solution prior to precipitation. Potassium was added to the catalysts in the form of potassium tertiary butoxide during the loading of the FTS reactor. The FTS was carried out in a 1 liter continuous flow stirred tank slurry reactor. The vapor phase products exiting the reactor passed through two traps in series. The first trap was maintained at 100~ while the second was maintained at 0~ The second trap served to condense almost all of the water present in the product stream. Further details of the reaction system and product analysis (both on- and off-line) have been reported previously [7]. A known amount (--70 g) of the catalyst was charged to the reactor along with 300 g of a hydrocarbon oil supplied by Ethyl Corp. (carbon number range of about ~5 - C30). The catalyst was pretreated with a continuous flow of CO at 2.0 NL/hr-gFe at 270~ 175 psig for 24 hours. Subsequently, the FTS was carried out at 270~ 175 psig with a synthesis gas having an H_,/CO ratio of 0.7 at a flowrate of 3.4 NL/hr-gFe. Both during pretreatment and synthesis, small amounts of catalysts (0. l to 0.2 g per sample) were removed from the reactor at various TOS. The aged catalyst samples were Soxhlet-extracted with xylene prior to analysis. These aged catalyst samples were examined to determine BET surface areas, elemental carbon and hydrogen contents and bulk catalytic phases. Elemental.carbon and hydrogen contents were determined using a Leco CHN analyzer wherein the catalyst sample was placed in a furnace at 1350~ under flowing oxygen. All of the carbon and hydrogen in the sample was converted to CO2 and water repectively which were then measured. Bulk catalyst phases were detemained by two methods: X-ray diffraction (XRD) and Mossbauer spectroscopy (MS). X-ray diffraction pattems were obtained using a Phillips APD Xray diffraction spectrometer equipped with a Cu anode and a Ni filter operated at 40 kV and 20 mA (CuKtz=l.5418 A).Mossbauer spectra of the catalyst samples were obtained with a constant acceleration spectrometer using a radioactive source consisting of 50 to 100 mCi of 57Co in a Pd matrix.

3. RESULTS AND DISCUSSION The conversion of synthesis gas is a measure of the FTS activity of a catalyst. The catalysts studied show varying rates of decline in synthesis gas conversion with TOS (Figure l ). Note that the conversions for different catalysts are compared at the same synthesis gas space velocity: which is defined as nonual liters of synthesis gas per gram of iron in the reactor. Tile addition of silica alone (100Fe/3.6Si) slows the decline in FTS activity as compared to a catalyst containing only iron (100Fe). In contrast, the addition of potassium alone increases the deactivation rate as compared to a catalyst containing only iron (100Fe). The catalyst containing both silica and potassium (100Fe/3.6Si/0.71K) exhibits the lowest deactivation rate. Hence, there is a synergisna in the maintenance of FTS activity with the addition of both silica and potassium.

529 The surface areas of the catalysts are substantially decreased during pretreatment (Figure 2). During synthesis, however, the surface area of the catalysts increases slightly. It is postulated that this slight increase may be due to the deposition of porous carbon on the catalyst surface: Catalysts containing silica exhibit higher surface areas on preparation, and during pretreatment and synthesis. Hence, silica appears to stabilize the surface area of the catalysts. Since the total surface area increases during synthesis, the presence/absence of sintering during FTS cannot be deduced from BET surface area measurements.

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The catalyst sample is exposed to air for a short time during XRD and MS. However, using similar procedures Mossbauer spectra have been previously obtained showing 97% of the iron in a catalyst as iron carbide and substantial (36%) amounts of metallic iron in another catalyst [8]. Hence, re-oxidation of catalyst phases during XRD and MS is negligible. The results from XRD indicate the following events occurring during pretreatment and synthesis: The catalyst as charged into the reactor is in the form of an oxide and/or an oxyhydroxide. During pretreatment in CO, the catalyst is rapidly converted to F~O4 which is subsequently partially converted in a slow step to a mixture of carbides (FesC 2 and F%2C ) as shown in Figure 3. During synthesis, the relative amounts of carbide and oxide change dependent on the catalyst composition. For the catalyst containing only potassium, the catalyst at the end of the run consists of mainly carbides (Figure 3) while for the other catalysts the oxide is the dominant phase at the end of the run (example of 100Fe given in Figure 3). These qualitative observations are consistent with quantitative results obtained by MS. Since the catalyst containing potassium alone has the highest deactivation rate. it is tempting to conclude that the iron oxide is the active phase, or at least more active than the carbide, for the FTS. However this is not true as the bulk composition of the catalyst, as shox~aa below, does not con'elate to the catalyst activity. Figure 4 shows quantitative results of the bulk composition of the catalyst obtained by MS as a function of TOS for the catalyst with the lowest

530 deactivation rate (100Fe/3.6Si/0.71K). The results are similar to those for the catalyst containing only iron but are different from those obtained for the catalyst containing iron and potassium but no silicon. The bulk composition of the catalyst changes greatly with TOS. At small TOS, iron carbide is the dominant phase while at long TOS the catalyst consists predominantly (80%) of iron oxide. Note, however, that synthesis gas conversion for this catalyst varied only slightly (between 80 and 90%) with TOS. Hence, the particular active phase of the catalyst (oxide or carbide) cannot be deduced from bulk catalyst compositions.

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Figure 3. X-ray difl~ction results for two catalysts (100Fe and 100Fe/0.71 K) during pretreatment and at the end of synthesis ((1) Fe304, (2) x-FesC2, (3) e'-Fez2C). u. ,,J

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Figure 4. Bulk composition of catalyst: 100Fe/3.6Si/0.71K from M6ssbauer spectroscopy as a function of time on stream. Negative TOS denote catalyst pretreatment; positive TOS denote Fischer-Tropsch synthesis. S.P. denotes super paramagnetic.

531 The carbon measured as a result of high temperature oxidation can be due to either Cl-Ix species on the catalyst surface, or the oxidation of iron carbides, or the oxidation of carbon deposits. The contribution of CH~ species can be estimated by elemental analysis of hydrogen and assuming the value of x to be 2. The remaining carbon measured on the catalysts (Figure 5) ranged from 10 to 30 wt.%. Note that if all the iron in the catalyst were converted to carbide the carbon amount would be 8 to 9 wt.% assuming the iron carbides are FesC., and F~_,C as shown by XRD and MS. Hence, substantial amounts of carbon deposits were present on the catalyst surface. The addition of potassium to the catalyst containing only iron (100Fe/0.71 K) increased the amount of carbon measured as well as its rate of increase. However, tbis may have been due to the high amount of carbide formed on this catalyst during synthesis as shown by XRD and MS. The catalyst with the lowest deactivation rate (100Fe/3.6Si/0.71 K)contained substantial amounts of carbon at the end of the run even though its activity was quite stable. Recall that this catalyst consisted mainly of iron oxide so that the carbon measured consisted mainly of carbon deposits. Hence, the amount of carbon deposits on the catalysts does not correlate with the rate of deactivation. As stated previously, almost all of the water present in the product stream is condensed in a trap downstream of the reactor maintained at 0~ It is further assumed that the uncondensed gases leaving the trap are saturated with water. The partial pressure of water is then calculated by measuring the total flow rate of the uncondensed gases and the composition of the other products and unconverted reactants. This procedure yields oxygen component balances of greater than 96%. The value of the partial pressure of water in the reactor depends on the presence/absence of potassium promoter (Figure 6). The partial pressure of water is low for

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532 catalysts containing potassium and decreases slightly with TOS. In contrast, the partial pressure of water is higher for catalysts without potassium and increases with TOS. Low water partial pressures for potassium-containing catalysts may be due to the higher rates of the water gas shift (WGS) reaction observed over these catalysts. A high partial pressure of water is known to deactivate the catalyst [6]. However, the results shown in Figure 6 indicate that the value of the partial pressure of water does not correlate to the rate of catalyst deactivation. For instance, both the catalyst with the highest and lowest deactivation rates (100Fe/0.71K and 100Fe/3.6Si/0.71K respectively) exhibit similar partial pressures of water. It has been postulated that changes in the amount of iron oxide on the catalyst surface (as opposed to the bulk followed by XRD and MS) can be followed by a comparison of the ratios of the rates of the FTS to that of the WGS with TOS [9]. It is postulated [9] that iron oxide is more active than iron carbides for the WGS than for the FTS and hence, a decrease in the ratio of the rate of the FTS to the rate of the WGS indicates an increase in the amount of iron oxide on the catalyst surface. However, none of the catalysts examined in this study show a decline in the reaction rate ratio with TOS (Figure 7). For catalysts containing potassium, the reaction rate ratio changes negligibly with TOS, while the reaction rate ratio actually increases with TOS for the catalysts without potassium. Further, the increase in the reaction rate ratio with TOS for catalysts without potassium may not be due to changes in surface phases. We have earlier shown [10] during kinetic studies that the ratio of the reaction rate of the FTS to that of the WGS increases with residence time (not TOS) or synthesis gas conversion for iron-based catalysts. Hence, the increase in the reaction rate ratio observed for two of the catalysts studied may be a consequence of decreasing synthesis gas conversion (due to deactivation) rather than a change in surface catalyst phases. o

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Time on Stream (hr)

Figure 7. Ratio of the reaction rates of the Fischer-Tropsch synthesis to the water-gas shift reaction as a function of time on stream.

4. CONCLUSION Increasing and decreasing deactivation rates of iron FTS catalysts are observed upon the addition of potassium and/or silicon respectively. There is a synergism in the maintenance of activity with the addition of both potassium and silicon leading to a low deactivation rate.

533 Changes in catalyst properties and reactor conditions with TOS depend upon the presence/absence of potassium and silicon. The addition of silicon appears to stabilize the surface area of the catalyst both on preparation and during pretreatment and synthesis. At the end of pretreatment, all of the catalysts studied consist of a mixture of iron oxide (Fe304) and iron carbides (FesC2 and Fez2C). During synthesis, the catalyst containing potassium alone is converted predominantly to iron carbides at the end of the run. The other catalysts, however, decline in the amount of iron carbide and increase in the amount of iron oxide during synthesis. The addition of potassium alone to an iron catalyst substantially increases the amount of carbon (from carbon deposits and carbides) measured by high temperature oxidation of aged catalyst samples. The partial pressure of water in the reactor is low and decreases with TOS for potassium-containing catalysts while it is higher and increases with TOS for the catalysts without potassium. Possible causes of catalyst deactivation have been investigated in this study. The surface areas of aged catalyst samples increase with TOS during synthesis possibly due to porous carbon deposits. Hence, information from total surface area on sintering as a possible mechanism of deactivation is inconclusive. Measurement of the amount of carbon deposits on the catalysts by high temperature oxidation is complicated due to the contribution from iron carbides. The bulk composition of the catalyst having the lowest deactivation rate changes greatly during synthesis whereas the synthesis gas conversion changes negligibly. This implies that the catalyst bulk composition does not correlate to the rate of activity decline. The value of the partial of water as well as its variation with TOS are similar for the catalysts with the highest and lowest deactivation rates. Hence, water partial pressure is not a reliable predictor of the catalyst deactivation rate.

REFERENCES 1. D.B. Bukur, D. Mukesh and S.A. Patel, Ind. Eng. Chem. Res., 29, 194 (1990). 2. D.B. Bukur, X. Lang, D. Mukesh, W.H. Zimnlerman. M.P. Rosynek and C. Li, Ind. Eng. Chem. Res., 29. 1588 (1990). 3. M.E. Dry, in J.R. Anderson and M. Boudart (Eds.), Catalysis-Science and Technology, Vol. 1,159-255, Springer, Berlin (1981). 4. J.P. Reymond, P. Meriaudeau and S.J. Teichner, J. Catal., 75, 39 (1982). 5. G.B. Raupp and W.N. Delgass, J. Catal., 58, 361 (1979). 6. C.N. Satterfield. R.T. Hanlon, S.E. Tung, Z. Zou. G.C. Papaefihymiou, Ind. Eng. Chem. Prod. Res. Dev., 25,407 (1986). 7. L. Xu, S. Bao. R.J. O'Blien, D.J. Houpt and B.H. Davis, Fuel Sci. Tech. Int., 12. 1323 (1994). 8. B. Jager and R. Espinoza, Catalysis Today, 23.17 (1995). 9. A.P. Raje and B.H. Davis, Catalysis Today, (accepted for publication)..

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9 Science B.V. All rights reserved Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

535

Effects of Reduction and Regeneration Conditions on the Activity of CuO-ZnO Catalysts C.E. Quincoces a, N.E. Amadeo b, M.G. Gonz~ileza a (UNLP-CONICET)- 47 N 257- (1900) La Plata-Argentina b PINMATE, Dpto de Industrias, Fac.Cs Exactas y Naturales, U.N. Bs. Aires. Argentina The influence of thermal and hydrothermal treatments and reduction processes on the activity of CuO-ZnO catalysts was studied. Temperature-programmed reduction (TPR) was used in the study of catalyst activation with different reduction mixtures. A mixture of H 2 and N 2 in the range of 500 to 573 K was the most effective reducing agent. High temperature, water vapor and a high partial pressure of hydrogen enhance Cu sintering and Cu-Zn alloy formation with a decrease in catalytic activity. XRD analysis of reactivated catalysts showed that their deactivation can be related to the formation of surface spinel species. Reoxidation at 623 K and subsequent reduction of the aged catalyst led to metal redispersion but not catalyst regeneration. 1.

INTRODUCTION

Catalysts based on CuO-ZnO are of great industrial interest because they exhibit high activity for the low temperature-pressure methanol synthesis and the water-gas-shift reactions. It is known that the activity and useful life of catalysts depend mainly on the activation process and the thermal history they experience during the operation. In the low temperature water gas shift (LTWGS) process, prior to reaction, the catalyst is activated by gas reduction to convert copper oxide into metallic copper [ 1]. It has been observed that reduction conditions affect the activity and the stability of Cu-ZnO catalysts. For instance, sintering and formation of alloys must be avoided in the reduction step because they deactivate the catalyst [2-3] for the water-gas-shift reaction. On the other hand, the possibility of reactivating aged Cu-ZnO catalyst must be considered. To restore the activity and stability of the catalyst the regeneration process must redisperse the metal while ensuring appropriate thermal stability. The most suitable method to produce these changes in the deactivated catalyst is to oxidize the copper at high temperature (623 K). The aim of this work was to analyze the structural changes accompanying deactivation and reactivation of Cu-ZnO-AI203 catalyst used in the low temperature water-gas-shift reaction. 2.

EXPERIMENTAL

A commercial LTWGS catalyst was studied in this work, using fresh and aged samples to analyze the effect of reaction conditions. Aged samples were taken from an industrial reactor after 8 months of operation at 473 K and 1.5 MPa. The fresh catalyst had the following characteristics: composition by weight percent of 43% CuO, 47% ZnO, 9.7% A120 3, a specific area of 41 m 2 g-1 and a pore volume of 0.09 c m 3 g-1. The reduction studies were performed using temperature-programmed reduction (TPR) in flow equipment, using a heating rate of 5 K/min until reaching the final temperature (500-773K)

536 which was maintained for two hours, so as to avoid abrupt temperature changes during reduction. The catalyst aged in the commercial reactor was oxidized in the laboratory in air at 623 K for 1 h, and then treated similarly. Catalyst composition and crystallite size were determined by x-ray diffraction analysis, the latter being determined by x-ray line-broadening [4-5]. Cu-Zn alloy formation was followed by the evolution of the Cu [ 111] lattice constant. The experimental measurements were carried out in a Philips PW 1740 x-ray diffractometer using CuK~x radiation at 40 kV and 20 mA. Specific surface area and pore volume distribution were measured by nitrogen adsorption in an Accusorb 2100E Micromeritics adsorption analyzer. The data were interpreted using the BET equation, assuming a cross-sectional area of 16.2/~2 for N 2. Catalytic activity was determined in a tubular packed bed isothermal reactor at 500 K and 1 atm. A gas mixture was fed to the reactor at 350 cm 3 min -1(CO: 3%, H20: 26% H2:48% N2:23% v/v); the catalyst weight was 0.04 g with a particle size of 0.177-0.250 mm. Reactants were analyzed by gas chromatography, using a thermal conductivity detector. Two packed columns were employed to analyze the reaction mixture. One was packed with 5A molecular sieve to separate hydrogen, nitrogen and CO, while CO 2was analyzed in a column packed with Porapak Q. Absence of diffusional control was experimentally verified by measuring the reaction conversion with catalyst particles of various sizes. The bed was diluted (D%=10% v/v) with inert particles to provide isothermal conditions. 3.

R E S U L T S AND D I S C U S S I O N

During the highly exothermic reduction step, LTWGS catalysts can lose activity [5] by sintering of the metallic phase, formation of a Cu-Zn alloy or changes in the textural characteristics of the support induced by temperature, the presence of steam and the reducing mixture composition For this reason, in this study we applied different thermal and hydrothermal treatments to analyze their effects on the structural and textural properties of the catalysts. Part of the catalyst was reduced in a N2-H2 (9:1) stream at programmed temperatures increasing linearly up to a final value between 503 and 773 K, for 1 h. A second portion was heated to 503 K in a pure hydrogen flow (50 cm 3 min 1) or in a H2-H20 mixture (prepared by introducing 1 gl pulses of water every 6 min. for 1 h). Table 1 lists BET surface areas and pore volumes of fresh and aged catalysts. These parameters are modified by both high temperature and reaction under industrial conditions. Surface area and porosity of the fresh catalyst were lowered as reduction temperature reached 773 K. However, losses were higher in samples taken from the industrial reactor (aged catalyst). The textural characteristics were not modified by oxidation at 623 K. Table 1 Effect of reduction conditions on the textural characteristics Catalyst Reduction Reduction mixture temperature (K) Fresh

Spent Reoxidized at 623K

Specific area m2

Pore volume cm 3 ~-1

~-1

N2-H 2 N2-H 2 N2-H 2 N2-H20 Ha

503 573 773 503 503

41.0 41.6 32.3 48.0 45.8

0.090 0.108 0.054 0.111 0.095

N2-H 2 N2-H 2 N2-H 2

503 503 773

14.4 18.2 12.8

0.028 0.042 0.036

537 The pore size distribution maxima (data not shown) in the flesh catalyst were observed at pore radii of 15 and 30/~, whereas in aged catalysts, pore distribution is shifted toward higher pore radii. While reduction with pure hydrogen at higher temperature shifts the distribution to higher pore radii for flesh and aged catalysts, the presence of water vapor in the reducing mixture favors smaller pore formation. These results show that the porous structure is more affected under industrial reaction conditions than by reduction. To elucidate the effects of reduction, reoxidation and industrial reaction conditions (temperature, presence of steam and composition of reactant mixture) on the structural properties, the Cu/ZnO-A1203 catalysts were characterized by XRD and TPR. Fig. 1 shows the TPR profiles of the different samples. For comparison purposes, the reduction profile of CuO was also included, which shows a [ single sharp peak at 503 K. The profile for the fresh CuO-ZnO-ml203 catalyst is also characterized by a single symmetrical sharp peak whose maximum is observed at a lower temperature (about 473 K). In the aged catalyst, the main peak is at 473 K but a shoulder on the lower temperature side (curve b) is observed. This shoulder becomes more pronounced for the aged catalyst reoxidized at 623 K. Its TPR profile is = characterized by two signals, a peak at about 500 K o and a shoulder at 473 K. These results show that a single reduction process occurs in the fresh catalyst cawhile two reducible copper species exist in the aged E and reoxidized samples. =tn From XPS studies Fierro et al. [6-8] have reported the presence of two different copper species in Cu/Zn oo catalyst. The high temperature peak was assigned to a small homogeneously dispersed copper species =: interacting with the ZnO with a reactivity similar to that of CuO, while the other peak was assigned to a more reactive copper species interacting strongly with the surface of ZnO particles. Thus, it appears that the ZnO plays the role of a support. Therefore, from our TPR results for the catlyst regenerated by oxidation at 623 K we infer that different copper compounds are produced during the reactivation step, each with a 273 473 K particular reactivity to hydrogen. In turn, the appearance of a TPR peak at high temperature indicates that a less reducible copper species is formed. The TPR peaks for different Cu-ZnO catalysts (Fig. Figure 1. TPR patterns (a) flesh 1) are about 30 K lower than the value found for pure catalyst, (b) aged catalyst, (c) aged CuO. These results suggest that ZnO promotes copper catalyst reduced at 773 K after oxidation reducibility, which may be attributed to a H 2 spillover mechanism. The TPR results show that copper reactivity toward reduction in these catalysts depends on the nature of Cu species, CuO particle size and Cu-Zn interaction. In order to determine Cu speciation, Cu crystal size and the extent of the Cu-Zn interaction, different samples were studied by XRD. Fig. 2 shows the XRD patterns obtained for fresh, aged and reactivated samples. For the catalyst reduced at 503 K (Spectrum a), only peaks from metallic Cu and ZnO are observed. It can be seen that during reduction of the fresh sample

538

2

'

1

I

I

.

1

d 4

4

Z

1 1

9 2

b

1

(1

1

60

6

1

2

50

3

I

40

,,

20

3o

Figure 2: X-ray difl%action pattern (a) fresh catalyst reduced in Nz-H2 at 503 K, (b) fresh catalyst reduced in N2-H2 -H20 at 503 K,(c) aged catalyst, (d) catalyst reoxided at 623 K and reduced at 773 K. References: 1= ZnO, 2= Cu, 3= CuO, 4= Cu A1204, 5= Cu-Zn 6= Cu A1202

539 (Spectrum b), the presence of water vapor does not affect catalyst composition nor crystal size, and that it produces a partially oxidized Cu only. In the catalyst aged at industrial conditions (Spectrum c), peaks for a "spinel surface" species and CuO are observed as sharp signals. The diagram also shows peaks at 20 = 42.31 and 44.51 which are assigned to a Cu-Zn compound [9]. It is evident that industrial conditions produce an interaction of Cu with the support and an increase in the crystal size. The sample reoxidized at 623 K (Spectrum d) has the same crystalline phases that are observed in the aged catalyst. Alloy formation was confirmed by the increase in the lattice constant of copper, determined from the Cu [111] line (Table 2), after reduction in pure hydrogen or in H2-H20 streams or treated at high temperature (773K). Since the lattice constant of Cu-Zn alloy is higher than that of Cu [ 10], the difference is ascribed to the formation of a Cu-Zn alloy. This is consistent with the presence of the peaks corresponding to Cu-Zn (20 = 42.31 and 44.51) in the XRD spectrum (Fig. 2c) of the aged catalyst. Kanai et al. [ 11] concluded that alloy formation involves the migration and dissolution of ZnOx from ZnO particles to Cu crystallites which is favored by catalyst reduction with H 2 and is accelerated at high temperatures, in agreement with our results. The effects of temperature, steam and composition of reducing mixture on Cu crystallite size are shown in Table 2. It is seen that crystallite size of the fresh samples is relatively constant with reduction temperature. Reduction in steam does not sinter Cu, although crystal size increases at industrial conditions. For the aged catalyst reduced at 503 K after being reoxidized at 623 K a decrease in crystal size is observed. Lo Jacono et al. [ 12] have proposed that Cu redispersion is produced by interaction between CuO and A1203 to form a "superficial spinel" similar to CuA102 which, during reduction, releases the metal forming small crystaUites of Cu. Friedman et al. [ 13] have reported that, at low calcination temperatures (T c <773K), the copper is found mostly as a "surface spinel" resembling the structure of CuA1204 Table 2 Structural properties of Cu/'ZnO-AI~Oa catalysts Catalyst

Fresh

Reduction temperature (K)

Reduction mixture

503 573 503 513

N2-H 2 N2-H 2 H2-H20 pure H 2

Crystal size (/~)

Lattice constant (A)

46 49 44 43

3.6170 3.6146 3.6260 3.6245

On the other hand, since Cu crystallite size increases when reoxidized samples are reduced in pure H 2 at higher temperatures, a relationship appears to exist between the increase of crystal size, higher reduction rates and the heat evolved at these experimental conditions. The thermal treatment at higher temperature leads to reduction of ZnO to Zn, so that the support loses the role of material spacer for copper crystallites.

3.1. Catalytic activity In order to correlate the performance and stability of Cu-ZnO catalysts with their textural and structural characteristics, catalytic activity was determined for the CO water-gas-shift reaction. The results are presented in Table 3 for various reduction conditions in the form of reaction rates in a differential reactor.

540 Table 3 Catalytic activi~ of Cu/ZnO-AlaO,~ catalysts Catalyst

Reduction temperature

Reduction mixture

Reaction rate (mmol min ~ g-l)

(K) Fresh

503 573 503 513

N2-H 2 N2-H 2 H2-H20 pure H 2

0.52 0.64 0.24

Spent Reoxidized Reoxidized

503 503 773

N2-H 2 N2-H 2 N2-H 2

0.44 0.20 0.18

Water gas shift reaction conditions: catalyst weight: 0.4 g, reaction temperature: 500K, total flow rate: 490.5 cm3 min1. Gas feed composition: Pco: 0.03 atm, PH2:0.48 arm, PH20:0.26 atm, PN2:0.23 atm. The results show that catalytic activity is higher when the catalyst is reduced in a N2-H2 mixture between 500 and 573 K and that it decreases when steam is present in the reduction mixture. For samples reduced in H2-H20 flow, treated at high temperature or taken from an industrial reactor, lower activity is evident for the LTWGS reaction. Such deactivation in the samples reduced in H2-H20 stream may be caused by the formation of a Cu-Zn alloy and by changes in the Cu oxidation state, as determined by XRD characterization. However, the deactivation observed here in the reoxidized-aged samples could not have been caused by the observed decrease in Cu crystal size. In this case, the decrease in activity may be related to a "spinel surface" species observed in the XRD pattern (Fig. 2d). Laine et al. [14] have proposed that an increase in dispersion accompanied by deactivation can be explained by the formation of a highly dispersed but less active phase, possibly a CuA1204 "surface spinel". For the catalyst taken from an industrial reactor, the observed decrease in the reaction rate is possibly related to the observed formation of an alloy, the increase in crystallite size and changes in textural characteristics. 4. C O N C L U S I O N S Higher activity is observed for fresh LTWGS catalysts when reduced in a N2-H 2 mixture between 500 and 573 K, but their activities decrease when steam is present during the reduction step. Deactivation of CuO/ZnO/A1203 can be related to thermal sintering and Cu-Zn alloy formation. XRD data suggest that the presence of water, high temperature, and high partial pressure of hydrogen partially reduce ZnO, enhancing formation of Cu-Zn alloy. In the aged-reactivated catalyst reduced at 500 K; Cu is redispersed but the catalyst has a lower activity and stability than the fresh samples. This observation can be related to "surface spinel" species observed by XRD analysis Fresh catalysts have a single reduction peak in the TPR studies whereas reactivated catalysts have two. These two peaks are ascribed to two different copper species, one which reduces at lower temperature and another which reduces at a temperature similar to that of CuO. The first phase is more reactive and thus possibly the active phase in the reaction.

541 In summary, the results show that deactivation of Cu-ZnO catalysts can occur by thermal sintering and by a solid phase reaction between Cu and the support. REFERENCES

1. 2. 3. 4. 5. 6. 7. 8. 9. 10. 11. 13. 14.

P.W. Young and C.E. Clark, Chem. Eng. Prog., 69. G. Sengupta, D.K. Gupta, M.L. Kundu, S.P. Sen, J. Catal, 67 (1981) 223. T. Van Herwijnen, W.A. DeJong, J. Catal. 34 (1974) 209. A.J. Alvarez, R.D. Bonetto, D.M.A. Guerin, Powder Diffract., 2 (1987) 220. G. Sengupta, D.K. Gupta, L.M. Kundu, S.P. Sen, J. Catal, 67 (1981), 233. G. Fierro, M. Lo Jacono, M. Inversi, P. Porta, F. Cioci and R. Lavecchia, Appl. Catal,137 (1996) 327-347. G. Moretti, S.De Rossi and G. Ferraris, Appl. Surf. Sci 45 (1990) 341. P. Porta, M.C. Campa, G. Fierro, M. Lo Jacono, G. Minelli, G. Moretti and L. Stoppa, J. Mater. Chem., 3 (1993) 505. Power Diffraction File - 2 , 1 9 -179 (1995), ICD Data, Pennsylvania, USA. W.B. Pearson, A Handbook of Lattice Spacing of Metals and Alloys, Pergamon Press, Oxford, 1973. Y. Kanai, T. Watanabe, T. Fujitani, M. Saito, J. Nakamura and T. Uchijima, Cat. Letters 27 (1994), 67-7812.- M. Lo Jacono, A. Cimino, M. Inversi, J. Catal, 76 (1982) 320. R.M. Friedman, J.J. Freeman and F.W. Lyttle, J. Catal., 55 (1978) 10. J. Laine, F. Severino, A. Lopez Agudo, J.L.G. Fierro, J. Catal. 129 (1991) 297-299.

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9

Elsevier Science B.V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

543

Effects o f Promoters and Supports on Coke Formation on Pt Catalysts During CH4 R e f o r m i n g with CO2 Susan M. Stagg and Daniel E. Resasco School of Chemical Engineering, University of Oklahoma, 100 E. Boyd St., Norman, OK. 73019, USA

The promotion of Pt catalysts by the addition of Sn has been investigated. The results show that the addition of Sn lowers the initial activity for the reforming of CH4 with CO2 but increases its stability. The effects of the support on the catalyst activity have also been studied. ZrO2 promotes the highest activity and selectivity. It is believed that the ZrO2 aids in the dissociation of CO2, which helps to remove carbon formed from the decomposition of CH4 and increases the CO production.

1. INTRODUCTION In recent years, CH4 reforming with CO 2 (dry reforming) has experienced renewed interest. Initially, environmental concerns made this an attractive reaction because both CH4 and COE are greenhouse gases [ 1-4]. However, permanent reduction of COE emissions could only occur if the synthesis gas produced is used to make materials other than fuels [5, 6]. One advantage over conventional steam reforming is the resulting 1:1 ratio of HE:CO, which is preferred for the production of oxo-alcohols, and other oxygenates [6, 7]. Excess Ha, formed by conventional steam reforming, can suppress chain growth and decrease the selectivity to higher hydrocarbons in the FT reaction [8]. Dry reforming is the preferred reaction for Chemical Energy Transmission Systems (CETS) [9] and for the simultaneous utilization of CH4 and CO2 from natural gas reservoirs with an abundance of CO 2 [10, 11]. The high temperatures and low pressures needed for this endothermic reaction are conditions conducive to deactivation due to coke deposition through the Boudouard reaction and the decomposition of CH4 [12]. Current research has focused on Group VIII metals [13] on a variety of supports. Most of the work has been on Rh and Ni supported on A1203, TiO2, SiO2, and MgO [1, 8, 14-17] with some interest in Pt and Pd [5, 10, 18-19]. High activity and low carbon deposition have been observed on Rh catalysts [20-23], but the limited supply of this metal when compared to others reduces the potential for it to be commercially feasible [6]. Many studies have been done on Ni [24-28] which show that the support has a profound effect on the activity and deactivation of the catalyst. Severe deactivation due to carbonaceous deposits was found to

544 occur on SiO2 and A1203, while La203 supported Ni maintained high activity [28]. Research investigating Pt catalysts has been limited due to the deactivation observed from coke formation. The use of modifiers could give Pt a chance to become an effective methane reforming catalysts. In this contribution, we have explored the effect of adding Sn as a promoter. Previous studies on dehydrogenation catalysts have shown that the addition of Sn results in the formation of bimetallic alloys, which are less susceptible to coke deposition [29]. The second method investigated in this paper involves the variation of the support. Whether through electronic modification or participation in the reaction, the support can alter the rate and type of carbon formed during the reaction. Recent studies have shown that using ZrO2 as a support for Pt results in high activity and selectivity [5-6, 30]. We have investigated the use of bimetallic Pt-Sn catalysts supported on both silica and zirconia under various reaction conditions and after regeneration in air. 2. EXPERIMENTAL Silica and zirconia supports were obtained from W. R. Grace & Co. (silica gel Grade 923) and Magnesium Electron Inc. (zirconium hydroxide MEI #617), respectively. To obtain ZrO2, the hydroxide was calcined at either 650 or 800~ for 4 hours in stagnant air, prior to impregnation of the metals. Pt and Pt-Sn catalysts were prepared by incipient wemess impregnation with aqueous solutions of H2PtC16"6H20 and SnCI2-2H20. Bimetallic catalysts were made by co-impregnation with a 1"1 molar ratio of Pt:Sn. The incipient wetness liquid/solid ratio, determined for each preparation, varied from 0.5 to 0.75 cm3/g. The Pt loading on all the catalysts was 1.5 wt%. All impregnated samples were dried overnight at 110~ calcined in air (30 cm3/min) at 400~ for 2 h, and finally reduced, in situ, in H2 (30 cm3/min) at 500~ for 1 h prior to reaction. The reforming reaction was performed in a packed bed reactor. A quartz tube reactor was used with an inner diameter of 0.4 cm and an outer diameter of 0.6 cm. In each run, 0.1 g of sample was loaded into the reactor. The majority of the reactions were performed at 650~ for 6 h. In those rtms, the sample was preheated in He (99.999% purity) up to 650~ and then exposed to the reaction mixture. In another set of experiments, the catalysts were exposed to the reaction mixture of CO2 and CH4 at 300~ the temperature was then increased at a rate of 3.9~ to 650~ and kept constant at 650~ for 5 hours. The CH4:CO2 ratio and the time on stream were varied, but the weight-hourly-space velocity (WHSV)was held constant for all the runs at 36 h l. The exit gases were analyzed using a quadrupole residual gas analyzer from MKS Instrument Inc. Each signal was normalized to the total pressure. Conversions of CH4 and CO2 were calculated from the difference between the signals obtained at the reactor outlet at zero conversion (room temperature) and the signal measured during the run, then dividing by the signal at zero conversion. The calculation of moles of CO and H2 produced was based on a calibrated mol/signal ratio. Temperature Programmed Oxidation (TPO) of coke deposits was performed on all spent catalysts. The samples were heated at a rate of 8~ in 5% O2:He (30 cc/min) up to 800~ The exit gases were analyzed by mass spectrometry, normalizing the data with the average of five 0.2 ml pulses of CO2.

545

3. RESULTS AND DISCUSSION 3.1. Bimetallic Pt-Sn/SiO2 catalysts Fig. 1a and 1b show C O 2 and CH4 conversions as a function of time on stream with a 1:1 feed ratio of CO2:CH4 (WHSV - 36 h l ) for Pt and Pt-Sn supported on SiO2. In the course of these tests, temperature was increased from 300~ to 650~ in 90 min.

(b)

(a) 12 ~10-

L

i--Pt

65OC

!

s \,;/

g

g8 6

g

g 4 c,.) 2

~Pt i ~'-Pt-Sn i

0 120 240 Time (min)

360

0

120 240 Time (rrm)

360

Figure 1. (a) CO2 and (b) CH4 conversion for a 1:1 feed ratio of CO2:CH4 on the silica-supported Pt and Pt-Sn catalysts. Reaction temperature ramp from 300~ to 650~ followed by 5 hours of reaction at 650~ Figure 1a shows that the conversion of C O 2 o n the monometallic sample was much lower than that on the bimetallic catalyst and decreased to almost zero after 360 min on stream (250 min at 650~ CO2 conversion on the Pt catalyst was first observed at 590~ while on the bimetallic sample CO2 conversion started at 490~ By contrast, the conversion of CH4 (Fig. l b) was initially the same for both samples, but after 200 min on stream (110 min at 650~ the conversion on the bimetallic catalyst dropped below that of the monometallic sample. Similar differences in activity were also seen when comparing the moles of CO and H2 produced. The monometallic sample produced less CO, and experienced more deactivation, but had slightly higher H2 production than the Pt-Sn sample. The ratio of H2:CO produced varied between 0.2 for the Pt-Sn catalyst to 0.33 for the pure Pt catalyst. These results indicate that the monometallic catalyst has a greater tendency to decompose CH4, thus generating hydrogen and leaving carbon on the surface. TPO performed on the catalysts after reaction runs showed that the Pt sample had much more carbon deposited during the reaction (17.7 mmol C/g cat.) than the Pt-Sn sample (5.0 mmol C/g cat.) (Fig. 3a). The carbon formed on the Pt sample was not completely removed until 650~ which is 100~ higher than the temperature required to clean the bimetallic catalyst. This indicates that the carbon deposited on the Pt sample is much more graphitic. Previous studies on Pt supported on SiO2 have shown that Pt experiences rapid deactivation in the presence of

546 hydrocarbons and that significant carbon formation occurs after short times on stream [29]. It appears that the formation of graphitic carbon on Pt during dry reforming also occurs very rapidly and results in rapid loss of the CO2 conversion. The bimetallic sample which experienced some loss of CO2 conversion also formed carbon during the reaction. However, the amount was much less than that on the monometallic sample and was removed at much lower temperatures. Previous studies conducted on the same Pt-Sn catalyst series as that of this work showed that the presence of Pt-Sn alloys increased the resistance to coke deposition [29]. Under the conditions of the previous study, the formation of Pt-Sn alloys was demonstrated by TPR and TEM analysis. Similar alloys may be present on the bimetallic catalyst in the present work. Therefore decreased carbon could be ascribed to the presence of Pt-Sn alloys. Activity studies were performed on the catalysts with the reaction temperature being held constant at 650~ Fig. 2a and 2b show the CO2 and CH4 conversions as a function of time on stream for the monometallic and bimetallic samples in a 1:1 ratio of CO2:CH4. The monometallic sample had a high initial CO2 conversion but, due to rapid deactivation, the conversion dropped below that of the bimetallic sample after only 25 min.

(b)

(a) 60

Pt ~Pt-Sn

50

II _ _

50

o-,,v. = 40

!

o~ 4 0 - -~

.~0 30

0 .,..~

_

- - P t

1

~-Pt-Sn [ o Pt-Sn Reg i

~" 30 ~D >

20

o 20 10

Ooo~,~ooooooooo

-

oo

_

10 0

0

5O

T~ne (rnin)

1O0

50

100

Trne (mm)

Figure 2. (a) CO2 and (b) CH 4 conversion for a 1"1 ratio of CO2:CH 4. Reaction temperature 650~ The addition of Sn reduced the amount of carbon formed, but did not eliminate it. Therefore regeneration of the catalyst is required and from the TPO results shown in figure 3, regeneration must occur at 650~ to remove all of the carbon that is deposited. Previous work by our group has shown that regeneration at high temperatures in air causes segregation of the bimetallic alloys, resulting in activity losses due to increased carbon deposition and sintering [29]. It is important to observe if regeneration of the reforming catalysts causes segregation and how this affects the activity and stability of the catalyst during subsequent reactions. Figure 2 includes the conversion obtained on a Pt-Sn sample which was regenerated by heating in 5% O2/He mixture up to 650~ at a rate of 8~ and then in pure air (30 cc/min) at 650~ for 30 min. The sample was then reduced at 500~ in H2 for 1 h and finally exposed to reaction at 650~ The same initial CO2 conversion was observed as for the fresh bimetallic sample, but it rapidly deactivated

547 to a conversion near that of the pure Pt sample. The same trends were observed when monitoring the formation of CO and H2. The differences in activity between the monometallic, bimetallic, and regenerated bimetallic sample can again be explained in terms of the coke formation.

(b)

(a) 0.08 -

.--, 0.06 .

r

-

i --Pt

Pt

0.06

t ~ - - Pt- Sn ,~

,...~

9

0.08

0.04

r

o r..)

0.02

9 ,

i Pt-Sn

0.04 0.02 _

300

400 500 600 Temperature (C)

700

300

400 500 600 Temperature (C)

700

Figure 3. TPO spectra for Pt and Pt-Sn with reaction temperature (a) ramp from 300~ to 650~ and (b) constant at 650~ CO2 signal in arbitrary units. Figure 3b shows the TPO of carbon deposits for the reactions under isothermal conditions. As described above for the reactions conducted with a temperature ramp, the amount and type of carbon formed on the pure Pt catalyst was very different from the bimetallic sample. The Pt catalyst contained more graphitic carbon and a larger amount of carbon than the Pt-Sn catalyst, which explains the monometallic sample's large decrease in activity with time on stream. Deactivation of the monometallic sample occurred rapidly, and the sample reached a constant conversion of approximately 15% within 30 min on stream. The regenerated sample underwent much faster deactivation with no CO2 consumption observed after 17 min on stream. This rapid deactivation is typical of unalloyed Pt suggesting that segregation of the alloy occurs during regeneration. We have seen that regeneration also results in sintering that changes the particle morphology. Indeed Pt-rich particles (150-200 A) are present on the regenerated catalyst along with Sn-rich particles (50-100 A). After regeneration, the concentration of alloy particles is greatly reduced. It is believed that both sintering and segregation play important roles in determining the activity after regeneration. Segregation increases initial activity but also effects faster deactivation. Sintering would also cause an increase in the rate of deactivation and a decrease in the initial activity. Due to the formation of the large particles, it is possible that smaller amounts of carbon are needed to completely deactivate the catalyst. This is supported by the fact that very little carbon formation was observed on the regenerated Pt-Sn sample. Comparison of the CO2 and CH4 conversions for the Pt and Pt-Sn catalysts under the two different reaction temperatures (linear ramp and constant at 650~ shows that the conversion of both species was much higher when the reaction temperature was held constant at 650~ Comparison of the TPO's for both the Pt and Pt-Sn/SiO2 catalysts under the two reaction conditions show that the runs in which the temperature was held constant at 650~ resulted in

548 carbon formation (6 and 1 mmol C/g cat., respectively) much lower than when a temperature ramp was employed (17.7 and 5 mmol C/g cat., respectively). The type of carbon formed 'on the pure Pt sample was the same under both reaction temperature conditions, however the types of carbon formed on the bimetallic sample were different. The carbon deposited when the reaction temperature was held constant was removed at high temperatures, indicating that it was graphitic. The TPO of the carbon deposited during the temperature ramp experiment shows that the carbon found on the bimetallic sample was removed at much lower temperatures. Since at this low temperature carbon is not deposited during the constant temperature runs, it can be postulated that the carbon is deposited during the ramping portion of the run. Therefore, when the temperature of the reaction reaches 650~ the catalyst is already partially deactivated. This deactivation explains the lower initial activity observed during the experiments with a temperature ramp. Fig. l b shows that the conversion of CH4 begins 125~ before the conversion of CO2 It is believed that the dissociation of CO2 acts to remove carbon which is formed on the surface of the metal during the decomposition of CH4. If the decomposition of CH4 occurs in the absence of CO2, the amount of carbon formed will be much greater. When the ratio of CH4 to CO2 was varied, the amount of carbon changed significantly. The amount of carbon deposited increased as the ratio of CH4 to CO2 increased. When no CO2 was present in the feed (1:1 ratio of He:CH4, WHSV=36 h l ) substantially more carbon was formed on the monometallic and bimetallic samples. When the same experiment was performed with a 1:1 ratio of He:CO2 (WHSV=36 h "l) no carbon deposition was observed for either sample.

3.2. Zirconia-Supported Catalysts Comparison of the activity and the amount of carbon formed during reaction at 650~ for the samples supported on SiOE and those supported on ZrOE shows that the latter have a much higher initial activity and stability (Fig. 4). Also, the HE:CO ratio in the products for the ZrOEsupported catalysts are about 0.8, as opposed to a much lower ratio for the SiOE-supported catalyst. When the feed ratio is increased to 2:1 CH4"CO2, a small amount of carbon is deposited on the Pt/ZrOE catalyst (Fig. 5), but no deactivation is observed. Figure 4 also shows the COE conversion when the reaction is performed at 800~ Under these conditions, the Pt/ZrOE catalyst exhibits significant deactivation. This loss of activity is accompanied by the formation of large quantities of carbon (35 mmol C/g cat), as determined from the TPO shown in Figure 5. The largest TPO peak occurs at 400~ the same position as the small amount of coke deposited at a reaction temperature of 650~ In addition, a small peak is observed at 650~ which indicates the formation of graphitic carbon. As shown in Figure 4, the bimetallic sample deactivates more rapidly than the monometallic sample, but exhibits greater long-term stability. After 6 h on stream the Pt-Sn catalyst reaches a constant activity while the Pt catalyst continues to deactivate. Activity measurements (not shown) after regeneration of the Pt-Sn sample, under the same conditions described previously, showed a decrease from initial conversion. However, after 6 h on stream the activity of the regenerated sample was the same as that of the fresh sample. In contrast with the TPO for Pt/ZrO2, the TPO of the Pt-Sn/ZrO2 after reaction at 800~ revealed negligible amounts of carbonaceous deposits.

549

(b) 0.08

(a) 100

i

650C i Pt/ZrO2

_ 0.06

~ 8o o 70 o

8ooci

~0.04 ~ ~

60

0

650 Pt 800 Pt o 800 Pt-Sn

0.02

o_i)_

50 0

120 240 Trim (,Tin)

i

360

300 i,

ii

400 500 600 Temperature (C)

700

Figure 4. C O 2 conversion for Pt/ZrO 2 at reaction Figure 5. TPO spectra for Pt/ZrO 2 reactions at temperatures of 650~ and 800~ 650~ and 800~ CO2 signal in arbitrary units.

It has been proposed by Lercher et al. [30] that the ZrO2 facilitates C O 2 dissociation which, when occurring near the metal particle, aids in the removal of carbon formed on the metal from the decomposition of CH4. Based on this model, our results indicate that at 800~ the rate of C deposition on Pt/ZrO2 exceeds the rate of removal by CO2 dissociation. The large amounts of C deposited at 800~ ( > 450 C/Pt molar ratio) suggest that C may start depositing on the metal but then migrates onto the support. The drastic reduction in the amount of C observed by TPO on the Pt-Sn/ZrO2 samples could be due to the ability of Sn to inhibit carbon deposition. However, a possible role of Sn in accelerating the carbon removal by CO2 dissociation cannot be ruled out. 4. CONCLUSIONS It is demonstrated by the data from this study that the dry reforming performance of Pt, which is highly susceptible to carbon deposition, can be enhanced with the addition of Sn. On a SiO2 support, alloy formation between the metal and promoter decreases coke formation, increasing the activity and stability of the catalyst. Regeneration of the bimetallic catalyst results in segregation of the alloy which decreases the activity and stability observed during the subsequent cycles. The use of ZrO2 as a support yields a much more efficient catalyst with resistance to deactivation at 650~ At 800~ the Pt/ ZrO2 catalyst deactivates due to the deposition of large amounts of carbon. With the addition of Sn, the initial rate of deactivation increases, but long-term stability is observed.

550

ACKNOWLEDGMENTS This work was partially supported by the Donors of the Petroleum Research Fund administered by the American Chemical Society. We gratefully acknowledge the National Science Foundation for a graduate traineeship for one of us (SMS).

REFERENCES 1. 2. 3. 4. 5. 6. 7. 8. 9. 10. 11. 12. 13. 14. 15. 16. 17. 18. 19. 20. 21. 22. 23. 24. 25. 26. 27. 28. 29. 30.

A. Takano, T. Tagawa, and S. Goto, J. Chem. Eng. Japan, 27 (1994)727. T. Osaki, T. Horiuchi, K. Suzuki, and T. Mori, Catal. Lett. , 35 (1995) 39. O. Takayasu, C. Soman, Y. Takegahara, and I. Matsuura, Stud. Surf. Sci. Catal., 88 (1994) 281. J . R . Rostrup-Nielsen, Stud. Surf. Sci. Catal., 81 (1994)25. K. Seshan, H. W. ten Barge, W. Hally, A. N. J. van Keulen and J. R. H. Ross, Stud. Surf. Sci. Catal., 81 (1994) 285. J . R . H . Ross, A. N. J. van Keulen, M. E. S. Hegarty, and K. Seshan, Catal. Today, 30 (1996) 193. S. Wang and G. Q. Lu, Energy & Fuels, 10 (1996) 896. S . B . Tang, F. L. Qiu, and S. J. Lu, Catal. Today, 24 (1995) 253. J. T. Richardson and S. A. Paripatyadar, Appl. Catal. , 61 (1990) 293. F. Solymosi, G. Kustan, and A. Erdohelyi, Catal. Lett. , 11 (1991) 149. Z. L. Zhang, V. A. Tsipouriari, A. M. Efstathiou, and X. E. Verykios, J. Catal., 158 (1996) 51. J. S. Chang, S. E. Park, and H. Chon, Appl. Catal. A, 145 (1996) 111. J. R. Rostrup-Nielsen and J-H. Bak Hansen, J. Catal., 144 (1993) 38. A. Erdohelyi, J. Cserenyi, and F. Solymosi, J. Catal., 141 (1993) 287. J. Nakamura, K. Aikawa, K. Sato, and T. Uchijima, Stud. Surf. Sci Catal., 90 (1994) 495. T. Uchijima, J. Nakamura, K. Sato, K. Aikawa, K. Kubushiro, and K. Kunimori, Stud. Surf. Sci. Catal., 81 (1994) 325. H. M. Swaan, V. C. H. Kroll, G. A. Martin, and C. Mirodatos, Catal. Today, 21 (1994) 571. F. Solymosi, A. Erdohelyi, and J. Cserenyi, Catal. Lett. , 16 (1992) 399. M. Masai, H. Kado, A. Miyake, S. Nishiyama, and S. Tsuruya, Stud. Surf. Sci. Catal. , 36 (1988) 67. D. Qin and J. Lapszewicz, Catal. Today, 21 (1994) 551. J. H. Edwards and A. M. Maitra, Fuel Processing Tech. , 42 (1995) 269. V. A. Tsipouriari, A. M. Efstathiou, Z. L. Zhang, and X. E. Verykios, Catal. Today, 21 (1994) 579. Y. Sakai, H Saito, T. Sodesawa, and F. Nozaki, React. Kinet. Catal. Lea., 24 (1984) 253. M . C . J . Bradford and M. A. Vannice, Appl. Catal. A, 142 (1996) 73. M . C . J . Bradford and M. A. Vannice, Appl. Catal. A, 142 (1996) 97. E. Ruckenstein and Y. H. Hu, J. Catal. , 162 (1996) 230. E. Ruckenstein and Y. H. Hu, Appl. Catal. A, 133 (1995) 149. Z. Zhang and X. E. Verykios, Catal. Lett. , 38 (1996) 175. S.M. Stagg, C. A. Querini, W. E. Alvarez. D. E. Resasco, J. Catal. (in press) J.A. Lercher, J. H. Bitter, W. Hally, W. Niessen, and K. Seshan, Stud. Surf. Sci. Catal., 101 (1996) 463.

