Fluid Catalytic Cracking V

Fluid Catalytic Cracking V Materials and Technological Innovations by Occelli M. L. Occelli, P. O'Connor

• ISBN: 0444504753 • Publisher: Elsevier Science • Pub. Date: April 2001

vii

List of Contributors

S. A1-Khattaf Chemical Reactor Engineering Centre Faculty of Engineering University of Western Ontario London, Ontario CANADA N6A 5B9

H. de Lasa Chemical Reactor Engineering Centre Faculty of Engineering University of Western Ontario London, Ontario CANADA N6A 5B9

S.-I. Andersson Chalmers University of Technology Department of Applied Surface Chemistry SE-41296 Gothenburg SWEDEN

M.A. den Hollander Industrial Catalysis Department of Chemical Technology Faculty of Applied Sciences Delft University of Technology Julianalaan 136 2628 BL Delft THE NETHERLANDS

A. Auroux Institut de Recherches sur la Catalyse CNRS 2 Av. A. Einstein 69626 Villeurbanne FRANCE R.A. Beyerlein National Institute of Standards and Technology Gaithersburg, MD 20899-4730 USA L.T. Boock Grace Davision 7500 Grace Drive Columbia, MD 21044 USA M. Castro Diaz University of Strathclyde Department of Pure and Applied Chemistry Glasgow G1 1XL Scotland UK A. Corma Instituto de Tecnologfa Qufmica UPV-CSIC Avda. de los Naranjos, s/n 46022 Valencia SPAIN

H. Eckert Institut ftir Physikalische Chemie Westf~ilische Wilhelms-Universit~it Miinster Schlossplatz 7 D-48149 Miinster GERMANY I. Eilos Fortum Oyj P.O. Box 310 06101 Porvoo FINLAND A.E. Fallick Scottish Universities Research & Reactor Centre East Kilbride Glasgow G75 0QU UK J. Frasch Laboratoire de Matrriaux Minrraux (CNRSENSCMu) 3 rue Alfred Wemer F-68093 Mulhouse FRANCE

viii

R. Garcia-de-Le6n Programa de Investigaci6n en Tratiamento de Crudo Maya Instituto Mexicano del Petr61eo Ejo Central L~aro C&denas 152 C.P. 07730 M6xico, D.F. MEXICO W.R. Gilbert PETROBAS R & D Center Process Division Rio de Janeiro, 21949-900 BRAZIL R. Gonz~ilez-Serrano Programa de Investigaci6n en Tratiamento de Crudo Maya Instituto Mexicano del Petr61eo Ejo Central L~aro C&denas 152 C.P. 07730 M6xico, D.F. MEXICO M.-Y. Gu Research Institute of Petroleum Processing (RIPP) SINOPEC Beijing CHINA N.J. Gudde BP, Oil Technology Centre Chertsey Road Sunbury-on-Thames Middlesex TW16 7LN UK P. Gullbrand Instituto de Tecnologfa Qufmica UPV-CSIC Avda. de los Naranjos, sin 46022 Valencia SPAIN P. Hagelberg Fortum Oyj P.O. Box 310 06101 Porvoo FINLAND

P.J. Hall University of Strathclyde Department of Pure and Applied Chemistry Glasgow G1 1XL Scotland UK M. He Research Institute of Petroleum Processing China Petrochemical Corporation Beijing 100083 P.R. CHINA F. Hern~dez-Belmin Programa de Investigaci6n en Tratiamento de Crudo Maya Instituto Mexicano del Petr61eo Ejo Central L~aro C~denas 152 C.P. 07730 M6xico, D.F. MEXICO J. Hiltunen

Fortum Oyj P.O. Box 310 06101 Porvoo FINLAND R. Hughes University of Salford Chemical Engineering Unit Salford M5 4WT UK A. Humphries Akzo Nobel Catalysts Inc. 2625 Bay Area Blvd., Suite 250 Houston, TX 77058 USA P. Imhof Akzo Nobel Catalysts Research Center Catalysts Amsterdam THE NETHERLANDS K. J/i/iskel/iinen Fortum Oyj P.O. Box 20 00048 Fortum FINLAND

R. Jonker Akzo Nobel Catalysts Research Center Catalysts Amsterdam THE NETHERLANDS M. Kalwei Institut fiir Physikalische Chemie Westf~ilische Wilhelms-Universit~it Miinster Schlossplatz 7 D-48149 Miinster GERMANY S. Katoh Kashima Oil Company Kashima JAPAN G.W. Ketley BP, Oil Technology Centre Chertsey Road Sunbury-on-Thames Middlesex TW16 7LN UK P. Knuuttila Fortum Oyj P.O. Box 310 06101 Porvoo FINLAND C.L. (Arthur) Koon University of Salford Chemical Engineering Unit Salford M5 4WT UK C.W. Kuehler Akzo Nobel Catalysts Houston, Texas USA A.A. Lappas Chemical Process Engineering Research Institut (CPERI) Department of Chemical Engineering University of Thessaloniki P.O. Box 361 57001 Thermi, Thessaloniki GREECE

B. Lebeau Laboratoire de Mat6riaux Min6raux (CNRS-ENSCMu) 3 rue Alfred Werner F-68093 Mulhouse FRANCE M.I. Levinbuk Gubkin Moscow Oil and Gas University 65 Leninsky prosp. Moscow 117917 THE RUSSIAN FEDERATION C.-Y. Li Research Institute of Petroleum Processing (RIPP) SINOPEC Beijing CHINA K. Lipiainen Fortum Oyj P.O. Box 310 06101 Porvoo FINLAND F. L6pez-Isunza Departamento de Ingenierfa de Procesos e Hidrfiulica Universidad Aut6noma MetropolitanaIztapalapa Av. Michoacfin y La Purisima sin Col. Vicentina Iztapalapa M6xico 09340, D.F. MEXICO E. L6pez-Salinas Programa de Investigaci6n en Tratiamento de Crudo Maya Instituto Mexicano del Petr61eo Ejo Central Lfizaro Cfirdenas 152 C.P. 07730 M6xico, D.F. MEXICO Y. Lu Research Institute of Petroleum Processing China Petrochemical Corporation Beijing 100083 P.R. CHINA

J. Majander Fortum Oyj P.O. Box 20 00048 Fortum FINLAND M. Makkee Industrial Catalysis Department of Chemical Technology Faculty of Applied Sciences Delft University of Technology Julianalaan 136 2628 BL Delft THE NETHERLANDS S.C. Martin

University of Strathclyde Department of Pure and Applied Chemistry Glasgow G1 1XL Scotland UK

J.C. Moreno-Mayorga Programa de Investigaci6n en Tratiamento de Crudo Maya Instituto Mexicano del Petr61eo Ejo Central L~aro C~irdenas 152 C.P. 07730 M6xico, D.F. MEXICO

J.A. Moulijn Industrial Catalysis Department of Chemical Technology Faculty of Applied Sciences Delft University of Technology Julianalaan 136 2628 BL Delft THE NETHERLANDS T. Myrstad Statoil's Research Centre N-7005 Trondheim NORWAY

C. Martfnez Instituto de Tecnologfa Quimica UPV-CSIC Avda. de los Naranjos, s/n 46022 Valencia SPAIN

M. Nakamura Nippon Ketjen Tokyo JAPAN

G.B. McVicker ExxonMobil Research & Engineering Co. Annandale, NJ 08801 USA

V.M. Niemi Fortum Oyj P.O. Box 310 06101 Porvoo FINLAND

V.B. Melnikov Gubkin Moscow Oil and Gas University 65 Leninsky prosp. Moscow 117917 THE RUSSIAN FEDERATION

S. Numan Gubkin Moscow Oil and Gas University 65 Leninsky prosp. Moscow 117917 THE RUSSIAN FEDERATION

E. Mogica-Martfnez Programa de Investigaci6n en Tratiamento de Crudo Maya Instituto Mexicano del Petr61eo Ejo Central L~aro C~denas 152 C.P. 07730 M6xico, D.F. MEXICO

M.L. Occelli MLO Consulting Atlanta, GA 30328 USA P. O'Connor Akzo Nobel Catalysts Amersfoort THE NETHERLANDS

J. Patarin Laboratoire de Mat6riaux Min6raux (CNRS-ENSCMu) 3 rue Alfred Wemer F-68093 Mulhouse FRANCE V.A. Patrikeev Salavat Catalyst Factory Salavat 453206 THE RUSSIAN FEDERATION M.L. Pavlov Ishimbai Catalyst Factory Ishimbai 453210 THE RUSSIAN FEDERATION A. Petre Institut de Recherches sur la Catalyse CNRS 2 Av. A. Einstein 69626 Villeurbanne FRANCE T.F. Petti

Grace Davision 7500 Grace Drive Columbia, MD 21044 USA Z.-H. Qiu Research Institute of Petroleum Processing (RIPP) SINOPEC Beijing CHINA J. R6pp~inen Fortum Oyj P.O. Box 20 00048 Fortum FINLAND W.L. Schuette (deceased) A.E. Schweizer ExxonMobil Refining and Supply Company Process Research Laboratories P.O. Box 2226 Baton Rouge, LA 70821-2226 USA

X. Shu Research Institute of Petroleum Processing China Petrochemical Corporation Beijing 100083 P.R. CHINA B. Skocpol Akzo Nobel Catalysts Amersfoort THE NETHERLANDS C.E. Snape University of Strathclyde Department of Pure and Applied Chemistry Glasgow G1 1XL Scotland UK J. Song Research Institute of Petroleum Processing China Petrochemical Corporation Beijing 100083 P.R. CHINA M. Soulard Laboratoire de Mat6riaux Min6raux (CNRS-ENSCMu) 3 rue Alfred Werner F-68093 Mulhouse FRANCE L.-W. Tang Research Institute of Petroleum Processing (RIPP) SINOPEC Beijing CHINA Z.A. Tsagrasouli Chemical Process Engineering Research Institut (CPERI) Department of Chemical Engineering University of Thessaloniki P.O. Box 361 57001 Thermi, Thessaloniki GREECE Y.R. Tyagi University of Strathclyde Department of Pure and Applied Chemistry Glasgow G1 1XL, Scotland

xii I.A. Vasalos Chemical Process Engineering Research Institute (CPERI) Department of Chemical Engineering University of Thessaloniki P.O. Box 361 57001 Thermi, Thessaloniki GREECE C.L. Wallace University of Strathclyde Department of Pure and Applied Chemistry Glasgow G1 1XL Scotland UK A. Wrlker Institut fiir Physikalische Chemie Westf/ilische Wilhelms-Universi~t Miinster Schlossplatz 7 D-48149 Miinster GERMANY S.-H. Yan Research Institute of Petroleum Processing (RIPP) SINOPEC Beijing CHINA S.J. Yanik Akzo Nobel Catalysts Singapore

Preface Catalyst production for the transformation of crudes into gasoline and other fuel products is a $2.1 billion/year business and fluid cracking catalysts (FCCs) represent almost half of the refinery catalyst market (M. MacCoy, Chemical and Engineering News, p. 17, September 20 (1999)). During the cracking reactions, the FCC surface is contaminated by metals (Ni, V, Fe, Cu, Na) and by coke deposition. As a result, the catalyst activity and product selectivity is reduced to unacceptable levels thus forcing refiners to replace part of the recirculating equilibrium FCC inventory with fresh FCC to compensate for losses in catalyst performance. About 1,100 tons/day of FCC are used worldwide in over 200 fluid cracking catalyst units (FCCUs). Today, the worldwide capacity to produce gasoline exceeds the 460 million gal/day. In addition, FCCs are used in the commercial synthesis of acrylonitrile, phthalic anhydride and maleic anhydridie and in the production of 45% of the world propylene (Chemical and Engineering News, p. 51, November 23, 1998). In recognition of the great technological importance of the FCC process, on November 3, 1998 the first commercial fluid bed reactor using catalytic cracking constructed at the Exxon Baton Rouge refinery, was designated a National Historic Chemical Landmark by the American Chemical Society. It is for these reasons that refiners' interest in FCC research has remained high through the years and almost independent of crude oil prices. However, recent oil company mergers and the dissolution of research laboratory, has drastically decreased the number of researchers involved in petroleum refining research projects. As a results the emphasis has shifted from new materials research to process improvements and this trend is clearly reflected in the type of papers contained in this volume. Modem spectroscopic techniques continue to be essential to the understanding of catalysts performance and several chapters in the book describe the use of 27A1, 29Si and ~3C NMR to study variation in FCC acidity during aging and coke deposition. In addition several chapters have been dedicated to the modeling of FCC deactivation, and to the understanding of contact times on FCC performance. Refiners efforts to conform with environmental regulations are reflected in chapters dealing with sulfur removal, metals contaminants and olefins generation In conclusion, as before we would like to express our gratitude to our colleagues for acting as technical referees. The views and conclusion expressed herein are those of the chapter authors whom we sincerely thanks for their time and effort in presenting their research at the Symposium and in preparing the camera ready manuscripts for this Volume. Mario L. Occelli and Paul O'Connor November 2000

Table of Contents List of Contributors Preface

1

Defect structure and acid catalysis of high silica, FAU-framework

1

zeolites: effects of aluminum removal and of basic metal oxide

3

addition The use of microcalorimetry and solid state nuclear magnetic resonance (NMR) to study the effects of post-synthesis treatments

2

41 on the acidity and framework composition of several HY-types zeolites The effects of steam aging temperature on the properties of an HY

3

59 zeolite of the type used in FCC perparations Effect of catalyst properties and feedstock composition on the

4

71 evaluation of cracking catalysts Study on the deactivation-aging patterns of fluid cracking

5

87 catalysts in industrial units

The improvement of catalytic cracking process through the

6

107 utilization of new catalytic materials

7

NExCC-Novel short contact time catalytic cracking technology

111

8

Effect of vanadium on light olefins selectivity

133

9

Reduction of olefins in FCC gasoline

141

Gasoline sulfur removal: kinetics of S compounds in FCC

10

153 conditions Development of a kinetic model for FCC valid from ultra-short

11

167 residence times Deactivation of fluid catalytic cracking catalysts: a modelling

12

187 approach Catalyst design for resin cracking operation: benefits of metal

13

201 tolerant technologies

14

Active site accessibility of resid cracking catalysts

209

Catalyst evaluation for atmospheric residue cracking, the effect of

15

219 catalyst deactivation on selectivity

16

Optimum properties of RFCC catalysts

227

An experimental protocol to evaluate FCC stripper performance in

17

239 terms of coke yield and composition Use of [superscript 13]C-labelled compounds to probe catalytic

18

251 coke formation in fluid catalytic cracking

19

Bifunctionality in catalytic cracking catalysis

263

Catalytic cracking of alkylbenzenes. Y-zeolites with different

20

279 crystal sizes On the mechanism of formation of organized mesoporous silica

21

293 that may be used as catalysts for FCC Catalyst assembly technology in FCC. Part I: A review of the

22

299 concept, history and developments Catalyst assembly technology in FCC. Part II: The influence of

23

fresh and contaminant-affected catalyst structure on FCC

311

performance Keyword Index

333

Studies in Surface Science and Catalysis 134 M.L. Occelli and P. O'Conner (Editors) 9 2001 Elsevier Science B.V. All rights reserved

Defect Structure and Acid Catalysis of High Silica, F A U - F r a m e w o r k Zeolites: Effects of A l u m i n u m R e m o v a l and of Basic Metal Oxide Addition Robert A. Beyerlein* and Gary B. McVicker* *National Institute of Standards and Technology, Gaithersburg, MD 20899-4730 *ExxonMobil Research & Engineering Co., Annandale, NJ 08801 The catalytic properties of ultrastable Y (USY) are directly influenced by the zeolite destruction that occurs during formation of USY and during subsequent hydrothermal treatment. Mildly steamed USY materials exhibit a secondary pore system (mesopores) of 550 nm dimensions, which are evident as light amorphous zones in Transmission Electron Microscopy (TEM). Combined high resolution electron microscopy (HREM) and analytical electron microscopy (AEM) investigations on hydrothermally deaiuminated USY materials have shown that, in regions of high defect concentration, mesopores "coalesce" to form channels and cracks, which, upon extended hydrothermal treatment, define the boundaries of fractured crystallite fragments. The predominant fate of aluminum ejected from lattice sites appears to be closely associated with dark bands, which decorate the newly formed fracture boundaries. A smaller proportion of ejected aluminum exists as "nonframework AI" within the zeolite cages. High silica Y materials, having little or no nonframework A1 exhibit poor catalytic activity for a large variety of acidity-dependent reactions. Investigations on mildly dealuminated zeolites suggest that the origin of the enhanced catalytic activity is a synergistic interaction between Br0nsted (framework) and highly dispersed Lewis (nonframework) acid sites. The enhanced cracking, isomerization activity associated with the presence of highly dispersed nonframework A1 species i) is not reflected in direct measures of solid acidity obtained, for example, by calorimetry or by NMR spectroscopy, and ii) is not consistent with a major increase in average acid site strength. Numerous structure/function studies indicate that the critical nonframework A1 species may exist as cationic species in the small cages of dealuminated H-Y. By contrast, partial exchange of high silica Y materials with monovalent cations, such as Na or K, leads to significant reduction in activity, presumably by poisoning acid sites. In prior studies of isobutane conversion over dealuminated H-Y, it was shown that the addition of sodium equivalent to 1/3 of the total framework A1 atoms completely eliminates catalyst activity. Extensions of these poisoning studies show that addition of potassium produces a much stronger poisoning effect, with one K + ion giving an activity suppression roughly equivalent to that produced by two Na § ions. Calcium addition gives rise to a poisoning effect intermediate between those of Na + and K + at low levels of exchange, ca. 10%, but is more mild than that of sodium as Ca ++ exchange levels exceed 20%. Previous correlations of isobutane conversion activity with framework composition support a direct dependence of carbocation-facilitated processes on framework aluminum (A1F), with a linear dependence of carbonium ion rates on A1F content. The observed linear

dependencies exhibited for Na or K addition show that the primary effect of poisoning, or of A1F removal, is a decrease in the number of active sites. Measured selectivities for carbocation products indicate a limiting site density of about A1F/ucell - 8 (out of a maximum 56 A1 among 192 tetrahedral framework sites for a starting zeolite Y), below which carbocation activity diminishes rapidly. Consistent with previous discussion of dual mechanisms, the results for formation of methane, a stable reaction product marker, show that the initiation step and the secondary carbocation processes are intimately linked over the entire range of acid site content, whether manipulated by dealumination or by permanent poisoning by basic alkali or alkaline metal oxides.

1. INTRODUCTION The importance of acid catalysis for the production of fuel and petrochemicals is underscored by recent environmental mandates calling for reformulated motor fuel that contains greater proportions of high octane, branched paraffins and oxygenates. Environmental concerns about the catalysts themselves, particularly the highly corrosive and toxic liquid acids, such as sulfuric and hydrofluoric acids, have created a need for stable, strongly acidic solid acids. Combined theoretical and experimental studies of the last decade have substantially improved our level of understanding of solid acidity in zeolites. The prospect for obtaining a detailed molecular level understanding of heterogeneous catalysts that could better guide the search for improved catalysts appears to be optimum for crystalline solid acids. It is the object of this paper to review our current understanding of the predominant solid acid catalyst, the family of protonated FAU-framework materials stabilized by hydrothermal treatment, originally designated ultrastable Y (USY) [1, 2] and commonly referred to as dealuminated H-Y (H-ultrastable Y). The catalytic properties of ultrastable Y are directly influenced by the zeolite destruction attending its formation and further modification by subsequent hydrothermal treatment. For ultrastable, high silica, FAU framework materials prepared by steam dealumination, interpretation of catalytic data is complicated by the presence of entrained, nonframework aluminum (NFA) species. Although the individual and collective roles of framework and nonframework aluminum species are not well understood, it is clear that the presence of some nonframework A1, presumably highly dispersed, is essential for the strong solid acidity exhibited by high silica H-Y [3-5]. While the critical nonframework species are not easily subject to direct observation, the existence of isolated, intracrystalline NFA species in dealuminated H-Y materials is not in doubt. The importance of certain nonframework A1 species for the ability of zeolitic solid acids to catalyze acidity-demanding reactions [6], such as alkane skeletal isomerization or cracking, is not limited to H-Y. An abundant literature shows a consensus that the development of "enhanced carbocation activity" in mildly steamed HZSM-5 is also critically dependent on

the presence of nonframework A1 [6-9]. Enhancement of carbocation activity, generally associated with BrCnsted acidity, has also been observed in mildly steamed mordenite [10] and in HZSM-20 [ 11 ]. Knowledge of framework geometry is essential for understanding overall reactivity patterns for hydrocarbon conversions over these open framework solid acids. The well-known features of "molecular traffic management" exhibited by these materials are not always limited to molecular sieving, that is reactant or product size exclusion effects. For example, at reaction temperatures of 400 ~ to 500~ dealuminated H-USY and dealuminated mordenite (large pore zeolites)each catalyze the isomerization of isobutane to n-butane [12, 13]. Under similar conditions, the medium pore system HZSM-5 produces relatively little n-butane, but instead yields much methane and propylene [13, 14]. The dramatically different product selectivities in the latter case are attributed to the more severe spatial restrictions of the medium pore ZSM-5, which tend to inhibit hydride transfer and oligomerization/backcracking processes involving bulky reaction intermediates. Ever since the rapid commercialization in the early 1960's of a zeolite-catalyzed process for gas oil cracking [15, 16], zeolites have comprised the predominant usage of solid acid catalysts. The characterization and application of high silica, protonated zeolites in fluid catalytic cracking has been reviewed by Scherzer [17]. A broader overview of the use of zeolites in hydrocarbon processing is found in Maxwell and Stork [ 18]. A large number of potential zeolite catalyst applications in the synthesis of intermediates and fine chemicals have been discussed by Hoelderich et al. [19-21]. An outstanding attribute of the acidic FAU H-Y materials is their ability to catalyze intermolecular hydride transfer reactions in numerous hydrocarbon conversions, which are all but missing in less strongly acidic amorphous solid acids. Despite high industrial and academic interest, the nature of the active site in solid acids remains largely unresolved. In systems of wide interest such as a protonated zeolite, a chlorided or fluorided alumina, or sulfated zirconia, we are unable to quantify the distribution or relative importance of BrCnsted and/or Lewis sites, the surface acid strength, or the concentration of acid sites [22]. Recent advances in physical characterization of sites have substantially improved our understanding of these issues. Both 1H MASNMR [8, 23-25] and 13C MASNMR [26] have been effective in formalizing the structure/function relationships for BrCnsted acid sites in HZSM-5, an especially favorable system for analysis owing to its low BrCnsted acid site density and high crystallinity. So far, only "clean framework" ZSM-5 has been reasonably well characterized. The extension of these and related spectroscopic studies to the more complex systems represented by mildly steamed, carbocation activity enhanced HZSM-5 or dealuminated H-Y comprises a significant experimental and data interpretation challenge.

In the following sections, we review advances in the understanding of (i) the evolution of defect formation in dealuminated H-Y, (ii) the critical role played by nonframework aluminum in the acid catalysis exhibited by these materials and (iii) the dual effects of aluminum removal and of basic metal oxide addition on catalyst activity and selectivity. The discussion of acid catalysis is cast primarily in terms of kinetic methods for classifying solid acidity. As noted in a recent review by Haag [27], the question of acid strength is much more problematic and complicated. Theoretical efforts to elucidate structure-acidity relationships have made great progress, but are primarily limited to the "clean framework" case. Direct measures of solid acid strength within a homologous series of solid acids, e.g., ZSM-5 or dealuminated H-Y, including "in situ" calorimetric and/or spectroscopic methods, will often, but not always correlate with catalytic properties. 2. D E F E C T S P R O D U C E D BY H Y D R O T H E R M A L T R E A T M E N T 2.1. B a c k g r o u n d - F o r m a t i o n of Ultrastable Y

The formation of Ultrastable FAU materials may be viewed as a two-step process in which steam-calcination of the ammonium (or H) form of the Y-zeolite at the approximate conditions T=500~

Pa2o = 1 atm, for 2 hours, leads to the expulsion of A1 atoms from

framework T-sites as indicated in Figure 1. Following aluminum expulsion from framework positions, the resulting vacancies are to a great extent, refilled by silicon atoms migrating from

Two-step process: 1 Calcination of the ammonium (or H) form of the Y-zeolite at T > 500~ PH20< 1 atm ==~>Expulsion of AI atoms from framework T-sites Si I 0 I Si--O-AII 0 I Si

H+ 0 I O-Si-O I 0

Si I 0 H

+

3H20

b,-

Si-OH

HO-SI H 0 I Si

+

AI(OH)3

G.T: Kerr, 1979

2 Healing / Si replacement- Resulting vacancies are refilled to a large extent with Si atoms Limited concensus

on this m e c h o n i s m

Figure 1. Schematic representation of formation of ultrastable Y materials

the collapsed portions of the crystal. If this "healing" did not occur, the entire zeolite crystal would transform into a predominantly x-ray amorphous phase. The resulting restructured FAU material, originally designated Ultrastable Y or USY [1, 2], displays a contracted unit cell size and increasing hydrothermal stability as framework Si/A1 increases. A typical reduction in unit cell dimension is from 2.470 nm for the starting zeolite Y to 2.456 nm for the product USY. Increasingly more severe steam treatments result in a higher level of dealumination, a more contracted unit cell, and an increasing fraction of crystalline zeolite destruction. In certain chemical methods of dealumination, notably the ammonium hexafluorosilicate (AHF) method of Skeels and Breck [28], the silicon for "healing" of vacancies resulting from framework A1 removal originates from an external source. In contrast, during the formation of USY, the inserted silicon originates from an internal source (the zeolite itself) that results from damage to the zeolite framework with the concomitant formation of mesopores. The mechanism of silica transport and insertion into vacancies was first suggested by Maher, Hunter and Scherzer [29]. According to this mechanism, the silica required to fill the framework vacancies originates in those parts of the zeolite crystal which collapse during the hydrothermal treatment. The silica freed from the collapsed framework migrates under high temperature steam toward the tetrahedral vacancies of the remaining framework and, by filling them, increases framework stability [30]. Prior sorption studies tend to support this interpretation. Based on sorption studies on USY zeolites, Lohse et al. concluded that entire sodalite units, or [3-cages, are destroyed during hydrothermal dealumination [31, 32], leading to the formation of a secondary pore system [30] or mesoporosity in the range of 5 to 50 nm. As shown schematically in Figure 2, the collapse of sodalite units, or even ensembles of them, generates mesoporosity and simultaneously provides the source of Si atoms for healing the framework sites vacated by A1. Lattice destruction of hydrothermally dealuminated (or thermochemically treated) FAU materials, leading to the formation of an amorphous silica phase, may be identified in the 29Si MASNMR spectra by a broadened peak or shoulder at about-110 ppm (TMS) [33]. Recent

29Si and 27A1 MASNMR

and microcalorimetry

investigations of the aging and regeneration of Fluid Cracking Catalysts (FCC) have shed new light on mechanisms of healing and on overall silica losses [34a]. Results of these studies indicate that the products of hydrothermal transformation of the kaolin clay and the alumina/silica gel (components of the catalyst matrix of the composite FCC particle) represent possible sources of silica. Upon aging, each of these components generate penta-, tetrahedrally- and octahedrally-coordinated A1 species, and at the same time, Si compounds, which can contribute both to healing and to net silica losses from the FCC particle [34a,b]. Measurements of silica/alumina composition for a fresh catalyst and for the corresponding equilibrium catalyst from a commercial Fluid Cracking Unit show a dramatic drop in the bulk

Figure 2. Schematic model for mechanisms of dealumination in FAU, showing Si for "healing," and the source of mesoporosity.

SIO2/A1203 ratio from 3.6 for fresh to 2.5 for the used or "equilibrium" catalyst, indicating a substantial silica loss from the FCC particle during use [34b]. 2.2. Nonframework Aluminum- Local Environment

Nonframework aluminum (NFA) is a catchall description for a wide collection of defects that are produced during the formation of USY and during subsequent hydrothermal treatment. NFA species are themselves composed of several different types, some isolated, some agglomerated, as outlined in Table 1. It has proven to be particularly difficult to characterize the many different NFA species in dealuminated H-Y. Structural studies using X-ray and neutron diffraction have indicated the presence of octahedrally coordihated microcrystalline aluminum species in the supercages [35], and isolated tetrahedral aluminum species in the small (sodalite) cages [36], but do not give much information about agglomerated noncrystalline species. A systematic study of the reduction in micropore volume [37] resulting from mild dealumination of H-Y, and of HZSM-5, indicated that the majority of NFA species go to the micropores accessible to N2, the supercages (c~-cages) in the case of H-Y, and to the channels or channel intersections in the case of HZSM-5. The combination of high resolution 29Si and of 27A1 solid state NMR has been effectively applied to studies of

Table 1 Types of aluminum a in ultrastable Y Types detect- Probable structure able by NMR description

Isolated or clustered

Most abundant

TF

T-site

isolated

mildly dealuminated

TNFA

A1 in small cage and/or

isolated

mildly dealuminated

TNFA

clustered alumina species in S-C b and/or surface enrichment

severely dealuminated

PNFA

intermediate between octahedral and tetrahedral

unknown

severely dealuminated

ONFA

alumina species in s-c

clustered

always present

s-c = supercage

[from Ref. 61]

a

T = tetrahedral, 0 = octahedral

b

hydrothermally dealuminated Y zeolites [38]. The former provides direct information on the composition and Si, A1 distribution of the tetrahedral framework, independently of the presence of non-framework A1 species, while the latter allows distinction between tetrahedral framework A1 (-60 ppm) and octahedral non-framework A1 (-0 ppm). The interpretation of 27A1NMR in terms of A1 species location or state of agglomeration is often ambiguous. Even at low levels of dealumination, the contribution of nonframework species to the tetrahedral resonance cannot be ruled out. The fact that a wide range of transitional aluminas exhibit a ratio of tetrahedrally coordinated A1 over total A1 content of about 0.4 [39] illustrates the need for caution in attributing the tetrahedral resonance in zeolites to framework aluminum. Extensively dealuminated samples typically show substantial broadening of the tetrahedral resonance [40, 41], as shown in Figure 3, only a small portion of which can be attributed to framework A1 [41, 42]. In addition, as higher magnetic fields and faster sample spinning have become more routine, a new resonance has been observed at 30 ppm (Figure 3), which has been attributed to either an aluminum in a highly distorted tetrahedral environment [41-44] or a penta-coordinated aluminum species [44, 45]. Application of the novel double-rotation (DOR) spinning technique [46, 47] to the study of

27A1in zeolites [48] has shown two

different tetrahedral A1 species for a commercial USY material, one framework and the other nonframework [48]. On the basis of comparison of single pulse

27A1MASNMR and 27A1CP

MASNMR taken on a steamed Y zeolite, Fripiat and co-workers [49] concluded that a substantial portion of the band near 60 ppm is contributed by tetrahedrally-coordinated NFA.

10

4X

T Acid wash, 0.1 N HCI . Steam 650~ 4 hr

5X

T ~, It

"

]\

2X

A

Acid wash x 2, 0.1 N Steam650~ 3 hr, NH4§ exch.,acid wash,

o.o33..

A

Calcine 538~ 2 hr Steam 600~ 3 hr

l ~ j ~

T Parent USY (LZ-Y82)

1X 9 =+0

~5o

A z~

~oo"5'o . . . .

o:sO-~'oo"-l'sOPp",

~

Figure 3. Evolution of deauluminated USY as tracked by 27A1 MAS NMR [42, 61 ]. 9 framework AI" •, nonframework A1. 2.3. Nonframework A l u m i n u m and Mesoporosity

Previous transmission electron microscope (TEM) studies of hydrothermal aging of neat USY materials [50-54] and also of USY cracking catalysts [55-58] have shown 5 to 50 nm defect domains, which were attributed to mesopores. Such features, more pronounced in the presence of vanadium [55, 56], are characteristic of extended hydrothermal treatment. Typical porosity analyses of mildly steamed USY materials show a distribution of mesopore dimensions in the range 5 to 50 nm that is skewed toward the smaller sizes [5 l, 53], supporting the association of the light amorphous zones observed by TEM with the secondary pore system characteristic of USY materials [3 "., 32, 50-53]. A new understanding of the formation and evolution of mesopores has emerged from combined high resolution electron microscopy (HREM) and analytical electron microscopy (AEM) investigations on hydrothermally treated USY materials [42, 59]. In contrast with

11 results of previous TEM investigations [50, 51, 54], HREM and AEM studies of a steam/acid treated neat USY material, and of a high-temperature steam-treated USY cracking catalyst [42, 59], gave clear evidence for an inhomogeneous distribution of mesopores (Figure 4), which occurs concomitantly with further zeolite dealumination. Such inhomogeneities were found to be more pronounced for the (high temperature) steam-deactivated USY cracking catalyst than for the (moderate temperature) steam/acid-treated neat USY material. It was concluded that the extent of inhomogeneity is driven by non-equilibrium processes represented by accelerated steam-aging treatments in the laboratory. In regions with high defect concentration, mesopores "coalesce" to form channels and cracks (Figures 4, 5a, 6), which ultimately define the boundaries of fractured crystallite fragments. At these boundaries, a dark band is often observed which is highly enriched in aluminum (Figure 5), while within the mesopore, aluminum appears to be deficient (Figure 7) [59]. Such dark bands appear to have been observed in prior studies [53], but their presence was not discussed. The predominant fate of aluminum ejected from lattice sites appears to be closely associated with the dark bands, which often decorate these newly formed fracture

Figure 4. HREM image of several USY grains in the steamed USY catalyst showing an inhomogeneous distribution of mesopores. A coalescence of mesopores indicating an evolving fracture is indicated by an arrow [42].

12

Figure 5. (a) TEM image of a few steam/acid treated USY grains. An inhomogeneous distribution of mesopores is seen within individual grains; some grains contain more mesopores than others. In regions with high mesopore concentration, the pores coalesce to form channels (indicated by arrows). (b) HREM image of steam/acid treated USY grains. Many mesopores are formed. Although localized disorder is observed within the pores, the connecting regions remain crystalline [59].

13

Figure 6. HREM image of a steam/acid treated USY grain. Cracks (as indicated by arrows) are formed from the evolution of the coalesced mesopores. Dark bands, which were found to be A1 rich, are seen along these cracks [59].

Figure 7. A1, Si analyses from a STEM image of the region adjacent to and also within a mesopore. Within a mesopores, A1 is slightly deficient [59].

14 boundaries (Figure 6). These features were observed both for the steamed USY cracking catalyst (Fig. 8) and for the steam/acid treated neat USY zeolite (Figs. 5a, 6), consistent with results of previous studies, which found the surface enrichment of A1 to persist through aqueous treatments that remove substantial amounts of aluminum [60, 63]. For the steam/acid-treated neat USY material, the associated development and evolution of nonframework A1 species was investigated by high resolution solid state 27A1 MASNMR (Figure 3) [42, 61 ]. From this parallel study, it was concluded that the extracrystalline phases represented by the dark bands revealed in the electron microscopy studies contribute the majority of the nonframework aluminum species, tetrahedral, penta-coordinate, and octahedral, that were detected by 27A1NMR [61 ].

Figure 8. HREM image showing the fracturing of a USY grain in the steamed USY catalyst. Each fractured crystallite is bounded by cracks evolved from coalescence of mesopores. The dark band seen along each crack is A1 enriched, similar to those observed in the steam/acid treated neat USY material [59].

15

Similar patterns of defect formation were found upon re-investigation by HREM [61 ] of the age-separated FCU fractions from an earlier study [58]. An FCU "young" fraction (Figure 9) proved to be similar to the lab-steamed samples except that defect patterns are more homogeneous than for the case of accelerated aging in the lab. An FCU "old" fraction (Figure 10) shows more destruction, higher concentrations of mesopores, more A1 enriched bands, and the presence of more highly fractured grains than are found in the laboratory-aged samples (Figs. 4, 5, 6, 8).

16

Figure 10. HREM image of an "old" fraction from a commercial Fluid Catalytic Cracking Unit. Extensive crystallite fracture and many prominent dark bands are observed [61].

3. S T R O N G ACIDITY IN U L T R A S T A B L E Y 3.1. Background - Some Nonframework Aluminum is Essential The central role of framework A1 content in defining the catalytic properties of dealuminated H-Y was discussed in a classic paper by Pine, et al [62]. These workers concluded that catalyst activity, selectivity, and octane performance are each correlated with unit cell size, which, in turn is directly proportional to the number of aluminum atoms in the framework. The result of this study is genetically summarized in Figure 11. With increasing dealumination and concomitant loss of framework A1, selectivity (towards naphtha octane and olefinic content) tends to increase, but at the expense of catalyst activity and naphtha yield. The importance of nonframework aluminum (NFA) was not revealed in studies such as this. The more obvious manifestations of NFA species, such as the agglomerated species found in the electron microscopy studies discussed in the previous section, typically lead to decreased activity. Such agglomerated NFA species are thought to contribute to increased coke make

17 and not to desirable products. X-ray photoelectron spectroscopy (XPS) studies on hydrothermally dealuminated H-Y zeolites have shown that a considerable enrichment of aluminum occurs near the zeolite surface during the process of forming USY, with further accumulation at the surface upon extended hydrothermal treatment [60, 63]. Haag [27] observed that this migration of nonframework A1 to the external crystal surface generally lowers activity, presumably due to NFA species that have migrated to the surface and neutralized some of the BrCnsted acid sites near the crystallite surface [27, 63]. Haag further associated the increase in catalytic activity following ammonium exchange of a steamdealuminated zeolite with the reversal of this surface site neutralization [27]. Prior to the mid1980s, NFA species were generally regarded as undesirable, pore-blocking, amorphous "debris" and were often referred to as "detrital" A1. This picture has been changed by more recent studies which show that high silica H-Y materials, having little or no nonframework A1, exhibit poor catalytic activity for acidity dependent reactions (See Ref. 3-5, 61 and Table 2). It is concluded that the presence of some nonframework A1 is essential for the strong catalytic activity exhibited by fresh and mildly steamed USY materials. It is now generally accepted (3, 5, 61, 64) that the activity and selectivity of Y zeolites in catalytic cracking are determined by an interplay of framework aluminum and nonframework aluminum species. The implicated NFA species, presumably well dispersed, may exist as isolated, cationic species in the small cages.

Si/AI 2O 15 i

C

o

!

i

10 i

5 i

!

i

i

i

24.5

c (D

E :15 24.4 fl)

c

~D

24.3

24.2

9

N~ , # Al/ucell

!

Figure 11. Relation between framework composition of USY and unit cell dimension, upon which is superposed a generic representation of the role of unit cell size as a unifying concept in the catalytic properties of USY as discussed by Pine, et al [62].

18 Table 2 Certain NFA species are essential to good catalytic performance a for ultrastable FAU materialsb Si/A1 Unit Cell (nm)

Isobutane conv. rates, mol/(h.g) x l 0 3 total

carb. ion

1. clean framework, high-silica Y

4.9

13.7

8.5

2. conventional USY c

5.1

2.456

35.8

28.7

80

3. USY formed from AHF-dealum. material d

8.1

2.441

43.5

33.1

76

2.454

% carb. ion 62

Catalytic performance of clean framework, high-silica FAU is poor. a 500oc, 1.0 atm, 0.25 atm of i-C4/He (200 cm3/min); pretreat 1 h, 500~ He. b Na level in all three materials is low, Na_< 0.15 wt% [Ref. 3] c LZ-Y82, obtained from Union Carbide Portion of Sample #1, steamed at 570~ (PH2o = IA atm, 2 h) and then ammonium-exchanged.

3.2. Direct Measures of Solid Acidity in Zeolites: (a) Lack of Correlation with Activity, (b) The Question of Superacid Sites The enhanced cracking, isomerization activity associated with the presence of dispersed nonframework A1 species is not reflected in direct measures of solid acidity. In a review of studies of solid acidity by adsorption microcalorimetry, Dumesic and co-workers [65] note that, for dealuminated H-Y, the calorimetric results obtained at room temperature do not correlate with the catalytic activity for cumene cracking at 573 K. Samples with high activity showed essentially the same values of heat of adsorption of ammonia as did samples exhibiting

substantially

lower

activity.

Gorte

and

co-workers

[66]

carried

out

microcalorimetry measurements of pyridine and of isopropylamine adsorption, and also measured activities for hexane cracking on a series of steamed and chemically dealuminated H-Y materials. The calorimetry investigations failed to detect evidence for superacidic sites, and there was no correlation between hexane cracking activities and heats of adsorption for the materials examined. These workers also found no evidence for a very small concentration of strong sites, the presence of which had been previously proposed from calorimetry studies on steam-dealuminated USY catalysts [67]. Sommer, et al. [68] used 1H and 2H N M R to compare zeolite Y-catalyzed versus superacid-catalyzed proton-deuterium exchange in alkanes. From investigations of the extent of H/D exchange during passage of different light alkanes over an acidic D20-exchanged USY material (Si/A1F-- 4.5), it was concluded that

19 dealuminated H-Y cannot be considered as a superacid capable of protonating ~-bonds in alkanes. The acid strength was compared to that of sulfuric acid, consistent with results of Umansky, Engelhardt and Hall [69], and also with recent investigations of 13C chemical shift measurements of mesityl oxide by Haw and co-workers [70], which indicate that the acidity of ZSM-5 is comparable to that of a solution of 70% sulfuric acid.

3.3. Theoretical Studies of Br~nsted Acidity in Zeolites Significant progress has been made in understanding the nature and energetics of BrOnsted acid catalyzed hydrocarbon conversion reactions through theoretical, quantum chemistry-based studies, many of which have been carried out in combination with experiment [71-83]. The results of these studies indicate that stable carbocations in acid catalysis by zeolites may be the exception rather than the rule. In fact, the only such species observed spectroscopically have been exceptionally stable (and bulky) cations such as the methyl indanyl [84] or trityl [85] cations. This consensus is consistent with the landmark contribution by Kazansky [77-80], who first presented a mechanism for protonation of alkenes that did not involve the formation of stable carbenium ions. Theoretical studies by the van Santen group [71-76], initially on reactions of H-D exchange, and more recently extended to a wide range of hydrocarbon elementary reactions [76], have demonstrated a closely-related mechanism by which a stable carbocation is avoided. "The results of calculations indicated that the relatively stable intermediates of hydrocarbon transformations in zeolites are not carbocations but alkoxy groups covalently bound to the zeolite lattice. The carbocations represent high-energy activated complexes or transition states [75]." For example, this group found transition states for ethane cracking that are similar to carbenium ions albeit with stabilization from the lattice [74]. In other words, zeolite catalysts do not stabilize free carbocations, but the transition states associated with reaction pathways do resemble carbocations, consistent with Kazanky's earlier findings [77-80] and with the suggestion by Kramer and McVicker in a 1986 review paper [ 14]. In their recent review of the combined NMR and theoretical studies of solid acidity by Haw and co-workers [25, 26, 81-84], Haw, Nicholas and Xu [83] discuss similar conclusions from their studies of zeolites, emphasizing the contrast with comparable studies on true solid superacids. These workers observed that the body of theoretical and complementary experimental studies of zeolite acidity "do not support the long-held interlocked assumptions that zeolites are solid superacids and that free carbenium ions are prolific in zeolites... One of the earliest results of our combined theoretical and experimental collaboration was to show that H-D exchange in benzene on zeolites could also proceed without a stable benzenium intermediate. In contrast to the results on zeolites, it proved to be very easy to generate carbenium ions on Lewis acids such as A1C13powder and

20 BrCnsted acids such as HB/A1Br3 ... the wealth of NMR observations of cations on true solid superacids provided a context for interpreting and understanding the negative observations on the zeolites [83]." In summary, zeolites, as a class of solid acids, appear not to comprise superacids that stabilize free carbocations. The quantum chemical calculations indicate that carbocation-like transition state pathways, not stable reaction intermediates are available in zeolites. These are thought to proceed via routes that require stabilization from the lattice, such as through formation of surface alkoxy groups.

3.4. Synergism between framework and nonframework sites The question remains "What is the nature of the apparent synergism between framework and nonframework aluminum species that gives rise to enhanced catalytic activity? Beyerlein, McVicker and co-workers [3,4] suggested that the increased catalytic activity exhibited by USY materials, in comparison with "clean framework" FAU materials with comparable Si/A1F, may involve a "synergism" between Br0nsted sites associated with framework A1 and Lewis sites associated with dislodged aluminum. Such a synergism is consistent with the previously proposed concept of superacidity [10, 86] - as distinct from "superacids." A conceptual model for such a synergism was discussed by Lunsford and co-workers [5] who suggested that polyvalent A1 ions in the small cages are responsible for withdrawal of electrons from the framework OH- groups ("bridging hydroxyls"), thus making the protons more acidic. This model appears to be at odds with most direct measures of Br0nsted acidity (Sections 3.2. and 3.3.) and also with the results of combined acidity and reactivity investigations of Soled, McVicker and co-workers on USY zeolites modified by fluoride (HCF3) treatment [87]. For Si/A1 framework ratios near 5.5, these workers observed a substantial activity enhancement (ca. x 1.5) at low fluoride dealumination l e v e l s - in comparison with conventionally steam-dealuminated ultrastable Y zeolites with comparable framework composition. Such an activity enhancement cannot be correlated with either the number or strength of acid sites. In fact, the most active fluorided zeolite, with a framework Si/A1 ratio of about 5.5 and an F content near 1%, contained fewer BrCnsted sites than nonfluorided USY zeolites with similar Si/A1 framework ratios. Temperature programmed desorption of NH3 and integral calorimetric adsorption heats of NH3 indicated that acid strength decreased upon fluorine incorporation [87]. What remains unclear is the possible presence and role of a very small concentration of strong acid sites [67]. Soled, McVicker and co-workers [87] found that integral heats of adsorption obtained from TG/DTA NH3 titrations,160 kJ/mol for USY and 149 kJ/mol for 0.9%F/USY, did not show an increase in acid site strength for the fluorided sample. In contrast, investigations of differential heats of titration with pyridine by Ahsan, Arnett and

21 McVicker [88] revealed a small number of strong sites on both the USY and the 0.9%F/USY samples. Initial heats of adsorption, 144 kJ/mol and 174 kJ/mol for the USY and 0.9%F/USY samples, respectively, suggested that a larger number of exceptionally strong sites are present on F/USY than on USY. At the same time, both the average acid site strength and the total acid site number for the fluorided sample showed a decrease in comparison to USY. These workers concluded that the high activity F/USY catalyst behaves as if there were a larger number of acid sites present capable of converting isobutane. At equal conversions, the USY and F/USY samples exhibited near-equivalent selectivities, both for primary products and for secondary (carbocation) products [88]. The results of these studies raise anew the question: Could a small number of strong sites be key in the strong catalytic acidity exhibited by mildly dealuminated USY catalysts? Fripiat and co-workers have carried out extensive investigations on the nature of Lewis sites in aluminas, dealuminated mordenite, and dealuminated H-Y using both 27A1NMR [49, 89] and FFIR on adsorbed CO [90, 91]. The high resolution 27A1NMR studies [89] indicated the presence of two kinds of Lewis sites within the nonframework A1 distribution in dealuminated zeolites - a tetrahedral site and a pentagonal site with isotropic shifts of about 53 ppm and 37 ppm, respectively. The FfIR-CO adsorption studies [90, 91] also revealed the presence of two types of Lewis sites associated with nonframework aluminum. In the case of dealuminated mordenite, it was further shown that these Lewis sites were highly dispersed. Investigations of the isomerization of n-pentane and of o-xylene over dealuminated mordenites [92] showed initial reaction rates to be proportional to the product of the number of BrCnsted and the number of Lewis sites, suggesting that the the high acidity of dealuminated mordenites derives from a synergistic interaction between BrCnsted (framework) and highly dispersed Lewis (nonframework) acid sites [90-92]. Brunner et al. [6] arrived at similar conclusions from their NMR investigations of acid sites in ZSM-5. Fripiat and co-workers have also applied high resolution 298i NMR REDOR as an advantageous way to study the 1H - 29Si interaction for the characterization of BrCnsted sites [93]. They arrived at the quantitative conclusion that a BrCnsted site is an OH bridging an aluminum to a silicon with only one aluminum neighbor. An OH bridging an aluminum to a silicon with more than one aluminum neighbor is a proton donor but not a BrCnsted site capable of donating a proton to NH3 to make NH4+. These workers also found that there are fewer BrCnsted sites than framework A1 sites and that the difference between the number of BrOnsted sites and the number of framework A1 sites increased with the amount of nonframework A1 [93]. It was suggested that some of the acidic OH groups have disappeared owing to their neutralization by reaction with NFA species, consistent with the suggestion of Haag [27].

22

3.5. Alternative explanations for enhanced activity in mildly steam-dealuminated zeolites It has recently been suggested that the enhanced catalytic activity exhibited by steamed H-Y zeolites may be an artifact of a diffusion-limited reaction, which is "enhanced" by the formation of structural defects during hydrothermal treatment [94]. In the absence of any direct measurements on micropore diffusion, a model is put forth that incorporates the assumption that most of the BrOnsted sites in HY are inaccessible to the "micropore diffusionlimited bimolecular and oligomeric cracking reactions." In this model, the enhanced activity arises from the increase in defects/external surface area associated with hydrothermal treatment rather than from any direct influence of nonframework aluminum. This model fails to explain a number of salient observations in prior literature, including: (1) the markedly enhanced activity in HZSM-5 where the hydrothermal treatment is exceedingly mild [6, 27], (2) the poor catalytic activity exhibited by "clean framework," chemically dealuminated (via ammonium hexafluorosilicate, i.e., AHF) FAU materials [3-5], notwithstanding the demonstrably increased external surface area associated with chemical dealumination via

Table 3 Influence of sample size on Isobutane Conversion: a Results for Ultrastable FAU materials Conversion rates Carb ion Convn., Si/A1F grams mol %

% carb ion act.

total mol/(h'g) x 103

5b

0.30

7.28

69

32.5

22.4

0.196

5

0.40

9.48

70

31.7

22.2

0.194

5

0.50

12.6

73

33.7

24.6

0.215

5

0.60

15.7

74

34.9

25.8

0.225

5

0.70

18.8

75

35.9

26.9

0.235

5

0.80

21.1

76

35.3

26.8

0.234

5

1.00

26.9

80

36.0

28.8

0.251

12.4 c

1.00

13.0

71

17.4

12.4

0.227

12.4

2.30

30.1

78

17.5

13.6

0.249

molec/(min'A1F)

1.0 atm, 0.25 atm of i-C4/He (200 cm3/min); pretreat 1 h, 500~ He. b LZ-Y82 obtained from Union Carbide. c High-silica FAU material prepared from LZ-Y82; 650~ steam, 2 h, PH2o = 1/3 atm; NH4+ exchange; 0.33 N HC1 extraction [Ref. 3].

a 500oc,

23 AHF [3], and (3) the dramatically enhanced activity exhibited by ultrastable FAU materials prepared from AHF-treated materials (see Table 2, Fig. 12 and Refs. 3, 96). In the latter study, Beyerlein, McVicker and co-workers [3] carried out additional isobutane conversion investigations using varying amounts of conventional ultrastable FAU materials with compositions Si/A1F = 5 and 12.4. The results, reproduced in Table 3, show that, over a wide range of sample sizes, total rates and selectivities are little affected. Carbonium ion activity [3, 12, 95] per framework aluminum atom is essentially independent of both sample size and framework aluminum content. Catalyst deactivation in these studies was minimal, barely detectable in the mass balance [3,12]. Such consistency furnishes compelling evidence against the presence of any diffusion limitations in these studies and strongly supports the contention that carbonium ion activity [95] is directly dependent upon framework aluminum content. As discussed previously by McVicker et al. [12], the experimental evidence strongly suggests that a bimolecular mechanism gives rise to the secondary "carbonium ion products" in isobutane conversion, namely, propane, n-butane and isopentane.

3.6. Importance of the concentration and type of nonframework species The dependence of catalytic properties of dealuminated H-Y materials on unit cell size, or equivalently, on framework A1 content, can be profoundly altered by the concentration and type of nonframework species. In the case of steam/mineral acid-dealuminated ultrastable Y materials with framework compositions Si/A1F of 5 or greater, both hexane cracking [5, 97, 98] and isobutane conversion [3, 4] investigations show a linear dependence of activity on framework aluminum content. However, "unconventional" ultrastable Y materials, prepared by mild steam treatment of a "clean framework," AHF dealuminated USY were found by Beyerlein, McVicker and co-workers [3] to exhibit enhanced carbonium ion activity [95] for isobutane conversion (Figure 12). In these studies, carbonium ion activity, as evidenced by skeletal isomerization and oligomerization and back-cracking of isobutane [12], was found to be directly proportional to framework A1 content. It was suggested that the enhanced activity exhibited by the these materials, in comparison with results from conventionally prepared USY materials of comparable framework composition, was associated with their relatively lower content of nonframework A1 species (A1NFA), A1NFA/A1F-- 0.4. By contrast, the conventionally prepared ultrastable materials, for which the ratio A1NFA/A1Franged from 0.66 to 2.2, contained higher levels of nonframework A1. Enhanced activities for similarly prepared materials were observed by L6nyi and Lunsford [98] in the course of hexane cracking investigations on high silica Y materials prepared by mild steam treatment of chemically (AHF) dealuminated (Na +, NH4+)-Y, and by Sun, Chu, and Lunsford [11] on mildly steamed ZSM-20 materials. (ZSM-20 is a high silica, hexagonal variant of FAU.)