9

Elsevier Science B.V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

551

Ni -based Catalysts for Methane Conversion with Carbon Dioxide. Catalyst pellet Deformation Induced by Coking and Gasification L. A. Rudnitsky Institute for Nitrogen Industry, 109815, Zemljanoi Val, 50, Moscow, Russian Federation, e-mail : [email protected]

The study of deformation of Ni-based catalyst pellets, induced both by carbon formation in pellet pores and by its gasification, was carried out in situ in flowing gaseous CH4 and CO2. Deformation intensity and dynamics were determined from parallel measurements of carbon content AG/G and pellet linear elongation AL/L during linear heating and cooling sequences. This allowed us to obtain (a) the temperature dependencies of AG/G and AL/L and (b) the relationship between AG/G and AL/L. The character of the latter relationship depends on the type of carbon reaction (formation or gasification). During linear heating, the carbon deposition proceeds in two steps. In the first step, deposition of carbon does not cause pellet deformation. Rather, the expansion of pellets commences only after a certain content of carbon has been exceeded (the second step). To a first approximation, AL/L is directly proportional to the content of carbon added during the second step. The linear deformation of the pellet over the entire range (from 5xl04up to 101) is completely irreversible. This parameter can therefore be regarded as a measure of the extent of pellet destruction. 1. INTRODUCTION One of the important causes of catalyst deactivation in hydrocarbon-reforming is the mechanical failure of pellets. It would be desirable to study this problem in situ under "unfavorable" conditions, namely those under which significant internal stresses may develop in the pellets, through formation of carbon or coke. Cycling through a range of conditions, under which carbon is either formed or gasified, induces changes in carbon content AG/G which, in turn, give rise to the deformation of pellets AL/L. The analysis of the relationships AL/L - T, AL/L - AG/G and consequently AG/G - T enables the study of the dynamics and intensity of pellet deformation under "unfavorable" conditions and the evaluation of the ability of the pellets to withstand internal stresses. This field of catalyst deformation has received little investigation. Conversion of methane with carbon dioxide over a supported Ni-based catalyst (NiC) accompanied by reversible carbon formation was selected as a test reaction for study of pellet deformation. This reaction is of special interest since catalyst deformation may be induced not only by carbon formation (CF), but also by its gasification (GAS).

552 Relative rates of CF and GAS are determined by a complex combination of both thermodynamic and kinetic factors [ 1]. The pellet deformation due to these processes can be effectively studied in cyclic, nonisothermal experiments.

2.

EXPERIMENTAL

In this study, changes of pellet mass and pellet linear deformation, were measured using a microbalance (Sartorius) and a quartz inductive dilatometer (Netch) in parallel experiments involving identical specimens, gas flow, and temperature programming regimes. The specimen was a 7 mm long x 5 mm diameter cylinder having a mass of about 600 mg which had been cut from either Ni/o~-A1203or mixed-catalyst pellets. The reaction of methane conversion was conducted in both units at 1 atm, from 400~ to 900~ and at a SV of 10000 h1 on catalysts previously reduced with linear heating up to 800~ in flowing H 2. Reduced specimens were cooled to 150-200~ after which they were exposed to a reaction gas mixture of 20% CH 4 and 80% CO 2. The cyclic thermal program, involving linear heating and cooling sequences, was then initiated. Heating rate was varied over a wide range but as a rule was 5 K/min. Reproducibility was provided by using "standard" test conditions (constancy of mixture content, heating rate k, NiC mass, reactor size and gas flow rate). 3. RESULTS

3.1. Dynamics of NiC CF and GAS during linear heating Previous results obtained in this laboratory [2-4] are necessary for the analysis and interpretation of data on NiC pellet deformation. On linear heating, CF from this study initiates at 400 - 420~ and proceeds to 670~ (Tin), at which point CF changes to GAS. The reverse transition of GAS to CF takes place at Tm during linear cooling. Thus, "the principle of equilibrated gas" [1] can be applied to our system and Tm can be taken as a thermodynamic equilibrium temperature. The dependence of AG/G on T obtained during linear heating therefore results in a peak ("carbon peak" or CP). The peak height I-Icp (which is equal to (AG/G)m- the maximum of carbon content) is determined by CF dynamics and its position (Tin) is determined by thermodynamics. The CP of some Ni/a-A1203 catalysts differing one from another in promoter content are presented in Figure 1. It is seen that the values of Hcp obtained on NiC linear heating vary over a wide range. Nevertheless, the comparison of CP normalized to the single value of Hcp shows them practically coinciding.. In Figure 2, the rates of CF and GAS obtained at ~ values from 2 to 12 K/min are shown. In these tests (AG/G)m, i.e., Hce was found to be inversely proportional to L. Hcp =

B/~,

( B is a constant )

(1)

Therefore, it is seen that rates of both reactions depend on temperature only, not on the carbon content. Zero order kinetics of CF is consistent with whisker carbon growth at constant whisker density[ 1]. Zero order kinetics of GAS correspond to whisker gasification. Active centers of whisker growth are thought to arise during reaction at relatively low temperatures [3,4]. This factor especially influences the dynamics of NiC CF studied during linear cooling.

553 3.2. Dynamics of NiC CF on linear cooling After transferring from linear heating to linear cooling at T > Tin, GAS continues and then finishes at Tin. With further cooling where T < Tm CF is reinitiated. However its rate may vary considerably. Let us denote the temperature at the end of heating (start of cooling) as Tf. The thermogravimetric curves of cooling which differ from one another in values of Tf are only given in Figure 3. T f varies here from Tm up to 830~ where we observe practically complete removal of carbon. For the sample with complete carbon removal (Tf = 830~ we observed the absence ofNiC coking on linear cooling from Tf to 350~ Apparently at very high temperatures gasification eliminates not only whisker carbon but whisker growth centers as well. We speculate that the whisker centers removed by GAS cannot be re-formed at high temperatures (at least, during the time of cooling).

0,12 0,10 0,08

1,5] 1,0

AG/G 2N

dG/d ~,rel.units ~ ~

0,02

0,0 -0,5 -1,0 -1,5 -2,0

0,00 400 500 600 700 800 Temperature, C

-2,5()0 ' 460 ' 660 ' 860 Tem peratu re, C

Figure 1. The CP of Ni/tx-A1203 Catalysts 1-3 differing one from another in promoter concentration. Curves 2N and 3N are CP of Catalysts 2 and 3 normalized to CP height of Catalyst 1. Heating rate ~ = 5 K/min.

Figure 2. The dependencies of' rates of CF (positive values) and GAS (negative values) on temperature obtained during linear heating of Ni/oc-A1203catalyst at 2, 3, 6 and 12 K/min.

0,06 0,04

3.3. The deformation of NiC produced by CF and GAS reactions The data obtained in parallel thermogravimetry and dilatometry tests during linear heating are shown in Figure 4. It is seen, besides the CP described above, that the deformation coking peak (DCP) is rising. The left branch of DCP shows the expansion produced by CF and the right branch shows the pellet contraction produced by GAS - near 800~ and higher temperatures by sintering. Peak temperature maxima (Tin) for both samples approximately coincide, but the shape of DCP differs significantly from that ofCP: (a) The temperature of DCP initiation is about 100 K higher than for CP and (b) The contraction of the pellet produced by GAS does not restore its original size despite nearly complete elimination of deposited carbon- this means that NiC expansion has an irreversible character. The dilatogram which was selected for Figure 4 has a very high value of linear deformation. Nonetheless the results are valid over a wide range of DCP height HDCP(HDcp equals (AL/L)m- elongation obtained on linear heating to Tin). Let us first describe DCP with relatively low and moderate HDCPvalues (Figure 5). Low DCP values are those whose HDCP values are of the order of thermal elongation (the low DCP values are presented without the contribution of the thermal elongation effect). The HDCP of smallest

554 DCP corresponds to value of pellet thermal elongation obtained upon heating from 20 to 550~ The Hocp - Z, relation is unique at these X values: the decrease from 15 to 10 K/min produces an increase in Hocp of 0.0035 to 0.03. In Fig. 5 and in further ones, the portions of the dilatogram obtained on linear cooling are represented as well. Upon heating NiC to 850~ - 900~ the dilatometric curves are nearly horizontal, indicating no pellet expansion. This effect can be explained by the absence of carbon deposition on linear cooling after high-temperature treatment of the pellet, as previously described. In Figure 6 the enormous DCP curves shown, are obtained at relatively low ~, values (4.5 and 6.5 K/min). These pellets discharged after cyclic treatment were, as a rule, not degraded. DCP curves obtained when pellets failed, as discussed later, have a different shape

0,04

Figure 3. Carbon content increase upon cooling for a Ni/a-A1203 catalyst during reaction after heating up to Tf. Tf values are marked near the curves. All cooling curves with one exception ( T f 830~ have been moved down to point A so that the carbon content at the CF initiation is equal to zero.

_ AG/G . . . . . . . . --670 C 9

0,03

i - u - n ~

A1

~n~ln

9 9 9 9n

745 C-"~,~

/

0,02

0,01

2o oooo

0,00

360 460 560'660 760'860 Temperature, C

0,16-ALl L

0,08

0,030 0,025 0,020 0,015

0,04

0,010

0,12

A

ALl L 10 13N

0,005

0,00

0,000 I

n

I

~

n

t

A

1

J

320 420 520 6'20 720 820 Tern perature, C

Figure 4. Comparison of the CP (thermogravimetric Curve G) with DCP (deformation Curve L) obtained on the linear heating of NiC, prepared by mixing. Both curves relate to the common axis y. ;L = 6.5 K/min. The microbalance and dilatometer specimens have been cut from the same NiC pellet.

I

400

J

I

~

I

600 800 Temperature, C

~

I

1000

Figure 5. Dilatometric curves (DCP) of linear heating and cooling of NiC prepared by mixing at low and medium values of Hocp. Z, varies from 10 to 15 K/min, Z, values are given near the curves. Curves 13N and 15N are DCP 13 and 15 normalized to the height of DCP 10.

555

from those in which no failure occurred. Apparently the stability of enormously deformed pellets can be explained by binding with carbon deposited in the catalyst pores. In Figure 5, no essential distinction between shapes of the lowest and other DCP curves after HDce normalization is evident. Moreover, they show the same irreversible character of linear deformation. The DCP obtained at 4.5 K/min (Figure 6) has an Hoce which is forty times larger than the Hoce obtained at 15 K/min. (Figure 5). Nevertheless, the character of these DCP curves is generally similar; i.e., the shapes of the peaks are close to one another before Tm as well as after it.

0,14 0,12 A L/L 0,10 0,08 0,06 0,04 0,02 0,00

~-~4.5

400 500 600 700 800 900 Temperature, C

Figure 6. Dilatometric curves (DCP) for linear heating and cooling of NiC prepared by mixing for with high HDCPvalues. Z, varies from 4.5 t o 10 K/min; its values are given near the curves. The DCP of precipitated Ni/tx-A1203 catalyst is marked by "Pr."

0,12r AG/G 0,10 0,08 0,06 0,04 0,02 0,00

300 400 . . . . . .500 . . . 600 760 800 . . . . .900

Temperature,C

Figure 7. Thermogravimetric curves of CF obtained with linear heating of NiC prepared by mixing. The catalyst and experimental conditions are identical to those given in the legend to Figure 6. The curve of DCP for the precipitated Ni/aA1203 catalyst is marked by "Pr".

The important result of this comparison is the close agreement between the values of relatively irreversible deformation (evident from the relation between values of AL/L and HDCP) attained at the end of the heatingprocess, i.e., 0.54 at 4.5 K/min and 0.58 at 15 K/min. The similarities between the different DCP shapes will be discussed later. Data given in Figure 6 allow us to estimate dependence of HDCP on ~,, which differs from Equation 1: HDCp = [( B / 2L) - C]* R = (Hop - C ) R

(2)

(where B, R and C are constants). This simple relation has two important consequences: 1. Under conditions of Z >B/C, DCP (e.g., pellet deformation) is negligible upon linear heating of NiC during reaction in spite of normal CF. As will be shown later, this conclusion is supported experimentally. 2. It can be supposed that C is the maximum content of carbon having no influence on pellet linear size. Thus the maximum content of active carbon deposited upon linear heating to Tm can be expressed as (Hcp - C). Accordingly, the maximum linear deformation is in direct

556

proportion to the maximum content of active carbon. More detailed information has been obtained from the direct study of the relationship between pellet deformation and carbon content. It is obvious that the values of HDCP correlate directly with the values for I-Icp. This result, following from Equation 2 is generally valid for different types of NiC. The numerical correlation fits only the experimental data for the specified catalyst. It is seen that the CP of the precipitated catalyst is the same as that of the mixed catalyst obtained at ~ = 10 K/min. There is no such coincidence for the corresponding DCP. The catalyst based on a-alumina support shows a significantly lower elongation than the promoted one does at the same carbon content. This means that the values of the constants in Equation 2 depend on NiC's nature.

0,14 _AL/L 0,12 0,10

0,10 GEl1

0,06 0,04 0,02 0,00

0,04

0,08

6.5

0,08

!\

0,06

AG,G

0,02 0,00'0,132'0,04'0,06'0,08'0, i0 A G/G, carbon content

Figure 8. Dependencies of catalyst pellet elongation on carbon content obtained for promoted NiC upon linear heating at different Z, values which are labeled near the curves. The dotted straight line connects the points, which correspond to parameters at Tm (maximum of the carbon content and elongation).

HE

0,00 |

4 0

1

500

i

1

600

i

I

700

Temperature, C

i

I

800

Figure 9. Coking and pellet elongation dynamics in the sequence of linear heating and cooling of promoted Ni/ctA1203 catalyst ( ~ , - 2.5 K/min). HE denotes end of heating, CE, end of cooling. A L / L and AG/G mark the dilatometric and thermogravimetric curves respectively.

3.4. The relationship between pellet deformation and carbon content The data given in Figures 6 and 7 allow us to obtain the dependency of pellet elongation on carbon content. Their dependency is presented in Figure. 8 which shows that there are three types of relationships between AL/L- AG/G. The first and the second relationships occur in the linear heating regime leading to CF (up to Tin). The initiation of corresponding curves takes place at the point of zero carbon content and zero deformation. They extend nearly horizontally up to a certain critical carbon content (AG/G)o- thus, the deposition of carbon does not produce any deformation before this point. Evidently, within this interval the growing carbon whiskers fill the pellet pores without developing essential inner stresses. After (AG/G)c is exceeded, a sharp transformation is evident. The second type of AL/L AG/G relation involves considerable expansion of the catalyst pellet up to Tm where CF ceases and GAS begins. The values of the maxima Hcp and HocP attained at this point are extremely important parameters which determine the intensity of pellet contraction on further heating during the GAS process (the smooth curves obtained in this case are of the third type of AL/L - AG/G relationship).

557

The relationships given in Figure 8 indicate that there are two steps for carbon deposition during linear heating which are separated by some critical carbon content (AG/G)~. This point is believed to correspond to a sharp rise of inner stress up to some critical level which produces microdestruction of the pellet. Accumulation of carbon proceeds over the interval from (AG/G)~ to (AG/G)m. As Figure 8 shows HDCP is in direct proportion to (AG2/G)m - the maximum of carbon content added at the second step of CF: Hoop = R (G2/G)m = R [(AG/G)m - (AG/G)o]

(2a)

This equation is identical to Equation 2; however, the sense of its terms is more clea~. C is (AG/G)c and 3 R is the relative value for volume corresponding to microdestruction.

3.5. The influence of CF two-step character on deformation dynamics It follows from Equations 2 and 2a that at a critical value of ~,, which equals B/(AG/G)c, HDCP approaches zero - the disappearance of DCP takes place in spite of CF actively proceeding. For some types of NiC the critical values of ~, are practically accessible. At this and greater ~, values the catalyst pellet can be heated from ambient temperature up to Tm without any deformation. As an example, the data obtained during the cyclic treatment of the promoted Ni/a-A1203 high porosity catalyst at ~ = 2.5 K/min are presented in Figure 9. On heating we obtained a "nomud" CP with an acceptable value of Hce. The heating was then interrupted at Tf = 770~ and GAS of the deposited carbon has not been completed. Thus, during the linear cooling we obtained the CF increase by cooling to 400~ At the same time, under identical heating condition we obtained the horizontal dilatogram (full absence of the pellet elongation). We obtained the elongation effect only after a certain time of CF during the cooling period. The latter allowed us to reach and exceed the critical value (AG/G) c. This test provides an opportunity to use NiC during CF without any pellet deformation. In addition, the two-step character of CF explains: (a) the delay of DCP initiation shown in Figure 4 (associated with the first step of CF) and (b) the interconnected character between ~, values and the position of DCP branches, the left of which is expected to shift in the direction of higher temperatures, while the fight remains unchanged despite the increase of ~,. The validity of the last effect is supported by the DCP given in Figure 10 obtained at different ~, values and normalized to a single value of HocP. 3.6. The sense of CP and DCP normalization to a single height The coincidence in the shapes of the peaks (CP as well as DCP), which significantly differ in their intensity, before recalculation to a single value for peak height (Figures 1, 5 and 10), indicates their mathematical similarity. This means that the peak curves can to be described by a single expression with precision for the numerical factors. The similarity of the peaks in each demonstrated Figure indicates the constancy of mechanism for the processes studied over a whole range of carbon content data. It is valid for mechanisms of deformation produced by CF as well as by GAS to stretch over a whole range of maximum deformation values: (HDce) from 0.0035 up to 0.14. This conclusion is supported by Figure 11, where values of all irreversible deformations, expansions as well as contractions, are in direct proportion to (AG2/G)m.

558 3.7. The irreversible deformation of pellets as a result of accumulation of microdestruction events

The enormous scales of elongation as well as contraction, which are produced by the second step of CF and GAS (Figure 6), indicate that these effects are completely irreversible. The first process produces pellet microdestruction events, each with a relative volume 3R, acaunulated during heating to I",. The pellet deformation recorded at this point is the result of their summation.

0,16 _ALl L 4.5 .... 0,12

0,16-

:

0,12

0,08 0 , 0 4 ~ 0,00

i

lO,Pr

pellet! failure 400 ' 600 ' 800 ' 10'00 Temperature, C

Figure 10. DCP of promoted NiC obtained at different 7~ values marked near the curves normalized to a single height HDcp. Original DCP curves are given in Fig. 6. DCP of precipitated NiC is denoted by Pr. DCP of pellet failure has been moved up to 0.02.

AL/L

0,08 0,04 : 0,00 -0,04 -0,08

1

'

i

0,02 0,04 0,06 0,08 O,iO Carbon content

Figure 11. The dependencies of the NiC pellet linear deformation on carbon content obtained from the data in Figure 6. 1 denotes Ht~p values; 2, final irreversible deformation; 3, values of contraction within the range TIn- 780~ 4, within the range T,, -880~ A is a calculated point.

As was shown in Figure 8, HDCPdepends linearly on the content of carbon added during the second step of CF. This also allows us, with crude approximation, to describe the AL/L AG/G relationships as linear. As noted before, the CF process has zero order kinetics in carbon content. Therefore, elongation kinetics are also zero order in carbon content. The second step of CF produces local microdestructions in the original structure. At the same

time the newly formed fragments of structure are linked, to some extent, by the deposited carbon. Carbon deposits acts as a binder which stabilizes the pulverized structure and keeps the catalyst pellet form unchanged, even over a region of enormous deformation. GAS eliminates the binding material (carbon) resulting in complete pellet destruction. The mechanism of contraction produced by GAS is not the reverse of elongation: it is an extension of pellet destruction by another mechanism. Also, its dynamics are not zero in carbon content. It is hypothesized that pellet contraction or failure produced by elimination of carbon may be described on the basis of a percolation model by numerical simulation method.

Thus there are two different types of catalyst pellet failure. The first takes place during elongation produced by carbon deposition. The second takes place during contraction produced by carbon elimination. The DCP of such failure is presented in Figure 10. Its shape is significantly different than that of the "normal" D C P .

559 4. CONCLUSIONS 1. Under conditions of linear heating, the dynamics of carbon deposition on Ni/o~-Al203, its gasification and pellet elongation are zero order in carbon content. 2. Under linear heating the deposition of carbon proceeds in two steps. The carbon deposited in the first step does not influence the pellet size. Pellet elongation is directly proportional to catalyst carbon content in the second step. 3. Variations in values for maximum pellet expansion and contraction over almost two orders (from 0.003 up to 0.1) do not change the dilatometric curve shape. Also, the mechanism of deformation remains constant over the whole deformation range. 4. Pellet deformation induced by the carbon formation as well as its gasification is a completely irreversible process over the whole deformation range. Both processes contribute to pellet degradation extent and its failure. 5. The two-step mechanism of carbon formation provides an opportunity for catalyst operation

in a range of carbon formation without pellet deformation.

REFERENCES

1. J.R. Rostrup-Nielsen, Catalytic Steam Reforming, Berlin, Heidelberg, 1984. 2. L. Rudnitsky, T. Soboleva, G. Korotkova and A. Alexeev, ACS. Petrol. Chem. Div., PREPRINTS, 40, No. 4 (1995) 567. 3. L. Rudnitsky, T. Soboleva, A. Alekseev, React. Kinet. Catal. Lett. 26, (1984) 149. 4. L. Rudnitsky, T. Soboleva, A. Alekseev, Ibid, 30, ( 1986 ) 113.

NOMENCLATURE CF- reaction of carbon formation CP - coking peak DCP - deformation coking peak GAS - reaction of carbon gasification HDCP - deformation coking peak height NiC - Ni catalyst, Hcp- coking peak height T - temperature Tf- temperature at which linear heating finishes Tm- temperature of thermodynamic equilibrium of CF - GAS reactions AG/G - the content of carbon deposited in NiC pores (AG/G)o - the critical content of carbon where at the transfer from the first step of CF to the second takes place (AG/G)m - the maximum carbon content attained on linear heating to Tm (AG2/G)m, - the maximum carbon content added during the second step of CF AL/L - the pellet linear elongation (AL/L)m - the maximum elongation attained on linear heating to Tm L - rate of temperature linear programming

This Page Intentionally Left Blank

9

Elsevier Science B.V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

561

C a r b o n F o r m a t i o n f r o m D e c o m p o s i t i o n o f C H 4 on S u p p o r t e d Ni catalysts R. LCdeng a, M. Barr~-Chassonnery b'c, M. Fathi b, O. A. Rokstad a, A. Holmen b aSINTEF Applied Chemistry, N-7034 Trondheim, Norway. bDepartment of Industrial Chemistry, Norwegian University of Science and Technology, N-7034 Trondheim, Norway. COn leave from: University of Poitiers, URA CNRS 350 40, Avenue du Recteur Pineau, 86022 Poitiers Cedex, France

Carbon formation during decomposition of CH 4 on unpromoted Ni/t~-A1203 and on Mgand Ca-promoted catalysts, was studied in a tubular microreactor using a high pressure microbalance for continuously measuring the carbon deposition. The reaction was studied at 873-1073 K and partial pressures of CH 4 in the range of 0.45-4.5 bar. At initial conditions, the reaction was found to be first order with an activation energy of 80 kJ/mol. Adding promoters to the Ni/(t-A1203 catalyst increased the observed rate of carbon formation. The effect of H20 was investigated at 843 - 923 K and 3.5 bar CH 4. The formation rate of carbon was increased by the addition of small amounts of H20. Ca was found to lower the limit of carbon formation (i.e. the H20 partial pressure above which carbon is not formed), whereas Mg increased the carbon limit. It was also observed that additions of even relatively high concentrations of H 2 increased the observed rate of carbon formation from CH 4. 1. I N T R O D U C T I O N C H 4 may be converted to synthesis gas by steam reforming or by reaction with 0 2 through secondary reforming and partial oxidation [ 1]. Steam may also be replaced by CO 2. As for all processes where hydrocarbons or carbon oxides are exposed to high temperatures, carbon deposition is a possible problem in the production of synthesis gas. In the production of synthesis gas, carbon may be formed by decomposition of methane (and higher hydrocarbons) or by the Boudouard reaction [2]: CH 4

2 CO

~ C + 2 H2 ~-~ C + C O 2

AH~ AH~

= 75 kJ/mol = -172 kJ/mol

(1) (2)

CH 4 decomposition is favored by high temperatures. Steam reforming is usually carried out on supported nickel catalysts. Supported nickel is also a typical catalyst for secondary reforming and for methanation and could even be used for partial oxidation. The carbon forming reactions described above are also catalyzed by nickel [2]. Well-defined carbon structures are documented to form on nickel surfaces and other transition metal surfaces [3,4]. Distinguishable morphologies include filamenteous (or whiskerlike) carbon, encapsulating carbon films and pyrolytic carbon [5]. Whisker growth causes no immediate deactivation, but it can destroy the catalyst mechanically [6].

562 Carbon formation may lead to blockage of the reactor tubes resulting in the development of hot spots or it may cause catalyst breakdown [7]. It is therefore important to avoid carbon formation by choosing suitable operating conditions. Carbon formation may also be reduced by proper formulation of the catalyst itself. The present work focuses on carbon formation from C H 4 and C H 4 mixed with H 2 0 or H 2 on a Ni]tx-Al203 catalyst and on the same catalyst promoted by Ca and Mg. 2. E X P E R I M E N T A L

Carbon formation was studied on ~-AI203 supported Ni catalysts prepared by impregnation using the nitrate salts (Ni, Ca, Mg), according to the procedure given by Demicheli et. al. [8]. The promoters (Ca, Mg) were added to the impregnated base catalyst. After impregnation the catalysts were calcined in flowing air for 10 h at 973 K. The low surface area alumina was supplied by Alcoa (24-48 mesh, 0.06 m2/g BET). The Ni content as measured by AAS, was 9 wt% for the unpromoted catalyst and 7.7 wt% for the promoted catalysts containing 3.1 wt% Ca or 2.3 wt% Mg, respectively. The BET surface area increased to 0.6 m2/g for the unpromoted catalyst and to 0.7 m2/g for the Ca and Mg promoted samples. The catalysts were characterized by H 2 chemisorption at 373 K using a pulse technique [9] and assuming a H/Ni stoichiometry of 1 [10]. The dispersions of the unpromoted and the Ca and Mg promoted catalysts were determined to be 0.062 %, 0.34 % and 0.56 %, respectively. Carbon deposition rates were measured in a microreactor connected to a Sartorius 4436 high pressure microbalance [11]. The catalyst (17-70 mg) was placed on quartz wool in a perforated quartz basket in a stainless steel reactor lined with an alumina tube (i.d. = 15 mm) and hung from one arm of the microbalance by a quartz fiber. Water was fed using a Lewa M3 pump. A flow of inert gas was always maintained through the microbalance. The composition of the product gas could be determined by on-line gas chromatography. The catalysts were first reduced by heating in a He/H 2 mixture (0.5 bar H2) from 298 K to 973 K at a rate of 2 K/min and held at 973 K for 10 h. The reaction was then studied as a function of temperature and partial pressure of C H 4 in mixture with He, He + H 2 or He + H20. 3.

RESULTS

AND D I S C U S S I O N

3.1. Carbon formation from C H 4 Prior to the main study the experimental conditions for performing kinetic measurements were established. C H 4 conversions of below 10% (reactants diluted with H20) could be obtained at temperatures below 923 K. The carbon limits, which correspond to the maximum partial pressure and minimum temperature under which formation of carbon could be observed as a weight increase by the microbalance, were also determined. In addition, test runs at the most severe conditions, i.e. 1073 K and 4.3 bar c a 4, revealed that the support as well as the quartz and the alumina surfaces are suff'lcienfly inert, and that gas phase activation of CH 4 is not important. The effect of reduction in H 2 was also investigated by comparing the formation rate of carbon on reduced and unreduced catalysts at 2.3 bar C H 4 and 973 K. Average carbon deposition rates for reduced and unreduced catalysts after 100 min were about 0.5 and 2 mg carbon/gcat.-min, showing the importance of starting with nickel in its reduced state. Carbon is in this case defined as the weight increase observed by the microbalance, which is the net result of carbon formation and gasification. The TPR results shown in Fig. 1 indicate that promoters change the reduction behavior of the catalysts.

563

The complete reduction of NiO to Ni ~ for the Ni](x-ml203 sample would correspond to 0 a weight loss of 1.7 mg. Experimental values of weight loss for the unpromoted ,~ -0.5 Ni(Mg)/A catalyst range between 1.5 and 1.9 mg == -1 == indicating that NiO is nearly fully reduced. ~ -1.5 The reduction proceeds rapidly within a L .m narrow temperature range near 513 K in agreement with previous studies [12]. For the Ca and Mg promoted catalysts the ......... j -2.5 400 500 600 700 800 900 1000 3OO reduction starts at higher temperatures and Temperature [ K] the Ni(Mg)/~-A1203 catalyst is apparently the most difficult to reduce. The slow reduction process starts at a high tempFig. 1. Reduction profiles; Ni(X)/~-A1203 erature (673 K), and for this catalyst the (X= Mg, Ca). Conditions as in text. NiO phase can only be partly reduced. This could possibly be due to the presence of strong interactions between the MgO and the Ni rich phases [13]. The Ni(Ca)/cz-A1203 catalyst loses more weight than the unpromoted Ni catalyst. This could be due to loss of nitrate, loss of water or partial reduction of CaO. In fact, Fig. 1 indicates that the TPR spectrum consists of at least two peaks. It is therefore difficult to reach a definite conclusion about the degree of reduction of NiO for the Ca-promoted catalyst. Carbon formation from C H 4 w a s studied over the reduced catalyst at 873 K, 973 K, and 1073 K and partial pressures of methane of 0.45, 2.3 and 4.5 bars. The reaction order and the

1600

............................

400

.......................... ,.r

1200

S)

b)

E.

/

"

=.

E /

7.

- -

400

o

~_

- ~

~1oo

-"

I " -

o

0

50

100 150 200 Time on stream [rain.]

250

300

0

50

100

150 200 250 Time on stream [min.]

25

80

,

300

350

....................

.-:..

; r S ==

~

•40

~0~20

0

~ 20 r

4.46 bar

c)

0

~2.28 t,ar ,-

~

"-

1O0

d)

-==mm =1= ,,~,,= ,,,.,.,~ J.""

E~.15 ~ f_~'~

200 Time on stream [rain.]

....... - ....,,,,,.-.,,,~ o-

D

2.3 bar ~ 0.45 bar m== ~ 300

400

2.28 bar

.-

0

0

5

10 15 20 Time on stream [rain.]

25

i 30

Fig. 2 Experimental results and calculated data based on Eq.(3) for formation of carbon from CH 4 over Ni]a-A1203 at (a) 1073 K, (b) 973 K, and (c) 873 K; (d) initial data at 873 K.

564 activation energy were determined to 1 (0.965_+0.03) and 80 (79.31+0.44) kJ/mole, respectively. Initial rate data were used because of catalyst restructuring at higher carbon levels. Figs. 2a-d compare the observed carbon accumulation for the Ni/tx-A1203 catalyst with predicted values based on a simple model giving the rate in mg carbon/g cat, min: d[carbon]/dt

= 2363 9 e (-79310 J/R(T) , [ C H 4 ] 0"965

(3)

For the unpromoted catalyst it was observed (Fig. 2) that the carbon formation followed an autoaccelerating behavior. The autoacceleration is very weak and hardly observable at 873 K, gradually becoming more pronounced at higher temperatures. A simulation function was introduced in the model causing the rate to double every 350 min at 973 K, and every 100 min at 1073 K. The autoacceleration behavior could possibly be explained partly by a gradual increase in the accessible Ni area. Duprez et al. [ 14] observed that the accessible fraction of the metal surface increased during the initial stages of whisker formation. A period of particularly rapid accumulation of carbon from CH 4 was observed during the initial 5 min at 873 K. This effect (Fig. 2d) is favored by higher CH 4 pressure. This phenomenon is not observed at 973 K or 1073 K. Buoyancy is observed as a weight loss at the moment when CH4 enters the reactor and is thus excluded as a potential explanation for the rapid weight increase. Calculations show that the Ni/C ratio is very close to 1 at the end of this initial period (at 4.5 bar) indicating that a rapid saturation of the nickel phase could possibly be of importance. The effects of adding promoters to the base catalyst are shown in Fig. 3. Adding Ca and in particular Mg to the catalyst increased the rate of carbon formation quite substantially. In the case of Ni(Mg)/(-A1203, the basket was filled up with carbon after a short time. The conversion of CH4 through the catalyst bed is usually below 5 % except for the Mg promoted sample at high temperature where the conversion is much higher. Fig. 3 shows that the effects of alkaline earth promoters follow the same trend with time at both 2.28 and 4.5 bar CH4. It is also evident that while carbon accumulation follows an autoacceleration behavior for the Ni/o~A1203 catalyst a trend of deactivation is observed for the Ca and Mg promoted samples. This observed deactivation could indicate that formation of encapsulating carbon is important. The differences in the initial rates of carbon formation can partly be explained by the difference in available surface area of Ni, as measured by H2 chemisorption. On a turnover frequency basis the rates for Ni and Ni(Ca) are nearly identical. In addition, the TPR results (Fig. 1) indicate that Ni in the Mg-promoted catalyst has possibly been only partly reduced. As already mentioned, the rate of carbon formation on the unreduced Ni catalyst is high. i~~.......i

j' /

Z

-

/

,t

i

~'400

i

F

Ni( ;a)/AI2C |

"-

9~ 100

0

0

20

40

, 60 80 100 Timo on 8thrum [min.]

, 120

140

160

0

50

100

150 200 250 Time on stream [rain.]

300

350

Fig. 3. The effect on catalyst coking at 973 K of adding Mg and Ca to the Ni/a-A1203 catalyst.

565

3.2. Effects of hydrogen Effects of adding H2 to the C H 4 (+He) feed were also studied and the results for the unpromoted catalyst are presented on Fig. 4. Surprisingly, H 2 causes the observed rate of carbon formation from C H 4 to increase up to rather high partial pressures of H2. However, partial pressures of H2 higher than 0.9 bar results in reduced carbon accumulation. The mechanism by which H2 enhances carbon formation rate is not obvious. H2 may possibly lower the formation rate of encapsulating carbon, leading to a higher steady-state Ni surface area favoring filamentous carbon formation. Changes in the coverages of surface compounds could also affect the driving forces for surface diffusion of carbon, carbon solubility, carbon diffusion through the Ni bulk, and precipitation of carbon structures.

140 ,-:~.20

~.200

r

,,

E~,150

=e (= ~I00 -~9

P_H2 p(I.9 bar

_ H2 = 0.61 b= r

i ~ ~

~

0

50

1O0

150

T i m on stream (mln)

.~

60

p H2:0"41 bar P H2 = !).3 bar

5o

0

80

200

,E

g,o

H2 i 250

P_H2 = 1.1 bar P_N~) : 1 s ~ar J~

J..d

'~ P_H2= 13bar

,,=--..,,r

"-'r-r'-

o

0

50

1O0 150 Time on stream [rain.]

200

Fig. 4 The effect of H2 on the carbon formation on Ni/A1203 with 1.28 bar C H 4

250

at

923 K.

3.3. Effects of H 2 0 Effects of H20 were studied at 843 K, 883 K and 923 K at 3.5 bar C H 4. The experiments were carried out at differential conditions in the catalyst bed. Some results are given in Fig. 5. The observed rate of carbon formation is strongly dependent upon temperature, partial pressure of H20 and catalyst composition. The carbon limit was found to be close to 0.8 bar H20 for the unpromoted catalyst. Ca lowers the limit to 0.4 - 0.5 bar, whereas Mg increases the limit to above 2 bars. Carbon deposition in the presence as well as in the absence of H20 was found to follow an autoaccelerating behavior. In general, the carbon deposition gradually increased during the whole experiment. Exceptions were found at partial pressures of H20 close to the coke limit. T e m p 9= 8 8 3 K

200

,~

~ 160

,-:,. 100 .................................................................................................................

P_CH4 = 3.5 b zr I P_H20 (1): 0 bar

/ / / / ~ / 3

E12O I~)~. N 17 hc=r = (3): 0128bar

:

( 4 ) : 0 . 3 1 bar , ~ l . ^ - , ^ ,-__

~: 8o

._

"E

o) "~ 40

0

50

100 Time o n stream [ r a i n 9

150

200

-~ 01

o

0.4

0 9 1.2 H 2 0 Partial Pressure [bar]

....................J

1.6

2.0

Fig. 5. (a) Effects of H20 on the observed carbon deposition from C H 4 o n N i ] ~ - A 1 2 0 3 with 3.5 bar C H 4 at 883 K; (b) Effects of Mg and Ca on initial carbon formation rate at 923 K.

566 In these experiments up to 2 weight percent carbon was formed at a high rate during the first 10 min. After this period, the deposition rate rapidly declined causing the carbon deposition profile to follow the characteristic deactivating behavior. Small quantities of H20 increased the observed rate of carbon formation. This behavior was reproduced at all three levels of temperature. The maximum carbon deposition rate was observed at about 0.1 - 0.2 bar H20. 4. CONCLUSIONS Carbon formation during decomposition of C H 4 was studied at 873-1073 K and partial pressures of C H 4 in the range of 0.45-4.5 bar over Ni catalysts. Carbon accumulation on unpromoted Ni, thought to be highly filamenteous, typically follows an autoaccelerating trend. Adding Mg and Ca as promoters to the Ni/o~-A1203 catalyst, increases carbon formation rates. H 2 at partial pressures of up to 0.9 bar (at 1.28 bar CH4) effectively increases the observed rate of carbon formation from CH4; at higher PH2 carbon formation rate decreases, although carbon is still formed at 2.1 bar H R. Small amounts of H20 also increases the observed carbon formation rate, although a rate maximum is observed at about 0.15 bar H20 for 3.5 bar CH 4. Addition of Ca lowers the carbon formation limit (the partial pressure of H20 above which no carbon is formed), whereas Mg increases the carbon formation limit; the latter observation is linked to a lower extent of reduction of Ni in the Mg-promoted catalyst. 5. ACKNOWLEDGMENTS The Norwegian Research Council, Statoil and Norsk Hydro ASA, are gratefully acknowledged for their support of this work. REFERENCES

1. 2. 3. 4. 5. 6. 7. 8. 9. 10. 11. 12. 13. 14.

Solbakken, Aa., Stud. Surf. Sci. Catal., 61 (1991) 447. Rostrup-Nielsen, J.R., Catal. Today, 18 (1993) 305. Bartholomew, C.H., Catal. Rev.-Sci. Eng., 24(1) (1982) 67. McCarty, J. G., Hou, P.Y., Sheridan, D., Wise, H., in: Coke formation on Metal Surfaces, eds.: L.F. Albright, Baker, R.T.K., ACS Symp. Ser. 202, ChaP. 13 (1982). Rostrup-Nielsen, J.R., Catalytic Steam Reforming. In: Catalysis, Science and Technology. Eds.: J.R. Anderson and M. Boudart, Springer Verlag, Berlin, 1984. Menon, P.G., J. Mol. Catal., 59 (1990) 207. Trimm, D.L., Wainwright, M.S., Catal. Today, 6(3) (1990) 261. Demicheli, M.C., Duprez, D., Barbier, J., Ferretti, O.A., Ponzi, E.N., J. Catal., 145 (1994) 437. Blekkan, E.A., Holmen, A., Vada, S., Acta Chem. Scand., 47 (1993) 275. Scholten, J.J.F., Pijpers, A.P., Hustings, A.M.L., Catal. Rev. - Sci. Eng., 27(1) (1985) 151. Holmen, A., Schanke, D., Sundmark, G., Appl. Catal., 50 (1989), 211. Zielinski, J., J. Catal., 76 (1982) 157. Ruckenstein, E., Hu, Y. H., J. Catal., 163 (1996) 306. Duprez, D., Demicheli, M.C., Marecot, P., Barbier, J., Ferretti, O.A., Ponzi, E.N., J. Catal., 124 (1990) 324.

9 Elsevier Science B.V. All rights reserved Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

567

O n L i m i t a t i o n s o f R e g e n e r a t i n g an H Z S M - 5 C a t a l y s t for the M T G P r o c e s s A.T. Aguayo, A.G. Gayubo, J.M. Ortega, A.L. Mor~in and J. Bilbao Departamento de Ingenierfa Qufmica, Universidad del Pafs Vasco, Apartado 644, 48080 Bilbao, Spain.

The aging and combustion kinetics of coke deposited on an HZSM-5 zeolite-based catalyst in the MTG process have been studied. The kinetic study of coke combustion in air was carried out at 500-550~ in a differential scanning calorimeter, by following the evolution of the combustion products with on-line FTIR analysis. The results provide evidence for limitations on coke reactivity that can be attributed to the combined effects of several circumstances (e.g. bad oxygen-coke contact and heterogeneous distribution of coke within the zeolite crystal). The need is demonstrated for a thermal aging treatment which equilibrates reproducibly the coke prior to combustion. The aging of coke is also limited by a peculiar coke deposit in the microporous structure of the zeolite. 1.

INTRODUCTION

In spite of the importance of the regeneration step in the MTG process, there is scarce information about it in the literature [ 1, 2]. From recent literature, it is known that below 400~ coke is inside the zeolite crystals and is heterogeneously distributed between the straight channels of the zeolite and the intersections between straight channels and zig-zag channels [3 ]. This is a light coke, highly soluble in pyridine, with a high H/C ratio (between 1.75 and 2.65), the properties of which depend on reaction conditions, mainly on temperature [3, 4]. From the results of combustion of the coke released from the catalyst [5] (by catalyst destruction following the technique proposed by Magnoux et al. [6]), it can be established that catalyst regeneration is dictated by two factors: (1) the highly hydrogenated nature of the coke, which makes it highly unstable. This coke instability is important for the design of the regeneration step carried out in situ in the reactor. In addition to taking into account the coke profile along the adiabatic reactor, coke equilibration must be carried out in order for the same combustion kinetics to be applicable to catalyst in different locations in the reactor; (2) the microporous structure of the zeolite crystals in which coke is heterogeneously distributed. In a previous work [5] it was demonstrated that the kinetics of coke combustion in the MTG process, once the coke has been released from the catalyst, are similar to the kinetics of combustion of coke on other catalysts with similar H/C ratio. Nevertheless, coke deposited on catalysts in the MTG process has lower reactivity. In this work these two factors were studied with the aim of determining: (1) the effects of the coke H/C ratio on combustion and kinetics and (2) the role of the porous structure of the HZSM5 zeolite in limiting combustion. In this study, coke equilibration conditions for reproducible regeneration were sought and the kinetics of coke combustion were determined. Aging and combustion of coke were studied in a differential scanning calorimeter, where the gas-solid contact in a fixed-bed reproduces the operational conditions of the regeneration step in the industrial reactor. 2.

EXPERIMENTAL

The preparation and properties of the catalyst (prepared by mixing of a HZSM-5 zeolite with bentonite and alumina)have been detailed in a previous work [5]. The automated reaction

568

equipment (with isothermal integral fixed bed) has been designed for the uninterrupted operation in reaction-regeneration cycles, and has been described in previous papers [7, 8]. For the coke and the state of deactivated catalyst to be representative of the operation in reaction-regeneration cycles, the regeneration of a catalyst deactivated after three successive reaction-regeneration cycles and an additional reaction step was studied. The operating procedure included: (1) reaction at 350~ at a space time of 0.052 gcat h (g of methanol) -1 for a time on stream of 6 h and (2) regeneration in air at 550~ for 1 h. It was demonstrated that under the conditions used in the reaction-regeneration cycles the catalyst completely recovers its activity and selectivity. The last reaction step was carried out with two values of time on stream, 2 and 12 h, in order to obtain two different deactivation states of the catalyst. Activities corresponding to two deactivation levels, a = 0.79 and 0.20, were calculated using a proposed deactivation kinetic model [7]. After reaction the catalyst was purged with He (100 cm 3 min -1) at 350~ Deactivated catalyst samples were obtained by homogenizing the bed of catalyst, because it has been shown that the amount, and possibly the structure, of coke varies along the bed [3]. Coke combustion was studied by differential scanning calorimetry (Setaram DSC 111) at constant temperature: 500, 525, 540 and 550~ Coke combustion was followed by on-line FTIR analysis of the gaseous products (CO, CO2 and H20) in a Nicolet 740 FTIR spectrophotometer. The evolution of coke H/C ratio throughout combustion and the value of this ratio for the initial coke were calculated from the measurement of the combustion products. After a period of kinetic study at each one of the previously indicated temperatures, additional combustion at 600~ was carried out to total combustion. In this way, coke conversion and coke H/C ratio are referenced to total coke content. ,

3.1.

RESULTS E f f e c t s o f a g i n g o n c o k e H / C ratio

In Table 1 values of H/C ratio of 8 catalyst samples deactivated with two values of time on stream (2 h for Samples 1-4, and 12 h for Samples 5-8) are shown. Samples correspond to different aging conditions: 500 and 550~ and purge times with He (100 cm 3 min-1), 20, 30, and 60 min. The aging treatment of Samples 1 and 5 was carried out by following a temperature ramp from a reaction temperature of 350~ up to the combustion temperature, 550~ It can be observed that the H/C ratio of the original coke, which is different for 2 h time on stream (H/C = 2.05, Sample 1) and for 12 h time on stream (H/C = 1.70, Sample 5), decreases with aging until reaching an equilibration value between H/C = 1.05 (Sample 8) and H/C = 1.15 (Sample 4). In a previous work [5] it was demonstrated that for the coke released from the catalyst the effect of aging is more noticeable and H/C ratio is stabilized at a value close to 0.5. This difference in the results must be attributed to: (1) the poor contact of the He stream with the coke deposited and which blocks the internal channels of the zeolite crystals and (2) the fact that the coke chemisorbed on the active sites of the zeolite may be less pyrolysable than the released coke.

Table 1. H/C ratio of samples of deactivated catalyst. sample number, reaction time, h sweeping temperature, ~ sweeping time, min H/C ratio

1

2

3

4

2 ramp 350-550 20 2.05

2 500 30 1.75

2 550 30 1.35

2 550 60 1.15

5

6

12 12 ramp350-550 500 20 30 1.70 1.50

7

12 550 30 1.20

8

12 550 60 1.05

569

3.2.

Coke c o m b u s t i o n kinetics

In Fig. 1, as an example, the results of evolution of combustion products (CO, CO2 and H20) at 500~ for Samples 1 and 4, are shown. Sample 1 corresponds to the catalyst deactivated at 2 h time on stream, and Sample 4 corresponds to the same catalysts subjected to aging during 1 h. In Fig. la, it can be observed that for the most hydrogenated coke, Sample 1, with H/C = 2.0 5, the beginning of the combustion and the maximum peak of production of CO and CO2 correspond to lower values of temperature than those corresponding to the combustion of coke subjected to aging (Fig. l b), Sample 4, with H/C = 1.15.

0.1

,

0.08 "7 1._ E "7, 0.00

0.04

650

a

/

t

Sample 4 6O0

s50 ~

0.06

6O0

'c ._ E

~..

550 U o

6

~. 0.04 j

-,

~ '~

'~

u 0

450 450 ~~

'\

0.02

0

650

Sample 1

"400

350 0

10

20

30

40

time, min

50

60

70

~9 0,02

/'

,\

',

450

E

E ~

400 0

350 10

2O

3O

4O

lime, rain

50

60

70

Fig. 1. Evolution of coke combustion products at 500~ for Sample 1 (without previous aging) and Sample 4 (with aging). Conditions: Air (15 cm 3 min-]); sample weight, 20 mg. In Fig. 1a it is observed that total combustion of coke is not reached in 40 min, although it is very hydrogenated. For this reason, a second combustion step following a temperature ramp up to 600~ is needed in order to assure total combustion, which at 500~ takes several hours. For the combustion of coke at 525~ and at 550~ the need for the second step (temperature ramp up to 600~ total combustion of coke is also observed. This result can be explained by the hypothesis that the coke remaining after combustion in the 500-550~ range will be the least hydrogenated fraction of coke, which consists of biaromatics and higher structures, and which will be located in the intersections between straight and zig-zag channels of the zeolite crystals. This interpretation is based on the results of a previous study [8] in which it was shown that the catalyst totally recovers its initial activity after regeneration at 550~ for 30 min, which is evidence of the inert nature of the remaining coke. Copperthwaite et al. [9] have also demonstrated that the catalyst recovers its activity without complete coke combustion being reached. From the results corresponding to coke combustion of the samples in Table 1, the values of coke remaining in the catalyst, Cc, as a function of combustion time were obtained. These results have been fitted to first order kinetics with respect to each one of the reactants, the solid (coke) and the gas (02): - dCc/dt = kr Cc PO2 ( 1) or in integrated form: In (I-X) = -kr PO2 t (2) where X is the conversion of coke: X = 1 - Cc/Cco (3) and kr = Ar exp(-Er/RT) (4) In Table 2 the calculated values of activation energy, Er, frequency factor, Ar, and rate constant, kr, at 550~ are shown. It is observed that as the hydrogenated nature of the coke

570 increases, its reactivity for oxidation increases, the activation energy is lower and the value of the kinetic constant at 550~ is higher. The results of the effect of the H/C ratio upon reactivity for oxidation are qualitatively similar to those obtained for combustion of the coke released from the catalyst [5]. Nevertheless, for combustion of the coke deposited in the catalyst the range of values of the kinetic parameters is narrower, as a consequence of the range of values of H/C ratio also being narrower due to the limitations to aging of deposited coke. Table 2. Kinetic parameters and values of Thiele modulus (within the micropores of the HZSM-5 zeolite crystals, (hc, and within the mesopores and macropores of the catalyst, (hm) for coke combustion for samples of deactivated catalyst. sample number, 1

Ar, atm-lmin -1 Er, kcal mol 1 kr at 550~ atm -1 min-1 ~c x 103 ~m x 102

2

3

1.19 104 16.9

1.23 105 20.8

0.387 1.81 1.35

0.368 1.76 1.31

4

1.57 105 7.98 105 21.3 24.0 0.346 1.71 1.27

0.337 1.69 1.26

5

6

1.16 105 4.85 105 20.7 23.1 0.369 3.09 2.08

0.355 3.03 2.04

7

8.85 105 24.2 0.331 2.92 1.97

8 1.39 106

25.0 0.329 2.91 1.96

Values of activation energy in Table 2 are between 16.9 and 25.9 kcal mo1-1. These values and those of the rate constant in the range of temperature studied, are significantly lower than those corresponding to combustion of coke in other meso and microporous catalysts. In Royo's review [ 10], activation energies are in the 2 5 4 5 kcal mo1-1 range. The limitation to combustion cannot be attributed to the structure of the coke because, as has been demonstrated [5], the coke released from the catalyst is of high reactivity towards oxidation in agreement with its highly hydrogenated nature, with values of Er between 25.6 and 40.6 kcal mo1-1. 3.3.

Study

of pore diffusional

resistance

The possibility that the limitation in the kinetics of coke combustion has its origin in a high pore diffusional resistance for 02 in the channels of the zeolite crystal has been studied. Bilbao et al. [ 11 ] have carded out an extensive experimental study of regeneration of amorphous silicaaluminas with different granulometry and under different states of deactivation and have established a value of Thiele modulus, (h = 2/3, as a limiting value for negligible limitation to diffusion in the kinetic study of coke combustion. Other authors establish the limiting value of (h = 0.1 as criterion. The Thiele modulus is given by: (l):-~[k rRTCc~ De Me

1/2

(5)

where the effective diffusivity of oxygen has been calculated by: De = 3.36 105 e 2 / ~ ) 1 / 2

rpore,ave

(6)

Taking into account the bimodal structure of the catalyst of this study, in which microporous crystals of zeolite are agglomerated with a binder (bentonite) and with alumina (inert charge), both of high mesopore proportion, the limitation to internal diffusion of oxygen in both regions (in series ) of the porous structure has been quantified. For calculation of the Thiele modulus in the zeolite crystals, r eq 5 was applied with the following parameters: zeolite crystal radius, Rp = 3.1 10 -4 cm; temperature, 550~ density of the zeolite crystals, Pc = 1.73 103 g 1-1; molecular weight of coke, Mc = 300 [4]; molecular weight of oxygen = 16. The calculation of values of porosity and mean pore radius were carried out from CO2 adsorption-desorption isotherm for Samples 1-4 (time on stream 2 h): e = 0.29

571 2.60 10-8 cm, and for Samples 5-8 (time on stream 12 h): e = 0.26 and rpore,ave = 2.65 10-8 cm. The same value of coke content, Cco = 0.008, was used for Samples 1-4, and for Samples 5-8, Cco = 0.019. The values of effective diffusivity coefficients calculated for the zeolite crystals are: for Samples 1-4, De = 3.93 10-3 cm 2 min -1" for Samples 5-8, De = 3.22 10-3 cm 2 min -1. These values of diffusivity calculated by using the properties of the catalyst deactivated by coke are similar to values determined experimentally [ 12,13]. The Thiele modulus for the mesoporous structure of the catalyst, Cm, was calculated using the following parameters: particle size, R p = 0.0137 cm; mean pore radius, rpore,ave = 2 0 10 -8 cm; catalyst porosity, e = 0.52; catalyst density, Pc = 1210 g 1-1. N2 adsorption-desorption isotherms were used for measurement. The calculated value of effective diffusivity coefficient in the mesoporous structure of the catalyst is De = 9.71 10-2 cm 2 min -1. This value is not affected by coke deposition. In view of the calculated values of r and t~m (Table 2) the limitation to internal oxygen diffusion under the experimental conditions is even lower in the zeolite crystal than in the catalyst particle. The Thiele modulus in zeolite crystals for a severe deactivation state corresponding to activity = 0.20 [7] is r = 3.09 10 -3 (Sample 5), for an initial coke content of 1.9 wt% (approximately 40 wt% of the coke content needed for blockage of the internal zeolite channels). This low value of Thiele modulus is evidence that oxygen-coke contact is not limited by internal diffusion in the deactivated catalyst. The results obtained in this work strengthen the hypothesis that the limitation to combustion is caused by limited catalytic surface available for oxygen-coke contact due to coke location. Coke is distributed in two different positions, which determine its reactivity" (a) coke in the intersections between channels (partially soluble in pyridine [3]), which is less reactive in combustion than the average coke, as it has a lower H/C ratio, but offers a higher surface for contact with oxygen accessing through straight channels and zig-zag channels. This coke only blocks the intersections for extensive deposits, and it can be expected that this coke, initially with a mono or bi-aromatic alkylated structure, will acquire an inert structure for high values of coke conversion, which explains the fact that the presence of this residual coke does not affect catalyst activity. (b) coke within the channels (soluble in pyridine [3]). A great part of this coke blocks the internal channels of the zeolite crystals and only allows contact with oxygen at two clean surfaces at the extremes of the coke filaments. a n d rpore,ave =

.

sample 4 ,

- - - sample 1

1.6

(,.)

1.2

0.8

0.4

,

i

5

~

,

10

L

i

15

i

20

,

I

25

time, rain

Fig. 2. Evolution of the coke H/C ratio throughout combustion.

The heterogeneity of the coke has a similar effect on the limitation of the pyrolysis (cracking) for the aging treatment and of the cracking that occurs in parallel with combustion, which explains the decrease in the H/C ratio of the coke remaining throughout combustion (Fig. 2). The data in Fig. 2 were determined from analysis of combustion gases. The decrease in the H/C ratio of the coke deposited on the catalyst is less noticeable than the decrease observed in the rate of combustion of the coke released from the catalyst [5]. Carlton et al. [14] have shown that the linkage between acidic sites of the HZSM-5 zeolite with ether and carbonyl groups of partially oxidized coke plays an important role in the oxidation of coke. These linkages stabilize the hydrogenated components of coke and reduce its reactivity in cracking.

572 4. CONCLUSIONS Kinetic parameters (activation energy and rate constants in the range between 500 and 550~ for the combustion of coke with air in the regeneration of HZSM-5 zeolites used in the MTG process are lower than the values of these parameters for other acidic catalysts. This result can be attributed (1) to the peculiar coke deposit within the channels of the zeolite crystals, (2) the distribution of coke among the channels and the intersections between channels, and (3) poor oxygen-coke contact. The limitation of internal oxygen diffusion in the internal zeolite channels is not important when the regeneration is carried out with a coke content of 40 wt% of that corresponding to channel blockage and when activity is 0.20. Coke aging is limited by the same causes as in coke combustion and the coke H/C ratio decreases to a value close to 1.0 only in the care of a severe treatment (1 h in He stream at 550~ During combustion there is a decrease in H/C ratio of the remaining coke, which can be attributed to coke cracking together with a phenomenon of selective combustion of coke. By following the afore-mentioned coke stabilization treatment, it has been demonstrated that for initially different cokes (corresponding to a range of reaction conditions) values of 25.0 kcal mo1-1 and 1.39 106 atm -1 min -1 for the activation energy and frequency factor apply. Thus, these values can be used for the design of the regeneration step in the MTG process, which is carried out in situ in the adiabatic reactor itself. 5. A C K N O W L E D G E M E N T S This work was carried out with financial support from the University of the Basque Country (Project UPV 069.310-EB004/92) and from DGICYT (Project PB93-0505). REFERENCES

1. 2. 3. 4. 5. 6. 7. 8. 9. 10. 11. 12. 13. 14.

D.M. Bibby, R.F. Howe and G.D. McLellan, Appl. Catal., 93 (1992) 1. M. Guisnet and P. Magnoux, Stud. Surf. Sci. Catal., 88 (1994) 53. P.L. Benito,, A.G. Gayubo, A.T. Aguayo, M. Olazar and J. Bilbao, Ind. Eng. Chem. Res., 35 (1996) 3991. A.T. Aguayo, P.L. Benito, A.G. Gayubo, M. Olazar and J. Bilbao, Stud. Surf. Sci. Catal., 88 (1994) 567. J.M. Ortega, A.G. Gayubo, A.T. Aguayo, P.L. Benito and J. Bilbao, Ind. Eng. Chem. Res., 36 (1997) 60. P. Magnoux, P. Roger, C. Canaff, V. Fouche, N.S. Gnep and M. Guisnet, Stud. Surf. Sci. Catal. 34 (1987) 317. P.L. Benito, A.G. Gayubo, A.T. Aguayo, M. Castilla and J. Bilbao, Ind. Eng. Chem. Res., 35 (1996) 81. P.L. Benito, A.T. Aguayo, A.G. Gayubo and J. Bilbao, Ind. Eng. Chem. Res., 35 (1996) 2177. R.G. Copperthwaite, G.J. Hutchings, P. Johnston and S.W. Orchard, J. Chem. Soc., Faraday Trans. 1, 82 (1986) 1007. C. Royo, Ph.D. Thesis, Universidad de Zaragoza, Zaragoza, Spain, 1994. J. Bilbao, J.M. Arandes, A. Romero and M. Olazar, Chem. Eng. J., 35 (1987) 115. J. Kiirger, H. Pfeifer, J. Caro, M. Btilow, H. Schlodder, R. Mostowicz and J. V61ter, Appl. Catal., 29 (1987) 21. M. Btilow, J. Caro, J. V61ter and J. Karger, Stud. Surf. Sci. Catal., 34 (1987) 343. L.C. Carlton, R.G. Copperthwaite, G.J. Hutchings and E.C. Reynhardt, J. Chem. Soc., Chem. Commun. (1986) 1008.

9 Elsevier Science B.V. All rights reserved Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

573

A s y m p t o t i c Behavior during Sintering of Supported Catalysts Gustavo A. Fuentes* and Elizabeth Salinas-Rodfiguez Department of Process Engineering and Hydraulics, Universidad A. MetropolitanaIztapalapa, A.P. 55-534, 09340 M6xico, D.F., MEXICO. E-mail: [email protected]

By using scaling analysis it is shown that the presence of a steady-state particle distribution is compatible with standard mechanisms of crystallite sintering in catalysts if fragmentation kernels are added. Analytical solutions are found when aggregation and fragmentation kernels are homogeneous, but the results can be extended to kernels with different properties. This provides a theoretical justification for the use of steady-state dispersions during kinetic analysis of sintering using the generalized power law equation.

1. INTRODUCTION Although the detailed molecular mechanism of catalyst sintering is not yet fully characterized, there are two well established models commonly used to explain this phenomenon. They are based on migration and coalescence of whole crystallites in one case [ 1] and of adatoms [2-4] in the other. These models are frequently considered to be different, but the underlying physical processes are in principle similar once the size of the migrating entities is accounted for. In reality, it is difficult to accept that only one mechanism exists. The actual process of sintering in a given case is very likely a combination of adatom and crystallite migration, and the relevance of each route depends on the particular metal-support pair as well as the atmosphere and temperature employed during pretreatment and reaction. Fuentes [5-7] has observed that the kinetic forms and/or distribution shifts predicted by both crystallite migration or Ostwald ripening models in general do not correspond to the actual measurements. He proposed a generalized kinetic equation for sintering (GPLE) and determined that most sintering processes reported in the literature were first or second order. The key assumption in the equation used was that structural rearrangements during sintering tend towards a steady-state configuration and a resulting non-zero dispersion. The possibility of a non-zero limiting dispersion was not considered during development of fundamental models for crystallite coarsening. That seems to be the reason why limited kinetic expressions were previously determined. We have analyzed the issue and found that steady state solutions were indeed possible both in the crystallite migration and Ostwald ripening models by introducing suitable assumptions. These results give validity to the use of the GPLE beyond simple kinetic fitting. In the following sections we present an analysis through scaling techniques of the structure of the steady state solutions based on the Smoluchowski equation.

574 2. KINETICS, FITTING, AND MECHANISM The usual analysis of sintering kinetics involves empirical fitting of mean particle size, mean diameter, or dispersion versus time-on-stream data in order to obtain kinetic expressions. At the next level of understanding it is expected to establish a clear correspondence between a basic model derived from a reasonably fundamental mechanism and the global kinetic parameters. In this case order and activation energy have been frequently used as the guiding parameters. It is fundamental for the empirical equation used to fit macroscopic data to satisfy criteria derived from pure fitting schemes, as well as to be related to the fundamental mechanism proposed by adequate generalizations or lumping schemes. As a result, not all empirical fitting equations allow an adequate analysis of the processes that a particular system is undergoing. It is frequently found that processes involving decay, such as chemical reactions or excitation of diverse physical variables, are best fitted using power law equations (PLE). The order of the process and the value of the activation energy help in discriminating between different mechanisms. In the case of physical phenomena, exponential decay seems to be prevalent. For it to be applicable, the process must involve loosely interacting or noninteracting moieties, as well as closely grouped characteristic times. If that is not so, nonlinear phenomena appear. In the case of deactivation or sintering phenomena there are added complexities because the system is initially in a metastable configuration, reached during low temperature synthesis and/or in the presence of non-reactive gases. Once the catalyst is placed under conditions where its structure is labile, it will evolve towards an equilibrium configuration that is not yet predictable via standard analyses because the fundamentals of the phenomena involved are only partially understood. As an example, use of the Gibbs-Thompson equation in the standard Ostwald ripening model, has implicit the assumption that in all cases, irrespective of system or conditions, the crystallite distribution must evolve first towards a stationary distribution and then continue evolving until it collapses into a single hemisphere encompassing all metallic atoms, i.e. a minimum in overall surface free energy. Experimental observations, on the other hand, show that crystallite distributions do not turn into monomodal narrow distributions, not to mention a single particle, but remain rather broad, usually tailing in a direction opposite to the predicted stationary distribution. In spite of their extended use in the analyses of sintering processes, the values of process order (n) and activation energy (Easn) derived while using power-law kinetic equations (PLE) are at odds with experimental data obtained through other means. The estimated Easn are usually exceedingly large [6,8] and have no clear relation with measured or estimated atom emission energies, nor with the values characteristic of atom capture by metallic clusters. Also, the expected energies associated with breakage of metal-support bonds are not as large as the calculated Eas,. The presence of time-dependent n is also questionable, and in fact is in the majority of cases a mathematical artifact. There are also problems with the predicted particle size distribution. As mentioned, Ostwald ripening leads to a quasistationary distribution that tails towards small particle sizes, but that is frequently at odds with experimental results that tend to follow normal or log-normal distributions. Fuentes addressed the issue of deactivation kinetics in general [5], and of sintering kinetics in particular [6,7], and found an important discrepancy between experimental results and use of the PLE as a fitting tool. He observed that in most cases metallic dispersion reached a quasistationary plateau after an initial stage of fast changes in dispersion. That was at odds

575

with the built-in PLE asymptotic dispersion (zero under all conditions) as t ~ o0. That inconsistency between the measurements and the mathematical model explains why the apparent PLE sintering order becomes a function of the actual level of the asymptotic dispersion. Fuentes also showed that sintering kinetics could be well represented by a modified first- or second-order equation having a non-zero asymptotic dispersion, i.e. the generalized power-law equation (GPLE). Of course, this equation includes as a particular case that of Sss=0, i.e. PLE kinetics. The GPLE has also been used to tackle problems where dual sintering and coking were occurring, as well as poisoning. Bartholomew [8] applied the equation proposed by Fuentes to a large number of published data and confirmed that first or second order GPLE provided an excellent fit of the experimental observations. He also confirmed that activation energies for sintering were low. The GPLE in dimensional form is dS

(1)

_k.~.m.(S-S,,) m

dt

where (2)

k s m - k ..... 9e x p ( - ~ )

Application to sintering data has shown that m = 1 or 2, depending on the sample under analysis [6,7]. In non-dimensional form, the integrated equation for either order becomes:

s s-s,,

Z

S O - S------~.~=

s~

-So - So 1 - S------~,.,= exp(-k'~l" t) - e x p ( - r )

(3)

for m=l

So

and 1 l+(So-S,:~).k,.:.t

.

1

. . . l+k,z.t

1

for m=2

(4)

l+r

The general solution for m ~ 1 is 1

1

1

Z - (I + ( m - I). ( S O - S~.,)m-l. k,.m 9t) ~-m - (I + k,*m 9t) ~-" - (I + r) ~-"

(5)

where ksm*, the inverse characteristic time, and x, the non-dimensional time - both defined for any value of m;~ 1 - are given by

- ( m - 1). (So - s , , ) ' - ' , k~,. t - k~'~. t

(6)

576 Analysis of sintering data with an equation similar to Equation 3 has been done by Dautzenberg and Wolters [9], although no generality was claimed. The mechanistic implications of the GPLE have been only partially discussed. Standard models cannot be used to justify the use of a steady-state distribution because they were developed using only aggregation kernels. However, there is no fundamental reason why steady-state configurations do not exist as shown by Fuentes and Gamas [6] based on an analysis of the surface free energy corresponding to a crystallite distribution. The fact that steady state particle distributions occur has also been noticed in fields ranging from biology to astrophysics and has been generally dealt with by analyzing coagulationfragmentation of particles with the Smoluchowski equation. In the next sections we will present such an analysis using scaling methods and restricted forms of the equation to determine under what conditions steady state solutions are possible in the limit as t-->oo. The results are applicable to the full problem.

3. T H E O R E T I C A L DESCRIPTION 3.1 General Equation for the Distribution Function Any event where a spontaneous creation or annihilation of particles takes place is strictly speaking a probabilistic event. It is then possible to formulate a complete description of the aggregation process in statistical terms, i.e. using a probability distribution. In particular, we are interested in the time evolution of the probable number of particles. We define its distribution function as Ns(s,x,v,t), i.e. the number of particles with size* s located in the spatial range x and with velocity v at time t. Here, x and v are the vectors containing the elements of the physical and velocity spaces, respectively. The balance equation for Ns(s,x,v,t) can be derived phenomenologically by following Boltzmann's ideas [ 10] about the kinetic theory of gases. The balance includes as sources of clusters the changes of N~(s,x,v,t) resulting from particle growth, the motion of particles into and out of the spatial element, and the change in the number of particles in the velocity element by their acceleration [ 11 ]. We also consider the rate of change of the distribution function caused by particle transport resulting from diffusion, where D is the diffusion coefficient, and we include collisions with other particles, i.e. coalescence and fragmentation, through a term named J. Letting ds/dt denote the rate of change of the size s of a particle, and F=dv/dt the force per unit mass on a particle located at (s,x,v,t), the result is

cTV~

3

ds

cgt' + ~(N~.cOs,-~-)+ v -

VrN s

+F.VvNs = DVZN,+J

(7)

This equation describes the time rate of change of the distribution function. It cannot be solved analytically, except for special cases. However, depending on the assumptions made it simplifies to the models of crystaUite migration or Ostwald ripening, both in continuous or discrete form, as is shown in a forthcoming paper.

- We are following the standard derivation, where size is equivalentto mass or volume of each particle. We will indicate those cases where particle radius or diameter are used instead.

577 As a particular case we can consider a mean field description, where the diffusion and convection terms disappear. That is equivalent to assuming that changes in the size of the particles are only caused by coalescence and aggregation processes; that also implies that ds/dt=0. The J term is obtained by summing the interactions between particles over all possible sizes. The resulting equation is the generalized Smoluchowski discrete equation [12], which describes the time evolution of the number of clusters of size s, N,, at time t oo

(8) Here, i and j define the size of the interacting particles, such that i+j = s. K(i,j) and F(i,j) represent the rates of coalescence (i+j --~ s) and fragmentation (s --~ i+j) and their form depends on the diffusive behavior of the clusters and on the details of their interactions. In order to solve the equation it is necessary to consider specific rates of aggregation (see section 3.2). Recall that the moments of the N, distribution are given by

h4. (t) - ~ s'Ns(t )

(9)

s

where M0(t), the zeroth moment, corresponds to the mean number of clusters; the first moment Ml(t) represents the total mass, that in our case is a constant. The mean cluster size is defined as S(t) - M2(t)/Ml(t).

3.2 Aggregation and Fragmentation Processes The first formal study of catalyst sintering kinetics involving use of the Smoluchovski rate equation was done by Ruckenstein and Pulvermacher [ 1], following the work of Friedlander on aerosols [13]. In that case irreversible coalescence was assumed to be the only cause of changes in the particle distribution, and different specific kernels were assumed. It was implicit that spatial fluctuations were not affecting the kinetics of aggregation. A power law kinetic equation was derived, and correspondence between sintering order and various limiting regimes was obtained by fitting S versus time data from specific simulation runs where the structure of the rate constants was modified. Based on the fits obtained it was indicated that 4 <_n _< 8 corresponded to control by diffusion on the surface, whereas 2 _
578 structure of the steady state distribution. The usual procedure involves the assumption that the kernels are homogeneous functions of their arguments, i.e.

K(ai, aj) = a ~K(i, j)

(1 O)

F(ki, kj) = k~F(i,j)

(11)

where 13 and a indicate the homogeneity degree and a and k are the coalescence and breakup constants. This scaling form is satisfied by most physically relevant kernels. Family et al. [15] extended the scaling assumption for the particle-size distribution of irreversible aggregation to cases involving also reversible aggregation. Using part of their derivation, let S(k,t) be the mean particle size defined by

S(k,t) ~ Z.~.sZN~.(k,t)

(12)

where the distribution was previously normalized to give Ml(k,t) = 1. The scaling behavior of Ns(k, oo) becomes a function of the mean particle size S(k, oo) if it is also supposed that

N~.(k,t)- s-z f (S(ks't) )

(13)

where fix) depends on the details of the process. It is expected that S(k, oo) has a dependence on the breakup constant k of the type s ( k , oo) - k -~

(~ 4)

because the average size of the clusters should decrease with increasing values of k. The scaling form of the steady-state particle size distribution hence becomes

N~.(k, oo) = s-Zf (sk y)

(15)

and the total number of particles scales as

N(oo)= ~-'~Ns(k, oo)- k y

(16)

s

Family et al. [15] proposed that the aggregation and fragmentation kernels satisfied the scaling forms given in Equations 10 and 11 with a = k = ~, and 13 = 2-c0. It was also assumed that the Smoluchowski equation was invariant under a scaling transformation k --+ )~'k and s --~ U's. As a result, in the steady state limit of the generalized Smoluchowski equation, i.e. if fragmentation is included, it follows that y = ( a + c o + 2 ) -1

(17)

579 This explicit form for y can be tested numerically. A consistent scaling function is

f (x) =

x 2-~

exp(-cx)

(18)

z is the crossover time when the balance between the aggregation processes is established. The validity of this approach was confirmed by Montecarlo simulations of the full equation using a constant coagulation kernel and a breakup probability equal to k(i+j) ~, where ct and k were adjustable parameters [ 15]. Spatial fluctuations were compensated by cluster breakup and the generalized Smoluchowski equation had a critical dimension dc < 1.

0 . 4

_

. . . .

I

'

'

'

'

I

'

'

'

'

I

. . . .

_

-

0 3 9

-

-

::5

9

-

-

0.1-;/ 0.0

-

'

.

-

0.2Z

. .

o~=O.Ol

" '

o~=0.05 o~=0.1

'

o~=0.5

"

.........

.

o~=1

9 //-\i

.\

.

.

.

.

i

i

i

~

0

5

10

15

.

.

20

Particle radius, a.u. Figure 1. Steady state particle size distributions obtained through scaling analysis. Equations 15-18. t0=0, k=0.0001, z=O, c=0.05 Figure 1 shows the shape of the steady state particle distribution in terms of an apparent particle radius. The predicted particle distribution is nearly symmetrical, although a small tail towards larger particle radii is present. This is not surprising given all the symplifying assumptions made in order to get an explicit solution. However, the shape is quite different from the Ostwald ripening stationary distribution, and is closer to the experimental distributions. It is important to remember that the solution given by Equations 15-18 corresponds to the asymptotic particle size distribution, meaning that sintering processes given by the Smoluchowski equation stop evolving in time once they reach that distribution. 3.3 Coagulation via Monomers The analysis presented in section 3.2 does not restrict the interactions between particles to any particular size range. That means that in principle all particles are allowed to move on the surface. Under some conditions that is not plausible, and only monomers or dimers may

580 diffuse. The analysis that we are presenting can also be applied to those cases. As an example, in the case where only monomers are mobile [ 17], we can base our derivation on Equations 10 and 11, supposing that for s = i+ 1,

K =K~.s p

(19)

F. = k. s q

(20)

where p and q are arbitrary exponents. From the definition of the zeroth moment, an equation of motion can be obtained [ 17] dM

o

dt

- ( s - 1) p N1N~_ 1 - Nss -q

(21)

In the steady state, M0(s,~) is constant, and a detailed balance condition results in

M,(oo), = sq[(s - I)!] p+qN;

(22)

The moments Ms are then given by M,,(oo), = ~-'~n"+q[(n - I)!] p+qN~"

(23)

n

The existence of solutions and their structure depends both on the sign of p+q and on the particular values of p and q. It can be demonstrated [ 17] that for p+q < 0 there is a steady state solution given an initial concentration of monomers M; whereas for p+q > 0 the solution diverges. For p+q = 0 and p=0, q=0, implying constant fragmentation and coalescence, and if M is large, there is a steady state solution given by S ---> 2 M v2

(24)

N,,. --> exp(- s / M '/2)

(25)

The case p+q = 0 with p < 1 also exhibits scaling behavior. The limit for large M, given also the condition MI - M, is

.

,

p

M. ~ [F(n + 1 -

">1 p ) / r ( 2 - p){(,+,-,)/(z-n}]. M{(,+,-,)/(2-n}

(26)

(27)

581 Figure 2 shows particle size distributions for different values of p. We determined that the shape of the distribution changes abruptly above p = 0.6 and that p > 0.9 results in instability in the solution. The shape is similar to the distributions shown in Figure 1.

12_~,, ,, i,,

|

t

|

i

i

i

i

I

I

p=0.1 1

10

L_

I

p=O.5 p=O.6

6-

V

Z

42,,,

I

l

0

I

I

I

2

~

I

4

i

i

I

i

I

I

I

6

8

Particle Radius, a.u. Figure 2. Particle size distribution when only monomer migration is possible. Equation 26 with p
InM

(28)

I. When M>>I

(29)

S --~ ]p + q[

and In

M~

N, -+ IP + ql)

P+q

(30)

582 4. DISCUSSION AND CONCLUSIONS The scaling approach for homogeneous aggregation and coalescence kernels gave different asymptotic solutions resulting in an equilibrium state in which the crystallite size distribution during catalyst sintering and its moments became independent of time. This equilibrium state for sintering processes was proposed previously by Fuentes and Gamas [5]. The results just presented show that steady state particle size distributions are possible for a wide variety of coagulation-fragmentation conditions. In the case of homogeneous kernels the shape of these distributions is a slightly distorted Gaussian, and it approaches experimental results. In spite of the fact that the analytical solutions derived in this work are necessarily restricted, it can be expected that solutions with shapes farther from normal distributions are possible once the fragmentation and aggregation kernels are modified. Although it was not shown here, a general cluster size distribution in equilibrium can be obtained using a different approach [18, 19]. It involves a stochastic description for the aggregation-fragmentation system given by the master equation of a probability balance. The equilibrium probability then follows from the detailed balance. That work is under way. There is the question as to whether and how the derivation presented here relates with actual processes resulting in catalyst sintering. To begin with, experimental reports do point out to the existence of steady-state particle distributions with varying shapes [20,21]. Interestingly, the aggregation-only crystallite migration model [1]cannot predict stable distributions because crystallites can only grow. The predicted invariant distribution in Ostwald ripening models [2,4] is extremely different to experimental distributions, a fact that is ot~en overlooked. Our results can be viewed as a first order approximation to the complete problem, because we simplified the kinetic structure by assuming the kernels to be homogeneous. Notwithstanding that, the predicted equilibrium distributions do resemble experimental data, validating our approach. The existence of fragmentation during crystallite sintering is not obvious based on the studies reported in the literature. The arguments used to include that phenomenon in the stochastic analyses employed in other fields basically stem from the fact that fluctuations become more important as the size of the aggregates increase. Under those conditions it is then possible for smaller fragments to leave the larger ensembles. The possibility of fragmentation is intimately related with the existence of a critical particle size. That corresponds to the minimum cluster size that is stable against fragmentation on the surface. Under ideal conditions the critical radius corresponds to one atom, but as the mismatch between the cluster and the support increases, as well as the temperature, that number increases significantly, reaching the thousands of atoms [22]. This results have been obtained during film growth in ideal systems, but apply to catalyst sintering. It is clear that particulars about cluster-support and cluster-cluster interactions necessarily control both kinetics and the equilibrium distribution in catalyst sintering. What support and metal are being studied would in principle determine the structure of the aggregationfragmentation kernels, including size effects upon mobility and particle-particle interaction. Unfortunately we are at a stage where our predictive capabilities are still being developed, and we have to infer the kernels from global measurements of kinetics and equilibrium distribution data, as is the aim of the present study. By trying to have coherence between the global analysis and fundamental models we expect to set new guidelines concerning the rate controlling steps and ways to control sintering.

583 A formal derivation of steady state particle size distributions during catalyst sintering has not been made before. Their existence adds validity to the use of steady state dispersions during the kinetic analysis of sintering data. At the same time, this provides an interesting perspective for reanalyzing the present ideas about sintering mechanisms.

ACKNOWLEDGMENTS

We acknowledge discussions with Prof. Rosalio Rodriguez and the support of UAMIztapalapa

REFERENCES

Ruckenstein, E. and B. Pulvermacher, AIChE J. 19 (1973) 356. Chakraverty, B. K., J. Phys. Chem. Solids 28 (1967) 2401. Flynn, P.C. and S.E. Wanke, J.Catal. 34 (1974) 390. Wynblatt, P. and N.A. Gjostein, Acta Metall., 24 (1976) 1165. Fuentes, G.A.,Appl.Catal. 15 (1985) 33. Fuentes, G.A. and E.D. Gamas, in Catalyst Deactivation 1991, vol. 68, Stud. Surf. Sci. Catal., C.H. Bartholomew and J.B. Butt, eds., Elsevier, Amsterdam (1991), 637. 7. Fuentes, G.A., A ctas XV Simp. Iberoamericano Catfilisis, C6rdoba, Argentina, E.Herrero, O.Anunziata y C.P6rez, eds., vol. 1, (1996), TI- 105-114. 8. Bartholomew, C.H., Appl. Catal. A.-General 107 (1993) 1. 9. Dautzenberg, F.M. and H.B.M. Wolters, J. Catal. 51 (1978) 26. 10. Harris, S., An Introduction to the Theory of the Boltzmann Equation, Holt, Rinehart and Winston, Inc., New York, 1971. 11. Friedlander, S.K., Smoke, Dust and Haze, J.Wiley, New York, 1977. 12. von Smoluchowski, M., Z.Phys. 17 (1916) 585. 13. Friedlander, S.K., J. Meteorol. 18 (1961) 753. 14. Barrow, J.D., J.Phys. A 14 (1981) 729. 15. Family, F., Meakin, P., Deutch, J.M., Phys.Rev.Lett. 57 (1986) 727. 16. Ernst, M.H. and van Dongen, P.G.J., Phys.Rev.Lett. 59 (1987)363. 17. Blackman, J.A. and Marshall, A., J.Phys. A 27 (1994) 725. 18. Hendriks, E.M., Z.Phys. B 57 (1984) 307. 19. Salinas-Rodriguez, E., Ph.D. Thesis, UAM-Iztapalapa, Mexico, 1992. 20. Grankvist, C.G. and Buhrman, R.A., J.Appl.Phys. 47 (1976) 2200. 21. Kuo, H.K. and DeAngelis, R.J., J. Catal. 68 (1981) 203. 22. Zhang, Z. and Lagally, M.G., Science 276 (1997) 377. 1. 2. 3. 4. 5. 6.

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Elsevier Science B.V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

585

Sintering and R e d i s p e r s i o n of S u p p o r t e d Metals" P e r s p e c t i v e s f r o m the Literature of the Past D e c a d e Calvin H. Bartholomew Department of Chemical Engineering, 350 CB, Brigham Young University, Provo, Utah 84602, USA Studies of the sintering of supported metals over the past decade have emphasized (1) the use of sophisticated spectroscopies and surface-science techniques in combination with industrially-relevant and model catalyst systems to better understand the mechanistic aspects and (2) the analysis of available kinetic data for sintering with more physically-meaingful models. The recent experimental work reveals the complexities of sintering and redispersion processes, which include for example, dissociation/emission of metal atoms from crystallites, diffusion of atoms and crystallites across support surfaces, wetting and spreading of crystallites in oxygen, rupture of films and nucleation of particles in H 2, coalescence and/or bridging of two particles, capture of atoms by particles, epitaxy of metal clusters with support surfaces, restructuring and faceting of metal surfaces by different atmospheres and the formation of compounds such as surface aluminates in steam or 02. Sintering rate data obtained over the past four decades have been reanalyzed using the new General Power Law Expression (GPLE), which provides the capability of treating these data in a consistent, unifying fashion such that quantitative comparisons of effects of reaction conditions and catalyst properties are possible for the first time.

1. INTRODUCTION, LITERATURE, APPROACH Supported metal catalysts are widely used in petroleum processing, chemical synthesis, and pollution control. In many of these applications these catalysts are subject to high temperatures during use or regeneration. At these high temperature conditions, the activity of a supported metal may decline as metal surface area and/or support surface area decrease in a process called sintering. Metal surface area can for selected supported metals be restored by a sophisticated chemical/thermal treatment called redispersion. From previous experimental studies of sintering [1-3] it is evident that sintering and redispersion are strong functions of temperature, time, atmosphere, and support. Sintering/redispersion rates are also significantly affected by choice of metal and/or promoter, metal loading, and catalyst preparation. Sintering of supported metals involves complex physical and chemical phenomena including dissociation/emission of metal atoms from crystallites, diffusion of atoms and crystallites across support surfaces, spreading of particles, wetting of the support by particles, nucleation of particles, coalescence and/or bridging of two particles, capture of atoms by particles, liquid formation, vaporization of metal atoms, and volatilization of metals as complexes. The importance of each of these different phenomena may vary with reaction condition (especially temperature and atmosphere), time, and catalyst formulation (e.g support and/or promoter). Thus, the treatment of sintering processes requires an understanding of surface diffusional processes, interfacial phenomena, solid-state reactions and the energetics of metal-metal and metal surface interactions. Most of these processes are only qualitatively understood. Accordingly, the investigation and simulation of sintering phenomena are difficult tasks.

586 Nevertheless, efforts to understand, treat, and model sintering/thermal-deactivation phenomena are easily justified. Indeed, catalyst deactivation problems greatly influence priorities in research, development, design, and operation of commercial processes. While catalyst deactivation by sintering is inevitable for many processes, some of its immediate, drastic consequences may be avoided or postponed. If sintering rates and mechanisms are known even approximately, it may be possible to find conditions or catalyst formulations that minimize thermal deactivation. Moreover, it may be possible under selected circumstances to reverse the sintering process through redispersion. The purpose of this review is to summarize briefly what has been learnedfrom experimental studies of supported metal catalysts regarding the mechanisms and kinetics of sintering. A monograph [ 1] and recently published reviews [2-4] provide more comprehensive analyses of kinetic data and mechanistic information obtained from model supported catalysts [2] and commercially-relevant real supported metal catalysts [3,4]. The discussion in this paper focuses on the effects of temperature, atmosphere, and catalyst properties on the kinetics of sintering of the latter group of catalysts.

2 . EXPERIMENTAL STUDIES OF SINTERING AND REDISPERSION R E A L SUPPORTED METAL C A T A L Y S T S

OF

2 . 1 Sintering of Supported Metals Sintering of supported metals typically involves (a) the loss of metal surface area due to growth/agglomeration of metal crystallites, (b) the loss of support and metal areas due to collapse of smaller pores, and (c) phase transformations such as reactions of the metal with the oxide support. Most of the previous work has focused on the growth of metal crystallites; studies have emphasized measurements of metal dispersion as a function of temperature (at fixed, short times, e.g. 1-2 hours) or as a function of time over long periods of time (e.g. 30500 h) at constant temperature. The former type of measurement provide qualitative data on the thermal stability of various supported metals in different atmospheres (e.g. metals are generally more stable in reducing atmosphere relative to oxidizing atmospheres, Rh/silica is more thermally stable that Pd/silica in H 2 atmosphere, and Pt/alumina is more stable in H 2 than is Pt/silica); these observations are discussed in recent literature reviews [ 1-4]. The latter type of measurement (dispersion versus time at constant T) provides useful rate data for sintering processes that can be used in design and modeling of catalytic processes. For many decades available dispersion versus time data were fitted to a simple power law expression (PLE)-d(D/Do)/dt = k s ( D / D o ) n, where D = metal dispersion at any time t, Do= initial dispersion, k s = the first order sintering rate constant, and n the sintering order. The PLE, however, does not provide an accurate representation of sintering data, since metal particle size and hence dispersion ultimately approach limiting values rather than zero at infinite time. Using the simple PLE, sintering orders vary from 2 to 17 as a function of time and temperature; thus, it is not possible to compare with any validity the kinetic parameters determined from this empirical equation. Nevertheless, recent studies [ 1-4] provide a summary of most of the previously published sintering rate data fitted to 2nd order kinetics based on a unifying general power law expression (GPLE) developed by Fuentes et al. [5,6] for which the rate equation is of the form -d(D/Do)/dt = ks (D/D o - Dea/Do) m (where Deq is the limiting dispersion at infinite time and m is the sintering order); the GPLE has been shown to accurately fit available sintering data [1-6]. Using these new kinetic parameters for sintering, it is possible to discuss quantitatively, the effects of reaction conditions such as temperature and atmosphere and of catalyst parameters such as metal, support, and pore size on sintering rate. Table 1 summarizes some of the important trends that follow from analysis of previous kinetic data by the GPLE. Sintering rates of a given material are generally more greatly affected by temperature than any other variable, although atmosphere (usually next in importance) is sometimes the dominant factor in determining rate and extent of sintering. Table 2 summarizes some of the important experimental observations regarding sintering of supported metals from studies performed in the past decade [ 1-4,7-28].

587 Table 1. Effects of important reaction and catalyst variables on sintering rates of supported metals based on GPLE data [ 1-6]. Variable ' Effect temperature i Sintering rates are exponentially dependent on T; Eact varies from 30 to 150 . kJ/mol. Eact decreases with increasing metal loading; it increases in the ..................................... kf~176176176

atmosphere .....................................

~. .......

metal

.........................................................................

i Sintering rates are much higher for noble metals in 02 than in H2 and higher noble and base metals in H 2 relative to N 2. Sintering rate decreases i supported Pt in atmospheres in the following order: NO, O), H), N). ,Observed order of decreasing thermal stability in H 2 is Ru > Ir -_- Rh > ,thermal stability in 02 is a function of (1) volatility of metal oxide and ,,.,.,,~ ..................................................................................................

,., .....................................

..................................... i..s.t.re.n~.t.~..~176176

support .....................................

..~ ............................

Pt; (2)

.......................................................................................

r

..........................

~, .....

~ ................

~ ..........................................................................................................................

..................................... !...s..r..~.e..`..`traP.P.!.n~..~a~enr.sZt~a~t..~e.crease..s.!.nte~.n.~..r.at.e.

pore size

~ .......

i Metal-support interactions are weak (bond strengths of 5-15 kJ/mol); with a few exceptions, thermal stability for a given metal decreases with support in the i following order: AlgOa > SiO9 > carbon. i Some additives decrease atom mobility, e.g. C, O, CaO, BaO, CeO 2, GeO2; others increase atom mobility, e.g. Pb, Bi, C1, F, or S. Oxides of Ba, Ca, or

~. .......................

promoters

,,, ........

for for

.......................................................................

i Sintering rates are lower for porous versus nonporous supports; they decrease i as cgstallite diameters approach those of the pores.

Table 2. Important recent experimental observations regarding sintering of supported metals. Catalyst i Atm./condition i Observation(s) i Ref. Ag/0~-A1203 i C2H4/O2, rxn.T i Growth of Ag and A1203 crystallites; no interaction; i 9 D

....................................... i ................................................ La~ ~.!..P.r.~176

Cu/Cr203 .......................................

~ ................................................

Fe/I~AloO~ "'Ve~'~'r~6~

:..e :.~.:..Ea.c...b..P_r.~176 .te..s in t..erin.~. 9.................. J................

i C2HsOH dehydr, i Highest activity and stability for Cr/Cu = 4/40. Metal i 10 at 250-3 IO~ i crystallite growth follows 2rid order kinetics. i Ho/No, 400~ i Surface AloO~ prevents, K promotes sinterin~, i 11 ~ .........................................................................................................................................

.............. i"W'~'S"r'x'n:'c~'ii'~/:'i

.... i"i'n'cr~"Tn"i~e~'"i3"~'r~'"gi'ze"~'a'"~i~'re~e'"in"~'~X"~v~r'i

9

.

i 400-450~

l ................

.-

..... "i'~ .....

i 150 h.

' ~ ~ ; c o r ~ i ~ ? ; i ~ i ~ i ~ ~

......... i"'~'i'n/'Sri'n~"occu~"~"~'i"~sia~iie"'~row'/~"'~a'"s'ui~fio~'i

..... "i'~ .....

.......glass ...................~................................................i...~e.a...~.~...s...s...w!t..h....b..uri~....~..f...Ni.i..n....t..h..e...s.u~p........~..~. .................................i................ Ni/SiO2 i H2/CO = 2, 230~ During CO hydrogenation metal surface area declines i 14 i -by migration of subcarbonyls; growth process is l " i sequential: monomodal Ni --4 bimodel Ni ~ large i i i faceted c~stal with [111] p..lanes, i ' ~ a ~ ~ ~ ~ .............T"~7"~5~"~e~ox"c~'-'" ....r'~'rn'~r~'n~'~'~'6a'"ai'"~Ci"'i3~slv~ea"'~'aiaa~si'"c~"i ..... i3 .... ....................................... ~..~.!in~.:...!.~:..3.~.S.c.. 9

Ni/MgO

.... L..b..e..re.~.!spe.......r.se.~...b..r...be.atin~...at..1....Og.....3..~?...C....in..~:

iO2, 400~ H2, ! 400-1000~

....................................... i ................................................

Pt/A1203, Pt/SiO 2 FJgi6~

...................... i ................

i Li addition increases sintering rate of NiO during i 16 i calcination; Li promotes reducibility of Ni; at Tred ->i i...6.~., o...c....~i..fac.!~it.at.e.s.~.~a.t.ic..c..o.~aP..se...o..f.~re.s:

..................... i ................

i O9, 573-873 C; i Pt oxidation limited to surface layers for all crystallite i i H2,O 2 redox, TPR i sizes; redn. of highly-dispersed oxidized crystallites i i i difficult due to lack of metal sites for H~ dissociation, i o

..................T6~;~iq;~6~

i 600~ 9 i

17

.....i~/i6;~~~iow ....~i~ii~/~ ....~'~'r'~G'/i""i'~ ....~'u'~'~'~'G'/i'i"i ........i"~'" i during c 2 n 4 oxidation Pt crystallites grow rapidly,

i

i possibly due to migration of volatile metal complexes i ". formed by reaction of free radicals with the metal. ".

588 . .

9

Pt/A1203, ]ZrO2,/CaO

, ,

,

! O2/N2/I-'I20, 500- i Sintering rate increases with increasing temp., Po2" i ' and P.2o; Pt stability decreases for supports in the l ! order ZrO2/La20~ >> CaO > A1203. i redox cycling in i Sintering at relatively low temperature is accelerated i C2H4/O2/N2, i by rapid cycling between O2-containing and O2-free i 300~ ! atms" mig.ration of PtOxCly. sp.ecies is sug.g.ested, i H2,OE,CO2,H20an Sintering rapid in 02, moderate in CO 2, slow in H E i 500-800~ i or H20; Pt/CeY sinters more rapidly than Pt/LaY in i i 07; sintering of Pt/LaY is more rapid in other gases, i Ar,H2,O2,CO,NO I Pt/y-A1203 stable in H 2 or Ar at 500-700~ for 30-60 i at 200-700~ i min; rapid sintering in O 2 at 600-700~ and in 0.1% i ,~NO at500~ rate order: N O > 02 > H 2 > Ar.i i Formation of volatile PtOx in O2 or NO postulated, i H 2, 350-500~ ~ D = 0.6-1.0; D increases with increasing support i 12-36 h .acidity; highest resistance to sintering for carbon i

.......................................................................................

Pt/Al203

, .................................................................................................................................

.........................................................................................

Pt/CeY,/LaY

, ..............

.......................................................................................

Pt/~-A1203

~ .......

.........................................................................................

1% Pt/carbon

............................................................

Pt/Rh/A1203/ ..c.eg

Pt-Rh gauze

~ .......................

~ .....................................................

~ ............................

. ..........................

~ ................

n .................................

g.r.ap..

t!.zat.!o.n.

.........................................

........................... L.ex..h.aust;..re

1% O2]N2, 1% H2/N 2 at 600800~

~ ..................

uc.e.

..c.ata!y

..st..is..m.or.e..a.ct!.v.e:.

24

.................................. i ..................

! i i i

25

................................... ! ..................

'

i...o.f.t.h..e..at.mo.sz.h.ere.:

NO/CO/He

23

i ..................

i ................................................

RNA1203 Rh/SiO 2

22

~ ..................

i Rapid loss of dispersion within 2 hours at 600, 700, i i and 800~ aging in 02 at same temperatures causes i :rapid and complete oxidation to Rh203; sintering i 9during high frequency reduction/oxidation is slower i i than in p.ure g.ases, i i Ce retards reductive agglomeration of Rh atoms to ~ 'Rh~ clusters; NO redisperses Rh, while CO causes! ~ volatilization. Microstructure is a complex function i

........................................................................................

21

i Sintering more severe in oxidizing versus reducing i

H2/O2 mixtures; at i Gas etches Pt-Rh gauze forming rough sponge-like 600~ 40-46 h i surface containing Rh203; there is no evidence for selective removal of either Pt or Rh; particle-like surface possibly formed by interaction of gaseous

Rh/ml203

20

,~, . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .

........................................................................................

Raney Ru

.~ . . . . . . . . . . . . . . . . . .

, ....................................................................................................

Simulated auto

........................... exhaust

.......................................

19

E900~

, ...................................................................................................................................

.......................................................................................

26

~ .................

27

i .................

" in vacuo or H20 ati Ru SA decreases from 66 to 16 m2g 1 when heated i ,200-600~ i to 673 K; A1 is oxidized to A1203 which migrates to i : :' surface.

28

Data from the more recent studies (Table 2) have emphasized the complexities of sintering processes; they provide evidence that growth processes and catalyst microstructures are complicated functions of temperature, atmosphere, catalyst components (metal, support, promoter), and catalyst preparation/pretreatment. Studies of sintering under reaction conditions have received particular emphasis. Another important quality of the recent work is the emphasis on the application of sophisticated methods and spectroscopies, e.g. H 2 chemisorption, temperature-programmed desorption and reduction (TPD and TPR), x-ray photoelectron spectroscopy (XPS), high-resolution transmission electron microscopy (HRTEM), x-ray diffraction (XRD), and small-angle x-ray scattering (SAXS), to study changes in surface area, surface chemistry and composition, and particle morphology. Thus, a much more detailed understanding of sintering of supported metal catalyst has been afforded.

589 3 . M O D E L C A T A L Y S T STUDIES The literature treating sintering and redispersion experiments with model-supported metal catalysts is significant. The results of over 35 such studies, reported in the 16 year period from 1973 to 1988, have been summarized in recent reviews by Baker et al. [ 1] and Bartholomew [2]. Relative to study of high surface area metal-support systems, the study of model supported systems has the advantages of: (a) the direct examination of metal crystallite location, size, shape, and even chemical structure as a function of time, temperature and atmosphere, (b) direct observation of crystallite migration, wetting, and spreading as it occurs, and (c) chemical analysis of the bulk and surface of individual metal particles. At the same time, there are clearly some important limitations to the use of model catalysts: (a) the surface structures of flat nonporous substrates and sometimes their surface chemistries are significantly, even greatly different from those of high surface area supports and (b) surface metal concentrations of model catalysts typically exceed those of commercial supported metals. Since support porosity, surface chemistry, surface roughness, and metal concentration are important factors in determining sintering rate, the results of model catalyst studies generally have qualitative rather than quantitative value. Instrumental techniques successfully applied to investigation of model supported catalysts include TEM, CAEM (controlled-atmosphere electron microscopy), TPD, EXAFS, AES, XPS, STM, XRD, and M6ssbauer Spectroscopy. Analytical HRTEM and same area studies have demonstrated great promise for advancing mechanistic knowledge in this area of science. Based on previously reported model studies of sintering it is evident that at a given temperature atmosphere has an overwhelming influence on sintering/redispersion behavior of supported base and noble metals. Sintering of metals is relatively slow in H2 or N2 atmosphere while more rapid sintering or dispersion may occur in O2-containing atmospheres depending on conditions. At higher temperatures (773-873 K) metals tend to spread under oxidizing conditions by forming an oxide f'dm while heating in H2 may cause rupture of these films and/or crystallite contraction. As in the case of commercially-relevant supported metals, sintering kinetics of model supported catalysts are generally correlated well by a GPLE of order 2. This result has important mechanistic implications since a number of fundamental processes such as emission of atoms from crystallites, diffusion of adatoms on a support, collision of crystallites, or recombination of metal atoms may involve second order processes. Based on quantitative GPLE treatments of sintering kinetics it is possible to define effects of metal, metal dispersion, metal concentration, and support thermal stability which are very similar to those observed (and discussed above) for commercially-relevant supported metal catalysts. Observations from model catalyst studies reported in the last 10 years are summarized in Table 3. They provide confirmation of many of the conclusions reached in earlier studies regarding the complex mechanistic nature of sintering and redispersion processes [ 1,2] as well as new direct insights into sintering and redispersion mechanisms. Some important fundamental observations and principles that follow from this work include the following: 1. Metal crystallite spreading, wetting of oxide support surfaces, and film formation in 02 at temperatures above about 773 K are widely recognized phenomena which occur in both noble and base metal-support systems. In general these phenomena are a result of formation of volatile or highly-mobile metal oxide species which migrate to and interact strongly with oxide supports. While halides such as chlorine speed migration of metal oxide species to the support, this migration nevertheless occurs in the absence of halides. Simultaneous crystallite and molecular migration are observed in oxygen leading to a bimodal crystallite size distribution with some crystallites larger others smaller than the original crystallites. 2. Particle migration, film breakup to small clusters, and crystallite contraction are generally observed in reducing or inert environments. 3. Combination of 1 and 2 under carefully controlled conditions can lead to metals redispersion. 4. At very high temperatures (500-1000~ in air or O 2, alumina-supported metals combine with the support to form surface and even bulk aluminates. At temperatures above 800~

590 Table 3. Important recent experimental observations regarding sintering of supported metals from model catalyst studies. Catalyst i Atm./cond./meth. i Observation(s) ~':Ref. Co, Fe, Ni/ i H2, H20, CH4; i Sintering and/or coking occur more rapidly in gas i 29 7-A1203 i 500-700~ mixtures (e.g. CH 4 + H20) relative to pure gases. ! i micro-TEM a i Under simulated steam-reforming conditions deacti-i .. vation occurs by carbon deposits (films, patches, and i :filaments), severe sintering, deformation of crystal-i 9

.......................................

i ................................................

.......................................

~ ................................................

FeCo, FeNi, i H2, CO; FeRu/MnCO 3 ! 400-900~ .9M6ssbauer 9 Pt/7-A1203

'~;Xi~i53

~ .........................

~ H2,O2; 9500-900~ i micro-TEM a i i _.

........... iiS~i~~i~i

.............................................

: ...............................................

: ........................

: ......................................

.

- ................

_-

".................... i"'gi'nieri'nj~'"oF'T"n'm'"c3jsi~ii't'es""~;rba'u'ces'"'a"'Bi~8"8~"i

! macro-TEM b i i

! ................

In H 2 alloy formation occurs below a critical particle i 30 i size; above this critical size segregation to the surface i i of an iron-rich phase occurs and may depend upon i metal-support interaction which decreases in the order i Ru, Ni, Co, Fe. i Observations suggest formation of a thin film of Pt on i 31 i the support in 02 atmosphere and contraction to 3D! i particles in H2; direct observation of short distance i i migration/coalescence of crystallites; evidence for i ! transfer of atoms between adjacent p.articles, i

..... ~'~ .....

i distribution consistent with combined atomic and cry- i i stallite migration; inteffacial tension and surface ana-i i lysis data indicated that surface Pt is oxidized and that i

....................................... i ................................................ i..t~s..m.etA~176

act..w.!tD~e..su~.P.~

...................... i ................