24 Interestingly, for framework compositions Si/AIF of 5 or greater, each of these studies showed a ratio A1NFA/A1F = 0.4, consistent with the results of earlier investigations by Beyerlein, McVicker and co-workers [3]. While such studies demonstrate the critical role of nonframework species in the development of strong acidity, no information is provided on their location. Carvajel, Chu, and Lunsford [5] showed that the presence of La 3+ in the small cages leads to a significant increase in hexane cracking activity over that shown by clean framework, high silica FAU materials. These results provided definitive evidence for the association of isolated cationic species in the small cages with the development of enhanced acidity. For a given framework A1 content, each La-exchanged material showed substantially increased activity over that of its clean framework parent material and somewhat lower activity than that of dealuminated H-Y (Figure 13) [5]. The results of this study, and also those from recent investigations of Lewis acidity in dealuminated zeolites by Fripiat and coworkers [89-92], provide compelling evidence that the critical nonframework A1 species are a) highly dispersed, and b) quite possibly exist as cationic species in the small cages of dealuminated H-Y, as indicated from earlier structural studies [36].

,so0utane

30

/I

/

n"~t-

x

---~

E~

20

. ~ x:: 0

~-~ lo o.... 10 20 30 AIF, Framework AI/uc

Figure 12. Carbonium ion rates from studies of isobutane conversion over high silica Y, ultrastable materials: conventionally prepared, O; prepared from materials initially dealuminated by using AHF, 4,. The single data point at the lower right, II, represents the carbonium ion rate over a low-sodium, AHFtreated FAU material [3].

240 ,,.-.,

._=

E 200 160 O

E ::I.

120

9~

80

o

<

40 0

I

I

I

i

a

10 20 30 40 50 60 70 80 Framework aluminum (AI/u.c.)

Figure 13. Hexane cracking activity over variously dealuminated high silica Y materials as a function of framework aluminum content [5] the effect of La3+ exchange. The intermediate activity curve shows the increase in activity following La 3§ exchange of clean framework materials (Those materials exhibiting lowest (near zero) activity are variously prepared clean framework materials). For the materials exhibiting highest activity: 9 Y-type zeolite dealuminated with SIC14; e, Y-type zeolite dealuminated by steaming [5].

25

3.7. Removal of Framework Sites by Poisoning - Comparison with Removal by Dealumination The arguments presented in the preceding section suggest that the exchange of high silica Y zeolite with trivalent cations, such as La 3+ or A13+, has important implications for solid acidity. By contrast, partial exchange of high silicaY materials with monovalent cations, as Na + or K +, leads to significant reduction in activity [4, 61, 66, 97]. The effect on isobutane conversion of controlled additions of Na + to a commercial ultrastable Y (Series 1) and to a further dealuminated ultrastable Y (Series 2) is shown in Figure 14 [4, 61 ]. It is apparent that Na addition suppresses activity much more rapidly than framework aluminum removal by dealumination. As indicated from the results of Na addition given in Figure 14 and by the replotting of these results in Figure 15, the addition of Na + equivalent to 1/3 of the total framework A1 atoms completely eliminates catalytic activity [4, 61 ]. Even lower proportions of added Na + were found to suppress hexane cracking activity in studies by Lunsford and coworkers [97]. It was concluded from these sodium poisoning studies that only a fraction of the framework A1 atoms are associated with strong acidity. This conclusion is supported by results of studies of isopropylamine desorption from dealuminated H-Y by Gorte and coworkers [66], which demonstrated that framework A1 content is significantly greater than BrCnsted acid site density. These workers suggested that each alkali ion affects only a single

Figure 14. Isobutane conversion studies of Na-poisoning [4, 61 ]. Effect on carbonium ion activity of controlled Na + addition to two USY parent materials with different Alp contents: o, Series 1, A1F = 32; A, Series 2, A1F = 21.6. Conditions: 500~ 1.0 atm, 0.25 atm of i-C4/He (200 cm3/min); pretreat 1 h, 500~ He.

26

40o o o

II Low Na, Dealuminated Materials

I"-"

x

E

30

133

Na + poisoned AIF = 32

,~

tO v

E

WP

"kll

20

m

c~ rr co

10

ii

.o

o

"k 0

--

m

0

I

5

.

I

10

I

I

.

I...

15 20 25 (AI F - 3Na)/Ucell

I

30

35

Figure 15. Effect of Na § addition on Isobutane Conversion - Comparison with dealumination: II, Low Na, dealuminated materials; ~, Na § poisoned materials, Alv= 32. active site and that framework A1 is not a good measure of acid site concentration. In their studies of the effect of NH3-poisoning on isobutane conversion on protonated zeolites (USY and ZSM-5 materials), W. K. Hall and co-workers [99] showed that the selective chemisorption of an NH3 equivalent of less than 10% of A1 tetrahedral [T] sites eliminated more than 90% of the activity. These workers concluded: "...only a small fraction of the total [A1T] sites is doing most of the work [99]." The recent high resolution 29Si NMR REDOR study of mildly dealuminated H-Y by Ffipiat and co-workers [93] has provided quantitative evidence that there are fewer Br0nsted sites than framework A1 atoms. This study showed that the difference between the number of BrCnsted sites and the number of framework A1 sites increased with increasing levels of nonframework A1. It was further suggested that some of the acidic OH groups have disappeared owing to their neutralization by reaction with NFA species [93]. The consistency of results from many different research teams using a variety of approaches furnishes compelling evidence that only a fraction of the framework A1 represents active sites. The notion that each alkali ion affects only a single active site has been called into question by results of hitherto unpublished studies which show that controlled addition of K + to dealuminated H-Y leads to a far more rapid suppression of isobutane conversion activity than does addition of Na + [100, 101]. The effect on isobutane conversion of controlled

27

additions of Na +, K +, and Ca ++ to a USY sample with Si/A1 = 5 is shown in Figure 16. Addition of potassium produces by far the strongest poisoning effect. Calcium addition gives rise to a poisoning effect intermediate between that of Na + and K + at low levels of exchange, ca. 10%, but shows an effect more mild than that of sodium as Ca ++ exchange levels exceed 20%. Comparison of the effect of K + addition on isobutane conversion with that of aluminum removal (Figure 17) shows that the addition of one K + gives an activity suppression roughly equivalent to the removal of 7 framework aluminum atoms. Alternatively, addition of one K § gives an activity suppression comparable to that of about two Na + ions. Similar trends have been reported by P.C. Dijkstra et al. [102] for alkali-ion-exchanged ZSM-5. This group showed that the extent of suppression of hexane cracking activity was directly related to cation size, with the effective degree of poisoning increasing with increasing ionic radius of the alkali metal cations. At equivalent levels of exchange, K + proved to be a more effective poison than Na + by a factor 1.5/1.2. It was concluded that aluminum sites in ZSM-5 are not isolated, but occur in clusters of two or three depending on aluminum content, consistent with suggestions of A1 pairs by Haag et al. [7, 27]. A quantitative assessment of the probability for occurrence of next-nearest neighbor (NNN) A1 pairs in ZSM-5 may derive from appropriate statistical analyses of the results of very recent synchrotron powder diffraction structural studies on Cs-ZSM-5 (Si/A1F = 15.5), which have provided definitive evidence for a nonrandom A1 distribution in ZSM-5 [103].

Figure 16. Effect of Cations, (Na § , K +, Ca ++) on Isobutane Conversion: ~:, Na + A, K § e, Ca ++

28

Figure 17. Effect of K§ addition on Isobutane Conversion- Comparison with dealumination: II, Low Na, dealuminated materials; A, K+ poisoned materials, A1F =32.

3.8. Effect of Acid Site Removal on Selectivity Previous correlations of isobutane conversion activity with framework composition [3] support a direct dependence of carbocation processes on framework aluminum, with a linear dependence of carbonium ion rates on A1F content. The demonstrated linear dependencies for Na or K addition (Figs. 15, 17) show that the primary effect of poisoning, or of framework A1 removal, is a decrease in the number of active sites. Utilizing the demonstrated linear dependencies for Na or K removal, the effect of decreasing acid sites on selectivity by aluminum removal is compared with that of Na addition in Figure 18, and with that of K addition in Figure 19. Carbonium ion selectivities (% carb ion activity) were estimated from the molar product ratio (n-butane + propane + isopentane)/total conversion products [95]. The measured selectivities (Figs. 18, 19) indicate a limiting site density of about A1F = 8, below which carbonium ion selectivity [95] diminishes rapidly.

29

Figure 18. Effect of decreasing acid sites on selectivity- Comparison of dealumination with Na + addition: II, Low Na, dealuminated materials; ~, Na + poisoned materials, A1F = 32.

Figure 19. Effect of decreasing acid sites on selectivity - Comparison of dealumination with K + addition: II, Low Na, dealuminated materials; A, K + poisoned materials, A1F =32.

30 3.9. Methane Formation - Tracking of Initiation Step in Carbocation pathways

Apparently, the likelihood of the initiation step for carbocation processes is sharply reduced in the vicinity of and below a limiting site density of about 8 A1F per unit cell. On the other hand, the results for the formation of methane (Figs. 20, 21), a stable reaction product marker, show that the initiation step and the secondary carbocation processes are intimately linked over the entire range of acid site content, whether manipulated by dealumination or by basic metal oxide addition. Comparison of the methane formation results in Figs. 20 and 21 with the carbocation selectivity results given in Figs. 18 and 19 shows that the initiation step is apparently independent of carbocation activity. These result are consistent with previous discussion of dual mechanisms [12, 14] and also with the results of acidity modification of USY zeolites by fluoride treatment [87]. The enhanced activity per A1F observed upon mild fluoride treatment was suggested to be due to an increased concentration of carbocation intermediates. It was concluded that the sites responsible for increasing the initiation rates, thus increasing the pool of carbocations, are not readily characterized by traditional chemical and physical acidity measurement probes [87, 88]. A combined EPR, NMR and product distribution study of oxidation sites in dealuminated mordenite has shown a strong correlation between the presence of nonframework A1 and the generation of radical cations [104]. Since Lewis sites in H-mordenite have been shown to be associated with nonframework A1 species [89-92], these results suggest a role of nonframework A1 species as electron acceptor sites for the generation of a radical cations, ultimately leading to the formation of olefins. Such a reaction sequence can serve as an initiation step for carbocation pathways, as proposed earlier [14]. Dumesic and co-workers [105] have shown that the reaction kinetics of the catalytic cracking of paraffins can be described by catalytic cycles that involve carbenium ion

initiation,

l~-scission,

oligomerimization, olefin desorption, isomerization and hydride ion transfer reactions. These workers observe that a model involving carbenium ions is kinetically equivalent to a model involving surface alkoxy species, provided that the surface coverages by reactive intermediates are low. Consistent with the results on methane formation discussed here, Dumesic et al. [ 105] conclude: "Initiation reactions are irreversible, insensitive to conversion and form olefins in conjunction with olefin desorption reactions...Olefin adsorption/desorption reactions are in quasi-equilibrium, but they have a significant impact on activity and selectivity by determining surface coverages of carbenium ions."

31 m Low Na Delauminated Materials

I

, m *m n-'~" 0.8 0

Na +

,mm"

C o .0

poisoned AI F 32

a3X

=

o~-E"~ 0.6 aC ~~

,m m

0.4

t'-4=-,

:~

m

0.2 0

"

0

I

I

5

10

I

I

I

15 20 25 (AI F- 3 Na)/Ucell

I

30

35

Figure 20. Methane formation rate vs. acid site concentration" m, Low Na, dealuminated materials; ~, Na + poisoned materials, A1F = 32.

Figure 21. Isobutane conversion rate vs. methane formation rate: m, Low Na, dealuminated materials; ~, Na + poisoned materials, A1F = 32; A, K+ poisoned materials, A1F =32; o, Ca ++ poisoned materials, A1F =32.

32 3.10. Severe versus Mild S t e a m T r e a t m e n t of D e a l u m i n a t e d H - Y - Effect on Acidity

Mild steam treatment (700 - 850 K, P~2o ~ 1) of ammonium-exchanged Y materials or of clean framework, high silica Y materials leads to sites of enhanced activity in acidity dependent reactions such as paraffin cracking or isomerization. By contrast, severe steam treatment (970 K to 1100 K, PH2o = 1) leads to rapid deactivation both by massive dealumination to framework compositions Si/A1F of 20 or more and by extensive crystalline destruction, as reviewed in the fluid catalytic cracking literature [17, 62,106, 107] and earlier in this paper. Investigations of differential heats of ammonia adsorption by Yaluris, et al. [108] have shown that, upon severe steam treatment of a USY catalyst (1060 K, Pn2o = 1, 2 hr) the acid site strength distribution shifts toward weaker sites. It has been suggested that the extensive dehydroxylation associated with severe steam treatment leads irreversibly to the formation of strong Lewis sites at the expense of BrCnsted sites [65]. For less severely dealuminated samples, the large scale migration and agglomeration of nonframework species discussed in Section 2 appears to have a retarding effect on activity, as a portion of the acidity and catalytic activity lost by steaming can be recovered by careful acid-leaching [3, 4, 109, 110], which removes both framework and nonframework A1 species. It has been suggested that the microporosity modifications associated with the generation of NFA species during hydrothermal treatment could change the initial pore diameter and thus the availability of acid sites [ 111 ]. Corma, et al [ 112] found that removal of some of the nonframework species from severely steam-dealuminated H-Y led to an enhanced alkylation activity and an extended catalyst life. In the case of mildly steamed HZSM-5, Lago, Haag and co-workers [7] attributed the generation of acid sites of enhanced activity (45 to 75 times more active than sites in clean framework HZSM-5) to the number of paired (NNN) A1 sites in the unsteamed parent. Upon severe steam treatment, or upon mild steam treatment of aluminum-poor HZSM-5 materials with one or less A1F/ucell (Si/A1 =100 or greater), enhanced activity is not observed. Measurements of isobutane conversion over extensively dealuminated H-Y indicate a limiting site density of about A1F/ucell = 8 (Si/A1F = 23 or greater), below which carbonium ion selectivity [95] diminishes rapidly [100]. HREM investigations of mesopore formation in USY materials subjected to different severities of steam treatment showed an inhomogeneous distribution of mesopores among different USY grains as well as within single grains [42, 59, 61]. Such inhomogeneities were found to be more pronounced for materials that were steam-treated at a higher temperature. These inhomogeneities, which encompass the presence of restricted regions with more severe dealumination and also regions with milder dealumination than the average, were attributed to the non-equilibrium nature of the accelerated steam-aging treatments in the laboratory. The

33 presence of numerous regions, less severely dealuminated than the average, can be expected to give a disproportionate contribution to the measured activity. As a result, the interpretation of measurements of solid acidity and/or of catalytic activity becomes increasingly problematic for severely steam-deactivated materials where such inhomogenities are most pronounced. In consideration of the inhomogeneities in defect concentration found by HREM [42, 59, 61 ], BrCnsted sites can be expected to contribute to measured activity even for severely steamdeactivated zeolitic catalysts where the average crystallinity is low. For the case of equilibrium cracking catalyst from the FCU where the average crystallinity is also low, the remaining BrCnsted acidity, and a disproportionate share of catalyst activity, is chiefly associated with the younger catalyst fractions, which have been shown to retain as much as 80% of the fresh catalyst crystallinity [58].

4. CONCLUSION The defect formation, which occurs reproducibly upon steam treatment of protonated zeolites, plays an important, although not well understood role in the acid catalysis exhibited by these materials. During the formation or further dealumination of a USY material by hydrothermal treatment, the collapse of small regions of the crystalline framework generates mesoporosity, simultaneously providing a source of Si atoms for healing the vacancies left by expulsion of A1 atoms from framework T-sites. A new understanding of the formation and evolution of mesopores during hydrothermal treatment has emerged from recent combined high resolution electron microscope (HREM) and analytical electron microscope (AEM) investigations on hydrothermally dealuminated USY materials [42, 59, 61]. The nonequilibrium nature of the accelerated steam-aging treatment in the laboratory gives rise to an inhomogeneous distribution of mesopores, which occurs concomitantly with further zeolite dealumination. In regions with high defect concentration, mesopores "coalesce" to form channels and cracks, which, upon extended hydrothermal treatment, ultimately define the boundaries of fractured crystallite fragments. The predominant fate of aluminum ejected from lattice sites appears to be closely associated with dark bands that often decorate these newly formed fracture boundaries, as observed by H R M . A smaller proportion of ejected aluminum exists as isolated [36] or agglomerated species [35] within the zeolite cages. Similar defect patterns, although less inhomogeneous, have been observed for age-separated (young) fractions from a commercial Fluid Cracking Unit [61 ], where the rate of deactivation is slower than for accelerated steam-aging in the lab. As a consequence of the pronounced inhomogeneities, including numerous regions with milder dealumination than the average, the interpretation of measurements of solid acidity and/or catalytic activity becomes increasingly

34 problematic for severely steam-deactivated materials where such inhomogeneities are most pronounced. High silica H-Y materials, having little or no nonframework A1, exhibit poor catalytic activity for acidity dependent reactions such as paraffin isomerization or cracking. Investigations of mildly dealuminated zeolites indicate that the origin of their high catalytic activity is a synergistic interaction between BrCnsted (framework) and highly dispersed Lewis (nonframework) acid sites. The enhanced cracking, isomerization activity associated with the presence of highly dispersed nonframework AI species is not reflected in direct measures of solid acidity obtained, for example, by calorimetry or by NMR spectroscopy. The question of how nonframework aluminum affects catalytic activity remains an experimental and modeling challenge. The enhanced cracking, isomerization activity of mildly steamed materials almost certainly involves some synergistic interaction between framework BrCnsted sites and highly dispersed nonframework species. The presence of this enhanced activity is not consistent with a major increase in acid site strength. BrCnsted sites are presumably more dilute than the distribution of framework A1 as only a fraction of the framework A1 appears to represent active sites. The critical nonframework A1 species quite possibly exist as cationic species in the small cages of dealuminated H-Y. By contrast, partial exchange of high silica Y materials with monovalent cations, such as Na or K, leads to significant reduction in activity, presumably by poisoning acid sites. In prior studies of isobutane conversion over dealuminated H-Y, it was shown that the addition of sodium equivalent to 1/3 of the total framework A1 atoms completely suppresses catalyst activity [4, 61]. More recent extensions of these poisoning studies show that addition of potassium produces a much stronger poisoning effect, with one K + ion giving an activity suppression approximately equivalent to that produced by two Na § ions. In analogy with conclusions form a recent study reporting similar trends in ZSM-5 [ 102], these results are consistent with the presence of next-nearest-neighbor (NNN) A1 clusters in the FAU framework, such that a single alkali ion can influence more than one acid site. It is noted that conclusions on partial ordering of aluminum in the FAU framework, based on analyses of

29Si MASNMR

intensity data, indicate that A1 NNN pairs tend to be avoided

[1131. The current body of theoretical and associated experimental studies indicate that zeolites, as a class of solid acids, do not comprise superacids that stabilize free carbocations [71-84]. The quantum chemical calculations show that the transition states associated with reaction pathways resemble carbocations. That is, carbocation-like pathways are available in zeolites, but these proceed via routes that require stabilization from the lattice, such as through formation of surface alkoxy groups. In contrast with the results on zeolites, where none other

35 than bulky carbocations have been spectroscopically observed, carbenium ions are easily generated on Lewis acids such as A1C13 powder and BrCnsted acids such as HB/A1Br3 [83]. Previous correlations of isobutane conversion activity with framework composition support a direct dependence of carbocation processes on framework aluminum (A1F), with a linear dependence of carbonium ion rates on A1F content. The demonstrated linear dependencies for Na or K addition show that the primary effect of poisoning, or of A1F removal, is a decrease in the number of active sites. Measured selectivities indicate a limiting site density of about A1F/ucell - 8 (192 T sites/ucell), below which carbonium ion selectivity diminishes rapidly. Consistent with previous discussion of dual mechanisms, the results for formation of methane, a stable reaction product marker, show that the initiation step and the secondary carbocation processes are intimately linked over the entire range of acid site content, whether manipulated by dealumination or by poisoning. The initiation step appears to be independent of carbocation activity, consistent with the conclusions from independent efforts to describe the reaction kinetics of the catalytic cracking of paraffins [105]. While these and related studies have provided new understanding of the initiation step for carbocation pathways on zeolite catalysts, significant challenges remain. The sites responsible for increasing the initiation rates (and thus increasing the pool of carbocations) are not readily characterized by traditional chemical and physical acidity measurement probes. The nature of the initiation step(s) for the cracking of paraffins remains a subject of research. The recent advances in understanding of the bifunctional nature of zeolite catalysts reviewed here have led to some notable advances in Fluid Catalytic Cracking technology. The use of amorphous alumina added to rare earth exchanged Y zeolite to improve FCC activity and conversion to gasoline has been reported by Exxon [114]. The added alumina is reported to increase the Lewis acidity of the catalyst providing a dehydrogenation function. The olefins produced due to the increased Lewis acidity lead to improved conversion on the BrCnsted acid sites of the zeolite. These findings resulted in the commercialization of a new FCC formulation by Exxon.

ACKNOWLEDGEMENTS The authors gratefully acknowledge many useful conversations with C. Choi-Feng, Werner Haag, Jan Hall, W. Keith Hall, George Kramer, Jack Lunsford, Mario Occelli, Lloyd Pine, G. J. Ray, Bill Schuette and Stu Soled. The expert assistance of B. J. Huggins, L. N. Yacullo and J. J. Ziemiak is also gratefully acknowledged.

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39 72. R. A. van Santen, in: Advanced Zeolite Science and Applications, eds. J. C. Jansen et al. (Elsevier, Amsterdam, 1994) p. 273. 73. J. A. Lercher, R. A. van Santen and H. Vinek, Catal. Lett. 27 (1994) 91. 74. S. R. Blaszkowski, M. A. C. Nascimento and R. A. van Santen, J. Phys. Chem. 100 (1996) 3463. 75. M. V. Frash, V. B. Kazansky, A. M. Rigby and R. A. van Santen, J. Phys. Chem. B 102 (1998) 2232-2238. 76. M. V. Frash and R. A. van Santen, Topics in Catalysis 9 (1999) 191-205. 77. V. B. Kazansky, in: Advanced Zeolite Science and Applications, eds. J. C. Jansen et al. (Elsevier, Amsterdam, 1994) p. 251. 78. V. B. Kazansky and I. N. Senchenya, J. Mol. Cat. 74 (1992) 257. 79. V. B. Kazansky, Acc. Chem. Res. 24 (1991) 379. 80. I. N. Senchenya and V. B. Kazansky, Catal. Lett. 8 (1991) 317. 81. J. F. Haw, J. B. Nicholas, Teng Xu, L. W. Beck and D. B. Ferguson, Acct. Chem. Res. 29 (1996) 259. 82. J. B. Nicholas, Topics in Catalysis, Vol. 4 (1997) pp. 157-171; from: Acidity in Aluminas, Amorphous and Crystalline Silico-Aluminas, Eds. J. J. Fripiat and J. A. Dumesic (Balzer Science Publishers, Amsterdam, 1997). 83. a) J. B. Nicholas, Teng Xu and J. F. Haw, Topics in Catalysis 6 (1998) 141-149; b) J. F. Haw and Teng Xu, Advances in Catalysis 42 (1998) 115-180. 84. Teng Xu andJ. F. Haw, Am. Chem. Soc. 116 (1994) 10188. 85. T. Tao and G. E. Maciel, J. Am. Chem. Soc. 117 (1995) 12889. 86. a). P. A. Jacobs, Carboniogenic Activity of Zeolites (Elsevier, Amsterdam, 1977). b) P. A. Jacobs, H. E. Leeman and J. B. Uytterhoeven, J. Catal. 33 (1974) 17. 87. S. L. Soled, G. B. McVicker, J. J. Ziemiak, G. DeMartin, L. N. Yacullo, K. Strohmaier and J. Millar, In Symp. on Chemically Modified Molecular Sieves, Amer. Chem. Soc., Chicago, August 22-27, 1993, ACS, Div. Petr. Chem. Preprints, Vol. 38 (3), 1993, pp. 546-551. 88. T. Ahsan, E. Arnett and G. B. McVicker, unpublished results. 89. D. Coster, A. L. Blumenfeld and J. J. Fripiat, J. Phys. Chem. 98 (1994) 6201-621!. 90. V. Gruver and J. J. Fripiat, J. Phys. Chem. 98 (1994) 8549-8554. 91. V. Gruver, Y. Hong, A. G. Panov and J. J. Fripiat, in: llth Int. Congr. on Catal., Stud. Surf. Sci. Catal., Vol. 101, eds. J. Hightower, et al. (Elsevier, Amsterdam, 1996) p. 741. 92. Y. Hong, V. Gruver and J. J. Fripiat, J. Catal. 150 (1994) 421-429. 93. A. L. Blumenfeld, D. Coster and J. J. Fripiat, J. Phys. Chem. 99 (1995) 15181-15191. 94. B. A. Williams, S. M. Babitz, J. T. Miller, R. Q. Snuff and H. H. Kung, A ppl. Catal. A: General 177 (1999) 161. 95. Carbonium ion selectivities were estimated with the molar product ratio (n-butane + propane + isopentane)/total conversion products. This ratio accounts for the major carbonium ion based products arising from isomerization and chain-cracking sequences. Carbonium ion activity was defined as the product of carbonium ion selectivity and total rate (See Ref. 12). 96. J. W. Ward and T. L. Carlson, U.S. Patent 4517073, May 14, 1985.

40 97. P. O. Fritz andJ. H. Lunsford, J. Catal. 118 (1989) 85. 98. Ferenc L6nyi and J. H. Lunsford, J. Catal. 136 (1992) 566-577. 99. W. K. Hall, J. Engelhardt and G. A. Sill, In Zeolites: Facts, Figures, Future, eds. P. A. Jacobs and R. A. van Santen (Elsevier, Amsterdam, 1989) pp. 1253-1262. 100. a) R. A. Beyerlein, G. B. McVicker, L. N. Yacullo and J. J. Ziemiak, "Acidity Studies on High Silica Zeolite Y Materials - Effects of Aluminum Removal and of Na, K, Ca Addition," Thirteenth North American Meeting of the Catalysis Society, Pittsburgh, PA, May 2-6, 1993; b) this review paper. 101. The controlled cation additions were carried out using reagents NaNO2, KNO3 and Ca(NO3)2. In each case, hydrated zeolite was combined with a properly diluted reagent solution. In order to ensure equilibrium distribution of the added cations, this wet impregnation step was followed by a programmed calcination with slow nitrite/nitrate decomposition and a final calcination temperature of 475~ For a series of materials coveting the range - 2.5% to 60% cation exchange- good agreement was obtained between nominal and analyzed exchange levels. 102.P.C. Dijkstra et al. Catalysis Letters 10 (1991) 375-382. 103. D.H. Olson, N. Khosrovani, A. Peters, and B. Toby, J. Phys. Chem., In press. 104. G. Harvey, R. Prins, R. Crockett and E. Roduner, J. Chem Soc., Faraday Trans. 92 (1996) 2027. 105. R. D. Cortright, J. A. Dumesic and R. J. Madon, Topics in Catalysis 4 (1997) 15-26. 106.J. Biswas and I. E. Maxwell, Appl. Catal. 63 (1990) 197-258. 107.D.J. Rawlence and K. Gosling, Appl. Catal. 43 (1988) 213-237. 108. G. Yaluris, J. E. Rekoske, L. M. Aparicio, R. J. Madon and J. A. Dumesic, J. Catal. 153 (1995) 65-75. 109. A. Auroux and Y. Ben Taarit, Thermochim. Acta 122 [ 1987] 63. 110.M. Krivfinek, N. Thiet Dung and P. Jim, P. Thermochim. Acta. 115 [1987] 91. 111.L. Kubelkova, S. Beran, A. Malecka and V. M. Mastikhin, Zeolites 9 (1989) 12. l12.A. Corma, A. Martinez and C. Martinez, Appl. Catal. A: General 134 (1996) 169. 113. A. W. Peters, in: Catalytic Materials: Relationship Between Structure and Reactivity, eds. T. E. Whyte, et al. (Amer. Chem. Soc., Washington D. C., 1984) pp. 201-217. 114.W.L. Schuette and A. E. Schweizer, EP 0 749 781 A3, 1997.

Studies in Surface Science and Catalysis 134 M.L. Occelli and P. O'Conner (Editors) 9 2001 Elsevier Science B.V. All rights reserved

41

The use of microcalorimetry and solid state nuclear magnetic resonance (NMR) to study the effects of post-synthesis treatments on the acidity and framework composition of several HY-type zeolites M.L. Occelli a, A. Aurou b , M. Kalwei c, A. W61kerc, and H. Eckertc

aMLO Consulting, Atlanta, GA 30328, USA blnstitut de Recherches sur la Catalyse, CNRS, 2 Av. A. Einstein, 69626 Villeurbanne, France Clnstitut fiir Physikalische Chemie, Westf~ilische Wilhelms-Universit/it Mtinster, Schlossplatz 7, D-48149 Mianster, Germany

The effects of post-synthesis treatments on HY-type zeolites have been investigated by X-ray diffraction (XRD), nitrogen porosimetry, 29Si and 27A1 MAS NMR together with microcalorimetry experiments at 150 ~ with ammonia as the probe molecule. Postsynthesis treatment of NaY crystals with ammonium salts and reaction of NH4Y with ammonium fluorosilicate solutions followed by calcination, can generate silicon-enriched faujasite frameworks (LZ-210s with Si/AI=3.1 and 4.4) which, in addition to a unique distribution of T[nSi,(4-n)A1] sites, are free from extraframework Al-species. Acid site density and the number of strong acid sites (B+L) in these crystals are greater than in the reference HY ; however, the strength of the strongest acid sites in LZ-210 is lower than in LZY-82. As the framework Si/A1 molar ratio in LZ-210 increases, the 29Si spectrum becomes dominated by the T(4Si,0A1) resonance and AI(V) and AI(VI) species can now be detected in the crystals 27A1 spectra. The decreased acid site (B+L) densities in these Sienriched LZ-210s, are attributed to losses of framework SiOHA1 units while the increased strength of the strongest sites is attributed mainly to the removal of most of residual charge compensating Na-ions. IR spectra of pyridine chemisorption have shown that extraframework Al-species in a LZ-210 (Si/AI=4.4) under study, form two types of weak L-sites and that the B-sites present are stronger than in the reference HY (LZY-82). In contrast to LZ-210s, the steam-stabilized ultrastable HY (USHY) has a 29Si spectrum that resembles the one for LZY-82 and its 27A1 spectrum contains resonances representing AI(IV) as well as AI(V) and AI(VI) species; its acidity is lowest owing to a 4.4% Na20 content. Removal of most of the residual charge-compensating Na-ions and

42 steaming produces USHY-DA crystals with increased acidity and framework Si/At = 11.0 that contain only minor amounts of extraframework AI(VI)species. I. INTRODUCTION Although at the time this work was undertaken the price of crude oil was below $11/barrel and now it exceeds the $30/barrel mark, the fluid catalytic cracking technology has steadily continued to evolve to meet the new challenges facing refiners. More than 370 fluid catalytic cracking units are in use worldwide generating the capacity to produce in excess of 460 million gallons of gasoline a day (1). The lead phaseout, reformulated gasoline, together with ever stringent automotive emission standards, provide the main incentive to continue to study the properties of fluid cracking catalysts (FCCs) for gasoline production (2). Today, the demand for olefin-rich gasoline is focusing the refiner attention on new FCC compositions in which acidity can be manipulated to generate materials with a low hydrogen transfer index (3,4). To impart the desired catalytic activity to the aluminosilicate framework, several postsynthesis treatments have been proposed to modify the acidic properties of HY type zeolites (5-11), the main cracking component of all commercially used FCCs. These efforts began almost 35 years ago when researchers at Davison realized that steam-aging NH4Y crystals at mild conditions (500 ~ generates faujasite crystals with a contracted unit cell size (resulting from losses of framework AI(IV) atoms) and greatly improved thermal and hydrothermal stability (5-7). The high (-4 % Na20) residual Na content of Davison's ultrastable Y zeolite (USHY) has a deleterious effect on the crystals acid site density and strength. However, repeated ammonium ion exchange and calcination cycles can yield HY crystals with properties similar to those of USHY and containing only minor (< 0.3 wt.% Na20) amounts of alkali ion impurities (8) as, for example in UOP's LZY-82. Acid-leaching framework Al-atoms with mineral acids (9), or repeated exchange reactions with ammonium salts coupled with thermal and hydrothermal treatments, allows the generation of siliceous faujasite crystals such as those described in this paper. Alternatively, de-alumination can be achieved by exposing HY crystals to gaseous SiCI4 (10), or by reacting a suspension of NH4Y with aqueous (NI-hhSiF6 solutions (11). When using ammonium hexafluorosilicate solutions, reaction conditions and stoichiometry can be manipulated to obtain HY crystals with the desired Si/A1 ratio (11), Si atom distribution and content of extraframework Al species. High resolution electron microscopy ( H R M ) images have shown that, in FCC containing HY type crystals, the extraframework A1 generated during aging migrates from the HY framework (12,13) leaving behind mesopore-containing faujasite crystals whose acidity and catalytic properties depend (in part) on the distribution of their Al species (14,15). The physicochemical properties of modified HY type zeolites have been recently reviewed by Beyerlein and McVicker (see ref. 16 and references therein). The purpose of this paper is to report the use of 29Si and 27A1 solid state MAS NMR data, together with microcalorimetry and FTIR results from pyridine thermodesorption, in the study of the framework composition and acidic properties of several faujasite

43 crystals modified by post-synthesis chemical, thermal and hydrothermal treatments. The utility of NMR and adsorption microcalorimetry to characterize zeolites has been reviewed and discussed elsewhere (17-19)

2. EXPERIMENTAL

2.1 Materials The reference HY (UOP's LZY-82) is a sample prepared by repeatedly exchanging NaY crystals with NH4NO3 solutions; it has a bulk SiO:/A1203 molar ratio of 5.6 and contains 0.16% Na20. The three silicon enriched HY crystals (UOP's LZY-210s) have been produced instead by reacting a slurry of NH4Y crystals with ammonium hexafluorosilicate solutions (11); the LZY-210 samples have a bulk SIO2/A1203 molar ratio of 6.2, 8.8 and 12.8 respectively. Sample LZ-10 is an ammonium exchanged and hydrothermally dealuminated HY(LZ-210) with a bulk SiO2/Al203 molar ratio of 6.0. The sample of ultrastable HY crystals (USHY) from Davison has been prepared by steamaging at 732 ~ (with 100% steam at 1 atm) HY crystals containing a residual 4.4% Na20. Sample USHY-DA has been prepared by reacting USHY with an (NH4)2SO4 solution at pH-2.9. After decomposing the NH4 ions, a second steam-aging and (NH4)2SO4-exchange step reduced the % Na:O level to 0.17 from 4.4. The properties of all the samples under study can be found in Table 1. 2.2 Surface Area Analysis Nitrogen sorption isotherms obtained at liquid nitrogen temperature were collected using a volumetric technique on a Micromeritics ASAP 2010 adsorption instrument equipped with version 3.0 software. Prior to analysis, samples weighing from 0.1-0.3 g were outgassed in v a c u u m at 400 ~ for at least 16 h. The total pore volume was derived from the amount of nitrogen adsorbed at a relative pressure close to unity (P/Po = 0.995) by assuming that all the accessible pores were then filled with liquid nitrogen. Surface area measurements have been performed using the BET equation (20). 2.3 Microcalorimetry Heats of adsorption of NH3 were measured using a heat-flow microcalorimeter of the Tian-Calvet type (from Setaram) linked to a glass volumetric line. Successive doses of gas were sent onto the sample until a final equilibrium pressure of 133 Pa was obtained. The equilibrium pressure relative to each adsorbed amount was measured by means of a differential pressure gauge from Datametrics. The adsorption temperature was maintained at 150~ Primary and secondary isotherms were collected at these temperatures. All samples were dried overnight under vacuum at 400~ before calorimetric measurements were undertaken. The accuracy of the data reported is +/- 4 kJ/mol or +/- 1 kcal/mol. 2.4

298i and 298i NMR

27A! MAS NMR Spectroscopy spectra were recorded at 59.6 MHz on a modified Bruker CXP300 spectrometer. Samples were spun in cylindrical 7mm zirconia rotors at spinning speeds

44 near 4 kHz. 90 ~ pulses of 6~ts length and 30s recycle delays were used in all cases. Chemical shifts were determined relative to tetramethylsilane as an external reference; spectra were deconvoluted into Gaussian lineshape components. 27A1 MAS NMR spectra were obtained at 130.2 MHz using a Bruker Avance (DSX) 500 spectrometer. Samples were spun in cylindrical 4mm zirconia rotors at a spinning frequency of 12 kHz. The spectra were recorded with solid 45 ~ pulses of 2~ts length and relaxation delays of 1s. Resonance shifts are reported using liquid samples of 1M aqueous solutions of AI(NO3)3 as an external reference standard. Prior to measurements, all samples were dried at 400 ~ in air overnight. Because the exact spectral parameters of the AI(IV), AI(V), and AI(VI) signal components are not exactly known (due to a distribution of quadrupolar coupling parameters) no 27A1 lineshape deconvolution was attempted.

2.5 IR Analysis Infrared (IR) spectra were obtained using a Nicolet 3600 spectrometer. Two spectra were obtained from each sample before accepting the results as being representative of the zeolite sample under study. Spectra were acquired with a 2 cm~ resolution (8192 data points) and apodized using the Happ-Genzel algorithm. Self-supporting wafers (4-8 mg/cm2 in density) were prepared by pressing samples between 25 mm diameter dies for one minute at 6,000-7,000 lbs. pressure. Prior to pyridine sorption, the wafers were mounted in an optical cell and degassed by heating at 200 ~ for 2h at 10-3 torr. The pyridine-loaded wafers were then heated (in vacuum) in the 200-500 ~ temperature range. 3. RESULTS AND DISCUSSION 3.1 29Si MAS NMR Results The 29Si MAS ~ spectra of the different HY samples under study are compared in Figure I. As expected, the 29Si spectrum of the reference HY (LZY-82), obtained from NaY (LZY-52) through NH4-exchange and calcination steps, yields a familiar peak pattern containing resonances at -I 06.9, -101.5, -95.8, -90.9, -87.6 ppm representing the different T[nSi,(4-n)Al] sites present, see Figure IA (19). The resonance near -I I0 ppm has been attributed to the presence of extraframework silica (2 I). The Si/AI ratio from this spectrum is 4.5; this value is, considerably higher than the value of 2.8 obtained from chemical analysis, Table 2. This discrepancy can be safely attributed to the presence of extraffamework Al-species which are clearly visible in the 27AI MAS spectrum shown in Figure 2A. In contrast, HY crystals prepared by reacting NH4Y with ammonium hexafluorosilicate solutions followed by calcination, produce the spectrum shown in Figure 1B, indicating a much stronger retention of framework AI and a completely different distribution of T-sites. In fact, whereas the spectrum for HY(LZY-82) is dominated by the T(4Si,0AI) resonance, the contribution to the total spectral intensity in LZ-210 with Si/AI=3.1, increases in the order: T(3Si, IA1) > T(2Si,2AI) > T(4Si,0AI)>> T(1Si,3A1); see Figure lB. As reported elsewhere (11), the hydrolysis of the fluorosilicate salt forms protons and fluoride anions that cause the dealumination of the

45 Table 1. Some physicochemical properties of HY type crystals after calcination in air at 500~ for 1Oh.; bulk Si/A1 molar ratios are from ehemi'cal analysis . . . . . . . . . HY- T)~_e_

Si/A1

ao (nm,)

LZY-82

2.8

2.4480

785

0.16

LZ-210 LZ-210 LZ-210

3.1 4.4 6.4

2.4581 2.4474 2.4377

607 625 629

2.44 1.17 0.01

LZ-10

3.0

2,4310

600

0.17

2.9 8.3

2.4530 2.4360

742 740 . . . . . .

4.41 0.17

USHY USHY-DA

SA(m2/_g)

%Na2Q

Table 2.

29Si chemical shifts (-ppm) and % relative intensity (in parenthesis) results;

some Dhvsicochemical properties of these HY-type zeolites, are described in Table 1 . . . . . . . . . -

-

HY type

Si/A1

Tf0Si,4A~ T(1Si,3AI) T(2Si,2A1) ~

T(4Si,0AI} SiOT. NMR Chem

LZY-82

87.6 (1.5)

90.9 (4.6)

95.8 (14.1)

101.5 (37.2)

106.9 (37.6)

110.8 (5.0)

4.5

2.8

LZ-210

--

8 9.0 (3.5)

94.6 (27.9) 94,3 (11.0) 95,8 (3.4)

100.1 (50.0) 100.3 (67,8) 101.9 (22.9)

106.2 (16.3) 106.1 (19.3) 107.5 (70.7)

110.1 (2,8) 111,6 (1.9) 111.7 (3.0)

3.4

3.1

4.4

4.4

13.1

6.4

101.5 (16.9)

107.5 (76.7)

110.6 (6.4)

24.0

3.0

101.3 (37.3) 101.3 (1.9.0)

t06.2 (32.2) 106.8

111.9 (9.7) 111.1

4.2

2.9

11.0

8.3

(64.0)

(10.2) . . . . . . . . .

LZ-210 LZ-210

LZ-10

USHY USHY-DA

--

90.3 (7.6)

95,2 (13,2) 95.4 (6.8)

46

A

D

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

_._ _L L_..~

i~

G o

-io

.......

-ibo ..... ~ - ~ o

Chemical

Shift

-

-ioo

[ppm]

Figure 1.29Si MAS NMR spectra of: A) HY (LZY-82) and LZ-210 with bulk Si/A1 ratio orB) 3.1, C) 4A, and D) 6.4. The spectra for LZ-10, USHY and USHY-DA are in E, F, and G respectively

47

A

B

120

80

40

0

.-40

Chemical Shift [ppm] Figure 2, ~TAlMAS NMR spectra of A) HY (LZY-82) and LZ-210 with bulk Si/Al ratio of B) 3.1, C) 4.4, and D) 6.4. The spectra for LZ-10, USHY and USHY-DA are in E, F, and G respectively. faujasite structure and the removal of some of the Al in the form of a soluble salt such as (NH4hAIFs. The rest of the AI is re-introduced, together with the Si, to occupy vacant defect sites in the dealuminated faujasite framework. As a result, the concentration of Tsites in which Si is coordinated to one or more A1 atoms increases yielding the spectrum

48 shown in Figure lB. The enhanced concentration of T(3Si, IA1) sites could be of particular importance to the cracking activity of these crystals during hydrocarbon conversion because work by Fripiat and co-workers has indicated that only T(3Si, IA1) groups are Bronsted sites strong enough to protonate ammonia and form NH4-ions (22). When the aforementioned post-synthesis treatment is modified to increase the bulk Si/AI ratio to 4.4 from 3.1 (while the %Na20 decreased to 1.17 from 2.44), the crystals produce a new 29Si spectrum indicating that contributions to the total spectral intensity by the different T[nSi,(4-n)Al] sites has changed; Figure 1C. The 27A1MAS NMR spectrum in Figure 2C contains a single resonance resulting from AI(IV) sites. Thus, as before, only framework A1 has been introduced in these faujasite crystals. It is for this reason that Si/AI values computed by NMR and chemical analysis are in good agreement, see Table 2. After repeated exchanges with NH4-ions and thermal treatments, the %Na20 in the LZ-210 sample decreased to 0.01 from 1.17 while its Si/AI ratio increased to 6.4 indicating dealumination and A1 removal from the crystals during the NH4-exchange reactions; Tables 1-2. As a result, the 29Si spectrum in Figure 1D becomes dominated by the T(4Si,0AI) resonance at -107.3 ppm and extraframework AI(V) and AI(VI) species appear in the 27A1 spectrum in Figure 2D. The large discrepancy between Si/A1 ratios from M R and chemical analysis shown in Table 2 is attributed to losses of framework A1. Results in Table 1 show that as the Si/Al molar ratios in these LZ-210 increases, the crystals exhibit the anticipated (36) unit cell contraction while their SA remains in the 600630 m2/g range. If after treatment with ammonium hexafluorosilicate solutions the silicon enriched HY crystals are steam-aged as in the LZ-10 sample, the T(4Si,0A1) resonance becomes sharper indicating increased local order, Figure 1E. In this figure, contributions of other T[nSi,(4-n)Al] sites to the 29Si spectrum are hardly noticeable indicating substantial losses of AI from the framework. The 29Si spectrum for Davison USHY in Figure 1F resembles the one in Figure 1A for LZY-82, however resonances are considerably less well resolved. The contributions to the total spectral intensity of the different T[nSi,(4-n)Al] sites, are in Table 2. The Si/A1 ratio value of 4.2 computed from the spectrum in Figure IF is higher than the 2.9 value obtained from chemical analysis. As before, this difference is attributed to the presence of extraframework M-species. In fact, the 27A1 spectrum in Figure 2F contains contributions from AI(V) and AI(VI) sites. By repeating the (NH4hSO4 exchange and calcination steps, USHY can be dealuminated to produce USHY-DA crystals with a much higher framework Si/AI ratio, see Table 2. The 29Si spectrum in Figure 1G is dominated by a single resonance representing T(4Si,0AI) sites and it closely resembles the one in Figure 1D for a siliceous LZ-210 sample. 3.2 27A1MAS NMR Results The 27A1MAS NMR spectra of the differem HY samples under study are shown and compared in Figure 2. The spectrum in Figure 2A for HY(LZY-82), has already been

49 discussed elsewhere (19,23). Surprisingly, the spectrum for LZ-210 with Si/AI = 3.1 contains a single sharp resonance near 54 ppm indicative of AI(IV) sites; see Figure 2B. Consistent with this result, the Si/A1ratio (3.4) from NMR is in excellent agreement with the value of 3.1 from chemical analysis. When the treatment with (NH4)2SiF6 solutions is modified to increase the bulk Si/A1 ratio to 4.4 from 3.1, the27Al spectrum remains practically unchanged suggesting that only framework AI has been introduced into the zeolite framework, see Figure 2C. However after reducing the %Na20 level to 0.01 from 1.17, extraframework A1 species are formed and the spectrum in Figure 2D contains a resonance near 0 ppm representing AI(VI) sites and a resonance near 54 ppm that overlaps another centered near 30 ppm attributed to AI(V) species (24) or to highly distorted AI(IV), (25). In the LZ-10 sample (Table 1), the resonance near 30 ppm becomes well defined and, together with the one near 0 ppm, increases in intensity, Figure 2E. In fact, the 27A1 spectrum of the steam-aged LZ-210 crystals shown in Figure 2E, contains resonances from AI(IV), AI(V), and AI(VI) sites. In Figure 2E, the AI(V)+AI(VI) resonances contribute more than 50% to the overall signal intensity. As a result the Si/A1 ratio from NMR is eight times larger than the one from chemical analyfis, see Table 2. Recent results obtained using the double-rotation (DOR) spinning technique (26,27) to the study of AI in zeolites, have concluded that extraframework A1 contains, in addition to AI(V) and AI(VI), some AI(IV) species. By comparing 27A1MASNMR with 27A1 CP MASNMR, it has been shown that in steam-aged HY a substantial portion of the resonance attributed to AI(IV) results from the presence of extraframework A1 (28). Since the 29Si spectrum in Figure 1E for LZ-10 is represented by a tingle resonance from T(4Si,OAI) sites, the AI(IV) resonance in the 27A1 spectrum shown in Figure 2E is attributed to the formation and retention of extraframework AI(IV) species. The spectrum for the USHY sample in Figure 2F resembles the one for HY(LZY-82) in Figure 2A; however it shows the presence of lower levels of AI(VI) species. After dealumination, the resonance attributed to AI(IV) becomes sharper and only trace amounts of residual extraframework AI(VI) contribute to the overall spectral intensity of the USHY-DA sample; see Figure 2G. Although present, the amount of extraframework A1 indicated in Figure 2G is much less than the one seen in Figure 2A for the reference HY (LZY-82). This result has been attributed to a more efficient removal of excess A1 from the faujasite micropores by the sulfate solution used during the NH4-exchange reactions. 3.3 Mierocalorimetry

Mierocalorimetry results have been reported in Tables 3-4 and in Figures 3-4. Sorption isotherms for ammonia chemisorption are shown in Figures 3A-3B respectively. Not shown are secondary sorption isotherms, that is, sorption isotherms for samples after adsorption of the probe molecule and degassing in vacuum at 150 ~ By subtracting the adsorbed volume of the secondary isotherms from the one of the primary isotherms at the same equilibrium pressure (p = 0.2 tort), it is possible to obtain V~, the volume of irreversibly chemisorbed NHs. It has been reported that this value is indicative of the total number (B+L) of strong acid sites present (29,30). The difference in NH3 sorption shown in Figure 3 does not correlate with the crystals BET surface area in Table 1. In fact the

50

NH8 uptake (wn~/g)

~1000

.~,

,

2500

,, , . ,

"u_,.,

u~,J

__,

. .......

:.

,,,.

-

L.Z-2'D-,e.I

-+-

-4-

L,Z-2D-,54

- . e - L,Z-10

A

I.Z..210";4,4

2OOO

1600 I(XX)

0

0.1

0.2.

0.8

0,4

0.8

~0

0.7

1400

B

1200 1000 8OO

0OO

0G ~

o

- ' - LZY-82 . . . . . . .

i .....

oa

: . . . . . .

I:. . . . .

eL=

+

USHY

__+ :

Jt . . . . . .

.

,

. . - i - USHY~)A. t_...

~a cx4 P (torr)

.....

s~.__.,.

o~a

_

_

,tr .

~e

.

.

.

.

.

.

c~7

Figure 3. Ammonia sorption isotherms. HY crystal properties are given in Table I. two LZ-210 samples free from extraframework AI (Si/AI=3.1-4.4, %Na20=l. 17-2.44) and with a surface area in 600-630 m2/g range, sort) almost twice as much NH3 than LZY-82 (Si/AI=2.8 and %Na20~. 16) with 785 m2/g surface area. Although the (B+L) number of strong acid sites (as detem~ed from V~ values) and acid site density (as determined from integral heats) are greater in the Si- enriched crystals, the strength of their strongest sites (L-sites) is 61-68 Ll/mol lower than the one in HY (LZY-82), Table 3. In addition, the two LZ-210 samples have a population of sites with stren~h in the 100-150 lcJ/mol

51

Table 3. Ammonia Chemisorption Data (+/-4kJ/mol) at 150 ~ and at p = 0.2 torr ; Si/A1 ratios .are from chemical analysis . . . . . . . . . . . Si initial Int. A1 Heat Heat NH3 (l,tmol/g) HY Woe kJ/mol (J/~) VT Virr 1. LZY-82

2.8

208

122

1020

680

2. LZ-210 3. LZ-210 4. LZ-210

3.1 4.4 6.4

140 147 191

220 211 94

1893 1919 799

1130 1316 525

5. LZ-10

3.0

190

35

373

188

6. USHY 7. USHY-DA

2.9 8.3

150 i 86

40 67

479 608

161 420

Table 4. Ammonia chemisorption data at 150 ~ (p = 0.2 torr). Population of sites with given strenffth is in ~tmo! NH3/g - bulk Si/AI molar ratios are from chemical analysis data HY Type

Si AI

<100

kJ/mol ~ 100-150

....... 150-200

LZY-82

2.8

280

657

72

LZ-210 LZ-210 LZ-210

3.1 4.4 6.4

341 435 241

1552 1484 433

. . . . 125

LZ-10

3.0

230

107

USHY USHY-DA

2.9 8.3

365 195

102 9 388

>200 11 . .

. . --

36

,

12 24 . . . .

---

that is more than twice as large as the one found in LZY-82. LZ-210 crystals with Si/A1=3.1-4.4, do not contain sites with strength greater thanl50 kJ/mol; see Tables 3-4. This moderate acidity has been attributed mainly to the greater levels of Na ions in these two samples; Tables 1-3. When the Si/AI molar ratio in LZ-210 increases to 6.4 from 4.4, initial heat values increase while, as expected, acid site density (as determined from integral heat values) decreases due to losses of framework A1, see Tables 3-4. The appearance of a population

52 Q (kJ/mol) 2oo r - - ' - ' - - - - - -

:

LZ-'210"e.1

-+"

LZ-'210-~4

-'~

LZ-'210-~4

- =

LZ-IO

A |

O 0

_

.....