Pt/a-A1203

iO2; 500~ AES, i Sintering of 1.8 nm crystallites produces a bimodal i 33 i micro-HRTEM a i distribution of particles having d > 4-10 nm and d < 1 1 .......................................i ................................................[...n..m-.i:.e:...d..!r.ect.evj..d...e...n..c.e...f..~176 ....................i................ Pt film i 02 plasma; SEM; i Pt/O 2 plasma and Pt catalytic etching processes are i 34 200-800~ i similar; formation of metastable Pt oxides by i i i decomposition of HO2 radicals is suggested, i Ni/a-A1203, i H2,O2; 500-900~ Surface Ni aluminate is formed in 02 at 500~ the 1 35 7-A1203 ~ analyticalHRTEMi aluminate covers metal crystallites formed during i .: reduction in H 2 at 500~ Reduced Ni microcrystals i 9 i are formed epitaxially to the (100), (110), (111), and i .......................................

,i ................................................

: ........................................................................................................................................

." . . . . . . . . . . . . . . . .

....................................... i ................................................ i...(..1.~ .~2.s.u.~a.ce s..o f..~:.~hga_._!.mv.~!.e.s..s.~.9..n.~..int.er.act.!..o..n..

.... i .............

Fe/7-M203

i H2,O2; 500-900~ In 02, Fe metal crystallites spread forming patches of I 36 ! micro-TEM a i multilayer films; in H 2 films collect and rupture to~ .......................................i ................................................[.p...r..o...d..u.ffe..sm~l.me.t...al...p......a~t...!.c!..e...s.: ..............................................................i................ Pt/SiO2,M203 i H2, N2, CO; i Crystallite shape and surface free energy depend on ~ 37 9850-1200~ i atm; high temp. annealing produces spherical particles i " SAM, SEM, ~ with(100) and (111) facets. N2 has little effect; in H2i ....................................... i..t~.e.~176176

................... i ................

metal surfaces i various adsorbates~ Under rxn. conds, surface is dynamic; adsorbates and i 38 .9under rxn conds.; i other adatoms rapidly restructure surface; surface i " LEED,XPD,STM i atom diffusion leads to slow faceting; thermal treat-i ". ~ ments cause slow changes in the distribution of l ....................................... ! ................................................ i..c~176176176

Pd/y-ml203

i 02; 300-900~ In situ TEM

" :.

...................... i ................

i Pd is converted to PdO above 325~ which wets the i 39 ~ support and spreads; above 600~ part of the PdO is! i converted to Pd metal and above 800~ to platelet and i ~ needle-like forms of Pd-A1 intermetallic.

591

9

M/AI203, M= i O2,H20; 1000~ Co,Cu,Ni,Fe i high-energy ion . beam analysis 9 . . . . . . . . .

. . . . . .

i Higher rate of spinel formation with ~,-A1203than with i 40 i a-A1203 and larger for Co and Cu oxides relative to i i Fe and Ni oxides; volatile hydroxides of Co, Cu, and i . Ni form at 1000~ in 0.3 atm steam. . . . . .

9analytical TEM, i Rh203; upon aging surface Rh aluminate is formed, i ": XRD, XPS i Rh accelerates transform, of ~/to ~5,e, and c~aluminas, i a Micro refers'to microscopic, "same area," sequential TEM experiments. b Macro refers to macroscopic, large area TEM experiments.

5.

6. 7.

8.

Pd may form a Pd-A1 intermetallic of platelet or needle-like morphology. This intermetallic has considerably lower activity for oxidation of CO and hydrocarbons. Growth of metal crystallites and interactions with the support occur more readily in reaction environments containing more than one gas, e.g. H 2, H20 and C H 4 relative to the separate gases. Formation of corrosive free radicals such as HO 2 may play a role in accelerating removal and transport of metal atoms from crystallites. Under high-temperature reaction conditions, metal surfaces are dynamic systems in which metal atoms or metal oxide species are highly mobile; high rates of diffusion lead to changes in surface structure to minimize overall surface energy. Surface adsorbates and adatoms serve to restructure surfaces, especially at high temperatures. At sufficiently high temperature, equilibrium crystallite shape and surface free energy vary with atmosphere. Metal clusters and crystallites are observed to form epitaxial layers on supports; this epitaxy can contribute to a stronger metal-support interaction. At elevated reaction temperatures, metal crystallites and promoters can accelerate support phase transformations leading to loss of surface area and even collapse of the support.

4 . CONCLUSIONS / THE FUTURE Studies of sintering and redispersion of commercially-relevant and model supported metal catalysts during the past decade have significantly advanced our understanding of these complex phenomena. Direct spectroscopic observations confirm spreading, film formation, and formation of metal-support solid solutions in oxygen-containing atmospheres, while film breakup to small clusters, crystallite contraction, epitaxy with the support surface and particle migration are observed in reducing atmospheres. At high temperature reaction conditions, metal surfaces are dynamic systems, the structures of which are locally rapidly modified by adsorbing, desorbing and reacting molecules; over longer periods of reaction, equilibrium metal crystallite shape is a function of atmosphere. From careful analysis of sintering rate data, effects of catalyst properties and reaction conditions on sintering rate have been quantified. There is a trend in increasing sophistication of spectroscopic tools used to study sintering and redispersion. In the next decade we might expect additional insights into atomic and molecular processes during reaction at the atomic scale using STM, analytical HRTEM, and other such powerful surface science tools. REFERENCES

1. 2. 3.

R.T. Baker, C.H. Bartholomew, and D.B. Dadyburjor (J.A. Horsley, Editor), Stability of Supported Catalysts : Sintering and Redispersion, Catalytic Studies Division, 1991. C.H. Bartholomew, "Model Catalyst Studies of Supported Metal Sintering and Redispersion Kinetics," Catalysis, Specialist Periodical Report, Royal Society of Chemistry, Thomas Graham House, Cambridge, UK, Vol. 10, 1992. C.H. Bartholomew, "Sintering Kinetics of Supported Metals: New Perspectives from a Unifying GPLE Treatment," Appl. Catal., 67 (1994) 1.

.....

592 4. 5. 6. 7. 8. 9. 10. 11. 12. 13. 14. 15. 16. 17. 18. 19. 20. 21. 22. 23. 24. 25. 26. 27. 28. 29. 30. 31. 32. 33. 34. 35. 36. 37. 38. 39. 40. 41.

C.H. Bartholomew," Sintering Kinetics of Supported Metals," in Catalyst Deactivation 1994, eds. B. Delmon and G.F. Froment, Stud. Surf. Sci. and Catal., Elsevier, 88 (1994) 1-18. G.A. Fuentes and F.A. Ruiz-Trevino, National AIChE Meeting, New York, Nov. 15-25, 1987. G.A. Fuentes and E.D. Gamas, in Catalyst Deactivation, 1991, C. H. Bartholomew and J. B. Butt (eds.), Stud. Surf. Sci. and Catal. 68 (1991) 637-644. C.H. Bartholomew and J. B. Butt (eds.), Catalyst Deactivation, 1991, Studies in Surface Science and Catalysis 68 (1991). B. Delmon and G. F. Froment (eds.), Catalyst Deactivation 1994, Studies in Surface Science and Catalysis, Elsevier, 88 (1994). G.B. Hoflund and D.M. Minahan, J. Catal. 162 (1996) 48-53. Y.J. Tu, C. Li, and Y.W. Chen, J. Chem. Tech. and Biotech. 59 (1994) 141-7. Z. Kowalczyk, S. Jodzis and J. Sroda, Appl. Catal. A, Gen. 87 (1992) 1-14. R.L. Keiski and T. Salmi, Appl. Catal. A, Gen. 87 (1992) 185-203. I. Fujiyoshi, J. Catal. 129 (1991) 544-545. M. Agnelli, M. Kolb, and C. Mirodatos, J. Catal. 148 (1994) 9-21. S.D. Mikhailenko, T.A. Khodareva, and E.V. Leongardt, J. Catal. 141 (1993) 688-99. F. Arena, A.L. Chuvilin and A. Parmaliana, J. Phys. Chem. 99 (1995) 990-8. R.W. McCabe, C. Wong, and H.S. Woo, J. Catal. 114 (1988) 354-367. N.L. Wu and J. Phillips, J. Catal. 113 (1988) 129-143. P. Duneau, Effect of Support on Stabilization of Platinum under Oxidizing Conditions, Ph. D. Dissertation, Univ. Pierre and Marie Curie (Paris), July, 1990. Y. Murakami, S. Komai and T. Hattori, in Catalyst Deactivation, 1991, C.H. Bartholomew and J.B. Butt (eds.), Stud. Surf. Sci. and Catal. 68 (1991) 645-642. S. Pikus, J. Catal. 136 (1992) 334-341. P. L66f, B. Stenbom, H. Norden, and B. Kasemo, J. Catal. 144, (1993) 60-76. F. Coloma, A. Sepulveda-Escribano, F. Rodriguez-Reinoso, J. Catal. 154 (1995) 299-305. I. Onal, in Catalyst Deactivation, 1991, C.H. Bartholomew and J. B. Butt (eds), Stud. Surf. Sci. and Catal. 68 (1991) 621-628. J.M. Hess and J. Phillips, J. Catal. 136 (1992) 149-160. D.D. Beck and C.J. Carr, J. Catal. 144 (1993) 296-310. J.M. Schwartz and L.D. Schmidt, J. Catal. 148 (1994) 22-29. K.-I. Aika, Y. Ogata, K. Takeishi, K. Urabe, and T. Onishi, J. Catal. 114 (1988) 200-205. S.H. Lee, and E. Ruckenstein, J. Catal. 107 (1987) 23-81. K. Nagorny and S. Bubert. J. Catal. 108 (1987) 112-134. I. Sushumna and E. Ruckenstein, J. Catal. 108 (1987) 77-96; J. Catal. 109 (1988) 433-462. A. Bellare, D.B. Dadyburjor, and M.J. Kelley, J. Catal. 117 (1989) 78-90. J.M. Rickard, L. Genovese, A. Moata, and S. Nitsche, J. Catal. 121, (1990) 14-152. C. H. Chou and J. Phillips, J. Appl. Phys. 68 (1990) 2415-2423. R. Lamber and G. Schulz-Ekloff, Surf. Sci. 258 (1991) 107-118. E. Ruckenstein, in Catalyst Deactivation 1991, eds. C.H. Bartholomew and J.B. Butt, Stud. Surf. Sci. & Catal. 68 (1991) 585-596. W.H. Lee, V. Petrova, K.R. VanLoon, J.B. Woodhouse, C.M. Loxton, N.L. Finnegasn, and R.I. Masel, in Catalyst Deactivation 1991, eds. C.H. Bartholomew and J.B. Butt, Stud. Surf. Sci. & Catal. 68 (1991) 597-603. G.A. Somorjai, "Surface Reconstruction and Catalysis," Annu. Rev. Phys. Chem. 45 (1994) 721-751. N.M. Rodriguez, S. G. Oh, R.A. Dalla-Betta, and R.T.K. Baker, in Catalyst Deactivation 1994, eds. B. Delmon and G.F. Froment, Stud.Surf. Sci. & Catal. 88 (1994) 417-424. P.H. Bolt, F.H.P.M. Habraken and J.W. Geus, in Catalyst Deactivation 1994, eds. B. Delmon and G.F. Froment, Stud.Surf. Sci. & Catal. 88 (1994) 425-432. A. Weng-Sieh, R. Gronsky, and A.T. Bell, Paper PB 37 presented at 14th North American Catalysis Society Meeting, Snowbird, Utah, June 11-16, 1995.

9

Elsevier Science B. V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

593

The Existence o f an U n u s u a l Reversible Deactivation P h e n o m e n o n Associated w i t h Preferential Surface Segregation in Bimetallic Systems R. T. K. Baker, A. Chambers, C. Park and N. M. Rodriguez Department of Chemistry, Northeastern University, Boston, MA 02115.

When samples of certain bimetallic catalysts were heated in a carbon containing gas mixture an unusual form of deactivation was observed that was completely reversible in nature. It was found that the activity of the bimetallic catalyst was extremely sensitive to the reaction conditions. At temperatures in excess of 725~ the system underwent appreciable deactivation and if hydrogen was added to the reactant the lifetime of the catalyst could be extended up to 800~ Previous studies have favored the argument that under these circumstances the formation of an encapsulating layer of graphite is responsible for blocking the active metal sites on the surface. In order to restore the activity of the spent catalyst it would be necessary to gasify the carbonaceous overlayer in an oxidizing environment at elevated temperature before re-introduction of the initial reactant mixture. In the present systems it was evident that this form of deactivation was not operative since the activity of the "spent catalyst" could readily be restored to its original level by merely lowering the temperature to a previously active regime, indicating the reversible nature of the deactivation process. It is suggested that this phenomenon is related to the preferential wetting action by one of the metal components of the solid carbon filament deposit and this behavior results in a concomitant enrichment of the exposed particle surface in contact with gas phase by the other metal. When this pattern of behavior results in the formation of a surface that does not favor dissociative chemisorption of ethylene or carbon monoxide, deactivation of the catalyst particles will be observed. 1.

INTRODUCTION

Catalyst deactivation is generally attributed processes that promote the formation of carbon deposits on the active metal surface, the loss of metal surface area due to particle sintering or poisoning reactions by sulfur, arsenic, phosphorus or heavy metal impurities in the reactant feed stream (1-6). Regeneration of coked supported metal catalysts is usually accomplished by gasification of the carbonaceous deposit in an oxygen-containing environment at temperatures of about 500~ or higher. This treatment invariably results in the concomitant growth of the metal particles, since such entities frequently participate in the carbon oxidation step and during motion they collide with particles and undergo coalescence (7). As a consequence, while efficient removal of the accumulated carbon is achieved by treatment in oxygen, metal particle sintering is unfortunately facilitated during the process. Restoration to the initial activity level requires that the agglomerated particles be subdivided by a suitable redispersion process that generally involves heating the system in an oxidizing environment containing a small amount of chlorine gas. The decomposition of ethylene over metal catalyst particles to form a mixture of gaseous products and filamentous carbon has been found to provide a very sensitive method

594 of probing the chemical nature of the particle surface in contact with the gas phase (8). Indeed, it was demonstrated that from the analysis of the total product distribution it was possible to obtain an estimate of the composition of the surface of a bimetallic surface as a function of temperature, time and nature of the reactant gas (9). This approach overcomes the ambiguities that are encountered in a post reaction chemisorption procedure since the measurements of the various product concentrations are taken under reaction conditions. In the current investigation we have investigated the catalytic behavior of unsupported iron-nickel during interaction with CO/H2 and C2H4/H2, respectively; coppernickel and copper-cobalt with C2H4/H2, over the temperature range 500 to 725~ A fascinating aspect to emerge from these investigations was the finding that when the bimetallics were heated in the reactant gas streams to temperatures in excess of 700~ there was a dramatic change in the catalytic activity, particularly with regard to the formation of filamentous carbon. In the latter two systems and that were iron-nickel was reacted in CO/H2 as the temperature was increased to 725~ a precipitous drop in catalytic activity was observed. However, if the temperature was subsequently lowered to a previously active regime, the activity could be restored to the same level as that of corresponding fresh sample that had not been treated at 725~ In sharp contrast, when iron-nickel powders were reacted in an C2H4/H2 environment then a complete opposite trend in activity characteristics was observed as a function of temperature. In this case, the conversion of ethylene was found to be relatively low when the reaction was conducted at 600~ whereas when the temperature was raised to 725~ copious amounts of filamentous carbon were generated. Furthermore, this switch in performance of the bimetallic could be achieved by cycling between these two temperatures. This behavior is rationalized according to the notion that preferential segregation of one of the components to the catalyst particle surface occurs at a certain critical temperature and this phenomenon is responsible for the observed unusual reversible deactivation. 2.

EXPERIMENTAL

2.1 Materials The various powdered mixtures of iron-nickel and copper-cobalt used as catalysts in this work were prepared by coprecipitation of the metal carbonates from the respective nitrate solutions mixed in the desired ratios using ammonium carbonate. The detailed procedure was similar to that reported by other workers (10,11). The precipitates were dried overnight at 110~ and then calcined in air at 400~ for 4 hours, followed by reduction in a 10% hydrogen/helium mixture at 500~ for 24 hours. Before removing the reduced samples from the reactor, they were initially cooled to room temperature in helium and then given a passivation treatment in a 2% oxygen/helium mixture for 1.0 hour. Following this step the bimetallic granules were subjected to mechanical grinding to produce powdered samples of about 1.0 gm average particle size. The ratio of the component metals present in each bimetallic powder was determined from XRD patterns of the samples. These measurements were performed in a Scintag diffractometer using nickel filtered CuKct radiation. Diffraction patterns were recorded over a range of 20 angles from 10 to 90 ~ and compared with the known X-ray powder files to establish the phase identities. The gases used in this work, ethylene, carbon monoxide, hydrogen, helium and oxygen were all 99.99% purity and were supplied by MG Industries and used without further purification. Reagent grade iron nitrate [Fe(NO3)2.3.9H20], nickel nitrate [Ni(NO3)2.6H20], cobalt nitrate [Co(NO3)2. H20] and copper nitrate [Cu(NO3)2.3H20] were obtained from Fisher Scientific for the catalyst preparations.

595

2.2 Apparatus and Procedures The apparatus used throughout this study has been described in detail elsewhere (8) Briefly, the system consists of a quartz flow reactor heated by a conventional Lindberg horizontal tube furnace. The gas flow to the reactor is precisely monitored and regulated by the use of MKS mass flow controllers allowing a constant composition of reactant feed to be delivered. Powdered catalyst samples (50 mg) were placed in a ceramic boat at the center of the reactor tube in the furnace. After reduction in a 10% hydrogen/helium mixture for 2.0 hours at the desired reaction temperature the system was flushed with helium for 1.0 hour. The reactant gas, CO/H2 (4:1) or C2H4/H2 (4:1) was then allowed to flow over the catalyst for periods of up to 90 mins. The reaction was followed as a function of time by sampling both the inlet and outlet gas streams at regular intervals and analyzing the reactants and products by gas chromatography using a 30-megabore (GS-Q) capillary column in a Varian 3400 unit. Carbon and hydrogen atom mass balances in conjunction with the relative concentrations of the respective components were employed to obtain the various product yields. At the termination of a given experiment, the gas flow was switched off and the reactor tube cooled to room temperature in helium and the amount of carbon deposited on the catalyst determined by weight difference. This value was always within 5% of that calculated for solid carbon from the mass balances. The structural features of the solid carbon deposit were established from examinations carried out in a JEOL 2000EXII transmission electron microscope. This instrument has a lattice fringe resolution of 0.14 nm. Suitable transmission specimens were prepared by ultrasonic dispersion of a small quantity of the carbonaceous deposit in isobutanol and then application of a drop of the supernate to a holey carbon film. Inspection of many areas of such specimens revealed that in all cases the major type of material generated in these reactions consisted of filamentous carbon structures. 3. R E S U L T S AND DISCUSSION

3.1. Effect of Temperature C2H4/H2 Reactions

on the Iron-Nickel-CO/He

and Iron-Nickel-

When a samples of iron-nickel powder was heated in the presence of a CO/H2 mixture the maximum catalytic activity in terms of the conversion of CO to solid carbon and CO2 was found to occur at temperatures between 500 to 600~ On continued heating to higher temperatures a precipitous drop in the activity of the catalyst towards the formation of these two products and at 725~ almost complete deactivation was observed. This behavior appeared to be most pronounced for the bimetallic powders containing a relatively high fraction of nickel. If, however, the sample temperature was simply lowered from 750 ~ to 600~ the performance of the catalyst was restored to its original high level. The data presented in Figure 1 shows the activity profile with respect to CO2 and solid carbon formation as the reaction temperature of an iron-nickel (3:7) catalyst being reacted in CO/H2 (4:1) is gradually lowered from 725 ~ to 600~ In further set of experiments the temperature was cycled between these two limits several times and the deactivation-activation phenomenon found to be entirely reversible in nature. High resolution transmission electron microscopy examinations were performed on catalyst particles that have treated at temperatures in excess of 725~ in the presence of CO/H2. Most of the bimetallic particles had short filaments associated with them, indicating that under these conditions the catalytic activity towards the growth of this form of carbon was relatively low. It was significant, however, to find that the particle surfaces on which gas phase adsorption and decomposition reactions occurred had remained completely free of any accumulated carbon residues. This observation is consistent with the argument that the

596 decrease in catalytic activity upon raising the temperature from 600 to 725~ is not linked to creation of a carbon overlayer at the metal/gas interface.

30

tO = m

_ , , , ,

I

. . . .

I

. . . .

t

. . . .

i

. . . .

_

o

25

720 9

-_

" 20 !._.

a

740

15

700

-4

680

"o

660

=

640

~

3

c

o

o !_.

.-e

10

62o ~ 5 0

__~._0

,

,

,

,

I

50

,

,

,

,

I

,

,

,

,

I

1 O0 150 Time (rain)

,

,

,

,

I

200

,

,

,

,

250

600

580

Figure 1. Reactivation behavior of an iron-nickel (3:7) catalyst during the decomposition of a CO/H2 mixture as the temperature is lowered from 725 ~ to 600~ O Carbon; ~1CO2. Treatment of the same bimetallic system in an C2H4/H2 (4:1) reaction mixture produced some very unexpected findings. In this case, the catalytic activity for the conversion of ethylene to solid carbon and gaseous products, methane and ethane exhibited a steady increase as the temperature was progressively raised from 600 ~ to 650~ and continued to rise in a smooth fashion on continued heating up to 725~ This dependence in catalyst performance with temperature is presented in Figure 2 for the interaction of iron/nickel (7:3) with C2Hn/H2 (4:1). The change in both solid carbon formation and ethane production are shown as a function of reaction temperature over the range 725 ~ to 600~ Furthermore, this trend was completely reproducible when the temperature was cycled between these two values, indicating that no permanent deactivation phenomenon was operative in this system. It was significant to find that this behavioral pattern was most prominently displayed by bimetallic mixtures in which iron was the major component. It is evident that iron-nickel exhibits diverse catalytic behavior as function of reaction temperature when the bimetallic is treated in CO/H2 and C2H4/H2 mixtures. It is also evident that the behavioral pattern in both these environments is extremely sensitive to the composition of the bimetallic powders. In the case of the former reaction it is clear that the observed deactivation that occurs at 725~ cannot be attributed to the formation of a carbonaceous overlayer since the system could not regain its catalytic activity by merely decreasing the temperature since the active sites would remain blocked. A more cogent explanation may preferential segregation of nickel occurs surface in this component, conditions that nickel is not a facile process (12,13).

lie in the possibility that at high temperatures and this results in an enrichment of the particle do not favor dissociative chemisorption of CO on Furthermore, the reactions leading to carbon

597

deposition on nickel are dependent to a large degree on the crystallographic orientation of the face on which the adsorption of the gaseous precursor takes place (14-16). There is evidence from surface science studies indicating that treatment of iron-nickel alloys at high temperatures results in the preferential segregation of nickel to the surface (17). Theoretical studies reported by Chelikowsky (18), who examined the segregation behavior of a number of bimetallic systems under reducing conditions, are also in accord with this finding.

40

IL--rQ, F \

,

,

I~

,

,

,

!

. . . .

I

. . . .

I

. . . .

740

_ _

35

720

': 30

700

e"

0

,,,

t

25

i5 20

680 660

~ 5

640

"

620 o

10

5 0

--~ 0

50

100 150 Time (rain)

200

250

60O

580

Figure 2. Variation in the production of solid carbon and ethane from the interaction of iron/nickel (7:3) with C2H4/H2 (4:1) as a function of reaction temperature. O Carbon; [~C2H6. A further factor that could have an impact on this segregation behavior are the events that occur at the metal/solid carbon interface. Dynamic studies carried out in the controlled atmosphere electron microscope have shown that there is a strong tendency for iron-nickel particles to wet and spread along the edges of graphite when such systems were heated in H2 at temperatures in excess of 600~ (19). This behavior is consistent with the formation of a strong interaction between the bimetallic and the graphite platelets constituting the filament structure. Under these circumstances it is not unreasonable to expect that any excess nickel atoms would tend to accumulate at the exposed faces of the catalyst particles. In contrast to the behavior of CO, the decomposition of ethylene is a facile process when performed on a nickel catalyst, but does not occur when the hydrocarbon is passed over iron. Based on these data we can rationalize the observed deactivation behavior observed in the present investigations according to the notion that at 725~ the surface of the bimetallic particles become enriched in nickel, a condition that favors decomposition of adsorbed ethylene molecules, but is inert with regard to catalyzed disproportionation of CO. Subsequent lowering of the temperature to 600~ results in the restoration of the original surface composition and the concomitant attainment of the initial catalytic reactivity pattern.

598

3.2. The Effect of Temperature on the Copper-Cobalt-C2H4/H2 Reaction In this set of experiments copper -cobalt (1:3) samples were exposed to a n C 2 H 4 / H 2 (4:1) mixture at temperatures over the range 400 to 800~ The variation in the percent of solid carbon formed after a period of 60 mins at each temperature is presented in Figure 3. Inspection of these data shows that at about 750~ the amount of this product that is formed is almost negligible. In some experiments, after reaction at 750~ the temperature was lowered to 600~ whilst maintaining the flow of reactant gas. From the plot it is evident that after a further 60 mins reaction at this temperature, the yield of solid carbon is very close to that obtained from a fresh catalyst sample treated at a maximum of 600~

70 o"

60

-_

O _

,, 5 0

-

o c 0

40--

= , ~

_

r

'e0

o

-

:30

-

20

-

-

_

~

-

10

-

0

!

300

!

400

500 600 Temperature (~

700

800

Figure 3. Variation of the amount of solid carbon formed from the copper-cobalt (1:3) catalyzed decomposition of C2H4/H2 (4:1) as a function of reaction temperature. The solid circle denotes the product yield from a specimen that was initially treated in the reactant gas at 725~ and then subsequently cooled to 600~ and reacted for a further 60 mins. Examination of reacted samples by transmission electron microscopy revealed that filamentous carbon appeared to be the exclusive form of deposit generated in al the systems. Furthermore, it was interesting to find that the characteristics of the filament structures produced on bimetallic powders that had initially been treated under conditions where the temperature had been reduced from 750 to 600~ were identical to those generated from a sample that was reacted directly at 600~ These findings indicate the high temperature treatment does not appear to produce any permanent changes in the crystallographic properties of the bimetallic particles, since such a transformation would be expected to be reflected in structural modifications in the characteristics of the carbon filaments (20). It is possible that the reversible deactivation phenomenon observed in this system is similar in many respects to that found when copper-nickel particles were reacted in the same gaseous environment (21). In the latter case additional evidence was obtained from a

599 combination of controlled atmosphere electron microscopy and in-situ electron diffraction techniques that demonstrated the tendency of the nickel component to undergo a wetting and spreading action with graphite at temperatures of about 750~ thereby leaving copper in the form of discrete particles. While this type of behavior was so pronounced in the coppercobalt system, it is likely that the cobalt component will form a stronger interaction with graphite and as a consequence, tend to segregate towards the carbon precipitating faces of the catalyst particle, thus creating an enrichment of copper at the metal/gas interface. 4.

CONCLUSIONS

The results of this investigation have highlighted the existence of an unusual deactivation phenomenon when certain bimetallic catalyst systems are reacted in either a hydrocarbon or carbon monoxide environment. The activation/deactivation behavior of the respective bimetallic catalysts appears to be completely reversible in nature and is found to be highly sensitive to both reaction temperature and of the composition of the reactant gas. Deactivation is attributed to the preferential segregation of the respective metal components to either the solid carbon or gas phase interfaces of the catalyst particles. When the temperature is subsequently increased or lowered to a previously active regime then catalytic activity is restored to a level comparable to that observed with a fresh sample.

Acknowledgments Financial support for this work was provided by the Department of Energy, Basic Energy Sciences Grant No. DE-FG02-93ER14358.

References 1. 2. 3. 4. 5. 6. 7. 8. 9. 10. 11. 12. 13. 14. 15. 16. 17. 18. 19. 20. 21.

C.H. Bartholomew Chem. Eng. 91, 96 (1984). "Deactivation and Poisoning of Catalysts" (Oudar,J. and Wise, H. eds.) Dekker, New York, 1985. J. R. Rostrup-Nielsen in "Catalysis, Science and Technology" Vol.5. Chap. 1. (Anderson, J. R. and Boudart, M. eds.) Springer-Verlag, Berlin,1983. D . L . Trimm in "Progress in Catalyst deactivation" (Figueiredo, J. L., ed.) NATO Series A Applied Sciences No. 54, p.3 Martinus Nijholf, London, 1982. E. Ruckenstein, in ,Metal-Support Interactions in catalysis, Sintering and Redispersion" (Stevenson, S. A. et al. eds.), Van Nostrand Reinhold, New york, p.140 (1987). R . T . K . Baker, and P. S. Harris in "Chemistry and Physics of Carbon", Vol.14, p.83 (Walker, P. L. Jr., and Thrower, P. A. eds.) Dekker, New York, 1978. R . T . K . Baker,, J. Adhesion 52, 13 (1995). M.S. Kim N. M. Rodriguez, and R. T. K. Baker J. Catal. 131, 60 (1991). A. Chambers N. M. Rodriguez, and R. T. K. Baker J. Phys. Chem. 99, 10581 (1995). R. J. Best, and W. W. Russell J. Am. Chem. Soc. 76, 838 (1954). J. H. Sinfelt J. L. Carter, and D. J. C. Yates J. Catal. 24, 283 (1972). N. M. Rodriguez M. S. Kim, and R. T. K. Baker J. Catal. 144, 93 (1993). A. Sacco Jr., F. W. A. H. Geurts G. A. Jablonski Lee, S. and Gateley, R. A., J. Catal. 119, 322 (1989). H. Leidheiser Jr., and A. T. Gwathmey J. Amer. Chem. Soc. 70, 1206 (1948). V. J. Kehrer, and H.Leidheiser Jr., J. Phys. Chem. 58, 550 (1954). G. D. Renshaw, C. Roscoe, and P. L. Walker, Jr., J. Catal. 22, 394 (1971). K. Wandelt, and G. Ertl J. Phys. F., Metal Phys. 6, 1607 (1976). J. R. Chelikowsky Surf. Sci. 139, 1197 (1984). R. T. K. Baker J. J. Chludzinski, and R. D. Sherwood Carbon 19, 99 (1981). M. S. Kim N. M. Rodriguez, and R. T. K. Baker J. Catal. 134, 253 (1992). N. M. Rodriguez M. S. Kim, and R. T. K. Baker J. Catal. 140, 16 (1993).

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Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

601

V a p o r i z a t i o n - A s s i s t e d D e g r a d a t i o n of High T e m p e r a t u r e C o m b u s t i o n Catalysts J. G. McCarty, K.-H. Lau, and D. L. Hildenbrand SRI International, 333 Ravenswood Avenue, Menlo Park, CA, 94025, USA Vaporization and transport can lead to serious thermal deactivation of catalysts during hightemperature combustion, partial oxidation, and steam reforming of light hydrocarbon fuels. These environments, with high concentrations of steam, and in some cases excess oxygen, at high temperatures and high pressures, represent perhaps the most severe conditions that catalysts must endure. While the evaporation of atomic metal and polyatomic metal oxide gaseous species has been generally understood, the volatilization of hydroxides and oxyhydroxides under these conditions has not been fully appreciated. This paper reviews vaporization thermochemistry and the influence of vapor transport on thermal degradation processes of supported metal and metaloxide catalysts. 1. I N T R O D U C T I O N Catalytic combustion is an environmentally-driven, materials-limited technology with the potential to lower nitrogen oxide emissions from natural gas fired turbines consistently to levels well below 10 ppm. Catalytic combustion also has the potential to lower flammability at the lean limit and achieve stable combustion under conditions where lean premixed homogeneous combustion is not possible. Materials limitations [1,2] have impeded the development of commercially successful combustion catalysts, because no catalytic materials can tolerate for long the nearly adiabatic temperatures needed for gas turbine engines and most industrial heating applications. Vaporization can contribute to the sintering of supported metal and supported metal-oxide catalysts through the Ostwald ripening process. The high concentrations of vapors produced at sustained high temperatures can also lead to the transport of significant amounts of catalytic materials to substrates, where they can react, or into the gas phase, where they are lost in the effluent gas stream. In addition, vaporization can cause the migration of corrosive components from substrates, supports, or deposited contaminants, and lead to an accumulation of deactivating components on the surfaces of catalytically active phases. Therefore, it is essential to examine the thermodynamics of vaporization in order to assess which materials are potentially unstable over very long periods (> 8000 h) at high temperatures (> 1000 ~ in combustion environments. 2. T E C H N I C A L A P P R O A C H

2.1 Estimation of Thermochemical Properties We have estimated the thermodynamic properties of a wide variety of gaseous metal hydroxides and oxyhydroxides by developing a correlation between the bond strengths of metal hydroxides and metal halides [3]. In particular, the bond strengths of metal hydroxides were found to be comparable based on correlations of known measurements to that of the bond strengths of the

602 associated metal chlorides. The bond strengths of metal hydroxides was also proportional to the bond strengths of metal fluorides with an addition of some 130 kJ/mol. After estimating the critical enthalpies for formation for a particular gaseous metal hydroxide and oxyhydroxide using these correlations for metal halide and oxyhalide data, and also by extrapolating a series of related compounds (same Periodic Table column), we were able to estimate the full thermodynamic properties of all relevant gaseous metal hydroxides and oxyhydroxides. Thermal functions included enthalpy of formation and heat capacity were calculated using estimated molecular constants, which include moments of inertial for rotational contributions, fundamental modes of vibration, symmetry, and electronic contributions, if any. After compiling these fundamental thermochemistry parameters, we were able to predict the volatility of relevant metals and metal oxides.

2.2 Knudsen Cell Measurements of Enthalpies of Formation In addition to the estimated properties, we measured the thermochemistry of several important vapor species. These measurements were conducted in a Knudsen effusion cell using special lineof-sight vaporization under subambient pressures with flowing O2 and H20 vapor mixtures [4]. The gaseous species over silica [5], manganese oxide [6], lanthana, alumina, and palladium metal were detected and relative partial pressures measured as a function of temperature. These vapor pressure measurements were calibrated by using the known metal atom or binary metal oxide volatility as a calibration source. Oxide species concentrations were measured relative to that of a reference compound, e.g., metal atom. The identification of oxide and hydroxide compounds was facilitated by the technique of threshold electron ionization [7]. These data were then evaluated using estimated entropy functions and the third law temperatures. 3. R E S U L T S Vaporization partial pressures have been predicted by combining the estimated and experimentally-measured thermochemistry. One of the findings is that silica is volatile in the presence of high concentrations of steam, e.g., parts-per-billion concentrations of volatile Sicomponents are reached at temperatures as low as 850~ (Figure 1). Unacceptably high levels of SiO(OH)2 (100 ppb) occur at temperatures as low as ll00~ A series of manganese oxide and oxyhydroxide compounds were also observed in the Knudsen cell experiments with Mn203 powders. Mn(OH)3 (manganese trihydroxide), which was estimated from halide data, and Mn(OH)2 (manganese dihydroxide), are predicted to be the most stable volatile compounds over Mn203 under combustion conditions. Unacceptable levels of vaporization occur for Mn203 at temperatures as low as 750~ (Figure 2). At temperatures approaching 1000~ these levels could reach as high as 100 ppb in 10-atm combustion environments. This degree of volatility of hydroxides is typical for the catalytically active transition metal oxides. Estimated vapor pressures for a series of metal containing species versus temperature show that of the noble metals, Pd and its oxides are superior with respect to vaporization (Figure 3). It is well known that platinum dioxide is volatile (Figure 4). Our estimated thermochemistry allowed us to also predict that Pt(OH)2 is as volatile as PtO2 although this was not verified by direct measurements. For temperatures below 1000~ Pd is followed by Rh, Ir, and subsequently, Pt, in order of increasing volatility (Figure 5). The first row transition metal oxides are also surprisingly volatile (Figure 6). Iron (primarily relying on estimates based on the relatively limited data for halides), appears to possess the most stable oxide with acceptable volatility to temperatures as high as 1000~ Ruthenium and copper oxides are extremely volatile and cannot be recommended as active components in combustion catalysts. The least volatile transition metal oxide is iron (based primarily on estimated enthalpies of formation). A trend is observed for the more active transition metal oxides of increasing volatil-

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606 ity with increasing specific activity. Cobalt may be too volatile to be useful at 1000~ without stabilization in a complex oxide such as LaCoO3. Copper appears to be unsuitable for combustion except at very low temperatures, as its volatility rises well above 100 ppm at 1000~ Materials which have extremely low volatility and the potential to serve therefore as nonreactive substrates and supports include zirconia, titania, and alumina. The rare earth oxides, represented by lanthana and ceria, also show very low volatility and can be expected to resist vapor-transport-assisted sintering and corrosion. Of the alkaline earth oxides, MgO is superior to all others, although the alkaline earth aluminate complex oxides can be much more stable (e. g., by a factor of 1000 for BaO) than the alkaline earth oxides. While the incorporation of transition metal oxides into complexes with materials such as alumina can lower their volatilities by factors from 10 (CuO) to 1000 (BaO) depending primarily upon the heat of reaction between the two oxides, it is also likely that formation of very stable complex metal oxides, such as aluminates, can also greatly lower the chemical activity of the transition metal. As mentioned above, Mn, Ni, and Co may require stabilization in complex oxides for long catalyst life, but the complex oxides generally have inferior activity. The most active transition metal oxides (Ru and Cu) may still have unacceptable volatility as relatively active complex oxides. As a consequence, there may be a technology-limiting trade-off between the catalytic activity of metals and metal oxides and their chemical and thermal stability in combustion environments. Trace contaminants such as air-born chlorides or catalyst impurities, can adversely affect the vaporization of metal oxides. Preliminary calculations of metal oxide volatility in the presence of oxygen water vapor and trace hydrogen chloride show that transition metal oxides (e.g., the oxides of Cr and Fe) form relatively stable oxychorides.

4. D I S C U S S I O N

4.1 Vaporization-assisted transport Levels of volatility that would lead to unacceptable rates of vapor transport-driven sintering, attrition of catalytically-active materials, or corrosion of catalytic materials or support oxides by transport from contaminants or substrate materials can be estimated given equilibrium vapor pressures and a few assumptions about evaporation rates and mass transport. In particular, the rate of condensation of a vapor species on its source solid phase at high temperatures is almost certainly non-activated and may show little configurational restriction. Therefore, using the principle of microscopic reversibility, we can take the rate constant for condensation to be approximately equal to the collision frequency. Presuming the catalyst to exist as a thin wash-coated layer of an oxide that supports the active phase, one can estimate the rates of vaporization-assisted sintering within the micropores of the catalyst particles of the bimodal porous washcoat with estimates of the effect of crystallite size on vapor pressure and mean separation between particles (the Ostwald ripening process). Similarly, using the appropriate effectiveness factor one can estimate the rates of vapor transport of an active component dispersed within the microporous particles into the macropores of the washcoat. Finally, one can estimate the flux of active component lost to the gas phase by transport from the macropores through the convection boundary layer near the catalyst wall. These analyses all indicate that roughly 1-ppb (10 -9) total metal vapor concentration approaches (within an order of magnitude or two) the threshold where long-term (8,000 hours) thermal deactivation can be expected.

607 4.2 Selection of Catalytic Combustion Materials Compilation of thermochemical properties allows us to predict equilibrium vapor pressures and vapor transport for catalytically-relevant metal catalyst component materials for given atmospheric conditions at temperatures from 727 ~ to 1727~ The vaporization thermochemistry indicates that of the supporting oxides, only alumina, zirconia, titania, magnesia, and the rare earth oxides, are recommended as support or substrate materials in combustion catalysts. Of the catalytically active materials, palladium is superior to platinum and other noble metals, and iron oxide is superior to other Group VIII transition metal oxides. The use of silica, higher molecular weight alkaline earth oxides, e.g., barium, and any alkali metal in combustion catalysts cannot be recommended because of long-term vaporization-assisted sintering and potential corrosion of active components. 4.3

Conclusions The data determined directly by Knudsen cell measurements, plus a strong correlation between the bond strengths of metal hydroxide bonds and metal halide (in particular, chloride and fluoride bonds) in the gaseous metal hydroxides and halides were developed and allow us to more reliably estimate the enthalpy of formation of many hydroxide and oxyhydroxide metal compounds whose values of thermochemical heat and formation were previously unknown. These thermochemical properties were then used to estimate volatility of various supporting oxide substrates and catalytically-active solids that were relevant for the fabrication of catalytic combustors. High volatility limits the selection of many otherwise useful materials. Use of silica, all alkali, and alkaline earth metals higher than magnesium as components in any substrate or supporting material cannot be recommended. Catalysts with the least amount of volatility include palladium and iron oxides. It may also be possible to form complex oxides with other transition metal oxides, such as Mn, Ni or Co, which could have an acceptable activity and stability. For example, the formation of transition metal aluminates could lower volatility and increase sintering resistance, but at the expense of decreased activity.

5. ACKNOWLEDGMENTS The authors gratefully acknowledge the support of the Gas and Research Institute, Osaka Gas Company, the Southern California Gas Company, Tokyo Gas Company, Toho Gas Company, and STATOIL, for a large, multi-disciplinary project conducted from 1991 through 1996 at SRI International. REFERENCES

1. H. Arai and H. Machida, Catal. Today, 10 (1991) 81. 2. M. F. M. Zwinkels, S. G. J/ir~s, P. G. Mennon and T. A. Griffin, Catal. Rev.-Sci. Eng., 35(3) (1993) 319. 3. D. L. Hildenbrand, K. H. Lau and J. G. McCarty, Thermodynamic Properties and Vaporization of Metal Oxides and Hydroxides at Elevated Temperatures, Gas Research Institute Report GRI-92/0521 (1992, available through GRI and NTIS). 4. D.L. Hildenbrand and K. H. Lau, J. Phys. Chem. 96 (1992)2325. 5. D.L. Hildenbrand and K. H. Lau, J. Chem. Phys., 101 (1994) 6076. 6. D.L. Hildenbrand and K. H. Lau, J. Chem. Phys., 100 (1994) 8377. 7. K.H. Lau and D. L. Hildenbrand, J. Phys. Chem. 86 (1987) 2949.

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9 Elsevier Science B.V. All rights reserved Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

609

Deactivation by Sintering and Coking of Sol-Gel NiO-A1203-TiO2Hydrogenation Catalysts J.C. Rodriguez I , T. Viveros 2 and A. Monz6nt. 1 Department of Chemical and Environmental Engineering. Faculty of Science. University of Zaragoza. 50009 Zaragoza. SPAIN. e-mail: [email protected]. 2 CBI-IPH. Area de Ingenierfa Quimica. UAM-Iztapalapa. 09340 M6xico D.F. Mt~XICO.

The influence of the reduction temperature, preparation method and type of support on the textural properties and on the activity, selectivity and coking formation on Ni-TiO2-A1203 catalysts, using acetylene hydrogenation as the test reaction, is investigated. We obtained that catalysts with nickel incorporated into the support, although exhibiting a higher initial values of specific surface area, also exhibited poorer stability during reduction. Also, the thermal stability of TiO2 supports is increased by the addition of A1203. On the other hand, the catalytic performance and the effect of coke deposition on the activity and selectivity, were found to be strongly dependent on the preparation procedure, on the reduction temperature, and on the type of support. Catalysts containing TiO2 provided higher ethene selectivities than Ni-A1 catalysts. 1. INTRODUCTION One of the main aspects that determine the properties of a given catalyst is the nature of the interaction between the oxide support and the dispersed active metal. The influence of this interaction on the activity, selectivity and stability of the catalyst is determined by factors such as the preparation method, the atmosphere and temperature of the calcination and reduction stages (1,2), and the specific metal-support system studied. The latter is especially important for catalysts based on transition metals supported on partially reducible oxides (3). Such catalysts displaying strong metal-support interaction (SMSI) exhibit suppression of chemisorption of H2 and CO (4-6) when reduced at temperatures up to 773 K. One of the most frequently used models to explain the SMSI phenomena is the "decoration model" (2,7,8) according to which a progressive reduction of the surface of the oxide support leads to the coverage of the metal particles, involving complex interaction at the interface such as the migration of supported metal into the lattice of the oxide support. However, the nature of this phenomenon is not entirely understood (9,10). Studies of TiO2 supported catalysts have shown that the anatase-rutile transformation of TiO2 also occurs around the temperature where the SMSI effect manifests itself (11-13). Furthermore, during phase transitions, sintering of the metal and support can occur. These facts tend to occur simultaneously, in such a way that the prime cause of the SMSI effect remains unclear, or possibly does not exist, the SMSI effect being a sum of all the individual facts mentioned.

610 On the other hand, the influence of SMSI upon the activity and selectivity are well known to be dependent on the nature of the catalyst, as well as the type of reaction. A decrease in the activity at reduction temperatures of 773 K is a commonly accepted fact. Nevertheless, this activity reduction seems to be more drastic for demanding reactions such as hydrogenolysis, and weaker for other non-sensitive reactions such as hydrogenations or dehydrogenations (8). In spite of the above, there are some aspects that are not clear at present such as the influence of the preparation method on the SMSI state, and the effect of SMSI phenomena on secondary reactions such as coking. In a previous work we examined partial aspects of the catalytic performance of NiO-AI203-TiO2 mixed oxides, prepared by the sol-gel route, in the hydrogenation of acetylene and phenylacetylene (13). In this work, we investigate the influence of the reduction temperature, preparation method and type of support on the textural properties and on the activity, selectivity and coking formation on these materials, using acetylene hydrogenation as the test reaction. 2. E X P E R I M E N T A L

2.1 Catalyst preparation In order to investigate the influence of the preparation method and of the type of support, four different catalysts were prepared, Ni-TiO2 sg (Ni-Ti), Ni-TiOz-AlzO3 sg ((Ni-A1-Ti)sg), Ni-AI203 sg (Ni-A1) and Ni-TiOz-AI203 imp ((Ni-A1-Ti)imp). In the first three catalysts, nickel was incorporated during the manufacture of the support, prepared by a sol-gel procedure. (Ni-Al-Ti)imp was prepared by wet impregnation with an alcoholic Ni(NO3)2 solution of a pre-manufactured TiO2-AlzO3 support, also obtained by the sol-gel technique (13). All the catalysts were prepared with a nominal nickel content of 10% wt, and the mixed supports contained equal proportions of alumina and titania. Calcination of the dried gels was performed in vacuum for one hour and then in air flow for two hours. 2.2 Characterisation Temperature Programmed Reduction (TPR) experiments were carried out using a total flowrate of 20 cm3/min containing 5% H2 in N2, under a ramp of 10 K/min. XRD spectra were obtained in a Rigaku/Max system powder difractometer, using the radiation K~ of copper (~,= 1.5418 A), and at a scanning rate of 0.03 ~ BET areas were determined on a Micromeritics 2100E apparatus using N2 adsorption at 77 K. 2.3 Reaction tests The reaction experiments were performed in a thermogravimetric unit (C.I. Electronics MK2), fitted with automatic controllers of flow (Brooks mass flow controllers), and temperature. Weight data were collected at regular time intervals with an accuracy of 10-2 mg, using a computer-controller data logging system. The balance was operated as a differential reactor, maintaining the conversion below 10%. During the reaction, coke was formed as a byproduct. The catalyst sample (66 mg) was introduced into an stainless steel gauze basket and hung from one of the thermobalance side-arms. In previous works the contribution of this basket to the catalyst activity was found to be negligible. The range of catalyst particle diameter used (160-673 ~tm) and total flow rate (750 cm3/min) ensured the absence of internal or external diffusional limitations. The catalyst samples were reduced under a Nz/H2 stream (total flow rate 100 ml/min, 50% H2) for 3 hours, at temperatures ranging from 573 K to 773 K. After reduction, the sample was cooled in N2 down to the reaction temperature (448 K).

611 The system was operated at atmospheric pressure, and the gas mixture had the following composition H2/C2H2/N2 = 60/15/25. The exit gases were analysed by an on-line gas chromatograph equipped with a FID detector. 3. RESULTS AND DISCUSSION 3.1 Characterization The XRD patterns presented in Figure 1 show a low degree of crystallinity in the Ni-A1, (Ni-A1-Ti)sg and (Ni-A1-Ti)imp samples, and only diffraction bands corresponding to 7-A1203 can be detected. In the case of the Ni-A1 catalyst, a slight displacement of the diffraction lines to greater angles is obtained. This could indicate the presence of NiA12Oa-like phases (Figure 1). The lack of crystallinity implies a well dispersed A1-Ti matrix, in cases of the mixed supports. In addition, in the Ni-Ti, (Ni-A1-Ti)sg and Ni-AI samples, the absence of NiO diffraction lines indicates a small nickel oxide particle size (commonly less than 5 nm), as a result of the catalyst synthesis procedure followed (13).

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On the other hand, in the diffraction pattern of the (Ni-A1-Ti)imp catalyst it was possible to identify the presence of NiO bulk. Finally, the XRD pattern of Ni-Ti revealed a higher degree of crystallinity of this sample with respect to the others. Brookite and Anatase (main specie) species were identified in this case. The BET specific surfaces were measured after calcination, and after reduction at 573 and/or 773 K, for all the samples studied (Table 1). After calcination, the (Ni-A1-Ti)sg and Ni-A1 catalysts exhibited the greater values, in contrast to the Ni-Ti sample. Nevertheless, the result obtained for Ni-Ti is similar to the reported values for commercial TiO2 supports (14). From the comparison of the values obtained with the (Ni-A1-Ti)imp and (Ni-AITi)sg catalysts, it is clear that the impregnation of nickel on the support caused a pore filling, and subsequently a

decrease in the area of the catalyst (14). The reduction process causes a loss of the surface area in the four catalysts studied. This fact became very marked in the case of Ni-Ti, after reduction at 773 K (Figure 2). On the othe hand, (Ni-A1-Ti)imp revealed itself as a very stable catalyst during reduction. Intermediate results were found in Ni-AI and (Ni-A1-Ti)sg. The results presented above allow us to conclude that the presence of NiO inside the support matrix leads to its sintering during the reduction stage, due to the shortening of Ni-A1 or

612 Ni-Ti distances (15). The sintering of the Ni-Ti catalyst is also confirmed for by XPS measurements (16). Finally, from the comparison of the percentages of area reduction in the Ni-A1, (Ni-A1-Ti)sg and Ni-Ti catalysts, it can be seen that the presence of A1203 plays a beneficial role in decreasing the rate of the support sintering. The lowest reduction temperature of NiO, and thus the highest reducibility, was obtained in Ni-Ti. The Ni-A1 catalyst was exactly the opposite. Once again, an intermediate behaviour was found in mixed supports. Regarding the reduction profiles: i) the TPR of Ni-Ti shows a narrow sharp peak at around 623 K, caused by the reduction of NiO, and a small shoulder at 773 K that is caused by the reduction of NiTiO3 intermetallic species, ii) the TPR of Ni-AI shows two peaks, the first (573-653 K) being due to the reduction of NiO, as in the former Table 1. Influence of temperature of reduction on the case, and the second (873 K) SBET. probably caused by the reduction S.~T (m2/a) of NiA1204 aluminates (see XRD Catalyst Calcined TR=573 K TR=773 K results), or at least, very well disNi-AI 295_+3.1 215_+0.6 188_+1.9 persed NiO particles on the Ni-Ti 132 +_2.0 124 _+ 1.6 35.5 -_4-_1.2 support matrix (13). iii) in the Ni-A1-Ti)imp 215+_2.5 204_+3.0 197_+1.8 case of mixed support based Ni-A1-Ti)sg 320 _+3.8 205 _+2.6 209 _+3.2 catalysts, the main H2 consumption peak appears at 773 K, and in the (Ni-AI-Ti)imp catalyst a shoulder at 613 K can also be detected. After reduction at 573 or 773 K, the TPR profiles changed substantially, and a comparison with the TPR's of the fresh samples allows us to estimate the quantity and "quality" of the metallic phase attained after each reduction treatment.