9

,

,,|

~,,

LI

I(XX)

i

i

|

|

. . . . . . . .

|

RtX)O

1500

250

-

t.~'412

"-I-- 081"~

200

-11- ~ ' O A

100

B 0 0

200

400

~10

,~nmon~ ~

800

1000

1200

1400

(uric/q)

Figure 4. Differential heat of ammonia adsorption as a function of ammonia coverage. (125 ~tmol NH3/g) of sites with strength in the 150-200 kJ/mol range has been attributed to the drastic reduction (0.01%Na20) of Na levels that has occurred in this sample. The sorption isotherms for the two LZ-210 samples free from extraframework AI, overlap over the entire pressure range investigated, Figure 3A. In this figure, the decreased sorption capacity of the LZ-210 sample with Si/AI = 6.4, has been attributed to losses of framework SiOHAI sites and to the deposition of most of the extraframework A1 thus

53 generated inside the crystals micropores. Additional ammonium exchange reactions and thermal and hydrothermal treatments of LZ-210 samples can produce a siliceous (Si/Al=24) framework with a unit cell of only 2.4310 nm that contains trace amounts of Na-ions ; the SA remains high as seen for the LZ-10 sample in Table 1. The low sorption capacity of these crystals (Figure 3A), is the result of a greatly reduced acid sites density induced by severe framework AI losses. However the initial heat in LZ-10 remains high (190 kJ/mol), Table 3. Thus, the dealumination of LZ-210 type crystaIs produces siliceous faujasite frameworks containing a reduced acid site density in which the strength of the strongest acid rites is enhanced and the population of a site with a given strength depends on the concentration of residual charge compensating Na ions and on the framework Si/A1 ratio. The deleterious effects of residual Na ions on acidity are well illustrated by the different microcalorimetry results obtained with USHY (with 4.4% Na20) and USHYDA (with 0.17 % Na20) samples ; see Figure 3B and Tables 3-4. The number of rites available to NH3 ehemisorption in USHY is the lowest measured in this set of HY samples', this result has been attributed to the high level of Na present. Although undergoing severe dealumination, the number of strong acid rites in USHY-DA is more than double the one in USHY. In addition, in USHY-DA there is a large increase in Bronsted type acidity and in initial heat values ; Tables 3-4. Differential heats of ammonia adsorption as a function of coverage are shown in Figure 4. Results for the reference HY (LZY-82) have been discussed elsewhere (23). The differential heat profiles in Figure 4A reveal that, as the isomorphous substitution of AI with Si increases, acid site density and population of sites with a given strength change. In agreement with the sorption isotherms in Figure 3A, the profiles for the two LZ-210 samples with Si/AI ratio of 3.1 and 4.4 are somewhat similar, Figure 4A. When Si/AI=3.1, the LZ-210 crystals contain a small (-142 ~tmol/g) population of sites with strength near 140 kJ/mol and a larger one (-611 lamol/g) with strength near 123 kJ/mol. This large population of sites is attributed to framework Si(OH)A1 groups (30-34) which according to 29Si NMR results (22) are associated mainly with T(3Si, IAI) sites. Sites with strength below 100 kJ/mol have been attributed to interaction of NH3 with weak L-sites or with silanol groups (30-34). When the Si/A1 ratio increases to 4.4 from 3.1, the distribution of T[nSi,(4-n)A1] sites change; see Figure 1C and Table 2. As a result, a population (-388 t,tmol/g) of acid sites with strength in the 128-135 kJ/mol range and second one similar in size (-383 lamol/g) but with strength near 113 kJ/mol can be observed in Figure 4A. At a higher Si/A1 ratio of 6.4, with the exception of a small population of sites with strength near 152 kJ/mol, differential heat of ammonia ehemisorption monotonically decrease with coverage indicating that the dealuminated framework contains an heterogeneous distribution of acid site strengths. In Figure 4A, the sharp decrease in acidity exhibited by the LZ-10 sample has been assigned to drastic losses of Si(OH)AI groups from the faujasite structure during thermal and hydrothermal treatments. The differential heat profiles in Figure 4B for Davison's USHY and USHY-DA, its dealuminated counterpart, provide another example of the deleterious effects of residual charge compensating Na cations on the distribution of acid site strength in the faujasite structure. At a coverage below 250 gmol/g, the acidity of USHY is similar to the one in

54 LZ-10, then as NH3 coverage increases, a small population of sites with strength near 70 kJ/mol (not seen in LZ-10 samples), appears. After dealumination, the crystals %Na20 decreases to 0.17 from 4.41 and although the acidity of the USHY-DA crystals remain lower than in LZY-82, the two zeolites exhibit similar differential heat profiles for ammonia uptakes up to 400 ~mol/g; Figure 4B. 3.4 IR Results The infrared bands of pyridine of interest to this study are the N-H and C-H stretching frequencies as well as the C-C and the C-N stretching frequencies at 15901660 cm1 and near 1500 cm-1 (37). IR spectra showing the C-C and C-N stretching vibrations of chemisorbed pyridine on LZ-210 crystals, are shown in Figure 5. Bands near 1545 cm~ represent protonation of pyridine by Bronsted (B) sites and the presence of coordinatively bonded pyridine on Lewis (L) sites is represented by bands near 1455 cm"~ (37). The strong band near 1490 cm ~ in Figure 5 is attributed to the presence of both B and L sites (37). IR results for LZ-210 with bulk Si/AI=4.4, have been collected in Table 5. In this table, B and L acid site densities as a function of degassing temperature have been obtained by dividing the integrated absorbance in the 1563-1523 cm"1 region and in

Table 5. Pyridine chemisorption data for a Si enriched faujasite (LZ-210 with Si/AI=4.4) and reference HY (LZY-82), calcined at 500~ Oh ; LZ-210 composition is in Table 1. (* = could not be measured with accuracy) .

T(C) 200 300 400 500

LZY-82 (Si/AI=2.7) B L B/L 0.59 0.13 4.54 0.47 0.09 5.22 0.27 0.05 5.40 0.10 0.03 3.33

LZ-210 (S...i/AI=4.4) B L B/L 1.44 0.035 41.1 1.24 * -0.74 * -0.12 * --

.

the 1475-1400 cm"~ region in Figure 5, by the sample density. Pyridine desorption data given in Table 5 indicate that, in contrast to what is generally reported (38) for the reference HY(LZY-82), the exchange procedure with ammonium hexafluorosilicate solutions generates Si enriched faujasite crystals that contain mostly weak L-sites from which thermal desorption is more facile than from the B-sites present. Figure 5A show that aider degassing at 200 ~ there is a broad asymmetric band in the 1400-1475 cm"1 region that appears to be the superposition of two bands. In fact, at 300 ~ this broad and intense band disappears and is replaced by two weak bands centered near 1425 and 1452 -1 O cm respectively, Figure 5B, the low frequency band is no longer visible at 400 C, Figure 5C. Thus it appears that two different types of L- sites are present possibly the results of the different types of extraframework Al-species in the crystals' micropores. In the 200 to 400 ~ temperature range, the amount of pyridine chemisorbed on B-sites in the LZ-

55

~--~

1600

~

~.

, -,.,,

,

,

"

w-

'-

'"

"'~

J ' r -'-~'''''

--'""

'~'1

1500 1400 WAVENUMBERS (1/cm)

Figure 5. Infrared spectra of chennsorbed pyndme on a LZ-210 sample with bulk Si/AI=4.4 The sample has been degassed at 200 ~ and then exposed to pyridine and degassed at" A) 200, B) 300, C) 400 and D) 500 ~ LZ-210 spectrum in Figure 5D closely resembles the one for HY published elsewhere (38). The different acidic properties of LZ-210 type crystals should be of particular interest to researchers investigating the modification and improvement of the reactivity and selectivity properties of fluid cracking catalysts (FCCs),

56 4. SUMMARY AND CONCLUSIONS 29Si and 27A1 MAS NMR spectra have indicated that the reaction of aqueous ammonium hexafluorosilicate solutions with NH4Y can yield HY crystals with a silicon enriched framework free from measurable non-framework A1 species in which T(3Si, IA1)/T(4Si,0A1)>I.0. Microcalorimetry results with ammonia as the probe molecule, have revealed that LZ-210 type crystals with Si/A1=3.1-4.4 have greater acid site density (as indicated by integral heat data) and contain more strong acid sites (as indicated by V~ data) than the reference HY (LZY-82); however the strength of the strongest sites is greater in LZY-82 probably because of its lower Na content. When the severity of the post-synthesis treatment increases to raise the framework Si/A1 ratio of LZ-210 type crystals to 6.4, the 29Si NMR spectrum becomes dominated by the T(4Si,0A1) resonance and AI(IV), together with AI(V) and AI(VI) species are detected by 27A1NMR. FTIR results from pyridine thermodesorption in the 200 ~ to 500 ~ temperature range, have indicated that these extraframework Al-species in LZ-210 are probably responsible for the broad and asymmetric band observed, in the 1400-1475 cmZ region, after degassing the pyridine-loaded crystals at 200 ~ This broad and intense band disappears at 300 ~ to be replaced by two much weaker bands indicating that different Lsites are present and that the strength of these L-sites is low. Replacement of A1 with Si in the faujasite framework, increases B acid site strength and acid (B+L) site density. The high (4.4% Na20) level of residual charge compensating Na ions is what quenches the acidity in USHY crystals. The :9Si NMR spectrum of these crystals resemble the one for LZY-82. After dealumination, the USHY-DA crystals obtained contain 0.17% Na20 and acidity reappears; these crystals yield a spectrum similar to the one for the Si-enriched LZ-210s with Si/Al = 6.4. In USHY-DA, only trace amounts of Al(VI) species can be detected by 27A1 NMR. Thus after repeated exchanges with (NH4)2SO4 solutions most of the Na as well as non framework Al-species are removed from the HY microporous structure indicating that the sulfate anions are more efficient than the nitrate ions (used in preparing LZY-82 type crystals) in removing non framework Al-species from the faujasite micropores.

Acknowledgements Special thanks are also due to Dr. P.S. Iyer for initial M R spectra, to Dr. P. Ritz for IR data and to Dr. A.E Schweizer (Exxon) for elemental analysis and XRD results. This work has been supported by NATO collaborative grant CRG-971497 to MLO and HE.

REFERENCES 1.Chemical and Engineering News, November 23, p.51 (1998). 2. Catalysis Looks to the Future, NRC Dept., Nat. Acad. Press, Washington, D.C. (1992) 3. U.A. Sedran, Catal. Rev.-Sci. Eng., 36, 3, p.405 (1994) 4. E. Dai, C.M .Tsang, R.H. Petty, and M.L. Occelli in "Proc. Int. Zeolite Conf. "Seoul, South Korea; Kodansha -Elsevier, vol. 105 B, p 981 (1996).

57 5. C.V. McDaniel and P.K. Maher in US Patent 3,292,192 (1966) 6. C.V. McDaniel and P.K. Maher in US Patent 3,449,070 (1969) 7. C.V. McDaniel and P.K. Maher in "Zeolite Chemistry and Catalysis", ACS Monograph 171, J.A. Rabo Ed.; Am. Chem. Sot., Ch. IV, Washington D.C. (1976) 8. D.W. Breck in Zeolite Molecular Sieves, Wiley, New York, N.Y., (1974) 9. J. Scherzer, J. Catalysis 54, 285 (1978). 10. H.K. Beyer and I. Belenykaja, Catalysis by Zeolites, B. Imeliket al., Eds.; Elsevier, Amsterdam, p.203 (1980). 11. GW. Skeels,, and D. Breck, in "Proc. 6th. Int. Zeolite Conf", D. Olson and A. Bisio Eds.; Butterworths, p.97 (1984). 12. C. Choi-Feng, J.B. Hall, B.J. Huggins, R.A. Beyerlein, J. Catal. 140, 395-405 (1993) 13. Beyeflein, R.A., Choi-Feng, C., Hall, J.B., Huggins, B.J., and Ray, G.J. "Fluid Catalytic Cracking Iii: Materials and Processes" M.L. Occelli and P. O'Connor Eds.; ACS, Washington D.C., 81 (1994) 14. R.A.Beyerlein, C. Choi-Feng, J.B. Hall, B.J. Huggins and G.J. Ray, Topics in Catalysis, 4, 27-42 (1967) 15. R.A. Beyedein, G.B. MeVicker, L.N. Yaeullo, J.J Ziemiak, J. Phys. Chem.., 92, 1967 (1988) 16. R.A. Beyerlein, G.B. MeVieker, (this volume) 17. A. Auroux, Topics in Catalysis 4, 71 (1997). 18. A. Auroux Catalyst Characterization : "'Physical techniques for solidMaterials ", Chapter 22 ; B. Imelik, and J.C. Vedrine, Eds.; Plenum press, N.Y. (1994). 19. G. Engelhardt and D. Michel, "High Resolution Solid-State NMR of Silicates and Zeo#tes", J.J. Wiley, New York, N.Y. (1987) 20. S. Brunauer, P.H. Emmett and E.J. Teller, J. Am. Chem. Soc.,60, 309 (1938). 21. M.L. Oeeelli in "Catalysts in Petroleum Refining and Petrochemical Industries" M. Absi-Halabi et al., Eds.; Studies in Surface Science and Catalysis, Elsevier, Amsterdarn,100, 27 (1995) 22. A.L. Blumenfeld, D. Coster, J.J. Fripiat, J. Phys. Chem. 99, 15181-15191 (1995) 23. M.L. Occelli, A. Auroux, M. Kalwei, A. WOlker, and H. Eckert, (this volume) 24. J.P. Gilson, G.C. Edwards, A.K. Peters, K. Rajagopalan, R. Wormsbecher, T.J. Roberie, M.P. Shatlock, J. Chem. Soe. Chem. Comrnun. 91(1987) 25. A. Samoson, E. Lippmaa, G. Engelhardt, D. Lohse, H.G. Jerschwitz, Chem. Phys. Lett. 134, 589 (1987) 26. A. Samoson, E. Lippmaa and A. Pines, Mol. Phys. 65, 1013 (1988). 27. G.J. Ray, and A. Samoson, Zeolites 13,410 (1993) 28. H. Hong, D. Coster, F.R. Chen, J.G. Davis and J.J. Fripiat "'New Frontier in Catalysis", L. Guczi Ed.; Elsevier, Amsterdam, 1159-1170 (1993) 29. A. Auroux "Catalyst Characterization : Physical Techniques for Solid Materials", Ch.22; B. Imelik, J.C. Vedrine, Eds., Plenum press, N.Y. (1994) 30. A. Auroux, Topics in Catalysis, Vol 4, 71-89 (1997). 31. Z.C. Shi, A. Auroux and Y. Ben Taarit, Can. J. Chem. 66, 1013 (1983) 32. A. Auroux, and Y. Ben Taarit, Therm. Aeta, 122, 63 (1987)

58 33. D. Chen, S. Sharma, N. Cardona-Martinez, J.A. Dumesic, V.A. Bell, G.D. Hodge, R.J. Madon, J. Catal. 136, 392-402 (1992) 34. N. Cardona-Martinez and J.A. Dumesic, J. Catal. 125,427 (1990) 35. M.L. Occelli, M. Kalwei, A. Wolker, H. Eckert, A. Auroux and S.A.C. Gould J. Catalysis (submitted) 36. L.A. Pine, P.J. Maher and W.A. Wachter, J. Catal. 85, 466-476 (1984) 37. E.P. Parry, J.Catalysis Vol 2, pp 371-379, (1963) 38. M.L. Occelli, A. Auroux, H. Eckert, M. Kalwei, A. W61ker and P.S. Iyer, Micro. and Meso. Materials 34, 1, 15-22 (2000).

Studies in Surface Science and Catalysis 134 M.L. Occelli and P. O'Conner (Editors) 9 2001 Elsevier ScienceB.V. All rights reserved

59

The effects of steam aging temperature on the properties of an H Y zeolite of the type used in F C C p r e p a r a t i o n s M.L. Occellia, A. Aurouxb, A. Petr b , M. Kalweic, A. Wolkerc, and H. Eckert c

aMLO Consulting, Atlanta, GA 30328, USA blnstitut de Recherches sur la Catalyse, CNRS, 2 Av. A. Einstein, 69626 Villeurbanne, France Clnstitut for Physikalische Chemie, Westf~ilische Wilhelms-Universit~it MOnster, Schlossplatz 7, D-48149 MOnster, Germany X-ray diffraction (XRD), N2 sorption, 29Si and 27A1 MAS NMR together with microcalorimetry experiments at 150 ~ with ammonia, have been used to investigate the effects of steam-aging temperature (at constant steam-aging time) on the properties of an HY zeolite of the type used in fluid cracking catalysts (FCCs) preparation. 27A1 MAS NMR spectra indicate that exposure for 5h to 100% steam at 760 ~ causes the anticipated removal of framework AI(IV) with the formation of AI(V) and AI(VI) species together with a contraction of the zeolitesurface area and unit cell (u.c) size. When the steaming temperature is raised to 788 and 815 ~ from 760 ~ lattice degradation increases and additional AI(IV) is removed from the lattice and converted (in part) into extraframework AI(V) and AI(VI) species; the presence of these species increases with steaming temperature. Evidence of framework dealumination during steam-aging can be easily observed also in the zeolite 29Si NMR spectra. After steam-aging at 760 ~ the HY crystals yield a 29Si NMR spectrum in which T(3Si, IAI) sites are still present. In contrast, when the steam-aging temperature is raised to 788 and 815 ~ only T(4Si,0A1) sites can be observed in the 29Si spectra suggesting that the AI(IV) detected by 27A1 MAS NMR is mostly non framework. In agreement with NMR results, microcalorimetry experiments with ammonia at 150 ~ indicate that during steam-aging the acid site density as well as the number of strong acid sites in the faujasite structure decrease drastically. However, in the temperature range investigated, the strength of the strongest acid sites remains practically unaffected by steaming.

60 I. INTRODUCTION The importance of framework composition on the activity and selectivity properties of faujasite crystals in a wide range of hydrocarbons conversion reactions has long been recognized (1-4). The specific role of framework and extraframework AI species on HY performance remains largely unexplained although the high catalytic activity of dealuminated HY crystals has been attributed to a synergistic interaction between Bronsted sites in the framework and Lewis acid sites in extraframework Al-species (5). The advent of high resolution solid state nuclear magnetic resonance (NMR) has allowed the almost routine tracking of AI species in solids (6-9). Specifically, 29Si NMR provides direct information on framework composition and on the Si atoms environment while 27A1NMR allows for a distinction between framework AI(IV) and extraframework AI(IV), AI(V) and Al(VI)-species (8,9). The modifications of HY type zeolites following steam aging at microactivity test (MAT) conditions are important because they mimic and explain in part, the initial decrease in cracking activity that HY-containing fluid cracking catalysts (FCCs) undergo when introduced into a fluid cracking catalyst unit (FCCU) of a typical refinery. The accelerated steam-aging procedure used by researchers in the lab reduces the surface area (SA) of a fresh FCC to values near those measured in the corresponding equilibrium FCC sample (10). Moreover, 29Si NMR spectra of steam-aged and equilibrium FCC samples are practically indistinguishable (10) and the 29Si NMR spectrum of fresh or aged FCCs can be considered a superposition of the spectra of their components (10). After aging, the 29Si NMR spectrum of a fresh FCC is reduced to one main resonance near -107 ppm representative of T[4Si,0A1] sites. In these samples, minor amounts of T[(4-n)Si,nAl] sites may also be present although these sites are not required for high cracking activity during gas oil conversion(10). Similar conclusions have been obtained from 27A1 MAS NMR spectra. However, significant differences exist between steam-aged and equilibrium FCC samples. In fact, steam-aged and equilibrium FCCs have different pore structure, initial heats of ammonia and pyridine chemisorption as well as different distribution of AI(IV), AI(V) and AI(VI) species (10). High resolution electron microscopy (HREM) images have shown that steam-aging give rise to an inhomogeneous distribution of mesopores in FCCs that occurs concomitantly with zeolite dealumination and crystallite fracture (11). Such inhomogeneities, which are more pronounced for higher temperature steam treatments, have been observed among different HY grains as well as within single grains (11). Microcalorimetry experiments with ammonia and pyridine have shown that after aging either at MAT conditions or in a FCCU, a fresh FCC undergoes severe losses in acid sited density while retaining most of the strength of its strongest Lewis acid sites. These rites, and the retention of an open micro and mesoporosity, are believed responsible for the aged equilibrium FCCs cracking activity. Heat flow microcalorimetry with pyridine as the probe molecule, has already been used to study the effects that steam-aging times have on a FCC acidity (12). Aider only 15 rain at 787 ~ (and in the presence of 100% steam at 1 atm), the FCC under study lost most of its total acidity and the main population of sites decreased with increasing steam-aging times (12). It is the purpose of

61 this paper to study the effects of steam-aging temperature (at a constant steam-aging time of 5h) on the framework composition of an HY zeolite of the type used in FCC preparations. 2. EXPERIMENTAL

2.1 Materials The reference HY (UOP LZY-82) is a sample prepared by repeatedly exchanging NaY crystals with NH4NO3 solutions followed by calcination; it has a bulk SiO2/A1203 ratio of 5.6 and contains a residual 0.16 wt.% Na20. The (as received) HY powder was pressed into a thin wafer that was then crushed to produce 80 x 150 mesh granules irregular in size and shape. Prior to steam-aging the granules were calcined in flowing air at 500 ~ Steam-aging was conducted for 5h with 100% steam at 1 atm. The temperatures used were: 760, 788 and 815 ~ The properties of the steamed HY crystals are listed in Table 1 2.2 Surface Area Analysis Nitrogen sorption isotherms obtained at liquid nitrogen temperature were collected using a volumetric technique on a Micromeritics ASAP 2010 adsorption instrument equipped with version 3.0 software. Prior to analysis, samples weighing from 0.1-0.3 g were outgassed in vacuum at 400 ~ for at least 16 h. The total pore volume (PV) was derived from the amount of nitrogen adsorbed at a relative pressure close to unity (p/Po 0.995) by assuming that all the accessible pores were then filled with liquid nitrogen. Surface area (SA) measurements were performed using the BET equation; PV and SA results are presented in Table 1.

and 27A! MAS NMR Spectroscopy 29Si NMR spectra were recorded at 59.6 MHz on a modified Bruker CXP300 spectrometer. Samples were spun in cylindrical 7mm zirconia rotors at spinning speeds near 4 kHz. 90 ~ pulses of 6~ts length and 30s recycle delays were used in all cases. Chemical shifts were determined relative to tetramethylsilane as an external reference; spectra were deconvoluted into Gaussian lineshape components. 29Si chemical shifts together with % relative intensity data are reported in Table 2. 27A1MAS NMR spectra were obtained at 130.2 MHz using a Bruker Avance (DSX) 500 spectrometer. Samples were spun in cylindrical 4mrn zirconia rotors at a spinning frequency of 12 kHz. The spectra were recorded with solid 45 ~ pulses of 2~ts length and relaxation delays of ls. Resonance shifts are reported using liquid samples of 1M aqueous solutions of AI(NO3)3 as an external reference standard. Prior to measurements, all samples were dried at 400 ~ in air overnight. Because the exact spectral parameters of the AI(IV), AI(V), and AI(VI) signal components are not exactly known (due to a distribution of quadrupolar coupling parameters) 27A1 lineshape deconvolution not was attempted. 2.3

29Si

62

2.4 Microcalorimetry Heat of adsorption of NH3 was measured using a heat-flow microcalorimeter of the Tian-Calvet type (from Setaram) linked to a glass volumetric line. Successive doses of gas were sent onto the sample until a final equilibrium pressure of 133 Pa was obtained. The equilibrium pressure relative to each adsorbed amount was measured by means of a differential pressure gauge from Datametrics. The adsorption temperature was maintained at 150~ Primary and secondary isotherms were collected at these temperatures. All samples were dried overnight under vacuum at 400 ~ before calorimetric measurements were undertaken. Microcalofimetry results have been collected in Tables 3-4. 3. RESULTS AND DISCUSSION Nitrogen porosimetry and XRD results for the steam-aged HY crystals, are listed in Table 1. As expected, when the severity of the thermal and hydrothermal treatment increases, the u.c. size of the crystals contracts. It decreased from 2.454 nm in the calcined HY, to 2.425 nm in the HY steam-aged at 815 ~ owing to a severe dealumination of the faujasite framework (13). Dealumination causes partial lattice degradation and the surface area (SA) and pore volume (PV) of the crystals decrease as indicated in Table 1. Table 1. Some physicochemical properties of HY-type crystals (LZY-82 from UOP) after calcination in air at 500 ~ followed by steam-aging at 760, 788, and 815 ~ for 5h with 100 % steam at 1 atm ao PV Surface Area, SA (m2/_g) Sample (nm) (cc/g) Total Micro. Meso. %Na~O 1. HY 2. HY-760 3. HY-788 4. HY-815

2.454 2.431 2.426 2.425

0.27 0.23 0.22 0.17

785 668 649 493

697 580 569 428

88 88 80 65

.

0.16 --,-

3.1 zgsi MAS NMR Results The 29Si MAS and 27A1MAS NMR spectra of HY crystals (LZY-82) steam-aged for 5h in the 760 ~ to 815 ~ temperature range, are compared in Figures 1-2. As expected, the 29Si spectrum of the reference HY (LZY-82), calcined in flowing air, yields a familiar peaks pattern containing five resonances at -106.9, -101.5, -95.8, -90.9, and 87.6 ppm representing the different T[nSi,(4-n)Al] sites present, see Figure 1A (6). The Si/A1 molar ratio from this spectrum computes to 4.5 which is a value considerably higher than the value of 2.8 obtained from chemical analysis results, see Table 2. This apparent discrepancy can be safely attributed to the dealumination of the faujasite structure and formation of extraframework AI(V) and Al(VI)-species that can be easily detected in the 27A1spectrum shown in Figure 2A and elsewhere (10). In Figure 1A, the resonance near -

63

q

A

.

B

A

D I

0

'

-'

'~'

"'--

9

"

-50

'

"

"'

"i"

'

"' . . . .

~'

-100

I

-150

Chemical Shift [ppm] Figure I. 29Si MAS NMR spectra of HY (LZY-82) after: A) calcination in air at 500 ~ followed by steam aging with 100% steam at B) 760, C) 788 and D) 815 ~ 110.8 ppm is attributed to extraframework silica (14) and was not included in the computation of Si/AI ratios. After steam-aging at 760 ~ the 29Si spectrum is reduced to a dominant sharp peak near -108 ppm representing a T(4Si,0AI) resonance. Deconvolution of the entire lineshape reveals an additional resonance at -102.0 ppm presumably representing T(3Si, IAI) sites while the second one at -110.4 ppm again indicates amorphous extraframework silica resulting from crystals degradation and fracture (15). The Si/A1 ratio from this spectnan is -- 45 indicating extensive losses of

64

100

50

0

-50

-100

Chemical Shift [ppm] Figure 2. 27A1MAS NMR spectra of HY (LZY-82) after ' A) calcination in air at 500 ~ followed by steam aging with 100% steam at I arm and at B) 760, C) 788 and D) 815 ~ framework AI from the faujasite structure. In these crystals, the T(3Si, IAI) sites present are the main source of Bronsted acidity. In fact, Fripiat and coworkers have reported that a Bronsted site is a bridging OH between an AI and a Si coordinated to three other Si atoms (16). When the steam-aging temperature is increased to 788 and 815 ~ from 760 ~ the Z9Sispectrum becomes sharper and T(3Si, IAI) sites disappear. The spectrum of the steam-aged HY at 815 C was deconvoluted into a sharp line at -107.7 ppm representing T(4Si,0AI) sites together with a broad contribution centered near -108 ppm

65 Table 2. 29Si chemical shifts (-ppm) for several HY-type zeolites steam-aged at 760, 788, and 815 ~ with 100% steam at 1 atm(* approximate value since the signal near -102 ppm is very weak) . . . . . . . . . . . . . . . . . Si/A1 . Sample T(0Si,4AI)T(1Si,3A1) T(2Si,2A1) T(3Si,IA1) T(4Si,0A1) ~ NMR Chem 1.HY 2.HY-760 3.HY-788

87.6 (1.5) . . . . . . . . . . . .

4.HY-815 .

.

.

90.9 95.8 I01.5 (4.6) (14.1) (37.2) . . . . . 102.1 . . . . . (6.6) . . . . . . . .

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

.

.

.

.

.

.

106.9 (37.6) 108.0 (69.5) 108.1 (100) 107.7

110.8 (5.0) 110.4 (23.9) ......

(72)

(28)

4.5

2.8

45*

2.8 2.8

108

---

2.8

....

attributed to amorphous silica generated during crystals degradation by 100% steam.

3.2 27A1 MAS NMR Results As expected the 27Alspectrum spectrum of the calcined HY (LZY-82) contains two main resonances near 0 ppm and 64 ppm representing Al(VI) and AI(IV) species respectively. Inspection of the spectrum in Figure 2A suggests that the resonance at 64 ppm overlaps another centered near 30 ppm attributed to AI(V) species or to highly distorted AI(IV), (17). Recent results obtained using a double-rotation (DOR) spinning technique (7,8) to the study of Al in zeolites, have concluded that extraframework A1 contains, in addition to AI(V) and AI(VI), some AI(IV) species. By comparing 27A1 MASNMR with 27A1 CP MASNMR, it has been shown that in steam-aged HY a substantial portion of the resonance attributed to Al(IV) results from the presence of extraframework Al-species (9). The 27A1spectrum of the stearn-aged HY at 760 ~ contains three resonances arising from the presence of four, five and six coordinated Al-species representing the dealumination of the HY framework resulting from thermal and hydrothermal treatments, Figure 2B. In the 27Al spectra in Figures 2C-2D, it becomes apparent that as the steamaging temperature is increased, additional Al(V) and AI(VI) species are formed at the expense of framework Al(IV). In fact, the resonance near 64 ppm moves toward lower chemical shift values while its intensity monotonically decreases with temperature. For HY steamed at 788 ~ or at 815 ~ the resonance near 64 ppm is attributed mainly to extraframework AI(IV) species; see Figures 1C-1D and Figures 2C-2D. 3.3 Mieroealorimetry Although microcalorimetry cannot distinguish betwen Bronsted and Lewis acidity, it can relate site strength to the nature of the sites present (12,18-20). In zeolites, the strong sites (> 160 kJ/mol) observed at low NH3 coverage, have been attributed to Lewis

66 Table 3. Ammonia Chemisorption Data at 150 ~ (p = 0.2 torr). Si/AI molar ratios are from 29Si NMR data Sample Si In. H. ...Int.H . ID. AI V~t____ Vr___ V~____ kJ/mol __~ HY (LZY-82) HY-760 HY-788 HY-815

2.8 45 ---

1020 319 222 174

340 155 109 98

680 164 113 76

208 158 156 157

122 33 20 15

Table 4. Site population distribution (in Ixmol NH3_/g );. Si/A1 values are from NMR data, Sample Si/AI Acid Site strength in kJ/mol . <100 100-120 120-150 150-200 >200 HY(LZY-82) 2.8 306 187 470 72 11 HY-760 45 108 27 92 42 0 HY-788 -51 25 57 28 0 HY-815 -49 15 14 23 0 centers (18,19) resulting from extraframework A1 species produced during thermal and hydrothermal treatments. Sites of intermediate strength of adsorption (i.e., 130-160 kJ/mol) have been associated with Bronsted centers, resulting mainly from bridging hydroxyls. Heat of adsorption in the 100-130 kJ/mol range represents weak Lewis sites and below 100kJ/mol what observed is the interaction of the probe molecule with weak L-sites or silanol groups (12,18-20). Microcalorimetry (21,22) results have been reported in Table 3 and in Figures 3-4. In Table 3, initial heats (In. H) are in kJ/mol. Integral heats (Int. H) are in J/g. Vt, Vr and V~ are the total NH3 volume (in txmol/g) adsorbed, re-adsorbed, and irreversibly sorbed at p = 0.2 torr respectively. Chemisorption isotherms for ammonia are shown in Figures 3-4. Not shown are sorption isotherms aider NH3 adsorption and degassing in vacuum at150 ~ By subtracting the adsorbed volume of the secondary isotherms from that of the primary isotherms at the same equilibrium pressure (p = 0.2 torr), it is possible to obtain Virr, the volume of irreversibly chemisorbed sorbate. This value is believed to correlate with the presence of strong acid sites (21,22). In Table 3, initial heats have been associated with the strength of the strongest acid sites present (21,22). On the other hand, integral heats represent the total heat of adsorption evolved at p=0.2 torr and are therefore associated with the solids acid site density. The differences in NH3 sorption shown in Figure 3 and in Table 3, reflect the loss of surface area and crystallinity that the calcined HY (LZY-82) experiences after thermal and hydrothermal treatments at MAT conditions. After steaming at 760 ~ the strength of the strongest acid sites and integral heats ofNH3 chemisorption decrease b y - 50 kJ/mol and 89 J/g respectively. In addition, the number of strong acid sites, as determined by Virr values (Table 3), decreases to 164

67

1400i

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

1200 1000 LZY'S2 8O0

" + - HY-]'80 HY-TU

0OO

-,0- HY-818

400 200 0

0,1

0,2

0,8 P (tort}

0.4

(15

0.6

Figure 3. Sorption isotherms for HY (LZY-82) crystals after calcmatlon m mr at 500 ~ followed by steam aging with 100% steam at 1 atm and at 760, 788 and 8 i5 ~

260

Q (kJ/mol) ....

_.

,,

,.

,

. ,

,

~,

,,

- :" LZY-82

, , . ,

~,

,,

~,

~, : _

,

,

,

.....

.,

-4-- HY-Te0

.....

,,

,.,

,

.

._

,. . . . . .

~::_~

- ' ~ HY-TU

-:

--=

. . . . . . .

:

.

~ . . . .

-.B- HY"818

[

1200

1400

200

100

0 t,

0

2O0

4OO ~nla

~X) 800 IOO0 uptak8 (llnol/o)

Figure 4. Differential heat profiles for HY (LZ-Y-82) crystals after calcination in air at 500 ~ followed by steam aging with 100% steam at 1 arm and at 760, 788 and 815 ~

68 l,tmol/g from 680 l,tmol/g in the calcined HY sample. However, when the steam-aging temperature is increased to 788 and 815 ~ from 760 ~ initial heats remain essentially unchanged while integral heats monotonically decrease probably because of losses of framework A1; see Table 3. Depletion of framework A1, following steam aging with 100% steam at 1 arm and at 760,788 and 815 ~ reduces the population of sites with a given strength in a manner shown in Table 4. The corresponding differential heats of ammonia adsorption as a function of coverage are shown in Figure 4. In addition to some strong Lewis sites with strength near 208 kJ/mol, the calcined HY contains two populations of sites, Figure 4. The larger population of sites with strength near 145 kJ/mol is attributed to framework Si(OH)AI groups (12,18-20). The smaller population of site with strength below 100 kJ/mol has been attributed to interaction of NH3 with weak L-sites or with silanol groups (12,18-20). After steam-aging at 760 ~ there is a sharp decrease in acidity that can be safely assigned to losses of Si(OH)A1 groups from the zeolite structure. In fact, the major population of sites with strength near 145 kJ/mol disappears and is replaced by a much smaller one with strength near 130 kJ/mol representing L-sites associated with extraframework Al-species. There are two smaller populations of sites with strength near 80 kJ/mol and 50 kJ/mol attributed to the presence of weak L-sites and silanol groups; see Figure 4. Steam aging at higher temperatures further decreases acid site densities leaving a faujasite structure containing a heterogeneous distribution of site strengths; the strength of these sites monotonically decreases with NH3 coverage; see Tables 3-4 and Figure 4. 4. SUMMARY AND CONCLUSIONS The effect of steam-aging temperature on the physical properties and structural characteristics of HY (LZY-82) have been studied in detail. In general, as the steam aging temperature is increased from 760 ~ to 788 and to 815 ~ there is a progressive decrease in the crystals surface area, pore volume and unit cell size that can be safely attributed to losses of framework AI(IV) from the faujasite structure. In fact, 27A1 NMR spectra have shown that as the severity of the hydrothermal treatment increases, more AI(IV) atoms are transformed into AI(V) and AI(VI) species. In conjunction with 29Si NMR spectral data, the substantial amounts of AI(IV) sites still detected in the 27A1M R spectra suggests the presence of extraframework AI(IV) sites. In agreement with NMR results, microcalorimetry data has shown that steam-aging causes a drastic decreases in acid sites density. Interestingly, even if the number of strong acid sites (as determined from Virr values) decreases with increasing steam-aging temperature, the strength of the strongest acid sites remains practically unchanged near 156-158 kJ/mol. These sites are believed responsible for initiating cracking reactions during gas oil conversion (10). The results of the present study are of relevance for developing an understanding of the equilibration and regeneration process of zeolite-containing fluid cracking catalysts (FCCs). While accelerated aging with steam at MAT conditions cannot reduce all of the physicochemical properties to those in the corresponding equilibrium FCC sample, it produces materials with similar structural characteristics as evidenced from 29Si NMR results (10). Details of this work will be published in a forthcoming paper (10).

69

Acknowledgements Special thanks are also due to Dr. P.S. Iyer for providing NMR data at the beginning of this research and to A.E Schweizer (Exxon) for elemental analysis and XP~ results. This work has been supported by NATO collaborative grant CRG-971497 to MLO and HE.

REFERENCES 1. R. Beaumont and D. Barthomeuf, J. Catalysis 26, 218 (1972) 2. J.W. Ward, "Zeolite Chemistry and Catalysis", J.A. Rabo Ed.; ACS Monograph No. 171, ACS, Washington D.C., 118 (1976) 3. M.L Poutsma, "Zeolite Chemistry and Catalysis", J.A. Rabo Ed.; ACS Monograph No.171, ACS, Washington D.C., 437 (1976) 4. R.A. Beyeflein, G.B. McVicker, L.N. Yacullo, J.J. Ziemiak, J. Phys. Chem., 92, 1967 (1988) 5. R.A. Beyerlein, C. Choi-Feng, J.B. Hall, B.J. Huggins and G.J. Ray, Topics in Catalysis, 4, 27-42 (1967) 6. G. Engelhardt and D. Michel, "High Resolution Solid-State NMR of Silicates and Zeolites", J.J. Wiley, New York, N.Y. (1987) 7. A. Samoson, E. Lippmaa and A. Pines Mol. Phys. 65, 1013 (1988). 8. G.J. Ray and A. Samoson, Zeolites 13, 410 (1993) 9. H. Hong, D. Coster, F.R. Chen, J.G. Davis and J.J. Fripiat "New Frontier m Catalysis", L. Guczi Ed.; Elsevier, Amsterdam, 1159-1170 (1993) 10. M.L. Occelli, M. Kalwei, A. Wolker, H. Eckert, and A. Auroux, J. Catalysis (accepted) 11. R.A. Beyerlein, C. Choi-Feng, J.B. Hall, B.J. Huggins and G.J. Ray, "Fluid Catalytic Cracking III: Materials and Processes" M.L. Occelli and P. O'Connor Eds.; ACS, Washington D.C., 81 (1994) 12D. Chen, S. Sharma, N. Cardona-Martinez, J.A Dumesic, V.A Bell, GD. Hodge, and R.J. Madon, J. Catal. 136, 392-402 (1992) 13. L.A.Pine, J. Catal. 125, 514(1990) 14. M.L. Occelli in "Catalysts m Petroleum Refining and Petrochemical Industries" M. Absi-Halabi et al., Eds.; Studies in Surface Science and Catalysis, Elsevier, Amsterdam, 100, 27 (1995) 15. J. Scherzer, Octane-enhancing Zeo#te FCC: Scientific and Technical Aspects; Marcel-Dekker. New York, N.Y., (1990) 16. AL. Bluemenfeld, D. Coster, J.J. Fripiat, J. Phys. Chem. 99, 15181-15191 (1995) 17. J.P. Gilson, GC. Edwards, A.K Peters, K. Rajagopalan, R. Wormsbecher, T.G Roberie, and M.P. Shatlock, J. Chem. Soc. Chem. Commun. 91(1987) 18. A. Auroux, and Y. Ben Taarit, Therm. Acta, 122, 63 (1987) 19. Z.S. Shi, A. Auroux, and Y. Ben Taarit, Can. J. Chem. 66, 1013 (1983) 20. N. Cardona-Martinez, and J.A. Dumesic, J. Catal. 125, 427 (1990) 21. A. Auroux, "Catalyst Characterization : Physical Techniques for Solid_MateriaN', Ch.22; B. Imelik, J.C. Vedrine, Eds., Plenum press, N.Y. (1994) 22. A. Auroux, Topics in Catalysis, Vol 4, 71-89 (1997).

Studies in Surface Science and Catalysis 134 M.L. Occelli and P. O'Conner (Editors) 9 2001 Elsevier Science B.V. All rights reserved

71

Effect of catalyst properties and feedstock composition on the evaluation of cracking catalysts A. A. Lappas a, Z. A. Tsagrasouli a, I. A. Vasalos a and A. Humphries b a Chemical Process Engineering Research Institute (CPERI) and Department of Chemical Engineering University of Thessaloniki, P.O. Box 361, 57001 Thermi, Thessaloniki, Greece b Akzo Nobel Catalysts, Inc., 2625 Bay Area Blvd., Suite 250, Houston, TX 77058 USA The objective of the present paper is to determine the effect of catalyst properties on the FCC product yields and to investigate the interaction between FCC catalysts and feedstocks. The work was carried out in a MAT unit using 12 new catalysts and various types of feedstocks. Main emphasis was given in the investigation of the effect of catalyst accessibility on FCC product yields. The study shows that catalyst accessibility is an important property that affects the catalyst activity and the selectivity of the FCC products.

1. INTRODUCTION Worldwide oil production has grown significantly in this decade and it is going to reach 3.5 billion tons in the first year of the millenium [ 1]. Transportation fuels are the main products from the refineries and they are going to develop further in the future. Furthermore, the new very strict requirements for transportation fuel specifications for the year 2000 and especially for the year 2005 puts the fuel market, supplied by refineries, into a new era. Refineries must increase profitability by processing more residual oil and by producing fuels of improved quality. The refinery of 2000-2005 must invest in new processes where new catalysts must be used. Among refinery processes Fluid Catalytic Cracking (FCC) plays a very important role. On the one hand FCC produces a large portion of the total gasoline and diesel and on the other hand these products contain a lot of undesirable components (sulfur, olefins, aromatics etc.). For these reasons research today emphasizes the development of a new FCC process and new FCC catalysts. The process modifications to FCC in the recent years include the conversion of the conventional FCC units to short contact time units (SCT). Currently, the revamping of many FCC units to SCT FCC means new riser termination technologies or/and improved feed nozzles or/and higher stripper efficiency [2]. FCC's in Europe employing some of the above modifications have increased to 18 by 1998 and are expected to reach 25 by the year 2002. This number is approximately one third of all FCC's in Western Europe [3]. The benefits of SCT-FCC are mainly better dry gas and coke selectivity, more conversion, more throughput and more resid in the feed [3]. However, the optimization of new SCT FCC requires new FCC catalysts [4,5,6]. The FCC catalysts are considered to be equally important to the various technological advances in FCC. Today's FCC catalysts are mixtures of functional components (zeolite, additive, binder, filler) customized to maximize FCC profitability within hardware constraints. The new SCTFCC catalysts are required to equilibrate to higher microactivities than FCC conventional

72 catalysts and are designed to have higher RE content, more zeolite content and more active matrix [6]. The new improved SCT-FCC catalysts can also play an important role in the production of reformulated fuels [7]. For these reasons the catalyst manufacturing companies design new zeolites with optimum aluminum topology, new types of zeolites and matrixes, optimum zeolite to matrix ratio and optimum matrix activity [7,8,9,10]. The target of all these modifications is the synthesis of new catalysts that give more selectivities and better stability. One catalytic feature that is especially important for the optimal performance of FCC catalysts is the accessibility of its active sites. This property has special importance as refineries process heavier FCC feeds. According to literature [9,11 ] accessibility is defined as the catalyst ability to have active sites accessible to large molecular structures which are supposed to interact with these sites within a certain time limitation set by the catalytic process. Catalysts with high accessibility maximize large molecule cracking and minimize secondary reactions such as gasoline overcracking, hydrogen transfer and coke formation. One more advantage of high accessibility is better strippability of catalyst and thus the reduction of coke yield. Nowdays refiners process a wide range of FCC feedstocks while moving toward heavier feeds. The hydrocarbon molecular types in the FCC feedstock are sometimes the most important variables in determining the basic yield structure in FCC [ 12]. For example, more resid in the feed or high aromatics in the feed affect strongly the gasoline and coke selectivity [13]. Sulfur content in the feed influences the final sulfur distribution in the FCC products [12,13]. Optimum catalyst performance can be obtained by matching specific catalyst components with selected feed characteristics [ 14]. The target of the present paper was to investigate the effect of different catalysts and feedstocks on FCC product yields and to establish the interactions between FCC catalysts/feedstocks. Twelve new Akzo Nobel catalysts with different properties and five different feedstocks were evaluated. Special emphasis was given to the effect of catalyst accessibility on the activity and selectivity using catalysts with otherwise equivalent properties but with different accessibilities. The experimental work was carried out in a MAT unit under a range of WHSV values. In the literature, different methods have been proposed for the evaluation of SCT catalysts. Although it is generally accepted that a circulating FCC pilot plant is the most suitable tool, various microscale reactors are suggested in the literature (fixed fluid beds or modified MAT units) [15,16,17,18,19]. 2. EXPERIMENTAL PROCEDURE

The present study was mainly performed in the MAT unit of Chemical Process Engineering Research Institute (CPERI) using a modified ASTM D 3907-80 procedure [20]. Two temperatures (T=521 and 566 ~ two catalyst residence times (tc=50, 20s) and varied C/O ratios were used for the MAT tests. For the validation of the MAT unit a comparison study of MAT results with CPERI FCC pilot plant results was also carried out. The CPERI FCC pilot plant is a continuous circulating FCC unit designed to simulate the commercial FCC performance. A full description of the CPERI FCC pilot plant is presented elsewhere [21]. Twelve new Akzo Nobel catalysts were evaluated in this study. These catalysts have a wide range of properties but all contain an active matrix. The catalysts 1,3 and 5 are low rare earth (LRE) catalysts with RE203=0.075 wt%. Catalysts 2,4 and 6 have similar Zeolite to Matrix surface areas (Z/M) to catalysts 1,3 and 5 but they have a high RE content (HRE, RE203=2.3, 4.3, 5.3 %wt for the three catalysts respectively). Catalysts 1 through 6 are low accessibility (LA) catalysts with Akzo Accessibility Index-AAI=4 while catalysts 7 through 12 have

73 similar properties (RE, Z/M) but they are high accessibility (HA) catalysts (AAI=26). Catalysts 1,2,7 and 8 have a relatively low Z/M ratio (LZ/M---0.7). Catalysts 3,4,9 and 10 have a high Z/M ratio (HTJM=2.6) while catalysts 5,6,11 and 12 a medium Z/M ratio (MTJM=I.7). Catalysts 1,2,7 and 8 have a low total surface area (LSA=160m2/g), catalysts 3,4 have a medium SA (MSA=185m2/g) while the remaining catalysts have a high SA (HSA>220m2/g). The catalyst properties are summarized below: Table 1 Catalyst properties RE Matrix Access Z/M SA

1 L L L L

2 H L L L

3 L L H M

4 H L H M

5 L L M H

6 H L M H

7 L H L L

8 H H L L

9 L H H H

10 H H H H

11 L H M H

12 H H M H

All catalysts have been deactivated at 795~ for 6 hours (100% steam, 1 atm) in Akzo Nobel laboratories using standard steaming procedures. The catalysts have zero metals loading and so all results presented in this work may differ with a resid feed. For comparison, one commercial catalyst (base catalyst) with low metals content was also used in this study. Five different types of FCC feeds were used in this work. A naphthenic feed (A) was used as standard (or base feed). Feed B is a high sulfur vacuum gas oil while feed HYD-B is the corresponding hydrotreated feed. Feed C is a highly aromatic feed and feed HYD-C is the corresponding hydrotreated feed. Some key properties of the feeds are given below: Table 2 Feedstock properties A B HYD-B C HYD-C API 18.7 20 26.8 13.1 17.3 CA* (ndm), %wt 24 22.3 10.3 49.5 26 Sulfur, %wt 2.51 2.14 0.106 0.78 0.184 *Aromatic carbon content (%0f total carbon) determined bythe method ASTM D-3238 3. RESULTS AND DISCUSSION 3.1. Comparison of MAT and pilot plant data In order to validate the MAT unit a comparison of CPERI FCC pilot plant results and MAT results was carried out. The CPERI FCC pilot plant has been in operation for the last 5 years and its performance has been validated with much commercial data. For the comparison of MAT and FCC pilot plant, the base catalyst was evaluated in both units using all feedstocks of Table 2. For this comparison a unit product factor was used for each FCC product and feed. This factor is defined at constant conversion and gives the ratio of each product yield obtained by a feedstock in relation to the same product yield from the base feed at the same unit. For example the gasoline factors for the two units are:

Gasoline MATfactor: gasoline yield for feedj /gasoline yield for base feed Gasoline Pilot plant factor: gasoline yield for feedj /gasoline yield for basefeed These factors are presented for each one of the five feeds in the Figures 1 and 2 for the C/O ratio and the gasoline yield using a parity plot of the corresponding factor from the two units.

74

C/O factor, Catalyst: Base, T=970~

3.0

at 60%wt conversion

xA liB

2.5

J

El HYD-B O .ca O

eC

2.0

o HYD-C

J

1.5

J

,i..a O ,..-, ~

1.0

Y

0.5

0.0 0.0

0.5

1.0

1.5

2.0

2.5

3.0

MAT factor Fig. 1. Comparison of M A T and F C C pilot plant for C/O factor

Gasoline factor, Catalyst: Base, T=970~

1.6 xA 1.4

at 60%wt conversion

'

LmB

/

O HYD-B

1.2

eC O

o HYD-C

1.0

O

0.8 O ~

jf

0.6 0.4 0.2 0.0

f/

J 0.0

0.2

0.4

0.6

0.8

1.0

1.2

1.4

1.6

MAT factor Fig. 2. Comparison of M A T and F C C pilot plant for gasoline factor

75 Figure 1 shows that with the exception of the heavy feed (C), the C/O factors from the two units agree. Satisfactory comparison between the two units exists for the gasoline (Figure 2), the other product yields and the olefinicities factors as well (not presented in figures). Large differences exist in dry gas factors. This was expected since the fixed bed (MAT unit) is not as suitable as the riser (pilot plant) is for dry gas yields predictions. The rank of the various products produced from all feeds in the two units is identical. For example the crackability order of all feeds is the same in the two units (HYD-B>B>A>HYDC>C). Generally, it can be concluded that the evaluation of the feedstocks is similar in the two units for the majority of the FCC products. In a future paper a similar comparison study using some of the new catalysts will be presented.

3.2. Effect of catalyst accessibility on zeolite retention Based on the measurements of zeolite surface areas (by t-plot method) of fresh and steamed catalysts, we concluded that the catalyst accessibility affects the zeolite retention. By zeolite retention we mean the percentage % of zeolite surface area (ZSA) that remains after steaming in relation to ZSA of the fresh catalyst. Figure 3 presents the difference in the zeolite retention between an HA and the corresponding LA catalyst after steaming (at the same high rare earth level). After deactivation, independently of the type of catalyst (HRE or LRE, HM or LM), the HA catalysts maintain more zeolite than the corresponding LA catalysts. Figure 3 also shows that the zeolite retention depends on the ZSA of catalysts and it is higher for catalysts with higher ZSA. Thus, after deactivation, a HA catalyst with a high RE content maintains 12 % more zeolite than the LA catalyst (with similar the other properties). This result is in accordance with previous studies in Akzo Nobel [9,11 ]. In those studies, using a more realistic catalyst deactivation protocol (metals deposition in a Cyclic Deactivation Unit), it was also concluded that a high accessibility catalytic system give the best zeolite protection. The better ZSA retention at higher AAI is not yet fully understood. It is believed that for the HA catalysts there is a large number of acid sites and hydroxyl groups available which provide alternate sites for poisons to attack, therefore affording the zeolite extra protection.

=

13 X

11 Q

9

f

f

J

7

1~

1~

180

2~

2~

3~

Zeolite SA of fresh catalyst (ma/g) Fig. 3. Effect of catalyst accessibility on zeolite retention (catalysts at same HRE )

76 3.3. Effect of catalyst properties on activity The effect of catalyst properties on catalytic activity (the required C/O for achieving a constant conversion) was studied in the MAT unit under two different experimental conditions: i) T=566~ (1050~ tc=20s (SCT-MAT conditions) and ii) T=521~ (970~ tc=50s (MAT conventional conditions). For the catalyst evaluation at the SCT-MAT conditions the standard feed (feed A of Table 2) was used. The experimental data from this study are presented in Figure 4. The eight most promising catalysts obtained by these results (Figure 4) were further evaluated using 3 feedstocks (A, HYD-C, C of Table 2) in order to establish catalyst/feedstock interactions (the work with the other feeds will be presented in a future paper). In this part of the study the MAT conventional conditions were used since the very refractory feeds (feed HYD-C and especially feed C of Table 2) created operating problems in the unit when the very low run time of 20s was used. Moreover, the use of the highly refractory feed C necessitates comparing yield data at a relatively low conversion level (55%wt). The evaluation of the most promising catalysts with these 3 feeds is presented in Figure 5. Figures 4 and 5 reveal that HA catalysts are always more active than the corresponding LA catalysts. For all types of catalysts (high matrix-HM or low matrix-LM) the HRE catalysts are more active. Catalysts with more active components (Z+M) or higher Z/M ratio are also more active than catalysts with lower active content or lower Z/M. The effects of catalyst accessibility and RE content on catalyst activity are summarized in Figure 6. It seems that at the same RE content the HA catalysts require about 0.5 less C/O than the corresponding LA ones. The effect of the RE content seems to be the same for both HA and LA catalysts. Above a certain RE concentration the effect of RE on catalyst activity is not so strong (Figure 6). Comparing the catalyst ranking from Figures 4 and 5 it seems that the catalyst activity ranking does not depend on the MAT conditions used in the present study. Figure 5 shows that the ranking of catalyst activity is the same for all feeds. Independently of the aromaticity of the feed, the rank of catalyst activity is: cat. 12>cat. 10>cat. 6>cat. 8>cat. 4>cat. 2>cat. 11 >cat.5 >cat. 9 >cat. 7>cat.3>cat. 1

For all feeds, HA catalysts are always more active than the LA catalysts. However, if we examine the difference in the C/O ratio (delta C/O) between a LA and a HA catalyst (with similar the other properties) we can see that catalyst accessibility improves catalyst activity if the feed is more aromatic. This result is summarized in the Table 3. The performance of an HA catalyst in relation to the corresponding LA is much better when the feed has a higher aromatic content. Table 3 Interaction of feedstock type and catalyst accessibility On catalytic activity Delta C/O (LA-HA) Cat.5-Cat. 11 Feed type Cat.6-Cat. 12 Cat.4-cat. 10 HYD-B 0.2 Not Available-NA 0.05 B 0.4 NA 0.4 A 0.7 0.4 0.3 HYD-C 0.7 0.6 0.3 C 1.1 1.3 0.3

Cat.9-Cat.3 NA NA 1.6 1.6 2.1

77

HA

HA

LA

HA

HA

LA

LA

LA

HA

HA

LA

LA

HRE

HRE

HRE

HRE

LRE

HRE

HRE

LRE

LRE

LRE

LRE

LRE

MZ/M

H Z / M M Z / M LZ/M M Z / M H Z / M LZ/M H Z / M H Z / M LZ/M

HSA

HSA

HSA

LSA

HSA

MSA

LSA

HSA

HSA

LSA

H Z / M LZ/M MSA

LSA

Fig. 4. Effect of catalyst properties on C/O ratio

HA

HA

LA

LA

HA

LA

HA

LA

HRE

HRE

HRE

HRE

LRE

LRE

LRE

LRE

HSA

HSA

HSA

M SA

HSA

HSA

HSA

M SA

MZ/M

HZ/M

MZ/M

HZ/M

MZ/M

HZ/M

HZ/M

HZ/M

Fi~. 5. Effect of catalyst/feedstock on C/O ratio

78 The above results are consistent with those from a similar literature study [ 14]. In that work as an FCC feed increased in molecular weight a catalyst with an active matrix become much more active (changes in the ranking were not mentioned). In the present study, since all catalysts have active matrixes while all feeds have similar molecular weights, we can conclude that there is a relation between catalyst accessibility and feed aromaticity. Thus, the high accessibility catalyst activity increases the most when feeds with higher aromaticity are used. The catalyst screening presented in Figures 4 and 5 revealed that the majority of the new catalysts perform better than the base catalyst. Moreover, the new catalysts have a wide range of activities. As discussed above, the magnitude of the catalyst activity depends strongly on the properties of feed being cracked. However, for all feeds of this study catalyst 12 is the most active. Catalyst 12 has a high matrix and a high zeolite content along with high RE content and high accessibility. MAT results, Feed: A, T---970~ tc=50sec, at i5%wt conversion ....