I

Thus, i) with the Ni-Ti samples, 3h of reduction at 573 K is enough to achieve about 100% of Ni metallic, ii) with Ni-AI, the 9 "" "0. . . . . . : :. same reduction treatment only w~. "~ . . . . . . . . . . :.. 9 allows us to reach 28% of the at~ ". tainable Ni ~ (this figure can be :~ - ~ "~ e . . Ni-AI easily calculated as 1 minus the A.. (Ni-,~m) "o quotient of the area below the TPR curve of the pre-reduced 10 ~ ~ . Ni-~ [ sample and of the fresh one). This , I i ~ ~ l , I percentage correspons to the reReduction Temperature (K) duction of nickel with weak interaction with the support, iii) Figure 2. Sintering of the catalysts after reducwith (Ni-A1-Ti)imp and (Ni-A1tion treatments. Ti)sg catalysts, reduction at 573 K also produces a small percentage of nickel reduction (21% and 17%, respectively), and iv) reduction at 773 K leads to an almost complete reduction of the Ni 2§ into Ni ~ (75%, 100%, 81% and 98% are the reduction percentages achieved in Ni-A1, Ni-Ti, (Ni-A1-Ti)imp and (Ni-A1Ti)sg, respectively). Only the more dispersed Ni § species firmly attached to the support remaining as non-reduced (e.g. NiA1204 in the case of Ni-A1). .

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613 Regarding the amount of accessible metal significant differences were found (the amount of metal that is possible to reduce at temperatures below 1173 K), in spite of the equal nominal metallic content (around 10% wt.). Thus, the higher values are obtained in the (NiA1-Ti)imp sample. With respect to this catalyst, Ni-A1, Ni-Ti and (Ni-A1-Ti)sg samples exhibited a relative H2 consumption of 80%, 43% and 42%, respectively. As expected, the amount of accessible metal is greater in the impregnated catalysts than in those in which nickel was incorporated during the support manufacture stage. In the case latter Ni 2+ is likely to be occluded inside the support lattice. It is important to note that this fact is more significant in those catalysts containing TiO2 as part of the support, probably due to the formation of NiTiO3 species (2,3) with a very high degree of dispersion, and therefore nondetectable by XRD. 3.2 Reaction tests

Figures 4 and 5 show the activity and selectivity results obtained with the different catalysts at the two reduction temperatures studied. 3) T-"red.=773 K Two opposing factors control the I=1 I I-' .' .' ' ' ' ' ' '-I activity levels. Reduction percentages increase as the reduction temperature is increased (Figure 3). On the other hand, metal and support sintering also becomemore important at the higher reduction temperatures, leading to a 6 ~ decrease in the available metal atom.s (Figure 4). At a reduction temperature of 573K, the Ni-Ti and (Ni.-A1-Ti)imp catalysts gave the highest conversion levels, due to their higher metallic surface content. The low activity of Ni-AI 500 625 750 875 1000 at this reduction temperature is Temperature (K) obviously due to the small percentage of metallic reduction achieved under these conditions. If the reduction temFigure 3. TPR results of calcined and reperature was increased to 773 K, all duced catalysts. the catalysts, with the only exception of Ni-Ti, increase their activity as a consequence of the greater metallic content attained. Although this is also true with Ni-Ti, the intense support sintering that takes place at this reduction temperature (Figure 2) strongly affects its activity. Regarding the stability of the different samples over time, for Ti or A1-Ti based catalysts with nickel incorporated during the manufacture of the support (Ni-Ti and (Ni-A1-Ti)sg samples), deactivation by coke deposition decreases the activity levels (Figure 4) and is accompanied by steadily increasing methane yields (results not shown). On the other hand, the (Ni-A1-Ti)imp catalyst exhibits very good stability and resistance to coking. The high coking resistance of Ni-A1203 catalysts in reactions such as selective hydrogenations as previously repol~ed (17,18) and suggested to be the result of whisker-like coke formation. Therefore, the L"

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614 deactivation pattern exhibited by the Ni-Ti and (Ni-A1-Ti)sg samples was unexpected. Obviously, this behaviour could have two causes: firstly, the presence of TiO2 (not detected in Ni-A1) and secondly, the method of preparation (not detected in (Ni-Al-Ti)imp). Two different patterns of selectivity were obtained. After reduction at 573 K, all the catalysts except (Ni-Al-Ti)imp, were shown to be initially very selective to C2H4, but coke deposition caused a progressive drop in selectivity until reaching a constant level (Figure 5). On the other hand, the (Ni-AITi)imp catalyst showed RO exactly the opposite behaviour. After reduction at 773 K, &O with the exception of Ni-Ti, all ZO o 4 the catalysts manifested low v (Ni-AI-Ti)I - 6.0 6.0 initial selectivity levels that 5.0 steadily increased over time. Three factors can be proposed 4.0 , as being responsible for the 3.0 3.0 evolution of the selectivity: the 2.0 size of the metallic particles, 2.0 the interaction between metal 1.0 1.0 and support, and the 0.0 0.0 interaction of reduced nickel O. 1 0.2 0.0 O. 1 0.2 0.0 Coke Content (mg/mg cat.) with the unreduced fraction (Ni 2§ (19). Clearly the first Figure 4. Influence of the reduction temperature upon two factors are related. Given the activity and coking resistance of the catalysts. that the reduction percentage attained in Ni-Ti with the 573 K treatment is about 100%, it does not seem probable that the decrease in the C2H4 selectivity over time is caused by the unreduced nickel fraction. Furthermore, the (Ni-A1-Ti)sg and (Ni-A1-Ti)imp samples have equal percentages of reduction and their selectivity patterns are quite different. Therefore, the size of the metallic particles and the metal-support interaction could be the cause of the individual behaviour of each catalyst regarding selectivity. This hypothesis is confirmed by the fact that after reduction at 773 K, the evolution of the selectivity of each catalyst (except Ni-Ti) is similar to that shown by (Ni-Al-Ti)imp after reduction at 573 K. Thus, it seems clear that large metallic particles and therefore weak metal-support interactions are responsible for the evolution in the selectivity to ethylene. This selectivity behaviour is found in the (Ni-A1-Ti)imp catalyst at reduction temperatures as low as 573 K, probably is caused by the method used to incorporate nickel, and is also found in the Ni-A1 and (Ni-A1-Ti)sg catalysts at higher reduction temperatures. This fact obviously is caused by the sintering of the metallic particles. Whatever the reduction temperature, catalysts containing TiO2 are shown to be more selective than catalysts based on A1203, and the following sequence regarding C2H4 selectivity can be proposed: Ni-Ti > (Ni-A1-Ti)sg --- (Ni-AI-Ti)imp > Ni-A1. It can be argued that the high 9

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selectivity of the Ni-Ti catalyst after reduction at 773 K could be due to the low conversion level, but, after reduction at 573 K this is the most selective catalyst and also the most active. Therefore, the presence of TiO2 in the formulation of the support leads to an improvement in the catalytic performance, in the sense of achieving a higher selectivity to the intermediate reaction product.

0.0

Figure 5. Influence of the reduction temperature on ethylene selectivity.

4. CONCLUSIONS As a result of the catalyst preparation methods used, highly dispersed and low crystallinity of oxide phases are obtained. No evidence of nickel oxide particles greater than 5 nanometers was observed in the XRD spectra of the samples in which metal was incorporated during the preparation of the support. On the contrary, XRD of catalysts in which nickel was impregnated clearly showed the diffraction bands corresponding to NiO. Catalysts with nickel incorporated into the support, although exhibiting a higher initial values of specific surface area, also exhibited poorer stability during reduction. In Ni-Ti samples, sintering and phase transition of the support caused a dramatic loss of BET area. Therefore, the thermal stability of A1203 supports is increased by the addition of TiOo. TPR results showed that the reducibility of the accessible metal strongly depends on the method of incorporation of nickel and on the type of support. The amount of accessible metal (metal that can be reduced) in the (Ni-A1-Ti)imp samples is greater than in the Ni-A1, Ni-Ti or (Ni-A1Ti)sg samples. On the other hand, while the accessible metal in samples containing TiO2 can be completely reduced at 773 K, samples with A1203 supports require higher reduction temperatures, because of the formation of NiA1204 species. The catalytic performance and the effect of coke deposition on the activity and selectivity, were found to be strongly dependent on the preparation procedure, on the reduction temperature, and on the type of support. At low reduction temperatures, the deposition of coke on NiA1, (Ni-Al-Ti)sg, and Ni-Ti samples decreased the selectivity to ethylene whereas on (Ni-A1Ti)imp, deposition of coke increased selectivity. At high reduction temperatures, with only the exception of the Ni-Ti catalyst, coke deposition increased the selectivity to the desired product. The pattern of deactivation by coke was also different for the different samples. In Ni-Ti and (Ni-A1-Ti)sg samples, coke formation strongly diminished their activity and simultaneously increased methane production. With Ni-A1 and (Ni-A1-Ti)imp samples, coke did not cause an significant deactivation or an increase in methane yield. Finally, catalysts

616 containing TiO2 (Ni-Ti, (Ni-A1-Ti)sg and (Ni-A1-Ti)imp) provided higher ethene selectivities than Ni-A1 catalysts. These results are a consequence of the differences in the interactions between Ni particles and the support.

Acknowledgments The authors wish to acknowledge the financial support of DGICYT (MEC, Spain) (Project PB94-0568) and of the Instituto de Cooperaci6n lberoamericana (MEC, Spain).

REFERENCES. 1. 2. 3. 4 5. 6. 7 8. 9. 10. 11 12 13 14. 15. 16. 17 18. 19.

G. Sankar, C.N.R. Rao and T. Rayment, J. Mater. Chem., 1 (1991) 299. T. Arunarkavalli, G.U. Kulkarni, G. Sankar and C.N.R. Rao, Catal. Lett., 17, (1993) 29. G. Sankar, K.R. Kannan and C.N.R. Rao, Catal. Lett. 8 (1991) 27. S.J. Tauster, S.C. Fung and R.L. Garten, J. Am. Chem. Soc., 100 (1978) 170. S.J. Tauster and S.C. Fung, J. Catal., 55 (1978) 29. T.J. Gardner, C.H.F. Peden and A.K. Datye, Catal. Lett., 15 (1992) 111. S. Chien, B.N. Shelimov, D.E. Resasco, E.H. Lee and G.L. Haller, J. Catal., 77 (1982) 301. D.E. Resasco and G.L. Hailer, J. Catal., 82 (1983) 279. G.B. Raupp and J.A. Dumesic, J. Catal., 97 (1986) 85. A.J. Simoens, R.T.K. Baker, D.J. Dwyer, C.R.F. Lund and R.J. Madon, J. Catal., 86 (1984) 359. M.G. Sanchez and J.L. Gazquez, J. Catal., 104 (1987) 120. B.A. Sexton, A.E. Hughes and K. Foger, J. Catal., 77 (1982) 85. J.A. Montoya, J.C. Rodriguez, I. Schifter, A. Monzon and T. Viveros, Stud. Surf. Sci. and Catal., 88 (1994) 531. S. Narayanan and G. Sreekanth. J. Chem. Soc. Faraday Trans., 85 (1989) 3785. J.M. Gallardo, V.Sanchez and G. Busca,, J. Mat. Chem., 5 (1995) 1245. J.C. Rodrfguez, E. Romeo, A. Monz6n, A., Borgna and A.J. Marchi, Proc. of XV Iberoamerican Symposium on Catalysis, C6rdoba (Argentina), 1996, pp. 909-914. J.C. Rodrfguez, C. Tellez, J.L.G. Fierro, A. Montoya, A. Monz6n and T. Viveros, in preparation (1997). J.A. Pefia, J. Herguido, C. Guimon, A. Monz6n and J Santamarfa, J. Catal. 159 (1996) 313. F. Medina, P. Salagre, J.E. Sueiras and J.L.G. Fierro, J. Chem. Soc. Farad. Trans., 90 (1994) 1455.

9

Elsevier Science B.V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

617

Sintering of Ni/SiO2 Catalysts Prepared by I m p r e g n a t i o n and D e p o s i t i o n Precipitation D u r i n g C O H y d r o g e n a t i o n G. Martra a*, H. M. Swaan b, C. Mirodatos b, M. Kermarec a, and C. Louis a aLaboratoire de R6activit6 de Surface, URA 1106 CNRS, Universit6 Paris VI, 4 place Jussieu, F75252 Paris Cedex 05, France. blnstitut de Recherches sur la Catalyse, UPR 5401 CNRS, 2 avenue Albert Einstein, F69626 Villeurbanne Cedex, France

Two Ni/SiO2 catalysts, prepared by impregnation and deposition-precipitation, exhibit different mechanisms of particle sintering during CO hydrogenation at 230~ The metal particle size distribution of the first catalyst remains monomodal but is shifted towards larger size whereas that of the second one becomes bimodal. In situ DRIFT experiments during the catalytic reaction, and CO adsorption at room temperature on both catalysts show that the mechanism of particle sintering may be explained in terms of different strengths of the metal-support interaction in the two catalysts. 1. INTRODUCTION Loss of metal surface area by poisoning, coking [1-3] or sintering constitutes a serious problem in a number of catalytic reactions involving Ni catalysts, such as CO hydrogenation or steam reforming. During these last decades, many studies have been carried out in order to understand the mechanisms of deactivation by carbon deposition and particle sintering and therefore to attempt to minimize these phenomena. Particle sintering during catalytic reaction is often explained by i) the particle migration and coalescence [4,5] or ii) the Ostwald ripening model of atomic species migration [6]. These processes are generally favored at high reaction temperatures (T > 500~ In the case of low temperature of CO hydrogenation reaction (200-250~ with Ni/SiO2 catalysts, the Ni particle growth involves a "chemical sintering", i.e., the migration of Ni subcarbonyl species extracted from low coordinated Ni atoms (comer, step or kink position) and their transport towards dense planes [7,8]. In this work, we have studied the catalytic behavior of two types of Ni/SiO2 catalysts, prepared by impregnation (INi) and deposition-precipitation (DPNi). They have been previously characterized before and after reduction [9-12]. An interesting feature is that the nickel metal particles of the DPNi samples exhibit thermal stability and resistance to sintering under H2 or inert gas at high temperatures (800-900~ higher than the INi samples [ 13-16]. The resistance to sintering would arise from stronger metal-support interaction [13,14,16]. It was proposed [10] that the interface between Ni(II) and silica is a Ni(II) brucitic layer bonded to silica, and that after reduction, the metal-silica interface still contains ionic nickel species, which would act as anchoring sites for the metal particles, leading to the stabilization of the metal particles onto the silica surface and to the resistance to sintering. ~ to whom correspondence should be addressed; present address: Dipartimento di Chimica IFM, Universith di Torino, via P. Giuria 7, 110125 Torino, Italy; fax: +39-11-6707855; e-mail: [email protected]

618 2. EXPERIMENTAL

2.1. Materials Ni/SiO2 catalysts (18 wt. % Ni), prepared by incipient wetness impregnation [9] and deposition-precipitation [10-12], were reduced by TPR up to 700~ (INi) and 900~ (DPNi) with a 5 % H2/Ar mixture, i.e., up to the minimum temperature required to obtain full nickel reduction [9,10]. For both catalysts, porous silica Spherosil XOA400 (Rh6ne Poulenc, A 400 m2.g -1) was used as a support. 2.2. CO hydrogenation. The reaction (H2:CO = 3:1) was carried out at 230 ~ and at atmospheric pressure in a differential fixed bed flow reactor during 75 h. Gas analyses were performed by on-line FID chromatography at the reactor outlet. The degree of conversion was maintained low enough (< 3 %) to ensure differential and isothermal conditions. It was checked that external diffusion did not limit the conversion. 2.3. Electron Microscopy. TEM images of the catalysts before and after the catalytic runs were obtained on a Jeol 100CXII microscope. Histograms of the metal particle size distributions were established by counting about 500 particles. 2.4. DRIFT spectroscopy. In situ DRIFT spectra were recorded with a Nicolet 550 instrument, equipped with a Spectra Tech Diffuse Reflectance Unit, which permits the recording of the spectra of the catalytic bed under the reaction conditions reported above. In this case, for practical reason, the catalytic runs were stopped after 15 hours. 2.5. Transmission FTIR spectroscopy. Catalyst powders as self-supporting pellets, were put into an IR cell connected to a conventional vacuum line. The catalysts were reduced in the IR cell in the conditions reported above. CO (99.99 %, Air Liquide) was adsorbed at room temperature after liquid nitrogen trapping. Spectra were recorded with a Bruker IFS 66V spectrometer, using a 4 cm- 1 resolution and 128 accumulating scans. 3. RESULTS

3.1. Catalytic reaction CH4 is the main hydrocarbon produced (S ~ 80 %) as usually observed with supported Ni catalysts. Fig.1 reports the curves of the normalized activities versus the time on stream. Normalized activity is defined as the ratio of the instantaneous activity At to the initial one A0. DPNi and INi exhibit almost the same initial activity A0:1.7 10-5 and 2.1 10-5 (molCO) s-l(gNi'l), respectively. The activity of the DPNi catalyst (Fig. la) slightly increases during the 20 first hours of the run and then slowly decreases to finally lose 10 % of the initial activity after 75 h of reaction. In contrast, a gradual decrease of the activity is observed from the beginning of the reaction for the INi catalyst, which undergoes a deactivation of about 25 % after 75 h (Fig. 1b). Stronger deactivations are reported in the literature for Ni/SiO2 catalysts during CO hydrogenation at 230~ but they are consistent with a lower H2:CO ratio = 2" 1 in the reaction mixture [7,8]. 3.2. Evolution of the metal particle size distribution (MPSD) TEM images of the two catalysts were taken before and after the catalytic run, and the respective MPSD are reported in Fig. 2. A monomodal MPSD is observed for both fresh catalysts (Fig. 2a and 2b).

619 A t/A 0 1.0

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Figure 1" Normalized activity versus time on stream for: (a) DPNi; (b) INi However, the INi sample exhibits a mean metal particle diameter (dm = 40 A) smaller than the DPNi one (dm = 57 A). Assuming the same particle shape, it may be deduced that the ratio of the number of low coordinated Ni atoms to the number of Ni atoms on surface planes is higher for the INi sample than for the DPNi sample. In other words, the particles of INi exhibit higher roughness than those of DPNi. After the catalytic reaction, the size distributions have changed in a different way for the two samples. An overall shift of the Ni particle sizes towards larger values is observed for the INi catalyst, still resulting in a monomodal MPSD ranging from 100 to 500 A, with a mean diameter of 265 A (Fig. 2al). In contrast, the DPNi sample exhibits two sets of particle sizes after 75 h of reaction: a set of smaller particles (17 - 60 A with dm = 38 A) (Fig. 2bl), and a set of much larger faceted crystallites (400 - 750 A with dm = 530 A) (Fig. 2b2). A rough evaluation of the relative amount of the two populations indicates that the second set represents about 1/1000 of the total number of the metal particles, corresponding yet to about 70% of the Ni weight. %25

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Figure 2: Normalized MPSD determined from TEM micrographs: (a) INi before reaction; (al) INi after 75 h of reaction; (b) DPNi before reaction; (bl) and (b2) DPNi after 75 h of reaction

620 TEM images were also taken after 15 h of reaction. For the INi sample, a monomodal MPSD is observed but already shifted towards larger sizes (dm= 78 A). The DPNi sample exhibits a similar MPSD (dm = 60 A) as before reaction (dm = 57 A) but very few faceted particles (600 A) are also observed, much less numerous than after 75 h of reaction. The metallic surface area of the INi sample, before and after reaction, was estimated from the micrographs, assuming that i) there is no loss of Ni and ii) the Ni particles are spherical. The metallic surface area was found to decrease from 153 m 2 (gNi) -1 before reaction to 21 m 2 (gNi) -1 after 75 h of reaction, i.e., an 85 % decrease. For the DPNi sample, the change in the metallic surface area cannot be easily calculated because of the presence of faceted particles after the reaction. 3.3 In situ DRIFT studies

To obtain information on the evolution of the metallic surface during CO methanation, in situ DRIFT measurements were also performed, even if, as reported in the experimental

section, the coupled catalytic runs were limited to 15 h. For both catalysts, the IR spectra (Fig. 3) exhibit bands in the range of linear/subcarbonyl (2100-2000 cm -~) and bridged carbonyl species (2000-1700 cm -1) [7,8,17-20]. Their broadness reflects the surface heterogeneity of the Ni particles. In both cases, the range of linear carbonyls is dominated by a component at 2067 cm -1, which may be assigned to subcarbonyl nickel species Ni(CO)n (n = 2,3) formed on low coordinated Ni atoms (comer, step or kink position) [ 19,20]. Besides these common overall features, the spectra of the two samples differ to some extent in the relative intensity and position of the various components. The ratio of the subcarbonyl (2067 cm -1) to the linear carbonyl (2030 cm -1) bands is higher for the INi catalyst (Fig. 3a) than for the DPNi (Fig. 3b). This is consistent with the smaller mean size of the INi particles, i.e. their higher roughness, as discussed above (.Fig. 2a). The main components of the bridged bands are observed at 1920 and 1890 cm "j for INi and DPNi samples, respectively. Shoulders around 1920 and 1820 cm -1 are also observed on the DPNi sample. These differences reflect different surface morphologies of the Ni particles in the two catalysts [18,19,22]. Indeed, the shoulder at 1820 cm "1 has been assigned to the multibonded Ni4CO species [21]. Furthermore, previous studies of CO adsorption on Ni single crystals [23-25] suggest that the band at 1920 cm -1 corresponds to Ni2CO on Ni (111) plane and the band at 1890 cm -1 corresponds to bridged CO on Ni (100) plane, possibly as a mixture of Ni2CO and Ni4CO. The presence of Ni4CO species on the DPNi sample gives evidence of the presence of a larger area of dense planes than in the INi sample at the very beginning of the reaction. 0.2

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Figure 3" DRIFT spectra of adsorbed CO (Pco= 190 Torr) during CO methanation reaction at 230~ on: (a) INi after 5 min; (a') INi after 15 h; (b) DPNi after 5 min; (b') DPNi after 15h

621 The intensities of the IR bands of the two samples evolve in different ways during the catalytic reaction without significant change in the shape of the spectra. In the case of the INi sample, a very slight decrease of the overall intensity is observed (Fig. 3a'), whereas the bands of CO adsorbed on the DPNi catalyst significantly increase in intensity (Fig. 3b'). For this latter sample, the increase in the intensity of the bridged CO bands indicates the development of the dense surface planes during the reaction and thus the smoothing of the metallic surface, since bridged carbonyls are preferentially formed on regular surfaces where sets of Ni atoms with the required geometry may be found [ 17]. Coupled measurements of the catalytic activity by monitoring the gas composition at the DRIFT cell outlet, indicate a slight decrease of the activity for the INi sample and a slight increase for DPNi within the 15 h of reaction, in agreement with the results obtained during the same period with the conventional catalytic test (Fig. 1).

3.4. FTIR study of adsorption of CO at R.T. "Chemical sintering" is known to increase with decreasing CO adsorption temperature and increasing CO pressure [7,8]. So, experiments of CO adsorption at room temperature were also carried out. Although these experimental conditions are different from those of the catalytic reaction, one may expect to get information on the change in the particle morphology through the extraction of low coordinated Ni atoms. The spectra of CO adsorption at RT on DPNi are reported on Fig. 4. After admission of 1 Torr of CO, the IR bands assignable to subcarbonyl (2080 cm'l), linear monocarbonyl (2040 cm "1) and bridged carbonyl (1960 cm "1) species are observed (Fig. 4a) [7,17]. The band positions differ to some extent from those obtained by in situ DRIFT experiments (Fig. 3). This probably arises from different types of carbonyl species and/or surface distribution, depending on temperature and CO partial pressure (also affecting the band intensities) and on the presence, in the case of the DRIFT measurements, of H and CHx coadsorbed species.

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Figure 4: FTIR spectra of CO adsorbed at RT on DPNi: (a) 1 Torr of CO; (b) 10 Torr of CO, just after admission onto the sample; (c) - (g) after 1, 3, 5, 8 and 13 h of contact. Inset: zoom of (b) and (g) in the 2000-2120 cm -l range showing the depletion of the 2080 cm -1 band

622 When the CO pressure is increased to 10 Torr, the subcarbonyl band increases, and a new peak at 2046 cm "1 becomes predominant (Fig. 4b). This band is ascribed to the T2 mode of Ni(CO)4 physically adsorbed on the silica surface [26], the formation of such species being expected when Ni particles are exposed to such a high CO pressure [ 18,27,281. By increasing the time of contact with CO (10 Torr), the band of Ni(CO)4 at 2046 cm "l progressively grows up (Fig. 4c-g), whereas that of subcarbonyl (2080 cm -1) undergoes a gradual depletion and almost completely disappears (Fig. 4d-g). The complex bridged carbonyl bands are also modified as follows: the stronger band at 1960 cm -1 increases and shifts down to 1945 cm -1 and the intensity of the shoulder at 1895 cm -1 increases as well (Fig. 4b-g). Subsequent cycles of CO adsorption (10 Torr) and outgassing were also performed. After several cycles, the subcarbonyl and the tetracarbonyl bands are strongly depleted whereas the intensity of the bridged CO bands continues to increase (spectra not shown for the sake of brevity). All these results show that i) the subcarbonyl species are the precursors for the Ni tetracarbonyl formation; ii) the Ni atoms preferentially extracted by CO are those in low coordination (comer, step or kink) [18-20]; iii) they are transported towards dense planes, which induces a smoothing of the metal particles, and possibly a change in their size. When CO is adsorbed on the INi sample and maintained (spectra not shown), the intensities of the subcarbonyl (~ 2080 cm "1) and tetracarbonyl (~ 2048 cm -1) components are noticeably lower than in the case of the DPNi sample (Fig. 4), whereas the bridged CO band at 1895 cm "1, which is not visible in the DRIFT spectra (Fig. 3a, a'), is higher than in the DPNi sample (Fig. 4). This experiment suggests a higher reactivity of the low coordinated Ni atoms for the INi sample. Indeed, one may proposed that they are instantaneously extracted as subcarbonyl and tetracarbonyl species and decomposed after migration over the surface of dense planes, as indicated by the high intensity of the bridged CO band, i.e., the smoothing of the Ni particles. These results are consistent with previous studies on the fragmentation of Ni single crystal, showing the reconstruction of the metallic surface through the disappearance of low coordinated Ni atoms and the subsequent development of dense (111) planes [29-31 ]. 4. DISCUSSION Different MPSD, monomodal and bimodal, were observed after CO methanation for the 1Ni and DPNi samples, respectively (Fig. 2). However, in both cases, a similar "chemical sintering" process via nickel subcarbonyl formation on low coordinated Ni atoms, extraction and migration towards dense surface planes is probably involved under these reaction conditions, in agreement with previous studies on CO methanation at 230~ with Ni/SiO2 catalysts [8]. The difference in the final MPSD has therefore to be related to the initial state of the catalysts, i.e., to the different metal-support interactions. As mentioned in the introduction, strong interaction between the nickel metal particles and the silica-support characterizes the DPNi sample [10,13,14,16]. It may be noted that the evolution of the MPSD of the DPNi catalyst is similar to that already observed for Ni/SiO2 catalysts prepared with the ammonia method [8]. This may be due to the fact that the same supported Ni(II) precursor, i.e., nickel phyllosilicate, is formed during these two preparations [10-12,32]. The Ni(II) precursor probably leads to the same type of Ni-silica interaction and to the same kind of Ni particles after reduction, and therefore to the same type of sintering. It may be postulated that the Ni atoms extracted from the DPNi particles would first cluster and grow up on the smallest metal particles in strong interaction with silica, even not visible by TEM. These particles may act as nuclei, and give rise to smaller particles than the initial ones. In contrast, the weaker metal-support interaction in the INi sample may explain that the MPSD shifts toward larger particles during the methanation reaction,

623 according to a simple Ostwald ripening process, which maintains the monomodal distribution [8]. Although there is no direct evidence for the faceting of the smallest particles of the DPNi sample, it may be speculated from the observation of the few large faceted particles that the sintering process would favor the particle growth along a preferential axis, as already demonstrated in the case of the Ni/SiO2 catalysts prepared by the ammonia method [8]. The DRIFT spectra of the DPNi and INi samples observed during CO methanation show that the intensity ratio of the Ni4CO bands (1890 and 1820 cm -1) to the Ni2CO one (1920 cm -1) is higher for the DPNi sample (Fig. 3). Moreover, the stronger increase in intensity of the multibonded CO bands for the latter sample, indicates the formation of larger dense surface planes. These planes are known to lead to stronger Ni-CO bonding and therefore to favor C-O breaking and in consequence the methanation reaction [8,21,22]. This would explain that the deactivation is weaker for the DPNi catalyst than for the INi catalyst. Although the micrographs of the 1Ni catalyst seem to exhibit spherical particles and the DRIFT results show very small changes during the methanation reaction, it is likely that these particles also ensure slight modifications of morphology towards smoothing. This is attested by the experiment of CO adsorption (10 Torr) at RT, which is known to accelerate the sintering process [7,8]. The results show that (i) the initial particles smaller than in DPNi, ensure a higher reactivity for the low coordinated Ni atoms, which are extracted through subcarbonyl Ni species and Ni(CO)4; (ii) these particles are smoothened, as evidenced by the strong decrease of the subcarbonyl band and the appearance of multibonded species corresponding to the formation of dense surface planes. Now, the catalytic results may be rationalized. The main feature of the DPNi catalyst is its catalytic stability over a long period of reaction (90 % of the initial activity after 75 h) with even a slight increase of activity during the 15 first hours of reaction consistent with an increase in the intensity of the IR bands in the DRIFT experiment (Fig. 3). This catalytic stability may be related to the strong metal-support interaction developing during the preparation and activation steps. By stabilizing the Ni atoms on the silica surface, this interaction tends to slow down the particle sintering process during the reaction and to favor the faceting of the particles growing along a preferential axis. The result of this process is the slow formation of new small particles (tending to increase the metallic surface area) but also of some large and faceted particles (tending to decrease the surface area), leading to the observed bimodal size distribution. The initial slight increase of catalytic activity followed by a slow decrease after 75 h on stream would therefore result from the combination of at least two processes: i) the slow changes in metallic surface area, and ii) the continuous smoothing of the nickel particles tending to increase the intrinsic activity of the catalyst via the formation of dense planes on which the CO dissociation can occur more easily. For the INi sample, the weaker metal-support interaction does not prevent a large and continuous sintering process to occur, leading to a drastic decrease of the metallic surface area (85 % after 75 h on stream). However, the particle growth involves a smoothing effect, so the simultaneous increase in intrinsic activity partially compensates the loss of surface, limiting the decrease of the overall catalytic activity to 25 % over the same ageing period. In addition to the sintering phenomena, the carbidization of the nickel particles and the formation of external carbon deposits, already revealed by temperature programmed hydrogenation carded out after the methanation reaction [33], may also contribute to the overall deactivation process. Further experiments aimed at quantifying this type of deactivation are in progress.

624 ACKNOWLEDGEMENTS

G. Martra thanks the Consiglio Nazionale delle Ricerche and NATO for the financial support for this study. REFERENCES

1. 2. 3. 4. 5. 6. 7. 8. 9. 10. 11. 12. 13. 14. 15. 16. 17. 18. 19. 20. 21. 22. 23. 24. 25. 26. 27. 28. 29. 30. 31. 32. 33.

J . R . Rostrup-Nielsen, in "Steam Reforming Catalysts", Danish Technical Press, Copenhagen, 1975. D.L. Trimm, Catal. Rev. Sci. Eng., 16 (1977) 155. C.H. Bartholomew, Catal. Rev. Sci. Eng., 24 (1982) 67. B. Pulvemacher and E. Ruckenstein, AICHE J., 19 (1973) 1286. P.C. Flynn and S. E. Wanke., J. Catal., 37 (1975) 432. P. Wynblatt and N. A. Gjostein., Prog. Solid State Chem., 9 (1975) 21. C. Mirodatos, H. Praliaud and M. Primet, J. Catal., 107 (1987) 275. M. Agnelli, M. Kolb and C. Mirodatos, J. Catal., 148 (1994) 9; M. Agnelli, Thesis, 1992, Universit6 C. Bernard - Lyon I. C. Louis, Z.X. Cheng and M. Che, J. Phys. Chem., 97 (1993) 5703. P. Burattin, Thesis, 1994, Universit6 P. et M. Curie - Paris VI. P. Burattin, C. Louis and M. Che, J. Chim. Phys., 92 (1995) 1377. P. Burattin, C. Louis and M. Che, J. Chem. Phys., submitted. M. Montes, C. Penneman de Bosscheyde, B. K. Hodnett, F. Delannay, P. Grange and B. Delmon, Appl. Catal., 12 (1984) 309. M. Montes, J. B. Soupart, M. de Saedeleer, B. K. Hodnett and B. Delmon, J. Chem. Soc., Faraday Trans. I, 80 (1984) 3209. J.W.E. Coenen, Appl. Catal., 75 (1991) 193. M.A. Keane and P. M. Patterson, J. Chem. Soc., Faraday Trans., 92 (1996) 1413 N. Sheppard and T. T. Nguyen, in "Advances in Infrared and Raman Spectroscopy", Vol. 4, P. J. Clarke and R. E. Mester (Eds), John Wiley, New York, 1978, p. 67. J.A. Dalmon, M. Primet, G. A. Martin and B. Imelik, Surf. Sci., 50 (1975) 95; J. A. Dalmon, M. Primet, G. A. Martin and B. Imelik, Surf. Sci., 50 (1975) 25. D.G. Blackmond and E. I. Ko, J. Catal., 96 (1985) 210. C.H. Rochester and R. J. Teller., J. Chem. Soc., Faraday Trans. I, 73 (1977) 609. G.A. Martin, M. Primet and J. A. Dalmon, J. Catal., 53 (1978) 321. A. Parmaliana, F. Arena, F. Frusteri, S. Coluccia, L. Marchese, G. Martra and L. Chuvilin, J. Catal., 141 (1993) 34. J.C. Bertolini, G. Dalmai-Imelik and J. Rousseau, Surf. Sci., 68 (1977) 539; J. C. Bertolini and G. Tardy, Surf. Sci., 102 (1981) 131. J.C. Campuzano and R. G. Greenler, Surf. Sci., 83 (1979) 301. K. Christmann, O. Shober and G. J. Ertl, J. Chem. Phys., 60 (1974) 4719. K. Moana Rao, G. Spoto and A. Zecchina, Langmuir., 5 (1989) 319. E. Guglielminotti, A. Zecchina, and F. Boccuzzi, in "Growth and Properties of Metal Clusters", J. Bourdon (Ed), Elsevier, Amsterdam, 1980, p. 165. W.H. Shen, J. A. Dumesic and C. G. Hill, J. Catal., 68 (1981) 152. D.B. Liang, G. Abend., J. H. Block and N. Kruse, Surf. Sci., 126 (1983) 392. W.A. Schmidt, J. H. Block and K. A. Becker, Surf. Sci., 122 (1982) 409. P. De Groot, M. Coulon and K. Dransfeld, Surf. Sci., 94 (1980) 204. O. Clause, M. Kermarec, L. Bonneviot, F. Villain and M. Che, J. Am. Chem. Soc, 114 (1992) 4709. G. Martra, et al, unpublished results.

9

Elsevier Science B. V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

625

An Experimental Protocol for Studying Kinetics and Catalyst Deactivation: Application to Heptane Reforming on Pt-Re/A1203 K. Liu ab., S. C. Fung b, T. C. Ho b, and D. S. Rumschitzkia a Department of Chemical Engineering, City College of the City University of New York, New York, NY 10031 b Exxon Research and Engineering Company, Annandale, NJ 08801

We propose an experimental protocol that can provide a diagnostic tool for probing the state of catalyst coking and its influence on product composition in an operating reactor throughout the lifetime of the catalyst. The experimental equipment includes a vibrational microbalance, a multi-outlet fixed bed reactor, and a temperature-programmed oxidation apparatus. We illustrates its use by applying it to n-heptane reforming on a commercial PtRe/A1203 catalyst. The results identify five-membered naphthenes (C5N) as the key source of coke and provide the data needed to develop kinetic models for both reforming and coking. From the measured or predicted C5N concentration, the models correctly predict the reactor coke profile and the reforming product composition as functions of on-stream time.

1. INTRODUCTION Catalyst coking is ubiquitous in hydrocarbon processing. It is highly desirable to develop a practical tool of general use for probing the state of catalyst coking in an operating reactor. The present work is an attempt to address this need. Specifically, we propose a relatively simple experimental protocol that can provide much of the needed information. We apply the protocol to n-heptane reforming on a Pt-Re/A1203 catalyst. The protocol consists of the following experiments. A vibrational microbalance identifies which of the hydrocarbon types amongst the reaction participants is the major coke source. With the thus-identified coke source, we measure the intrinsic coking rate and develop a mechanistic coking kinetics model by using the microbalance and a temperature-programmed oxidation (TPO) apparatus. The latter provides information on the nature of the coke. The next step is to develop a lumped reaction network for the main reaction that contains the coke source as one of its lumps. For discriminating among rival lumped reaction networks, the microbalance operated in differential mode probes the relative importance of individual lumped reactions and helps eliminate kinetically insignificant reaction pathways. Once a network emerges as the leading candidate, the microbalance gives fairly fight estimates of the rate constants for selected subsets of the network by using the reaction intermediates as the feed. * The current address of Ke Liu is Mobil TechnologyCompany, Paulsboro, NJ 08066

626 The remaining data are amassed in an integral multi-outlet fixed-bed reactor (e.g., 12 conversion vs. space time data in just three runs). These data allow estimation and testing of the remaining rate constants by using the already-determined reaction paths and rate constants. The resulting kinetic model obtained from this approach has a minimum number of truly adjustable model parameters. Finally, the coking and main reaction models are combined to predict reactor coke profiles and the evolution of product composition throughout the lifetime of the catalyst. The present paper gives an overall perspective of our experimental techniques and modeling approach. Details of the work will be published in a series of forthcoming papers.

2. THEORETICAL BASIS

Since the n-heptane reforming rate is much faster than catalyst deactivation rate, the reaction system considered, with appropriate assumptions, is represented by the following quasi-steady state mass balance equations: c7Ci

3t

= & (c~, c: ;.. c,, ...)r

= r(C 7)~b(Ce)

)

i = 1, 2, 3,.-.

(1)

(2)

Here C~ are the concentrations of reaction participants, Ck the concentration of coke on catalyst, and C; (a subset of C~) the concentrations of those reaction participants which are coke source. Rj, x, t, and r are the main reaction rate, space time, on-stream time, and coking rate, respectively. The deactivation function ~ is scaled such that ~ = 1 at t = 0. An assumption here is that the main reaction and coking have the same ~. This can be relaxed, however. We use the protocol to determine R~, r, and ~. The majority of the intrinsic coking kinetic data derive from the vibrational microbalance operated in differential mode. In a given microbalance experiment, both the temperature and hydrocarbon pressures remain essentially constant. In this environment, the only quantity that varies significantly with t is Ck. Then

dCe = rof~(Ce) dt

(3)

where ro is the initial coking rate at t = 0 and is time invariant. Our aim is to predict C~(t, x) and Ck(t, x) from measured or predicted C/. Note that Ck at the end of catalyst life is an initial condition for modeling and control of the catalyst regeneration process. In developing the coking kinetic model for nC7 reforming, the data lead us to consider that coke can deposit on both active sites and on already-coked sites. This consideration, as will be seen, gives simple explicit expressions for r as a function of Ck or t.

627 3. EXPERIMENTAL 3.1. Vibrational microbalance. We have previously described the newly developed vibrational microbalance reactor system which eliminates many undesirable features (e.g., gas bypassing, buoyancy/drag effects, etc.) of conventional thermogravimetric analyzers [1, 2]. Figure 1 shows that the system consists of a flow-through tapered element oscillating microbalance, a valving control feeding unit and a product analysis unit with a HP5890 GC and a 50-m capillary column coated with cross-linked methylsilicone gum. Since the catalyst is in a packed bed at the tip of a hollow tube, all the feed gas flows through the bed and hence the space velocity is precisely known. For a vibrating assembly designed to hold 0.1 g catalyst, it detects a mass change of 10.5 g. FEED

=

PURGE GAS

',

~

_q'l

"

r

TAPERED ELEMENT ........ i'"

.,olo o l = /

FEEDBACK

SYSTEM

In

n I

I~ ~1

! ../ / ,/:

CATALYST BED

/

,!' I,

!o----i

TOCONTROLJ UNIT

I

/ i"

=

/,,

, TO GC

~

-- FEED BYPASS

-- CATALYST

GRAVIMETRIC MICROBALANCE

~

-- NO FEED BYPASS

- - - CATALYST

VIBRATIONAL MICROBALANCE

Figure 1. Flow diagram of the vibrational microbalance 3.2 Multi-Outlet Fixed-Bed Reactor As detailed in [3], this multi-outlet reactor is fabricated from a 1.27 cm OD stainless steel tube with an ID of 1.02 cm. It has four outlets along its wall, which permit sampling of the gas at four bed heights for product analyses by GC. This gives data at four different space times for a given flow rate and catalyst loading. The flow rate of the sampling stream is controlled to below 5% of the total flow so that the catalyst downstream of the sampling point experiences no significant change in space velocity. In this way we can efficiently amass conversion vs. space time data by either changing the catalyst loading in different sections, by varying the feed rates, or both. The reactor bed is isothermal via dilution with inert particles. 3.3. Temperature-Programmed Oxidation The catalyst in each reactor section can be unloaded without mixing and its coke content determined by a highly sensitive TPO technique [4], using a modified Altamira temperatureprogrammed unit (Model AMI-1). In this modification, the gas exiting the reaction cell enters a methanator where CO2 and CO are converted to methane over a Ru catalyst with a constant supply of hydrogen. The methane formation rate is measured by an FID detector.

628

3.4. Materials The gases used are H2 (99.95%), He (99.95%), and NE (99.95%). They all have gas purifiers to remove traces of oxygen and water. The liquid feed n-heptane (nC7) , methylcyclopentane (MCP), toluene, ethyl-cyclopentane (ECP) and 2-methyl-hexane (2MH) are all analytic pure grade (99 mol%). MCP and ECP are two of the C5N that hexane and heptane reforming generate and 2MH is a model compound for isoheptanes. The Pt-Re/AI203 catalyst (1/16" OD, 3/16" maximum length) in the fixed-bed reactor contains 0.3 wt% Pt, 0.3% Re and 0.9 wt% CI, and has a N2 BET surface area of 200 m2/g. Two Pt/AI203 catalysts with 0.3 and 0.6 wt% Pt and 0.9 wt% C1 were used in selective experimnents to determine the nature of the coke deposits. We neither presulfide the catalyst nor add sulfur to the feed during the run. Catalyst particle sizes are 177-250 ~tm (60/80 mesh) in the microbalance experiments.

4. IDENTIFICATION OF COKE SOURCE Following a number of suggestive fixed-bed experiments, we determine which of the hydrocarbon types is the major coke source by comparing the coking rates of different hydrocarbon feeds in the vibrational microbalance at 207 kPa, 750~ and H2/hydrocarbon = 3.0. Figure 2 shows the results for the following feeds: toluene, nC7, 96% toluene + 4% ECP, 96% toluene + 4% MCP, 50% toluene + 50% MCP and pure MCP and ECP. Evidently, MCP and ECP produce the highest amounts of coke and, in fact, much more than nC7 or toluene. We also compare the coking rate of a 50/50 MCP-toluene feed mixture with that of pure MCP. This is motivated by a prior proposition [5] that the intercondensation between toluene and C7 cyclopentadienes is responsible for most of the coke. As Fig. 2 shows, the coking rate of the MCP-toluene feed is much slower than that of MCP under our conditions, contrary to what one would expect if the intercondensation were dominant. In fact, as will be seen later, the initial coking rate ro is linear in the partial pressure of MCP. The coking rate of toluene in Fig. 2 is negligible compared to that of nC7, MCP or ECP. Under the conditions used, the nC7 feed produces a total of-~4 wt% C5N at the microbalance exit. The 4% C5N + 96% toluene feed mixture shows a coking rate quite similar to that of nC7, suggesting that C5N coking alone accounts for essentially all of the nC7 coke and thus further confirming that toluene contributes very little to coke formation. Finally, the coking rates of MCP and ECP are almost indistinguishable from each other. This indicates that at least two, and may be most C5N have similar coking rates. These results, together with those in [3, 6, 7], indicate that coke in paraffin reforming derives mainly from the C5N intermediates. We next use readily available MCP as the model compound for C5N to study coking kinetics. 4.1. Deactivation Function

Figures 3 and 4 display Ck (g coke/g cat.) vs. t for different partial pressures of MCP (PucP) and of H: (PH) at 717~ and 50 WHSV, respectively. Increasing PucP increases coking while PH has the opposite effect. These microbalance data can best be represented by ~ = exp(aCk) [8], which upon integrating Eq.(3) gives Ck = (1/a)ln[1 + arot] for data fitting. The ro obtained from fitting the data in Figs 3 and 4 shows the correct dependence on coking

629 conditions [8] and the resulting a value, 56.8 g cat/g coke, is essentially independent of operating conditions including temperature [8], as it should be.

0"06t

'

,

~

to o C~ t~

0.04

To1+50% MC

C5N

a; O

~,~

0.02 To1+4% MCP

0.00

Tol

0

20

40 Time

60

t, hr

Figure 2. Microbalance data of Pt-Re/A1203 on coke vs. time for different hydrocarbon feeds 7.5 I . . . . . . . . . . PMcp=15

A

o

1

5.0.

v

o o

(9 x

o

2.5

0

0.53

0.0

0

10

20

30

40

50

60

Time (hr)

Figure 3. Ck vs. t for different PMCPat 717~ and 50 WHSV with Pt-Re/AlzO3

The foregoing result has a simple physical interpretation if site coverage is the sole cause of catalyst deactivation. Suppose that a small protocoke molecule forms on an active site. Suppose further that rather than automatically depositing itself there, it samples both the

630

A r

r.3

PH=15

4~-

O) v

19.5

O

o

X (D

o

O

30

,

0

10

20

30

40

Time (hr) Figure 4. Ck vs. t for different P/~ at 717~ and 50 WHSV with Pt-Re/A1203 uncoked and coked sites and then deposits randomly on one of them and quickly polymerizes there. We write the rate of active site loss as 1

dS

--

dt

S dC k = -fl-- ~

(4)

So dt

where So and S are the active site densities at time zero and t, respectively, and 13 is a unit conversion factor from coke content Ck to site density S. The site-coverage mechanism implies that + = S/So. Then Eq.(4) upon integration gives the experimentally observed result

q~= S/So = exp(-otCk)

with ot - [3/So

(5)

which from Eq.(3) gives ~ = 1/(1 + arot ) and shows a is independent of reaction conditions. Equation (5) is an example of the Poisson process in probability theory [9]. S decays slowly with t, so that enough sites survive even at long times to continue forming new coke. This simple model lends itself naturally to the distinction of coke depositing on clean sites, monolayer coke concentration Cm, and coke depositing on already-coked sites, multilayer coke concentration CM. That is, Cu (0 = Cm(t) + C ~ t ) at all times. A simple site balance yields 1 (So_S)

Cm(t) - -~

= 1 (l_e_~q)_

-~

rot l +arot

! rot CM(t ) = --" In(1 + COot) ct 1 + COot

1 Ref.[8] gives a more rigorous derivation based on a pseudo steady state approximation.

(6)

(7)

631 When t >> 1/(aro), Cm approaches 1/a, corresponding to the coke level for a complete monolayer coverage of the active sites, excluding those sites that are lost during the initial hydrocarbon adsorption and coking of the most active sites. The monolayer coke level then is C m , m a x = Co + 1/6gwhere Co, the initial coke buildup, is 0.006 g coke/g cat. for the data shown in Figs. 3 and 4. Figure 5 shows the time evolution of Cm and CM for MCP coking at PMcP= 3 psi, PH = 22.5 psi, WHSV = 50 and 717~ Note that Cm(t) is concave, while Cu(t) is convex for small t and concave beyond the inflection point at t = t* = 1/(Otro), the characteristic time scale of coke buildup. Initially, there is no monolayer coke and thus no multilayer coke. As the monolayer coke begins to build up, multilayer coke formation accelerates. But aider a significant portion of the active sites are coked, the rate of multilayer coke buildup slows down. It bears emphasizing that both the rates of monolayer and multilayer coke buildup depend on active sites (more on this later).

0.06

v

t~ O O) r

measured total coke

0.04 multilayer coke

(I)

o

L)

0.02 monolayer coke

0.00

0

25

50

75

100

Time (hr)

Figure 5. Monolayer and multilayer coke vs. time on Pt-Re/A1203 4.2 Nature of the Coke The above development proposes that both metal and acid sites in the Pt-Re/A1203 collect coke. We now reconcile this proposition with the general belief that Ck increases with metal loading, since metal sites produce coke precursors via dehydrogenation [10]. For this purpose, we use catalysts that contain Pt without Re. The reason is that we do not want to complicate the metal site density measurement by including Re since it is not possible to quantify the reduction of the dehydrogenation activity of Pt by Re [3]. Figure 6 shows the microbalance Ck vs. t traces for two Pt/A1203 catalysts of different Pt loadings, along with that for bare A1203. Fitting these data gives the ot and ro values in Table 1. Indeed, 1/or and ro both increase with Pt loading.

632 Table 1 Values of ro and ot for different Pt loadings Catalyst

ro. 103, g coke/g cat. h 18.02 20.84

0.3%Pt/A1203 0.6%Pt/A1203

1/or. 103, g coke/g cat. 14.5 18.4

Monolayer Coke wt% 2.05 2.44

0.08 .~

=

0.6%Pt/AL

0.06-

0.3%Pt/AI203

O 13)

o 0

0

0.04.