[] LA catalysts9 HA catalysts-

[]

\

L)

3

A v

0

2

4

REzO3 of fresh catalyst (%wt) Fig. 6. Effect of catalyst accessibility and RE content on C/O ratio

3.4. Effect of catalyst/feedstock properties on FCC product selectivities

Gasoline selectivity The gasoline selectivities (gasoline yields at the same conversion) for all catalysts are presented in Figure 7. As expected, the gasoline yield from HRE catalysts (gasoline catalysts) is always higher than the corresponding LRE catalysts (octane catalysts) with similar properties. This happens for both LA or HA, HM or LM catalysts. The most important result from this study is that the HA catalysts are always more gasoline selective than the corresponding LA ones. The difference in gasoline selectivity between a LA and a HA catalyst is higher when the two catalysts are octane catalysts (LRE). Among catalysts (HRE or LRE, HA or LA) with the same matrix more gasoline selective are those with higher zeolite content (higher ZA4 ratio). The zeolite SA of a catalyst is a key parameter for increasing gasoline selectivity. The above effects are summarized in Figure 8 where it seems that for catalysts

79 with the same zeolite SA the HA and HRE catalysts give the maximum gasoline selectivity. It is noticeable from Figure 7 that all the new catalysts present higher gasoline selectivity than the base catalyst. Among all catalysts, cat. 12 has not only the maximum catalyst activity but the maximum gasoline selectivity as well. The effect of feedstock type on catalyst evaluation with respect to gasoline selectivity is presented in Figure 7. The type feedstock influences strongly the gasoline yield. For the same catalyst, the more refractory the feed, the less gasoline is produced. The highly aromatic feed C gives very low gasoline yields for all catalysts tested. The effect of FCC feed hydrotreating is obvious from Figure 7. For all types of catalysts, hydrotreating an FCC feed causes a strong increase (about 30%) in gasoline yields. From Figure 7, it is also concluded that, for the feedstocks tested, the type of feed does not influence the ranking of catalysts. Catalyst evaluation with the base feed (feed A) and the more aromatic hydrotreated-C (feed HYD-C) indicates that when a more aromatic feed is used the significance of catalyst accessibility on gasoline yield increases. However, going to the results with the very aromatic non hydrotreated feed C it seems that with this feed the catalysts present very small variations in gasoline selectivity. This means that the discrimination of catalyst with respect to gasoline selectivity is much more difficult when very heavy feeds are used. For these heavy aromatic feeds the proper choice of catalyst is more strongly governed by other refinery objectives and unit operation [20].

HA

LA

HA

LA

HA

LA

HA

LA

HRE

HRE

HRE

HRE

LRE

LRE

LRE

LRE

HSA

HSA

HSA

MSA

HSA

HSA

HSA

MSA

MZ/M

MZ/M

HZ/M

HZ/M

MZ/M

HZ/M

HZ/M

HZ/M

Fig. 7. Effect of catalyst/feedstock on gasoline selectivity

80 MAT results, Feed: HYD-C, T=970~ tc=50sec, at 55%wt conversion 40.0

_

!

39.5 ~ [] LA catalysts 9 HA catalysts 39.0

~_~..~-~

/ i

-~.

_

38.5 o~

S

38.0 37.5

o

jJ

37.0

r,r

36.5 -

36.0 35.5

j

Z

i..a

35.0 1~

160

2~

220

2~

2~

Zeolite SA of fresh catalyst (m2/g) Fig. 8. Effect of catalyst properties on gasoline selectivity

Coke and dry gas selectivity The effect of catalyst properties on coke selectivities (coke yields at a constant conversion) is given in Figure 9. The HRE catalysts are always less coke selective than the corresponding LRE ones (with similar the other properties). Hydrogen transfer reactions in HRE catalysts force the production of coke from aromatic compounds. By comparing catalysts with the same active content (Z+M), the HM catalysts give more coke. This result is valid for HRE or LRE and HA or LA catalysts. Going to the other heavier feeds it seems that the coke selectivity of each catalyst is feedstock dependent. The degree of feed aromaticity has a very strong influence on coke yields. For example, feed C produces two times more coke than the lighter feeds A and HYDC. Figure 9 shows also that by hydrotreating an aromatic FCC feedstock 50% less coke is produced (feed HYD-C vs. feed C). The role of feed aromaticity in determining the coke selectivity or the coke-to-bottoms selectivity has been recognized by other researchers [14]. By examining the results of coke yields vs. HCO yields of the various catalysts and feedstocks used in this study (not presented in figures), it seems that the catalysts tested show very small variations when the feed C is used. In contrast, the light feed HYD-C differentiates better the catalysts. According to literature [ 14] feedstock effects on coke selectivity are clarified when the catalysts used contain matrices with different activities and when paraffmic feeds are used. In our case all catalysts have active matrix and thus the differentiation of catalysts in respect to coke selectivity is better for the lighter HYD-C feed. The evaluation results for the dry gas (H2, C1 and C2's) yield (not presented in figures) concluded that the RE content of catalysts is the most important parameter that influenced the dry gas yields. Independently of Z/M ratio and the catalyst accessibility the dry gases from HRE catalysts are always lower than that from the catalysts with LRE content. Among the dry gases, the most important effect of RE content is on H2 yield. The hydrogen from a high RE catalyst is about 60-70% lower than that from a low RE one. The corresponding reduction of

81 the total dry gases is 20%. The effect of catalyst accessibility on dry gases was not clear in the present study since the differences between the catalysts in dry gas selectivity were very low (2.2-2.6%wt). The feed effects on dry gas yield were more pronounced than the catalyst effects. The hydrotreating of FCC feed minimizes the dry gas yields. The study of the effect of feedstock type on catalyst ranking, in respect to dry gases, shows that there are similarities in the evaluation data between the two lighter feeds A and HYD-C. However, it seems that the more aromatic feed HYD-C discriminate better the accessibility of the catalyst in relation to the feed A. For the HYD-C all the HA catalysts show lower dry gases in relation to the corresponding LA catalysts (with similar the other properties).

LA LRE

HA LRE

HA HRE

HA LRE

LA LRE

HA HRE

LA HRE

LA HRE

HSA HZ/M

HSA MZ/M

HSA MZ/M

HSA HZ/M

MSA HZ/M

HSA HZ/M

MSA HZ/M

HSA MZ/M

Fig. 9. Effect of catalyst/feedstock on coke yield

LPG gases selectivity In contrast to gasoline, LPG ( C 3 ' S, C 4 ' s) selectivity (not presented in figures) is higher for LRE catalysts. As expected, the C3 and Ca olefinicities from the LRE catalysts are higher than from the HRE ones. The effect of type of feedstock on the olefinicities shows that for all type of feedstocks the HRE catalysts give lower olefinicities. The accessibility property of the catalysts affects also the olefinicities. For all feedstocks the HA catalysts give higher olefinicities in relation to the LA ones. The above results are summarized in Figure 10. Figure 10 shows that by increasing the RE content of a catalyst the olefinicities decrease. However, for a constant RE content the HA catalysts give more olefins. It is clear that the high pore volume of the HA catalysts assist for less hydrogen transfer reactions. These reactions contribute to lower olefin production. The ranking of catalysts in relation to olefinicities is not feedstock dependent at least for the LRE catalysts. For the HRE catalysts the values of olefinicities are low and general results for the feedstock effects can not be concluded.

82 MAT results, Feed: A, T=970~ tc=50sec, at 55%wt conversion

0.70 0.66

l rn LA catalysts]

0.62 9

l " HA catalysts]

0.58 ._o

\

0.54 0.50

\

O

0.46 w

0.42

[]

[]

0.38 0.34 0.30 0

2

4

6

RE203 of fresh catalyst (%wt) Fig. 10. Effect of catalyst/feedstock on Cn olefins

LCO selectivity The results for LCO yield, at a constant conversion, are presented in Figure 11 for the three feeds investigated. Figure 12 summarizes the effects of two catalytic properties (accessibility and Z/M ratio) on the LCO yield for the base feed-A. This figure refers to HRE catalysts although the results are also valid for LRE ones. Among catalysts with the same Z/M ratio and RE content, the HA catalysts give higher LCO yields. This effect of accessibility is valid independently of the type of catalyst (HRE or LRE, HZ/M or LT_/M). Figure 12 illustrates also the impact of Z/M on LCO yields. Among catalysts with the same accessibility and RE content the catalysts with the lower Z/M produce higher LCO yields. The RE content of the catalysts influences also the LCO yield. For both HZ/M and LZ/M catalysts the HRE ones produce more LCO (Figure 11). The study of the effect of type of feedstock on LCO yield (Figure 11) shows that there is a strong influence of feedstock aromaticity on LCO yield. An" increase of the aromaticity results in a reduced production of LCO. A hydrotreated aromatic FCC feed (HYD-C) gives about 20% more LCO than the raw feed (feed C). The effects of catalyst properties on LCO yield, discussed above, are valid for all feedstocks. However, it seems that an interrelation exists between catalyst Z/M ratio and feed aromaticity (Table 4). It is concluded that the effect of high matrix (low Z/M) on LCO yield is more pronounced in the cases where more aromatic feeds are treated. Table 4 Interaction of feed aromaticity and catalyst Z/M ratio on LCO yield Feed type Delta LCO yield (LTJM-HZ/M catalyst) A 1.4 HYD-C 2.2 C 3.2

83

HA HRE HSA MZ/M

HA HRE HSA HZ/M

HA LRE HSA MZ/M

LA HRE HSA MZ/M

HA LRE HSA HZ/M

LA HRE M SA HZ/M

LA LRE HSA

LA LRE M SA HZ/M

HZ/M

Fig. 11. Effect of catalyst/feedstock on LCO yield MAT results, Feed: A, T=970~

22 21

I

9 HM catalysts [] L M catalysts

20 19

tc=50sec, at 5 5 % w t conversion

.

.

.

.

.

.

.

18

17

9 C)

16 15 14 13 12 4

26

AKZO Access~ility Index (AAI) of fresh catalysts Fig. 12. Effect of catalyst accessibility on LCO yield (HRE catalysts)

84

3.5. Development of short form models Using the MAT experimental data of this work, short form models were developed, based on regression analysis studies. By short form models we mean mathematical correlations which can predict FCC conversion and the main FCC product yields as a function of the experimental parameters and catalyst properties. The zeolite surface area, the matrix area, the Z/M ratio and the UCS/RE ratio are properties, which, as it was discussed previously, affect the FCC product yields. Catalyst accessibility was recognized also, in this work, as a very important property that affects the activity and selectivity. However, catalyst accessibility should be taken into account quantitatively, for the development of the short form models. For this reason, and in order to discriminate the accessibility level of each individual catalyst, an attempt was performed to correlate the accessibility with the pore size distribution (PSD) and the pore volume of the catalysts. Based on the measurements of pore size distribution (PSD) of all catalysts (not presented in figures) the differences between a HA and the corresponding LA catalyst were established. LA catalysts present a very sharp peak at about 40A. However, although 3 different types of LA catalysts (cat. 1, cat.3, and cat.5) have different pore volume they have similar PSD and thus a discrimination between them can not be carried out based only on PSD measurements. HA catalysts show a very wide peak at a pore size that seems to be depended on the type of catalysts. Cat.7, cat.8 have a maximum at 60A while the other HA catalysts have the maximum at about 100A. Cat.9 that has the maximum pore volume has the maximum peak at about 150A. Comparing the pore size distribution diagrams for all HA catalysts it was concluded that it could be possible to correlate a degree of high accessibility between them based on PSD and pore volume measurements. However, the development of this correlation needs further investigation. Akzo Nobel has developed a proprietary test by which they can measure the relative rates of mass transfer of key hydrocarbons within FCC catalysts at nonsteady state conditions. The result is known as the Akzo Accessibility Index (AAD. This index was very helpful in understanding deeper the catalyst design parameters [22]. However, for the simplicity of this study and in order to introduce the accessibility property quantitative in the short form models two common accessibility factors (ACF) were defined for all the LA and all the HA catalysts respectively. By assumption it was considered that all the LA catalysts have ACF=I. The ACF of the HA catalysts will result from the regression analysis and will give a quantitative estimation of the effects of accessibility on FCC products. For the determination of the short form models, regression analysis studies were performed and suitable software was developed based on the Levenberg-Marquard algorithm. It is clear that this analysis depends on the choice of the function, which will be used. In this study many different mathematical functions were applied but linear functions were finally selected since they were adequate and simpler than the non-linear models. For the determination of the parameters only the statistically significant variables were considered while for the selection of the best model the F-test was applied at 95% confidence level. In the following the finally selected functions are given for the conversion and gasoline: Conversion (%wt)/(l OO-conversion %wt) = 0.00225 *SAz +O.OO448*SAm+O. 03399*100*RE/USC] *WHSV ~ *(C/O)~ , A CF= 1.09 Gasoline yield (%wt)= [119.562 +1. 703*100*REMJSC+ l.207*ZZM] *WHSV ~ A CF=1.055 It was found that the high accessibility catalysts are always more active than the LA catalysts and have an ACF=I.09. In addition the HA catalysts give higher gasoline yield than

85 ACF=l.055. The validation of the above correlations is generally difficult since there are not available data in literature. However, in the reference [ 11 ] two catalysts, a LA and a HA (with same other properties) are compared. From this comparison the above-defined ACF is estimated as 1.098 which is very close to that suggested in the present study. 4. CONCLUSIONS Twelve new SCT-FCC Akzo Nobel catalysts were evaluated in the present study using a MAT unit. The MAT unit was validated by comparing its results with FCC pilot plant results. The objective of the work was to establish catalyst/feedstock interactions on FCC product yields. It was concluded that catalyst accessibility is a key catalytic property affecting catalyst activity and FCC product selectivities. Under the same MAT operating conditions a high accessibility (HA) catalyst gives 1.09 times more conversion than a LA catalyst one. For similar RE content and Z/M ratio, the HA catalysts are always more gasoline selective than the LA ones. The difference in gasoline selectivity between a LA and a HA catalyst is higher when the two catalysts are low RE catalysts. The LCO selectivity is also influenced by catalyst accessibility. Independently of the type of catalyst (HRE or LRE, HZ/M or LZ/M) HA catalysts give higher LCO yields. Another effect of catalyst accessibility was found on the C3 and Ca olefinicities. At the same RE content, the HA catalysts give higher olefinicities than the LA ones. The rare earth content and the Z/M ratio are two other significant catalytic properties. Among catalysts with similar properties the high RE ones are more active and more gasoline and LCO selective than the LRE catalysts. The rare earth content affects also the dry gases and especially the hydrogen selectivities. A HRE catalyst gives considerably less hydrogen than a LRE one. The catalyst Z/M ratio has a significant impact on LCO yields. Among catalysts with the same accessibility and RE content the catalysts with the lower Z/M produce higher LCO yields. The catalyst evaluation was also performed with a variety of different feedstocks. The type of feedstock has a strong influence on catalyst activity and product selectivity. Generally the more aromatic feed has a lower crackability and gives less gasoline and more coke. However, the catalyst ranking in respect to activity and gasoline selectivity seems to be independent of type of feedstock. The effect of catalyst accessibility on catalyst activity is found to be more pronounced when more aromatic feeds are used. AKNOWLEGMENTS The authors express their appreciation to Dr. G.A. Huff of BP-Amoco for providing samples of the standard catalyst and the three feedstocks.

REFERENCES 1. 2. 3. 4.

Courty, P.R., Chauvel, A., Catalysis Today, 29, 9, 3 (1996). Upson, L., AIChE Spring Meeting, N. Orleans (1998). Refining catalyst News, Grace Davison refining Catalyst Europe, Issue No.2, April 1998. Hemler, C.L., Sexton, J.A., Schnaith, M.W., Sexson, P.A., Bartholic, D. AIChE Spring Meeting N. Orleans (1998). 5. Hunt, D.A., AIChE Spring Meeting, Houston (1999).

86 6. Humphries, A., Skocpol R.C., Smit, C.P., O'Connor, P.O, Akzo Nobel Catalysts Symposium, The Netherlands (1998). 7. Humphries A., Yanik, S.J., Gerritsen, L.A., P. O'Connor, P., Desai, P;.H., Hydrocarbon Processing, April (1991). 8. Diddans, P., Nee, J., Paloumbis, S., Grace Davison Europe FCC Technology Conference, Heidelberg (1996). 9. O'Connor, P., Veflaan, J.P.J., Yanik, S.J., Catalysis Today, 43, 305-313 (1998). 10. Haave, H., Diddams, P.A., Grace Davison Europe FCC Tech. Conference, Lisbon (1998). 11. O'Connor, P., Humphries, A.P., 206 th ACS Meeting, Chicago, IL, August 22-27 (1993). 12. Fisher, I.P., Applied catalysis, 65, 189-210 (1990). 13. Lappas, A.A., Iatridis, D.K., Vasalos, I.A., Catalysis Today, 50, 73-85 (1999). 14. Harding R.H., Zhao, X.Deady J., NPRA Annual Meeting, San Antonio (1997). 15. Pimenta, R., Quinones, A.R., Imhof, P., Akzo Nobel catalysts Symposium, The Netherlands (1998). 16. Pearce, J., Keyworth, D., Humphries A., Quinones A., AIChE Spring Meeting, N. Orleans (1998). 17. Wallenstein, D., Harding, R.H., Witzer, J., Zhao, X., Appl. Cat. A: Gen., 167, 141 (1998). 18. McLean J.B, Andrews, R.W. AIChE Spring Meeting, N. Orleans, March 10 (1998). 19. Stockwell, D.M., Wieland, W.S., Himpsl, F.L., AIChE Spring Meeting (1998). 20. Lappas, A.A., Patiaka, D.T., Dimitriadis, B.D., Vasalos, I.A., Applied Catalysis, A:General, 152, 7-26 (1997). 21. Vasalos, I.A., Lappas, A.A., Iatridis, D.K., Voutetakis, S.S., 5th International Conference on Circulating Fluidized Beds Beijing China 28-31 May (1996). 22. Hodgson, M.C.J., Looi, C.K., Yanik, S.J., Akzo Nobel catalysts Symposium, The Netherlands (1998).

Studies in Surface Science and Catalysis 134 M.L. Occelli and P. O'Conner (Editors) 9 2001 Elsevier Science B.V. All rights reserved

87

Study on the deactivation-aging patterns of fluid cracking catalysts in industrial units F. Hern~mdez-Beltran, E. L6pez-Salinas, R. Garcfa-de-Le6n, E. Mogica-Martinez, J. C. Moreno-Mayorga and R. Gonz~tlez-Serrano Programa de Investigaci6n en Tratamiento de Crudo Maya, Instituto Mexicano del Petr61eo, Eje Central Ldzaro Cdrdenas 152, C.P. 07730, M~xico, D.F., M~xico e-mail" [email protected]; Density graded fractions obtained from three equilibrium catalysts were studied. The zeolite (REUSY) in such catalysts was the same thus allowing the study of their deactivation and aging under different conditions. The deactivation pattern of these catalysts could be depicted by correlating the loss of aluminium (unit cell size shrinkage) of the zeolite with the surface area loss. This correlation has the advantage of being independent of the time on stream or other operating variables (e.g. the Ni content in the feed or the fresh catalyst makeup rate) in the unit and gives good directional influence of the particular operating conditions in a FCCU upon catalyst aging and deactivation. The zeolite framework collapse and zeolite dealumination were considered from the mechanistic and the kinetic point of view and discussed in terms of V poisoning and of the influence of operating conditions in the FCCU. Properties of laboratory deactivated samples were systematically compared.

1. INTRODUCTION Deactivation and aging are important phenomena observed in fluid cracking catalysts during their industrial usage. Due to the importance of these phenomena catalyst, laboratory-testing methods incorporate deactivation and aging protocols to get a more accurate assessment of the catalytic properties and to the potential benefits of catalysts. Laboratory deactivation protocols have been mainly developed by taking the properties of equilibrium catalysts (Ecat) as a reference. An Ecat is actually a complex blend of catalyst particles at different stages of deactivation or of different ages. This blend is produced as a consequence of the periodic addition of fresh catalyst to the catalyst inventory to maintain conversion level. The changes of catalyst properties while aging mainly concern loss in activity, decrease in pore volume and increase in metal levels. Although the laboratory deactivation protocols currently reported in the literature may emulate the average properties of the catalyst at equilibrium, the effects of aging are more difficult to reproduce. This can be explained since the active

88 components (zeolite and matrix) as well as the contaminant metals in the catalyst, age at different rates and following different mechanisms. To better approach the properties of an Ecat it has been recommended to simulate the "age distribution" profile by producing composite samples from samples deactivated at different severities [1,2]. McLean et al. [1] showed that 10% fresh catalyst added to steamed catalysts simulated in a very accurate manner the catalytic properties of an Ecat. The method mentioned previously might fit well for low metals catalyst deactivation. However, metals poisoning simulation is more complicated. It demands features addressing metals profile, metal-catalyst interactions and metals age distribution [3]. The cyclic deactivation methods [3,4] focus on all those features that in many instances are strongly related to the FCCU operating conditions (i.e. metal content in the feed, regeneration temperature and regeneration mode, etc.). Moreover, there is a strong perception that each industrial unit will deactivate the same catalyst in a very different way. Boock et al. [5] focused on the effects that contaminant metals produce in full combustion and partial combustion operations while developing the Cyclic Propylene Steaming (CPS) method. In spite of the advances reported, one of the main questions still remaining concerns the reliability of properties exhibited by artificially deactivated-aged catalysts in representing the properties of the Ecats. These properties can vary in a large extent according to the varying operating conditions found in the units. Therefore, leaming more about deactivation and aging of catalysts directly from industrial units is wise. The analysis of the chemical and physical properties of Ecats can provide rough information related to the performance of the catalyst but hardly give details about the deactivation and aging processes. A more detailed picture can be drawn by collecting information of the properties of Ecats at different stages of deactivation by grading catalyst particles according to differences in density [6,7]. Palmer and Cornelius [6] determined the age of density graded fractions of resid Ecats in a pilot plant according to their Ni content assuming that Ni is deposited quantitatively on the catalyst. Calculations were based on the Ni content in the feed, the feed rate and the catalyst inventory. Compared to a pilot plant operation, the Ni content in the feed and the feed rate to the unit are hardly constant. Therefore, this impacts the accuracy of data that are used to estimate the age of an Ecat and its fractions from their Ni content. The aims of this work were to identify patterns of deactivation and aging in a REUSY catalyst used in industrial units and to determine how such patterns change with operating conditions. Moreover, we have tried to develop a method to acquire information from Ecats that avoids inaccuracy in estimating the age of Ecat fractions from the Ni content. Catalysts deactivation and aging were tracked by simply correlating the relative change of zeolite dealumination and zeolite surface area losses. This correlation has the advantage of comparing catalyst properties on a time independent scale. In order to compare the effects produced by laboratory deactivation protocols we have also studied samples deactivated by steaming, CPS and a modified cyclic method.

89 Understanding deactivation and aging in industrial units and their relationship to operating conditions will help in designing more effective laboratory deactivation protocols for a better assessment of the catalytic properties of fluid cracking catalysts in view of their industrial application.

2. EXPERIMENTAL 2.1 Catalysts Three Ecat samples were obtained from either full (Kellogg Orthoflow-F) or partial combustion (Kellogg Ultra-Orthoflow) units. Table 1 gives the main properties of those samples and the fresh catalysts samples. The full combustion unit operated comparatively at higher regenerator temperature (RegT). Table 1. Properties of the fresh and equilibrium catalysts and data of operating variables from the FCC units. Catalyst 1 Fresh

Catalyst 2 PC1

Surface Area (m2/g)

FC1

Fresh

PC2

.

Zeolite

176

119

123

167

106

Matrix

83

41

39

60

43

Ni (ppm)

n.a.

611

344

n.a.

212

V (ppm)

n.a.

2505

1497

n.a.

991

Fe (ppm)

n.a.

4600

5400

n.a.

3714

Na (%)

0.28

0.19

0.26

0.27

0.23

MAT second order rate constant, K* UCS (nm)

9.4

2.01

1.89

8.1

2.03

2.4580

2.4245

2.4270

2.4510

2.4268

Combustion

--

Partial

Full

Partial

Catalyst inventory (ton)

"-

180

260

180

Daily fresh catalyst makeup (%inventory)*

__

1.5

1.8

1.7

Ni content in feed (ppm)* Feed rate (BPD)*

__

0.23

0.21

0.054

""

36,100

34,900

38,000

Regenerator -686 708 temperature (~ *defined in section 2.5; ** refers to an average of 3-4 months

665

90 The catalysts studied (coded 1 and 2) contained a partially rare earths exchanged ultrastable Y zeolite (REUSY). The rare earths oxide (REO) content was 1.7 wt% and 1.0 wt%, respectively, for catalysts 1 and 2. The type of matrix (silica-sol plus kaolin clay) and the alumina content (33 wt%) were essentially the same in both catalysts. The zeolite to matrix ratio (Z/M) was 2.1 and 2.8, respectively, for fresh catalysts 1 and 2. Catalysts underwent equilibration to very similar surface area and catalytic activity (MAT second order rate constant, K). PC1 showed a lower unit cell size value. The vanadium content on the Ecats ranged from ca. 1200 ppmw to 2500 ppmw. The Ni and Na contents were comparatively lower.

Table 2. Properties of density graded fractions of Ecats (catalysts 1 and 2) obtained from partial combustion (PC) and full combustion (FC) units. Sample

wt%

Ni days

Metals (ppmw) Ni V

ZSA (m2g1)

K (MAT)* UCS (nm)

PC1 A B C D

19 29 28 24

130 85 60 31

950 625 437 229

3372 2875 2212 1612

53 120 140 152

1.03 1.89 2.28 2.53

2.4220 2.4238 2.4245 2.4275

FC1 A B C D E

20 24 21 18 17

143 80 59 50 38

657 368 269 230 175

2372 1595 1235 1060 880

42 128 149 154 161

0.84 1.99 2.51 2.38 2.76

2.4243 2.4257 2.4269 2.4277 2.4286

PC2 A B C D

34 27 28 11

173 97 82 33

313 175 148 59

1523 920 775 303

80 130 138 146

1.36 2.32 2.34 2.65

2.4240 2.4259 2.4282 2.4301

*see section 2.5

2.2 Density graded fractions The samples of Ecats were separated into four to five density graded fractions by means of a modified sink-float method [7] using mixtures of acetone/methyl-diiodide. The starting mixture was adjusted to 1 g/cm 3 density. Step-wise additions of acetone produced mixtures 0.95, 0.91, 0.86 and 0.81 g/cm 3. The floating portion of the catalysts was separated in each operation after centrifugation. The samples recovered from each operation were coded from A to E from the most dense to the

91 less dense. Table 2 shows the main properties of the fractions obtained. The age of fractions was calculated and reported in terms of the "Ni days" [6] that is calculated from the averages of the feed Ni content, the feed rate and the catalyst inventory. PC1 exhibited the highest Ni deposition rate (7.33 ppm/day), followed by FC2 (4.75 ppm/day). It was much lower for PC2 (1.81 ppm/day). The progressive age of fractions matched with the progressive decrease in ZSA and the increase in metals content. The pore volume varied from 0.17 g/cm 3 to 0.21 g/cm 3 for the most aged fractions. It was observed that the matrix surface area was essentially constant at 35-45 m2/g, the lowest values occurred for the most aged fractions.

2.3 Laboratory deactivated samples In order to compare the effect of lab deactivation protocols samples of fresh catalyst 1 were steamed (sample coded STM) under 100% steam at 788~ for 4 h or deactivated by cyclic deactivation (sample coded CD) according to a method reported elsewhere [8]. This sample contained 680 ppmw Ni and 2930 ppmw V. A sample deactivated by the Cyclic Propylene Steaming (CPS) method was kindly provided by Grace Davison (W.R. Grace Research Center, Columbia, MD) and contained 1030 ppmw Ni and 2160 ppmw V.

2.4 Characterization techniques The chemical composition of catalysts and fractions was determined by atomic absorption and inductive coupled plasma analysis. The V and Ni content in fractions were used to calculate a vanadium mobility index according to a reported method [7]. Solids were characterized by X-Ray diffraction (Siemens D-500) using ASTMD3942- 85 to calculate the unit cell size (UCS) of the zeolite. Total surface area and zeolite surface area (ZSA) were calculated from nitrogen adsorption (Micromeritics ASAP-2405) data according to ASTM-D3663 and ASTM-4222 methods involving the t-plot method and the Harkins-Jura equation [9]. This approach is currently used for determining the ZSA in fluid cracking catalysts. ZSA was taken as criteria of crystallinity of the zeolite whereas the UCS was related to the zeolite framework aluminium according to the following equation [10a]: A! atoms per UC= 115.2 (UCS-24.191 )

(i)

where UCS = unit cell size (A). Samples CPS, CD and fraction A of Ecat PC1 were studied by Electron Paramagnetic Resonance (EPR). A sample of USY zeolite ion-exchanged (V acetate solution) with ca. 3000 ppmw V was used as a reference. EPR Spectra were recorded at room temperature in a Bruker X-Band equipment at a frequency of 106 Hz. Samples were previously calcined under dry air at 923 K for 2 h.

2. 5 Microactivity test (MAT) The catalytic activity and product selectivity of Ecats and their density graded fractions were studied in a fixed bed AUTOMAT unit from Xytel using 4 g of calcined (550~ for 2 h) catalyst at C/O=5, 520~ WHSV=16 h1 and 45 sec for feed injection. The feed used was a mixture of atmospheric and vacuum gas oils with

92 properties of 23.35~ 0.51%CCR, 11.9 KUOP, 81~ aniline point, 2.0% S and 896ppmw basic nitrogen. The catalytic activity was expressed in terms of a second order rate constant [K = conversion/(100-conversion) ]. The reaction products were analyzed by Gas Chromatography using two Hewlett Packard (HP) 7960's configured respectively for Refining Gas Analysis and Simulated Distillation according to HP applications. 3. RESULTS AND DISCUSSION 3.1. General features Figure 1 stresses the importance of the relative contribution of the most aged fractions to the calculated average activity of the Ecats' fractions. The cumulative MAT activity in Figure 1 is defined as the percentage of the overall activity of the Ecat that the mixture of catalyst fractions would show. More than 50% of the catalytic activity is provided by fractions possessing more that the half-age of the Ecat. The shape of curves in Figure 1 could be attributed to deviations with respect to average values of by the Ni deposition rate. Deviations are produced by either the feed rate or the feed nickel content variations during the period of time considered. Comparatively PC1 showed a less variable Ni deposition rate while FC1 and PC2 showed somewhat higher nickel deposition rates with respect to average in the front end period thus overestimating the Ni age.

,A

o~" 1 0 0

,t,,

m

._z, ._ 80

~ .__

40

-5 E

20

0

I

0

0

50

100

150

200

Ni age (days) Figure 1. Cumulative MAT activity of fractions from catalysts ( . ) PC1, (B)FC1 and (A) PC2 as a function of their nickel age. The high Ni deposition rate average on PC1 includes a high vanadium deposition rate as observed in Figure 2. Deviations of plots from linearity are essentially related to the V migration phenomena. Figure 2 allowed us to classify the three units. PC1 is a medium to high metal -low temperature operation, whereas FC1 is a medium

93 metal -high temperature operation and PC2 is a low metals -low temperature operation. The MAT activity expressed as the second order rate constant K of the Ecats fractions shows a strong dependence on the zeolite content. (Figure 3a). The correlation using the values of total surface area (Figure 3b) was comparatively poorer thus indicating a greater influence of the matrix component especially for the younger fractions. On the other hand, the laboratory deactivation methods closely reproduced this property as shown by the data from samples STM, CPS and CD. 4000 E 3000 c 2000 C O O

>

1000

0

50

100

150

200

Ni age (days) Figure 2. V content on fractions with respect to their Ni age, ( . ) PC1 (B)FC1 (A) PC2 3.2 Deactivation kinetics. For practical purposes the deactivation rate constant of a catalyst in a FCCU can be estimated from the inverse of the fractional fresh catalyst makeup (T, in days), the MAT activity (A) of the Ecat and the MAT activity (Ao) of the fresh catalyst [11,12] according to the equation"

A = Ao / (kd T +1)

(2)

where A and Ao are expressed as the second order rate constant, K. These numbers are reported in Table 1. The calculated deactivation rate constant kd was 0.055, 0.073 and 0.049 days 1 for PC1, FC1 and PC2 respectively. Therefore, deactivation was faster for FC1 compared to PC1, while PC2 deactivates at a very slow rate. The deactivation trend of fractions can also be observed in Figure 4. The ZSAK number is based on the concept by Mott [13] who calculated a SAK number from the ratio of the total surface area divided by the second order rate constant, K. This number is inversely proportional to the specific activity of the catalyst and thus can

94 be used to differentiate catalysts and deactivation related to aging from discrete events occurring in the unit (metal poisoning, hydrothermal shocks). In this work, the ZSAK number only considered the zeolite surface area as this parameter seems to be closer related to the catalytic activity compared to the total surface area.. Clearly the ZSAK number in Figure 4 tends to a constant with only a slight decrease for the oldest particles for PC1 and FCI. The apparent increase in activity in those fractions can be explained by an increase in the coke yield (see section 3.7) resulting from their higher metal content.

3.5

3.02.5 K

2.0 1.5-

A

1.0

II

0.5 ]

2OO

100

30

ZSA(m2/g) 3.5 3.0

II

2.5

K

2.0 1.5

A

1.0

II

0.5 I

0

100

i

200

300

TSA(m2/g) Figure 3. MAT second order rate constant (K) with respect to the a) zeolite surface area (ZSA) and b) total surface area (TSA) in fractions from ( . ) PC1 (II)FC1 (A) PC2 and samples (+) STM ([]) CPS and (X) CD.

95 70 60

J

5O ~" < 40 o~ N 30

i

20

i

10 0

0

r

i

i

50

100

150

200

Ni age (days) Figure 4. ZSAK number for fractions of catalyst (~) PC1 (m)FC1 and (A) PC2.

An approach for calculating an apparent deactivation constant according to the exponential equation relating the MAT activity expressed as K (A') and the Ni age expressed in days" r= A' exp (-k t)

(3)

1.2

-

~

0.8

9

0.4 0.0

-

-0.4 0

I

I

I

50

100

150

200

Ni age (days) Figure 5. Ln function of the second order MAT constant K with respect to the Ni age of fractions expressed in days, (~) PC1 (n)FC1 and (A) PC2.

Plots shown in Figure 5 as a function of the Ni age had regression coefficients of 0.92-0.98 with slopes being the deactivation rate constants. The values of 0.009, 0.012 and 0.005 days1 for PC1, FC1 and PC2 respectively were calculated. This confirmed the trends previously observed by using the fresh catalyst makeup calculation. However, differences in the absolute values of the constants obtained by the two methods were observed. This could be mostly ascribed to the fact that the

96 fresh catalyst MAT activity used in the makeup rate approach, is comparatively very high thus producing higher values of the apparent deactivation rate constant.

3.3 Aging kinetics It is well known that catalyst deactivation by aging in FCC units occurs mainly by dealumination and crystallinity loss of the zeolite component. The effect of aging is shown by lower surface areas and UC values. As a general rule high RegT, high steam or oxygen partial pressures and high vanadium contents increase dealumination and ZSA losses. Using a method already published [6] we studied the UCS and the ZSA of the Ecat fractions as a function of their Ni age in days. The linear regression of plots in Figures 6 a) and 6 b) allowed calculation of apparent rate constants (KZSA and KDA) related to ZSA losses and zeolite dealumination. The numbers obtained (Table 3) from the linear regression analysis indicate that the zeolite structure collapses at a similar rate in FC1 and PC1. The zeolite appears to be less destroyed in PC2. This can be ascribed to a combined effect produced by the lower RegT and the lower metal deposition rate in PC2 compared to PC1 and FCI. Table 3 indicates that the UC shrinkage is faster in PC1 and much slower in FC1 and PC2. In this case it seems that the zeolite dealumination rate is more sensitive to a high rate of metals deposition than to a high RegT. The values for KZSAand KDA for FC1 indicate that the relative high RegT in the full combustion unit accelerated the ZSA loss but had no major impact upon zeolite dealumination. The high dealumination rate and ZSA loss rate constants in PC1 could be ascribed to its relative high vanadium content. It is important to outline that KZSA is in close agreement with the apparent deactivation rate constant (Figure 5) leading us to conclude that the ZSA is the main parameter related to the catalytic activity. Using the same calculation basis, KDA for the REUSY zeolite studied in this work was one order of magnitude lower than the values calculated for CREY or REY zeolites reported elsewhere [6]. However, the vanadium contents in Ecats were very high (ca..1 wt%) in these cases. _

O E~ <5

O3 N r .J

o

",,~,

4

v

c_J

Or 0

100 Ni age (days)

200

0

100 Ni age (days)

200

Figure 6. Rate of a) ZSA losses and b) Aluminium atoms per unit cell losses for catalysts (~) PC1, (m) FC1 and (A) PC2.

97 Table 3. Regression constants calculated from the plots in Figure 6 from the equations ZSA = AZSA exp (- K ZSA t) and N AI atoms/UC = ADA exp (-K DA t) where t is the Ni age (days). Catalyst

ZSA loss AzsA K

PC1 FC1 PC2 CREY* REY*

246 290 185 -. . -. .

ZSA

0.011 0.013 0.005 . . . .

(days 1)

R2 0.88 0.95 0.89

. .

Dealumination ADA K DA (days "1)

R2

12.7 12.2 15.2 62.3 31.4

0.98 0.89 0.96 n.r. n.r.

0.012 0.006 0.007 0.062 0.068

* Reference [6] n.r.= not reported

The numbers in table 3 show that the ZSA at zero Ni days (AZSA) values are overestimated while the AI/UC (ADA) values are underestimated compared to those corresponding for the fresh catalysts. This can be explained by the fact that the data used for the regression analysis are restricted to the branch of the deactivationaging pattern (see section 3.4) where the rate of ZSA loss is accelerated and the zeolite dealumination rate is hindered. In order to get more information it would be necessary to study catalyst fractions corresponding to earlier stages of deactivation. Nevertheless, our results lead us to conclude that the UC shrinkage in REUSY zeolite is the most rapid process (compared to ZSA loss) occurring during the first cycles the catalyst spends in the unit. Therefore, we propose that ADA represents the AI atoms per UC for the initial stabilization of the zeolite UCS during the very first operating cycles the catalyst spend in the unit. It should be expected that ADA is affected by the operating variables. We ascribe the lower ADA numbers obtained with PC1 and FC1 to the relative high RegT compared to PC2.

3.4 Deactivation and aging patterns Zeolite stabilization under thermal and hydrothermal conditions like those prevailing in a regenerator of a FCCU occurs by a complex mechanism involving crystallinity losses and zeolite dealumination. Both share an initial step that involves aluminium expulsion from the zeolite lattice as illustrated in Scheme 1. This step produces structure vacancies and extraframework alumina (EFAL). Zeolite dealumination occurs in a second step by a mechanism involving silicon insertion to the vacancies left by the aluminium expelled from the framework [10b,14,15]. The silicon, which is transported as Si(OH)4 at elevated temperatures under water vapor atmosphere, is probably produced by zeolite decomposition [10b,15]. Silicon insertion and crystallinity losses are consistent with UC shrinkage and ZSA decreases, respectively. From Scheme 1 it becomes evident that the ZSA loss and the UCS shrinkage are interrelated processes. It should be expected that the rate constants K1 and K2 depend on a number of parameters, whereas K3 would depend to a large extent on the Si(OH)4 partial pressure in the regenerator. Therefore the ratio of K1 or K2 to K3

98 would vary mainly because of K1 or K2 variations and would explain the relative stability of the zeolite. This means also that the same catalyst will probably respond in a different way to the operating conditions in the unit. It follows that ZSA losses (or crystallinity losses) would be accelerated over zeolite dealumination (UCS shrinkage) under full combustion modes, higher temperatures and high V contents Since the UC shrinkage and ZSA loss are close related processes, it was interesting to study them togehther according to the operating conditions and the catalyst properties. Plots in Figure 7 indicate that the REUSY catalysts initially underwent a very fast dealumination process with respect to the ZSA loss. The pattern showed a marked ZSA loss below 10 AI/UC. From that point on, a striking difference among catalysts was observed. Catalyst FC1 and PC2 lost their ZSA at higher aluminium contents.

Scheme 1. Reaction pathways for zeolite dealumination- crystallinity loss. Si

SiO2

6I

AI-(SiO)4 ' ~

EFAL +

Si-O-H

H I

0

I Si

(collapse)

H-O-Si

~

Si 0 Si-O-Si-O-Si 0 Si

Data reported in the literature are also plotted in Figure 7. A REHY catalyst [16] steamed (815 ~ at varying time and a CREY resid catalyst [6] tested in pilot plant displayed quite different patterns with a fast ZSA losses with respect to dealumination. Comparatively, the ultrastable zeolite retained higher surface area values at lower UCS. In addition, Figure 7 shows that samples obtained from the laboratory deactivation protocols produced a pattem of lower ZSA and higher aluminium contents compared to equivalent Ecat fractions. It can be concluded that in order to reproduce more closely, the properties of Ecat a less severe steaming of the fresh REUSY catalyst should be used. The aging pattern is quite different between REUSY and high alumina zeolites like CREY or REHY. Initially, there is AI extraction from the lattice with no great destruction of the zeolite framework. Then the collapse of the zeolite structure becomes extremely sensitive to AI losses. It is well known that in order to produce USY zeolites conditions that favor dealumination upon structure collapse are needed. It appears that these conditions lead to dealumination in specific sites throughout the zeolite framework thus making

99 stabilization possible. Similarly, it could be assumed that the number of vacancies that are created in our REUSY catalysts under the FCCU operating conditions seem to be relatively few leading to further dealumination with a very low ZSA losses. However under c.a. 10 AI/UC the zeolite structure collapses much faster upon dealumination. It is concluded that the aluminium sites involved in the zeolite structure under c.a. 10 AI/UC seem to be essential for avoiding structure collapse. 200 160 -

"

~ ::

E~

"~ 1 2 0 E

<

o~

///

80-

/ :l

N

400

9a t"

./,"

.;;:.;'.:-'"

~:~'"

.-"

I

1

10 AI atoms/UC

100

Figure 7. ZSA-AI atoms/UC dependence. Fractions from catalysts ( - ) PC1, (m) FC1, (A) PC2, (+) STM, (1-1) CPS, (X) CD, (O) fresh catalyst 1 and (.) fresh catalyst 2. Dotted lines are for: a) CREY catalyst from reference [6]; b) REHY catalyst from reference [16].

The results obtained would lead to rule out the idea that dealumination of REHY catalysts in industrial units could produce an equivalent stabilization toi USY. In order to get a better knowledge on the relative importance of the two phenomena involved the ratio of the incremental differences of ZSA and AI/UC were plotted as a logarithmic function of the number of AI/UC remaining for each fraction (Figure 8). As the zeolite becomes more stable upon dealumination the extent of the relative change of ZSA would depend on the severity of operating conditions that finally produce the zeolite collapse. Thus, the higher the severity of the operating conditions in the FCCU zeolite will collapse at higher AI/UC values and the higher the slope values in Figure 8. The calculated values were 0.62, 0.81 and 0.44 UC/AI for PC1, FC1 and PC2, respectively, while the regression coefficients were 0.880.90.

3.5 The ZSA-AI/UC relationship to catalyst deactivation. The value of the slope of the linear regression functions (Figure 8) is related to the trends and patterns in Figure 7, and possibly this value could correlate with catalyst deactivation. Table 4 summarizes the values of the deactivation rate constants calculated for PC1, FC1 and PC2 according to the method discussed in section 3.2. Since the ZSA-AI/UC ratio functions are calculated on a time independent scale it was necessary to use a reference in order to compare them.

lo0 relative values referred to PC1 were calculated. The ZSA-AI/UC function predicted the trends and orders of the relative deactivation rates in good agreement with the "Ni days" and "makeup" based methods (see section 3.2). Relative numbers are more consistent with deactivation calculated from the "makeup" method. In the industrial units considered in this work, fresh catalyst makeup data were collected in a daily basis whereas the Ni content in the feed was analyzed on a weekly basis

~

6

<

4-

m,

Z

~

2-

N

0-

v

._.1

-2

-

I

0

I

5 10 AI atoms / UC

15

Figure 8. ZSA-AI atoms/UC differential changes in fractions obtained from catalysts ( . ) PC1, (m) FC1 and (A) PC2.

Table 4. Deactivation rate constants and relative rates obtained from different calculation basis. Ni days based [6] r = A' exp (-k t)

Makeup based [11,12] A = Ao / (kd T +1)

ZSA-AI/UCS based

k (days1)

relative

k~(days")

relative

relative

PC1

0.0072

1.0

0.055

1.0

1.0

FC1

0.0117

1.63

0.063

1.32

1.31

FC2

0.0046

0.64

0.049

0.89

0.72

Catalyst

The results obtained suggest that deactivation of the REUSY catalysts could be assessed by just looking at the ZSA-AI/UC or UCS ratio pattern in Ecats. This is possible because of the high dependence of catalytic activity upon its ZSA (regardless of the AI content in the zeolite) and because the reduction of the AI/UC occurs concurrently with the ZSA losses.

101

Coke yield (wt %) 8.0 . . . . . .

LPG olefinicity 0.9

7.0

0.8

6.0

X

-

5.0

-

4.0

-

0.7

|+m

D

41,

x"

"G

0.6 0.5

Gasoline yield (wt %) 60-

C4's olefinicity 0.9 0.8

~ Io

50-

9

41,

0.7 0.6

40-

m~+ gm

X

0.5

4)

30-

0.4

LPG yield (wt %) 20-

H2/CH4 (wt/wt) ratio 0.4

15

1,1

0.3

[]

XE] +

10

0.2 0.1 I

0

1

2

u

0.0

I

3

K

I

0

+ nl I

1

2 K

Figure 9. Product yields and selectivity obtained in MAT with fractions of catalysts (4,) PC1, (11) FC1 and samples (+) STM, (El) CPS, and (X) CD. MAT conversion is expressed as a second order rate constant K.

3.6 The zeolite exchange loss capacity. Contrary to what was shown for CREY and REY catalysts in resid pilot plant operations [6] the combined ZSA rate loss and dealumination constants failed to agree with the deactivation rate constant in REUSY. This lack of agreement could be explained by the fact that the CREY and REY catalyst used to create the correlation, contained larger amounts of AI and thus the effect of losing framework aluminium had a great impact on the acidity and on cracking activity, it also could be expected that other parameters in addition to the unit cell (i.e. mesoporosity and

102

extraframework alumina) would modify or even explain the catalytic activity as it was shown for dealuminated HY zeolites [17]. Nevertheless, our results indicated that the activity decay of the REUSY catalysts studied correlate very well with the ZSA.

3.7 Product selectivity in MAT Figure 9 shows varying aspects of product distribution obtained in MAT for fractions of PC1 and FCI. They were compared to those obtained with samples STM, CPS and CM samples. No appreciable differences in LPG, gasoline or coke yields were observed between fractions of PC1 and FC1 fractions. Some differences were observed in C4s' and LPG olefins suggesting that the lower UCS in PC1 may contribute to these yields. Higher H2/CH4 ratios were produced with PC1 fractions possibly because of their high metal content. The reason for the high coke yield in the most aged fractions of PC1 and FC1 is less clear. High coke yields can be associated with relatively enhanced non selective cracking on acid sites created by zeolite collapse [17]. Compared to Ecat fractions possessing the same MAT activity, laboratory deactivated samples showed a remarkable agreement in gasoline yield but not in LPG yields. STM produced less propylene thus lowering LPG olefinicity. CPS produced a very high coke yield and H2/CH4 ratio. CD showed lower coke yield and lower H2/CH4 ratio but still high compared to the corresponding Ecat fractions. These features reflect the fact that the metal content in these samples are higher than in the Ecat fractions of similar activity. Also, are metals much more dispersed throughout the catalyst surface because of the method they were impregnated [8].

tJ

1.8

LU

1.2

p, ~ 0.6 -z

o.0

0.8

0.9

1

Density of fraction (g/cm3) Figure 10. Plot for VMI calculation using fractions obtained from catalysts ($) PC1, (m) FC1 and (A) PC2.; VMI is proportional to the slope of the straight line.

3.8 Vanadium poisoning The full combustion unit considered in this work operates at high RegT and excess 02. These conditions would produce vanadium with a higher oxidation state leading to the formation of mobile species, harmful to zeolite stability [18].

103

An attempt to quantify vanadium migration (expressed as a vanadium mobility index or VMI) in FCCUs has been recently reported [7]. Figure 10 shows a plot used to calculate a VMI for the fractions of the three catalysts studied in this work as defined in reference [7] using the V and Ni contents in the Ecats and fractions. The VMi is proportional to the slope of the regression lines and was higher for PC1 (20.1) than FC1 (6.6) and FC2 (1.8) despite the fact that the former operates in a partial combustion mode. Boock et al. [7] studied a large number of FCCUs and found trends for higher values of VM! associated with full combustion regeneration. This lack of clarity indicates that vanadium mobility is affected by other parameters than combustion conditions. In the present study it could be suggested that the higher vanadium deposition rate leads to a high vanadium content in the catalyst inventory thus increasing vanadium mobility in PC1 compared to FCI. The low V content and low RegT in the FCCU of PC2 agree with the low V mobility found in this unit. FC1 showed lower ZSA than PC1 at the same vanadium content as can be observed in Figure 1la. In spite of its higher V content, PC1 deactivates at a slower rate than FCI. The high VMI in PC1 would indicate that vanadium moves at a relatively higher rate from the older to the younger catalyst particles. The results obtained would lead one to conclude that since this "active" vanadium does not accumulate in the older particles, it will be kept away from the zeolite and thus the ZSA will be retained in the older particles. Figure 11b shows that dealumination of PC1 and PC2 follows the same decay trend with the V content. However the fact that PC1 showed a high AI/UC compared to PC2 for the same V content indicates that the time effect is also very important, PC2 being 2 to 3 times older than PC1. On the other hand, Figure 1lb shows that the V activity for dealumination is lower for FCI. As discussed in Scheme 1, structural collapse and zeolite dealumination involve concurrent reactions. The V action at high RegT would produce such a high number of structure vacancies as a consequence of AI ejection that stabilization of the zeolite framework by silicon atoms insertion becomes more difficult. Therefore, the ZSA losses increases at the expenses of zeolite dealumination. 6 5 < 4 oo3

41111

9 A 11

9

9

"~2 IE

N

,,

~,,,,,,,,,,~

c2

..J

1 0 -1

c_.1

0

1000 2000 3000 V content (ppm)

4000

0

I

I

I

1000

2000

3000

4000

V content (ppm)

Figure 11. a) ZSA and b) AI/UC dependence on the vanadium content in fractions of catalysts (~) PC1, (m) FC1, and (A) PC2

104 Noteworthy, the intercept of the AI atoms/UC is not the same at zero V in Figure 1 lb as it is for zero Ni (Figure 6b). This is explained by the fact that the V content this correlation is based on is not necessarily related to the catalyst age as it is in the Ni case. High V migration tend to normalize V contents on fractions. Therefore high VMI increases the slope value, which is in fact observed for PC2. V species involved in migration and in zeolite structural collapse were studied recently [18]. V exchanges as V ~vinto the cationic positions of the zeolite, preceding destruction of the zeolite framework. EPR can detect such species [18]. EPR spectra are presented in Figure 12. In a pure V-exchanged USY sample the EPR spectrum shows the typical pattern of isolated tetravalent vanadium. The clear hyperfine splitting peaks indicate that the unpaired electron of V 'v is interacting with only one nucleus, pointing out the high dispersion of V species. In other words, in this stage V 'v species may be present as (V=O) 2§ vanadyl ions compensating the negative charge of the zeolite framework.

Figure 12. EPR spectra of V ~'in catalysts and V exchanged USY.

The EPR spectra of laboratory deactivated samples and fraction A of catalyst PC1 are more difficult to interpret due to the presence of a broad signal overlapping the V 'v species' signal. However, some hyperfine structure from V 'v species can be still observed mainly in lab-deactivated samples. The broad signal may be ascribed to

105

amorphous V204 or V205. In order to confirm the nature of the true species in such catalysts, simulations should be performed. Nevertheless, from the results obtained in this work we conclude that V cationic species are produced in a large extent in laboratory deactivated samples because of the fast rate of deposition [8]. The absence of an EPR signal on the Ecat might be simply associated with a very low concentration of such species in the catalyst. 4. CONCLUSIONS This study provides good insight into the effects of operating variables in deactivaton and aging of REUSY containing catalysts in industrial units. The REUSY catalyst studied in this work showed deactivation-aging patterns made up of an initial stage where a fast dealumination takes place and a second stage occurring below 10 AI/UC with great destruction of the ZSA. The most aged catalyst particles are major contributors to the catalytic activity of Ecats. The relative rates of ZSA losses and dealumination depend on the operating conditions in the FCCU and on the metal deposition rates. The results obtained also suggest that deactivation of REUSY catalysts can be assessed from the ZSA-AI/UC or UCS ratio pattern in Ecats. This is possible because of the high dependence of the catalytic activity on the ZSA (regardless of the AI content in the zeolite) and because the reduction of the AI/UC occurs concurrently with the ZSA losses. Even though separation of catalyst particles is needed, this approach would facilitate assessment since it is independent of time and average data from the FCCUs. MAT activity-selectivity studies did not show any significant dependence on the particular deactivation-aging process followed by the two catalysts. Metal deposition and steaming severity seemed to be key issues for laboratory deactivation protocols to reproduce typical Ecat properties. The effects of V poisoning could be related to the V content in the catalyst inventory and to the operating conditions in the units. V effects are strongly time dependent. The study of the ZSA-AI/UC pattern produced in a particular unit for a REUSY catalyst could be used as reference to reproduce Ecat properties in fresh fluid cracking catalysts. REFERENCES

1. McLean, J.B. and Moorhead, E.L., Hydroc. Proc., February (1991) 41. 2. Keyworth, D.A., Turner, J. And Reid, T.A., Oil & Gas J., 14 March (1988) 65 3. O'Connor P., Brevoord, E., Powels, A.C. and Wijingaards, H.N., Ch. 10 in Deactivation and Testing of Hydrocarbon Processing Catalysts, P.O'Connor, T. Takatsuka and G. L. Woolery, Eds., ACS Symp. Ser. 634, Am. Chem Soc., Chicago (1996) 147. 4. R.N. Cimbalo, R.L. Foster and S.J. Wachtel, Oil & Gas J., 70 No. 20 (1972) 112. 5. Boock, L., Pretti, T., Rudesill, J., Ch. 12 in Deactivation and Testing of Hydrocarbon Processing Catalysts, O'Connor, P, Takatsuka, T. & Woolery, G., Eds. Acs Symp. Ser. 634, Washington (1996) 171.

106

6. Palmer, J.L. and Cornelius, E.B., Appl. Catal., 35 (1987) 217. 7. Boock, L., Deady, J., Foon Lim T. and Yaluris, G., Catalyst Deactivation 1997, Bartholomew, C. & Fuentes G. Eds., Elsevier Sc. (1997) 367. 8. Hernandez, F., Garcfa de Le6n, R., Mogica, E., Moreno, J.C., Garcfafigueroa E., and Gonz~lez, R., Catalyst Deactivation 1997, Bartholomew, C. & Fuentes G. Eds., Elsevier Sc. (1997) 455. 9. Leofanti, G., Padovan, M., Tozzola, G., Venturelli, B., Catal. Today 41(1998) 207. 10. Breck, D.W., "Zeolite Molecular Sieves, Structural Chemistry and Use", John Wiley; N.Y. (1974) a) Ch. 2; b) Ch.6. 11. Lee, W., ind. Eng., Chem. Proc. Des. Develop., V. 9, No.l, January (1970) 155. 12. Leunberger, E.L., Oil & Gas J., Technology, July 15 (1985) 125. 13. Mott, R.W., Paper AM-91-43, NPRA Anual Meeting, March 17-19, 1991. 14.Kubicek, N., Vaudry, F., Chiche, B.H., Hudec, P., Di renzo, F., Schulz, P. and Fajula, F., Appl. Catal. A: general 175 (1998) 159. 15. Beyerlein, R.A., Choi-Feng, C., Hall, J.B., Huggins, B.J. and Ray, G. J., Topics in Catalysis, J.J. Fripiat and J.A. Dumesic, Eds., J.C. Baltzer A.G., Sc. Pub., 4 (1997) 27. 16. de la Puente, G., Sedran, U.A., Micr. Mat. 12 (1997) 251. 17. Corma, A., Forn#s V., Martinez, A. and Orchilles, A.V., Ch. 35 in "Perspectives in Molecular Sieves Science", ACS Symp. Ser. 368, Am. Chem. Soc., Chicago (1988) 541 18.Trujillo, C.A., Navarro, U., Knops-Gerrits, P., Oviedo, L.A., and Jacobs, P., J. Catal., 168 (1997) 1.

Studies in Surface Science and Catalysis 134 M.L. Occelli and P. O'Conner (Editors) 9 2001 Elsevier Science B.V. All rights reserved

107

THE IMPROVEMENT OF CATALYTIC CRACKING PROCESS THROUGH THE UTILIZATION OF NEW CATALYTIC MATERIALS M.I.Levinbuka, V.B.Melnikov a, Samir Numan a, M.L.Pavlov b and V.A.Patrikeev c "Gubkin Moscow Oil and Gas University, 65 Leninsky prosp., Moscow 117917, The Russian Federation. blshimbai Catalyst Factory, Ishimbai 453210, The Russian Federation. CSalavat Catalyst Factory, Salavat 453206, The Russian Federation. 1. INTRODUCTION For the 50 years that the catalytic cracking exists, the development of both its equipment and its catalysts has been caused mainly by the continuous increase in the proportion of heavy fractions in feed and changes in requirements for the quality and yield of products to be obtained (LPG, LCO, gasoline). But the cracking chemism remained unchanged even after zeolites had come to be used in FCC catalysts. Since the object of the design of the equipment of any catalytic process is to attain the maximum potential yield of the target product possible with the use of a particular catalyst, in recent years nearly all the tasks facing cracking have been solved by improving the production equipment of cracking units. Now, there exist two major way to develop the cracking process: first, to modify the reactor unit in such a way that would allowed the "Short Contact Time" technology to be applied to increase the gasoline yield when processing heavier feed, and, second, to either adapt the reactor vessel to operation at higher temperatures or to build new facilities producing ZSM-5 zeolites, which are the basis for the manufacturing of catalytic additives to the existing commercial catalyst with ultrastable Y zeolite so as to raise the yield of light olefins in LPG (with a decrease in the gasoline fraction yield), which are used as feed for growing capacities of alkylation, MTBE and TAME units. Since cracking gasoline is the major source of olefins in commercial gasolines, the possibility of adjusting the proportion of aromatics to olefins in the chemical composition of the gasoline fraction is very important, because a growths in the reaction temperature or an addition of ZSM-5 zeolite to the main catalyst increases the gasoline olefinicity; this way, though, is restricted by requirements imposed on reformulated gasolines. Both the above-mentioned directions of cracking development as well as the trend to modify the catalyst regeneration unit (the two-step regeneration) are connected chiefly with the fact that the possibilities of enhancing properties of FCC catalysts containing ultrastable Y zeolite to adjust to new tasks facing cracking, are limited. Besides, both manufacturers and consumers of catalysts have to make great investments to implement such ideas in industry. Therefore, the introduction of two new components, low-alumina LAY-zeolite and constant or variable valence metal oxides can provide a solution to many of problems facing cracking

108 process, without any considerable modification of the equipment of existing units required. LAY zeolites as against ultrastable and conventional NaHY zeolites allow catalytic properties to be improved in terms of both the residual Na20 content and the molar SIO2/A1203 ratio in the zeolite framework [ 1]. Variable and constant valence metal oxides constitutes an "oxygen trap" transferring oxygen from the regenerator into the reactor to partially oxidize hydrocarbons in feed (reagent catalyst). This changes the cracking chemism and greatly affect the range of products obtained [2]. 2. EXPERIMENTAL It is to be noted, that the independent existence of two catalytic cracking technologies (TCC and FCC) has always promoted the accelerated introduction of new catalytic components previously tested in TCC catalysts into FCC catalysts (for example, X and Y zeolites, CO complete combustion promoters, ZSM-5 zeolites etc.). Low capacities of TCC units and the flexible technology of TCC catalyst manufacturing makes it easy to include various additions in them. Therefore, LAY zeolites and constant and variable valence metals were first introduced into TCC catalysts and subjected to pilot and industrial tests, and only then were they used in laboratory and pilot tests of FCC catalysts. The preparation of TCC and FCC catalysts containing LAY zeolites and the conditions of their tests in vacuum gas oil cracking on laboratory MAT units (a density of 0.915 g/cm 3, a Conradson carbon of 0.52 wt. %, a sulfur content of 1.52 wt. %, the IBP of 321~ the EBP of 519~ are indicated in [1]. The condition of FCC catalyst testing on a pilot unit are specified in [2]. Oxides of constant and variable valence metals of groups IV to VII were introduced into TCC and FCC catalyst following a technique [3]. Physical and chemical methods of investigation of LAY zeolites and TCC and FCC catalysts are given in [ 1]. Samples of commercial DA-250 (20 wt. % of CREY zeolite) and Emcat-Extra (12 wt. % of RE-USY zeolite) were used for comparison with samples of TCC and FCC catalysts under investigation. 3. RESULTS AND DISCUSSION Major results of the introduction of RELAY zeolites (10 wt. %) into FCC catalyst as compared with CREY zeolites (DA-250) are represented in Tables 1 and 2. It follows from Table 1 that when FCC catalysts are tested in conditional MAT mode at a constant contact time, that target products (gasoline and coke) of vacuum gas oil cracking forms more selectively over FCC catalysts containing only 10 wt. % of RELAY zeolite than over DA-250 (20 wt. % of CREY zeolite). These results also shows that RELAY zeolite combines both the merit of REHY zeolite (a high activity) and that of RE-USY zeolites (a low coke yield). Over 12000 metric tons of Z-10 TCC catalyst based on LAY zeolite was manufactured in Russia from 1994 to 1998. Tests of a TCC catalyst containing 10 wt. % of RELAY zeolite as compared with Emcat-Extra containing RE-USY zeolite under laboratory and industrial conditions at a constant space velocity (which corresponds to actual operational conditions of the catalysts in cracking units) discovered that the gasoline fraction yield has a maximum at lower catalyst to feed ratios over the samples under investigation as against commercial TCC catalysts [1]. The same results are also observed in the case of FCC catalysts if tests are carried out following the MAT method at a constant space velocity (Table 2). It follows from Table 2 that a lower catalyst to feed mass ratio and, consequently, a longer contact time is required to reach a similar feed conversion at a constant space velocity in the case of a LAY-

109 zeolite-containing FCC catalyst as against DA-250 (not to speak of USY-zeolite-based FCC catalysts). Table 1 Tests of FCC catalyst samples in vacuum gas oil cracking through MAT method at a constant contact time of 75 s, and a reaction temperature of 500~ Product Catalyst to feed mass ratio = 4.4 Catalyst to feed mass ratio = 7.0 yields, wt. % DA-250 Sample with DA-250 Sample with 10 wt. % of RELAY 10 wt. % of RELAY Gasoline 40.6 43.8 43.0 48.0 Gas 16.1 16.2 20.5 19.0 LCO 27.1 24.3 19.1 18.7 HCO 6.7 8.9 4.5 4.5 Coke 9.5 6.8 12.9 9.8 Conversion 66.2 66.8 76.4 76.8 The results obtained testify to the necessity of carrying out special researches so as to establish whether it is possible not to adapt cracking units to a process with a shorter contact time if FCC catalysts with RELAY zeolites are used. Table 2 Tests of FCC catalyst samples in vacuum gas oil cracking through MAT method at a constant space velocity of 16 h l , and a reaction temperature of 500~ Product Catalyst to feed mass ratio = 9.0 Catalyst to feed mass ratio = 6.0 yields, wt. % Sample with 10 wt. % of RELAY DA-250 Contact time 37.5 s Contact time 25.0 s Gasoline 49.0 43.2 Gas 20.0 24.1 LCO 19.0 20.0 HCO 7.0 7.0 Coke 5.0 5.7 Conversion 74.0 73.0 Metal oxides were introduced in cracking catalysts by a special technique ensuring that the oxides would not chemically react with the framework and that oxygen would easily be bound and then released in the catalytic reaction-regeneration cycle. One of the major purposes of introducing metal oxides ("reagent catalysts") into cracking catalysts for partial oxidation of hydrocarbons in feed is to acquire the possibility of controlling the chemical composition of gasoline, LPG and LCO without changing the reaction temperature and other process conditions (which eliminates the necessity to modify the equipment of existing units). The potential of reagent catalysts for changing the chemical composition of the gasoline fraction by changing the ratio of zeolite to the metal oxide content in catalyst is demonstrated by Table 3, where the yields of products of vacuum gas oil cracking over TCC catalysts containing oxides of IV group metals are indicated. The content of oxygen-containing hydrocarbons in gasolines did not exceed 0.5 wt. %. It follows from Table 3 that the olefin content in gasoline decreases, and the aromatics content increases with the growths of the zeolite to metal ratio in the reagent catalyst. The same dependence is

110 characteristic of olefin content in LPG, too. However, the product yields over reagent catalysts have a complicated dependence on the contents of pure metal oxide and zeolite in samples, rather than on the zeolite to metal oxide ratio. Table 3 also shows that the use of LAY zeolite in cracking catalysts results in an increase in the aromatics content and a decrease in the olefin content in gasoline. The results obtained were also substantiated by FCC trials on a pilot FCC unit in comparison with a commercial DA-250 catalyst. So, the introduction of LAY zeolites and metal oxides, which ensure the partial oxidation of feed hydrocarbons in the reactor, in cracking catalysts, allows the yields and the chemical composition of target products to be changed within a wide range, without substantial alterations of operation conditions. Table 3 The effect of the zeolite to metal oxide ratio in TCC reagent catalysts on the yields of vacuum gas oil cracking products and the chemical gasoline composition (reaction temperature 460~ space velocity 1.5 h ~, catalyst to feed ratio 3,5)

'

1. Product yields, wt. % Gasoline Gas Light Gas Oil Coke 2. Gasoline composition, wt. % Paraffins Naphthenes Monoolefins Dienes and cycloolefins Arenes The total amount of olefins

o

0

35.2 10.2 27.2 5.8

40.0 12.5 26.5 5.0

32.3 9.0 32.1 4.8

- 36.4 12.0 28.7 6.7

26.5 8.7 29.5 5.0

36.9 8.5 29.2 7.2

41.1 19.7 22.1 8.8

39.7 31.0 11.1 18.2 11.1

37.4 30.4 5.6 26.6 5.6

30,3 17.9 16.2 13.6 22.0 29.8

30.0 24.1 12.4 11.3 22.2 23.7

29.5 30.7 16.2 23.6 16.2

30.2 29.8 8.3 31.7 8.3

33.4 25.3 6.7 34.6 6.7

4. CONCLUSION The development of new catalytic systems based on the executed researches into catalytic and physical and chemical properties of LAY zeolites and metal oxides incorporated into FCC catalysts shall be continued to create an alternative to the idea that catalytic cracking can only be enhanced by improving the equipment. REFERENCES

1. M.I. Levinbuk, M.L. Pavlov et al., Applied catalysis A; General, 172, (1998) 177 2. M.I. Levinbuk, V.B. Melnikov, V.I. Vershinin, USSR Patent #1812785 (1990) 3. M.I. Rustamov, G.T. Farhadova, S.A. Dorofeev, Chem. Tech. of Fuels and Oils, 6 (1991) 6

Studies in Surface Science and Catalysis 134 M.L. Occelli and P. O'Conner (Editors) 9 2001 Elsevier Science B.V. All rights reserved

111

NExCC TM -Novel short contact time catalytic cracking technology J. Hiltunen a, V.M. Niemi a, K. Lipi~iinena, I. Eilos a, P. Hagelberg a, P. Knuuttila a , K. J~i~iskelainenb, J. Majander b and J. R6pp~inenb aFortum Oyj, P.O. Box 310, 06101 Porvoo, Finland bFortum Oyj, P.O. Box 20, 00048 FORTUM, Finland ABSTRACT The demand for reformulated gasoline requires extending the flexibility of cracking operations to maximise the yield of light olefins and aromatics, both of which are to be used as building blocks for petrochemicals. Olefins are widely used for producing alkylates and oxygenates for the gasoline pool while aromatics are suitable for producing a large number of organic intermediates and chemicals. Fluid catalytic cracking combined with various hydrorefining technologies will be the main means used to meet the coming challenges for the refining industry. Demand for a short and exact contact time, higher temperature and a high catalyst/oil ratio, to produce more light olefins, has led to a new NExCC TM process incorporating a totally new reactor type. The NExCC TM design makes it possible to construct large equipment with a small height-to-diameter ratio. Compared to the conventional FCC, the NExCC TM product spectrum is much more olefinic. The process produces a higher conversion and less heavy components (>221 ~ In a suitable setting NExCC TM can also serve as a source of increased propylene production. The gas velocity in a NExCC TM reactor is lower than in a FCCU riser. The flow type in the regenerator is totally different compared with that of the FCC and the reaction time is shorter. The NExCC TM process utilises multi-entry cyclones for separating catalyst from gas flows both in the reactor and the regenerator. The separation efficiency of the multi-entry cyclone is usually better than that of a conventional cyclone, and it is optimal for handling gas flows with high particle concentrations, such as in the NExCC TM process. Fortum Oyj has been developing multi-entry cyclone technology by using cold and hot model tests with different testing units and by using computational fluid dynamics (CFD) for analysis test results and for detailed cyclone design. For a scale-up of the process, models which predict with reasonable accuracy the flow field and chemical reactions both in the risers and the cyclones are required. In order to study the effect of the reactor operation conditions on the product yields and the product quality a kinetic model has been developed. The model takes into account the gasoline PONAcomposition and includes eight product lumps with eight cracking reactions. A micro scale pulse reactor has been developed to obtain the parameters for the kinetic model.

112 1. FUTURE DEMANDS FOR FCC PROCESS There are three global, partly diverging trends that are transforming the refining industry: * 9 9

Available crudes will be increasingly sour and heavy. (Heavy oil constitutes 13% of current production and is expected to exceed 19% by 2003). Regulations are becoming more stringent concerning sulphur, aromatics, and other product and production based emissions. Demand for lighter, high quality transport fuels is increasing.

The demand for reformulated gasoline requires extending the flexibility of cracking operations to maximise the yield of light olefins used as building blocks for petrochemicals and for producing alkylates and oxygenates for the gasoline pool. The aromatics produced are suitable for a large number of organic intermediates and chemicals. Propylene supply may become short as ethane gas replaces naphtha as steam cracker feedstock. Fortum believes that the demand for polymers based on propylene and butylenes will grow faster than that for ethylene. The use of propylene as an alkylate feedstock will accompany the growing demand for high-octane gasoline. Also, the requirement for olefins for gasoline oxygenates (MTBE, ETBE, TAME) or alternatively for iso-octane will be expected to increase worldwide. Developing countries especially, are continuously looking for more profitable value-added conversion applications for cheaper and heavier feedstocks. Fluid catalytic cracking combined with various hydrorefining technologies will be the main means by which to meet the coming challenges for the refining industry. Table 1 Scenarios forecasting the increasing demand for lighter products and transportation fuels. Million toe 1973 1995 2000 2010 Transportation fuels 1010 1600 1870 2320 Petrochemicals 125 192 250 300 Other non energy uses 155 192 215 250 Heating and industrial fuels 1510 1216 1265 1430 TOTAL 2800 3200 3600 4300 Share of transportation 36% 50% 52% 54% Source: Bonnifay, P. and Vidal, A., (IFP), CatCon'98, June 3-5, 1998, Houston, TX. The fluid catalytic cracking process is the heart of the modern petroleum refinery. The flexibility of FCC enables the refiner to maximise the yield of desired light products and utilise diverse low value feedstocks.

113

2. PROBLEM AREAS OF PRESENT FCC TECHNOLOGY 2.1. Product properties More stringent specification gasoline will be needed in the future. However the components produced by FCC-units are not optimal. FCC gasoline has a high content of aromatics and olefins and a low octane number compared with the reformate. Nearly all the sulphur in the final product comes from the FCC heavy gasoline fraction. Hydrotreating this fraction for sulphur removal will lower the octane number, because the olefins in this fraction will also be hydrogenated. It is known that by using more severe process conditions, FCC could produce more light olefins for alkylation and etherification.

2.2. Investment A typical level of investment for a new FCC unit in the US Gulf Coast area based on 30,000 bpsd feed, including reaction and regeneration systems and product recovery but excluding offsites, power recovery and flue gas scrubbing, is about $2100-2800 per bpsd ($63-84 Million/unit). Industry has optimised FCC technology for large refineries and that is why the conventional FCC cannot be economically downscaled. Fortum believes that in most cases a 30,000-bpsd unit is the minimum economic size for the present FCC technology. 2.3. Erosion One of the most severe problems with the conventional FCC process is the erosion of construction materials, especially in process areas where high temperatures and high velocities are necessary. Mechanical reliability is paramount in FCC construction. Erosion of the components by the high velocities of the hard catalyst particles is inevitable. Practically every surface of an FCC unit contacting catalyst particles will need surface protection against erosion. The degree of protection required will depend on the severity of the conditions. Erosion of the walls of the components is a safety concern because, in the case of a leak, a substantial amount of hot hydrocarbon gases could escape into the environment. 2.2. Catalyst attrition Catalyst attrition also causes a dust problem. High velocities, especially in the riser reactor and cyclones, tends to grind catalyst particles down into micron size dust which is difficult to separate from vent gases and product streams. Particles are usually separated from the product stream in sedimentation tanks. Dust escaping from the catalyst regenerator into the atmosphere is one of the environmental problems of conventional FCC technology. It can be expected that environmental legislation will be enacted to reduce particulate emissions with a possible target level of 50 m g ~ m 3. Also, refiners using power recovery need to remove the fines in flue gases to avoid turbine blade erosion. Regenerator dust emissions under 100 m g ~ m 3 are difficult to achieve without Third Stage Separation (TSS) equipment. These typically involve multiple, relatively small cyclones.

114 3. D E S C R I P T I O N OF T H E NExCC TM T E C H N O L O G Y

3.1. Process description Demand for a short and exact contact time, an increased temperature and a high catalyst/oil ratio to produce more light olefins, has led to the NExCC TM process which incorporates a totally new reactor type. Both reactor and regenerator are operated in the fast fluidized flow regime type and have direct-coupled cyclones. To increase the catalyst regeneration time, part of the regenerated catalyst is circulated back to the bottom of the regenerator. The regenerator and reactor cyclone diplegs are connected directly to the reactor and regenerator risers without large catalyst vessels (Fig. 1). The reactor is operated a higher temperature, a shorter contact time and a higher catalyst/oil-ratio (see Table 2) than conventional FCC reactors. This is to increase conversion, yield of light olefins and the octane number of the gasoline. It is possible to double or triple the feed to propylene, MTBE, TAME and alkylation processing compared with the possibilities in existing refineries. The regeneration time is shorter than that of bubbling bed FCC regenerators, because mass transfer for coke burning is faster. Table 2 Main process parameters for NExCC TM and FCC

Regenerator Reactor

Temperature Residence time Temperature Residence time .. Catalyst/oi! ratio

NExCC TM

FCC

680- 720~ App. 30 s 550 - 620~ 0.7 - 2.2 s 1 0 - 20..

680- 720~ app. 240 s 520- 550~ 2-5s 5-7

TC=temperature control LC=level control LI=level indication PdI=pressure difference indication

Fig. 1. Flow and control principle of NExCCrU-reactor.

115

3.2. Construction description A multi-entry cyclone is more efficient for handling a higher cyclone catalyst loading with lower velocities and a lower pressure drop than the usual single port cyclones. Risers connected to multi-entry cyclones are optimally ring shaped. This type of geometry makes it possible to design a totally new reactor configuration, a demonstration unit of which has been operated since summer 1997. In the NExCC TM unit the reactor and regenerator have been built concentric, whereby the riser channels have an axially annular cross section. Used and regenerated catalyst flows are controlled by special valves, which are located inside the pressure vessel (Fig. 2).

Fig. 2. Schematic construction of NExCC TM reactor.

3.3. Construction benefits The need for a higher cracking temperature, a shorter contact time and a higher catalyst/oil ratio has led to the conclusion that the mechanical restrictions of existing FCC-units prevent the optimisation of the process for maximum olefin production. The greatest advantage of the NExCC TM reactor is that it facilitates construction of more compact units than the current combinations of reactor and regenerator, thus offering a much smaller footprint and easier installation. Insulation linings are used only in the outer shell to keep the temperature of the pressure shell below a metallurgical limit. All other walls between the channels and most of the other reactor parts are made of heat resistant steel. Only in a limited number of parts is more

116 erosion resistant material is required. The location of potential erosion points and the erosion rate has been studied with the pilot unit over more than two years. The NExCC TM design makes it possible to construct large equipment with a small heightto-diameter ratio, whereby the footprint/space requirements of the equipment are minirnised and problems associated with thermal expansion are reduced. NExCC TM reactors have a light construction; even to the extent that they can be designed portable with a modular construction (Fig. 3). 3.4. Investment benefits A compact, economical design with minimum refractory surfaces means savings in investment costs for equipment, piping and support. Preliminary economics for the NExCC TM process compared to conventional FCC units show that capital investment for a NExCC TM reactor is less than half that of the traditional FCC reactor in a 240 t/h (40 000 bpsd) scale. Due to the single vessel reactor / regenerator system, requirements for piping, foot space, steel structure, insulation and the amount of civil work are minimised, which results in most of the investment advantage.

Fig. 3. NExCCrM-reactor size compared to traditional FCC-reactor.

117 3.4. Product benefits Gasoline quality can be significantly improved by increasing the production of oxygenates and alkylate. Blending these components will lower the content of aromatics and olefins in the final product. The production of light olefins for these processes is possible by the catalytic cracking of heavy distillates with more severe conditions than applied in FCC. Moreover, the amount of olefins in the heavy gasoline fraction is decreased and hydrotreating this fraction has only a minor effect on the octane number. Fig. 4 describes how the NExCC TM process can be integrated into a refinery focused on the production of reformulated gasoline and its oxygenate and alkylate components. Alternatively, iso-butylenes can be converted to valuable iso-octane using the NExOCTANE process. Compared to the conventional FCC, the NExCC TM product spectrum is much more olefinic, producing higher conversion and less heavy (>221~ components. In a suitable setting the NExCC TM can also serve as a source of increased propylene production. While gasoline production is reduced by 40%, compensation comes from increased propylene, ether, and alkylate production. Multiple reactors can be operated in parallel with different feeds, with optimum catalyst and optimum process conditions. The construction of the cyclone offers a short and exactly controllable reaction time as the catalyst enters the cyclone simultaneously from each point of the riser top. In the NExCC TM reactor, the catalyst inventory is located in the diplegs. The inventory is much smaller compared with the FCC-units of equal capacity, which means a rapid response to market needs by enabling a fast catalyst change.

Fig. 4. NExCCrM-unit integrated into a refinery.

118 4. NExCC TM PILOT R E A C T O R RESULTS The main function of the pilot reactor was to demonstrate mechanical operability and to be an erosion test unit. Thus, chemistry studies have mainly concentrated on the rnicroreactor tests. Since its construction the pilot unit reactor has been operated for 7600 hours, which includes 2500 hours on feed. The maximum feed rate so far has been 950 kg/h. The only erosion encountered has been found on the cyclone vanes, but this can be easily handled by proper material choice. 4.1. Results with different feeds Two types of feed have been used; hydrotreated vacuum gas oil (VGO) and untreated VGO. The feed properties are shown in Table 3. Table 3 Feed properties. Hydrotreated VGO 910

VGO 908

wt-% ppm

0.1

1.4 1100

Aromatics

wt-%

35

50

GC distillation I]3P 20% 50% 80% FO3P

~ ~ ~ ~ ~

160 385 446 495 577

151 314 379 438 559

Density

kg/m 3

Sulphur Nitrogen

.

.

.

.

.

.

.

.

.

As can be seen from Table 3, the densities of the feeds are almost the same despite the more aromatic nature of the untreated VGO. An explanation for this can be found from the distillation curve, which shows that untreated VGO is lower boiling. For example the LCO fraction (221-343~ content is about 27% in the untreated VGO, whereas it is only 8% in the hydrotreated feed. The NExCCVM-reactor cracking results with the different feeds are compared to "high severity" FCC-yields (base case) in Table 4. Of course a direct comparison is difficult but we think that the yield tendencies between the FCC and NExCC TM and between two types of feeds in the NExCCTM-reactor are quite obvious. Despite the higher riser temperature in the NExCCTM-reactor the conversion is only about two per cent higher than in the FCC. Another feature, as expected, is that conversion of the untreated feed under the same process conditions is lower than that of hydrotreated feed. However, despite the reduced crackability of the untreated VGO the yields of the LPG fraction are almost comparable, hence the lower conversion is reflected primarily in the lower gasoline yield. It is also clear that the higher LPG yield is mostly a consequence of gasoline overcracking.

119 Table 4 NExCC TM results withdifferent feeds compared to typical FCC. .............. FCC Pilot Reactor Partly hydroHydrotreated VGO treated VGO VGO Base +50 +50 Riser temperature (~ Conversion (%)

base

+2.2

-2.2

Yields (wt-%) Dry gas Propylene i-Butane i-Butene C4-alkanes C4-olefins LPG Gaso line LCO+HCO

base base base base base base base base base

+ 1.2 +5.2 -1.0 +2.5 -1.6 +5.2 +8.4 -7.4 -2.2

+2.0 +4.7 -0.6 +2.2 -1.0 +4.7 +8.4 - 12.6 +2.2

.

.

.

.

.

.

.

.

.

,,,,

~_

.

.

.

.

.

.

.

~ .

.

.

.

.

.

.

,

.

.

.

.

,l

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

4.2. The effect of oil partial pressure

The effect of oil partial pressure on the product distribution was tested by changing the ratio of prefluidization gas (nitrogen) and feed oil (hydrotreated VGO). Table 5 shows quite clearly that at the lower oil partial pressure hydrogen transfer reactions (olefin+naphthene --) paraffin+aromatic) are minimised. For example the LPG olefins yield can be increased using dilution gas. Also the i-butane/i-butene ratio can be improved for the feed to post-treatment units (e.g. alkylation, etherification). Table 5 Effect of dilutio n on theproduct distribution:. . . . . . . . . . . . . . . . . . . . . . D ilutant/feedoil (kg/kg)_ _ !.0 . . . . . . . . . . . . . 0:6 LPG fraction ratios Propane/propylene i-B utane/i-butene

0.10 0.60

0.17 1.07

Gasoline fraction yields Paraffins+aromatics 44% 70% - 01efins+Naphthenes .......................5 6 % . . . . . . . . . . . . . . . . 30 % ...... 4.3. Post-treatment without octane loss

Tighter fuel specifications may require reducing the sulphur content in gasoline to a level less than 50 ppm, and it seems that this trend may go down to 10 ppm in the near future. Hence the FCC gasoline, a major source of sulphur in the final product, needs more and more stringent post-treatment to reduce the sulfur content. The sulphur content of FCC gasoline is reduced significantly by removing the heaviest gasoline cut, but this reduces the gasoline

121 5. MULTI-ENTRY CYCLONE TECHNOLOGY The NExCC TM process utilises multi-entry cyclones for separating catalyst from gas flows both in the reactor side and the regenerator side. Fortum Oyj has been developing multi-entry cyclone technology by using cold and hot model tests with different testing units and by using computational fluid dynamics (CFD) to analyse test results and for detailed cyclone design. Flow to the multi-entry cyclones comes through several inlet ports. The flow is directed into a rotating motion by vanes that are placed in a ring on the cyclone circle (Fig. 2). A typical gas entrance velocity to the multi-entry cyclone is at a range of 2 to 10 m/s that results in much less particle attrition and wall erosion rates compared to conventional cyclones.

5.1. Performance of multi-entry cyclones The separation efficiency of the multi-entry cyclone is normally as good as and usually better than the conventional cyclone. In the NExCC TM process there exist high particle concentrations in the cyclone inlet. At high particle concentrations the flow field in the cyclone is dominated by particle flow instead of fluid flow. Therefore high loading ratios should be taken into account in the cyclone design. The separation efficiency has also been estimated also by a model developed by Liesmaki et al. [2] for highly loaded multi-entry cyclones. Typical test results are presented in Fig. 6. The test results are compared with estimated results. The model gives reliable results for particles over 10 l.tm, for smaller particles the air humidity and temperature have a great effect on their flow behaviour. 100

-

"".

.

.

.

.

.

m

--

95 90 o~ 85 >:, c 80 .~_ ._o 0

--B--measured 2.2 kg.cat(kg air

~

==:

|r

75

~

I

.o a~

, ..,

estimated 3.2 kg.cat/kg air

--at measured 3.2kg.cat/kg air i

70

00 65 60

i'

J

55 50

9

0

i

i

5

10

~ "

i

'

,

i

15 20 25 Particle size, microns

'

i. . . . . . .

30

i

35

. . . . . . . . . . .

40

Fig. 6. Measured multi-entry cyclone separation efficiency compared to estimation by the modified Dietz model.

122 The test cyclone is presented in Fig. 7. The test cyclone diameter was 1.6 m. The tests were carried out with a typical FCC-equilibrium catalyst with the particle size distribution shown in Table 6. The air temperature was 55~ an inlet gas velocity of 2.8 m/s and the solids' concentrations were 2.2 and 3.2 kg.catalyst/kg.air. Table 6 Tested solids particle size distribution. Dp (~tm) Distribution % dp (~tm) Distribut!pn % 0,9 0,10 6,0 0,01 1,1 0,03 7,5 0,02 1,3 0,03 9,0 0,03 1,5 0,03 10,5 0,03 1,8 0,02 12,5 0,04 2,2 0,03 15,0 0,08 2,6 0,01 18,0 0,18 3,1 0,01 21,0 0,23 3,7 0,00 25,0 0,24 4,3 0,01 30,0 0,08 5,0 0,00 36,0 0,25

Fig. 7. Multi-entry cyclone test device.

dp (~tm) Distribution % 43,0 1,97 51,0 6,10 61,0 13,18 73,0 20,06 87,0 22,57 103,0 18,95 123,0 12,26 147,0 3,45 175,0 0,00

123

5.2. Computational Fluid Dynamic in the NExCCrM-reactor development During the last decade computational fluid dynamics (CFD) has proven to be a powerful design tool in many fields of engineering [3-8]. Fortum has utilised CFD tools in engineering since 1990. In NExCCrM-reactor development, CFD is used to analyse and assist in the design and construction of the multi-entry cyclone. The benefits of CFD-based modelling lie in identifying the essential mechanisms as well as giving estimates of some process information, which is difficult to measure. Furthermore, the user is able to identify trends and cause and effect relationships. In scaling the process, CFD can direct the design more effectively than the correlation based approach especially with a new type of process solution. Before using CFD in the detailed design, the utility of the physical models was ensured. The focus was on the general behaviour of the flow field. The results were checked against NExCCTM-reactor cold model tests and the measured efficiency of the multi-entry cyclones. In the second step the results have been successfully used in the detailed design of the reactor. Simulations have been used to identify and minimise erosion hotspots caused by the collision of catalyst with the reactor walls, to predict the efficiency of second stage cyclone constructions and to analyse the pressure drops over the NExCCTM-reactor

Fig. 8. Calculated particle concentration profiles.

124 6. MICROSCALE ACTIVITIES IN NExCC TM TECHNOLOGY Gas velocity in the NExCC TM reactor is lower than in the FCC riser, the flow in the regenerator differs from that in the FCC and the reaction time is shorter. To scale-up the process, models must predict with reasonable accuracy the flow field and chemical reactions. To get data for modelling purposes all major variables temperature, C/O-ratio and contact time have to be well controlled and adjusted one at a time. Therefore microscale tools must be fast and accurate enough to measure catalyst performance and deactivation. The advantages of the NExCC TM technology pulse reactors are that the catalyst activity can be considered to be constant due to the very short contact times and the temperature is constant due to the small amounts of injected oil. Furthermore, measurement of coke deposited on the catalyst allows one to determine the energy balance in the cracking reaction as well as deactivation of the catalyst. This also gives an opportunity to specify the reaction kinetic parameters of the catalytic cracking reactions reported by Lipiainen et al. [9].

6.1. Microseale tools of NExCC TM technology A novel piece of equipment for testing catalysts, catalytic reactions and a catalyst deactivation with short contact times has been developed. The reactor is operated in a pulse mode and the products are analysed by on-line gas chromatography. The equipment consists of a glass-tube reactor furnished with a special kind of fast and small Au-film furnace, temperature controller, mass-flow controllers and on-line gas chromatography (Fig. 9). When measuring performance of the catalyst, feed is injected into the upper section of the reactor either manually or by a liquid or gas injection valve. Catalytic reactions occur in the middle section and on-line samples are taken from the bottom section of the reactor with capillary tubing probes. The main variables; temperature, catalyst to oil ratio and contact time can be adjusted over a wide range. The results consist of compound analysis of product from C1 to C~3 and hydrogen, fraction analysis based on GC-distillation, pulse-mode detection of products and quantitative analysis of the coke amount based on detection of regeneration products. Coke measurements can be carried out either with the pulse regenerator (Fig. 10) using oxygen/nitrogen pulses or in a separate unit as a function of temperature (TPO) under continuous oxygen/nitrogen flow (Fig. 11). Due to the small amount of catalyst sample and the extremely small amount of coke deposited on the catalyst regeneration products, carbon monoxide and carbon dioxide are converted into methane over a Ni/~,-A1203 catalyst to achieve high sensitivity for coke measurement.

125

Fig. 9. Construction of the pulse reactor.

126

Fig. 10. Construction of the pulse regenerator.

127

Fig. 11. Construction of the TPO-regenerator.

128

6.2 Major features of the equipment The combination of the reactor and oven with the GC for testing of catalysts gives several superior features compared to other systems with respect to the process variables and analytical techniques. Major features of the pulse reactor, pulse and TPO-regenerators are summarised in Table 7. Table 7. Major features of pulse-reactor and coke measurement equipment. Feature Pulse reactor Pulse regenerator TPO regenerator Temperature range Up to 800 ~ Up to 800 ~ Up to 830 ~ C/O-ratio 0-45 g~t/goi~ Residence time 0 . 0 0 2 - 2.4 s 20-75 s 20-75 s Phase of the feed Liquid or gas 0-20 vol% 02 0-2 vol% 02 Mass of the catalyst 5-500 mg 5-500 mg 5-500 mg S amp 1e On-line On-line On- line Coke analysis CO, CO2, O2 and N2 CO and CO2

6.3. Experimental cracking, deactivation and coke measurement examples The example experiments were carried out over a zeolitic catalyst using middle distillate oil as a feed. Common reaction conditions were 20 mg of catalyst, 0.5 lxl of feed and a residence time of 0.055 s. Cracking reactions in the deactivation test were carried out at 600 ~

10

100

~ e-

80

0

too

/

60

68

13. _.i

* Gasoline A Dry Gas o Coke

O

._~

9 Conversion 9 LPG

(.9

40

40

P,

.~_

I

~ 2o P,

o

r

.

b

0 ~0

4~

500

5~

600

6~

Reaction Temperature, ~

Fig. 12. Results from cracking reaction of middle distillate over a zeolitic catalyst as a function of temperature.

129

Fig. 13. Deactivation of zeolitic catalyst as a function of injection times.

Fig. 14. Coke amount burned off from catalyst as a function of oxidation pulses.

130

Fig. 15. TPO-chromatogram of catalyst regeneration as a function of temperature.

7. KINETIC MODELING

In order to study the effect of the reactor operation conditions on the product yields and the product quality a kinetic model that takes into account the gasoline PONA-composition was developed by Hagelberg et al. [10]. The experiments were carried out in the previously described laboratory reactor with a wide range of C/O-ratios, temperatures and residence times. The kinetic model presented in Fig. 16 includes eight lumps with eight cracking reactions.

kl GAS OIL

"-GASOLINE PARAFFINS

k3

GASOLINE OLEFINS GASOLINE NAPHTHENES

COKE

DRY LPG GAS Fig. 16. Kinetic model with eight lumps.

GASOLINE AROMATICS

131 The kinetic parameters in Arrhenius' rate law were determined by a nonlinear parameter estimation program minimising the square sum of the differences between the estimated and the experimental product yields. The measured and the estimated PONA yields in the gasoline fraction (C5-221 ~ as a function of temperature are shown in Fig. 17.

25Aromatics

Paraffins

20

(

]

~

15

~

A

a

[

Olefins / ~- 10

5 ~

<

Naphthen .~s 0 400

450

500

550

600

650

Temperature, ~

Fig. 17. Measured (symbols) and estimated (curves) yields of paraffins (CJ), olefins (A), naphthenes ( ~ ) and aromatics (O) in the gasoline fraction as a function of temperature. The conversion in Fig. 17 varies from 44% (T = 400 ~ to 86% (T = 650 ~ In this range the gasoline aromatics increase continuously and the olefins and the paraffins reach a maximum after which they decrease rapidly. The decrease in the paraffin yield above 540 ~ is caused by the higher selectivity of gas oil cracking to olefins than to paraffins. The decrease in the olefin yield above 540 ~ is caused by secondary reactions where the olefins are cracked to LPG. A more simple 5-lump kinetic model was implemented in a FCC riser simulation model, where a core/annulus flow region was assumed [11]. The model calculated gas velocity, voidage and temperature profiles as well as the product composition in the axial direction of the riser. The kinetics of coke combustion were evaluated from the experiments made by the TPOmethod with different heating rates and different oxygen partial pressures. Kinetic parameters were determined for both a simple power-law kinetic model of coke combustion to CO and to CO2 and a more complicated semi-mechanistic model [ 12]. An example of a TPO experiment is shown in Fig 18.

132 50

40 E Q. 0 0 d 0

co

30 C02 20

10

9

500

i

600

~

700

!

]

800

900

i ''-~L-

1000

11 O0

Temperature, ~

Fig. 18. Experimental (thick line) and model output (thin line) of coke combustion with power-law kinetics with TPO-method. REFERENCES

1. S.W. Shorey, D.A. Lomas and W.H. Keesom, NPRA Annual Meeting, San Antonio, TX, 1999, AM-99-55. 2. M. Liesm~.ki, R.Karvinen, I. Eilos, Circulating Fluidized Bed Technology VI (Ed. J. Werther), DECHEMA, Germany, 1999, pp. 769-774. 3. K. Savolainen, P. Dernjatin, E. M~ki-Mantila and K. J~i~iskel~iinen, 1998 American Japanese Flame Research Committees International Symposium, Oct 11-15, 1998. 4. S. Korhonen, T. Jacobson and K. J~skel~iinen, American Japanese Flame Research Committees International Symposium, Oct 11-15, 1998. 5. R. D. Lonsdale, Basel World User Days CFD 1992, Conference Proceedings, May 24-28, 1992. 6. M. Forde, Basel World User Days CFD 1992, Conference Proceedings, May 24 -28, 1992. 7. I. Kulmala, VTT Publications 307, 1997, ISBN 951-38-5052-8. 8. R. Lundborg, 1998 Fluent Nordic Users Seminar, September 28-29, 1998, G6teborg, Sweden. 9. K. Lipi~iinen, P. Hagelberg, J. Aittamaa, I. Eilos, J. Hiltunen, V.M. Niemi, and A.O.I. Krause, Appl. Cat A, 183 (1999) 411-421. 10. P. Hagelberg, I. Eilos, J. Hiltunen, K. Lipi~iinen, V.M. Niemi, J. Aittamaa and A.O.I. Krause, submitted for publication, 1999. 11. K. Penttila, P. Hagelberg, I. Eilos and J. Aittamaa, Circulating Fluidized Bed Technology VI (Ed. J. Werther), DECHEMA, Germany, 1999, pp. 213-218. 12. J.M. Kanervo, A.O.I. Krause, J.R. Aittamaa, P. Hagelberg, K. Lipi~iinen, I. Eilos, J. Hiltunen and V.M. Niemi, submitted to ISCRE16, 1999.

Studies in Surface Science and Catalysis 134 M.L. Occelli and P. O'Conner (Editors) 9 2001 Elsevier Science B.V. All rights reserved

133

Effect of v a n a d i u m on l i g h t o l e f i n s s e l e c t i v i t y C.-Y. Li, S.-H. Yan, Z.-H. Qiu, L.-W. Tang and M.-Y. Gu Research Institute of Petroleum Processing (RIPP), SINOPEC, Beijing, China

Abstract The effect of vanadium on FCC catalyst activity has been widely described, but the effect of vanadium on light olefins selectivity has not been fully discussed. In this paper the effect of vanadium on light olefins selectivity was investigated using catalyst capable of enhancing light olefins yields. Keywords: vanadium contamination, light olefins selectivity, cracking catalyst 1. INTRODUCTION The major products from the modern fluid catalytic cracking (FCC) operation are LPG, gasoline and LCO. Periodically commercial or environmental constraints put pressure on the process to modify the composition and yield of these products. For example, environmental regulations coupled with lead phase out have shifted the focus of the FCC from that of an octane barrel producer to that of a light olefins generator. The light olefins are the necessary feedstock for reformulated gasoline blending components such as MTBE, TAME and alkylate. Therefore there is a need to improve the FCC process and catalyst in order to process more residues and generate more light olefms. The Research Institute of Petroleum Processing (RIPP) has developed processes and related catalysts to meet these market demands (1-3). However, the high metals content of residue feedstocks has a deleterious effect on the performance of cracking catalyst. The effect of vanadium on catalyst activity has been widely reported, but the effect of vanadium on light olefins selectivity has not been fully described. In this paper the effect of vanadium on the light olefins selectivity of FCC catalysts having a component enhancing the light olefins yield was investigated using the Micro Activity Test (MAT), with supplemental data being provided by physicochemical characterization and commercial operation.

134 2. EXPERIMENTAL Catalyst samples with ultra stable zeolite Y and proprietary shape selective zeolite ZRP, a rare earth-containing high silica zeolite having penta-sil type structure (4,5), were used in the present study. Some catalyst properties are listed in Table 1. Vanadium was deposited on the samples using a procedure similar to that described in the literature (6). Samples with and without vanadium are referred to as Cat. AV and Cat. A respectively. In some cases vanadium trapping materials based on modified alumina were involved. Cat. AT and Cat. A designate samples with and without vanadium trap respectively.

Table 1. Main properties of catalyst sam_ples Sample Na~O, A120~,

RE:O3,

Surface area

Pore volume

wt%

wt%

wt%

m~/g

ml/g

Cat. A

0.12

48.4

1.1

281

0.19

Cat. B

0.14

39.6

0.26

237

0.16

Cat. C

0.18

46.5

0.56

227

0.17

RFC

0.16

50.4

0.79

233

0.19

The evaluation of catalyst samples was performed using a MAT unit. Before testing samples were deactivated with 100% steam at 800~ for 4 hours. The reaction temperature was 520~ Catalyst-to-oil ratios of 3-6 were used. The weight hourly space velocity (WHSV) was 16 h 1, and a VGO feed (20~ density 0.8652 g/cm 3, boiling range 227 ~ 475~ was used. Surface area was determined by the BET method (ASTM-4641-87). The vanadium content was determined by a wet chemical method. 3. RESULTS AND DISCUSSION Figure 1 shows the influence of vanadium on catalyst activity. With the increase of vanadium deposited on the samples the conversions decrease sharply and the yields of unconverted heavy oil (HCO) increase. Figure 2 indicates the gasoline yields decline on increasing the vanadium contents on the catalyst samples. These results agree with reported literature data (7). Figure 3 shows the influence of vanadium on light olefins selectivity. Increasing vanadium seems to increase the light olefins selectivity. Previous investigations indicate that the vanadium deposited on the catalyst is initially located on the matrix. Steaming mobilizes vanadium to attack both matrix and zeolite via aluminum removal (8). The catalyst samples contain a

135

80.

" ~ 75. ,, "-070.

o~ 10 " ~

Q)

~

2884

o-

=~

-r-

r

r

m

0

55

24

,

o

~oo

'

,o'oo

'

26.

20

,o;o

1o'oo'~o'oo '~o'oo ',o'~

0

V, ppm

V,

Figure 1 Effect of V content on conversion and HCO yield

Figure 2

'

Effect of V content on gasoline yield

~ ~ A J

0.60

o

0,16 :"~ ~>

o

9-,--' 0.55 0 (l.)

(1) 0.1400

O0 0.50 II If)

o

II LO 0.45

0.12 .Q~

I II

II 0.40i 0") 0.10 .('~

O.35 '

0

Figure 3

I

~o0o

'

4o'~ V, ppm

'

~0'~

i

'

Effect of V on light olefins selectivity

8000

,

~o'oo ,o'oo ,ooo

ppm

0.18 0.65

',

136

10

J

<..~

/ II

II

.r 8 II

o-..

,

I

0

~

1000

i

2000

~

I

9

3000

I

J

4000

I

,__J

5000

6000

1000

7000

2000

3000

V, ppm

4000

5000

6000

V, ppm Figure 5

Figure4 Effect of V on light olefins yield

Effect of V on iC4 = 8z i C 5 =

yield

.••Cat

CT2V

9 /

3500pl~ V

.CatCT2V

..e...a

o~68

3~pm V IItit')

r

. o 64

133, II

(:]1..)

r

t---

0

o 60

0

'

I

3

9

9

4

'

I

5

'

I

6

Cat./Oil ratio Figure 6 Conversion as a function of C/O ratio

,

.

~

3400ppmV

I

60

'

'

I

65

'

I

/0

'

I

'

/5

C on v e rsi0 n, wt% Figure 7 Light olefins yield as a function of conversion

I

~11)

7000

137

certain amount of shape selective zeolite ZRP, which has much higher silica to alumina ratio and it is more acid resistant than zeolite Y. Therefore, it is more stable to attack by acidic V§ species. The proportion of shape selective zeolite, ZRP, in the vanadium-containing catalysts of the present study may actually increase, encouraging the gasoline fractions to be cracked into lighter molecules. Figure 4 and Figure 5 show the yields of light olefin and isoolefin on increasing the vanadium contents on the catalyst samples. Below 5000ppm vanadium the yields of light olefins increase with the increase in vanadium, in spite of the decrease of gasoline yields (Figure 2). This indicates that the shape selective cracking reactions increase while the hydrogen transfer activity (HTA) decreases. Since the sum ofisobutene and isobutane is essentially unchanged for certain reaction conditions, the ratios of isobutene to isobutane are frequently used to indicate such hydrogen transfer. It can be seen clearly from Table 2 that with the increase of vanadium deposits the iC4=/iC4~ ratios increase, indicating the reduction of HTA. It is certain that the reduction of HTA is caused by the increase of relative proportion of the shape selective zeolite and the decrease of unit cell size of zeolite Y by the effect of vanadium deposits (9). Therefore the selectivity for light olefins and iso-olefins increases steadily with the increase of vanadium deposits (Figure 3). However, above 5000ppm vanadium the surface areas of the catalyst drop about 50%, indicating partial destruction of the zeolite structure. The activity of the catalyst decreases sharply and not enough gasoline is formed as feed for cracking to light olefins, therefore the yields of light olefins decline. In order to maintain light olefins yield at high vanadium contents, it is necessary to suppress the deactivation of zeolite Y. Table 2 Hydrogen transfer activity of catalyst with various V content Catalyst* V, ppm Surface area, m2/g Cat A

iC4=hC4 ~

0

151

0.39

Cat AV1

3000

117

0.88

Cat AV2

4900

93

1.92

Cat AV3

6600

72

6.80

0

141

0.66

Cat BV1

3000

115

1.06

Cat BV2

5900

98

1.70

Cat BV3

6700

74

3.41

Cat B

*After 800~

100% steam treatment

138 Adding vanadium-trapping materials can retain cracking activity by transporting vanadium from the main active components to these traps. However, the properties of the trapping material might influence the selectivity for light olefins. Table 3 indicates that conventional trapping material T1 improves gasoline selectivity, but has a negative effect on the selectivity to light olefins. A novel vanadium trapping material T~. with improved activity and selectivity to light olefins was selected. The test results are shown in Figure 6 and Figure 7. At similar vanadium contamination levels (3500ppm), the catalyst with vanadium trap % exhibits higher conversion as well as higher Ca: ~ C5= yields. Therefore, T2 is a suitable component for a vanadium resistant catalyst for enhanced light olefins. This can be explained by the proper coordination of main catalytic components and the modified vanadium trapping material. Table 3 Yields of light olefins with and without vanadium trap T1 (MAT results at 75 wt% Conversion) V, ppm

Gasoline, wt%

Ca=~ C5--,wt%

iC4: + iCs:, wt%

Cat CV

3400

31.2

39.0

11.3

Cat CT1V

3700

34.4

36.7

10.5

Catalyst*

*After 800~

100% steam treatment

Table 4 Results of RFC commercial trial in refinery L Feed Properties Product Yields, wt% H2 - Ca Density (20~ g/cm 3 0.8898 Ca ~ C4 CCR, wt% 3.2 C5+ Gasoline Ni, ppm 2.66 LCO V, ppm 5.74 Coke Operation Condition Loss Pressure, MPa 0.15 Conversion, wt% Temperature, ~ 531 Ca-+ C4=, wt% C/O ratio 9.55

2.76 30.23 40.75 18.07 7.66 0.53 81.93 23.57

Based on these results a RFC catalyst was developed and commercialized. The commercial application of RFC catalyst was carried out in a 0.8 Mt/a FCC unit (3) and the main commercial trial results are indicated in Table 4. A blend of 20 wt% vacuum residue with vacuum gas oil was used as feedstock. The yield of LPG + gasoline + LCO is over 88 wt%, and the yield of propene and butene is

139 more than 23 wt%. It should be noted that the vanadium content on the E-cat is around 3500ppm, yet the surface area is still over 100 m~/g and the microactivity is higher than 60 wt%, indicating that the catalyst is vanadium resistant at this vanadium level. 4. CONCLUSIONS This study on vanadium contamination of light olefins-enhancing FCC catalysts confirms that although vanadium has a deleterious effect on catalyst activity, the effect on light olefins selectivity is positive due to the reduction of hydrogen transfer activity. Vanadium trapping materials can increase catalyst activity and gasoline yield, while the selectivity for light olefins depends on the coordination of the various trapping materials and the main catalytic components.