O.02-

0.00~

o

.

io

9

io

9

60

Time (hr) Figure 6. Ck.vs. t for two Pt loadings and A1203; MCP feed, 717~

50 WHSV

Let us roughly estimate So (the sum of initial metal and acid sites). The numbers of ~tmole of Pt per gram of the 0.3% and of the 0.6% Pt catalysts are 15.4 and 30.8, respectively. The number of acid sites based on the 0.9 wt% C1 loading is 254 ktmole/g cat., which is much greater than the number of metal sites. Suppose that all the metal and acid sites participate in forming the monolayer coke and that a change in the metal loading would not affect the number of participating acid sites. Then the theoretical ratio of monolayer coke buildups for the two catalysts is roughly (254 +15.4)/(254 + 30.8) = 0.94 - in fair agreement with the experimental value of 2.05/2.44 = 0.84. It thus seems reasonable to propose that both metal and acid sites collect coke, although it is the former that produces the olefin coke precursors. The acid sites collect most of the coke because they significantly outnumber the metal sites. The TPO results depicted in Fig.7 further illustrates the plausibility of the above supposition in that an increased metal content indeed enhances the overall coke deposition. It is generally accepted that the low-temperature peak (~ 370~ corresponds to coke deposited on, or in contact with the metal site. The intermediate temperature peak, ~ 445~ is associated with the coke on the acid sites close to the metal. The 570~ peak is produced by coke on the acid sites far away from any metal site [4, 11-13 ]. Since most of the coke deposits on acid sites and at a high coke level, the shoulder of the acid-coke peaks in Fig. 7 obscures the peaks for the metal-site coke, but one can see that the coke on acid sites near or far away from metal sites increases with increasing metal loading. The absence of a distinct metal-coke

633 peak in Fig.7 may also be due to the likelihood that most of the Pt sites get buried by the multilayer coke, since the oxidation associated with this peak is believed to be catalyzed by Pt. The above considerations may give us some insight on the coke structure. Assuming that at low Pt loadings the metal-site number increases in proportion to the Pt content whereas the acid site number remains constant, then the number of the participating acid sites can be estimated from the increase of the monolayer coke due to increased metal loading. Let X be

0.6%Pt/AI 2 0 x

~

c r

cl LL

|

200

.... i

400

600

Temperature (~

Figure 7. TPO spectra for catalysts of different Pt loadings and A1203 the number of ~tmole of participating acid sites per g of catalyst. Then the 0.3% and 0.6% Pt catalysts' monolayer coke ratio is proportional to the ratio of their total active sites; that is., (30.8 + X)/(15.4 + X) = 2.44/2.05 = 1.19. This gives X = 65.6 ~tmole/g cat, which says that for the 0.6% Pt catalyst, So ,-' 30.8 + 65.6 = 96.4 ~mole/g cat. Neglecting coke's hydrogen content, the number of ~tmole of carbon in the monolayer is (0.0244/12).106 = 2033 ~tmole/g cat. Thus, the number of carbon atoms per active site is 2033/96.4 = 21, suggesting that an average coke molecule has a highly condensed polycyclic structure containing 5-6 aromatic rings. This coke molecular size is similar to that reported for HZSM5 and USHY zeolites by Guisnet and Magnoux [ 14]. That 65.6 lamole/g cat. (from the monolayer coke data) << 254 ~tmole/g cat. (from catalyst's chloride content) seems to suggest that only one out of four acid sites participates in making monolayer coke. This may be rationalized as arising from a nonuniform distribution of acid sites on A1203. Depending on the crystallographic plane, the surface aluminum ions have widely different coordination numbers. The aluminum ions with a low coordination number tend to interact with the added chloride. This non-uniform distribution of the acid sites increases the local acid site density per unit area in regions where they cluster, thus resulting in more than one acid site covered by a single coke molecule. 4.3 Coking Kinetics This section proposes a mechanism for ro that is consistent with the random coke deposition model, the microbalance and TPO results, and the metal/acid dichotomy discussed

634 above. Conceptually, it posits the following sequence of events. The first is an equilibrated adsorption and dehydrogenation of MCP to form olefinic coke precursors on metal sites. Dehydrogenations are known to be much faster than the other reforming reactions. The precursors adsorbed on metal sites quickly equilibrate with the same precursors adsorbed on acid sites, either via readsorption or by direct surface spillover. These precursors, whether on metal or acid sites, then slowly form protocoke whose size is probably comparable to that of MCP. Once formed, and before progressing to coke, this protocoke molecule would be capable of quickly sampling the catalyst surface - both coked and clean sites (metal and acid sites) - deposit randomly on many of these sites and complete its propagation/termination to its final, more refractory coke form there. If the initiation step that produces the protocoke is first order in the precursor concentration and rate limiting, then the dependence of ro on Pn and PMcP turns out to take the form [8]

(8)

dCk : roq~ = ro e -~k =(Ir + tc 2 /Pn)Pucv e-~k

at

1o.

Here K~ and K2, which depend on temperature and catalyst preparation, are defined as l( 1 "-- ( k l [ S m o ] --[- kl' g 5 g 3 [ S a o ] ) g

1

I( 2 = (k2[Srao] -.[- k2' g 6 g 4 [ S a o ] ) g l g

2

(9)

Sao and Smo are the initial acid and metal site densities. The Kj, k{ and ki are the adsorption equilibrium constants and initiation rate constants discussed in [8]. Equation (8) says that the prodigious coke producers are methylcyclopetadienes and methylclopentenes. The latter predominate at high Pn with ro ~ 1~Pro while the former predominate at low Pn with ro ~ (1/PI~) 2. Equation (9) is consistent with the non-proportional increase in 1/or with increased Pt loading in Table 1. We mention in passing that a long paraffin can produce a multitude of C5N molecules. Let n be the number of hydrogen molecules extracted from the most dehydrogenated. C5N whose concentration is not negligibly small. Then the generalized form of Eq.(8) is d C k = exp(-aCk) dt

PcsN

(1 O)

Data obtained from the microbalance (ro v s . PMCP)confirm that the order of ro with respect to PMcP is indeed unity. The values of K~ and K2for 750~ based on both microbalance and fixed bed data coveting a wide range of operating pressures, are [8]: ~r

-

0.016 + 0.004 g coke g cat. h

tr 2 = 0.82 + 0.1 g coke psi g cat. h

(11)

The present model, although based solely on the MCP data, correctly predicts the coking behavior of C5N other than MCP [8]. When PH < 50 psi, as in most microbalance runs, the asymptotic kinetics ro ~ ~:(1/Pn)2 is quite accurate with a pseudo rate "constant" i r tclPH + Ir 2 . The temperature response of ~r shows an "activation energy" of 39.7 kcal/mole. Figure 8 shows that the model can predict the coke profile generated from nC7

635 reforming in the fixed-bed reactor under conditions quite different from those in the microbalance. The predicted coke profile is obtained based on the measured concentrations of C5N at the four sampling points along the reactor. This is of practical significance, since coke profile is an initial condition for modeling and control of the catalyst regeneration process. (9 L,. :3 r

CD L.

n

r 1::

Measured

.

Predicted ~ Co

t

a

.=,..

n

z

to

0

"0

.e-

CD ..~ O

O

6

2's

io Bed %

Figure 8. Measured and predicted (line) reactor (Pt-Re/Al203) coke and MCP profiles; t = 75 h

5. R E F O R M I N G N E T W O R K AND KINETICS Under the conditions used, the catalyst essentially stabilized its activity within 40 hours. Although the catalyst continued to change after the lineout period, its rate of change became much slower than that of reforming reactions. The product concentrations at the twelve spacetimes used in the development of the reforming kinetic model are from 3 separate experiments (each experiment provides 4 space time simultaneously) run under the same reaction conditions and at about the same on-stream time. We thus carried out kinetic experiments without considering catalyst deactivation after the lineout period. For the present system, experiments that only feed nC7 are not enough for developing a robust kinetic model. We thus use a parameter estimation procedure that requires multiple experiments, each of whose design makes its result sensitive to a particular parameter or to a small subset of the parameters. The major products of nC7 reforming are toluene, iC7 and cracking products (C~-C6). As justified elsewhere [15], the kinetics R, are all pseudo first order. Due to C5N's dominant contribution to coke formation, we segregate out all of the C5N as a separate lump, despite its relatively low concentration. We perform specific experiments that target individual reactions between the basic lumped species. Each experiment attempts to ascertain if a certain reaction occurs at a relevant rate and if so, what its rate constant is. Only after doing a number of such investigations, do we put together a lumped reaction network (Fig. 9) and examine the integral data for nC7 feed to determine the remaining parameters. In so doing, the number of truly

636 adjustable parameters reduces from seven to two. Figure 9 shows the reactions nC7 --> iC7 and are reversible. The rate constants determined are consistent with the thermodynamic equilibrium constants estimated from two independent sources [ 15]. nC7 ---> C5N

nC7

kl

=

C5N k7

it7

Tol

Figure 9. n-Heptane reforming network over Pt-Re/A1203 In contrast, a literature model [16] contains 14 reactions, since it allows almost all possible (reversible) reactions between lumps. By feeding toluene and hydrogen to the microbalance, we find that toluene is virtually unreactive, indicating that there are no reverse reactions from toluene to nC7 and iC7. Consider another literature model in which there is neither direct nC7 dehydrocyclization to toluene (k2 - 0) nor direct nC7 cracking [5]. Using this literature model, we can find a set of k's that would fit our fixed-bed data over a limited range of conditions. However, the new k6 and k7 do not even come close to describing the ECP-feed differential microbalance reactor data. The model derived from the above procedure compares quite favorably with experiment, see Fig. 10. We are now ready to combine the model with the coking kinetic model to predict the observed changes in the reforming rate constants during the lifetime of the catalyst.

1.0

'I

I

I

I

'

I

aCT. k'

II

: C.SN

0.8 t/)

.o_ to tO

0.6

iC7

Tol

t.,.

LL (9 ,,.....

iC 7

0.4

C1-C s

o

Tol

0.2 C5N

0.0

0'.0

" .

'

' 0.1

9

0

'.2

,,-,-) 9

'

013

9

I

0.4

'

9

'5

O.

S p a c e T i m e x, g cat-hr/g feed Figure 10. Model prediction vs. experimental data of Pt-Re/A1203 at 750~ and 517 kPa.

637 5.1. T i m e evolution of reforming rate constants

If each of the catalytic reactions involves only one active site, the random coke deposition model gives the following activity-time relation for nC7 reforming

k,o

k~ (t) = k~oS =

(12)

(1 + C~rot )

Here kio (g feed/g cat/h) are the rate constants determined at time zero t o which is the time required for the initial transients to disappear (t o ~ 40 hours). With the already-determined kio , 0% and ro, Eq.(12) gives k~ at any on-stream time after t o. Figure 11 shows excellent agreement between the predicted and measured time evolutions of the rate constants k; at 517 kPa, 750~ and H2/nC7 = 3 [ 15], thus permitting the prediction of the time evolution of the product compositions. Moreover, the assumption that the reforming reactions and coking have the same ~) is reasonable. Although not shown here, with k~o , Eq. (8), and Eq.(12) we have been able to correctly predict the reactor coke profile based on model-predicted C5N concentration profile. This is detailed elsewhere [ 15]. We thus complete the circle by providing an overall kinetics package that can predict not only the time evolution of the reforming products but also the coke profile along the catalyst bed.

12L_

.c

~n t-"

10

8

o

6

(9 r n"

4

C

0 r

nr"

2-

m

I

0

'

"

I

50

'

I

1 oo

On-Stream Time

"

I

150

t,

I

200

hr

Figure 11. Reforming rate constant changes vs. time at 750~ and 517 kPa (Pt-Re/Al203).

638 6. CONCLUDING REMARKS Our purpose here has been to propose a simple experimental protocol for studying the coupling between catalyst coking and the main reaction. This protocol is quite useful from the standpoint of practical applications. We illustrate its use by giving a fairly complete, quantitative picture of n-heptane reforming and coking over a commercial Pt-Re/AI203 catalyst. The present protocol should find applications in many other areas in both petroleum refining and chemicals manufacturing. ACKNOWLEDGMENT DSR would like to thank the National Science Foundation for support under PYI grant #CTS-8658147.

REFERENCES

1. S.C.Fung, C. A. Querini, K. Liu, D. C. Rumschitzki, and T. C. Ho, Catalyst Deactivation, Studies in Surface Science and Catalysis, B. Delman and G. F. Froment, Eds., Elsevier, 1994. 2. K.Liu, S. C. Fung, T. C. Ho, and D. R. Rumschitzki, J. Catal., in press, 1997. 3. C.A.Querini and S. C. Fung, J. Catal., 141 (1993) 389. 4. S.C.Fung and C. A. Querini, J. Catal. 138 (1992) 240. 5. P.A.Trimpont, G. B.VanMarin and G. F. Froment, Ind. Eng. Chem. Res. 27 (1988) 51. 6. S.C.Fung and C.A. Querini, C. J. McCoy in Catalyst Deactivation, C. Bartholomew and J. Butt, Eds., Elsevier, 1991. 7. u and G. M. Panchenkov, Kinetika i Kataliz, 21 (1980) 776. 8. K.Liu, S. C. Fung, T. C. Ho, and D. R. Rumschitzki, Ind. Eng. Chem. Res. accepted for publication, (1997). 9. W.Feller, An Introduction to Probability Theory and Its Applications, vol. 1, 2nd edition, John Wiley & Sons, NY, 1957. 10. G.A.Myers, W. H. Lang and P. B. Weisz, Ind. Eng. Chem., 53 (1961) 299. 11. J.N.Beltramini, T. J. Wessel, R. Datta, AIChE J., 37 (1991) 845. 12. N.S.Figoli, J. N. Beltramini, E.E. Martinelli, M. R. Sad, and J. M. Parera, Appl. Catal., 5 (1983) 19. 13. J.Barbier, P. Marecot, N. Martin, L. Elassal, and R. Maurel, in Catalyst Deactivation; Delmon and Froment, Eds.; Elsevier: Amsterdam, pp 53-62, 1980. 14. M.Guisnet and P. Magnoux, Appl. Catal., 54 (1989) 1 15. K.Liu, S. C. Fung, T. C. Ho, and D. R. Rumschitzki, in preparation, (1997). 16. K.R.Clem, Ph.D. Dissertation, Louisiana State University, 1977.

9

Elsevier Science B.V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

639

C h a r a c t e r i z i n g the F r a m e w o r k D e m e t a l l a t i o n o f E n v i r o n m e n t a l l y R e l e v a n t Zeolites U s i n g IR, N M R and N e u t r o n Diffraction T e c h n i q u e s Mark T. Paffett, Jfinos Szanyil, Richard M. Jacubinas, Kevin C. Ott, Robert VonDreele, Craig D. Hughes 2, and William L. Earl Los Alamos National Laboratory Los Alamos, NM 87545 Results of characterization studies monitoring the sequential chemical bond breaking events, local site symmetry, and long range structural modifications of specific zeolites (HZSM-5, TS-1) during hydrothermal treatment of these catalyst materials are described. These characterization techniques include infrared spectroscopy of selected probe molecules, magic angle spinning NMR spectroscopy, and powder neutron diffraction. Information regarding selected examples from each of these techniques is presented and the inherent strengths of each is discussed. The experimental insight into the chemical and structural modifications of high surface area microporous catalyst materials as a function of deactivation conditions (hydrothermal conditioning) is highlighted. 1. INTRODUCTION Extensive efforts have been expended on understanding and modifying the stability of zeolite catalysts to high temperature water vapor. Deactivation of zeolite based catalysts is particularly important in novel automotive catalysts which have remarkable catalytic activity in the selective reduction of NOx from emission sources [ 1,2]. However, high temperatures and water vapor are implicit in automotive exhaust and result in rapid catalyst deactivation. The details of deactivation and methods of circumventing this problem remain a technical challenge. Several mechanistic descriptions have been put forth describing the global deactivation of metal loaded zeolites. Loss of Brtnsted sites contributes to loss of ion exchange capacity, increased mobility of entrained metal ions, sintering of metal ions into metal and metal oxide particles, and in extreme conditions loss of framework structural integrity [3,4]. Furthermore, the reversible transformation of other framework metal ions such as Ti and V in a steam atmosphere has generated a considerable debate regarding the active state of such species as catalysts of the selective oxidation of alkanes and alkenes [5,6]. In this work we examine the role of water vapor in the hydrothermal dealumination and coordination chemistry of the non-siliceous atoms of several zeolites with the MFI structure (e.g., ZSM-5, TS-1) through the use of conventional and non-conventional techniques that include: infrared spectroscopy of adsorbed probe molecules, NMR spectroscopy, and neutron powder diffraction techniques. The purpose of these studies is to exploit the strengths of each technique to arrive at an understanding of the chemistry of water molecules acting on the framework substituted non-siliceous atoms. Central to this concept is an understanding what factors control reversible adsorption (the suggested coordination of water to Ti centers in titanium silicalite-1 [5]) versus bond breaking activity initiated by water molecules in the alumino-silicate zeolite materials. Understanding the 1 2

Currentaddress: PPG Industries, Pittsburgh, PA. Current Address: Morton Thiokol Corporation, Salt Lake City, Utah.

640 thermodynamic and kinetic factors underlying the reactivity and deactivation of these materials will certainly aid in the development of novel higher temperature materials more resistant to extreme environments. Furthermore, as part of this effort, we describe and demonstrate the first use of a novel in situ neutron diffraction reactor assembly in examining the structural aspects of steam modification of titanium silicalite-1 (TS-1).

2. EXPERIMENTAL Steamed H-ZSM-5 samples were prepared in a shallow bed flow reactor using an H20/N2 gas mixture with 7 Torr H20 partial pressure at temperatures ranging from 400-850oC. The samples were further characterized by X-ray powder diffraction (XRD) and BET surface area measurements. The BET measurements, carried out using nitrogen as the adsorbent, showed no significant change in the surface area as a result of the steaming conditions used in this study. The XRD analysis of the fresh and steamed samples provided no evidence for the collapse of the zeolite framework upon low temperature hydrothermal treatment and only a small extent of structural degradation for the moderate temperature steamed samples with the longest steaming time. TS-1 was prepared according to literature [7] hydrothermal procedures in an autoclave at 175oC. The starting material was characterized by conventional XRD, and diffuse reflectance UV-Vis spectroscopy. Elemental analysis of the H-ZSM-5 and TS-1 was performed using atomic absorption spectroscopy and x-ray fluorescence. The Ti/Si ratio was nominally 0.033 and nonframework Ti was determined to be below our detection limit of 0.002. Infrared experiments were performed as described in Ref. 8 using a Mattson Cygnus 100 FTIR spectrometer operated at a typical frequency resolution of either 2 or 4 cm -1. The 27A1 magic angle spinning (MAS) NMR analyses of aluminum were performed using a Varian Unity 400 NMR spectrometer equipped with a Varian 5 mm MAS probe. In-situ neutron diffraction of zeolite samples was accomplished using a reactor cell specifically designed to fit inside the High Intensity Powder Diffractometer (HIPD) beam line at the Manuel Lujan Neutron Scattering Center at LANL. A more detailed description of the overall time-of-flight neutron diffraction specific intrumentation and technique is contained in Ref. 9. The time averaged neutron flux on sample is ca. 107n/s/cm 2 and is derived from a pulsed proton beam striking a target normally operating at ca. 70 microamps. A diffraction pattern from all four detector banks on a zeolite sample required approximately 4 hours to acquire. The essential elements of the in situ reactor cell are shown in Fig. 1. The in situ reactor is a conventional plug flow reactor using a thin wall vanadium metal reactor tube (wall thickness 0.015", o.d. 0.375") welded to conventional stainless steel fittings (VCR4 or swagelock). A porous stainless steel frit is captured at the interface between the V and stainless steel and provides a convenient means of positioning the powdered sample in the neutron beam path. Alternatively, a quartz wool plug could be inserted in the bottom of the flow through tube. The V tube reactor was positioned inside an annular boron nitride (hBN) furnace block capable of temperatures to 800~ The BN furnace block was radially ported to allow neutron beam irradiance, neutron diffraction detector access at + 90 ~ + 40 ~ , and a neutron beam cut out in the back of the reactor (and support rod) to decrease secondary scattering events. In addition, the BN furnace block was drilled to accommodate four 1000W cartridge heaters and six themocouple wells. The BN furnace block and V tube reactor rests on a Cu pedestal ported for water cooling to reduce the temperature following high temperature excursions. The entire experimental assembly is suspended on a stainless steel post attached to an ASA 11 flange and is inserted in the vacuum well of the HIPD. The neutron detector elements are entirely contained within this vacuum well. All electrical, gas and water connections were made through KF hardware fittings and in operation vacuum integrity to 10-5 Torr was achieved. An additional benefit of the BN furnace block is an extremely low background neutron scattering signal arising from secondary scattering events that are screened by the BN because of the very large absorption cross section of ~~ The technical advantages of the in situ neutron powder diffraction reactor include examining the structural changes occurring in the catalyst under

641 reaction conditions without the necessity of exposing a potentially air sensitive catalyst to atmosphere. In addition, the active state (structure and phase composition) of the catalyst may be substantially different under reactive conditions than in an ex-situ non-reactive environment.

--1

ASA 11 flange

Support Post for Furnace Block

hBN furnace block (cartridge heaters inserted in block)

Fig. 1: In Situ Neutron Diffraction Cell; Functional items described in text. Gas, electrical power, and thermocouple feedthrus reside in KF flanges atop the ASA11 flange and are not shown for visual clarity. 3. RESULTS AND DISCUSSION In-situ FTIR of CO (and other molecules as site specific probes) was studied on the Hform and hydrothermally treated samples of H-ZSM-5. These results have been more thoroughly reported elsewhere [ 10] but are summarized here for comparison with experimental data from other techniques. Zeolites have been traditionally examined using infrared spectroscopy of N-containing adsorbates such as ammonia and pyridine to assess Br6nsted acidity. The use of weaker Lewis bases allows a more discriminating approach to assessing the strength and quantity of Brtinsted sites arising from the partial and sequential dealumination of framework A1 in H-ZSM-5. The spectrum of H-ZSM-5 in the O-H stretching region in the absence of CO displays three absorption features with widely differing intensities. The most intense peak, centered at 3614 cm -l corresponds to the O-H stretching of the bridging Si(OH)A1 groups [the Brtinsted acid sites of the protonic form of H-ZSM-5 structure]. The sharp feature at 3747 cm -I arises from the O-H stretching of the terminal SiOH groups. Its intensity can qualitatively be used as a measure of the continuity of the zeolite framework. The least intense peak is present only as a shoulder on the high frequency tail of the bridging Si(OH)A1 feature and has a frequency of 3665 cm -l. This peak is assigned [13, 14] to A1OH species originating from a certain extent of hydrothermal degradation during the sample calcination process. Upon CO adsorption at low temperature (T <150K) there are dramatic changes in the O-H stretching region of the IR spectra of H-ZSM-5. The 3614 cm -I peak completely disappears and a new, very broad, intense feature develops with 3310 cm -~ frequency at its maximum intensity. This feature arises from the bridging Si(OH)A1 groups interacting with CO molecules. The integrated intensity of this peak is several times higher than that of the corresponding OH groups in the absence of CO. This is a direct consequence of the higher polarization of the O-H bond in the presence of CO than in the absence of CO. There is no change in either the peak position or the integrated

642 intensity of either of the other two high frequency O-H features. These observations imply that under the experimental conditions applied (T>-120~ PCO < 200 Torr) CO does not adsorb at either the silanol groups or the OH groups of extra-framework AI species, while all the bridging Si(OH)A1 groups interacted with CO. Increasing the sample temperature from -120~ to -105~ results in the desorption of most of the weakly held, physisorbed CO while only a very small amount of CO desorbs from the zeolitic proton sites. Further sample temperature increases result in the complete loss of the feature attributed to physisorbed CO and the gradual decrease of the proton-held CO feature at 2173 cm -~. Concomitantly, the IR features of the zeolite OH groups interacting with CO decrease while the intensity of the CO-free bridging Si(OH)A1 feature increases. There is no measurable change in the intensity of either the 2230 or the 2195 cm -~ peaks (CO bound at extra framework species) and the intensities of the O-H stretches of the silanol remained unchanged through the experiments. The effect of steaming on the H-form of ZSM-5 was examined following two steaming temperatures (600 and 850~ Steaming at 600~ resulted in the appearance of two new, intense IR absorption features in the O-H vibrational region at 3665 and 3785 cm -~. The relative intensity of the very high frequency O-H vibration increases with steaming time while that of the bridging Si(OH)A1 hydroxyls decreases. This very high frequency O-H vibration has been observed on A1203 by KniSzinger and Ratnasamy [13] and was attributed to OH groups coordinatively bonded to a single A1 cation (-A1-OH). The hydroxyl groups represented by the IR feature at 3665 cm -~ are also products of the hydrothermal treatment and are suggested [ 14] to be AI-OH groups where the A1 atoms have not been completely removed from their zeolite framework positions but where the -Si-O-A1- bridge has been partially hydrolyzed: -Si-O-A1-+ H20 ..... >-Si-OH +HO-A1It is also possible that the 3785 and 3665 cm -1 peaks are correlated and they belong to two different kind of hydroxyl groups located on the same Al-containing species such as the A12(OH)6 dimer shown below. This species is entrained within the pores of the host

OH

\A1/ OH\A1/

/

OH

~OH

/

OH

~X

OH

zeolite and recent quantum mechanical calculations have strongly suggested that this is a likely species accounting for the observed IR data [ 15]. Note that the existence of this A1 species is significantly different from the partially coordinated A1-OH species normally suggested in hydrothermal demetallation of framework A1 [14,16]. For this chemical entity the very high frequency absorption feature would correspond to the terminally bound OH groups while the 3665 cm -1 feature would represent bridging OH groups. High temperature steam treatment of the H-form of the ZSM-5 for 2, 4, and 8 hours at 600~ gradually decreased the spectral signature of the bridging Brtinsted hydroxyls and resulted in the growth of features corresponding to extraframework alumina and/or A1-OH species which are formed during the hydrolysis of the Si-O-A1 bridge but are still in the zeolite framework. The apparent heat of adsorption of CO was measured in the steamed H-ZSM-5 materials and determined to be approximately 4 kjoule mol -l lower than that for the parent H-ZSM-5 zeolite (32.2 kjoule mol~). The heat of adsorption of CO was noted to be less sensitive to changes (or losses) in the

643

framework A1 atoms than the spectral lineshape changes. The extra lattice aluminum is also shown experimentally to transform into stable forms of Al-oxide species under specific hydrothermal conditions as indicated from spectral lineshape shifts. These Al-oxides present in extra-framework positions can also act as Lewis acid centers for the adsorption of CO. Zeolites are amenable to extensive characterization by high resolution NMR techniques. In principle, 29Si NMR can give information about framework aluminum and its siting because there are separate peaks for silicon with zero, one, two or more aluminums as next nearest neighbors. Since the MFI structure has 12 or 24 different T sites (depending on the crystalline polymorph) the 29Si MAS NMR spectrum is too broad to obtain quantitative information on framework aluminum. On the other hand, high resolution 27A1 NMR can distinguish between tetrahedral and octahedral aluminum and has been used to characterize dealumination in acid catalysts [ 17]. However, this analysis is not entirely straightforward. If the aluminum lsite~ is very asymmetric there is extensive second order quadrupolar broadening of the resonance which can cause the signal to vanish. In fact, very dry, H-form zeolites often have a loss in the 27A1 NMR intensity because of distortions about the tetrahedral aluminum site. We have developed sample preparation and NMR procedures that allow us to circumvent these difficulties. The samples are very carefully hydrated at room temperature prior to NMR analysis. The octahedral (non-framework) aluminum is comprised of a heterogeneous set of aluminum oxide and hydroxide monomers and oligomers in the pores. These give very broad NMR signal(s) which cannot be normally integrated for practical analytical purposes. However, the tetrahedral (framework) aluminum gives a sharp symmetrical resonance near 55 ppm (Fig. 2). Sapphire

(17.1B ppm)

Tetrahedral Octmhedral

U~L! w I 200

. . . . aL. _ ~ -- v . -" ,

,

I

m I 1 O0

t ,

,

I

w I 0

,,

I

I

I

- 1 O0

J

,

,

w I -200

Figure 2. 27A1MAS spectrum of the H-form of ZSM-5 after partial dealumination. The broad signal intensity in the baseline is due to aluminum which has been removed from the structure.The sapphire and tetrahedral (framework) resonances are indicated on the figure. A calibrated sapphire chip, glued to the cap of our NMR rotor was used as an internal standard. The intensity of the framework aluminum resonance can be integrated relative to the sapphire resonance and the loss Of framework aluminum as a function of steam treatment can be followed. An estimate of the global energy of activation can be obtained from the loss of framework aluminum from samples steamed for different times and temperatures [ 18]. In addition to these conventional methods of examining catalyst deactivation we have examined the effects of hydrothermal conditioning (using D20) of TS-1 in the in situ neutron powder diffraction cell capable of flow studies described above. The TS-1 sample was steamed using D20 saturated He at progressively higher temperatures for successive periods of an hour

644 A

I

I

I

I

I

I

O c,,l al O

0

~o

I

I

I

I

I

I

0.0 0.2 D-spacing, A

0.4

0.6

0.8

1.0

I 0 0 0.2 D-spacing, A

J 0.4

I 0.6

I 0.8

I 1.0

I

I

I I

1.2

I

XIOE

1.4 1

XIOE

, 1.4 1

B

~ l

I 1.2

Fig. 3. A) Neutron powder diffraction pattern recorded from TS-1 following dehydration and B) following steaming with D20 at 400~ (deuterated solvents are used in place of proto species because of the exceptionally large incoherent cross section and large absorption cross section of H). For brevity we show data recorded after dehydrating the TS-1 sample under a steady flow of dry He and the neutron powder pattern recorded following a 400~ steaming operation (Fig. 3A and B, respectively). The neutron powder spectrum was recorded simultaneously at all four detector banks and for brevity we display the data recorded at the +40 ~ detector bank. The data are displayed as normalized scattering intensity versus d spacing. It is immediately apparent from inspection of the two powder patterns that essentially no structural changes are observed below 5A. Above 5,~, however, very significant changes are observed in the d-spacings of the TS-1 sample following steaming. These include the nearly complete disappearance of the multiple d spacing features centered about 6A and loss of the prominent strong reflections occurring at 10.1 and 11.3,i~. A complete description of the structural transformations and methods for determining structure accounting for these changes are beyond the scope of this article. A preliminary interpretation of the observed differences suggest that the D20 coordination to Ti centers within

645 TS-1 results in loss of site symmetry within the channel structure. More explicitly stated, substantial adsorption of D20 within the silicate framework (in addition to the framework Ti centers) upon hydrothermal treatment is indicated from the nearly complete loss of the higher dspacing reflections of the diffraction patterns shown in Fig. 3. The data display the intrinsic strength of the technique in that both long and short range structural integrity are probed by the technique. The technique does not however probe specific bond making or breaking events or local symmetry about the Ti atom in this simplistic rendering of the acquired data. More sophisticated structural refinements can, under specific conditions, enable that level of detail to be obtained [ 19]. 4. S U M M A R Y

The combination of complementary experimental techniques provides an invaluable understanding of the details of water coordination in several representative zeolite catalyst materials. Specifically, from infrared measurements of weakly coordinated Lewis bases a discrimination in partially dealuminated species can be readily observed in materials that have been progressively dealuminated. From an examination of both the spectral position, its appearance and shift as a result of coordination versus partial pressure of adsorbate, and the relative change in the integrated band intensity versus the ubiquitous silanol functionalities present in most commercial forms of zeolite a picture of the nature and distribution of the partially framework coordinated species can be determined. Coupled with the 27A1MASNMR measurements that specifically probe the distribution of specific site symmetries within the hydrothermally treated zeolite a detailed picture of chemical species developing within the framework can be corroborated. Finally, the changes in long range order have been demonstrated from a Ti containing structural congener, TS-1, following hydrothermal processing. The emerging picture is that the structural transformations resulting from hydrothermal treatments disrupt site symmetry and order within the long range channels of the zeolite framework and may have a significant impact in utilizing the restrictive site requirements thought to control the very selective catalytic chemistry occurring in such microporous materials.

REFERENCES ~

2. 3. 4. 5. 6. ,

8.

9. 10. 11. 12. 13. 14. 15. 16.

Shelef, M., Chem. Rev., 95 (1995) 209. Amiridis, M.D., Zhang, T., and Farrauto, R.J., Appl. Catal. B:Env., 10 (1996) 203. Kharas, K., Robota, H., and Datye, A., ACS Symposium Series, 552 (1994) 39. 1994 Monroe, D., DiMaggio, C., Beck, D., and Matekunas, F., SAE Paper 930737 (1993). Notari, B., Adv Catal., 41 (1995) 253. Bellusi, G., Carati, A., Clerici, M.G., Maddinelli, G., and Millini, R., J. Catal., 133 (1992) 220. Clerici, M.G., Bellusi, G., and Romano, U., J. Catal., 129 (1991) 159. Ballinger, T.H., Wong, J., and Yates, J.T., Jr., Langmuir, 8 (1992) 1676. VonDreele, R.B., Reviews in Minerology, 20 (1989) 333. Szanyi, J. and Paffett, M.T., Microporous Materials, 7 (1996) 201. Loeffler, E., Lohse, U., Peuker, Ch., Oehlmann, G., Kustov, L.M., Zholobenko, V.L., and Kazansky, V.B., Zeolites, 10 (1990), 266. Delaval, Y., Seloudoux, R., and deLara, E.C., J.Chem.Soc.Fara.Trans. 1,82 (1986) 365. Kn/Szinger, H., and Ratnasamy, P., Catal.Rev.-Sci.Eng., 17 (1978) 31. Mirsojew, I., Ernst, S., Weitkamp, J., and Knrzinger, H., Catal.Lett., 24 (1994) 235. Brand, H., Redondo, A., and Hay, P.J., J. Mol. Catal. A: Chem., in press, 1997. Zecchina, A., Bordiga, S., Spoto, G., Scarano, D., Petrini, G. Leofanti, G.,

646

17. 18. 19.

Padovan, M., Otero-Arean, C., J. Chem. Soc. Faraday Trans., 88 (1992) 2959. Englehardt, G., Stud. Surf. Sci. Catal., 58 (1991) 285. Campbell S., Bibby, D., Coddington, J., Howe, R., Meinhold, R., J. Catal., 161 (1996) 338. Bish, D.L., and VonDreele, R. B., Clays and Clay Minerals, 37 (1989) 289.

~ Elsevier Science B.V. All rights reserved Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

647

S t u d y o f the Deactivation o f an H Y Zeolite Pellet using 129Xe N M R S p e c t r o s c o p y and ~H N M R I m a g i n g T. Domeniconia, J.-L. Bonardet ~, M.-A. Springuel-Huet ~, J. Fraissard ~ and J.-M. Dereppe b aLaboratoire de Chimie des Surfaces URA CNRS 1428, case 196, Universit6 P. et M. Curie, 4 place Jussieu, tour 55, F-75252 Paris Cedex 05, France bLaboratoire de Chimie-Physique et de Cristallographie, Universit6 Catholique de Louvain, Louvain-la-Neuve, B- 1348, Belgium Combined use of 129Xe~ and IH-imaging NMR shows that coking of an HY zeolite pellet results in a heterogeneous distribution both at the nanometric and macroscopic scales. This distribution depends upon the direction of the flow of reactant. It was also observed how diffusion of dimethylpentane occurs through both uncoked and coked sections of the pellet.

1. INTRODUCTION The obstruction of micropores or the blocking of the active sites of a catalyst by coke formation during a chemical transformation is one of the major problems encountered in zeolite deactivation. Moreover, catalysts are usually used in extruded form and the problem then is to know which part of this extrudate is active. For this reason we have undertaken the study of the deactivation of a Y zeolite in the form of a model pellet, firstly by the NMR of adsorbed 129Xe and then by 1H NMR imaging. The 129Xe nucleus has spin 89 its natural abundance of 25% makes it easy to detect and its voluminous electron cloud is very polarizable. This is the reason why Fraissard and Ito had the idea at the beginning ofthe 80s [1] of using it as a probe for microporous systems. A first ~H NMR imaging study of the distribution of a coke deposit on an alumina pellet was performed by Cheah et al. in 1994 [2]. After saturating their sample with water they used its protons as the M R probe. The superposition of a magnetic field gradient in a given direction on the signal detection sequence makes it possible to obtain the profile of the resonant spins in this direction. In the case of coked samples, since it is impossible to study the ~3C spins directly (sensitivity too low), the image obtained is that of the distribution of the water protons, i.e. the "negative" of the image of the coke distribution. We have reused this technique with 2,3dimethylpentane (DMP) as the adsorbate; this alkane has 14 almost magnetically equivalent protons, given the signal width (several tens of Hz) and, moreover, its kinetic diameter (0.6 nm) is quite close to the diameter of the opening of the hexagonal windows of Y zeolite (0.78 nm). After coking, this must decrease, which allowed us to probe the diffusion process(es) of the DMP at ambient temperature. The use of ~H-imaging technique allowed us to complete, at the macroscopic scale, information obtained by the xenon NMR at the nanometer scale.

648 2. EXPERIMENTAL 2.1. Materials A NH4Y zeolite (LZ 54 from UOP) was used. The zeolite was compressed, at a pressure of 1.5 tonne/cm2 in the form of a cylinder 12 mm long and 6 mm in diameter. This pellet was heated under nitrogen flow up to 720 K, then coked by cracking n-heptane at the same temperature for periods between 2 and 10 hours to obtain various coke contents: 2.5, 7.5 and 10 % w/w. The coke content is measured by weighting the sample before and after coking. After coking and prior to DMP (Aldrich Chem. Co., purity 95%) or xenon (Air Liquide, purity > 99.95%) adsorption, the pellets were outgassed at 673 K under vacuum overnight. Gas adsorption was performed using a home-made volumetric apparatus at 300 K and gas equilibrium pressures between 0 and 1200 Torr. For the study of DMP diffusion by ~H NMR imaging we used liquid DMP in equilibrium with its vapor so that diffusion occurred at constant pressure (the saturation pressure of DMP being 70 Torr at 300 K). 2.2. Nuclear magnetic resonance ~29Xe NMR. Ito et al. have shown that in the case of a HY zeolite and in the absence of divalent or paramagnetic cations [3] the observed chemical shift 5 is the sum of two terms; one

6s is characteristic of Xe-wall interactions and, therefore, dependent on the pore dimension and on the intracrystalline diffusion of the xenon, and the other, ~Xe-Xe, depends on Xe-Xe interactions in the zeolite and is proportional to the xenon concentration nxe: = ~S + ~Xe-Xe = ~s + k nxe

(1)

where k is characteristic of the internal free volume. All the information concerning the filling or the blocking of the micropores by coke is contained in the variation of 8s and k with the amount of coke in the sample. 129Xe NMR spectra were recorded on a Bruker MSL 400 spectrometer operating at a frequency of 110.7 MHz and using a simple 90 ~ pulse sequence with a repetition time of 1 s. ~H NMR imaeine. The Hahn echo pulse sequence is shown in Figure 1. To compensate field inhomogeneity, the pulse n along y, after time x must be composite; it is followed and preceded by a n/2 pulse along the x axis of the frame. If it is assumed that the field gradient is applied in the z direction, which is also the symmetry axis of the pellet, the magnetic field applied to the protons is then: B(z) =B0 +g0.z

(2)

and the signal intensity for an echo time x and a given value of z is: S(x) = ~z ~x~yp(z)exp[2inv'(z)x]exp[-x/T2]dz

(3)

where p(z) is the spin density in the xy plane for a given z, v'(z) = ?g0z/2~ (? gyromagnetic ratio of the proton) and T2 is the transverse relaxation time of the protons in the xy plane. The 9signal intensity depends not only on the number of spins but also on the relaxation time T2; the

649 real profile is that extrapolated to zero echo time. The evolution of this profile with x makes it possible to determine S(0)z and T2 for a given z: S(x)z

=

S(0)z[exp(-2x/T2z)])

(4)

and therefore to deduce qualitatively the more or less polyaromatic (graphitic) character of the coke (the more graphitic, the shorter T2). Simultaneous application of the field gradient in x and/or y directions of the pellet should make it possible to obtain a 2D or 3D image of the coke distribution. ~H NMR imaging was performed on a Brt~er MSL 300 spectrometer operating at a frequency of 300.13 MHz.

Bo

(2)x

(2)X(X)Y (2)x

~kz

,'1 IVYv''x

k So

'x'l l~V " -- t

I

I_.,ml

,~

y

gz

1;'

Figure 1. Schematic representation of a 1H NMR imaging experiment.

3. RESULTS AND DISCUSSION 3.1. Xenon adsorption and 129XeNMR Xenon adsorption isotherms are given in Figure 2. At low equilibrium pressure there is an increase in the amount of adsorbed xenon due to stronger interaction with the coke as compared to that with the zeolite surface. This effect is all the more important when the coke content is high. Conversely, at high equilibrium pressures, there is a decrease in the amount of adsorbed xenon due to a decrease in the pore volume by coke deposit and partial pore blocking. For the lightly coked pellet (2.5% w/w) the NMR spectra show two lines (Figure-3-a) whose chemical shift variations depend linearly on the xenon pressure Pxe (Figure 4). The fact that there are two distinct signals shows that on the NMR time-scale the Xe atoms do not exchange between two types of adsorption sites. The variation of the chemical shift of the high field signal (signal _a, lowest 8) is practically the same as that of the reference sample (noncoked pellet) which proves that this signal is due to xenon adsorbed in non-coked (or at least very little coked) regions of the pellet. The low field signal (signal b, highest 8) comes from

650 xenon adsorbed in zeolite crystallites containing coke deposit. In this latter case the slope d~5/dPx~ (nx~ is proportional to Px~), is identical to that of signal a but the straight line is shifted downfield by about 12 ppm.

1,2E+21 1E+21

01

8E+20

"~x6E+20 C 4E+20 2E+20

0

200

400

600

800

1000

1200

1400

Pxe ( t o r r s )

Figure 2. Xenon adsorption isotherms of HY pellets coked at various rates: (x) 0%, (•) 2.5%, (a) 7.5%, (O) 10%.

b a

l

I

200

100

I p pm

0

Figure 3.129XeNMR spectra of xenon adsorbed in a 2.5% coked nv pellet; the n-heptane flow coming from the bottom of the pellet (a, Pxe--200 Torr) and from the top (b, Px~=1000 Torr). These results are consistent with those of Barrage et al. [4] for lightly coked HY powder catalysts, which show that the coke is formed initially at the hexagonal windows of the acages, reducing the diffusion of Xe (;5,'1') without affecting the pore volume (d~5/dnxe= const.). Moreover, signal b is much broader than signal a; this is certainly due to a distribution of coke deposits in the pellet. Indeed, the colour of the pellet is seen to vary along its z axis, the lower

651 part being darker than the top when the flow was from the bottom through the frit. In order to obtain more homogeneous coking, the gas flow was reversed. In this case there are still two NMR lines but the width of signal b is much smaller (Figure 3). For a set of samples with different coke levels (2.5, 7.5, 10% w/w), we always observe two signals, but the intensity of signal a decreases when the coke loading increases, while the chemical shift of signal b increases. The 8 = f(nxe) variations are seen in Figure 5. For coke loadings up to 7.5 % the slopes of the curves do not change when 8s increases. This is consistent with coke deposit inthe openings of the supercages; this hinders diffusion of xenon but does not affect the micropore volume very much. For the more coked sample (10%), ~Ssis the same as for the 7.5 % sample but the slope is higher; this is in agreement with the results of Barrage et al. [5] and reveals that part of the zeolite is no longer accessible to xenon. This can be explained by a partial blocking of the zeolite pores, the part still accessible to xenon remaining practically unchanged.

120

110 100 [i ga,

90 80

6o! 70i-

r

50 0

200

r

9

400 600 Px. (torrs)

9

800

1000

Figure 4. 129XeNMR chemical shi~s: (A) non-coked pellet, ([3) a and ( I ) b signals of a 2.5% coked HY pellet. 3.2.

1H N M R

imaging

Figure 6-a displays the distribution profile of the spins of the proton probe along the vertical axis z (1D imaging) of a pellet containing 7.5% of coke and saturated with DMP (9.10 20 molecules.g-'). We observe first of all that the signal intensity and therefore the shape of the profile depends markedly on the echo time x imposed on the pulse sequence. If this is high there is considerable loss of information, and the coke profile (the negative of that of DMP) appears to be homogeneous. This result is spurious and shows that it is the short relaxation time T2 which controls the information. In fact, as the echo time is decreased the signal intensity increases, regardless of z, but for a given x there is a practically monotonic increase in the signal, from the begining of the pellet (z0) to the end (zf) (the discontinuities observed can be attributed to structural heterogeneities due to pelleting). Determination of the real profile by

652 extrapolation to zero echo time shows that it is homothetic with that observed for the shortest time x applied (x = 576~ts) compatible with the equipment. These results indicate that the coke distribution at the macroscopic level is very heterogeneous and that the most active part in the reaction is that which is first exposed to the reagent flow. 130 120 110

100 g~

~, 90

r

80 70 60 50 0

5E+20

1E+21

1,5E+21

n (XO/g

Figure 5. 129XeNMR chemical shifts ofthe b signal" (O) 2.5 %, (D) 7.5% and (~)10% coked HY pellet; (X) non-coked pellet Measurement of T2 from the slope of the linear transform of equation 4, shows that it is very short (225
Lll--

i

250

150

2001-

r

= o

't'',~

a !_

i

Zf

<

z ~ Z o

Zo -

....Z

~- Zf

Figure 6. Spin distribution along the z axis of 7.5% (a) and 10% (b) coked HY pellets Figure 6-b shows the profile of the DMP probe for a 10% coked pellet and saturated with adsorbate (7.1020 molecules.g-l). As previously, it is observed that when x decreases the signal intensity increases but, in contrast to the previous sample, and down to the lowest echo time (x

653 = 576~ts) the profile obtained along the axis of the pellet shows a certain homogeneity of the distribution of the resonant spins and therefore of the coke. The profile extrapolated to zero echo time displays marked discontinuities attributable partly to structural heterogeneities but also to the imprecision of the measurements. The constant but very low relaxation time, T2, (2201as) and the lower concentration of spins lead to a greater experimental error. Nevertheless, these results show: - that there is always a coke concentration gradient in the pellet but that it is lower than in a less coked sample; - that the quality of the coke, which is very graphitic (T2 very short), is chemically homogeneous (T2 constant). These conclusions confirm those of several authors on highly coked powder catalysts [4,6]. Finally, Figure 7 shows an attempt to obtiain a 2D image of a lightly coked sample (2.5%). The transversal sections (xy plane of the pellet) for different values of z show a very heterogeneous distribution of the coke from z0 to Ztotal/2. The coke is located essentially on the face attacked by the reagent flow and on the external surface of the pellet. This image is in agreement with the observation of two signals for adsorbed 129Xe.

Figure 7. 2D (xy) spin profiles for differem z values (z0 (top left) to z0/2 (bottom right)). 3.3.1H NMR

diffusion

To complete this study we performed a dynamic experiment on the diffusion of DMP in the 10% coked pellet. It should be recalled that the kinetic diameter of the probe is 0.6 nm; after heavy coking it is reasonable to think that the diameter of the hexagonal openings (0.78 nm) of the Y zeolite is seriously reduced, leading to a decrease in the rate of diffusion of the molecule in the pellet, it being possible then to study this diffusion by 1-D NMR imaging. At time t - 0 the pellet is placed in contact with the DMP vapour at its saturation vapour pressure (70 mm Hg) and the spin profile along z is recorded for a constant echo time, at different times (1 min < t < 150 min). The shapes of the profiles are displayed in Figure 8. It can be seen that from t = 1 min to t = 30 min, the signal intensity increases, the maximum moving from the edge of the pellet towards the centre. Up to 70 min the maximum continues to shift towards the lower face of the pellet; finally, from t = 70 min to 100 min, the signal intensity decreases over the whole pellet and becomes steady for times greater than 100 min.

654 This evolution shows that there are two competing diffusion processes; one concerns intercrystalline diffusion corresponding to the movement of the adsorbate front from the outside to the inside of the sample, the other, slower process being attributable to intracrystalline diffusion of the probe from uncoked or lightly coked zones towards heavily coked zones which act as strong physical adsorption centres. A mathematical model is being developed to quantify this phenomenon and to measure the inter- and intracrystaline diffusion coefficients. |

9" = e s t .

Z

tm

dSO _100 _ 70

-50 - 30 - 20

.10 '9' '

.......

| ...............................

I'-

Zo

Zf

"1

........

Figure 8. Evolution with time (in minutes) of the spin profiles during the DMP adsorption.

4. CONCLUSIONS

129XeNMR is particularly helpful to study the coking of HY zeolite pellets. The presence of two signals in the spectra clearly shows the heterogeneous distribution of coke at the nanometric scale. This heterogeneity depends on the direction of the flow of the reactant. IH-imaging NMR using the protons of adsorbed hydrocarbon molecules (DMP) has been used to determine the distribution of coke in the pellet. The microscopic heterogeneity observed by 129XeNMR has been seen also, at the macroscopic scale, by imaging NMR. The study of the di~sion during the adsorption of DMP reveals competing processes. REFERENCES 1. T. Ito and J. Fraissard in: Proceedings of the Fifth International Conference on Zeolites (Rees L.V.C. Ed.) p 510, London, Heyden (1980) 2. K.Y. Cheah, N.C. Chiaranussati, M.P. Hollewand and L.F. Gladden, Appl. Catal., 115 (1994) 147 3. T. Ito, J. Fraissard, J. Chem. Phys., 76 (1982) 5225 4. M.C. Barrage, J.L. Bonardet and J. Fraissard, Catal. Lett., 5 (1990) 143 5. J.L. Bonardet, M.C. Barrage and J. Fraissard, J. ofMol. Catal., 96 (1995) 123 6. R.H. Meinhold and D.M. Bibby, Zeolites, 10 (1990) 2532

9

Elsevier Science B.V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

655

A F M Study o f Carbon F o r m a t i o n on a M a n g a n e s e Oxide Catalyst N. Batina, L.M. Ioffe and Y.G. Borodko Universidad Autonoma Metropolitana-Iztapalapa, Department of Chemistry, A.P. 55-534, Mexico D.F., 09340, Mexico

Formation of carbon coke on a manganese oxide (MnOx) catalyst surface during the exposure to methane at 700~ was studied by means of atomic force microscopy (AFM) imaging. Images collected before and aider the reaction, showed significant changes in the catalyst surface morphology due to carbon coke formation. A carbonaceous deposit was successfully visualised at different stages of this formation. The early stages were in particular interesting in the case of the NaP-modified MnOx catalyst, which appears to be less sensitive towards deactivation due to carbon coke formation than the pure MnOx. Results from X-ray photoelectron spectroscopy (XPS) and Fourier transform infrared spectroscopy (FTIR) showed that the catalyst coke mainly consists of CHx deposits and manganese carbide species.

I. INTRODUCTION Sodium-phosphorous promoted manganese oxides were found to be effective oxidative coupling catalysts for converting methane to C2 hydrocarbons [ 1-3 ]. Manganese oxide usually operates in a redox mode, which uses lattice oxygen from metal-oxides to perform the oxidative coupling of methane in the absence of gaseous oxygen. This approach was realized by researchers at Union Carbide and ARCO [1,2]. In redox mode, manganese oxide is an oxidative reagent, so available lattice oxygen was being depleted during long time contact with methane. It was shown that atter 6-10 min exposure to methane (GHSV=600 hr1, 800~ the surface layer of manganese oxide was reduced and the reaction of oxidative coupling of methane was complete [2]. Although modified MnOx catalysts have already been recognized as active catalysts for oxidative coupling of methane, there is a lack of information concerning catalyst deactivation. So far, preliminary studies have shown that reduction of the surface Mn(III) ions to Mn(II) leading to oxygen depletion and poisoning by means of carbonization of the catalyst surface, can be pointed as a major cause of the catalyst deactivation [4]. Carbonization is a result of a thermal decomposition process (coking). The mechanism, elemental composition and structure of the carbonaceous deposit have been studied extensively on metal catalysts [5-7]. Recent development of scanning probe microscopy techniques (SPM), such as scanning tunneling microscopy (STM) and atomic force microscopy (AFM), have brought new possibilities to studies the surface morphology of the carbonaceous deposit. SPM techniques provide real time images of the real surface regardless of the surface order, and can operate in a variety of experimental conditions, e.g. high pressure, solutions, air and vacuum [7].

656 The aim of our study was a characterization of the surface morphology (structure) and transformation of the carbon coke on the pure and NaP-modified MnOx-catalysts. An AFM was used to monitor the progress of the carbonaceous deposit formation (coking) over the catalyst surface, from the early stages up to the later stages when the surface is completely covered by carbon coke. XPS and FTIR were employed to characterize the chemical composition of the carbonaceous deposit on the MnOx surface.

2. EXPERIMENTAL Pure and NaP-modified MnOx-catalysts were used in our study. Due to easy visualization by AFM, the MnOx layer was placed on a Si-wafer substrate (1 cm x 1 cm plate), by a reactive deposition technique. The sample preparation was carried out in a vacuum installation equipped with an resistance evaporator. Metallic manganese (99.8 %) as a source and a Si wafer with a surface orientation (111) and resistivity of 7.5 ohm/cm as support, were used. During MnOx deposition, an oxygen partial pressure of ca 10-4 torr, in dynamic mode, was maintained. Before used for the catalytic purpose, MnOx samples were calcined in air at 700~ for 60min. In order to prepare the NaP-modified catalyst, the MnOx samples were impregnated in a diluted Na4P207 solution (5 wt %), dried and finally calcined at 500 ~ C, in air during 30 min. The interaction with methane was performed in a quartz reactor in a methane atmosphere at 700 ~ + 5~ C. In order to monitor the kinetics of the carbonaceous deposit formation, exposure to the methane atmosphere (5s - 30min) and the pressure of the methane (70-350 Torr) in the reaction chamber were varied as well. A new set of samples was used for each experiment. After exposure to methane, the samples were cooled either in the methane atmosphere or in vacuum. The samples were analyzed by means of AFM imaging before and after reaction. The visualization was performed by an ex-situ AFM, ( NanoScope III, Digital Instr.) in air, at room temperature. All measurements were performed in contact mode, using silicon nitride probes, with a scan rate lower than 2 Hz. Images were collected at four different places within each sample. No destruction of the sample surface was noticed during and after AFM imaging. All images are presented in the topview-height mode. XPS and FTIR surface analysis of the NaP-modified MnOx catalyst (powder) were performed in conditions and by procedures described before elsewhere [3-4].

3. RESULTS AND DISCUSSION

3.1. XPS and FTIR analysis of NaP-modified MnOx catalyst. XPS and FTIR have been used for surface characterization in order to prove the formation of carbonaceous material on the NaP-modified MnOx catalyst surface. The results obtained offer clear evidence about formation of hydrocarbon deposits on the catalyst surface during the methane treatment. XPS showed a carbon line at 285 eV, whose intensity increased with a longer treatment in methane, this suggests that the amount of the carbon deposit increased. Intensity of Auger Na(KLL) line at ca. 264 eV was used as reference, since both ESCA and AFM studies showed that the surface of NaP particles remained uncovered by coke for a longer time than the manganese oxide surface The fine structure of the XPS line of C 1s

657

CHxAC 1 S

[A \

t

ul v

c

Na(KLL)

CxMn

z i

290

,

,

280

I

2"/0

' 260 ' '

290 280 Eb, eV

Figure 1. XPS spectra of carbon C ls and Auger Na(KLL), background corrected and deconvoluted, for NaP-MnOx catalyst before (A) and after (B), (C) exposure to methane (5 min., 800~ 100 torr of CH4). PHI-551 Spectrometer, Eb C Is, 285 eV, as a reference.

e.; z 60

!.

A

B

'~

'

I

3000

2000

1500

I

1000

500

,

Figure 2. FTIR spectra of NaP-MnOx after exposure to methane (100 torr) at 800~ min.; (B) for 30 min. Transmittance mode, %T. FTIR Spectrophotometer PE-1000.

t

f0.26

~

"

._..

(A) for 5

CH3

~

CH2

. o

3000

I

2950

I

2900

I

2850

Vicm "~

Figure 3. FTIR spectra in the range 2800-3000cm ]. Transmittance mode, absorbance unit. Typical positions of the C-H stretching vibrations for CHx fragments are shown above.

658 electrons for initially oxidized catalyst showed two lines with Eb at ca. 285 eV and 288 eV (Fig. 1A). They may be assigned to the surface hydrocarbon species (CHx) and carbonate ions respectively. After exposure to methane, the C 1s line becomes broader with a shoulder at the low energy side (Fig. 1B). The C ls line for the used catalyst was deconvoluted on three components: 285 (CHx), 288.5 (CO3) and 282.5 eV (Fig. 1C). The low-energy component may be assigned to a carbidic state of carbon, since this region of binding energy (281-283 eV) is typical for carbon in metal carbides [8]. After exposure to methane the IR spectra of manganese oxide showed absorption bands, which are characteristic of the C-H stretching vibrations (CH3: 2962, 2872 cm~; CH2: 2926, 2853 cm1 and CH: 2890 crn"1) [9]. The intensity of the C-H bands increased, if the exposure time to methane increased (Figs. 2,3). The intensity of the band at 1050 crn~, which is assigned to v3(Si-O) of silica was used as an internal reference (Fig. 2). the intensity ratio for the CH2 and CH3 groups estimated for samples after 30 min reaction with methane, was found to be about five [ 10]. Thus XPS and FTIR surface analysis showed that carbonaceous material formed on the MnOx catalyst surface consists of CHx hydrocarbon deposits and manganese carbide species.

3.2. AFM analysis of MnOx catalyst. A set of AFM images in Fig. 4, illustrates the surface morphology change observed for the MnOx catalyst. Before an exposure to the methane (Fig. 4a), the catalyst top surface layer appeared to be fiat. AFM revealed uniformly big (diameter 0.12 ~tm to 0.19 ~tm) clusters of the MnOx surface. Atter the reaction with methane, the surface morphology of the MnOx was found to be significantly different. Images in Figs. 4b and 4c, show a catalyst surface atter the methane treatment. At micro level the catalyst surface seems to be smooth, but individual clusters of MnOx could not be further resolved. Instead, the surface was found to be covered with fibril-like features (rows), which were ca. 1 lam elevated from the rest of the sample. The shape of these rows was found to be a function of the exposure to methane. Fig. 4b, shows the catalyst surface atter 15s, at 700 ~ C with 120 torr of partial pressure of methane. At similar condition (300 torr, 700 ~ C) but a~er 10 min reaction, a surface with a very characteristic and well-defined pattern was developed (Fig. 4c). AFM images revealed numerous rows (fibril-like features), which propagate almost parallel over the surface at a mutual distance of 1.4 ~tm to 1.57 ~tm. Most of the observed rows were about 3 ~tm to 4 ~tm long and 0.5 ~tm wide, with variation in the height from 0.89 l~m to 1.30 ~m. Two neighboring rows usually merge, and continue as a single row, atter 3 ~tm to 4 ~tm. Since the AFM images were obtained atter methane treatment, they actually show topography of the carbonaceous deposit on the catalyst surface. Image in Fig. 4d, shows few MnOx rows covered by carbonaceous material, with a small portion of the middle row only partially covered by the coke deposit. At the carbon coke free surface, an underlying MnOx substrate, organized in a row, with a typical cluster structure can be recognized. It clearly shows that observed rows appear because of a change in the MnOx substrate morphology. Most of the surface was found covered by a thick coke film (ca. 2 nm to 6 nm) in the form of individual clusters, which fused into a "flower" type pattern (agglomerates). Note the clear domain boundaries around each carbon "flower" pattern. Regarding the shape of the observed carbon features it could be concluded that during formation, carbon coke clusters possess some kind of limited mobility. These coke deposits possibly grow by spreading over the catalysts surface via a branching mechanism. Fig. 4e shows an AFM image recorded at the

659

Figure 4. AFM images of MnOx Catalyst surface before (a) and after exposure to CH4: (b) 15s, 700~ 120 torr; (c) 15 min., 700~ 300 torr. The structure of carbon coke deposit visualized by AFM upon deactivation (d) and additional reactivation (20 min., 500~ air) (e). Image size: (a, d) 6~tm x 6~tm; (b) 5~tm x 5~tm; (c) 15~tm x 15~tm and (e) 940~tm x 940~tm.

660 carbon coated MnOx sample after an additional 20 min treatment in air, at 500 ~ C. The treatment was applied in order to remove carbon coke from the catalyst surface and to reactivate the catalyst. AFM images revealed that the treatment was rather successful. Most of substrate clusters appear to be coke flee. Remains of the carbonaceous material could be recognized mainly around smaller MnOx clusters, in particular, around cluster edges. Viceversa, it indicates that formation of the coke started preferentially around the substrate cluster edges. Large scale AFM images (10 ~tm and more, not presented here) revealed that the samples did not return to its initial flat stage after cleaning in air and removal of the carbon coke. This confirms our assumption that observed fibril-like structures and significant perturbation in the catalyst surface morphology, are related to the MnOx layer perturbations. A thermal drift (different thermal expansion coefficients) between substrate layers and the carbon coke adlayer has been considered as the responsible mechanism for compressive stress failure [ 11 ]. Further work in this direction is in progress. 3.3. AFM analysis of NaP-modified MnOx catalyst. Interestingly, but not surprisingly, the NaP-modified MnOx system appeared to be in many ways different than the MnOx one. NaP-modified samples appeared to be less contaminated by carbon coke, or in another words less deactivated than pure MnOx samples, after similar treatment in CH4 atmosphere. This is in good agreement with previously reported findings concerning the higher activity of NaP-modified MnOx catalysts [2]. Fig. 5a shows an AFM image of the NaP-MnOx catalyst surface, before reaction with methane. In comparison with the surface morphology of the pure MnOx catalyst, the average particle size appears to be different (bigger). Two typical particle types with an average diameter of 0.08 ~tm and 0.283 ~tm were observed. The bigger particles, with rectangular shape were somehow segregated and expelled toward the top position on the sample surface. We suppose that such big clusters are crystals of the NaP-modifier which gradually diffused to the catalyst top layer. The AFM image in Fig. 5b, shows a surface of the NaP-MnOx sample after reaction with CH4 (15 min, 300 tort of CH4). Similar to the case of the carbon coated MnOx surface, we also found that the catalyst surface exhibits a fibril like structural pattern. However, the individual carbon clusters among the existing rows could be recognized easily even at the large size images. Finer details such as the average cluster size (ca. 018~tm) and shape of individual clusters can be recognized in the high resolution images in Figs. 5c and 5d. Regarding the image presented in Fig. 5d, it looks like the carbon clusters have a tendency to fuse into a more compact carbonaceous film. In order to visualize very early stages of the coke formation we exposed NaP-MnOx samples to CH4 for only 5s, at 700 ~ C with a partial pressure of methane of 10 tort. The result was surprising since the catalyst surface was found to be completely covered by carbon coke (see a high resolution image in Fig. 5e), consisting of clusters with individual size about 0.04 ~tm (diameter). This is almost twice smaller than the size of the smallest substrate clusters and leads to the conclusion that each substrate cluster possesses two or more active sites for the coke formation process. For samples treated in CH4 a bit longer (30s) at a CH4 pressure of 170 tort, further changes involving a fusing of carbon clusters into non-shaped agglomerates (Fig. 5f), were noticed. During this transformation, which accompanied by a significant material transport over the catalyst surface, some of the substrate particles become again coke free (see bright features randomly spaced in the dark carbon film).

661

Figure 5. AMF images of NaP-MnOx catalyst surface before (a) and alter different exposure to methane: 15 min., 700~ 300 torr (b, c, d); 5 s, 700~ 10 torr (e) and 30 s, 700~ 170 torr (f). Structure of the catalyst surface covered partially by carbon coke deposits visualized after cooling in the methane atmosphere (g, h). Image size: 6gm x 6gm (a); 17gm x 17gm (b); 3gm x 3~tm (c, h); 4gm x 4gm (d); 1.75gm x 1.75gm (e, f); 12gm x 12gm (g).

662 Interestingly, as could be seen in Fig. 5g, there was no growth of the carbonaceous material around the rectangular substrate particles which we tentatively associated with NaP crystals. The same animosity between big substrate particles and coke is also visible in the image in Fig. 5h. Since both images were recorded at the very late stages of the coke growth (15 min, 700 ~ 300 torr), the catalyst surface was expected to be completely covered by carbon coke (Figs. 5b to 5d). However due to a specific cooling procedure in the methane atmosphere, the catalyst surface finished being covered partially. This means that changes in the surface morphology continue beyond the methane treatment during the cooling phase.

4. CONCLUSIONS In summary, our study shows that the surface of the MnOx catalysts during exposure to methane at high temperature undergoes significant morphological changes that are mostly related to the formation of the carbonaceous film on the catalyst surface. The growth and characteristics of the carbonaceous film have been monitored and described from the early stages (partial coverage) up to a thick film formation. The structure of the carbon coke at low exposure to methane is determined by the appearance of small size aggregates, which latter fuse in to bigger aggregates. After 30 min. exposure to CH4, the catalyst surface was found completely covered by carbon coke film, showing a typical fibril like pattern. This is associated with differences in the thermal expansion between the catalyst substrate and the coke adlayer. MnOx and NaP-MnOx catalysts were found to differentiate during the early stages of the coke formation. The NaP-MnOx catalyst appeared to be less contaminated (deactivated) by carbon coke after the identical treatment in methane. The stability of a carbon coke deposit was also found to be sensitive towards the cooling procedure in the post-reaction period. Removal of carbon coke from the catalyst substrate by the air treatment at 500 ~ C was demonstrated. AFM should be considered as a very useful technique in the probing of deactivation and reactivation processes of solid catalysts. XPS and FTIR surface analysis has shown that the carbonaceous deposit formed on the MnOx catalyst surface consists of CHx coke and manganese carbide species.

ACKNOWLEDGEMENTS We thank Dr. J. G. Ibafiez from Depto. de Ingenieria y Ciencias Quimicas, Universidad Iberoamericana, Mexico, who allowed us to use the AFM facilities and Dr. Robin Sagar for helpful discussion. Financial assistance was received from CONACYT. Projects: 0913E-P; Chtedra Patrimonial de Excelencia and L0081 -E9608 for Dr. N. Batina, and 0232P-A9506 and 400200-5-0440PA for Dr.Y.G. Borodko.

REFERENCES 1.

2.

G.E.Keller and M.M. Bhasin, J. Catalysis, 73 (1982) 9. J.A.Sofranko, C.A. Jones, J.J. Leonard, J. Catalysis, 103 (1987) 311; 103 (1987) 302.

663 3. L.M.Ioffe and Y.G. Borodko, Catalysis Today, 13 (1992) 597; L.M. Ioffe, T. Lopez, Y.Borodko, R. Gomez, J. Mol. Catalysis, 98 (1995) 25. 4. L.M.Ioffe, P.Bosch, T.Viveros, H.Sanchez, Y.G.Borodko, Mater. Chem. Phys, presented 5. C.H.Bartholomew, Catalysis Review- Sci. & Eng., 24 (1982) 67. 6. Studies in Surface Science and Catalysis. Catalyst Deactivation 1991, C.H. Bartholomew and J.B.Butt., Eds., Vol 68, Elsevier, Amsterdam, 1991. 7. B.J.Mclntyre, M. Salmeron and G. A. Somorjai, J. Catalysis, 164 (1996) 184.; B.J. Mclntyre, M. Salmeron and G. A. Somorjai, Catal.Lett., 14 (1992) 263.; B.J. Mclntyre, M. Salmeron and G. A. Somorjai, Rev. Sci. Instrum., 64 (1993) 687. 8. L.Ramqvist, K. Hamrin, G. Johanson, U. Gelius and C. Nordling, J. Phys. Chem.Solids, 31 (1970) 2669 9. N.Sheppard, G. De la Cruz in Advances in Catalysis, Eds.: D. D. Eley et al., Vol 41, Acad. Press, San Diego, 1996. 10. D.C.Harris, M.D. Bertolucci, Symmetry and Spectroscopy, Dover Publ., N.Y., 1989. 11. D.L.Smith, Thin-film deposition, McGraw-Hill, N.Y., 1995. Physics of Thin Films, M.H. Francombe and J.L. Vossen, Eds., Acad. Press Inc., Boston, 1993.

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~

Elsevier Science B.V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

665

N o v e l R e g e n e r a t i o n M e t h o d for D e a c t i v a t e d N o b l e M e t a l Catalysts L.A. Camacho a, C. Park b and N.M. Rodriguez b, Department of Materials Science, The Pennsylvania State University, University Park, PA16802

a

b

Department of Chemistry, Hurtig Hall, Northeastern University, Boston, MA-02115

The deactivation and regeneration of supported noble metal catalysts is a major problem encountered in a number of commercial processes. These types of catalyst systems are highly susceptible to loss of activity due to the accumulation of carbonaceous residues on the metal particle surfaces that necessitates the periodic utilization of a high temperature "decoking" step that is generally performed in air. In the current investigation we have explored the potential of an alternative method to remove deposited carbonaceous species from a supported metal catalyst system that is based on a "surface cleaning" operation performed by adsorbed CO. The impetus for this study stems from the fact that there is a growing interest in the use of various forms of carbon as supporting media and the traditional methods of removing deposits from the catalyst would also result in the concomitant gasification of the support. As a consequence, we have been directed to search for regeneration methods that can be carried out at lower temperatures. The concept is predicated on the knowledge that the adsorption of CO on noble metals is relatively strong and such an action could possibly result in the preferential displacement of carbon species pre-adsorbed on the catalyst surface. The impact of this approach on the catalytic behavior of both carbon nanofiber and silica supported platinum, palladium and rhodium particles is discussed. 1. I N T R O D U C T I O N The deactivation and subsequent regeneration of noble metal catalysts continues to be a major problem in the petroleum, petrochemical and pharmaceutical industries. These catalysts are highly susceptible to deactivation due to the collection of carbon deposits on their surfaces during hydrocarbon conversion reactions resulting in costly regeneration procedures and prolonged plant shutdown (1,2). A great deal of research work has been devoted to the development of methods of minimizing the effects of carbon deposition on metal particles and in this regard, one of the more successful methods in combating this form of catalyst deactivation is the notion of introducing another metal into the catalyst system. It is well known for example, that the addition of Re, Sn or Ir, to Pt not only stabilizes the activity of the noble metal catalyst, but also exerts a significant influence on the coking characteristics of the catalyst during a hydrocarbon conversion reaction (3-6). At this point in time, however, there is no clear understanding as to the reasons why the addition of Re or Ir tends to diminish carbon deposition, whereas Ge and Sn appear to have the opposite effect on the level of carbon accumulation on Pt/A1203 catalysts (3,7-9). The addition of low levels of nonmetallic adatoms, such as sulfur and chlorine, have also been studied in these systems with a considerable degree of success. The time on-stream for a reaction being substantially increased when the catalyst is either pretreated or exposed to a constant concentration level of these adatoms. It should be recognized, however, that the beneficial effects of these

666 particular additives are only realized at a critical concentration level, above such a value they act as poisons and at very low concentrations they exert only a weak influence on the coking and selectivity of the catalyst for the reaction under investigation (10,11). Eventually after a prolonged period of time on stream even these systems require to be regenerated due to the build up of carbonaceous deposits. The type of carbon generated from the metal catalyzed decomposition of hydrocarbon containing reactant mixtures has for the most part, merely been considered an undesirable biproduct of the reaction. Indeed, fundamental studies aimed at the understanding of the key factors involved in the growth mechanism of such material have been driven by the need to develop effective treatments to inhibit its formation (12) In the last decade, however, since it has become evident that certain forms of filamentous carbon possess a set of unique chemical and physical properties that make the structures highly desirable for a number of applications. As a consequence, efforts are now being directed at optimizing the growth of carbon filaments and controlling the structural characteristics of the material using a variety of metal catalysts and hydrocarbon sources. Details of both the growth procedures and characterization studies of carbon nanofibers (13-15) have been given in recent publications. Previous studies from this laboratory have indicated that carbon nanofiber supported nickel particles can exhibit an entirely different activity selectivity behavior to that found when the same metal is dispersed on more traditional carriers and reacted under the same conditions (16-18). These aspects were particularly evident in the hydrogenation of hydrocarbons, including butene and butadiene. Impregnation of Ni onto carbon nanofibers resulted in significant changes in the catalytic behavior of the metal compared to its performance on either active carbon or 7-alumina. Certain types of carbon nanofibers have been found to be highly graphitic in nature and possess a set of unusual structural conformations. By careful choice of the catalyst and reactions conditions it is possible to produce such materials in a form where the graphite platelet component is aligned in a direction perpendicular to the fiber axis. In this arrangement only the prismatic faces are exposed and as such, present unique surface sites for the subsequent location of metal catalyst particles. A further property found with these nanofiberous materials is that they possess high electrical conductivities; a feature that might induce some unexpected catalytic behavior in supported metals. It was therefore anticipated that by supporting noble metals such as Pt, Rh and Pd on these materials significant differences in selectivity and activity might be observed for hydrogenation of hydrocarbons when compared to that obtained with more traditional catalyst supports, SiO2 and 7-A1203. Using carbon nanofibers as a catalyst support for these hydrocarbon reactions has, however, one drawback, the reaction and regeneration procedures of these particular catalysts have to be conducted at relatively low temperatures in order to avoid gasification of the catalyst support. Normal regeneration operations would cause irreparably damage to carbon nanofiber supported metal systems and therefore an alternative method to reactivate spent catalysts is a critical factor when considering the practical applications of materials. The adsorption of CO on noble metals is known to be relatively strong. Based on this fact, we have designed a program designed to explore the potential of using periodic doses of CO to a reactant gas mixture as an in-situ catalyst "cleaning agent". If successful, this concept would have the advantage of prolonging the active life of the catalyst without the concomitant drawback of creating any permanent detrimental effects. The adsorption of CO on noble metals is believed to be strong enough to displace the carbonaceous species adsorbed on the catalyst surface during the reaction. With this in mind a series of experiments was undertaken where the noble metals supported on selected carbon nanofiber conformations and more traditional catalyst supports were intentionally deactivated under conditions where the formation of carbon is known to occur. Once the catalyst had been suitably deactivated CO was subsequently dosed into the gas feed stream and changes in the

667 gaseous product spectrum monitored as a function of time. This sequence of events was repeated for several cycles in an attempt to establish the validity of the treatment for restoring catalyst activity over an extended period. 2. 2.1

EXPERIMENTAL Materials

The catalysts used throughout this study were prepared by a standard incipient wetness impregnation technique to realize a metal loading of 0.3wt.%. The noble metal of choice; Pt, Pd or Rh, was impregnated in a solution containing the respective nitrate precursor, no chlorine containing precursors were used in order to avoid any interference in determining the product distributions. The carbon nanofibers used in this work were supplied by Catalytic Materials Ltd.; and were highly graphitic in nature and had a N2 BET surface area of 184.0 m2/g. The characteristics of the carbon nanofibers were such that the graphite platelets constituting the structures were aligned perpendicular to the fiber axis. Prior to use the metal particles associated with the carbon nanofibers were removed by dissolution in 1.0 M hydrochloric acid over a period of 3.0 days. The efficiency of this procedure the was checked by performing X-ray diffraction analysis on the demineralized nanofibers, which showed the complete absence of any metallic components. In addition to carbon nanofibers, SiO2 and T-A1203 were also utilized as catalyst supports in these investigations. The SiO2 was supplied by the Cabot Corporation and the T-A1203 supplied by the Degussa Corporation and these oxides had N2 BET surface areas of 32.1 and 91.2 m2/g, respectively. Reagent grade noble metal nitrates used in this investigation were purchased from Aldrich Chemicals. The gases used in this study, ethylene (99.95%), hydrogen (99.999%) and helium (99.99%) were purchased from MG Industries, Inc. and were used without further purification.

2.2.

Apparatus and Procedures

Ethylene hydrogenation reactions were carried out in a vertical flow reactor system that was connected to an on-line Varian 3400 gas chromatography unit equipped with 30-m megabore columns (GS-Q). The reactor system consisted of a vertical Pyrex tube reactor (25 mm i. d. and 40 cm long) fitted with a frit at the mid-point and onto which, 100 mg of the catalyst was placed. In this arrangement the hydrocarbon/hydrogen (1:1) reactant mixture was introduced at the top of the reactor and after passage through the catalyst bed exited at the bottom. Gas flow rates to the reactor were regulated by mass flow controllers (MKS). In all cases during a typical experiment, the catalyst samples were initially reduced in a 10% hydrogen/helium at 400~ for 1.5 hours and then the temperature lowered to 100~ for the subsequent hydrogenation reaction. A predetermined composition of reactant gas containing hydrocarbon, hydrogen and helium, with the former two components being maintained at a 1:1 ratio, was introduced into the system at a flow rate of 100 cc/min and the reaction allowed to proceed for selected periods of time, with gas product samples for analysis being taken at regular intervals. The relative activity of each catalyst sample was determined from the percent total conversion of ethylene to ethane. In order to determine the effects of the addition of a small amount of CO on the hydrogenation of ethylene, a series of experiments were undertaken in which the activity and selectivity of supported noble metal catalysts was followed as a function of time on-stream and with respect to the nature of the support material. Following the completion of a given reaction the reactant gas was replaced by helium and the sample allowed to cool to room temperature. Before removal from the reactor the sample was passivated according to the previously described procedure and then stored for later characterization studies.

668 An attempt was made to determine metal particle sizes by standard chemisorption procedures, however, this approach was fraught with problems and only gave meaningful data for the noble metal particles dispersed on silica and ),-alumina. This result was not entirely unexpected since previous work from this laboratory has shown that a number of gases tend to be chemisorbed on the carbon nanofibers (19) and this behavior obscures the contribution that would be exhibited by any metal particles present on the material. In contrast, transmission electron microscopy examination of the samples proved to be a much more suitable method and provided information not only on the relative particle size distributions, but also an insight into the variations in morphological characteristics displayed by the particles on the various support media. These examinations were performed using a JEOL 2000EXII transmission electron microscope (lattice resolution 0.14 nm) that was equipped with a closed-circuit high resolution television system. Suitable transmission specimens were prepared by ultrasonic dispersion of the catalyst samples in iso-butanol and application of a drop of the suspension to a carbon support film. Images of these specimens were displayed on a TV monitor and the high magnification appearance of catalyst particles recorded on video-tape for subsequent direct transfer to a Mitsubishi printer unit. The size distribution of particle s on the different supports was determined from measurements of over 600 particles from a number of diverse regions of each specimen. In addition, morphological details of the particles and the manner by which the nature of the support influenced such features were obtained from careful inspection of many high resolution electron micrographs. It should appreciated that the majority of images obtained from conventional transmission electron microscopy are 2-dimensional in nature and as such does not allow one to determine the topographical characteristics of supported particles i.e. overall shape and thickness. There are, however, exceptional situations where it is possible to observe the profile of particles located on edges of the support and under these circumstances determination of the topographical features can be accomplished. 3. 3.1

R E S U L T S AND DISCUSSION Flow Reactor Studies

The hydrogenation of ethylene was selected as the probe reaction as deposition of carbon is known to be a facile process under appropriate conditions leading to relatively rapid catalyst deactivation. The hydrogenation of ethylene to ethane over these noble metal catalysts occurs readily at 100~ implying that the adsorbed surface species are easily hydrogenated. Land and coworkers (20) and Zaera and coworkers (21) proposed that in this particular temperature range ethylidyne was the adsorbed intermediate species during the ethylene hydrogenation reaction. It is important to allow these catalyst systems to deactivate readily in order for the experiments to be carried out in a meaningful time scale, permitting the effect of any regeneration studies be followed in a systematic fashion. The hydrogenation of a C2Hn/H2 (1:1) reactant mixture was followed as a function of the time on stream for various supported noble metal catalysts at 100~ It was apparent from these experiments that the nature of the support had a profound effect on the behavior of the noble metals. Figure 1 shows the dramatic effect that the CO purge exerts on the decomposition of ethylene over carbon nanofiber supported Rh at 100~ At each designated point 10% CO was added to the reactant feed for a period of 30 minutes and then removed from the system in order to determine the impact of this treatment on the performance of the catalyst. Upon introduction of CO to the reactant feed an immediate effect was realized, seen in the the decomposition of ethylene, which exhibited a sharp decrease. This result is not entirely unexpected since CO is predicted to strongly adsorb on the catalyst surface displacing any hydrocarbon species that are present. As a consequence, the decomposition of ethylene will

669 be suppressed as the surface coverage of CO increases. Upon removal of CO from the reactant mixture, it was observed that the conversion of ethylene increased significantly to a level that was 15% higher than that found before the CO treatment and was comparable with the activity of a flesh catalyst surface. This behavior appears to indicate that the presence of CO effectively reverses the deactivation process by performing a "cleaning" of the active catalyst surface. The conversion of ethylene decreases with time reaching a steady level after each CO treatment. As can be seen from Figure 1 the catalyst activity can be restored to a high value merely by addition of CO to the reactant feed for a period of time prior to the stage where the catalyst started to exhibit a gradual deactivation. An identical sequence of events was observed when similar loadings of either Pt or Pd on carbon nanofibers were subjected to the same reaction sequence. A very interesting feature to emerge from the more detailed product analysis was the finding that following a CO treatment there was a measurable increase in C3 and higher hydrocarbons. Furthermore,the distribution of these higher hydrocarbon species was highly dependent upon the both the noble metal and the nature of the support used in the reaction.

CO added

= 100 o 90 ,.Q 80 E 70 60 50 40 o 30 20 10 0

\

CO added Carbon Methane Ethane ---o-- Ethylene

0

20

40

60

80 100 120 140 160 180 200

Time (hours) Figure 1.Effect of CO addition on the activity of a 0.3wt.%Rh/carbon nanofiber catalyst on the decomposition of a C 2 H ~ 2 (1" 1) mixture at 100~ It should be stressed that this novel regeneration phenomenon is not restricted to noble metals supported on carbon nanofibers. Indeed, a similar effect is achieved when these metals are supported on either y-AI203 or SiO2. An example of this behavior can be seen from Figure 2, which shows the behavior of a 0.3wt.% Pt/SiO2 catalyst during exposure to a C2H4/H2 (1:1) reactant mixture at 100~ Once again, the ability of the catalyst to hydrogenate ethylene decreases dramatically upon the introduction of 10% CO to the reactant feed. When the CO is subsequently removed from the feed the hydrogenation activity of the catalyst is fully restored to its original level indicating that the "cleaning" procedure is a general phenomenon. A closer inspection of the data presented in Figure 2 reveals that when the Pt/SiO2- C2H4/H2 reaction was allowed to proceed for a prolonged period of time the catalyst readily deactivated at a much faster rate than the corresponding carbon nanofiber supported Pt catalysts. Nevertheless, upon each subsequent addition of 10% CO to the deactivated catalyst, the activity was rapidly restored to its original level of activity. As with the carbon nanofiber supported metal catalysts, the regeneration step of this system was also accompanied by a significant increase in the C3 and higher hydrocarbons products, immediately following the CO treatment. Regeneration of the catalyst by the CO

670 treatment did not appear to exert any detrimental effect on the catalyst, the activity could be repeatedly recovered to its initial level, regardless of the degree of deactivation. An identical pattern of behavior was found when Pt, Rh or Pd was supported on ),-alumina and the system treated according to the same sequence of reaction conditions.

.=.

100 90

CO added

0 added

I ~

80

7o

+Carbon

\,

.m 60 50 40 30 O d: 2 0 ~

----0-- Methane ---/k--Ethane +

Ethylene

1o 0

5

10

15 20 Time (hours)

25

30

Figure 2. The effect of CO addition on the activity of a 0.3wt.%Pt/SiO2 catalyst on the decomposition of a C2Hn/H2 (1" 1) mixture at 100~ The efficiency of the CO cleaning performance of the carbon contaminated catalyst surface can, to a certain degree, be controlled. By careful manipulation of the CO concentration and the treatment time, the catalyst can be either regenerated in stages, or completely in a single step. This aspect is highlighted in Figure 3, where the addition of CO for a period of 10 minutes to a Pd/carbon nanofiber catalyst that has been undergoing continuos reaction in C2H4/H2 at 100~ for about 19 hours, only partially restores the catalyst activity. When, however, a second CO treatment is allowed to proceed for 30 minutes, at a stage where the hydrogenation reaction has continued for a further 3.0 hours, under these circumstances, the catalyst activity is fully regenerated. In addition, the desorption of C3 and higher hydrocarbons is significantly higher when CO is added for a longer period of time suggesting a more effective cleaning operation of the surface is achieved These findings suggest that the CO coverage of the catalyst surface is a key factor in controlling the "cleaning" process and thus the regeneration of these supported metal systems. The desorption of hydrocarbon species from the catalyst surface indicates that CO is able to chemically clean the catalyst surface, thus re-instating the high activity associated with a fresh catalyst. A comprehensive series of experiments performed by Burke and Madix (22) demonstrated that on an Fe(100) surface at high coverages, CO could displace ethylene and intermediate species from the surface. It is possible that the presence of CO on the catalyst surface may also induce a partial reconstruction of certain metal sites accounting for the observed increase in activity following such a treatment. By comparing the amount of C4 and C5 products several very notable differences are observed when different catalysts are subjected to this CO treatment. The product distribution of these species may lead to a greater understanding of the carbon species that responsible for the deactivation of these catalysts. Further work on this aspect is in progress.

671

_-

100

c o added

CO added

I.----0---Carbon --0.--- Methane AI - - / ~ Ethane ---0-- Ethylene

80 60

4o

~

o 0

5

10

15

2O

25

Time (hours) Figure 3. The effect of the addition of CO on the activity of a 0.3wt.%Pd/carbon nanofiber catalyst on the decomposition of a C2H4/H2 (1:1) mixture at 100~ 3.2.

Transmission

Electron

Microscopy

Examinations

Examination of the supported Pt, Rh and Pd catalysts in the transmission electron microscope revealed the existence of some dramatic differences in both the sizes and morphological characteristics of the metal particles in the three systems. From a survey of many areas of each specimen it was evident that the metals were evenly distributed over the carbon nanofiber surface and in general the particles adopted a well defined hexagonal shape. In many cases it was possible to discern features of the underlying support through the particles, indicating that they were relatively thin. Furthermore, the electron density across any given particle appeared to be constant suggesting that they were flat structures. It was significant to find that the particle size distributions of the metals were considerably more compact on both the oxide supports than those measured on the carbon nanofiber support. From these distributions it has been possible to derive the average metal particle sizes on the three supports and the respective values are as follows: carbon nanofibers, 7.1 to 10.2 nm; silica 1.0 to 2.4 nm; and y-alumina, 0.6 to 0.8 nm. In all cases, rhodium tended to produce the smallest particles, palladium the largest and platinum exhibiting an intermediate size between these two extremes. Examination of catalysts that had been treated for prolonged periods of time under hydrogenation conditions at 100~ did not appear to exhibit any major changes in either size or morphological characteristics indicating that particle sintering was not an issue under these conitions. 4.

CONCLUSIONS

The results of this investigation clearly demonstrate that treatment of supported noble metal particles with periodic doses of CO to a hydrocarbon reactant feed is a novel, inexpensive yet highly effective method of regenerating previously deactivated hydrogenation catalysts. The major advantage is that the reactor does not need to be shut down for long periods of time to allow for costly regeneration procedures. The CO treatment is expeditious and would appear to be applicable for a wide range of catalyst systems. Supporting noble

672 metals upon the unique conformations of carbon nanofibers results in a significantly lower rate of deactivation than the corresponding SiO2 and A1203 catalysts for this model hydrogenation reaction. It would therefore be beneficial to use these novel catalyst supports as an alternative to the more traditional materials given their increased resistance to deactivation for many reactions. In this regard, the CO cleaning procedure would have major advantages over the current de-coking methods. At this point in time, however, we cannot comment on how such a treatment would impact on the formation of graphitic types of carbon, that tend to accumulate over relatively long reaction periods. REFERENCES

1. 2. .

4. 5. 6. 7. 8. 9. 10. 11. 12. 13. 14. 15. 16. 17. 18. 19. 20. 21. 22.

J.B. Butt, and E.E. Petersen, "Activation, Deactivation and Poisoning of Catalysts", Academic Press, San Diego (1988). B.C. Gates, J.R. Katzer, and G.C. Schuit, "Chemistry of Catalytic Processes", McGraw-Hill, New York (1979). C. Park, Ph.D. Thesis, The University of Glasgow (1994). L.W. Jossens, and E.E. Petersen, J. Catal., 76, (1982) 265,. R.J. Bertolacini, and R.J. Pellet, "Catalyst Deactivation", (Delmon, B. and Froment, G.E., Eds.), Elseiver, Amsterdam, p.73, (1980). Y.X. Li, K.J. Klabunde, and B.H. Davis, J. Catal. 128, (1991) 1. N.S. Figoli, M.R. Sad, J.N. Beltramini, E.L. Jablonski, and J.M. Parera, Ind. Eng. Chem. Prod. Res. Dev. 19, (1980) 545. J. Barbier, E. Churin, and P. Marecot, J. Catal. 126, (1990) 228. S.R. de Miguel, J.A. Martinez Correa, G.T. Baronetti, A.A. Castro, and O.A. Scelza, Appl. Catal. 60, (1990) 47. R.J. Verderone, C.L. Pieck, M.R. Sad, and J.M. Parera, Appl. Catal. 21, (1986) 239. C.R. Apestiguia, and J. Barbier, J. Catal. 78, (1982) 352. R . T . K . Baker, and J. J. Chludzinski, J. Catal. 64, (1980) 464. M.S. Kim, N.M. Rodriguez, R.T.K. and Baker, J. Catal. 131, (1991) 60. N.M. Rodriguez, J. Mater. Res. 8, (1993) 3233. N.M. Rodriguez, A. Chambers, and R.T.K. Baker, Langmuir 11, (1995) 3862. N.M. Rodriguez, M.S. Kim, andR.T.K. Baker, J. Phys. Chem. 98, (1994) 13108. C. Park, N.M. Rodriguez, and R.T.K. Baker, 21 lth National Meeting ACS, New Orleans, LA (1996). A. Chambers, T. Nemes, N.M. Rodriguez, and R.T.K. Baker, Appl. Catal. in press. C. Park, A. Chambers, R.T.K. Baker, and N.M. Rodriguez, submitted to Science T.A. Land, T. Michely, R.J. Behm, J.C. Hemminger, and G. Comsa, J. Phys. Chem. 97, (1992) 6774. F. Zaera, D. Godbey, G.A. Somorjai, and R. Yeates, Surf. Sci. 167, (1986) 163. M.L. Burke, and R.J. Madix, J. Am. Chem. Soc. 113, (1991) 1475.

9 Elsevier Science B.V. All rights reserved Catalyst Deactivation 1997 C.H. Bartholomew and G.A. Fuentes, editors

673

E s t i m a t i o n o f R e v e r s i b l e a n d Irreversible C o k e b y T r a n s i e n t E x p e r i m e n t s Mikael Larsson, Niklas Henriksson and Bengt Andersson Department of Chemical Reaction Engineering, Chalmers University of Technology, S-412 96 G6teborg, Sweden, mikael @cre.chalmers.se, http://www.cre.chalmers.se

Catalyst deactivation by coke formation can occur through a more or less reversible mechanism. We have applied a transient approach to model the reversible behavior of the deactivation, and to separate the deactivation from the main reaction kinetics. The deactivation of a Pt-Sn/A1203 catalyst was studied during propane dehydrogenation. The gas composition and temperature were varied during the experiments, which allowed us to model the deactivation by assuming one reversible and one irreversible type of coke. It was found that the deactivation increased with the propene concentration but was independent of the partial pressure of propane. Hydrogen decreased the deactivation rate and could even activate the catalyst by removing reversible coke.

1. INTRODUCTION Although many phenomenological models that include steps with reversible coke have been proposed, only a few attempts have been made to model this mechanism quantitatively [1,2]. Our goal in this study has been to model the reversible behavior observed during propane dehydrogenation by changing the inlet gas composition during runs. We wanted to find a method to separate the deactivation from the kinetics. Unfortunately there is no obvious way to do the latter. Many processes, such as adsorption, desorption, reaction and deactivation, occur simultaneously. For a hydrocarbon reaction deactivated due to coke formation, we can divide these processes into four types: (1) Processes much faster than the reaction rate (milliseconds), i.e. processes in equilibrium. (2) Processes in the same range as the reaction rate (seconds), i.e. processes at steady-state. (3) Relatively fast deactivation processes (minutes - hours). (4) Slower deactivation processes (hours - years). The system chosen here is the dehydrogenation of propane on Pt-Sn/A1203. This catalyst is used commercially because of its relatively high stability against coke formation and high selectivity during dehydrogenation reactions and naphtha reforming [3-7]. We have earlier measured the effect of altered reaction conditions on the amount of coke and activity [8] on this reaction system. We found that only a minor part of the coke caused the deactivation of the catalyst. This deactivating coke is formed in parallel with the non-deactivating coke. In the present work we will focus on the deactivating coke and model it by a transient method, in which the reaction conditions are changed a number of times during a run.

674

2. EXPERIMENTS About 30 mg of a Pt-Sn/A1203 catalyst were used in each run. The particle size were chosen small enough to avoid internal mass transfer limitations. The catalyst was prepared by sequential impregnation of a commercial "t-alumina support with aqueous solutions of SnC12 and H2PtC16 [9]. Characterization data are summarized in Table 1. The experiments were conducted in a fully automated flow reactor. The gas compositions were analyzed by a GC (see [8] for details). The reactor was fed with a mixture of propane, propene, hydrogen and nitrogen to achieve differential conditions. This was desired to avoid changes in the gas compositions along the bed leading to a non-uniform coke profile, and also to reduce the computational effort. The gas flow rates were controlled by mass flow controllers. The temperatures were 780, 800 and 820 K, and the total pressure was 1.5 bar. The inlet gas composition was changed in accordance with an extended 23 factorialdesigned experiment [ 10]. For each temperature, the experiment was divided into two parts to avoid very long experiments with severe deactivation. Reference conditions were repeated three times during each run. The total experimental plan contained six runs, each with 10 experimental points. The order of the points and the length between the steps (143 - 176 min) were randomized. The experiments conducted at 780 K will be named A1 and A2, while those performed at 800 and 820 K will be called B 1, B2, C1 and C2. In one experiment, the temperature experiment, the temperature was changed at constant gas composition, the same gas composition as the reference conditions in the main runs. This experiment was done in order to correlate the runs at different temperatures with each other. The catalyst was reduced in a flow of 5 ml/min of HE and 10 ml/min of N2 (flows at 273 K and 1.0 atm) at 1.5 bar for 180 min at 850 K. After that, the temperature was adjusted to the reaction temperature and allowed to stabilize for 30 min before the flow of hydrocarbons was introduced. The inlet gas composition was changed in the following range: 30 - 46 kPa C3H8, 4.5 - 7.5 kPa C3H6, and 6.8 - 11.2 kPa HE. The total flow trough the reactor was always 60 ml/min. The first GC analysis was taken after 5 minutes on stream, and the following ones every eleventh minute. Table 1. Catalyst Properties Catalyst

Particle size

Pt loading I Sn loading I

Dispersion 2 BET surface

Pt-Sn]A1203

0.05-0.14 mm

0.74 wt %

29%

1.53 wt %

172 m2/g

~By atomic absorption spectrometry. 2By hydrogen chemisorption in a volumetric system assuming H:Pt=l:l.

3. R E S U L T S Figure l a shows the gas compositions and the turnover frequency (TOF) in an experiment performed at 780 K (A1). From the TOF curve, one can see that the deactivation rate, i.e. the derivative of the TOF, changes with altered reaction conditions. The activity may even increase with time in some experimental points (not shown here). This demonstrates that it is

675 necessary to include a reversible part in a coke model. The results from the experiments will, in the following analysis, be evaluated as outlined in Figure 2. 3.1. Estimation of the deactivation function

In order to use the transient experiments to calculate the surface coverage of coke, we first introduce the continuous deactivation function. We define the deactivation function as the activity related to the first analysis after 5 minutes and compensated for the changes in reaction conditions. The compensation is carried out by assuming that the degree of deactivation immediately before and after a change in reaction conditions is the same. To be able to do so, a few assumptions have to be made (see Discussion): 1. The coke coverage varies only slowly when the reaction conditions are changed. 2. A specific coke coverage causes the same degree of deactivation independent of gas composition and reaction temperature. If these assumptions are valid, the degree of deactivation will be the same immediately before and after a change in gas composition. We have seen earlier that this is true for a change in temperature at constant gas composition for the present reaction [8]. Using these assumptions, the deactivation function for each experiment was calculated by extrapolating the reaction rates to the time for a gas change and equating the degrees of deactivation before and after the change. Figure lb shows the deactivation functions for the curve in Figure 1a. To examine the quality of the deactivation function, the relative activities in the repeated experiments were compared with the deactivation function. If the deactivation function would describe the degree of deactivation entirely, the deactivation function and the repeated experiments would agree. The agreement shown in Figure lb is good. However, slightly larger deviations occur in some of the experiments. This deviation was incorporated by adjusting the deactivation function linearly in order to achieve total correlation between the deactivation function and the repeated experiments. 1.4

:~

"-

~

0.8

~

0.6

b

0.20 c:

f

0.2 -

0

1.0

0.25

1.0

o.4 "

13)

propane,._

a

I~ ~ )

0

~

~ -

-

.... ~ ' ~ ~

5

o.,o

, hydrogen~

10

,~

,in"

~

~

~-~:;:~:~.~,~--':----~,-"

15

~-. . . . . . 20

Time, hours

o.05

"' . . . . . . . . . .

:~*-- ~: 25

30

~ 0.4 o~ m

O.2

.~I 0_ "~ oo ne

~

i'o

1's

Time, hours

~o

~s

~o

Figure 1. (a) The turnover frequencies (TOF) and gas compositions for one of the experiments performed at 780 K (A1). TOF is based on hydrogen chemisorption on a fresh catalyst [9]. The numbers in the figure indicate: (1) the first analysis after 5 minutes, (2) from where the modeling began (3) the last analysis during the second repeated experimental point with reference conditions. (b) The deactivation function and calculated total surface coverage for the same experiment (A1). The plus signs in the figure are repeated experiments with the same reaction conditions. The point-dotted section of the surface coverage curve was not used in the modeling.

676 I Malnexperiments~ A1,A2,I]I,B2,CI andC2

Determination H ofTOF

Estimation ofthe L deactivationfuncfloni ~ Temperature ~Wl experiment |

Estimationof the ~,~ coverage

surface

Modeling

]

Figure 2. Schematic picture of how the experimental data was treated.

3.2. Estimation of the surface coverage In the this paragraph we will relate the activity to changes in coke coverage. We assume that the deactivation function is proportional to the free metal surface area, i.e. the metal surface not covered by either reversible or irreversible coke. To relate the activity to the surface coverage by coke is however problematic. It is not possible to extrapolate to time zero and assume zero surface coverage, because the surface to a large extent is covered almost immediately [9,11]. Neither can we say that the initial surface coverage is the same in all experiments, since it may change with reaction conditions. Instead, the activity has to be correlated with the surface coverage by independent methods. Lin et al. [ 11 ] pulsed n-hexane on a Pt-Sn/A1203 catalyst and found by hydrogen chemisorption that about 70% of the metal surface rapidly became covered by coke, and that additional hydrocarbon pulses only slightly reduced the free metal surface. Larsson et al. [9] ran propane dehydrogenation on a PtSn/A1203 catalyst and found that the activity for the reaction between H2 and D2 was only about 15% of the activity for a fresh catalyst after 10 h on stream. In the present study, we will, in accordance with these findings, assume that the surface coverage of coke is 80% after 16 h on stream in the temperature experiment, and will then relate the other runs by comparing the activities. We assume that when the same reaction condition is applied, i.e. the same gas composition and temperature, differences in activity are only due to the degree of deactivation. Therefore, we can estimate the surface coverage for points with the same gas composition but from different runs given that the surface coverage is known in one run. Data from the temperature experiment was used to estimate the effect of changed temperature at constant coke coverage on the reaction rate. The relative activities were 0.584, 1 and 1.47 for 780, 800 and 820 K respectively. As described above, we assumed that the total coke coverage was 0.8 in the temperature experiment after 16 h, and estimated the surface coverage for all the experiments by using the activities at the end of the sixth experimental point (see number 3 in Figure l a). In this point the gas composition was the same in all the experiments. Figure 1b shows the estimated surface coverage for experiment A1. 3.3. Modeling of the surface coverage The following model of the deactivation has been made as simple as possible. A few fundamental assumptions have been used: 1. Only the coke covering the active sites is involved in the modeling. 2. Two types of coke that cause the deactivation exist, one reversible and one irreversible. 3. The reversible type of coke can be removed, be converted into a harmless type of coke, or form irreversible coke on a time-scale slow enough to be detected by the GC analysis. The amount of reversible coke is changed by modifications in the reaction conditions. 4. The irreversible coke cannot be removed in the time-scale concerned in the experiments.