REFERENCES 1. Z. T. Li, W. Y. Shi, R. N. Pan, and F. K. Jing, 206 th National Meeting, ACS Symposium, PREPRINTS, 581 (August, 1993). 2. Z. B. Chen, Y. Q. Huo, L. S. Zhong, Z. Y. Wang and X. Q. Wang, 210 th National Meeting, ACS Symposium, PREPRINTS, 773 (August, 1995). 3. X. X. Zhong, et. al., NPRA Annual Meeting, March 16-18, 1997, AM-97-30. 4. W. Fu, X. T. Shu, H. H. Zhu and M.Y. He, Preparation of High-Silica Zeolite Having Pentasil Structure, CN Patent No. 1058382A (1992). 5. X.T. Shu, et.al., Rare Earth Containing High-Silica Zeolite Having Pentasfl Structure and Process for the Same, U.S. Patent No. 5,232,675(1993). 6. B. R. Mitchell, Ind. Eng. Chem. Prod. Res. Dev., 19, 209 (1980). 7. M. L. OcceRi, Catal. Rev. Sci. Eng., 33, 241 (1991). 8. G. L. Woolery, A. A. Chin, etc., Fluid Catalytic Cracking: Role in Modern Refining, ACS Symposium Series, 375,215 (1988). 9. E. Tangstad, M. Bendiksen, and T. Myrstad, Appl. Catal. A., 150, 85 (1997).

Studies in Surface Scienceand Catalysis 134 M.L. Occelliand P. O'Conner(Editors) 9 2001 ElsevierScienceB.V. All rights reserved

141

R e d u c t i o n of Olefins in FCC Gasoline Shizuka Katoh, Kashima Oil Company, Kashima, J a p a n Makoto Nakamura, Nippon Ketjen, Tokyo, J a p a n Bob Skocpol, Akzo Nobel Catalysts, Amersfoort, The Netherlands

Abstract: The UOP side-by-side Fluid Catalytic Cracking unit (FCCU) at Kashima Oil was revamped in 1996, installing UOP-designed modifications. The unit was altered to new design features which included a short contact time (SCT) riser with a directconnected cyclone riser termination, improved atomization feed nozzles, a catalyst cooler, an expanded air blower and an expanded gas compressor. The revamp allowed the FCCU at Kashima Oil to increase throughput from 20,000 bpsd to 28,500 bpsd with a higher intake of residual oil. The modified unit and altered feed rate and composition produced FCC gasoline with a higher olefinicity. Kashima Oil was sometimes required to blend less FCC gasoline into the refinery gasoline pool as the higher olefinicity caused excessive olefins in the gasoline pool. In 1997, Kashima Oil switched to an FCC catalyst designed by Akzo Nobel to reduce olefin content while maintaining the RON of the gasoline. This new technology for the control of gasoline olefins succeeded in reducing olefins concentration in FCC gasoline by more t h a n 5 vol% without significant RON decrease. This technology is marketed under the name Total Olefins Management (TOM) option. F C C Unit Revamp H i s t o r y in K a s h i m a Oil Co., LTD, Kashima Refinery Kashima refinery in J a p a n was built to provide fuel to the Petrochemical and Power plant in the Kashima Petrochemical Industrial Area. The FCC unit was built in 1969 as a UOP side-by-side, semi-bed cracking unit with 14,000 bpsd capacity. Later, due to changing economic conditions, the objectives of Kashima Refinery changed to increased gasoline and diesel fuel. In line with the change of objectives,

142 the FCC unit was revamped to increase gasoline production and maximize atmospheric residue feed. In 1986, the capacity was increased to 20,000 bpsd by changing the feed injection nozzles to the shower head type and implementing allriser cracking. This also allowed the Desulfurized Atmospheric Residue (DSAR) in the FCC feed to be increased to 30%. In 1996, the latest UOP technology for RFCC, i.e., lift gas, elevated Optimix TM feed nozzles, direct connected cyclones, and a catalyst cooler, were incorporated during the revamp. The air blower capacity was increased and the wet gas compressor was revamped. The capacity of the unit was thereby increased to 28,500 bpsd. The results of the performance evaluation test for the nozzles are as follows. Table 1

P e r f o r m a n c e e v a l u a t i o n t e s t for t h e n o z z l e s

OPTIMIX'"

Wye-Pre-MIX

A

C H A M B E R A U G - i AUG-6

F E E D , KL/DAY

4311

4376

RX T E M P ~

515

515

VOL%

WT%

VOL%

DELT

0 WT%

VOL%

F u e l Gas

5.4 foe

5.38

4.2 foe

4.42

-1.2

C3's

8.3

4.69

7.6

4.31

-0.7

C4's

13.7

8.79

113.3-

8.58

-0.4

FCC G a s o l i n e

55.8

44.78

59.4

47.55

3.6

L i g h t C y c l e Oil

19.7

9 19.94

20.2

20.50

0.5

10.24

7.8

8.83

-1.3

112.5

100.76

0.4

!

i

! S l u r r y Oil

9.1

I TOTAL

112 1

!

:100.62

i i

Normally when DSAR was treated by the FCC unit, the gasoline yield went down due to lower conversion. However, it was demonstrated that UOP Optimix TM feed nozzles help to maintain conversion.

143

The E f f e c t o f R e s i d u e o n G a s o l i n e O l e f i n s When the a m o u n t of DSAR in the FCC feed is increased, and at the same time the total feed rate is increased, the olefin content also increases.

Figure 1

FCC Gasoline Composition vs Feed Rate and Resid Intake

In the K a s h i m a FCC unit, DSAR is the incremental feed. Therefore it is impossible to separate the effects of increased resid from increased total feed rate. Figure 1 shows t h a t increased resid and/or feed rate leads to an increase in gasoline olefinicity. For the 3 data shown, the catalyst type, metals level, and FCC h a r d w a r e are the same. R a t h e r t h a n being due to the resid itself, the m e c h a n i s m m a y actually be related to decreased contact time between catalyst and cracking products at high feed rates. The reduced contact time would lead to less hydrogen transfer, so fewer gasoline olefins would be s a t u r a t e d 1. Whatever the mechanism, the FCC gasoline olefinicity went up significantly.

144

RON, MON, and Olefins in Japan Throughout Japan, only Research Octane Number (RON) is measured, not Motor Octane Number (MON), yet the industry is aware that both RON and MON are needed for good drivability. This may be one of the reasons that Japanese companies usually put a limit on olefins in the gasoline pool. The explanation for this may be that olefins are known to have a larger delta between RON and MON t h a n paraffins or aromatics. By limiting the olefins content of gasoline, a sufficiently high MON is ensured, even when only RON is measured. Following this line of reasoning, the limit on gasoline olefins in J a p a n was originally simply a way to ensure good gasoline drivability. Whatever the exact reason, the limit on olefins in J a p a n predated the current environmental concerns about gasoline olefins which are driving the regulation in North America and Europe. At Kashima refinery, the RON target for FCC gasoline is around 92.5, based on gasoline component balance. Reaction temperature needs to be kept between 515~ and 525~ to obtain the required RON.

The P r o b l e m of I n c r e a s i n g FCC G a s o l i n e Olefins Three changes at the FCC contributed to an increase in gasoline olefins in the refinery pool: 9

The increase in feed capacity at the FCC unit

9

The new hardware, which gave higher yields of gasoline

9

The increased olefins due to increased DSAR in the feed

Staying within the gasoline pool olefins limit became a problem. Operating changes at the FCC to reduce olefins, such as reducing reactor temperature, were not acceptable due to the simultaneous reduction in RON. Therefore, FCC throughput had to be reduced.

145 Catalytic Solution

The catalyst manufacturer proposed a new catalyst formulation, called TOM Cobra. TOM stands for "Total Olefins Management." This technology works by 3 mechanisms: 9

Saturating the olefins by increased hydrogen transfer

9

Selectively cracking the gasoline olefins to LPG

9

Increased branching of saturates and/or olefins

The increased hydrogen transfer is achieved by using a special zeolite technology. In part this is accomplished by increased rare earth 2, but other proprietary technology is also employed. The selective cracking of gasoline olefins is accomplished with a ZSM-5-based additive, which cracks gasoline olefins to make C3-C4 olefins. The ZSM-5 also helps to recover the octane lost in the saturation of the olefins 3,4. Increased branching may occur by the use of zeolites with a high silica-to-aluminaratio (SAR) 2. High SAR is a feature of the ADZ (Akzo Nobel Developed Zeolite) used in the formulation. By using the combination of these features, the TOM Cobra can reduce olefins at constant RON. TOM as a M o d e l

TOM technology can actually be thought of as a model which predicts the effects of a combination of catalyst and operating changes. It can be used to optimize the combination of technologies used to reduce gasoline olefins or to accomplish other desired changes in the yields of C3 or C4 olefins. In this case, the target was a 5 LV% reduction of olefins in gasoline for Kashima's total FCC gasoline. R e s u l t s o f Trial of TOM

Figure 2 shows that the new catalyst gives ~7% lower olefins in the light gasoline at the same RON. The light gasoline 90% distillation point is approximately 70~ and is essentially the same for both catalysts.

146

Figure 2

Light Gasoline Olefins with Base Catalyst and TOM Cobra

The reduction in olefins in light gasoline for the new catalyst is confirmed by the fact that the light gasoline is slightly less dense, for the same boiling range. Olefins are denser t h a n paraffins. The reduced density indicates reduced olefins. This can be seen in Figure 3.

Figure 3

Light Gasoline Density vs. Distillation

147 Figure 4 shows that at the same RON, the new catalyst gives - 9 LV% lower olefins for the full-range gasoline (defined as debutanizer bottoms, full-range includes both light and heavy gasoline). The full-range gasoline 90% distillation point is approximately 170~ and is essentially the same for both catalysts.

Figure 4

Full_Range gasoline vs. Octane

The new catalyst met and exceeded the target of 5 LV% less olefins. This allowed the fraction of FCC gasoline in the refinery gasoline pool to be increased. This led to significant economic benefits for the refinery. Analysis of the gasoline by FIA shows that an increase of both saturates and aromatics accompanies the decrease in gasoline olefins. Figures 5-6 s h o w - 3 % increase in aromatics a n d - 6 % increase in saturates in the full-range gasoline with the new catalyst, at constant RON.

148

Figure 5

Full_Range Gasoline Aromatics vs. Octane

Figure 6

Full_Range Gasoline Saturates vs. Octane

149 Figures 7-8 show that the 7% olefins removed from the light gasoline by the new catalyst are replaced b y - 1 % aromatics a n d - 6 % saturates.

Figure 7

Light Gasoline Aromatics vs. Octane.

Figure 8

Light Gasoline Satulates vs. Octane.

150 Figures 2 and 4-8 can be s u m m a r i z e d as follows: the new catalyst gave the following shifts in composition at constant RON: Light gasoline: 7% olefins

Full-range gasoline: olefins

9

-

9

-9%

9

+1% aromatics

9

+3% aromatics

9

+6% saturates

9

+6% saturates

The increase in aromatics in light gasoline is quite modest; in fact as you might expect for a light gasoline with a low endpoint, all of the samples showed less t h a n 1.5 LV% aromatics, with or without the new catalyst. This suggests t h a t the new catalyst m a y be causing isomerization leading to increased branching of light gasoline molecules, helping to compensate for the lower octane of the saturates. This cannot be proved, however, because detailed gasoline composition, including branching, is not available. Further Capacity Increases The FCC unit capacity was further increased in 1998 and the FCC unit is now operated at 28,500 bpsd with 50% DSAR in feed. Both the feed increase and the increase in residue contributed to an increase in coke yield and hence air rate in the now seriously undersized regenerator. The m a g n i t u d e of the coke increase over time can be seen in Figure 9. Even with the i m p l e m e n t a t i o n of p a r t i a l combustion, regenerator gas superficial linear velocity is now close to 4 ft/sec. Therefore, K a s h i m a Refinery looked at catalyst fine-tunings and new catalyst technologies which would result in a lower coke yield. After commercial trials of 3 different low-coke catalysts, they chose a low-coke version of the TOM Cobra catalyst as the best. The low-olefins benefits of the original catalyst are still continuing with the new fine-tuned version.

151

Figure 9

Wt% Coke Yield vs. Time

Conclusion The FCC unit is central to refinery economics. FCC catalysts need to do more t h a n just cracking; they also need to provide technology to control the composition of the products. In this example, the olefins in the gasoline pool increased due to higher FCC feed rate, higher yield of FCC gasoline, and higher % olefins in gasoline, all of which were a direct or indirect result of the successful revamp of the FCC hardware. Olefins in the gasoline pool began to limit the FCC feed rate. A combination of new catalyst technologies collectively known as TOM (Total Olefins Management) was successfully applied, reducing the FCC gasoline olefins by -7-9% at constant Research Octane. The result was a higher feed rate allowed at the FCC and improved refinery economics.

152 References

(1) Scherzer, J., "Octane-Enhancing, Zeolitic FCC Catalysts: Scientific and Technical Aspects, VIII. Reaction Mechanism of Hydrocarbon, D. Reactions over Octane FCC Catalysts", Catal. Rev.-Sci. Eng., 31(3), pp. 300-307 (1989). (2) O'Connor, P., "Advances in catalysts for FCC octanes," Akzo Nobel Catalysts Symposium, paper F-1 (1988). (3) Buchanan, J.S., "Reactions of Model Compounds over steamed ZSM-5 at simulated FCC reaction conditions," Applied Catalysis, 74, pp 83-94 (1991). (4) Biswas, J., and L. E. Maxwell, "Octane Enhancement in Fluid Catalytic Cracking, I. Role of ZSM-5 Addition and Reactor Temperature," Applied Catalysis, 71, pp 1-18 (1990). A u t h o r s : Shizuka Katoh, Kashima Oil Company, 4 Towada, Kamisu, Kashima-gun, Ibaraki, 314-0198 Japan Makoto Nakamura, Nippon Ketjen, 1-5-9 Shiba, Minato-ku, Tokyo, 105-0014 Japan Bob Skocpol, Akzo Nobel Catalysts, Stationsplein 4, P.O. Box 247, 3800 AE Amersfoort, the Netherlands

Studies in Surface Science and Catalysis 134 M.L. Occelli and P. O'Conner (Editors) 9 2001 Elsevier Science B.V. All rights reserved

153

Gasoline sulfur removal: kinetics of s compounds in FCC conditions A. Corma, P. Gullbrand, and C. Martinez Instituto de Tecnologia Quimica, UPV-CSIC, Avda. de los Naranjos, s/n, 46022Valencia, Spare. By cracking a vacuum gas oil (1.4 %S) as well as pure sulfur compounds such as thiophene, tetrahydrothiophene, 2-methylthiophene and benzothiophene on Fluid Catalytic Cracking (FCC) catalysts, and following selectivity curves, a reaction scheme for the reactive sulfur compounds has been established. It is shown that H2S is a secondary unstable product, mercaptans are primary unstable, thiophene is primary stable, alkylthiophenes are primary unstable, and benzothiophene and C 1-, C2- and C3-benzothiophenes are primary unstable plus secondary products. Kinetic rate constants for the formation of all primary products have been calculated, and the influence of reaction temperature and catalyst composition on the sulfur product behavior is presented. 1. INTRODUCTION High levels of sulfur in gasoline have been linked to poor performance and shortened life of catalytic converters, as sulfur emissions poison the active noble metal in emission control catalysts. Thus, lower sulfur results in a more effective catalytic converter. Accordingly, gasoline sulfur content is being restricted by legislation both in the European union and in the USA. FCC gasoline represents about one third of the gasoline pool but accounts for almost 90% of the total gasoline sulfur. Then, in order to effectively reduce gasoline pool sulfur, a decrease in the sulfur content of FCC gasoline becomes mandatory [ 1, 2]. Among different possibilities for removing sulfur from FCC gasoline, the most cost effective would be to reduce the sulfur directly in the FCC process by means of a sulfur removing catalyst additive [2-4]. In order to develop such a catalyst additive, it is important to understand the reactivity of the sulfur compounds during the FCC process. The objective of this work is to make a kinetic study on the reactivity of sulfur compounds present as reactants or products in the catalytic cracking process, with the aim of producing a global reaction scheme of the sulfur compounds present during FCC conditions. 2. EXPERIMENTAL 2.1 Catalysts In the first part of this work, a commercial equilibrium FCC catalyst which contains 1 wt% of rear earth (RE), with a surface area of 170 m2/g and a unit cell size of 2.43 5 nm, was used to process a vacuum gas oil (VGO) feed (Table 1). In this way the selectivities to the

Corresponding author. Tel." 34-96-3877800, Fax: 34-96-3877809, e-mail: [email protected]

154 different reaction products were obtained at constant catalyst/oil (C/O) ratios and increasing times on stream (TOS). Table 1 Properties of the Arabian Light Vacuum Gas Oil. Density 288 K (g/cc) Aniline Point (wt %) Sulfur (wt %) N2 (ppm) Average molecular weight Viscosity (c.s. at 373 K) Refraction index at 340 K CCR (wt%) K(UOP) Distillation curve D-1160 in wt% (K): IBP 5 10 30 522 604 633 690

0.9072 91.2 1.4 890 438 6.29 1.488 0.32 12.03 50 721

70 751

90 812

In the second part of the work, pure sulfur compounds were cracked under the same experimental conditions as VGO. A series of well-designed catalysts were used for these experiments. Pure 2-methylthiophene was reacted on the commercial equilibrium FCC catalyst with 1 wt% RE described above. Thiophene, tetrahydrothiophene, 2-methylthiophene and benzothiophene were cracked on a pure high unit cell size USY zeolite. Finally pure thiophene was converted on three USY zeolites with different UCS (see Table 2). Table 2 Unit cell size (UCS), framework (FAL) and extraframework (EFAL) aluminum content of the USY zeolites. Catalyst Zeolite U.C.S (nm) FAL EFAL USY-1 2.448 28 25 USY-2 2.446 26 7 USY-3 2.426 3 50

In the last part of this work, H2S was reacted with 1-butene on the commercial 1 wt% RE FCC catalyst. 2.2 Reaction Procedure

Study of sulfur product selectivity in catalytic gas oil cracking. The catalytic cracking reaction of the VGO was performed at two different reaction temperatures, 783 and 873 K, in an automated Micro Activity Test (MAT) unit, designed in accordance to the ASTM D-3907. This unit can be programmed to perform cyclic experiments (reaction, stripping and regeneration) in a wide range of experimental conditions. The experiments were performed at different catalyst to oil (C/O) ratios (g catalyst/g feed) with a set of different reaction times or times on stream (TOS) for each C/O, in order to obtain different degrees of conversion.

155 The amount of feed (0.5 g) was maintained constant in all the experiments and the amount of catalyst was changed in the different C/O series. The purpose of using a small amount of feed was to achieve, with an acceptable backpressure, very short times on stream (TOS = 4 and 5 s). Before each experiment the reactor was purged with a 30 cc/min N2-flow for 30 min at the actual reaction temperature. Then, N2 flow is stopped and the gas oil was fed. When the reaction was finished, the catalyst was stripped for 15 min using a N2 flow of 30 cc/min. The catalyst was regenerated in an 100 ml/min air-stream for 3 hours at 813 and 903 K for reaction temperatures of 783 and 873 K respectively. Coke was determined in gravimetrically during this regeneration step. The combustion gases pass through a copper reactor at 623 K and after total oxidation the CO2 was adsorbed on an ascarite CO2-trap, while the H20 was adsorbed with a drierite trap.

Reaction with pure sulfur compounds. 2-Methylthiophene (2-MT, >98%), thiophene (TIO, >99%), tetrahydrothiophene (THT, >99%) and benzothiophene (BT, >95%), all supplied by Aldrich, were fed as pure compounds to the MAT unit, at 783 K, at 30 s TOS and at C/O of 3-6, and cracked over the different catalysts. The amount of catalyst used was 5 g in the case of the 1%RE FCC catalyst and 3 g in the case of the USY zeolites.

Reaction of H2S with 1-butene. These reactions were carried out in a fixed bed reactor, at atmospheric pressure, and using N2 as a carrier gas. Blank tests were performed using 5 ml of carborundum granules instead of catalyst. The influence of contact time and olefin/H2S ratio was studied at a reaction temperature of 510~ 0.4 and 2.0 g of catalyst were used respectively for the catalytic tests performed at the two different contact times, W/F=27 and 134 h (W/F = g catalyst/mass flow of H2S in g/h). In all cases the catalyst was diluted with carborundum up to a constant bed volume of 5 ml. The olefin/H2S ratios used were 10 and 20 mol/mol, values close to those found for (C 1-C4)/H2S ratios in the gas products at the outlet of the MAT unit. Temperature influence was also studied, at the highest contact time and at a 1-butene/H2S ratio of 10.

2.3 Analysis of reaction products The gases were analyzed using a Varian 3600-GC equipped with three detectors: a Thermal Conductivity Detector (TCD) for analysis of H2 and N2 which are separated in a 15 m molecular sieve column, a Flame Ionization Detector (FID) for C1 to C6 hydrocarbons separated in a 30 m Plot/A1203 column, and a Pulsed Flame Photometric Detector (PFPD) for the sulfur compounds separated in a 30 m GS-Q megabore column. The signal detected by the PFPD, after linearization, is proportional to the sulfur concentration independently of the nature of the sulfur compound. Simulated distillation of the liquids was carried out following the ASTM-2887-D procedure. Cuts were made at 150.8~ for light gasoline, 216.3~ for heavy gasoline and 359~ for LCO. In addition to this the sulfur compounds in the gasoline fraction and in the light diesel range (up to a boiling point of 286.8~ were separated, identified and quantified using a Petrocol-100 fused silica column installed in a Varian 3400-GC, equipped with two detectors (FID and a PFPD) connected in parallel at the outlet of the capillary column.

156 3. R E S U L T S AND DISCUSSION

3.1-Study of sulfur product selectivity in catalytic gas oil cracking In this part of the work, a VGO is cracked at two temperatures (783 and 873 K) and several C/O ratios, in order to have different levels of conversion. Different times on stream were studied at each of the constant C/O ratios, and in this way the Optimum Performance Envelopes (OPE's) were obtained for each product [5]. The curves obtained at increasing TOS for each of the C/O ratios are the constant C/O loops, and the OPE's are the curves tangent to the constant C/O loops when TOS=0. Thus, they represent the selectivity one would obtain in absence of deactivation. The OPE's can be obtained in a theoretical way with an infinite C/O ratio, or in an approximate way by working at short TOS (4 s), conditions where the deactivation of the catalyst is small. The selectivity curves for gases, gasoline and coke obtained at 783 K are given in Figure 1. It can be seen there that gases are primary plus secondary products, gasoline is primary and unstable and coke can be considered as a secondary product. 90

.,

20

9

18

8O 70

: "....

oo

..!i::

'-

A

so

A

T- z0

9~ 40

8 6

30 o

20

_

t,9

;

4

10

2

0 0

20

TO40(s)S

0

60

20

40

60

80

MAT Co nversion (wtO/o)

lot

5O

14

A

.0

3s 9u

~12

/~AA

3O

,~ >. 25

"O 8

e-

20

"--

15

~6

m (,9

10 5

2

0 0

20

40

60

MAT Conversion ( w t % )

80

zoo

0 0

20 40 60 80 MAT Conversion ( w t % )

loot

Fig. 1. Activity (a) and selectivity curves for gases (b), gasoline (c) and coke (d) obtained for catalytic cracking of gas oil in a MAT unit, at 783 K, on a l%RE-equilibrium catalyst. ( I ) C/O=0.1 9 ( e ) C/O=0.25 9 ( o ) C/O=0.5; ( A ) C/O=2.0; ( ~ ) C/O=3.0. Figure 2 shows the selectivity curves obtained at the same temperature for the sulfur compounds in the gasoline and light diesel range (up to Tb=559.8 K). In figure 2-a we can see that H2S is mainly a secondary product. The low initial slope obtained for this compound is a

157 sign of this secondary behavior. Perhaps some primary H2S is formed but the amount is very small. In any case, the primary or secondary nature of H2S will depend very much on the characteristics of the feedstock, since it could be formed as a primary product only by cracking mercaptans or sulfides that in some cases can exist as alkyl chains of the aromatic components. The proportion of these type of sulfur compounds in the feedstock will depend on the maturity of the feed. However, most of the sulfur can be expected to be present as thiophenes and benzo and multi-ring thiophenes (1). It interesting to note that at increasing conversions H2S presents an unstable character, indicating that when it is formed it can further react and be consumed. We will come back later to this point. Mercaptans (figure 2-b) are primary and unstable products, as they can crack very easily giving H2S and an olefin. H2S obtained in this way appears as secondary. Thiophene (figure 2-d) and alkyl-thiophenes (see C l-thiophenes in figure 2-e) are all primary products. However thiophene is a stable product, while the alkylthiophenes are unstable and can further react by hydrogen transfer. This reaction will be easier for alkylthiophenes since they can form a tertiary carbocation instead of a secondary carbocation, which would be formed by thiophene.

Thiophene

THT HT

Me-Thiophene

Me-THT

Alkylthiophenes can also dealkylate to give thiophene or can transalkylate with another alkylthiophene. In any case, reactivity increases when a side chain is present and this agrees with the unstable behaviour of C 1- to C4-thiophenes. As the length of the side chain increases (C3- and C4-thiophenes) an additional secondary character is observed, although these compounds are still unstable. Tetrahydrothiophene (figure 2-c) appears as a secondary and unstable product, since when formed it can crack very rapidly in our reaction conditions, in the same way as mercaptans do. Figure 2-f shows the selectivity curve obtained for benzothiophene. The curves obtained for the alkyl-benzothiophenes are very similar to that of benzothiophene and the behavior of these compounds is very interesting. They appear as primary and unstable at lower levels of conversion, but when severity is increased they appear as secondary and stable products. The same study has been performed at 873 K, and the selectivity behavior obtained for the different products is practically the same. However, it is interesting to see that an increase in the reaction temperature produces a decrease in the yields of H2S and THT, both secondary products that involve in some step hydrogen transfer reactions (figure 3). Moreover, at the

158 higher reaction temperature their formation is delayed in some way, and they start to be important at higher conversions. 1E-02

2.0E-04

9E-03-

1.8E-04

,~, 8E-O3-

1 6E-04

7E-O36E-O3-

9~ 1.2E-04

5E-O3-

'~'-- 1.0E-04

4E-O3-

,~ 8.0E-05

~'6.0E-05 -

3 E - 0 3 - ~ 2E-03-

A

1E-03-

.A

OE+O0

2.0E-OS

...,

0

~

4.0E-05 O.OE+O0

20

40 60 80 MAT Conversion ( w t % )

100

4.5E-05

5.0E-05

4,0E-05

,~, 4.5E-05

~

, 3.5E-05

v

20

0

40 60 80 MAT Conversion ( w t % )

100

d

4.0E-05 3.5E-05

3.0E-05

9~ 3.0E-05

2.5E-05

'~, 2.5E-05

"~ 2.0E-05 '~, 1.5E-05 i1 1.OE-05

~ 2.0E-05 1.5E-05

,... ~" I.OE-O5 5.0E-06

5.0E-06

O.OE+O0

O.OE+O0 0

20

40

60

80

20

100

40

60

80

10C

80

10C

MAT Conversion ( w t % )

MAT Conversion ( w t % ) 6.0E-04

2.0E-04

~

~ 1.6E-04 9~ 1.4E-04

5.0E-04

' 4.0E-04

9~ 1.2E-04

.~ 3.0E-04

1.0E-04 8.0E-05

~" 2.0E-04

,~ 6.0E-05 4.0E-05

1.OE-04

,,~ 2.0E-05 UO.OE+O0 .

0

.

.

20

.

.

40

O.OE+O0

.

60

MAT Conversion ( w t % )

80

100

0

20

40

60

MAT Conversion ( w t % )

Fig. 2. Selectivity curves for sulfur compounds obtained for catalytic cracking of gas oil in a MAT unit, at 783 K, on a l%RE-equilibrium catalyst: (a) HzS, (b) Mercaptans, (c) tetrahydrothiophene, (d)thiophene, (e)Cl-thiophenes, (f)benzothiophene. (m)C/O=0.1; ( e ) C/O=0.25; ( o )C/O=0.5; ( A )C/O=2.0; (~) C/O=3.0. On the other hand, yields to thiophene and benzothiophene increase, and both products are obtained by direct cracking of molecules in the feed. The effect of the temperature on sulfur product distribution can be explained taking into account that cracking has a higher activation energy than hydrogen transfer reactions. Thus, when increasing reaction temperature cracking will be more favored than hydrogen transfer with the corresponding increase in the selectivity to thiophenes and benzothiophenes and the decrease in H2S and THT.

159

1,OE-02 9. OE-03 -

i

,"

8. OE-03 -

,:

7. OE-03 -

4.5E-OS 4. OE-O5 -

J

,,,"""

3. SE-OS -

..."^"

;"

6. OE-03 -

F

5. OE-03 -

:"

4. OE-03 -

.-"

'~'i 2.0E-OS 2.-31"0E0S S5E-OS .E-OS

_--

,'--;~~~'~'7;" ";'9"" Y'"

/ ~ ' ~ ~

3. OE-03 2.0E-03 -

1.OE-O5 -

1.OE-03 -

S.OE-.06 -

~.OE+O0 -.=--=-...... "'"" . 20

. . 40 60 MAT Conversion (wtO/o)

i~I

.

80

|,,... ~.OE+O0 . ,' 0

9

, 20

~

~ A

'~."

, , 40 60 MAT Conversion (wtO/o)

, 80

100

7.0E-04 9

9.0E-05 8. OE-05 -

.//~A ~

6.0E-04 -

7.0E-05 -

b

.

~

5.0E-04

~ 6.0E-05 4.0E-04 -

5.0E-05 4.0E-05 .~ 39

"~ 3.0E-04 -

-

w

,,... .......... "

1.0E-05 O.OE+O0 0

~ 2.0E-04 -

20

, , 40 60 MAT Conversion (wtO/o)

1,OE-04 O.OE+O0 0

, 80

. 20

.

. 40

. 60

80

100

PlAT Conversion (wtO/o)

Fig. 3. Effect of reaction temperature on initial selectivity to sulfur compounds'9 OPE curves at 783 ( ...... ) and 873 K (---- ). (a) H2S, (b) tetrahydrothiophene, (c) thiophene and (d) benzothiophene. ( i ) C/O-0.1; (o) C/O=0.25 9 ( o ) C/O=0.5; ( A ) C/O=2.0; ( ~ ) C/O=3.0. In order to produce more quantitative results, a kinetic study was undertaken. To do this, experiments were performed at different TOS for a set of catalyst to oil ratios ranging from 0.1 to 3, and the results obtained were fitted to the kinetic three-lump model for gas oil cracking [5, 6]: Gas oil

[

k3 This model assumes a first order rate expression for the conversion of each of the hydrocarbons found in the feed. Then, the gas oil conversion into gases, gasoline and coke can be expressed as dC dv

160 where "t is the spatial time over a fixed reaction bed, CA is the concentration of gas oil in the gas phase and k is the global rate constant associated with the cracking mechanism. This constant can be expressed as a function that includes the kinetic rate constant at zero time on stream and a catalyst deactivation function ~bthat is a function of the time on stream [5]: (2)

k = k o . qb

Considering that the initial feed composition changes during the reaction, and it becomes more refractory due to an increase in the concentration of polyaromatic compounds, the rate expression was changed accordingly by introducing a refractory index W [6]

k-

ko. # 9

cAlw

~Tf~o

(3)

where CAo is the initial concentration of gas oil. Substituting the modified rate constant given by (3) in the initial rate expression (1) and rewriting in terms of conversion: 1+ W

dX dr

A .~ KT

1- x A !

.(1 + Gt ) - "

(4)

I-+ EX A )

where e accounts for volume expansion caused by cracking. This factor is estimated to be 2.5 [5]. The spatial time can be expressed as: WcCAo

r - ~

(kg cat.s.m 3)

F~o

r - car,~ oil. TOS./9vapo P~por

MW. P R. T

r

(5) (6) (7)

with M W = average molecular weight of the feedstock A C30H62hydrocarbon was used in this work. P = pressure (atm) R = 0.082 (atm.l.mol-l.K 1) T - temperature (K) Since the conversions measured in the MAT unit are averaged over time, the instantaneous fraction converted is used to calculate the theoretical average conversion, in order to compare the experimental and theoretical results.

Xtheor = - 1 i

X~dt

tf ~o

(8)

161 The differential equation has been solved using a fourth order Runge-Kutta technique. To use this method, conversion must be zero when spatial time is zero, and the values for spatial time v have been considered to be much smaller than the reaction time t. In this way, for every constant TOS the differential equations were solved for different values of v. The optimization was carried out by minimizing the relative average error, the F Fisher parameter and the ~-Exner parameter [7, 8], and the final results are shown in figure 4-a for 783 K and 4-b for 873 K. 0.9 ,~

0.7 c 0

9~

11

_~ 0.9 -I 0.84 0.7 q 0,6 -t

0.8-~

~"~

0.6 0.5

> 0.4c

~

0.2-

I'~

o.s~

~e

0.4-t

'i

0.3t 0,2 0.1

I

o

0.120

m

40

60

40 T O S (s)

l

A

9

TOS (s)

~

0.7

0.6

~.~~,~....~_~_

0.5 .2

0.4 0.3

o

u

0.2

lE 0.I i

0

20

i

40 T O S (s)

i

60

80

Fig. 4. Experimental (symbols) and simulated (lines) average conversions obtained over a l%RE-equilibriumcatalyst at different TOS and Cat/oil ratios. (a) T=783 K; (b) T-873 K; (c) T=873 K, thermal conversion withdrawn. (~) C/O=0.1; ( e ) C/O=0.25; ( A ) C/O=0.5; ( ~ ) C/O-3.0. The statistical parameters are given in Table 3 together with the kinetic parameters, in where it can be seen that the fitting was satisfactory, and that reliable values for the kinetic and decay constants were obtained. It was observed that the model describes better the cracking process at 783 than at 873 K (figures 4-a and 4-b). This can be due to the larger contribution of thermal cracking at the higher temperature. In fact, if we define a "kinetic" conversion as the global conversion to gases, gasoline and coke, minus the thermal conversion, the model described by equation 15 gives better results (figure 4-c). However, as thermal and catalytic cracking are not independent phenomena they can not be treated separately. Thus, for further discussion we will consider the kinetic parameter obtained at 873 K with values of the global conversion. As expected, the kinetic rate constant is higher for the highest temperature. The refractory index W has a higher value at 783 K. This can be justified considering that W

162 represents the resistance to cracking of the feedstock when increasing conversion. Then, an increase in temperature will decrease this resistance. Table 3. Kinetic parameters obtained by fitting of the experimental results to a three lumps model. k0=kinetic rate constant; Kd=decay constant; m=order of decay, where G=(m-1)-I~ and N=l/(m-1), and W=refractoriness index [5, 6]. Kinetic parameters T = 783 K T = 873 K k0 (s -~) 3'02 23.62 G (S"1) 7.3 64.37 N 1.01 0.94 W 4.29 1.4 Kd (s 1) 7.36 60.52 m 1.99 2.06 Error parameters: SD 2.35.10 2 6.95-10 .2 FFISHER 488.47 32.21 qbEXNER 0.110 0.381 Using the kinetic rate constants (ko) obtained with the three lumps kinetic model we have determined the kinetic rate constants for the formation of the primary sulfur products present in the gasoline using equation (9) where the initial selectivity is calculated from the slope at the origin of the OPE curve and ki is the kinetic rate constant for the formation of a particular primary product. ( I n i t i a l . Selectivity), = k, / k o

(9)

The kinetic rate constants have been compared with those for formation of gases and gasoline, taking into account the initial concentration of sulfur in the feed. In the case of the sulfur containing compounds, we have considered that the feed has 1.4 wt% of sulfur, and therefore we have multiplied their kinetic rate constants calculated from (9) by 98.6/1.4. The kinetic rate constants obtained in this way are given in Table 4. Table 4. Kinetic rate constants for prim~try products formed during catl lytic cracking of gas oil. Product Selelctivity Kinetic rate constant Ki Initial Selectivity T=783 K T=873 K Character a T=783 K T=873 K 2.72 18.66 Gasoline PU 0.90 0.79 0.24 4.72 Gases P+S 0.08 0.20 Mercaptans PU 4.6.10 .3 0.5.10 -3 15.0.10 .3 14.2.10 .3 Thiophene P+S 0.03-10 .3 0.06.10 "3 0.01.10 .3 0.003" 10.3 C 1-thiophenes PU 2.9.10 "3 2.6-10 .3 0.9.10 .3 0.1.10 .3 BT + ABT b PU+S 79.2.10 .3 53 7.9.10 .3 24.6.10 .3 21.3.10 .3 a P=primary product, S=secondary product, P U=primary unstable product b Benzothiophene (BT) and alkylbenzothiophene (ABT).

163 It can be seen there that the highest rates of formation for primary sulfur products are those of the benzothiophene and alkyl-benzothiophenes, followed by mercaptans. The one corresponding to thiophene appears to be the lowest. When the results are calculated at 873 K it can be seen that despite the large increase in the gases selectivity, the selectivity to the primary sulfur compounds is similar to that obtained at 783 K or even lower. Nevertheless, the kinetic rate constants increase for thiophene and benzothiophene derivatives, remain constant for methylthiophenes and decrease tbr mercaptans. It should be pointed out that the kinetic rate constant value for mercaptan formation is strongly affected by the fast cracking of these compounds that depletes its concentration even at low levels of conversion. Most remarkable is the high increase in the rate of formation of benzothiophenes indicating a very high activation energy for this reaction.

3.2-Reactivity of pure sulfur compounds under MAT conditions. 2-Methylthiophene (2-MT), thiophene (TIO), tetrahydrothiophene (THT) and benzothiophene (BT) were fed as pure compounds to the MAT unit, at 510~ during 30 s TOS and at catalyst to oil ratios in the range of 3-6 g.g-1. The products obtained when 2-MT was cracked on the commercial FCC catalyst were total gases (C1-C4 gases plus H2S), liquid products (C5+ fraction) and coke. The C5+ fraction was the most abundant, and within the C5+ products it was seen that thiophene, which was produced via by dealkylation, was the main one, followed by 3-MT, which was formed by methyl shift. C1-C4 gases and H2S were also tbrmed by hydrogenation of the alkyl-thiophene followed by cracking. Finally, large amounts of coke were also observed (Figure 5). This coke has a C/S ratio o f - 1 0 mol.mo1-1, indicating that the olefins formed during cracking also contribute to coke formation.

Fig. 5. Yields to the main products obtained by cracking of pure 2 methylthiophene in a MAT unit. T=783 K, TOS=30 s. (R) Gases; (o) C5+; ( A )Coke; ( ~ )H2S.

Fig. 6. Cracking of different pure sulfur compounds on a high unit cell size USY zeolite (2.448 nm). T=783 K, TOS=30 s.

Besides the commercial FCC catalyst, three USY zeolites with different unit cell sizes (Table 2) were used as catalysts for cracking pure sulfur compounds. When 2-MT, TIO and THT were cracked on a high unit cell size USY zeolite, USY-1, THT was seen to be the most reactive compound, giving mainly gases (figure 6). Within these total gases H2S and C4

164 olefins were the products formed in a larger extent. Conversion of 2-MT was also very high, but it gives a large amount of coke. Finally thiophene reacts in a lower extent, giving very high yields of coke. In Table 5 it can be seen that thiophene reacts to a larger extent on high unit cell size USY zeolites (USY-1 and USY-2) giving mainly coke, but when the catalyst is a USY zeolite with low unit cell size (USY-3), it converts less with lower selectivity to gases and coke and higher selectivity to products (including sulfur compounds) in the C5+ fraction. Thus, in order to eliminate sulfur from the gasoline range it is important to have a catalyst able to promote hydrogen transfer reactions. However, much of this sulfur will end up as coke if coke is the only hydrogen source. Finally benzothiophene was also cracked under the same experimental conditions. In this case conversion was very small, and the only products observed were coke and some sulfur containing compounds in the light diesel range such as alkylated benzothiophenes. Table 5. Cracking of pure thiophene in a MAT unit on USY zeolites with different unit cell sizes (U.C.S.). T=783 K, TOS=30 s. Catalyst USY- 1 USY-2 USY-3 Thiophene Conversion (wt%) 62.12 58.31 27.52 Yields (wt%): Total gases 16.62 16.89 5.25 C5+ 5.37 0.91 21.12 Coke 40.13 40.52 2.15 H2S 14.00 13.79 5.16 3.3-Reaction of H2S with 1-butene. With the results obtained up to now it has been possible to confrm the reactivity of thiophene, alkylated thiophenes and benzothiophene derivatives. In order to complete the reaction network it is necessary to determine the cause of the instability of H2S and to confirm if this compound can react with other products present in the reaction media, such as olefins, giving mercaptans, and maybe cyclizing and dehydrogenating in a second step to give thiophenic compounds. To do this, H2S was reacted with 1-butene at different olefin/H2S ratios, at the reaction temperature used for the previous MAT experiments, 783 K. From the results obtained it appears that at the high reaction temperatures used for cracking, the reaction between H2S and olefins occurs to a minor extent, since thermodynamically the reverse reaction should be favored. In fact, the amount of sulfur containing products was so low that quantification was not possible. When the reaction temperature was decreased, conversion increased, but it was still very low. This is not surprising if one takes into account that the reaction between H2S and olefins is favored at low reaction temperatures. In this way, the Elf-Atochem process reacts iso-olefins and olefins with H2S at 60-100~ and 10-15 bar [9], conditions that are very different from those used in catalytic cracking, but that should be taken into account in the downstream processes. 4.- Conclusions. In conclusion, we can establish the following reaction network for the sulfur products present in gasoline:

165

SulphurCompounds inthefeed[ ~

[Mercaptans[-

I

A

I

/

k

y

l

-

(D

T

t

~ ~-~ I "

I--*t +

B

~

I

(D +

I

1

~" Coke 4-

Moreover, it should be considered that the nature of the catalyst (zeolite unit cell size) has an impact, not only on sulfur removal via hydrogen transfer-cracking, but also on coke formation from sulfur containing reactants. The study presented in this paper shows that, in order to increase the sulfur reduction by converting gasoline sulfur containing compounds to hydrogen sulfide and a hydrocarbon, hydrogen transfer reactions have to be considered, since partial or total saturation of these sulfur compounds is a necessary step before cracking. The extent of secondary reactions, such as hydrogen transfer, will depend not only on the catalyst used, but also on the operation conditions and on the reactor design. One of the major drawbacks associated to the MAT unit is the long contact times used as compared to commercial operation. Thus, one could expect that hydrogen transfer and sulfur removal are overestimated when performing the catalytic test in a fixed bed MAT unit. In fact, when reactions are carried out in a Micro Invers-Riser (MIR) unit, at lower contact times of 500 to 800 ms, hydrogen transfer is lower and gasoline sulfur content is considerably higher. The sulfur removal capacity of a Zn impregnated alumina was seen to decrease from a 30 wt% in the case of the MAT to a 10 wt% in the case of the MIR unit.

REFERENCES. 1. Cheng, W.C., Kim, G., Peters, A.W., Zhao, X, Rajagopalan, K, Catal. Rev.-Sci.Eng., 40(1&2), 39 (1998). 2. Gatte, R.R., Harding, R.H., Albro, T.G., Wormsbecher, R.F., Fuel Chemistry Symposia, ACS NationalMeeting, San Francisco, CA, April 1992. 3. Harding, R.H., Gatte, R.R., Whitecavage, J.A., Wormsbecher R.F., A CS Symposium Series, 552, 286 (1994). 4. Alkemade, U., Dougan, T.J., Catalysts in Petroleum Refining and Petrochemical Industries 1995, M. Absi-Halabi et al. (Editors), 1996. 5. Wojciechowski, B.W, Corma, A., Catalytic Cracking. Catalysts, Chemistry and Kinetics, Chemical Industries 25, Marcel Dekker Inc., 1986. 6. Kemp, R.R.D., Wojciechowski, B.W, Ind. Eng. Chem. Fundam., 13, 332 (1974). 7. Himmelblau, D.M., Process analysis by statistical methods., J.Wiles & Sons Inc., New York, 1977. 8. Snedecor, G.W, Cochran, W.G., StatisticalMethods., Iowa State Univ. Press, Iowa, 1964. 9. E Labat. btformation Chimie, 367, 115 (1995).

Studies in Surface Science and Catalysis 134 M.L. Occelli and P. O'Conner (Editors) 9 2001 Elsevier Science B.V. All rights reserved

167

Development of a Kinetic Model for FCC valid from Ultra-Short Residence Times M.A. den Hollander*, M. Makkee*, and J.A. Moulijn Industrial Catalysis, Department of Chemical Technology, Faculty of Applied Sciences, Delft University of Technology, Julianalaan 136, 2628 BL Delft, The Netherlands Experimental data from Hydrowax cracking using the microriser equipment have been used to develop a kinetic model that adequately describes the experimental data from 0.05 to 5 s residence time. To account for different time scales of the reactions, different activity functions were used for coke formation and for conversion to the other products. The activity for coke formation was described by an exponentially decreasing function of the residence time with a very short characteristic time of 0.02 s. The activity for formation of the other products was also described by an exponentially decreasing function, in this case of the coke content of the catalyst. The model was used to predict the residual activity of the catalyst in a second experiment with flesh Hydrowax feedstock. Using the assumption that in the second run no further deactivation of the catalyst took place, the model predictions agreed well with the experimental data. The applicability of the model to data from other feedstocks, a hydrotreated flashed distillate and a vacuum gas oil, was successfully demonstrated. Keywords: riser reactor, modeling, deactivation, coke, residual activity 1. INTRODUCTION An industrial FCC unit is operated in a heat-balanced mode; in the regenerator coke is burnt from the catalyst thereby providing the heat that is needed for the evaporation and cracking of the feed in the reactor. Therefore, an important operating parameter is the amount of coke deposited on the catalyst in the reactor. Generally, coke formation in FCC can take place in two different time scales: firstly, deposition of coke upon the first contact of the catalyst with a feedstock that contains coke precursors (thermal coke, soaking coke from liquid droplets, extremely reactive hydrocarbons, etc.), and secondly, the catalytic formation of coke as a side-product of cracking due to hydrogen transfer, condensation, and dehydrogenation reactions [1,2]. To be able to discriminate between these two coking regimes, the experimental set-up should have the ability to discriminate between reactions on a millisecond time scale (deposition of coke) and reactions on a longer time scale (cracking reactions with coke as a side product). t Current address: Shell Global Solutions International, P.O.Box 38000, 1030 BN, Amsterdam, The Netherlands Corresponding author; E-mail: [email protected] ,

168

Figure 1. Outline of the microriser equipment. 1: Normal feedstock injection point. 2: Injection point ultra-short reactor. The microriser is a laboratory-scale entrained flow reactor, in which the reaction time can be varied by changing the reactor length [3-5]. The outline of the equipment is shown in Figure 1. The microriser reactor is placed inside a fumace and can be operated completely isothermally and with ideal plug flow for gases and catalyst. Nitrogen is used as a transporting gas for the catalyst from the catalyst feeder through a preheat oven to the reactor entrance. The reactor entrance is defined at the point where preheated feed is injected. Heating times are sufficient to ensure isothermal contacting of catalyst and feedstock. At the end of the reactor the catalyst is separated from the reaction products by a high loaded cyclone (entrance separation efficiency >99.99%). Stripping of the catalyst takes place post-run and it has been verified that the stripping procedure does not influence the (catalytic) conversion and product yields. Only experiments with a mass recovery of 97+2 wt% are considered successful (due to liquid hold-up in the product collection system, mass recoveries of 100% are impossible). Using the microriser, cracking experiments have been performed at different residence times. The shortest residence time measured was 0.05 s. It was found that catalytic cracking of a Hydrowax feedstock could be represented by two processes with different time scales. The first process, with a time scale <0.05 s, was characterized by a relatively high rate of conversion and coke deposition on the catalyst. On the second time scale (0.05-4.5 s) no significant additional coke formation was found. However, formation of the other products, LCO, gasoline, and gas, continued. The reactions occurring on the second time scale could be described with a four-lump kinetic model using first order reaction kinetics and a constant catalytic activity. This four-lump kinetic model was based on the five-lump model of Corella and Franc6s without the lump representing coke [3,6]. In this paper the kinetic model will be further developed to describe the reactions starting at zero seconds residence time. Coke formation will be included in the model and the different

169

time scales will be taken into account by using different activity functions; one for the formation of coke and one for the formation of the other products. The data from the Hydrowax cracking experiments will be used in the development of the model. Finally, the application of the model to other feedstocks, a hydrotreated flashed distillate (HTFD) and a vacuum gas oil (VGO), will be demonstrated. 2. THEORY 2.1 Model assumptions When deriving the model the following assumptions were made: 9 Reactions are all catalytic, i.e., thermal conversion is not taken into account; 9 The five-lump model, shown in Figure 2 and Table 1 describes the reaction network; 9 All lumps react according to first order kinetics [3]; 9 Coke is only produced from HCO and LCO and is the only product in these reactions, i.e., no formation of gaseous products is directly correlated with the formation of coke. The formation of coke from gasoline is not considered because it was found that coke is mainly formed initially, at the first contact catalyst/feedstock contact and the feedstock does not contain any gasoline-range hydrocarbons; 9 The rate of coke formation and the rate of formation of the other products are described by different activity functions, ~coke and (I)conv,respectively. As a result, the reaction rates of the different lumps (HCO, LCO, gasoline, gas, and coke) are represented by the following equations.

rHCO = ~conv "(kl +k2 +k3)'YHCO +~coke "k7 "YHCO

(1)

rLCO = ~ conv" (- k 1" Y HCO + k4. Y LCO + k 5" Y LCO ) + (I) coke" k 8 9YLCO

(2)

rgln = (I) conv" (- k2 "YHCO - k5 "YLCO + k6 "Y gin)

(3)

rgas = ~ c o n v ' ( - k 3 " Y H c o - k 4 " Y L c o - k 6 " Y g l n )

(4)

Coke, Y~oke HCO, YHCO

'

~ LCO, YLCO

Table 1 Definition of lumps HCO LCO

gasoline, Ygln Figure 2.

6

~ gas, Y0as

Five-lump model [6].

gasoline gas coke

composition/boiling point range b.p. >643 K b.p. 494-643 K C5 - b.p. 494 K

H2, C1-C4 carbon deposit, CxH•

170

rcoke = @coke "(-k7 "YHCO - k 8 " YLCO)

(5)

2.2 Activity function for coke formation Experimentally, coke formation was only found to take place on a time scale of milliseconds. This can be represented by an activity function that decreases exponentially with time"

coke : e-C~~

(6)

For high values of the activity parameter ot Equation 6 describes a strong decay of the activity for coke formation within a milliseconds time scale.

2.3 Activity for conversion to all products except coke Since coke is the actual deactivating species in the cracking reactions, it is preferred to use an activity function for conversion that is a function of the coke content of the catalyst. However, it is difficult to measure the local coke content of the catalyst. On-line measurement under realistic conditions of the coke content is not possible. Off-line measurement is only possible after a stripping procedure. The stripping procedure, however, can influence the amount and type of coke deposited on the catalyst [7-11], so it is not likely that the coke measured after stripping is physically the same as it was before stripping. Moreover, coke is a general term referring to carbonaceous deposits that can have different origins of formation, different compositions, and different effects on the catalyst activity[2,12,13]. For the model development described in this paper, it was assumed that the coke measured after stripping is the same as the coke responsible for the deactivation of the catalyst in the reactor. To describe the activity for conversion, different functions of the coke content of the catalyst from literature will be evaluated. The first type of activity function is based on the general equation:

d(@c~ d cc

= _kd. (~

conv

(Cc)) mcc

(7)

The physical background of this (semi-empirical) activity function is the decrease of activity due to the deposition of coke, where the order of deactivation, mcc, represents the number of active sites involved in the controlling step of the deactivating reaction[ 14]. Other assumptions involving this activity function are the existence of a homogeneous, or averaged, catalytic activity and deactivation by coke [ 15,16]. Depending on the value of the deactivation order m~, the following functions for the activity as function of catalyst coke content can be derived. For mcc=l (a number of authors use this equation; for example [17-19]):

@conv,l(Cc) = e-kd,l"Cc

(8)

171 For mcc=2 [ 19]: 1 conv,2 (Cc) = 1 + kd,2 "Cc

(9)

For mcc=3: 1 conv,3 (Cc) = .j1 + 2 .k d,3 9c c

(10)

A second type of activity function was used by Bernard et al. [20,21 ]. This semi-empirical function describes the decrease in activity resulting from both coverage of active sites and pore blockage:

kdA +1

e(kd,4.Cc)

@conv,4 (Co) :

(l 1)

kdA + The deactivation c o n s t a n t kd, 4 represents the sum of the contributions of site coverage and pore blockage to the total deactivation, whereas kdArepresents the decrease in activity due to pore blockage relative to the deactivation by site coverage. (Note that whenkda<
172 Table 2 Feedstock properties Hydrowax density (288 K) (kg-m3) 850 n.m. ~a) viscosity (373 K) (mPa.s) 0.09 Conradson Carbon Residue (wt%) 12 basic nitrogen (ppm) n.m. ~a) total nitrogen (ppm) 0.01 sulfur (wt%) __(b) vanadium (ppm) __(b) nickel (ppm) Boiling point composition (c): initial boiling point (K) 607 10 wt% (K) 649 30 wt% (K) 679 50 wt% (K) 704 70 wt% (K) 734 90 wt% (K) 784 final boiling point (K) 823 a) n.m.: not measured b) below detection limits (< 0.5 ppmwt) c) measured by simulated distillation according to ASTM

HTFD 891 5.2 0.10 22 337 0.09 O. 1 0.1

VGO 850 5.77 0.18 198 960 0.08 __(b) __(b)

522 619 664 696 730 773 872

488 609 666 706 744 798 852

D-2887

Table 3 Properties of the equilibrium catalysts used with the different feedstocks feedstock used apparent bulk density (kg.m 3) average particle size (~tm) surface area, matrix (m2.g-1) surface area, total (m2.g-1) A1203 (wt%) total RE content ~a~ (wt%) Na (wt%) Fe (wt%) Cu (ppm) Ni (ppm) V (ppm) Sb (ppm) a) total content of rare earth oxides

Hydrowax 820 70 n.m. 152 30.0 1.23 0.61 0.72 39 3800 550 2 103

HTFD 940 81 44 167 39.6 1.7 0.2 0.4 -1190 740 --

VGO 870 69 n.m. 150 33.8 0.61 0.22 0.62 15 90 108 31

determined in one fitting procedure the solution obtained would be highly dependent on the initial guesses of the parameters. Therefore, a special procedure has been followed to obtain realistic starting values for the parameters.

173 Table 4 Reaction conditions for the different experimental series feedstock used.. . rreactor (K) Tstrip (K) rcat preheat (K)

rfeed preheat

(K)

Preactor

(105 Pa)

~2,reactor

(mg's -1) (mg.s -1) (mg's 1)

(~lqq2,strip (Dfeed

"~ (s) CTO (kgcat'kgfeed"1)

Hydrowax 798 798 798 673 1.1 19 9.5 65 0.05-4.5 0-8

HTFD 798 798 798 673 1.1 19 9.5 65 0.05-4.8 0-7

VGO 788 788 788 673 1.1 19 9.5 65 0.3-5.6 2-5

The overall conversion of HCO (to coke and one overall lump of the rest of the products) together with different activity functions was used to determine the best activity function to describe the experimental data. The parameters determined in this way were:kHco, (the sum of kl, k2, and k3), k7, a, and kd for each of the Equations 8-11. The significance of the determined parameters and the sum of squared residuals (SSres) were used to discriminate between the different models and to select the most adequate activity function. After selection of the most adequate activity function and using the calculated rate constants ~Hco, ct, and kd) as fixed (in-put) parameters, realistic estimations for the other reaction rate constants were determined. Subsequently, the best fit of the experimental data was obtained using the previously calculated estimations of the parameters as starting values. In the following sections, only the results from the first step, selection of the activity function for conversion, and the last step, the best description of the experimental data by fitting all parameters, will be discussed. One of the objectives of this model was to adequately describe the relatively high conversion of HCO initially found. Therefore, in determining the best activity function, more weight was given to an adequate description of the HCO fraction at short residence time. Usually, to obtain an adequate description of the coke production, the relative error of the coke fraction was minimized instead of the absolute error used for the other products. 4. RESULTS 4.1 Selection of activity function for conversion The activity functions (I)conv,l_4(Cc) given in Equation 8-11 have been used to calculate the parameters shown in Table 5. The sum of squared residuals, SSres, was calculated for the HCO fraction. To prevent the fitting procedure from finding a local minimum, fitting was using a broad range of starting values for the parameters. In Table 5 it can be seen that most of the parameters determined are not significant, since their confidence interval contains the value zero. Only the parameters resulting from the model with activity function ~conv,l(cc) (Equation 8, activity as a first order decreasing function of the coke content of the catalyst) have a significant value. From the SSres it follows that activity function ~conv,2(cc) (Equation 9) describes the experimental data equally well. It has been verified, by fitting all parameters, that this model, using~conv,2(cc), was less adequate

174 Table 5 Parameters, with 95% confidence interval, determined using different activity functions for conversion as function of coke content of the catalyst

kHco k7,HCO-coke kd kda ot SSresHco

9 conv,l(cc) eq.8 4+2 0.3 + 0.1 6.0 + 0.9

(I)conv,2(Cc)eq.9 5+6 0.10 + 0.03 57 + 63

~co.~v,3(cc)eq.10 2 + 24 0.4 + 0.4 76 + 2" 103

48 + 19 0.10

15 + 4 0.09

65 + 62 0.23

(~)conv,4(Cc)eq.11 4+5 0.3 + 0.1 9+4 4 + 16 48 + 20 0.13

than the model using l~)conv,l(Cc). The models using (I)conv,3(Cc) and (I)conv,4(Cc) not only resulted in a poorer fit of the experimental data (higher SSres), but also in a number of parameters with insignificant values.

4.2 Fitting all product yields Estimations of the rate constants for the products LCO, gasoline, and gas were obtained by keeping the parameters given in Table 5 constant in the fitting procedure. As discussed above, only the first order activity function for conversion as function of coke content of the catalyst, Table 6 Reaction rate parameters, with 95% confidence interval, resulting from fitting of the model to experimental data obtained with different feedstocks. kl,HCO-LCO (g'gcat'l"S"1) kz,HCO-gln (g'gcat'l's "1) k3,HCO-gas (g'gcat-l"s1) kHco(a) (g'gcat'l'S "l)

k4,LCO-gas

(g'gcat'l's "1)

ks,LCO-gln

(g'gcat'l's "l) k6,gln-gas (g'gcat'~'s "l) k7,HCO-coke (g'gcat-l"S1) ks&cO-coke (g'gcat'l's "l) kd (wt% "l) (s "l) S Sres/(n-p) SSres PE LF

kl/kHco k2/kHco k3/kHco a) kHCO- kl+k2+k3

Hydrowax 4.1 + 0.5 9.2 + 0.3 1.6 + 0.2 14.9 + 0.6

HTFD 1.9 + 0.6 7.5 + 0.5 1.5 + 0.4 10.9 + 0.9

VGO 1.2 + 0.3 1.3 + 0.3 0.31 + 0.05 2.8 + 0.4 ,,

0 (b)

0 (b)

0 (b)

7.1 + 1.3 0.5 + 0.2 0.16 + 0.04 1.5 + 0.5 10 + 3 62 + 3 2.7.10 .4 0.075 0.044 0.031 0.3 0.6 0.1

1.0 _+0.7 0.3 + 0.3 0.21 + 0.05 0.5 + 0.1 8.2 + 0.2 46 + 5 7.3" 10 -4 0.047 0.013 0.035 0.2 0.7 0.1

0.9 _+ 0.6

b) fitting result <10 -6, therefore set to zero

0 (b) 0.059 + 0.006 0.02 + 0.02 11.9 + 0.3 11 + 1 8.9" 10 -5 0.0089 0.022 <0 0.4 0.5 0.1

175 given by Equation 8, was used. With these estimations as starting values, all parameters were fitted to find the best description of the experimental data. The results are shown inTable 6. This procedure resulted in a model that adequately describes the experimental data. The values of the parameters for HCO conversion and the activity functions given inTable 6 are all larger than determined before (Table 5), giving a better description of the high conversion obtained within the first second of the residence time. All parameters, exceptk4,LCO_gas,have a significant value. (Typical fitting results for k4 were 10-6-109). The description of the conversion and product fractions by this model is shown in Figure 3 and Figure 4. In this work, conversion is defined as the amount of HCO-range hydrocarbons converted according to Equation 12: Conversion =

YHCO,f - YHCO

9100 wt%

(12)

YHCO,f

The conversion, Figure 3, is well described over the complete residence time range and for all CTOs (catalyst-to-oil ratios). This is also valid for the description of the LCO and gasoline fractions shown in Figure 4; although the maximum in the LCO fraction is not described perfectly, the trends are most satisfactory. The description of the gas and coke fractions is somewhat poorer. The model describes well the trend of the observed coke fraction as function of residence time. From the value of the deactivation constant ot the characteristic time for coke formation can be determined from otl and is equal to 0.02 s. The absolute value of the coke yield calculated by the model, however, is not described perfectly. At low CTOs the calculated coke fraction is too low; at the highest CTO the model calculates a too high coke fraction.

1O0

C

~

} 60 c

._ > 40

c O O

20 0,,

0

I

I

1

2

-

I

I

3

4

residence time [s] Figure 3. Description of Hydrowax conversion by the most adequate model. Parameters from Table 6. D: CTO=2; A: CTO=4; o: CTO=6.

176 0.2

0.8

LCO

I

o

[]

gasoline

CTO=2

,,.~0.15

C T O = ~

v--

=

,~ 0.6 .~.

t-"

,~ o.4 L..

o

0

" 0.05

~, o.2

0 0

I 1

0.2

I I 2 3 4 residence time [s]

0 5

J i

0

gas

1

I i

I i

i i

2 3 4 residence time [s]

0.05

5

coke

0.04

0.15

,...-, "7

'0)

e~

CTO=6

0.03 - t -

I,-1,oo o - - - o

0.1

0.02-~A 8

~CTO=2

0.05

t-~o

-o

9

9

9

9

[]

[]

o

o

0.010 0, 0

I 1

,. I I I 2 3 4 residence time [s]

5

0

1

2 3 4 residence time [s]

5

Figure 4. Description of the products from Hydrowax cracking by the most adequate model. Curves were calculated with the parameters given in Table 6.

Figure 4 shows that the calculated gas fraction is too low for the shortest residence time. However, the variation in the experimental values of the gas fraction is relatively high compared to that of the other products, especially at residence times shorter than one second. 5. DISCUSSION The activity for coke formation was described by a function of the residence time, so it would seem likely to model the activity for conversion also by a (similar) function of the residence time, although this neglects the mechanism of deactivation by coke formation. It was found that, in this case, no adequate description of the experimental data could be

177 obtained. Clearly, the different time scales of the processes of coke formation and of formation of the other products, require the approach described in this paper. A more detailed discussion of the model development, including activity functions for conversion as function of residence time, will be given in [22].

5.1 Model assumptions The use of different activity functions for coke formation and the conversion to other products in the five-lump kinetic model with first order reaction rates resulted in a model that adequately describes the conversion of HCO and formation of LCO, gasoline, and coke. The formation of gas is reasonably described except for the high initial gas production experimentally found. This is probably more a result of inaccuracy of the experimental data because the ultra-short reactor used for these experiments is likely to behave less ideal than the normal reactor. In the ultra-short reactor the oil is injected in the horizontally flowing stream of catalyst and nitrogen just before the entrance of the cyclone, while in the normal reactor injection takes place into the fully developed down-flow regime. Due to the non-ideal behavior of the ultra-short reactor more back-mixing and entrainment of hydrocarbons to the stripper can take place, resulting in a higher gas production. The conversion of the feedstock by thermal cracking reactions, measured in experiments without catalyst, was at most 6 wt%. This was measured at the longest reactor length, 21.2 m. In catalytic cracking experiments the total conversion at this reactor length was 76-97 wt%, depending on the CTO. So, in the worst case, the contribution of thermal reactions to the total conversion is 8%. Due to the assumption that no thermal cracking takes place, the reaction rate constants determined will be somewhat too high (in the worst case 8%). Because the gas lump contains the fuel gas products (H2, CH4, C2H4,and C2H6),that partly result from thermal reactions, the rate constants related with the formation of gas, k34, are most influenced by this assumption. The assumption that coke formation is not accompanied by the formation of other products is not completely valid; due to stripping a transformation of coke deposited on the catalyst takes place. The coke is made more aromatic by the removal of small, hydrogen-rich molecules during stripping [7-9,22,23]. Separation of the gas lump into several lumps according to their mode of formation would give a more adequate description of the chemical origin, but this would also make the model more complex. Moreover, for the objectives of this work, describing the different time scales of coke formation and conversion, the five-lump model is satisfactory. The activity function for conversion used in the model, first order decrease as function of the coke content of the catalyst, represents deactivation due to covering of active sites. The deactivation order met equal to one should be interpreted as one site being involved in one deactivating event, or single site coverage by coke. The activity function used is by approximation (Taylor's series) equal to:

@conv, l(Cc) = e (-kd'l"cc) = 1 - kd, 1 .c c =

k-1 d,1 - C c k-1 d,1

(13)

178 This activity function can be interpreted as follows: the activity is equal to the fraction of total amount of active sites present that is not covered by coke, with kd,1-1 representing the total amount of active sites (kd,l-l=0.1 wt% in this case) [ 14,24]. In earlier work it has been concluded, from the high rate of (initial) coke formation in combination with the Wheeler-Weisz criterion for diffusion limitation, that the formation of coke was highly diffusion limited and that, therefore, egg-shell deposition of the coke was likely to take place [3]. Typically, at relatively high concentrations of coke this could lead to severe pore blockage or deactivation according to the so-called percolation model. Nevertheless, in this work, the activity of the catalyst could be described with the activity function representing (single) site coverage. Moreover, the activity function that accounted for both site coverage and pore blockage, ~conv,4(Cc), did not adequately describe the observed conversion. The explanation probably is that the coke deposition levels are low, lower than the so-called percolation threshold [25-27]. This implies that pore blocking is not significant under these conditions. 5.2 Prediction of the residual activity of the catalyst in a second experiment The model has been used to predict the performance of the coked (and stripped) catalyst in a second experiment with fresh feedstock. The initial activity of the catalyst was supposed to be equal to the final activity in the first run, calculated by the model. To account for possible coke formation and additional deactivation of the catalyst in a second run, two extreme cases for the coking and deactivation reactions in the second run have been evaluated:

A:

no additional coke formation or catalyst deactivation in the second run; the activity of the catalyst for conversion will be constant throughout the reactor;

B:

additional coke formation and catalyst deactivation according to similar kinetics as with the regenerated catalyst.

The schematic representation of the changes in catalyst activity for conversion of these two cases is illustrated in Figure 5. A comparison between the model calculations according to the cases A and B and the experimental results ('c=4.8 s) is shown in Figure 6a and Figure 6b, respectively. For case B, where additional deactivation by coke is considered, Figure 6b shows that the predicted residual activity is much too low; no significant yields of gasoline, LCO, or gas are predicted. Only the additional production of coke is predicted accurately by model B; the coked catalyst produces similar amounts of coke as the regenerated catalyst. It is, therefore, surprising that the model according to case A, without additional deactivation of the coked catalyst, describes the experimental data better (Figure 6a) than the model of case B, with additional deactivation. Apparently, the additional coke deposited on the coked catalyst does not further decrease the catalytic activity for conversion of the feedstock. A simple explanation would be that the additional coke was deposited solely on the coke already present on the catalyst. A coke content of 0.5-1 wt% approximately agrees with only 1-2 vol% of the catalyst porosity, but since deposition takes place in an egg-shell profile [3], a relatively large amount of the additional coke will indeed be deposited on the coke that was already present. The total amount of coke is not enough to reach the percolation threshold and apparently, it does not

179 s e c o n d run (coked catalyst)

first run ( r e g e n e r a t e d catalyst) c~=0 kHcd(~conv=14.9

I I I I I I

1 I I I I I I

o;

II I

cc=0.56

k,,

kHcddi~onv=O.06

! I I I I

I I

0 0.05

5

A)Co=0.56 kHco (I)conv=O.06

I I

,,, -

,

,,

B,Cc=l.1 )kilo6 (I)c~

LI

I i

0" 2,: 10-3

I

0 0.05

5

Figure 5. Schematic representation to describe the catalyst activity profile; in the first run according to the developed model (left) and in the second run (right) according to two different models: A) constant activity, so no additional deactivation; B) additional deactivation similar to that of the regenerated catalyst. cause severe deactivation by pore blockage explaining that the catalyst activity is hardly affected by the second amount of coke. 5.3 Application of the model to the results from other feedstocks

The applicability of the model to the results from other feedstocks will be demonstrated using data from a hydrotreated flashed distillate, HTFD, and a vacuum gas oil, VGO (properties in Table 2). The resulting parameters are given in Table 6. The parameters almost all have significant values, and from the low values of the normalized SSres it is also clear that the model has been fitted adequately to the experimental

1 00

100

--

(HCO) 80 -~

80HCO

" 60 ]-

6O line

"o

"o m ._e 40

9~, 4 0 -

20

,4~ n ~_.------=as..~,_T_ ~ '

0 0

a)

20 --

LCO 9

,

- i

i

2

4

CTO [kgcat'kgfeed "1]

9

zx (LCO)

coke

0 0

6

b)

[] ~-; 2

[] "

, 4 CTO [kgcat'kgfeed -1]

[].. f c o k e "

Figure 6. Performance of the coked catalyst in Hydrowax cracking predicted by different models, a) (adequate) prediction by model A; b)(unsatisfactory) prediction by model B. Symbols represent experimental data; x: HCO; I . LCO; A. gasoline; []" gas; *" coke.

180 data of the different feedstocks. The 'lack of fit', LF, has been obtained from the difference between the SSres and the pure error, PE, representing the accuracy of the experimental data. The description of the results from the HTFD feedstock has a lack of fit comparable to that of the Hydrowax results. For the results from the VGO feedstock, the lack of fit is negative, i.e., the deviation between the model curves and the experimental data is smaller than the experimental error. In the experiments different feedstocks and different catalysts were used (properties in Table 3). Therefore, the values of the reaction rate parameters are not intended to be used to draw conclusions about the relative reactivity of the feedstock or the activity of the catalyst as such. The conclusion that can be drawn from the results is that the model derived in this work is not limited to the description of Hydrowax cracking alone, but is generally applicable. Still, from the values of the deactivation constant, kd, used in the activity functions for conversion, it can be seen that in all cases the deactivation of the catalyst by deposited coke is similar or, according to the analysis of Equation 12, that for all catalysts the (initial) amount of active sites is comparable. The significantly lower value of the deactivation constant for coke formation, a, of the VGO system suggests that in this case coke formation takes place on a longer time scale, not only at the initial contact of feedstock and catalyst. The characteristic times for coke formation determined from ~-1 are 0.02 for the Hydrowax and HTFD, and 0.09 for the data for the VGO. However, this is merely the result of a lack of data obtained at residence times shorter than 0.3 s; for this feedstock no ultra-short reactor experiments have been performed. So, also for the VGO system the majority of the coke formation was observed to have taken place within the first part of the reactor. This illustrates the importance of collection of experimental data at a relevant time scale, although it cannot be certified that the time scale of the ultra-short reactor experiments has been short enough.

100

1~176 1

01/fcTo_4

CTO=5

~ 60 ' tO m L

>~ 40

tO

o

CTO=4 20

2O

"'-

0'.

0 0

Figure 7.

1

2 3 4 residence time [s]

5

Description of HTFD conversion.

0 " C T 0 = 2 ; A" CTO=4; o" CTO=6.

0

Figure 8.

CTO=2

i 2 4 residence time [s]

6

Description of VGO conversion.

D" CTO=2; A" CTO=4; x: CTO=5.

181 In all reaction systems, gasoline was formed both primary from HCO and secondary vi~ LCO. From comparison of the rate constants k2 and ks, it can be seen that, intrinsically, in the Hydrowax and VGO system the two routes are similar, while in the HTFD system the primary production of gasoline is the dominant route. The description of the conversion of the HTFD and VGO is illustrated in Figure 7 and Figure 8, respectively. These graphs confirm that the model can be used to adequately describe the experimental data from the different feedstocks. The description of the product fractions LCO, gasoline, gas, and coke is shown inFigure 9 for the HTFD data and in Figure 10 for the VGO data. Clearly, not only the conversion of

0.3

0.7 9

CTO=4

gasoline

[]

0.6 "7 "7

a~ 0.5

:TO=2

.~ 0.2 O3

C TO

=6 0.4

cO ,m O L.

~ 0.:3

oo 0 . 1

~ 0.2

..J

g

0.1

LCO

0 0

1

0.3

2 3 4 residence time [s]

0

1

2 3 4 residence time [s]

0.05

gas

5

coke

0.04 "7 ~

"7 {33

0.2

"~0.03

CTO=6

o

0

o

0

p~ 0.1 t

~-0.02 (D

o~A_.~d~o=4

A--

0.01 0

0 0

1

2 3 4 residence time [s]

"--

CTO=4

~"

0 0

5

CTO=2

u

0

k 1

I I I 2 3 4 residence time [s]

5

Figure 9. Description of the products from HTFD cracking. Curves were calculated with the parameters given in Table 6.

182 HCO (Figure 7 and Figure 8), but also the different product yields are adequately described by the models. Regarding the products from HTFD cracking shown in Figure 9, only a significant deviation between the calculated and the experimental data of the gas fraction is observed. Initially, the calculated gas fractions are too low and at the longest residence time the calculated gas fractions are too high. This is similar to the results that were observed for the Hydrowax feedstock; experimentally a higher gas yield was found at the lowest residence times, ascribed to the non-ideal behavior of the ultra-short reactor. The products from VGO cracking shown in Figure 10 are all equally well described by the model. As was also concluded from the 'negative' value of the lack of fit by the model, typically the deviation between the experimental and modeled results is insignificant. 0.35

0.4

CTO=5

,~:~ 0.3

v'--' a~0.25

ET)

CTO=2

c

gasoline

l'-"

0.21

0 ".~

. 0g

~ 0.2

~ ~0.15 9 0. 0.1

l-O

~0.~

0.05

LCO I

0

O.1

O~ 0

I

2 4 residence time [s]

6

gas

I

!

2 4 residence time [s]

0.03 -]-

6

coke

/

0.08

'a~0.02.4.

'03

~0.04

9

~ 0.01

g

CTO=5

If

[] []

0.02 0..

I

0

i

2 4 residence time [s]

0

6

0

2 4 residence time [s]

6

Figure 10. Description of the products from VGO cracking. Curves were calculated with the parameters given in Table 6.

183

6. CONCLUSIONS A five-lump kinetic model has been developed that adequately describes the experimental data obtained from different feedstocks. To account for the experimentally observed different time scales for the formation of coke and for the formation of the other products, different activity functions for these two processes were included in the kinetic model. The activity for coke formation was described by an exponentially decreasing function of the residence time with a characteristic time of about 0.02 s. The activity for the formation of the other products could be described most adequately using an exponentially decreasing function of the coke content of the catalyst. The model could adequately describe the activity of the coked catalyst in a second cracking experiment using fresh Hydrowax feedstock when it was assumed that the second amount of coke deposited on the coked catalyst did not further decrease the activity of the catalyst. The applicability of the model to the results from other feedstock-catalyst combinatiol~ has successfully been demonstrated: fitting of the model parameters to experimental data from a hydrotreated flashed distillate and a vacuum gas oil also yielded adequate descriptions.

Acknowledgment The authors are grateful to ir. Albert Plesman, ir. Jolanda van der Kamp, and ir. Marike Wissink, for performing parts of the experimental work. Dr. ir. J.M.H. Dirkx is acknowledged for the stimulating discussions. Shell International Oil Products is acknowledged for the supply of feedstock and catalyst and for financial support. 7. NOTATION Cc

CTO

kd kdA m n P p ?.

T Y

coke content of the catalyst (wt%) catalyst-to-oil ratio (kgcat'kgfeed-1) deactivation constant in ~conv related to coke content of the catalyst (wt% 1) deactivation constant representing the ratio between decrease in activity for conversion due to pore blockage and due to site coverage by coke (-) reaction rate constant of reaction j (gfeed'gcat"I'S'I) order of deactivation number of observations number of parameters (Pa) pressure (gi'gcatq's "l) reaction rate (i~) temperature (gi'gfeed"1) product fraction

Greek

(I)coke l~eonv

deactivation constant in (I)coke with respect to residence time mass flow rate activity function for coke formation activity function for conversion of HCO to all products except coke residence time

(S "l) (g.s "1)

(-) (-) (s)

184

Subscripts cat catalyst catalyst coke content CC conv conversion lump; HCO, LCO, gasoline, gas, coke i (reaction) number j,l,2,.. feed feedstock stripping section strip Abbreviations gasoline gin heavy cycle oil HCO hydrotreated flashed distillate HTFD light cycle oil LCO lack of fit LF not measured n.m. pure error, sum of squared errors of experimental data PE sum of squared residuals SSres vacuum gas oil VGO REFERENCES 1. E. Brevoord, F.P.P. Olthof, H.N. Wijngaards, and P. O'Connor, in P. O'Connor, T. Takatsuka, and G.L. Woolery (eds.), Deactivation and Testing of Hydrocarbon-Processing Catalysts, ACS Symp.Ser., Vol. 634, ACS, Washington, D.C., 1996, 340. 2. T.C. Ho, Ind. Eng. Chem. Res., 31 (1992) 2281. 3. M.A. Den Hollander, M. Makkee, and J.A. Moulijn, Appl. Catal. ,A, 187 (1999) 3. 4. M.P. Helmsing, M. Makkee, and J.A. Moulijn, in P. O'Connor, T. Takatsuka, and G.L. Woolery (eds.), Deactivation and Testing of Hydrocarbon-Processing Catalysts, ACS Symp.Ser., Vol. 634, ACS, Washington, DC., 1996, 322. 5. M.P. Helmsing, M. Makkee, and J.A. Moulijn, Chem. Eng. Sci., 51 (1996) 3039. 6. J. Corella and E. Franc6s, in M.L. Occelli (ed.), Fluid Catalytic Cracking I1: Concepts in Catalyst Design, ACS Symp.Ser., Vol. 452, ACS, Washington, DC., 1991, 165. 7. C.L. Koon, R. Hughes, Y.R. Tyagi, M. Castro Diaz, S.C. Martin, P.J. Hall, and C.E. Snape, in P. O'Connor and M.L. Occelli (eds.), Preprints Symposia, ACS, Div. Petr. Chem., Washington, DC., 1999, 478. 8. A.A.H. Mohammed, B.J. McGhee, J.M. Andresen, C.E. Snape, and R. Hughes, in M.L. Occelli and P. O'Connor (eds.), Fluid Cracking Catalysts, Chem.Ind., Vol. 74, Marcel Dekker, Inc., New York, 1998,279. 9. J.R. Bernard, P. Rivault, D. Nevicato, I. Pitault, M. Forissier, and S. Collet, in M.L. Occelli and P. O'Connor (eds.), Fluid Cracking Catalysts, Chem.Ind., Vol. 74, Marcel Dekker, Inc., New York, 1998, 143. 10. J. CorelM and E. Franc6s, in C.H. Bartholomew and J.B. Butt (eds.), Catalyst Deactivation 1991, Stud.Surf.Sci.Catal., Vol. 68, Elsevier Sciences, Amsterdam, 1991, 375. 11. E. Furimsky, Ind. Eng. Chem. Prod. Res. Dev., 18 (1979) 206. 12. T. Takatsuka, S. Sato, Y. Morimoto, and H. Hashimoto, Int. Chem. Eng., 27 (1987) 107.

185 13. I.P. Fisher, Fuel, 65 (1986) 473. 14. J.W. Beeckman and G.F. Froment, Ind. Eng. Chem. Fundam., 18 (1979) 245. 15. J. Corella and M. Men6ndez, Chem. Eng. Sci., 41 (1986) 1817. 16. J. Corella, R. Bilbao, J.A. Molina, and A. Artigas, Ind. Eng. Chem. Process Des. Dev., 24 (1985) 625. 17. A. Gianetto, H.I. Farag, A.P. Blasetti, and H.I. De Lasa, Ind. Eng. Chem. Res., 33 (1994) 3053. 18. L.L. Oliveira and E.C. Biscaia, Jr., Ind. Eng. Chem. Res., 28 (1989) 264. 19. G.F. Froment and K.B. Bischoff, Chem. Eng. Sci., 16 (1961) 189. 20. I. Pitault, D. Nevicato, M. Forissier, and J.R. Bernard, Chem. Eng. Sci., 49 (1995) 4249. 21.M. Forissier and J.R. Bernard, in C.H. Bartholomew and J.B. Butt (eds.), Catalyst Deactivation 1991, Stud.Surf.Sci.Catal., Vol. 68, Elsevier Sciences, Amsterdam, 1991, 359. 22. M.A. Den Hollander, Ph.D thesis, Catalytic Cracking in a Microriser. The Different Time Scales of Conversion and Coke Deposition, Delft University of Technology, Delft (2000). 23. P. Turlier, M. Forissier, I. Rivault, I. Pitault, and J.R. Bernard, ia M.L. Occelli and P. O'Connor (eds.), Fluid Catalytic Cracking III: Materials and Processes, ACS Symp.Ser., Vol. 571, ACS, Washington, DC., 1994, 98. 24. J.W. Beeckman and G.F. Froment, Chem. Eng. Sci., 35 (1980) 805. 25. J.-P. Janssens, A.D. Van Langeveld, S.T. Sie, and J.A. Moulijn, in P. O'Connor, T. Takatsuka, and G.L. Woolery (eds.), Deactivation and Testing of Hydrocarbon-Processing Catalysts, ACS Symp.Ser., Vol. 634, ACS, Washington, D.C., 1995, 238. 26. J.-P. Janssens, B.J. Bezemer, A.D. Van Langeveld, S.T. Sie, and J.A. Moulijn, in B. Delmon and G.F. Froment (eds.), Catalyst Deactivation 1994, Stud.Surf.Sci.Catal., Vol. 88, Elsevier Sciences, Amsterdam, 1994, 335. 27. G.F. Froment, in C.H. Bartholomew and J.B. Butt (eds.), Catalyst Deactivation 1991, Stud.Surf.Sci.Catal., Vol. 68, Elsevier Sciences, Amsterdam, 1991, 53.

Studies in Surface Scienceand Catalysis 134 M.L. Occelliand P. O'Conner(Editors) 9 2001 ElsevierScienceB.V. All rightsreserved

Deactivation Approach

187

of F l u i d C a t a l y t i c C r a c k i n g C a t a l y s t s : A M o d e l l i n g

Felipe L6pez-Isunza Departamento de Ingenieria de Procesos e Hidr~ulica Universidad Aut6noma Metropolitana-Iztapalapa Av. Michoacan y La Purisima s/n, Col. Vicentina, Iztapalapa M~xico 09340, D.F, MEXICO. A mechanistic model has been developed to study the deactivation of fluid catalytic cracking (FCC) catalysts, by describing interphase and intraparticle mass transfer interactions with cracking reactions, considering the influence of the concentration of active sites, the sizes of pore length and radius, the interface and diffusion mass transfer coefficients, and the mean residence time. The model describes an isothermal ideally-mixed fluidised bed, where surface reactions occur inside the cylindrical pore in a single pellet (microsphere) cracking catalyst, in which deactivation by coke deposition is described in terms of a L a n g m u i r Hinshelwood expression. The kinetics consider 4 lumps, adsorbed pseudo-species, the fraction of free active sites, and coke. Model simulations are performed to assess the role of different mechanisms and model parameters on coke deposition.

1. INTRODUCTION The systematic study on deactivation of cracking catalysts by coke deposition using different types of feedstocks, catalysts and reaction conditions, was initiated by Voorhies [1] in 1945. In this study the amount of coke formation in fixed and fluidised bed reactors was given by an empirical relation in terms of time on stream, ie. Coke=At n. This expression, together with the lumping of chemical species [2] to describe the kinetics of gas-oil cracking reactions, has been applied in several FCC reactor modelling works from i n d u s t r y and academia [2-6]. Both approaches have also been used in the analysis of experiments from micro-activity reactors for catalyst evaluation, translation of information between micro-activity, pilot and industrial reactors, or for kinetic p a r a m e t e r estimation [7- 9]. However, w h a t is suitable for modelling the behaviour in commercial reactors is not necessarily useful for the understanding of c a t a l y s t deactivation by coke deposition nor for e v a l u a t i n g c a t a l y s t performance. Deactivation by coke deposition has been directly related to coke itself, and not to time [10, 11], as coke formation is linked to the operating conditions, the nature of the feedstock and the type of catalyst. On the other hand, studies of the interactions between active sites, cracking reactions, coke deposition and multicomponent mass transfer are still missing. This type of research becomes important when the evaluation of catalyst performance using laboratory scale reactors is the main objective, particularly when very short

188 contact times are normally used in industrial reactors (5-10 s). The accessibility [12] to active sites within catalyst pores, a concept used in the manufacture of industrial catalysts, could be assessed using this kind of approach to complement any experimental programme with model molecules. FCC catalysts are made of several different components to produce the necessary cracking activity and selectivity with regular micro-porous structure given by the zeolite component, but u n k n o w n overall pore size distribution. Typically, FCC catalyst microspheres have an average diameter of 70 ~m and 30 .~ pore size [9]. For short contact times catalyst particles may experience diffusion limitations during the cracking of gas-oils, particularly in the case of bottoms conversion. The mechanisms by which coke is formed during the cracking of hydrocarbons are not well known, however, catalyst deactivation has been attributed to coverage of active sites by coke deposition, and pore blockage by coke growth [13, 14]. Assuming both mechanisms result from cracking reactions, they can be described by Langmuir-Hinshelwood kinetics coupled to intraparticle diffusion. In this way, all pseudo-species (or lumps) interact with active sites to form cracked products and coke which grows on the catalyst surface. As discussed by K r a m b e c k [15], the use of lumps to describe cracking reactions of multicomponent mixtures of hydrocarbons, imposed serious limitations on our ability to elucidate the mechanisms of FCC catalyst deactivation, and on adequately predicting product distribution and gasoline characteristics (MON and RON). A more fundamental approach, so-called "single-event kinetics", considers the description of the complete network of elementary steps of carbenium ion reaction intermediates on Lewis or Br0nted sites [16], but makes the analysis like the one presented in this work extremely complex. Another interesting approach considers an intermediate molecular description by defining lumps based on the chemical functions within each traditional cut, and accounts for all important reactions when modelling data from micro-activity reactors using different feedstocks with detailed product analysis [6]. This work deals with the cracking of a vacuum gas-oil in an isothermal, ideally mixed continuous fluidised bed reactor, using a lumped-kinetic scheme. The study considers a model for a cylindrical pore contained in a catalyst particle, where interphase and intraparticle mass transfer interact with a set of surface (cracking) reactions, involving four lumps and coke. The work analyses the relative contributions of transport and kinetic mechanisms on catalyst deactivation by coke deposition via model simulations. Different parameters and quantities are considered like: the magnitude of the average pore length and diameter, the m e a n gas residence time, diffusion coefficients, and the concentration of active sites which take into account the zeolite/matrix ratio. Because of the comparative sizes of the dispersed zeolite crystals within the catalyst pore to the (larger) molecules in lumps A and B, the former rather act as an active site. Moreover, the values of the effective diffusivities used in this work reflects the fact that diffusion takes place within the pore and not through the zeolite crystallite. 2. THE K I N E T I C S

The kinetic scheme considers the cracking of lump A (vacuum gas-oil) into lumps B (light cycle oil), G (gasoline), D (light gases, C1-C4), and coke which is

189 deposited on the catalyst, in t e r m s of reactions b e t w e e n active sites S, and adsorbed species according to" 2

B

1 A+S

~

G

~

AS

4

6 B+S

~-

1'

~

B S ~

6'

D

Coke

Coke D G+S

lO 10'

GS

/ ~ 12

Coke

3. T H E R E A C T O R M O D E L

The t r a n s i e n t model describes an isothermal fluidised bed reactor, where the feed of lump A is fully vaporised, and the heterogeneous cracking reactions take place inside the porous particles. For all species, diffusion takes place in the Knudsen region. The dimensionless interparticle mass balance for lump n in the gas phase is given by:

dxo~ 1 ----(XGn 0 --XGn)--5(XGn --Xn) dt

(1)

where ~ is the gas-phase m e a n residence time, and 5 is the ratio of interfacial mass t r a n s p o r t relative to the diffusional m a s s flux of A across the external catalyst surface. The mass balance for lump n in the gas-phase inside the cylindrical pore in the catalyst particle is"

G)t -- ~'n ~ 0~2 "~-ar ~ Oq~-2 "1-~'-~"

(2)

with initial and boundary conditions: at t=0:

XGA(0) : 1 ; XGn(0)-- 0

for n c A ;

Xn(0,~,~) = 0

(3)

at the closed end, ~=0, and at the radial axis, ~=0, of the cylindrical pore, respectively" &____~n= 0~. = 0

(4)

190 at the external catalyst surface, {=1, boundary condition relate interphase mass transport to diffusional mass flux of lump n. g/

a~

(5)

=Bi~(x~-xz~)

The model considers t h a t the heterogeneous reactions take place at catalyst walls inside the pore, then boundary conditions at catalyst surface, ~=1, couples mass flux of lump n to its surface reaction given in t e r m s of L a n g m u i r Hishelwood expressions. (6)

= D---~astn(O_,x)

where: Da is the D a m k 6 h l e r number; 7. is the ratio of effective diffusion coefficients; ar is the aspect ratio; Bin is the mass Biot number for lump n; 9~,(_0,x) is the overall rate of reaction for lump n at catalyst surface, a function of adsorbed (_0) and gaseous (x) species. The rate for coke deposition is given by: dOc = Das~Rs~ dt

=

Das(ksO A + k90 ~ + k120~)

(7)

Dasis the Damk6hler n u m b e r referred to the catalyst surface, and is related to Da (see Notation). The balances for lump A in the gas-phase and adsorbed on catalyst are given by: dX A : -Da~R A : -Da(lqxAO - klOA ) dt

(8)

dOA : Das~RsA : DasllqxAO~r dt

(9)

- ( k , ' + k 2 + k 3 + k 4 + k,)OAl, j

Catalyst activity, 0s, is expressed as the fraction of free active sites, which at any time is given in t e r m s of the total n u m b e r of adsorbed species and coke on catalyst: NS

0s = 1 - ~ 0 m

(10)

m=l

In this way, the model describes deactivation as due only to irreversible coke deposition. On the other hand, coke also grows on catalyst surface with a local rate given by:

dt

I, p~Rp

dt

(11)

191 Initially, at t=0, the model considers e i t h e r no coke on catalyst, or the typical a m o u n t of coke from a r e g e n e r a t e d equilibrium catalyst. A coke layer of thickness )~c builds up w i t h time according to E q u a t i o n (10), however, the model does not consider any negative influence of)~c on the diffusion coefficients. The n u m e r i c a l solution of the t r a n s i e n t model w a s p e r f o r m e d u s i n g the m e t h o d of orthogonal collocation [17] for E q u a t i o n s (2-6), using seven and one interior collocation points in the axial and r a d i a l co-ordinates, respectively. The use of up-to fourteen i n t e r n a l points in the axial direction gave practically the same results as with seven points, using shifted Legendre polynomials (a=[3=O). The resulting set of nonlinear ordinary differential equations was integrated with a s t a n d a r d fourth order R u n g e - K u t t a method. 4. R E S U L T S A N D D I S C U S S I O N

This s t u d y considers spherical pellets qf 70 lam in d i a m e t e r c o n t a i n i n g a cylindrical pore of 35 lam in l e n g t h and 30 A in radius. N u m e r i c a l s i m u l a t i o n s were performed considering a reference case w i t h p a r a m e t e r values given in Table 1, as well as their variation within plausible values. The reactor volume is 100 cm 3, with an a p p a r e n t bulk catalyst density, pp=890 Kg/m 3, operating at 793 K. Table 1. P a r a m e t e r values used in the study Parameter Reference case Pore length 35 lam Pore radius Concentration of active sites Effective diffusion of lump A Gas flow rate Interphase mass transfer coefficient

Range of parameter values 20 - 70

30/k 4.5 x 10.9 mole/m 2 lxl0 -8 m2/s

7.5 - 90 2.5 x 10.9 - 45 x 10.9

25 cm3/s

12.5 - 100

5 x l O -3 m / s

5 x l O - 3 - 5 x l O -6

l x l 0 -7 - 1 x l 0 -1~

4.1 T h e R e f e r e n c e C a s e F i g u r e 1 shows the e v o l u t i o n of a v e r a g e p o r e - v o l u m e d i m e n s i o n l e s s concentrations for the reference case. It is i m p o r t a n t to point out t h a t activity and coke on catalyst are linearly related by Eq. (10), and are given as an average fraction of active sites in the catalyst pore. For the p a r a m e t e r values used in this case, it can be observed t h a t it t a k e s a l m o s t 60 seconds of operation to achieve typical gasoline yields like those observed in MAT reactors.

192

~0

1

CO

0.8 E "0 v

.

0.6

!

cO

-G

n

0.4 c" 0 cO

0.2

0 '

0

10

20

30

40

Time

(sec)

-

50

-

D

60

70

Figure 1. Evolution of (volume average)concentration profiles with time on stream for the reference case. 4.2 T h e R o l e o f M a s s T r a n s p o r t The role of both interracial and diffusional mass transports on cracking kinetics is presented in Figures 2 to 4, which show average pore-volume dimensionless concentrations after 60 seconds of reactor operation. Figure 2 shows a region defined for values of interfacial mass transfer coefficients, kg>5 x l 0 4 m/s, where no limitation to interfacial mass transport exists. Otherwise, for lower values (kg<5 x l 0 4 m/s) conversion and coke yield are affected, having little impact on the yields of gasoline and light gases. To analyse the effect of intraparticle diffusion (accessibility), model simulations were performed considering three different aspects. In the first case, for effective diffusion coefficients in the range shown in Fig. 3, for DAeff<10 9 m2/s, mass transfer limitations exist. For larger values there is no effect of diffusion on conversion or product distribution. Second, Fig. 4 shows the combined effect of pore length and diffusion, with an important negative effect on gasoline yield when the effective diffusivity is two orders of magnitude smaller, for pores with lengths between 20 to 70 ~tm. This shows that large particles had apoor performance when diffusion limitations exist. In this respect, this explains, in part, why catalyst attrition improves gasoline yield as observed from experiments [18]. The third aspect is related to the time on stream, the physical dimensions of the catalyst pellet and the magnitude of diffusion inside the pore. Figure 5 shows pore accessibility along the time on stream, where it can be seen that even under no reaction conditions, the pore will need more than 60 seconds to be completely filled.

193

vn . 7_

u~
E

.

.

.

.

.

, .

.

0.6 ~ .................... ~

.

.

.

.

.

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

,,,, ~,,~.,+

~-cti-vi-ty ........

O.5

0 . 4 .......................~

:

......... ':':'. ........ '. .......................... C ~

.....

v

0

c(I) 0 cO

o . 3 ............................... i................................ ::.................................................................

o.~

) ............................. ::............................... !........................................................ B..~

(D

.....

t

. . . . . . .

, i1

F

!

0

"?

. . . . . . . .

0.0001

,+

,

,

,

, ,.L,

0.001 Interface

Mass

,

0.01 Transfer

A

,

,

,

~

LJ~

0.1 Coeficient

1

xlO e

(m/s)

Figure 2. Product distribution and activity for selected values of the interphase mass transfer coefficient.

0.7 " ~CD"

,

......

~----,---

,

0.6

t--

.o CD

0.5

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

~

~G

c"

(D

E

._

0.4 0.3

.............................. ~

................................... i c o k e ...............

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

~. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .

0.2 0.1

I

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

~

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

'.r~'~-',

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

,

i D

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

0 1 0 -7

10 -6 10 -5 0.0001 E f f e c t i v e Diffusi6n C o e f f i c i e n t x 1 0 4 ( m 2 / s )

Figure 3. Product distribution and activity for selected values of effective diffusivity.

0.001

194

0.54 !

0.52

D e f f ~ l O -8 m 2 / s

~ ..............................................................................

0.5

i

;

~

:

.

:

:

"1:3 m

>-

0.48

(D C

0

0.46 0.44

~_..................................::....................................................~....................................

0.42

!. . . . . . . . . . . . . . . . . . . . . . .

0.4 10

20

30

40

50

60

70

80

Pore Length (microns)

Figure 4. Concentration of lump G for selected values of pore length and effective diffusivity.

%" cat)

1

-

,

,

:

.

.

.

.

.

.

.

.

.

;

'

.o c

,

.

,

~ :

. . . .

T .

. . . .

~

. . . .

!

i 7

Only 0.8

(D

E "1:3

<

0.6

0.4

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

"5 0.2

o

(.3

0

......................................................~.................................. -J 0

10

20

30 Time

40

50

60

70

(sec)

Figure 5. Evolution of concentration profiles in the pore for selected values of diffusivities and different conditions.

195 0.6 Deft= 10 .8 tO

{Js-.--.-

'

. . . .'. . . . . . . . . . . . .,. . . . . . . . . . . . .

c-

0.4

E v" o

Gasolin;

...........................~

'

Deff=10l~ .........

co

~ o co

0.2 . . . .

o

0.1

.

0 -o ._

~0

....................i..................... i..................... i-..................... ~..................... i............

c--

iVaccum!Gas ~

,

'

i

.

.

.

.

.

.

.

.

.

.

.

.

.

.

.

.

.

.

.

.....................................D.~ff_l.0..s.__

~

,,i

0.55

O i l i /

-10

i , l , , ,

,,,

I~,

,~i

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

o.

g

0.5

~ o o0 9

0.45

O:~_-

~S o>., i--.o

o

De

ff = 1 0 a

m2/s

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

u.,4

0.35

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

o 11

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Figure 6. Axial profiles for vacuum gas-oil, gasoline and coke for selected values of effective diffusivities. Figure 6 show the axial profiles corresponding to Fig. 5, where it can be observed t h a t coke will block the pore faster for lower values of diffusivities, as

196 sharper profiles are created at the pore-mouth due to the larger difference in the relative rates between cracking reactions and diffusion.

4.3 A c t i v e Sites, P o r e R a d i u s a n d Gas F l o w - R a t e Model simulations showed t h a t the concentration of active sites (Ns), the pore radius (Rp), and the gas flow rate also affect conversion, yield and activity. Increasing Ns, as shown in Figure 7, will increase conversion and yield of G and D, while slightly reducing coke on catalyst, due to the larger n u m b e r of free active sites available. This case simulates increasing the ratio zeolite/matrix within the pore.

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197

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F i g u r e 9. P r o d u c t d i s t r i b u t i o n in t h e p o r e as a f u n c t i o n of t h e g a s flow rate.

120

198 5. C O N C L U S I O N S A mechanistic model has been developed to s t u d y deactivation of FCC catalysts, based on the description of interactions between mass transfer and a set of heterogeneous reactions occurring inside a cylindrical pore in cracking catalyst microspheres, in which coke deposition is given in terms of a LangmuirHinshelwood expression. The importance of the concentration of active sites, and the pore length and radius have also been examined. This kind of mechanistic model could be useful to correlate information between micro-activity and riser reactors, and to evaluate the performance of cracking catalysts.

Acknowledgements: The a u t h o r gratefully acknowledge the Instituto Mexicano del Petroleo and Consejo Nacional de Ciencia y Tecnologia (M6xico) for supporting this research.

6. NOTATION at: Bi.: Da: Das:

~j: Lp: Ns:

Rp: ~s.: t: trefi XGn : Xn :

aspect ratio=Lp/Rp Biot number for interfacial mass transfer. Damk6hler n u m b e r referred to the gas-phase = klNsRp~aez DamkShler number referred to the catalyst surface = Lp2klCaO/DAez rate coefficient for jth reaction. Pore length. Concentration of active sites (mole/m2). Pore radius. Overall reaction rate for lump n relating gas species. Overall reaction rate for lump n relating adsorbed species. time (sec). reference time = Lp2/Dnez dimensionless concentration of species n in the interparticle gas-phase. dimensionless concentration of species n in the intraparticle gas-phase.

Greek letters: ratio of effective diffusion coefficient for lump n with respect to lump A.

~tn:

5:

ratio of interfacial mass transport to diffusional mass flux of A at ~=1. radial co-ordinate of pore.

On: X~:

dimensionless concentration of adsorbed species n. thickness of coke layer. axial co-ordinate of pore.

Pp: a::

apparent bulk density of FCC catalyst in the reactor (Kg/m3) dimensionless time =

t/tre f

199 7. R E F E R E N C E S

1. A. Voorhies, Jr., Ind. Eng. Chem. 37 (1945) 318. 2. V.W. Weekman Jr., AIChE Monograph Series, No.11, 75 (1979) 1. 3. M. Forissier and J.R. Bernard, in C.H. Bartholomew and J. B. Butt (eds.), 1991Catalyst Deactivation, Elsevier, Amsterdam 1991, p. 359. 4. M.A. den Hollander, M. Makkee and J.A. Moulijin, Catal. Today 46 (1998) 27. 5. F. L6pez-Isunza, Comp.Chem. Eng. 16 (1992) $139. 6. F. Van Landeghem, D. Nevicato, I. Pitault, M. Forissier, P. Turlier, C. Derouin and J.R. Bernard, Appl. Catal. A: Gen. 138 (1996) 381. 7. A. V. Sapre and T. M. Leib, in M. Occelli (ed.) Fluid Catalytic Cracking. II. Concepts in Catalyst Design, ACS Symp. Ser. 452 (1991) 144. 8. U.A. Sedran, Catal. Rev.-Sci. Eng., 36 (1994) 405. 9. J. Ancheyta-Juarez, F. L6pez-Isunza and E. Aguilar-Rodriguez, Appl. Catal. A: Gen. 177 (1996) 227. 10. G. F. Froment, in G. Bond, P. Wells and F.C. Tompkins (eds.) Proc. Sixth Int. Cong. Catal. 1976. The Chem. Soc., London, 1977, p. 10. 1 1 . G . F . Froment, in C.H. Bartholomew and G. Fuentes (eds.) Catalyst Deactivation 1997, Elsevier, 1997, p. 53. 12.P. O'Connor and F. van Houtert, in H.Th. Rijnten and H.J. Lovink (eds.) Ketjen Catalyst Symposium, The Netherlands, 1986, Paper F-8. 13.J.B. Butt and E.E. Petersen, Activation, Deactivation and Poisoning of Catalysts, Academic Press, San Diego, 1988. 14.J.W. Beeckman and G.F. Froment, Chem. Eng. Sci. 35 (1980) 805. 15.F.J. Krambeck, in G. A s t a r i t a and S.I. S a n d e r s (eds.) Kinetic and Thermodynamic Lumping of Multicomponent Mixtures, Elsevier, Amsterdam, 1991, p. 111. 16.W. Feng, E. Vynckier and G.F. Froment, Ind. Eng. Chem. Res. 32 (1993) 2997. 17.J. Villadsen and M.L. Michelsen, Solution of Differential Models by Polynomial Approximation, Prentice Hall, Englewood Clifs, NJ, 1978 18.P. O'Connor, P. Imhof and S.J. Yanik, in P. O'Connor and M.L. Occelli (eds.) Fifth Int. Symp. Adv. Fluid Catal. Cracking, Preprints 218 th Nat. Meet. ACS, 1999, p. 466.

Studies in Surface Science and Catalysis 134 M.L. Occelli and P. O'Conner (Editors) 9 2001 Elsevier Science B.V. All rights reserved

201

.Catalyst Design for Resid Cracking Operation: Benefits of Metal Tolerant Technologies Lori T. Boock and Thomas F. Petti Grace Davison 7500 Grace Drive Columbia, MD 21044 410-531-4173

Abstract: FCC cracking operations that process resid-type feedstocks are often faced with the dilemma of managing significant levels of contaminant metals on their FCC catalyst inventory. Ni and V are of primary concern because of the ability of these metals to impair catalyst activity and to negatively impact catalyst yields. Typical practices for higher metals operations have included the use of equilibrium catalyst to "flush" contaminant metals from the unit, while adding minimal fresh catalyst just to maintain activity. This approach forces the unit to operate at a metal level that is often lower than the optimum, to counteract the contaminant metals brought in with the equilibrium catalyst and the potential negative yields associated with them. Modern catalysts, which incorporate metal tolerant technologies that minimize the effects of Ni and V, can be used to improve the overall selectivity of a particular unit's inventory. Fresh catalyst addition rates of these new technologies can often be much lower than the combined addition of "flushing" catalyst and fresh catalyst, with improvements in yields. This paper will examine the differences between these methods of catalyst management.

Introduction Refiners that process heavy FCC feedstocks are often faced with decisions concerning the management of contaminant metals.

These contaminants are deposited on the FCC catalyst

when organometallic complexes are present in the FCC feed [1]. Nickel, and to a lesser extent vanadium, act as dehydrogenation catalysts that increase the yields of unwanted byproducts, coke and hydrogen.