677

5. The reversible and irreversible types of coke occupy the same number of active sites per carbon atom. The irreversible coke is formed only from the reversible coke. 6. A pseudo steady-state for the reversible coke is reached at the end of the first point in the experiments (after about 5 hours). The last assumption leads to the following equation:

d( Orev,O 1--Oirr,O )

dt

=0

(1)

where E)rev,0 and Oirr,0 are the surface coverages of reversible and irreversible coke, respectively, .when the modeling begins. We did not involve the first five hours of the experiments in the modeling because the deactivation process in the beginning of a run is different from the following region in the experiment [8,12]. The modeling was done by minimizing the sums of squares from the residuals between the model and the estimated surface coverage functions. This non-linear regression was solved with a Levenberg-Marquardt routine in the MATLAB software package. A number of different models that fulfilled the six assumptions above were fitted to the experimental data. Equation (2) describes the model that gave the most satisfying fit. The same equations have been used by Wolf and Petersen [1] to describe the reversible deactivation behavior during methylcyclohexane dehydrogenation on Pt/)'-A1203.

dOrev

dt -k,. PC3H6.~ Pn2 --Orev-- Oirr )-k2 . PH. 20rev . .

dOirr dt

l('3, ~)rev

(2)

- k3 "Orev

However, problems in identifying physically acceptable solutions occurred. We found for the two series at 800 K (B 1 and B2) that, after 5 h, the initial deactivation was not finished and subsequently, the pseudo steady-state assumption was probably not fulfilled. Hence, in experiment B1 and B2, the next point was omitted and the modeling began from there. Difficulties with convergence still existed and, after comparing the runs, it was found that the deactivation behavior differed in run B2. Thus the latter was omitted from the analysis, and it was now possible to find an acceptable solution. Results from the modeling are shown in Figure 3 and the parameter estimation, with linearized 95% confidence intervals, is given in Table 2. The highest correlation between the parameters was 0.87. One can see that the activation energies have relatively large confidence intervals. This can be explained by both the correlation between the parameters and the relatively small temperature interval in which the experiments were conducted. The validity of the model was checked by studying the residuals. No patterns indicating effects that had been neglected were found.

678 (9

1

•r 0.8

(9 1 A2

A1

r.)

(9 O 0.4 't~'0. 2

r~J

~

~ 0.8

0.6

0.8

(9 ~ 0.4

0.4

5

1'0 l's 2'0 2s 30 ~...

/

~ ~

,5

o;

Time, hours

9 1 C2

r~o.6

~176

"

B1

0to~~~

~~:

'0.2

1 C1

8

/"-

~ 0.8

~0.6

00

(9 1

,,ao~/(]-o~,)

1'0 1'5 2'0 ~5 3'0 Time, hours

0

O~ 1.o

1'0 1'5 2'0 ;~5 3'0 Time, hours ..-

0to

~ 0.4

Orev r~ 0"2

1'0 1'5 ;~0 25 30 Time, hours

~ 0.7

= ~5 o.

5

1'0 1'5 2'0 2'5 3'0 0.5o.s o'.B 0'.7 ola 0'.9 Ii0 Time, hours Experimental 0total

Figure 3. The results from the modeling: reversible, irreversible and total coverage of coke. The experimental results are indicated with dashed lines and the model with solid lines. The ratios between the coverage of reversible coke and the surface not covered with irreversible coke are also shown. The predicted versus the experimental surface coverages are shown in the last graph. Table 2. The parameter values and linearized 95% confidence intervals Parameter Value 1.199" 10 -4 -i- 1.748" 10-5 Al (s -l) Pre-exponential factor l A2 (s-lPa -l) Pre-exponential factor 1 1.372" 108 + 4.794" 10.9 A3 (s l ) Pre-exponential factor 1 6.178.10 -5 + 8.968.10 -4 Ehl (kJ mol l ) Activation energy 119.5 + 43.8 EA2 (kJ mo1-1) Activation energy 68.5 + 103.3 EA3 (kJ mol -~) Activation energy 41.9 + 42.0 IEstimated from the centered Arrhenius expression ki=Ai'exp(Ehi/R'(1/T-I/T0)), T0=800 K.

4. DISCUSSION Equation (2) explains the transient deactivation with a model describing reversible and irreversible coke. We can see that the partial pressure of propane in the reactor does not influence the deactivation. This has also been demonstrated in an earlier study of the same system [8]. This observation is consistent with kinetic models for propane dehydrogenation proposed by Loc et al. [ 13]. They suggested that the rate-determining step is the dissociative adsorption of propane. From this mechanism it follows that the deactivation will be

679 independent of the propane partial pressure if the deactivation is caused by coke formed from intermediates in the reaction. The formation of reversible coke in Equation (2) can be explained by a model where a coke precursor first is formed from adsorbed propene by dehydrogenation. This precursor is in equilibrium with the propene and hydrogen in the gas phase, and it will form coke through a reversible mechanism. Finally irreversible coke is formed from the reversible coke. Figure 4 shows a mechanism for the formation of deactivating coke during propane dehydrogenation. The mechanism is derived from Equation (2) and from the propane dehydrogenation mechanism proposed by Loc et al. [13]. The coke formation part is in every sense consistent with a model proposed by Wolf and Petersen [ 1]. The definition of reversible coke is not clear in the literature, but is usually said to be the portion of coke that can be removed relatively easily by hydrogen treatment. The definition in this study has been, in effect, the portion of the coke that can be altered by changing the reaction conditions. This portion will vary according to how the experiments are carded out. More extreme changes in reaction conditions would allow a larger part of the coke to be modified. Two transient processes are superimposed when performing transient experiments on a deactivating reaction system. The difficulty in separating the two processes was overcome by running the reaction at differential conditions and using the deactivation function. In order to convert the turnover frequency curves to deactivation functions, it was assumed that the degree of deactivation did not change when the gas composition or the temperature was changed. This assumption is essential for the whole modeling. The basic idea is that what we define as coke coverage is changed only slowly and does not change during the short time it takes to change the reaction conditions. Repeated experiments were performed to verify the assumption and the small differences detected were corrected for. It is also necessary that the reaction mechanism does not change when the reaction conditions are altered. Our preliminary kinetic model of the main reaction kinetics indicates that this assumption is valid [ 14]. It was difficult to fulfill the objective of achieving differential conditions and producing accurate measurements when the activity was decreased by sometimes more than 90%. We had to admit non-differential conditions in the beginning of some runs (the first few hours). Because of the objective above, the variation in the variables had to be relatively limited. Therefore the parameter values are only valid in this limited range. The problems can be solved by changing the space velocity during the run and/or using a reactor that can be described as a stirred tank reactor. Hofmann and Kolb [15] and Beirnaert et al. [ 16] discuss the benefits of using a recycle reactor in deactivation studies. The surface coverage was assumed to be 80% after 16 h in the temperature experiment. This choice is not important for the qualitative analysis of the experiments, but will of course affect the parameter values. The method could be improved if the deactivation were correlated with the surface coverage by some independent method. Rivera-Latas et al. [17] have compared several methods to estimate the active surface on a deactivated catalyst and recommended using titration of preadsorbed 02 by H2.

C3H8

C3H7.+ H*

/

Coke Precursor + H2 r2 l~rl +H 2

C3H6 -~ ~ C3H6-+ H* Reversible Coke

r3

Irreversible Coke

Figure 4. A model for the formation of deactivating coke during propane dehydrogenation.

680 5. CONCLUSIONS The experiments clearly indicated the existence of reversible deactivation on the catalyst. It was possible to explain this behavior by a simple model including reversible and irreversible types of coke, and a mechanistic model based on the results was proposed in Figure 4. The transient approach, in which the reaction conditions were changed a number of times during each run, was powerful and yielded abundant information in a limited number of runs. Transient experiments were necessary in order to distinguish between the two types of deactivation. However, interpretation was difficult and it requires extensive numerical evaluations.

ACKNOWLEDGMENT This work was supported by the Swedish Research Council for Engineering Sciences (TFR).

REFERENCES 1. E.E. Wolf and E. E. Petersen, J. Catal., 46 (1977) 190. 2. J. Corella, J. Adanez and A. Monz6n, Ind. Eng. Chem. Res., 27 (1988) 375. 3 . L. C. Loc, N. A. Gaidai, S. L. Kiperman, in M. J. Phillips and M. Ternan (Editors), Proc. 9th Int. Congr. on Catalysis, Calgary 1988, Vol. 3, Chem. Inst. Canada, Ottawa, 1988, p. 1261. 4. P. Praserthdam, U. Panchareon and J. Tscheikuna, Kenkyu Hokoku - Asahi Garasu Zaidan, 61 (1992) 313. 5. R. Srinivasan and B. H. Davis, Platinum Metals Rev., 36 (1992) 151. 6. F.M. Dautzenberg, J. N. Helle, P. Biloen and W. M. H. Sachtler, J. Catal., 63 (1980) 119. 7. R. Burch and L. C. Garla, J. Catal., 71 (1981) 360. 8. M. Larsson, M. Hult6n, E. A. Blekkan and B. Andersson, J. Catal., 164 (1996) 44. 9. M. Larsson, B. Andersson, O. A. Bari~s and A. Holmen, Stud. Surf. Sci. Catal., 88 (1994) 233. 10. D. C. Montgomery, "Design and Analysis of Experiments", 2nd ed., John Wiley & Sons, New York, 1984. 11. L. Lin, T. Zhang, J. Zang and Z. Xu, Appl. Catal., 67 (1990) 11. 12. L. K. Loc, N. A. Gaidai, S. L. Kiperman, N. M. Podkletnova and S. B. Kogan, Kinetics Catal., 29 (1988) 989. 13. L. K. Loc, N. A. Gaidai, S. L. Kiperman, H. S. T'huoang, N. M. Podkletnova and S. B. Kogan, Kinetics Catal., 32 (1991) 61. 14. M. Larsson, N. Henriksson and B. Andersson, in preparation. 15. H. Hofmann and G. Kolb, Stud. Surf. Sci. Catal., 88 (1994) 257. 16. H. C. Beirnaert, R. Vermeulen and G. F. Froment, Stud. Surf. Sci. Catal., 88 (1994) 97. 17. F. J. Rivera-Latas, R. Dalla Betta and M. Boudart, AIChE J., 38 (1992) 771.

9

Elsevier Science B. V. All rights reserved

Catalyst Deactivation 1997 C.H. Bartholomewand G.A. Fuentes, editors

681

D e a c t i v a t i o n Studies o f Ni/A1203 Z S M - 5 Catalysts: Effect o f N i c k e l Incorporation Jolanta R. Grzechowiak a, Aleksandra Masalska a, Leszek Kqpifiski b, Jacek Rynkowski c alnstitute of Chemistry and Technology of Petroleum and Coal, Technical University of Wroctaw, 50-344 Wrodaw, 7/9 Gdafiska Str., Poland blnstitute of Low Temperatures and Structure Research, Polish Academy of Science, 50-422 Wroctaw, 20k61na Str., Poland Clnstitute of General and Ecological Chemistry, Technical University of L6d~, 90-924 L6d~., 36 Zwirki Str., Poland

Alumina-zeolite supported nickel catalysts were investigated for the conversion of n-hexane and dewaxing of diesel oil fraction. The method of nickel incorporation influenced the susceptibility of nickel to reduction and catalyst activity. Investigated also was the effect of reoxidation temperature on the accessibility of the metallic surface. Differences in the activities of the catalysts can be attributed to the morphology of the deposits.

1. INTRODUCTION Deactivation of metal-supported catalysts is affected by metal-support interactions. The effect of the support (alumina and silica) on the sintering of nickel (15 % Ni) was described elsewhere [1]. The results for sintering in hydrogen at 800 ~ showed that the extent of sintering for any given period was substantially greater for Ni/silica. The higher thermal stability of nickel on alumina with respect to silica can partly be attributed to a stronger metalsupport interaction. The data reported by Bartholomew et al.[2, 3] for Ni/alumina catalysts show a relationship between the structure of A1203 and the thermal stability of nickel. A number of techniques, including XRD [4,5], XPS [6], EXAFS [7] and TPR [5, 8-9], were applied to study the complex nature of nickel oxide-support interactions. The deactivation of acid zeolite catalysts is mainly due to deposition within the pores and on the outer surface of the crystallites of heavy secondary products. The deactivation and regeneration of zeolites was discussed in more detail by Guisnet et al. [ 10]. The formation of coke on metal zeolite-alumina catalysts depends on the distribution of the acid sites, metal-toacid ratio and of the texture. This paper describes the investigation of the deactivation of N i O / A 1 2 0 3 + N i , H - Z S M - 5 catalyst systems and the contribution of the method of nickel incorporation to the performance of the catalysts. We made use of 7-A1203, because with that support the rate of aging in Ni/alumina catalysts was found to be the lowest [1]. The activity experiments described here involved n-hexane conversion and diesel oil dewaxing.

682 2. EXPERIMENTAL

2.1. Catalysts preparation Investigated were catalytic systems containing 50 wt.% of zeolite Ni,H-ZSM-5 and 8 wt.% of NiO. Thus, in the preparation of Catalyst AZ-1 and Catalyst AZ-2 nickel was incorporated at the forming stage by adding Ni(NO3)2 to the peptisate before extrusion. The preparation of Catalyst AZ-1 included aging of the aluminium hydroxide/zeolite mixture in aqueous solution for 24 h. Catalyst AZ-3 was prepared differently - one half of the anticipated nickel amount was incorporated by impregnation. The catalysts extrudates were dried at 110 ~ and calcined at 480 ~ for three hours. The methods of preparation can be described as follows: Catalyst AZ-1 AI(OH)3 + Ni,HZSM-5 ~ aging Ni(NO3)-2 forming ~ drying ~ calcination Catalyst AZ-2 AI(OH)3 + Ni,HZSM-5

Ni(NO3)2~

forming

~ drying

~ calcination

Catalyst AZ-3 AI(OH)3 + Ni,HZSM-5 Ni(NO3)2 forming ----,drying --* calcination Ni(NO3-);2 impregnation

2.2. Catalysts characterization Acidity and acid strength distribution were evaluated by temperature-programmed desorption (TPD) of ammonia. Acid sites were classified as: weak (W) Tats < 300 ~ medium (M) 300 < Tdes < 450 ~ and strong (S) 450 < Tdes < 550 ~ The structure of the samples was studied by transmission electron microscopy (TEM), electron diffraction (ED) and X-ray diffraction (XRD) - CuKa radiation. Temperature-programmed experiments (TPR-TPO) involved the following cycles : up to 900 ~

up to 600 ~

up to 900 ~

up to 900 ~

up to 900 ~

The mixtures of 90 vol.% and 10 vol.% H2 or 02, respectively, were used at a gas flow rate of 50 cm3/min and a linear temperature ramp of 20 deg/min. Carbon deposits were characterized by TPO (10 deg/min in a stream of 10 vol.% 02 + 90 vol.% He).

2.3. Catalytic activity measurements Catalytic activity was studied in laboratory and bench-scale reactor systems. Laboratory catalytic tests (n-hexane conversion) were carried out in a through-flow reactor under the following conditions: T=200-450 ~ LHSV=7 h l. Bench-scale tests were run in a continuous -flow high-pressure microreactor unit, using diesel oil as a feedstock. Catalysts were reduced and then sulfided at 330 ~ for 5 h. The process was studied at T=300-330 ~ LHSV=3 h 1, p=3.5 MPa and H2:CH=500 Nm3/m3. The criterion for catalyst activity assessment was the freezing point (-20 ~ the measurements were done for the stabilized products. 3. RESULTS AND DISCUSSION The method of nickel incorporation had little effect on the pore structure of the catalysts, which were characterized by a comparable pore size up to 100 nm (about 0.30 cm3g"1) and by

683 a comparable pore size distribution in the range of up to 1.5 nm (about 0.04 cm3g-]), 1.5 to 3.0 nm (about 0.1cm3g'l), 3 to 5 nm (about 0.1 cm3gl), and 5 to 10 nm (about 0.05 cm3gl). Of the three contacts Catalyst AZ-1 had the smallest pore volume (0.03 cm3g-1) in the range of 10 to 100 nm. As it may be inferred from TPDNH3 data, the method of nickel introduction did not affect acidity or acid site distribution. Medium-strength acid sites were predominant (50 % of total acidity which varied from 0.81 to 0.89 mmol NH3/g ) for the catalysts tested (Fig. 1).

r

0,6"

Z

0,5.

.

~60

-

" 4o

o

EE 0,4.

-[80 ...

-

.I 80

0,2.

.

gO . r ~ ~ W

M

S

Figure 1. Acid site distribution; W (weak), M (medium), S (strong).

2oo

3oo

-,

~

!

.80

~"'~162 "'""e -

160

-

.14o

,,.4o -

9

9

"160 40

"

0,3.

0,1

- 8o

20

d r

~oo'

3bo'-,~x)~,~p,,~

Figure 2. Conversion of n-hexane and product distribution over AZ-1 (o), AZ-2 (x) and AZ-3 (.).

The results of activity tests show that in the temperature range 200-300 ~ conversion of n-hexane proceeds faster over AZ-1 than over AZ-2 or AZ-3 (Fig. 2). When nickel was incorporated by the method used with Catalyst AZ-1 the formation of products with boiling points above that of n-hexane (HBF) was shifted to lower temperatures. The results of bench-scale tests showed also that the method of nickel incorporation does affect the catalytic behaviour (Fig. 3). There was a rapid deactivation of Catalyst AZ-1 compared to the rate of deactivation for AZ-1 and AZ-2. Fresh samples did not very much differ in morphology (Fig.4). All catalysts consisted of a mixture of zeolite crystallites (0.5-1 ~m) and aggregates of fine crystalline fibrous alumina. Electron diffraction showed (Fig. 4) that the zeolite grains were single crystals giving sharp spot patterns, while alumina aggregates produced pattems with broad rings characteristic of ,/-alumina. No nickel or nickel containing phase could be detected by TEM and ED. The XRD spectra of fresh samples (Fig. 5a) contained lines which were characteristic of the Ni-form of the ZSM-5 zeolite (splitting of the strongest line at 23.10~ Again, only broad peaks from ),-alumina were detected, and Ni was not. For the spent catalysts differences in microstructures were observed. Zeolite grains preserved their shape and crystal structure, and so did alumina grains. First of all, the zeolite grains of spent catalysts contained dark particles of various size (5-20 nm) (Fig. 6); microdiffraction made it possible to identify them as metallic Ni. The migration of the active species from the lattice to the external surface of the zeolite (cf. Figs. 4 and 6) was also reported by Zhang et al. for ZSM-5 exchanged with Pd [ 11]. We could not detect the presence of metallic Ni on alumina grains.

684

"1~I

-20

--~ - ~ -

-30 -40 -50

t

,

i

Ca~alys~ A2-I

.....

,

I

I

I

J

--20

'~I

-ao -40

~

_

I

ra~,,,,,s~ AZ-2

-10 -20

-30 -40

-50

cata jst AZ-S 15

30

i

45

|

60

l

75

1

90

l

I

105 Time, h

Figure 3. Freezing points of stabilized products versus duration of the process at reaction temperatures of 300 ~ (.), 310 ~ (a), 320 ~ (x), and 330 ~ (o). The process of Ni migration may also have contributed to the observed modification of the XRD diffractogram of the samples (Figs. 5a and 5b). The diffractograms of the spent catalysts are characteristic of the H-form of ZSM-5 [ 12]. The edges of the zeolite grains in spent samples were sharp and clean and exhibited a clear lattice image, so we could exclude massive coke deposition (however, the existence of a thin, one- or two-monolayer thick, carbon deposit cannot be excluded). No massive carbon deposition on alumina aggregates could be seen. In the AZ-1 catalyst we found, however, that a large amount of carbon in the form of tube-like filaments of varying length was deposited on some alumina grains (Fig. 7).

L ~.o0

Figure 4. Electron micrograph and electron diffractogram of fresh catalysts.

20

30.00

Figure 5. XRD diffractograms of fresh (a) and spent (b) catalysts.

685 " I

Figure 6. Electron catalysts.

micrograph of spent

Iiiii: i I I "~

I I ~ I II

,

"lll~ ,

Figure 7. Filamenteous carbon deposit on alumina grain in spent AZ-1 catalyst.

The reduction of 8 % NiO/AlzO3+Ni,HZSM-5 catalysts (TPR I) occurs in two stages with peak maxima at 420-450 ~ and 780-810 ~ (Fig. 8). The first peak is attributed to the reduction of highly dispersed NiO, and the other to the reduction of Ni ions that reacted with the support to form unstoichiometric Ni-A1 oxides and/or Ni 2+ located in the zeolite. Catalyst AZ-1 showed the lowest susceptibility to reduction. The TPR II profiles (after TPO to 600 ~ revealed the presence of easily reducible species on the catalyst surface. The maximum below 300 ~ reflects the reduction of Ni (referred to in literature as "free" NiO) which does not interact with the support [13, 14]. The hydrogen consumption in the temperature range 330630 ~ can be ascribed to the reduction of a more dispersed NiO form, interacting (but not

~o= ~)

~

AZ-3

o~

b)

o

400

8oo "c

Figure 8. TPR pattems of catalysts (a) TPR I, (b) TPR II, (c) TPR III. (See reduction/oxidation cycles).

6;0

,c

Figure 9. The TPO profiles of spent catalysts, (a) AZ-2 and (b) AZ-1 (b) after diesel oil dewaxing.

686 chemically bound) with the support. It can be assumed that the first reduction process partially sintered the Ni particles to produce large agglomerates which were then reoxidized in the course of the TPO I process, yielding easily reducible NiO species. TPR III profiles of catalysts after reoxidation up to 900 ~ show that those severe reoxidation conditions lead to solid-state diffusion of Ni 2+ ions into the AI20 3 lattice, thus contributing to the formation of Ni-AI spinels. The reduction of these species requires elevated temperatures (TPR III peak at about 900 ~ [15]. The TPO results for spent catalysts after diesel oil dewaxing reveals only one temperature form of the deposit on Catalyst AZ-2 and two for catalyst AZ-1 (Fig. 9). In the lower activity catalyst AZ-1, the high-temperature carbon form accounted for less than 50 % of the total. It was also found that the amount of the deposit on Catalyst AZ-1 was greater than on Catalyst AZ-3 (7.6 wt. % and 5.9 %, respectively), even though the time of deposition for Catalyst AZ-1 was only one-fourth that for Catalyst AZ-3. 4. SUMMARY

The results show that during diesel oil conversion over catalysts containing high-silica zeolites deactivation depends not so much on the method by which nickel has been incorporated as on the method of aluminium oxide incorporation into the zeolite. This finding was confirmed by the selectivity of the catalysts to coke precursors (n-hexane conversion). Catalytic tests indicated that in the lower activity Catalyst AZ-1 the HBF was obtained at lower reaction temperatures (200-300 ~ in the Catalyst AZ-1 sample alone, tubelike filaments of various length produced blocking of the channel openings, thus impeding the reaction. The presence of low-temperature peak (which is characteristic of the coke deposited on metal sites) can also explain the low activity of Catalyst AZ- 1 [15, 16]. Catalyst AZ- 1 was also found to be slightly less susceptible to reduction. For all the spent catalysts there was a migration of nickel species from the lattice to the external surface of the zeolite.

REFERENCES

1. C.H. Bartholomew, Studies in Surface Science and Catalysis, 88 (1994) 1. 2. C.H. Bartholomew and W.L. Sorensen, J. Catal., 81 (1983) 131. 3. C.H. Bartholomew, R.B. Pannell and R.W. Fowler, J. Catal., 79 (1983) 32. 4. B. Scheffer, J.J. Heijeinga and J.A, Moulijn, J. Phys Chem., 91 (1987) 4752. 5. S.L. Chen, H.L. Zhang, J. Hu, C. Contescu, J.A. Schwarz, Appl. Catal., 73 ( 1991 ) 289. 6. M. Wu and D.H. Hercules, J. Phys, Chem., 83 (1979) 2003. 7. G.U. Kulkami, G. Sankar and C.N.R. Rao, J. Catal., 131 (1991) 491. 8. A. Kadldaodayan and A. Brenner, J. Catal., 117 (1989) 311. 9. B. Scheffer, P. Molhoek and J.A. Moulijn, Appl. Catal., 46 (1989) 11. 10. M. Guisnet and P. Magnoux, Studies in Surface Science and Catalysis, 88 (1994) 53. 11. Z. Zhang, B. Lerner, G.D. Lei and W.M.H. Sachtler, J.Catal., 140 (1993) 481. 12. JCPDS Powder Diffraction File No. 37-0359. 13. J. Zieli6ski, J. Catal., 76 (1982) 157. 14. J.M. Rynkowski, T. Paryjczak and M. Lenik, Appl. Catal., 106 (1993) 73. 15. J.M. Grau and J.M. Parera, Appl. Catal., 70 (1991) 9. 16. A. Masalska, J.R. Grzechowiak, Studies in Surface Science and Catalysis, 88 (1994) 561.

687 Author Index

Aguayo, A.T. Alfonso, J.C. Amadeo, N.E. Andersson, B. Apestegu'a, C.R. Aranda, D.A.G. Baba, K. Bacherikova, I. Baker, R.T.K. Barbier, J. Barbosa, M.N.M. Barre-Chassonnery, M. Bartholomew, C.H. Batina, N. Becerra, M.J. Beck, D.D. Bernier, A. Bilbao, J. Biscaia Jr., E.C. Bogutskaya, L. Boitiaux, J.P. Bonardet, J.-L. Boock, L.T. Borgna, A. Borodko, Y.G. Brown, T.C. Buglass, J.G. Butt, J.B. Camacho, L.A. Cameron, C. Carmona, E. Castro, A.A. Catani, R. Cerqueira, H.S. Chambers, A. Chen, D. Chocron, M. Comelli, R.A. Corro, G. Costa, L. Cosyns, J. Datye, A.K. Davis, B.H. de Deugd, R.M. de Jong, K.P.

567 335 535,311 673 495 335 259 207 99,593 327,359,421 415 561 517,585 655 399 21 267 567 303 207 447 647 367 183,495 655 383 127 69 665 447 351 191 463 303 99,593 159 311 139 359 231 447 501 501,527 283 127

de Miguel, S.R. Deady, J. Del Angel, G. Delmon, B. den Hollander, M.A. Dereppe, J.-M. Didillon, B. Domeniconi, T. Earl, W.L. Eliason, S.A. Espinosa, G. Falabella S.-Aguiar, E. Faraci, G. Fathi, M. Figoli, N.S. Finelli, Z.R. Flego, C. Fraissard, J. Frankel, K.A. Frety, R. Froment, G.F. Fuentes, G.A. Fung, S.C. Galasso, L. Gamas, E.D. Gamero, M.P. Garcia de Leon, R. Garcia, L. Garciafigueroa, E. Garetto, T.F. Gayubo, A.G. Giannetto, G. G6mez, E. G6mez, S.A. Gonzales, R. Gonzalez, M.G. Goto, Y. Grenfell, J. Grzechowiak, J.R. Guimon, C. Guisnet, M. Guzman, M.O. Haber, J. Henriksson, N. Henriques, C.A.

191 367 421 39,267 295 231,647 447 647 639 517 421 303,427 479 561 139,415 139 479 647 239 335 53 225,573 625 479 119 375,391 455 399 455 495 567 399 351 225 455,375 535 439 167 681 183 1

375 207 673 427

688 Hepola, J. Hernandez, B.F. Hildenbrand, D.L. Ho, T.C. Holmen, A. Hughes, C.D. Ihm, S.-K. Ioffe, L.M. Isnard, P. Jackson, N.B. Jackson, S.D. Jacubinas, R.M. Jaikaew, B. Jang, B.W.-L Janssens, J.P. Kawashima, H. Kepinski, L. Kermarec, M. Kharlamov, A. Kim, D.-W. Kim, M.H. Kim, M.S. Kim, Y.G. Kunatippapong, S. Laborde, M. Larsen, G. Larsson, M. Lau, K.-H. Lee, D.-K. Lee, K.-H. Li, C. Lim, N. Lim, T.F. Liu, H. Liu, K. Lodeng, R. Lotero, E. Louis, C. Maarsen, P.K. Magnoux, P. Makkee, M. Maldonado, M.C. Mansker, L. Marchi, A.J. Marecot, P.

471 455,391 601 625 159,561 639 343 655 267 501 167 639 153 239 283 259 681 617 207 343 213 99 213 153 311 147 673 601 343 463 383 153 367 199 625 561 147 617 127 1

295 375,391 501 183 327,359,421

Martin, D. Martinez, A. Martra, G. Masalska, A. Massoth, F.E. Matheson, I.M. McCarty, J.G. Miglio, R. Minh, C.L. Mir6, E.E. Mirodatos, C. Mogica, E. Mojica, M.E. Moljord, K. Mongkhonsi, T. Monteiro, J.L.F. Montoya, J.A. Monzon, A. Mooiweer, H.H. Moran, A.L. Moreno, J.C. Moulijn, J.A. Munro, S. Murta Valle, M.L. Nabity, M. Nagai, M. Nam, I.-S. Navarrete, J. N6dez, C. Nufiez, M. O'Brien, R.J. Oelker, P. Omi, S. Ortega, J.M. Ott, K.C. Padr6, C.L. Paffett, M.T. Papa, J. Parera, J.M. Park, C. Peng, X.D. Petersen, E.E. Petkovic, L. Petunchi, J.O. Pieck, C.L.

1

225 617 681 275 167 601 463 383 231 617 455 375 159 153 427 391 183,609 127 567 455,375 283,295 167 427 147 439 213 391 251 351 501,527 267 439 567 639 191 639 399 407,433 99,593,665 175 87 147 231 327,407,415,433

689 Praserthdam, P. Querini, C.A. Quincoces, C.E. Raje, A.P. Rangel, M.C. Raval, R. Ray, J.-L. Rebo, H.P. Resasco, D.E. Roa, E. Roberts, G.W. Rodriguez, J.C. Rodriguez, N.M. Rodriguez-Otal, L.M. Rokstad, O.A. Rudnitsky, L.A. Ruiz, P. Rumschitzki, D.S. Rusic, D. Rynkowski, J. Salinas-Rodriguez, E. Sanchez-Rubio, M. Santos, F. Sarkany, A. Sarrazin, P. Sasuga, H. Scelza, O.A. Schifter, I. Schmal, M. Sie, S.T. Simell, P. Smith, J.M. Spivey, J.J. Springuel-Huet, M.-A. Stagg, S.M. Stoch, J. Suarez, C. Swaan, H.M. Szanyi, J. Tabata, T. Toledo, A.J.A. Toseland, B.A. Underwood, R.P. Uzio, D. van Langeveld, A.D. Varela, S.

153 139,147,407 535 527 415 167 251 159 199,543 407 239 183,609 99,593,665 319 561 551 267 625 487 681 573 319 399 111 447 439 191 119,421 335 283 471 199 239 647 543 207 351 617 639 259 391 175 175 447 283 427

Vargas, E.A. Vazquez, R.A. Vidotto, S. Viveros, T. VonDreele, R. Webb, G Wolf, E.E. Xu, L. Yaluris, G. Yanez, F. Yaremenko, V. Zazhigalov, V. Zrncevic, S.

391 391 479 319, 609 639 167 463 527 367 399 207 207 487

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STUDIES IN SURFACE SCIENCE AND CATALYSIS Advisory Editors: B. Delmon, Universit~ Catholique de Louvain, Louvain-la-Neuve, Belgium J.T. Yates, University of Pittsburgh, Pittsburgh, PA, U.S.A.

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Preparation of Catalysts I.Scientific Bases for the Preparation of Heterogeneous Catalysts. Proceedings of the First International Symposium, Brussels, October 14-17,1975 edited by B. Delmon, P.A.Jacobs and G. Poncelet The Control of the Reactivity of Solids. A Critical Survey of the Factors that Influence the Reactivity of Solids, with Special Emphasis on the Control of the Chemical Processes in Relation to Practical Applications by V.V. Boldyrev, M. Bulens and B. Delmon Preparation of Catalysts I1. Scientific Bases for the Preparation of Heterogeneous Catalysts. Proceedings of the Second International Symposium, Louvain-la-Neuve, September 4-7, 1978 edited by B. Delmon, P. Grange, P.Jacobs and G. Poncelet Growth and Properties of Metal Clusters. Applications to Catalysis and the Photographic Process. Proceedings of the 32nd International Meeting of the Soci6t6 de Chimie Physique, Villeurbanne, September 24-28, 1979 edited by J. Bourdon Catalysis by Zeolites. Proceedings of an International Symposium, Ecully (Lyon), September 9-11, 1980 edited by B. Imelik, C. Naccache, Y. Ben Taarit, J.C. Vedrine, G. Coudurier and H. Praliaud Catalyst Deactivation. Proceedings of an International Symposium, Antwerp, October 13-15,1980 edited by B. Delmon and G.F. Froment New Horizons in Catalysis. Proceedings of the 7th International Congress on Catalysis, Tokyo, June 30-July4, 1980. Parts A and B edited by T. Seiyama and K. Tanabe Catalysis by Supported Complexes by Yu.l. Yermakov, B.N. Kuznetsov and V.A. Zakharov Physics of Solid Surfaces. Proceedings of a Symposium, Bechyhe, September 29-October 3,1980 edited by M. LazniEka Adsorption at the Gas-Solid and Liquid-Solid Interface. Proceedings of an International Symposium, Aix-en-Provence, September 21-23, 1981 edited by J. Rouquerol and K.S.W. Sing Metal-Support and Metal-Additive Effects in Catalysis. Proceedings of an International Symposium, Ecully (Lyon), September 14-16, 1982 edited by B. Imelik, C. Naccache, G. Coudurier, H. Praliaud, P. Meriaudeau, P. Gallezot, G.A. Martin and J.C. Vedrine Metal Microstructures in Zeolites. Preparation - Properties-Applications. Proceedings of a Workshop, Bremen, September 22-24, 1982 edited by P.A. Jacobs, N.I. Jaeger, P. Jin3 and G. Schulz-Ekloff Adsorption on Metal Surfaces. An Integrated Approach edited by J. Bdnard Vibrations at Surfaces. Proceedings of the Third International Conference, Asilomar, CA, September 1-4, 1982 edited by C.R. Brundle and H. Morawitz Heterogeneous Catalytic Reactions Involving Molecular Oxygen by G.I. Golodets

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Preparation of Catalysts III. Scientific Bases for the Preparation of Heterogeneous Catalysts. Proceedings of the Third International Symposium, Louvain-la-Neuve, September 6-9, 1982 edited by G. Poncelet, P. Grange and P.A. Jacobs Spillover of Adsorbed Species. Proceedings of an International Symposium, Lyon-Villeurbanne, September 12-16, 1983 edited by G.M. Pajonk, S.J. Teichner and J.E. Germain Structure and Reactivity of Modified Zeolites. Proceedings of an International Conference, Prague, July 9-13, 1984 edited by P.A. Jacobs, N.I. Jaeger, P.Ji~, V.B. Kazansky and G. Schulz-Ekloff Catalysis on the Energy Scene. Proceedings of the 9th Canadian Symposium on Catalysis, Quebec, P.Q., September 30-October 3, 1984 edited by S. Kaliaguine and A. Mahay Catalysis by Acids and Bases. Proceedings of an International Symposium, Villeurbanne (Lyon), September 25-27, 1984 edited by B. Imelik, C. Naccache, G. Coudurier, Y. Ben Taarit and J.C. Vedrine Adsorption and Catalysis on Oxide Surfaces. Proceedings of a Symposium, Uxbridge, June 28-29, 1984 edited by M. Che and G.C. Bond Unsteady Processes in Catalytic Reactors by Yu.Sh. Matros Physics of Solid Surfaces 1984 edited by J. Koukal Zeolites: Synthesis, Structure, Technology and Application. Proceedings of an International Symposium, Portoro~-Portorose, September 3-8, 1984 edited by B. Dr~aj, S. Ho~evar and S. Pejovnik Catalytic Polymerization of Olefins. Proceedings of the International Symposium on Future Aspects of Olefin Polymerization, Tokyo, July 4-6, 1985 edited by T. Keii and K. Soga Vibrations at Surfaces 1985. Proceedings of the Fourth International Conference, Bowness-on-Windermere, September 15-19, 1985 edited by D.A. King, N.V. Richardson and S. Holloway Catalytic Hydrogenation edited by L. Cerven# New Developments in Zeolite Science and Technology. Proceedings of the 7th International Zeolite Conference, Tokyo, August 17-22, 1986 edited by Y. Murakami, A. lijima and J.W. Ward Metal Clusters in Catalysis edited by B.C. Gates, L. Guczi and H. Kn6zinger Catalysis and Automotive Pollution Control. Proceedings of the First International Symposium, Brussels, September 8-11, 1986 edited by A. Crucq and A. Frennet Preparation of Catalysts IV. Scientific Bases for the Preparation of Heterogeneous Catalysts. Proceedings of the Fourth International Symposium, Louvain-la-Neuve, September 1-4, 1986 edited by B. Delmon, P.Grange, P.A. Jacobs and G. Poncelet Thin Metal Films and Gas Chemisorption edited by P.Wissmann Synthesis of High-silica Aluminosilicate Zeolites edited by P.A. Jacobs and J.A. Martens Catalyst Deactivation 1987. Proceedings of the 4th International Symposium, Antwerp, September 29-October 1, 1987 edited by B. Delmon and G.F. Froment Keynotes in Energy-Related Catalysis edited by S. Kaliaguine

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Methane Conversion. Proceedings of a Symposium on the Production of Fuels and Chemicals from Natural Gas, Auckland, April 27-30, 1987 edited by D.M. Bibby, C.D. Chang, R.F. Howe and S. Yurchak Innovation in Zeolite Materials Science. Proceedings of an International Symposium, Nieuwpoort, September 13-17, 1987 edited by P.J. Grobet, W.J. Mortier, E.F. Vansant and G. Schulz-Ekloff Catalysis 1987. Proceedings ofthe 10th North American Meeting ofthe Catalysis Society, San Diego, CA, May 17-22, 1987 edited by J.W. Ward Characterization of Porous Solids. Proceedings of the IUPAC Symposium (COPS I), Bad Soden a. "Is., April 26-29,1987 edited by K.K. Unger, J. Rouquerol, K.S.W. Sing and H. Kral Physics of Solid Surfaces 1987. Proceedings of the Fourth Symposium on Surface Physics, Bechyne Castle, September 7-11, 1987 edited by J. Koukal Heterogeneous Catalysis and Fine Chemicals. Proceedings of an International Symposium, Poitiers, March 15-17, 1988 edited by M. Guisnet, J. Barrault, C. Bouchoule, D. Duprez, C. Montassier and G. Perot Laboratory Studies of Heterogeneous Catalytic Processes by E.G. Christoffel, revised and edited by Z. Paal Catalytic Processes under Unsteady-State Conditions by Yu. Sh. Matros Successful Design of Catalysts. Future Requirements and Development. Proceedings of the Worldwide Catalysis Seminars, July, 1988, on the Occasion of the 30th Anniversary of the Catalysis Society of Japan edited by T. Inui Transition Metal Oxides. Surface Chemistry and Catalysis by H.H. Kung Zeolites as Catalysts, Sorbents and Detergent Builders. Applications and Innovations. Proceedings of an International Symposium, Wfirzburg, September 4-8,1988 edited by H.G. Karge and J. Weitkamp Photochemistry on Solid Surfaces edited by M. Anpo and T. Matsuura Structure and Reactivity of Surfaces. Proceedings of a European Conference, Trieste, September 13-16, 1988 edited by C. Morterra, A. Zecchina and G. Costa Zeolites: Facts, Figures, Future. Proceedings of the 8th International Zeolite Conference, Amsterdam, July 10-14, 1989. Parts A and B edited by P.A. Jacobs and R.A. van Santen Hydrotreating Catalysts. Preparation, Characterization and performance. Proceedings ofthe Annual International AIChE Meeting, Washington, DC, November 27-December 2, 1988 edited by M.L. Occelli and R.G. Anthony New Solid Acids and Bases. Their Catalytic Properties by K. Tanabe, M. Misono, u Ono and H. Hattori Recent Advances in Zeolite Science. Proceedings of the 1989 Meeting of the British Zeolite Association, Cambridge, April 17-19, 1989 edited by J. Klinowsky and P.J. Barrie Catalyst in Petroleum Refining 1989. Proceedings of the First International Conference on Catalysts in Petroleum Refining, Kuwait, March 5-8, 1989 edited by D.L. Trimm, S. Akashah, M. Absi-Halabi and A. Bishara Future Opportunities in Catalytic and Separation Technology edited by M. Misono, u Moro-oka and S. Kimura

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New Developments in Selective Oxidation. Proceedings of an International Symposium, Rimini, Italy, September 18-22, 1989 edited by G. Centi and F. Trifiro Volume 56 Olefin Polymerization Catalysts. Proceedings of the International Symposium on Recent Developments in Olefin Polymerization Catalysts, Tokyo, October 23-25, 1989 edited by T. Keii and K. Soga Volume 57A Spectroscopic Analysis of Heterogeneous Catalysts. Part A: Methods of Surface Analysis edited by J.L.G. Fierro Volume 57B Spectroscopic Analysis of Heterogeneous Catalysts. Part B: Chemisorption of Probe Molecules edited by J.L.G. Fierro Volume 58 Introduction to Zeolite Science and Practice edited by H. van Bekkum, E.M. Flanigen and J.C. Jansen Volume 59 Heterogeneous Catalysis and Fine Chemicals II. Proceedings of the 2nd International Symposium, Poitiers, October 2-6, 1990 edited by M. Guisnet, J. Barrault, C. Bouchoule, D. Duprez, G. Perot, R. Maurel and C. Montassier Volume 60 Chemistry of Microporous Crystals. Proceedings of the International Symposium on Chemistry of Microporous Crystals, Tokyo, June 26-29, 1990 edited by T. Inui, S. Namba and T. Tatsumi Natural Gas Conversion. Proceedings of the Symposium on Natural Gas Volume 61 Conversion, Oslo, August 12-17, 1990 edited by A. Holmen, K.-J. Jens and S. Kolboe Volume 62 Characterization of Porous Solids II. Proceedings of the IUPAC Symposium (COPS II), Alicante, May 6-9, 1990 edited by F. Rodriguez-Reinoso, J. Rouquerol, K.S.W. Sing and K.K. Unger Volume 63 Preparation of Catalysts V. Scientific Bases for the Preparation of Heterogeneous Catalysts. Proceedings of the Fifth International Symposium, Louvain-la-Neuve, September 3-6, 1990 edited by G. Poncelet, P.A. Jacobs, P. Grange and B. Delmon Volume 64 New Trends in CO Activation edited by L. Guczi Volume 65 Catalysis and Adsorption by Zeolites. Proceedings of ZEOCAT 90, Leipzig, August 20-23, 1990 edited by G. (~hlmann, H. Pfeifer and R. Fricke Dioxygen Activation and Homogeneous Catalytic Oxidation. Proceedings of the Volume 66 Fourth International Symposium on Dioxygen Activation and Homogeneous Catalytic Oxidation, Balatonf~red, September 10-14, 1990 edited by L.I. Simandi Structure-Activity and Selectivity Relationships in Heterogeneous Catalysis. Volume 67 Proceedings of the ACS Symposium on Structure-Activity Relationships in Heterogeneous Catalysis, Boston, MA, April 22-27, 1990 edited by R.K. Grasselli and A.W. Sleight Catalyst Deactivation 1991. Proceedings of the Fifth International Symposium, Volume 68 Evanston, IL, June 24-26, 1991 edited by C.H. Bartholomew and J.B. Butt Volume 69 Zeolite Chemistry and Catalysis. Proceedings of an International Symposium, Prague, Czechoslovakia, September 8-13, 1991 edited by RA. Jacobs, N.I. Jaeger, L. Kubelkovd and B. Wichterlova Volume 70 Poisoning and Promotion in Catalysis based on Surface Science Concepts and Experiments by M. Kiskinova

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Catalysis and Automotive Pollution Control II. Proceedings of the 2nd International Symposium (CAPoC 2), Brussels, Belgium, September 10-13, 1990 edited by A. Crucq New Developments in Selective Oxidation by Heterogeneous Catalysis. Proceedings of the 3rd European Workshop Meeting on New Developments in Selective Oxidation by Heterogeneous Catalysis, Louvain-la-Neuve, Belgium, April 8-10, 1991 edited by R Ruiz and B. Delmon Progress in Catalysis. Proceedings of the 12th Canadian Symposium on Catalysis, Banff, Alberta, Canada, May 25-28, 1992 edited by K.J. Smith and E.C. Sanford Angle-Resolved Photoemission. Theory and Current Applications edited by S.D. Kevan New Frontiers in Catalysis, Parts A-C. Proceedings of the 10th International Congress on Catalysis, Budapest, Hungary, 19-24 July, 1992 edited by L. Guczi, F. Solymosi and R Tetenyi Fluid Catalytic Cracking: Science and Technology edited by J.S. Magee and M.M. Mitchell, Jr. New Aspects of Spillover Effect in Catalysis. For Development of Highly Active Catalysts. Proceedings ofthe Third International Conference on Spillover, Kyoto, Japan, August 17-20, 1993 edited by T. Inui, K. Fujimoto, T. Uchijima and M. Masai Heterogeneous Catalysis and Fine Chemicals III. Proceedings ofthe 3rd International Symposium, Poitiers, April 5 - 8, 1993 edited by M. Guisnet, J. Barbier, J. Barrault, C. Bouchoule, D. Duprez, G. Perot and C. Montassier Catalysis: An Integrated Approach to Homogeneous, Heterogeneous and Industrial Catalysis edited by J.A. Moulijn, P.W.N.M. van Leeuwen and R.A. van Santen Fundamentals of Adsorption. Proceedings of the Fourth International Conference on Fundamentals of Adsorption, Kyoto, Japan, May 17-22, 1992 edited by M. Suzuki Natural Gas Conversion II. Proceedingsof the Third Natural Gas Conversion Symposium, Sydney, July 4-9, 1993 edited by H.E. Curry-Hyde and R.F. Howe New Developments in Selective Oxidation II. Proceedingsof the Second World Congress and Fourth European Workshop Meeting, Benalmadena, Spain, September 20-24, 1993 edited by V. Cortes Corberan and S. Vic Bellon Zeolites and Microporous Crystals. Proceedings of the International Symposium on Zeolites and Microporous Crystals, Nagoya, Japan, August 22-25, 1993 edited by T. Hattori and T. Yashima Zeolites and Related Microporous Materials: State of the Art 1994. Proceedings of the 10th International Zeolite Conference, Garmisch-Partenkirchen, Germany, July 17-22, 1994 edited by J. Weitkamp, H.G. Karge, H. Pfeifer and W. H61derich Advanced Zeolite Science and Applications edited by J.C. Jansen, M. St6cker, H.G. Karge and J.Weitkamp Oscillating Heterogeneous Catalytic Systems by M.M. Slin'ko and N.I. Jaeger Characterization of Porous Solids II1. Proceedings of the IUPAC Symposium (COPS III), Marseille, France, May 9-12, 1993 edited by J.Rouquerol, F. Rodriguez-Reinoso, K.S.W. Sing and K.K. Unger

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Catalyst Deactivation 1994. Proceedings of the 6th International Symposium, Ostend, Belgium, October 3-5, 1994 edited by B. Delmon and G.F. Froment Catalyst Design for Tailor-made Polyolefins. Proceedings of the International Symposium on Catalyst Design for Tailor-made Polyolefins, Kanazawa, Japan, March 10-12, 1994 edited by K. Soga and M. Terano Acid-Base Catalysis II. Proceedings of the International Symposium on Acid-Base Catalysis II, Sapporo, Japan, December 2-4, 1993 edited by H. Hattori, M. Misono and Y. Ono Preparation of Catalysts VI. Scientific Bases for the Preparation of Heterogeneous Catalysts. Proceedings of the Sixth International Symposium, Louvain-La-Neuve, September 5-8, 1994 edited by G. Poncelet, J. Martens, B. Delmon, P.A. Jacobs and R Grange Science and Technology in Catalysis 1994. Proceedings of the Second Tokyo Conference on Advanced Catalytic Science and Technology, Tokyo, August 21-26, 1994 edited by Y. Izumi, H. Arai and M. Iwamoto Characterization and Chemical Modification of the Silica Surface by E.F.Vansant, R Van Der Voort and K.C. Vrancken Catalysis by Microporous Materials. Proceedings of ZEOCAT'95, Szombathely, Hungary, July 9-13, 1995 edited by H.K. Beyer, H.G.Karge, I. Kiricsi and J.B. Nagy Catalysis by Metals and Alloys by V. Ponec and G.C. Bond Catalysis and Automotive Pollution Control III. Proceedings of the Third International Symposium (CAPoC3), Brussels, Belgium, April 20-22, 1994 edited by A. Frennet and J.-M. Bastin Zeolites: A Refined Tool for Designing Catalytic Sites. Proceedings of the International Symposium, Quebec, Canada, October 15-20, 1995 edited by L. Bonneviot and S. Kaliaguine Zeolite Science 1994: Recent Progress and Discussions. Supplementary Materials to the 10th International Zeolite Conference, Garmisch-Partenkirchen, Germany, July 17-22, 1994 edited by H.G. Karge and J. Weitkamp Adsorption on New and Modified Inorganic Sorbents edited by A. Dqbrowski and V.A. Tertykh Catalysts in Petroleum Refining and Petrochemical Industries 1995. Proceedings of the 2nd International Conference on Catalysts in Petroleum Refining and Petrochemical Industries, Kuwait, April 22-26, 1995 edited by M. Absi-Halabi, J. Beshara, H. Qabazard and Ao Stanislaus 1lth International Congress on Catalysis - 40th Anniversary. Proceedings ofthe 1lth ICC, Baltimore, MD, USA, June 30-July 5, 1996 edited by J. W. Hightower, W.N. Delgass, E. Iglesia and A.T. Bell Recent Advances and New Horizons in Zeolite Science and Technology edited by H. Chon, S.I. Woo and S. -E. Park Semiconductor Nanoclusters - Physical, Chemical, and Catalytic Aspects edited by RV. Kamat and D. Meisel Equilibria and Dynamics of Gas Adsorption on Heterogeneous Solid Surfaces edited by W. Rudzinski, W.A. Steele and G. Zgrablich Progress in Zeolite and Microporous Materials Proceedings of the 1lth International Zeolite Conference, Seoul, Korea, August 12-17, 1996 edited by H. Chon, S.-K. Ihm and Y.S. Uh

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Hydrotreatment and Hydrocracking of Oil Fractions Proceedings ofthe 1st International Symposium / 6th European Workshop, Oostende, Belgium, February 17-19, 1997 edited by G.F. Froment, B. Delmon and P. Grange Natural Gas Conversion IV Proceedings of the 4th International Natural Gas Conversion Symposium, Kruger Park, South Africa, November 19-23, 1995 edited by M. de Pontes, R.L. Espinoza, C.P. Nicolaides, J.H. Scholtz and M.S. Scurrell Heterogeneous Catalysis and Fine Chemicals IV Proceedings of the 4th International Symposium on Heterogeneous Catalysis and Fine Chemicals, Basel, Switzerland, September 8-12, 1996 edited by H.U. Blaser, A. Baiker and R. Prins Dynamics of Surfaces and Reaction Kinetics in Heterogeneous Catalysis. Proceedings ofthe International Symposium, Antwerp, Belgium, September 15-17,1997 edited by G.F. Froment and K.C. Waugh Third World Congress on Oxidation Catalysis. Proceedings of the Third World Congress on Oxidation Catalysis, San Diego, CA, U.S.A. 21-26, September 1997 edited by R.K. Grasselli, S.T. Oyama, A.M. Gaffney and J.E. Lyons Catalyst Deactivation 1997. Proceedings ofthe 7th International Symposium, Cancun, Mexico, October 5-8, 1997 edited by C.H. Bartholomew and G.A. Fuentes

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