Vanadium also serves to deactivate FCC catalysts through the destruction of

active zeolite sites. The most common method of managing contaminant metals is to adjust the catalyst addition rate to maintain the level of nickel and vanadium on equilibrium catalyst (Ecat) at a pre-

202 specified level [2]. Usually this level is determined from past experience, and is set so that the production of contaminant coke and hydrogen remains at a comfortable level. This approach, however, is often based on outdated perceptions of reasonable contaminant levels. For example, Ecat trends show that there has been little increase in the average ECAT nickel level (about 1000 ppm on ECAT) in North America since the early-1980's [3, 4]. The average vanadium level has increased somewhat (from the 1300 ppm range to the 1800 ppm range). In contrast however, 8% of the FCC units worldwide operate with over 4000 ppm vanadium, and 10% of the units operate with over 3200 ppm nickel. By utilizing modem catalyst technologies, these high metals units are able to minimize the contaminant byproducts. Another approach to control the level of contaminant metals in a unit's inventory (and reduce fresh catalyst cost) is the use of purchased Ecat. A combination of fresh catalyst and purchased Ecat is used in an effort to maintain in-unit activity and selectivity while controlling the level of contaminant metals (and contaminant coke and gas). In practice however, purchased Ecat is not just an active metal sponge. This catalyst has been used at least once, and therefore brings some level of metals into the unit that can have negative yield impacts. Modem FCC catalyst zeolite and matrix technologies have greatly expanded the range of contaminant metals over which a refiner can successfully operate.

Matrix materials, such as

Grace Davison's LCM (Low Coke Matrix) have been commercially proven to passivate the effects of the nickel. Additionally, advances in zeolite technology (such as Grace Davison's Z17) have resulted in excellent tolerance to vanadium deactivation. The use of rare earth-based vanadium traps, is also gaining widespread use. These technologies have allowed refiners to maintain catalyst activity in the face of high metal feeds, without resorting to very high catalyst addition rates or the use of purchased Ecat to "flush" the metals.

203

Yield Impacts of Purchased Equilibrium Catalyst In order to explore the effects of purchased Ecat use, we have conducted evaluations targeted at comparing the performance of various catalyst addition strategies and purchased Ecat properties. The Ecat data plotted in Figure 1 is an example of the potential negative effects of adding purchased Ecat.

In this chart, we show the Ecat coke factor as a function of nickel

equivalents for two different FCC units. One unit uses an all-fresh catalyst addition policy, while the second unit runs about 50% purchased Ecat. Both units use the same fresh catalyst. The data clearly show that the all-fresh unit has a coke selectivity advantage at all metals levels.

Figure 1"

Comparison of ECAT coke factor versus equivalent nickel for two FCC units using the same base fresh catalyst.

We have also looked at operating data from a unit that has run both an all-fresh catalyst addition policy and a combination of fresh and purchased Ecat. This particular unit does not run very high metals, but decided to run purchased Ecat as a cost reduction measure. Table 1

204 summarizes the operation during the two periods of interest. Notice that when the unit ran with purchased Ecat as 1/2 of its additions, metals increased. This increase in metals had a negative impact in coke selectivity, lowering conversion under their air limit.

Fresh Catalyst Yield Advantages at Higher Metals The above examples show that at any given level of contaminant metals, a refiner can expect better performance from an all-fresh policy than a blend of fresh and purchased Ecat (including metals from the purchased material). However, the more interesting consideration is that many units would observe selectivity benefits from an all-fresh policy even at a lower catalyst addition rate than the blended catalyst approach. In these instances, the use of metals tolerant catalyst technologies can deliver improvements in selectivity and activity maintenance despite operating at higher total metals.

Table 1" Average Daily Commercial Data Comparing All-Fresh to Purchased Equilibrium Catalyst Addition Policies All-Fresh 89 Fresh Policy 89 Pur.ECAT Unit ECAT MAT 69.3 67.8 Ni, ppm 418 783 V, ppm 1173 1821 Feedrate, MBPD Riser Temp, ~ Air Rate, MSCFM Catalyst Adds, TPD

26.9 1001 50.1 2.0

26.6 1006 49.9 2.5

Conversion, iv% LCN, Iv%

74.5 55.7

73.3 53.5

The ability to take advantage of an all-fresh catalyst policy and higher contaminant metals depends on the catalyst technology employed and the level of feed metals. Grace Davison metals tolerant catalyst technologies have had success to nickel levels over 10,000 ppm and

205 vanadium levels over 7000 ppm. Surprisingly however, many of the units in North America that run some fraction of purchased Ecat are not extremely high metals operations. Of the FCC units that run purchased Ecat, over half run below 3000 ppm vanadium and all run below 2800 ppm nickel. At these low metals levels, the fraction of metals brought in with purchased catalyst is often a large percentage of the total ECAT metals. Many of these units would be better served to use an all-fresh catalyst addition policy. By leveraging metals tolerant technologies, they should be able to reduce the total catalyst additions and improve the selectivity (and profitability) of the unit, without significantly increasing the daily catalyst cost. We have conducted several pilot plant studies to explore the trade-off between running purchased Ecat and using all-fresh (metals-tolerant) catalyst at somewhat higher metals level. In this first example, we made a 50/50 blend of a refiner's typical purchased Ecat and fresh catalyst, and then metals deactivated the blend to the appropriate Ecat level. For comparison purposes, we calculated the feed added metals and deactivated a fresh catalyst with 1.5 times the incremental metals. The catalyst properties are summarized in Table 2. These two samples were run in our Davison Circulating Riser (DCR) pilot plant using a reduced crude feedstock. The results are summarized in Table 3.

Table 2:

Chemical: A1203, wt% RE203, wt% Ni, ppm V, ppm

Catalyst Properties for Riser Study Comparing All-Fresh to Purchased Equilibrium Blend 50/50 Fresh All-Fresh w/1.5x Cat/Pur.ECAT Incremental Metals 47.4 1.5 1166 530

51.7 1.8 1677 570

206

Physical: Total SA, m2/gm Zeolite SA, mZ/gm Matrix SA, m2/gm Unit Cell Size, A~

Table 3:

171 108 63 24.25

176 113 63 24.25

DCR Pilot Plant Yields at Constant Conversion Evaluation on Reduced Crude Feed

Conversion, wt%

50/50 Fresh Cat/Pur.ECAT 72

All-Fresh w/1.5x Incremental Metals 72

Cat/Oil Ratio C2-, w t % Total C3s, w t % Total Cgs, wt% Cs-Gasoline, wt% LCO, wt% Bottoms, wt% Coke, w t %

8.8 2.9 5.9 10.3 48.4 22.0 6.1 4.3

5.9 2.8 5.7 10.2 49.1 21.9 6.2 4.0

The data suggest that a coke and gas selectivity advantage can be realized for the all-fresh policy, despite the higher metals level. We see nearly a 10% reduction in coke yield at constant conversion. Additionally, the all-fresh approach is considerably more active, since it requires much lower cat/oil to achieve the 72% conversion.

In practice, the 1.5 times higher metals

would correspond to a catalyst addition rate of 2/3 the current total level. Our second pilot plant evaluation looks at a unit that runs to a higher contaminant level and a much higher percentage of purchased Ecat than our previous example. In this case, we examined all-fresh catalyst addition policies at 2 and 3 times the incremental metals level, using a metals tolerant catalyst technology. This would correspond to addition rates of 1/2 and 1/3 the current total in practice.

Table 4 summarizes the catalyst properties used in the DCR study.

These samples were run in our DCR as described above using a blended resid feedstock. The

207 constant conversion results are summarized in Table 5. Notice that the all-fresh catalyst with twice the incremental metals had to be interpolated at a higher conversion level then the other two catalysts because this material was significantly more active. Once again these results show advantages for the all-fresh approaches. Although the 3x incremental metals case does not show a large selectivity advantage, there is a significant difference in activity compared to the purchased Ecat case. The 2x incremental case, however, shows large selectivity and activity benefits.

Table 4: Catalyst Properties for 2 nd Riser Study Comparing (Metals Tolerant) All-Fresh to Purchased Equilibrium Blend Fresh Cat/ Pur.ECAT

All-Fresh w/3x Incremental Metals

All-Fresh w/2x Incremental Metals

39.8 1.6 1533 3100

52.4 1.9 2632 7170

52.4 1.9 1791 4990

146 114 33 24.26

109 71 38 24.21

167 109 58 24.29

Chemical:

A1203,wt% RE203, wt% Ni, ppm V, ppm

Physical: Total SA, m2/gm Zeolite SA, m2/gm Matrix SA, m2/gm Unit Cell Size, ~

Table 5:

2 nd DCR

Pilot Plant Study Yields at Constant Conversion Evaluation on Blended Resid Feed Fresh Cat/ Pur.ECAT

All-Fresh w/3x Incremental Metals

All-Fresh w/2x Incremental Metals

Conversion, wt%

70

70

77

Cat/Oil Ratio C2-, w t % Total C3s , wt% Total C4s , wt%

8.8 4.3 5.7 8.9

6.2 4.2 5.2 8.3

5.9 3.6 7.4 12.1

208 Cs-Gasoline, wt% LCO, wt% Bottoms, wt% Coke, wt%

45.5 22.4 7.6 5.3

46.6 22.7 7.3 5.4

48.3 18.0 5.0 5.4

Conclusions Modem catalyst technologies have had a major impact on the levels of contaminant metals that can be handled in an FCC unit. However, many refiners choose to run some fraction of purchased Ecat in an effort to "flush" contaminant metals from the unit while minimizing daily catalyst cost. Despite the advances in metals tolerant catalyst technology, many units that run purchased Ecat operate at fairly moderate metals level. These units in particular have the most to gain by re-thinking their catalyst management philosophy.

The use of an all-fresh

catalyst policy designed with appropriate metals tolerant technologies can provide selectivity benefits at a lower catalyst addition rate and therefore, higher ECAT metals level. These benefits may be possible with very little impact on daily catalyst costs.

References

[1]

P.B. Venuto and E.T. Habib, Jr., Fluid Catalytic Cracking with Zeolite Catalysts, Marcel Dekker, Inc., New York, 1979.

[2]

R.W. Mott, "New Technologies for FCC Resid Processing," Paper No. AM-91-43, presented at the NPRA Annual Meeting, San Antonio, TX, 1991.

[3]

D.W. Cotterman, "Trends in Equilibrium Fluid Cracking Catalyst Properties: Part 79," Grace Davison Catalagram | No. 83, 1992.

[4]

M.T. Smith, "Worldwide Equilibrium Fluid Cracking Catalyst Properties Part 81," Grace Davison Catalagram | No. 86, 1998.

Studies in Surface Science and Catalysis 134 M.L. Occelli and P. O'Conner (Editors) 9 2001 Elsevier Science B.V. All rights reserved

209

Active Site Accessibility of Resid Cracking Catalysts Yong Lu, Mingyuan He, Jiaqing Song, and Xingtian Shu Research Institute of Petroleum Processing, China Petrochemical Corporation, Beijing 100083, P. R. China It is well understood that pore structure is important in designing FCC catalysts. For residue cracking, the pore structure of catalyst matrix should be accessible for precracking the large hydrocarbon molecules so that the resulting smaller molecules can transfer into the zeolite channels and can be converted to value-added products over the zeolite sites. However, pore parameters such as pore volume and pore size distribution are insufficient for selecting an appropriate matrix. A method for testing the active site accessibility of resid cracking catalysts has been developed on currently widely used MAT equipped with the unstripped hydrocarbon burned off in place, with an online measurement of the CO2 and H20 that are produced. The effect of pore distribution of the FCC catalyst matrix on the active site accessibility of several catalysts has been tested tentatively in terms of bottoms conversion and strippable coke yield. The results showed that the yields of HCO and strippable coke increase when the matrix small pores increase, and decrease when the catalyst matrixes are highly accessible. These test observations are in agreement with commercial results. Key Words: resid cracking, catalyst, active site accessibility, strippability 1. INTRODUCTION Maintaining high conversion level and proper yield structure for a resid FCC unit (RFCC) is essential for maximizing the profitability of a refinery. The catalyst choice is certainly a key factor in the operation. Unlike the conventional FCC, where cracking reactions are usually not diffusionally limited, there are observable diffusion restrictions for large size molecules within zeolite channels and activity loss in RFCC units [ 1]. Recently, Khouw et al. [2] tested a vanadium contaminated equilibrium catalyst from a RFCC unit. The results have shown that the conversion level for a feedstock with high CCR content drops dramatically. Zhao [3] investigated the deactivation of FCC catalysts by sodium in commercial operations and found that the surface area of the matrix is more strongly affected than that of the zeolite. Mann et al [4] studied the mass transfer into catalyst particles using a special technique on SEM. The results illustrated the fact that the accessibility of a typical catalyst is far from perfect. Hence, for residue cracking, the pore structure of the catalyst matrix should be accessible

210 for precracking the large molecules so that the resulting smaller molecules can be converted to FCC products within the zeolite channels. However, pore parameters such as pore volume and pore size distribution are insufficient for selecting an appropriate matrix. To measure the catalyst accessibility, Akzo Nobel has developed a method based on the non-steady state diffusion of hydrocarbons into FCC catalyst particles, which measures a relative rate of mass transfer called the AAI (Akzo Nobel Accessibility Index) [5]. Also, the cracking of TIPB has been used as a test reaction to study the activity and accessibility of zeolites. The results showed that both seem to control the rate of reaction [6]. In the present paper, a method for testing the active site accessibility of resid cracking catalysts has been developed through modification of currently widely used MAT equipment. The effect of the pore size distribution of the catalyst matrix on the active site accessibility of several catalysts, has been tentatively tested in terms of bottoms cracking and strippable coke yield. 2. EXPERIMENTAL

2.1. Catalysts Two catalysts, designated as M-1 and M-2, were prepared with equal zeolite contents and with matrixes having different pore structures. For this purpose, A-1 (7-A1203, SA: 275m2/g, PV: 1.22ml/g, Crystallinity: 85%) and A-2 (7-A1203, SA: 266m:/g, PV: 0.286ml/g, Crystallinity: 85%) were used as the matrix for M-1 and M-2 respectively. The pores in A-1 with diameter of >80A share ca. 88% of the total pore volume, but the pores in A-2 with diameter of <80A share ca. 85% of the total pore volume. The typical preparation is as follows. The catalyst M-1 was prepared by mixing 6.0 g (dry base) of A-I, 10.5 g (dry base) of kaolin clay, 14.1 g of A1 sol-gel (21.34 wt% of A1203) and 46 ml of distilled water, and stirring the mixtures for 10 minutes. The pH of the mixture was adjusted to ca. 3.0 by slowly adding 2.0g of HC1 (19% by weight) and stirred continuously for 0.5 hour. 6.0g of USY and 3.0g of REUSY were added to the resulting slurry, followed by vigorously stirring for 0.5 hour. The final slurries were dried and calcined at 523 K in air for 3 hours. Four commercial fresh catalysts (A, B, C, D) and three corresponding commercial equilibrium catalysts (designated as E-A, E-C and E-D), were also employed in this study. The fresh catalysts were deactivated at 1073 K for 4 hours with 100% steam in a fixed bed reactor and the equilibrium catalysts were calcined at 823 K for 2 hours under air before testing. 2.2. Characterization Nitrogen adsorption-desorption analyses were performed for all catalysts to measure the surface area and pore structure. Table 1 shows that the matrix surface areas of the catalysts are in the range of 33-6 lm2/g. The micropore volumes are between 2.2-6.1 x 10.2ml/g. In-situ IR analyses for adsorbed pyridine were conducted for measuring the quantity of

211 Bronsted and Lewis acid sites. Table 2 shows that the concentrations of the acid sites measured by IR spectra of pyridine adsorbed at temperatures 473 K (623K) ranks as follows" Bronsted acid site, A>C-B (A-B>C); Lewis acid site, B>C-A (A>C>B). On the whole, there are no significant differences in the acidity of the tested catalysts. Table 1 The Pore structure of the catalysts Total surface Matrix surface Catalyst . area, m2/g area, m 2 / g M-1 185 61 M-2 187 57 A 122 63 B 122 45 C 182 57 D 125 46 E-A 95 48 E-C 101 33 E-D 96 35 .

.

.

.

.

.

.

.

.

.

.

.

Total pore volume, cm3/g 0.253 0.127 0.180 0.156 0.139 0.135 0.161 0.128 0.100 .

.

Microp0re volume, cm3/g 0.052 0.061 0.028 0.036 0.059 0.037 0.022 0.032 0.028

Probable pore size, A 108 76 121 40/109 39/82 39/83 112 40/109 40/88

,

Table 2 The acid sites concentrations of the commercial FCC catalysts ' Catalyst A B C D

473 K B acid, ~tmpl/g 35.59 28.81 28.81 n.d.*

L acid, gmol/g 26.19 30.95 26.19 n.d.

6 2 3 I( B acid, lamol/g 11.86 11.86 10.17 n.d.

" " L acid, ~tm.ol/g 15.48 13.10 14.23 n.d.

* not detected. 2.3. Reactor and Procedures The cracking experiments were carried out in a downflow MAT (microactivity test) unit designed for handling heavy feeds. The unit consists of a fixed-bed reactor made of stainless steel, a syringe pump, a hot box, a liquid product collector, a liquid displacement column and a wet test meter. Catalyst loading was always 5 grams. The amount of heavy oil feedstock (the properties were summarized in Table 3) injected was 1.56 grams. Cracking was performed at a reactor temperature of 773 K and a time on stream of 70 seconds. The stripping time was between 30 and 900 seconds, using high purity N 2 as stripping gas with a flow rate of 30 ml/min. The liquid products were characterized by high-temperature gas chromatograph distillation (GCD). Gas compositions were analyzed by GC.

212 Table 3 The properties of the heavy oil feedstock Density, (g/cm 3, 293 K) Viscosity, (mm2/s) 353 K 373 K Solidifying Point, (K) Aniline Point, (K) Acidity, (mgKOH/g) Basic Nitrogen, (ppm) CCR, (wt%) Composition, (wt%) Parraffin Aromatics Resins Distillation, (K) 5% 50% 90% ,

0.8916 9.629 6.312 320 366.5 0.81 404 0.28 63.6 29.2 7.2 616 715 791

The amount of coke formed on the catalysts was determined by in-situ combustion, using a CO2 IR Carbon Detector. An on-line CO oxidation reactor (873 K) packed with cobalt oxide catalyst was added between the reactor and CO2 IR Carbon Detector to ensure the analyzed product was CO 2. Hydrogen content of coke (designated as H) was determined by combustion. Before the combustion of the coke was started, two weighed glass tubes packed with Mg(C104) 2 particles, were connected to the reactor outlet and the CO oxidation reactor outlet respectively. The product HzO formed during coke combustion could be completely adsorbed by Mg(CIO4)z and its amount determined by weighing. By this way, H (hydrogen content of coke) was calculated. Table 4 shows how reproducible the H measurements are. Table 4 The replicate measurements of the hydrogen content of coke (/-/) over catalyst A ,

,

Test number 1

....

2 3 4

,m

Hydrogen content of coke (/-/) Commercial spent A Spent A in MAT test 8.4 7.6 8.2 7.8 8.3 7.4 8.4 7.5

a

a After a 15rain-Nitrogen-Stripping To calculate the yield of strippable coke, an intrinsic parameter for the strippability of FCC

213 catalysts, the coupled equations (Equation 1 and Equation 2) given below was set up on the basis of the mass balance of C and H in coke. Hx Y = Hsc x Ysc + Hrc x Yrc Y= Ysc + Yrc

1) 2)

where Y, Ysc and Yrc represent yields of total coke, strippable coke, and reaction coke (catalytic coke + contaminative coke + feed CCR coke) respectively; H, Hsc and Hrc represent the hydrogen content of corresponding cokes respectively. According to literatures [7,8], Hsc and Hrc were assigned values of 12 and 5 respectively. 3. RESULTS AND DISCUSSION 3.1. Strippability of tested catalyst Table 5 shows the yield of strippable coke with different stripping time lengths. The yield of strippable coke closely increases in the following sequences: M-l, M-2; A, B-C, D; E-A, E-C, E-D. If the yield of strippable coke characterizes the strippability of an FCC catalyst, then the strippability of the tested catalysts can be rank-ordered. The results can be approximately correlated with the total pore volume and average pore size, but not with the total matrix surface areas as shown in Table 1. In the catalyst group of A, B, C, and D, the catalyst A has the largest matrix surface area, but gives the lowest yield of strippable coke.

Table 5 The yield of strippable coke after stripping for different time lengths Catalyst Yield of Strippable Coke After Stripping For Different Time Length, wt% 30 s 480 s 900 s M1 12.72 4.27 2.96 M2 11.52 3.90 2.90 A 9.05 2.39 1.52 B 9.36 4.15 2.83 C 10.60 5.37 1.89 D 13.42 5.73 2.89 E-A -1.90 1.15 E-C -2.89 1.58 E-D ~ 3.97 2.34 3.2. Effect of pore size distribution of catalyst matrix on reduction of strippable coke Fig. 1 shows the relationship between the yield of strippable coke and the pore diameter distribution of the tested catalysts. In the tested catalysts, the specific surface areas in pores in 20-50 A range increase in following orders: MI<M2; A
214 pores in the 50-200A and 200-1000A range, the order (except for B) is reversed. The yield of strippable coke generally increases when the area in the small pore (20-50A range) increases, and generally decreases when large pores (>50A) in the catalyst matrix increase. Yanik et al. [9] reported that the initial soft coke increased with the increase of zeolite content and the content of small pores in the matrixes. Their results also show that a catalyst based on a moderate zeolite content and a highly accessible large pore matrix system can reduce the quantity of adsorbed hydrocarbons and hence the soft delta coke.

20 r 18

7

t

16i

~j 5

~a~14 1

O O

~E 12

..Q

~. lO o

8

3 . ~i,..

=

6

i 2 ""ao .e_ 1 >-

4

2 0

0

' M-1 M-2

A

B

C

D

E-A E-C E-D

Catalyst ,n I~

SA in 20-50 A range pores + stripping for 900s

stripping for 480s

1 1

60

I 7

50

16

J

40

~

O

~ 3o 3.~

N 2o

2 ~-a

, M-1 M-2

A

B

C

D

I,}ll,ll

!

o

E-A E-C E-D

Catalyst n

SA in 50-200A range pores

i

i "-- sb'ipping for 480s

SA in 200-1000A range pores stripping for 900s

Fig. 1. Relationship between the yield of strippable coke and the pore size distribution of catalyst matrixes

215 Fig. 1 also shows that although catalyst B has fewer small pores than catalyst C, its strippability is poorer. Table 2 shows that catalyst B possesses more Lewis acid sites than catalyst C, which may outweigh the effect of pore structure. Pavel [10] has noted that the acidity of a catalyst is an important factor in its strippability. 3.3. Effect of pore size distribution of catalyst matrixes on heavy oil cracking ability Table 6 shows the tested catalysts activity and selectivity for heavy oil cracking. The heavy oil cracking ability of the tested catalysts decreases in the following order: M-I<M-2; A
a

,

,, ,,

Specific Coke b 23.40 20.60 2.16 2.20 2.46 2.78 2.58 2.67 5.06 5.24 5.33

216 3.4. Commercial data Table 7 shows the commercial performance of catalysts A and C. Clearly, when resid FCCU is using catalyst A (with more accessible pore system), a high yield of higher-value FCC products is achieved with a slight reduction of HCO yield and with a remarkable decrease in hydrogen content of the generated coke.

Table 7 The Commercial results of FCC catalysts Catalyst A Density of Feed, g/cm 3 0.9043 E-cat. Property Conversion, wt% 56 Ni on catalyst, ppm 3885 V on catalyst, ppm 8700 Cat. Addition, tons/day 1.35 Yield structure, wt% Dry gas 3.97 LPG 9.03 Gasoline 43.11 LCO 29.81 HCO 4.60 Coke 9.48 Hydrogen content of coke, wt% 8.89

C 0.9108 57 3500 8700 1.40 4.31 10.22 43.13 28.08 4.92 9.34 9.90

4. CONCLUSIONS A method for testing the active site accessibility of resid cracking catalysts has been developed through modification of currently widely used MAT equipped with the unstripped hydrocarbon burned off in place, with an online measurement of the CO2 and H20 that are produced. The bottoms conversion and the strippable coke yield, which are characteristics of the active site accessibility of the FCC catalyst, can be determined by the proposed method. Test results show that the pore size distribution of the catalyst matrix is a key factor in improving active site accessibility of an FCC catalyst. With the increase of the surface area or pore volume in the matrix small pores, the strippable coke yield and the bottoms yield increases. However, the catalyst with a highly accessible large pore matrix exhibits good performance in reducing the strippable coke yield and in enhancing bottoms conversion. Test results are well in agreement with commercial results of the catalysts with respect to bottoms cracking and hydrogen content of coke. However, for optimizing the active site accessibility of FCC catalysts, a better understanding of the interactions between the catalyst pore parameters and heavy oil molecules is required. Moreover, the development of novel porous materials is right now the

217 subject of further thinking and experimentation. REFERENCES

1. P. O'Connor and A. P. Humphries, Am. Chem. Soc. Div. Petrol. Chem. Preprints, No. 38(3) (1993) 598. 2. F. H. H. Khouw, M. J. R. C. Nieskens, M. J. H. Borley, and K. H. W. Roebschlaeger, NPRA 1990 Annual Meeting, Paper No. AM-90-42 (1990). 3. X. Zhao and W. Cheng, in: P. O'Connor and T. W. Takatsuka (eds.), ACS Symposium Series 634, American Chemical Society, Washington, D C, Chap 11 (1995) 159. 4. R. Mann, K. Khalaf, and A. Lamy, in: P O'Connor and T. W. Takatsuka (eds.), ACS Symposium Series 634, ACS, Washington, D C, Chap 3 (1995) 42. 5. A. Humphries, R. P. Fletcher and J. R. Pearce, NPRA 1999 Annual Meeting, Paper No. AM-99-63 (1999). 6. S. Falabella, E. Aguiar, M. L. Mur-Valle, E. V. Sobrinho and D. Cardoso, in: Bonnevoit, S. Kaliaguine (eds.), Zeolite: A Refined Tool for Designing Catalytic Sites, Elsevier, Amsterdam, (1995). 7. H. C. Kliesch, et al., Paper presented at the Ketjen catalysts Symposium, (1987). 8. J. L. Mauleon and W. S. Letzseh, Paper Presented at the 5th Katalistiks Annual FCC Symposium, Chap 7 (1984). 9. S. J. Yanik and P. O'Connor, NPRA 1995 Annual Meeting, Paper No. AM-95-35 (1995). 10. S. K. Pavel and F. J. Elvin, NPRA 1998 Annual Meeting, Paper No. AM-98-42 (1998).

Studies in Surface Science and Catalysis 134 M.L. Occelli and P. O'Conner (Editors) 9 2001 Elsevier Science B.V. All rights reserved

219

Catalyst evaluation for atmospheric residue cracking, the effect of catalyst deactivation on selectivity W.R.Gilbert PETROBRAS R&D Center, Process Division, Rio de Janeiro, 21949-900,Brazil

In December 1999 PETROBRAS's new residue cat-cracker, in the RECAP refinery, Sao Paulo, came on stream. After a number of preliminary evaluations done with MAT, a set of five catalyst candidates from different vendors was chosen for pilot riser evaluation with atmospheric residue feedstock. A sixth catalyst with lower metal loading was used as base case. Mitchell method (1) impregnation followed by CPS deactivation (2) was used for Ecat simulation. The atmospheric residue cracking results showed relatively high activity (conversiort>75w% at catalyst/oil (c/o) ratio 7) and very high coke yields (15-25w%). The effect of deactivation severity on selectivity was then investigated as a possible explanation for the high coke yields. One of the catalysts of the previous study, loaded with the same level of nickel and vanadium was treated at different CPS deactivation times, from 4h to 60h, and the resulting catalysts used for MAT evaluation. CPS deactivation did not show any improvement of hydrogen selectivity with run length. Calculations done with the MAT data using a 4 Lump model and subsequent correlation of the kinetic constants with catalyst surface area showed that selectivity could be significantly improved (lower coke yield and higher gasoline yield) if the catalyst had been deactivated at a higher severity than the standard CPS procedure.

1. Introduction: It had long been suspected, in catalyst pilot riser evaluations done in the PETROBRAS research center, that the standard catalyst deactivation procedures used were too mild for the very high activity fresh catalysts used in PETROBRAS refineries, leading to higher coke yields and more gasoline overcracking than seen in commercial FCC units. Surface areas of deactivated catalysts would average 30m2/g higher than commercial Ecat with the same metal loading, differences in conversion at constant c/o and coke at constant conversion of 10w% and 3w% respectively would also be common. In a recent evaluation done for a new residue FCC unit being built in RECAP refinery, Sao Paulo, the pilot riser coke yield absolute values for the catalysts tested were far above the acceptable levels, even for atmospheric residue cracking.

220

2. Catalyst evaluation: 2.1. Experimental Section: Table 1 RECAP refinery FCC feedstock properties: Sp.gravity g/cc 0.975 Sulfur w% 0.65 Visc.,cst @60~ 2660 Sim.dist ~ T5% 385 T50% 592 Nickel ppm 20 Vanadium ppm 18 Carb.residue (ASTMD524) w% 9.7 Basic Nitrogen ppm 2000 Five catalysts were chosen for comparison after preliminary MAT screening. The fresh catalysts were loaded with 6000ppm Ni and 5000ppm V (40% of the expected Ecat levels) using the Mitchell method (1), a sixth catalyst loaded with 2500/2100ppm Ni/V (simulating a higher make-up rate) was used as base case. The six catalysts were then submited to CPS (cyclic propylene steaming) deactivation (2). The evaluation was done with RECAP ATR described in Table 1,680g/h feed rate, 20w% inert gas, 540~ riser temperature and atmospheric pressure in a circulating pilot riser. C/O was varied between 5 and 10 for each of the catalysts. Coke was calculated from the catalyst regenerator flue gas on line cromatographic analysis. Gasoline, LCO and HCO were determined by cromatographic distilation of the liquid product, using C5-216~ 216~176 and 344~ as cut points. Catoil ratio was calculated by correlation from the stripper-regenerator transfer line temperature and checked with the carbon on spent catalyst analysis. A high metal Ecat, different from the catalysts tested, was run in the pilot riser at the same conditions to see what the pilot riser base-line coke yield would be.

221 2.2. Results: Table 2 Deactivation results: Catalyst name A B C D ...... E F ....G Fresh catalyst BET surface area (m2/g) 258.4 273.6 291.3 319.6 306.0 368.1 External surf.area (m2/g) 90.8 146.1 94.9 83.0 99.0 96.0 Micropore vol. (cc/g) 0.074 0.060 0.081 0.106 0.093 0.122 Deactivated catalyst BET surface area (m2/g) 164.6 164.5 155.3 147.1 153.1 205.0 127.1 External surf.area (m2/g) 45.0 75.7 73.5 97.0 97.4 68.0 26.9 Micropore vol. (cc/g) 0.054 0.040 0.038 0.045 0.045 0.062 0.045 Nickel (ppm) 2461 6200 6000 6200 6000 6200 5452 V (ppm) 2170 5000 4900 5000 4900 4900 4494 Surface area retention (%) 63.7 60 53 46 50 56 Micropore vol retention (%) 72 66 47 43 48 51 Catalyst G was the equilibrium catalyst used to check the pilot riser base line coke selectivity. Typical of catalysts used in PETROBRAS's refineries, which process high basic nitrogen feeds, surface areas were large for all the catalysts tested. In spite of the high vanadium concentration used in the impregnation, surface area retentions were relativly high. Table 3 Riser evaluation results*: Catalyst name Conversion at c/o=7 (w%) Conversion (w%) Catoil ratio Hydrogen (w%) Dry gas (w%)

A 76.4 75 6.5 0.5 5.0

B 78.7 75 5.9 0.9 4.5

D 81.4 75 5.0 0.9 6.7

LPG (w%) 14.0 14.7 11.3 Gasoline (w%) 40.8 36.9 34.9 LCO (w%) 15.8 16.2 14.0 HCO (w%) 9.2 8.8 11.1 Coke (w%) 15.2 19.0 22.1 ACoke (w%) 2.3 3.3 4.4 *Coke yield for catalyst C was so high that stable pilot riser

E 75.8 75 6.8 1.0 6.0 13.9 35.5 15.2 9.8 19.7 2.9 operation

F 82.1 75 4.8 1.4 6.6

G 65.4 75 11.3 0.6 4.9

14.7 16.7 33.0 36.7 14.7 15.4 10.3 9.6 20.8 16.7 4.3 1.5 was not possible.

222 Although a clear ranking could be established from the results in Table 3, with catalyst B coming as champion, the very high coke yields and delta coke led to some discomfort, so it was decided that the effect of catalyst deactivation severity on coke yield should be further investigated. 2.3. Effect of d e a c t i v a t i o n run t i m e on catalyst selectivity:

Theoretically, submitting the metal loaded catalyst to redox cycles, as in the CPS procedure, would passivate the nickel and vanadium on the catalyst surface, reducing their activity as coke and hydrogen promoters. The number of cycles to which the catalyst was exposed should then have an effect on coke and hydrogen selectivity. To test these concepts, a fresh batch of catalyst E was impregnated with the same level of nickel and vanadium as before and submitted to the same CPS deactivation conditions. At varying run lengths, catalyst samples were drained from the deactivation reactor and separated for surface area determination and MAT evaluation. When the MAT hydrogen yield results were plotted against activity (figure 1) no difference in hydrogen selectivity with run length could be seen. The MAT coke results were not so clear (figure 2) with the 4h deactivated catalyst giving a slightly higher coke yield followed by the 60h and 20h ones. To better clarify other effects of catalyst deactivation on selectivity, kinetic constants for a 4 Lump model (3,4> were determined (figure 3) and correlated with surface area, generating a new model showing how conversion, gasoline and coke yield change with deactivation severity for catalyst E (tables 4 and 5). MAT results from the original catalyst E (old batch) were also used in the model calculations. 3

......

25 1

j

S j~-

I

~,

,

8 lo

-

5-t

0

I

0

....

i

2

"

!

4

6 Activity

Figure 1 - Hydrogen versus activity for catalyst E (m 4h deactivation, A 20h, <> 60h )

0

2

4

6 Activity

Figure 2 - Coke versus activity for catalyst E

223

I

lDr,,,Oas+LPO+Coke I

......., Io

oiino I

[Drye~s+LpO I I Coke I

Figure 3- 4 Lump model used. Table 4 Chan~e in catalyst E properties with CPS Batch Run length (h) 4 BET surface area (m2/g) 209.7 Micropore vol. (cc/g) 0.063 Ko 173.1 K1 109.3 K2 1377 K4 37.5 Ks 27.0

deactivation ru n length. New 20 60 166.6 125.0 0.050 0.03,7 106.0 53.8 76.6 39.7 10.68 653.4 15.18 7.8 12.2 5.2

Old 20 153.1 0.045 70.2 51.0 850.0 12.9 4.7

Table 5 Correlations obtained between the lump kinetic constants and surface area (sa) K Equations R2 Ko Ko=l.47(sa)-147.4 0.96 K1 KI= Ko (-1.3e-3(sa)+0.92) 0.86 K2 K2= K1 (-5.0e-2(sa)+23.1) 0.80 K4 K4= Ko (7.76e-4(sa)+0.05) 0.61 K5 1(5= Ko (8.4e-4(sa)-0.03) 0.62 The kinetic constants for the differential equations derived from the diagrams in figure 3 were calculated so as to minimize the differences between calculated and experimental yield results (Table 4), and then correlated with the deactivated catalysts surface area for varying deactivation run lengths (Table 5). Figure 4 shows the comparison between predicted and observed results for gasoline, coke and conversion. The new models average error for conversion, gasoline and coke yields was 3.2, 1.4 and 1.7 w%.

224

4s.o ]

~

15

r,/,. ,

37.5

9 O0~SO 00

80

,25 r 40.0

25-

I

70

A

.

,

j

37.5

40.0

42.5

160 ; 45.0

5

!

60

70

80

|

5

15

25

Figure 4 - Model prediction (derived from equations in table 5) versus experimental results for gasoline (left), conversion (middle) and coke (right). The equations from Table 5 were then used to generate the selectivity curves as a function of c/o and surface area displayed in figures 5 through 8. 55 7

100

8O

,..

t

._. 45

60

.~ 35

0 .~

m O

I1)

~ 4o

(5

0

o

25

20

15 100 100

140 180 Surface Area (m2/g)

220

Figure 5- Model generated curves for conversion vs surface area at different c/o's.

140

180

220

Surface Area (m2/g)

Figure 6 - Model generated curves for gasoline vs surface area at different c/o's.

225

45

40 1

!

Ic/~ I /

30

//

~2o 8 lO 0

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

100

!. . . . . . . . . . .

140

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

180

!

220

Surface Area (rn2/g)

Figure 7 - Model generated curves for coke vs area at different c/o's.

15

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

0

,

10

,. . . . . . . . . . . . . . . . . .

Coke (w%)

20

30

Figure 8 - Model generated curves for gasoline vs coke at different c/o's.

3. Conclusions:

What figures 5 to 8 show is that the best gasoline selectivities for catalyst E are around 120m2/g. Insufficient catalyst deactivation lead to formation of unwanted by-products, such as coke. As activity dropped, there was a slight increase in gasoline yield, with the loss in activity being more then compensated for by the rise in c/o ratio. Under 120m2/g the drop in activity is too steep and it is no longer possible to recover the gasoline yield and conversion by raising c/o ratio. One may infer that a similar behavior would be expected of the other catalysts tested and a more realistic equilibrium simulation would had been achieved if the deactivation procedure had been more severe. Studying selectivity for a given catalyst at two different deactivation severities may be a way of avoiding deactivation mistakes, at least for small scale testing, such as M A T . The extra cost of the additional tests could be offset by the use of a FCC model, such as the one presented, which would allow a safer interpolation of experimental yields with a smaller number of tests. 4. References:

1. B.R. Mitchell,, Ind.Eng.Chem.Product Research Development 19, 209 (1980) 2. L.T.Boock, T.F. Petti, J.A. Rudesill, ACS Div. Petr. Chem. 40, (3), 421-426, (1995). 3. D.M. Nace, S.E. Voltz, V.M. Weekman, Ind.Eng.Chem., 45, 1186 (1953). 4. L.C. Yen, AIChE Spring National Meeting, Session #84, (1989).

227

Optimum properties of RFCC catalysts Sven-Ingvar Andersson a and Trond Myrstad b aChalmers University of Technology, Department of Applied Surface Chemistry, SE-41296 Gothenburg, Sweden bStatoil's Research Centre, N-7005 Trondheim, Norway In order to test and evaluate residue fluid catalytic cracking (RFCC) catalysts and find the optimum catalyst for the UOP RCC unit at the Statoil Mongstad refinery, Statoil has an ongoing research activity that is more than ten years old. In this research activity the modified ARCO pilot unit at Chalmers has been used for the pilot tests using a North Sea long residue as feedstock. For a short contact time residue catalyst, the accessibility to the acidic sites and the density of these sites is very important. This has been indicated by measuring the surface areas and the pore size distributions of the catalysts tested. All investigations in the pilot unit have been performed at constant coke yield. The results indicate that optimum gasoline yields are obtained when the zeolite surface area of the catalyst is as large as possible and when the matrix surface area is as small as possible, but not below a certain minimum value that is determined by the feed used. Pilot unit results have indicated that most of the precracking of the large feed molecules takes place on the mesopores surface area of the catalyst. More over it is also necessary to have some areas in the macropores for cracking the very large metal-containing feed molecules. The results indicate that the two different groups of catalysts tested showed the same naphtha trends with respect to variations in catalyst surface parameters. 1. INTRODUCTION The residue FCC unit at Statoil's refinery at Mongstad, Norway, was started in 1989. The unit, a UOP RCC design, had originally a design capacity of 250 t/h [ 1], but the rated capacity has gradually been increased to 325 t/h [2]. The feed to the unit is 100 % atmospheric residue (375 ~ mainly from the North Sea. Originally the RFCC unit was designed for atmospheric residues with API gravity of 17-21 ~ UOP K-factor of 11.69-11.85, 4.3-5.0 wt% Conradson Carbon and 7-18 ppm of metals [ 1]. The high boiling atmospheric residue feed to the RFCC unit challenges the catalyst more than conventional vacuum gas oil feeds. One challenge is that metals present in the feed are continuously deposited on the catalyst. The feed metals cause the metal levels on the residue equilibrium catalyst to be much higher than on a vacuum gas oil equilibrium catalyst. In order to prevent deactivation of the residue catalyst, it must have high tolerance towards deposited metals. Another challenge to the catalyst is the high level of coke formed by the high boiling

228 residue feed. The high coke yield may upset the regenerator heat balance and the water partial pressure in the regenerator. To handle coke-related problems the regenerator has to be designed properly and the catalyst must have a high hydrothermal stability as well as the ability to withstand high regenerator temperatures. To find the optimum catalyst for the RFCC unit at the Mongstad refinery, Statoil has an ongoing test program for residue FCC catalysts for more than ten years [3]. The driving force for this program has been to optimize the naphtha and distillate components and to improve their quality [4]. To realistically test the catalysts, North Sea (375 ~ atmospheric residues from the Mongstad refinery have always been used as feeds in the test program, both in MAT and in ARCO pilot units. This is very important because the ranking of the different catalysts has been shown to be dependent on the feed used [3,5]. Another challenge for the catalyst is the possibility of transport limitations on the cracking of feed to products within the catalyst. Diffusion of the feed molecules into the catalyst depends on the size of the molecule [6]. The larger the feed molecules are, the more difficult it is for the feed molecules to diffuse into the catalyst pores. Large feed molecules cannot directly enter the zeolite super cage whose opening is only 7.4 A in diameter [7]. These large molecules must be precracked on the matrix surface area first [8]. As soon as molecules small enough to enter into the zeolite structure have been formed, they are selectively cracked to naphtha. Various residues, however, have different molecular shapes and may therefore make different demands on the catalyst. Molecules in an aromatic residue have, for instance, a more voluminous shape than molecules in a paraffinic residue [9]. As a result it is important to match the pore structure of the catalyst with the molecular size of the residue feed used. This will influence the accessibility to the active sites on the matrix. The accessibility should be as good as possible but this is not enough for optimum performance of the selected catalyst [ 10]. The strength of the acidic sites also must be optimized. However, it is difficult to measure the strength of the acidic sites in an FCC catalyst matrix though many methods have been proposed [ 11,12]. One method used in practice has been published by Ashland [ 12]. So far most of the efforts in selecting optimum residue catalysts have been focused on the accessibility to the acidic sites. The total surface area, as well as the zeolite and matrix surface areas of the catalyst are easy to measure. The pore size distribution of the catalyst can also be measured and the mesopore surface area and the mesopore volume can be calculated from the data collected. Valuable information about the catalyst can then be obtained by plotting the yields as a function of the zeolite to matrix surface area ratio (7_/M) [ 13]. In an earlier investigation it was shown that the Z/M surface area ratio should be as large as possible for optimum performances of the catalyst, i.e. maximum naphtha yield, when paraffinic North Sea long residue was used as feed to the RFCC unit [4]. The results indicated that at pilot unit conditions, maximum naphtha yield was obtained when the zeolite surface area was as large as possible and the matrix surface area as small as possible. The behavior of the catalysts indicated, however, that for proper function in the pilot unit, the matrix surface area of the catalysts could not be below a certain minimum value. In the present work we have studied how two different groups of commercial catalysts respond to changes in the Z/M surface area ratio. We also show that valuable additional information about the catalysts can be achieved by plotting the yields as a function of the zeolite and matrix surface areas. The usefulness of

229 the total surface area and the pore volume for optimization of the catalysts have also been investigated.

2. EXPERIMENTAL 2.1. Feed The feed used in this investigation was a paraffinic North Sea atmospheric residue (375 ~ see Table 1. This North Sea long residue is representative for the feedstock to the RFCC unit at the Mongstad refinery. Table 1

). Density, kg/1 CCR, wt% Aniline point, ~ Sulfur, wt% Nickel, ppm Vanadium, ppm Nitrogen (total), ppm

0.922 2.8 89.7 0.48 1.7 2.7 2000

Distillation 0% 5% 10% 20% 30% 40% 50% 60% 70% 75%

~ 256 341 373 412 438 458 481 508 544 567

2.2. Metal Impregnation and Deactivation of the Catalysts The catalysts were first calcined, 600 ~ 2 hours, and then impregnated with nickel and vanadium naphthenates according to the Mitchell method [ 14]. The total metal level was 3000 ppm. The nickel to vanadium ratio was 2 to 3. Steam deactivation was performed with 100 % steam at 760 ~ for 16 hours.

2.3. Catalysts characterization Two groups, A and B, of commercially available catalysts were used in this investigation. Catalysts in group A were supplied by two vendors but showed similar characterization data and product yields at an overlapping point [4]. Catalysts in group B were from one vendor. All catalysts characterization were performed on metal impregnated and steam deactivated catalysts except for the pore volume measurements that were performed on fresh catalysts. The total surface area (B.E.T. area) and the pore size distribution were measured by a Micromeritics ASAP 2010 unit. The matrix surface area and the zeolite surface area were calculated by the t-plot method [15,16] and the mesopore area was caculated from the adsorption pore size distribution measurement. The mesopore range used was 3 0 - 350 A. For data see Table 2. The pore volumes of the catalysts were determined by adding water in small portions to a sample of the catalyst until the catalyst fluidity was lost. The catalyst pore volume was defined as the amount of water added per gram of catalyst sample. The measurements of the pore volumes were performed on fresh catalysts. For data see Table 2.

230

Zeolite unit cell size (UCS) was determined by X-ray diffraction (XRD) according to the ASTM-D-3942-80 standard at SINTEF (SINTEF Applied Chemistry, P.O.Box 124 Blindern, N-0314 Oslo, Norway). For data see Table 2. Rare earth content (RE) was determined by X-ray fluorescence (XRF) at SINTEF. For data see Table 2. Table 2 ....Catalysts characterization Catalyst Pore Mesopore Volume area cc/g m2/g 30-350 A A1 0.31 21.6 A2 0.40 22.0 A3 0.35 23.1 A4 0.37 24.7 A5 0.22 70.5 A6 0.30 75.5 B1 0.29 29.2 B2 0.38 64.3 B3 0.45 89.9 .

.

.

.

.

.

.

.

.

.

.

.

.

.

.

.

.

.

BET area m2/g

.

.

.

.

151 144 155 164 138 193 142 115 171 .

.

.

.

Matrix area

.

.

.

RE wt%

m2/g

Zeolite area m2/g

Unit Cell Size /k

30 35 36 40 75 86 46 73 92

122 109 119 124 63 108 97 42 79

3.31 1.64 2.37 2.34 0.79 0.59 1.62 2.12 2.07

24.31 24.28 24.26 24.31 24.31 24.31 24.31 24.33 24.29

.

.

.

.

.

.

.

.

.

.

.

2.4. Pilot Plant Unit Test The pilot plant unit tests were performed in a modified ARCO pilot unit at Chalmers described previously [3,17]. The reactor temperature was 500 ~ and the regenerator temperature 700 ~ Tests were run at 4 different catalyst to oil (C/O) ratios for each catalyst. Flue gases and product gases (C4-) were continuously collected for analysis with a refinery gas analyzer. Liquid products were analysed by simulated distillation. The breakpoints used were C5 to 216 ~ for gasoline, 216 ~ to 344 ~ for LCO and 344 ~ for HCO. After calculation of mass balance, product yields and conversion, yields as a function of conversion were established by linear regression for tests with a mass balance between 95 and 99 wt%. 3. RESULTS AND DISCUSSION How does one select FCC catalysts for evaluation? Several parameters could be used to predict the catalysts performance with paraffinic North Sea long residue. One such parameter is the pore volume of the catalyst. The literature indicates that an atmospheric residue catalyst should have a pore volume not significantly less than 0.30 cc/g [ 18]. Figure 1 shows that catalysts with pore volumes as low as 0.20 cc/g and high matrix surface areas performed well with North Sea long residue. However, one catalyst with a pore volume of 0.34 cc/g was not able to crack the North Sea long residue [ 19]. The pore volume and the matrix and mesopore surface areas correlated but not the pore volume and the zeolite surface area. This seems quite obvious because the mesopore and macropore volumes represent most of the volume within the catalyst particle. As a result it is very difficult to use the pore volume of a single catalyst to predict its performance as an atmospheric residue catalyst due to the lack of correlation

231 0.5 A

~0.4 i

i

0.3

o 0.2 m ~0.i o

0

25

50

75

i00

IV~trixSurfac~Area (m2/g)

Figure 1 Pore volume as a function of matrix surface area ( II = Catalysts group A, A = Catalysts group B) between the pore volume and the zeolite surface area. No correlation was found between yields and the total surface areas.Since the total surface area consists of both the zeolite and matrix surface areas it is not a useful parameter for optimizing catalysts. The zeolite-to-matrix surface area ratio (Z/M), however, has shown to be useful for optimizing catalysts both for vacuum gas oils [13] and North Sea long residues [4]. Regression of the yields as a function of the zeolite surface area and the matrix surface area also gives valuable additional information about the catalysts. When testing FCC catalysts, it has been common to compare catalysts at a constant conversion bases. However, in commercial operation a FCC unit in most cases is operated at a relatively constant coke level. As a consequence we have chosen to compare results from the ARCO unit at constant coke basis in this work. A constant coke yield of 6 wt% in the ARCO unit was selected for the regression analysis. The catalysts tested covered a wide range of RE-contents as can be seen in Table 2. However, there was little variation in UCS between different catalysts, and no relationship between RE-content and UCS could be found, see Table 2. Due to the little differences in UCS between the catalysts, effects of RE and UCS are not discussed in this work. The results showed that the naphtha yield, see Figure 2a, increased when the Z/M ratio increased. This indicated that the catalyst should have a large zeolite surface area and a small matrix surface area for optimum naphtha yield when processing North Sea long residues. This was also confirmed by plotting the naphtha yield as a function of the zeolite surface area, see Figure 2b. The figure also visualizes that it is favorable to have a high zeolite surface area when North Sea long residues are cracked. This seems quite obvious because for optimum cracking of the feed the zeolite has to be able to further crack all the precracked products which are small enough to enter the zeolite channels. When plotting the naphtha yields as a function of the matrix surface area, see Figure 2c, it is shown that the matrix surface area should be as small as possible for maximum naphtha yield and for maximum coke selectivity. However, the matrix surface area must have a minimum size in order to be able to precrack the large molecules present in the feed [4]. This is in good agreement with the earlier vacuum

232

gas oil investigation and with other data in the literature [ 13,20,21 ]. As can be seen in Figures 2a, 2b and 2c, different groups of catalysts showed the same tendency but had different regression lines. 55

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Figure 2a Naphtha yield (wt%) at constant coke as function of the zeolite to matrix surface area ratio 55

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Figure 2b Naphtha yield (wt%) at constant coke as function of the zeolite surface area

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Figure 2c Naphtha yield (wt%) at constant coke as function of the mesopore and matrix surface areas (11 = Catalysts group A, A = Catalysts group B) A more selective cracking to naphtha often results in a decrease in the LCO yield. This was also the case when LCO yields were plotted as a function of the Z/M surface area ratio, see Figure 3a. When 7JM increased, the LCO yield decreased for both groups of catalysts. The same was observed for LCO yields as a function of the zeolite surface area, see Figure 3b. Likewise it is observed that the LCO yields increased when the matrix surface area increased, see Figure 3c. These results are in good agreement with the fact that LCO yields should decrease when the coke selectivity is increased. This is also in good agreement with the observation that the cracking of the feed to naphtha was more selective when the matrix

233

surface area decreased [4,20]. Moreover catalyst B 1 showed difficulties to crack the residue part of the feed and some of its high naphtha yield was obtained by cracking LCO to naphtha. This could explain the somewhat low LCO yield for this catalyst. 20..

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Z e o l i t e t o M ~ t r i x Surface Area P a t i o

Figure 3a LCO yield (wt%) at constant coke as a function of the zeolite to matrix surface area

Figure 3b LCO yield (wt%) at constant coke as a function of the zeolite surface area

.-.18 dO

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Figure 3c LCO yield (wt%) at constant coke as a function of the mesopore and matrix surface areas (11 = Catalysts group A, A = Catalysts group B)

The HCO yields at constant coke decreased when the conversion was increased by increasing the coke selectivity of the catalyst. Thus the HCO yield decreased when the Z/M surface area ratio increased for catalysts of group A, as shown in Figure 4a. The HCO yield at constant coke also decreased when the zeolite surface area increased, see Figure 4b, for Group A catalysts. Figure 4c shows that the HCO yield at constant coke increased when the matrix surface area increased for catalysts of group A. For catalysts of group B the HCO yield at constant coke slightly decreased when the matrix surface area increased. This indicated that

234

catalysts of group A and group B had different matrices. While a high Z/M surface area ratio was preferable for catalysts of group A, the opposite was the case for catalysts of group B. A possible explanation might be that catalysts of group B needed a larger matrix surface area to be able to crack the paraffinic North Sea long residue efficiently.

15

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Figure 4a HCO yield (wt%) at constant coke as a function of the zeolite to matrix surface area ratio

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Figure 4b HCO yield (wt%) at constant coke as a function of the zeolite surface area

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Figure 4c HCO yield (wt%) at constant coke as a function of the matrix surface area (11 = Catalysts group A, A = Catalysts group B)

The LPG yield usually increases when the conversion or the coke selectivity of the catalyst increases. For the group A catalysts this was the case, see Figure 5a, when the ZAVl surface area ratio increased. This result is also supported by the observation that the LPG yield increased when the zeolite surface area increased, see Figure 5b, and that the LPG yield decreased when the matrix surface area increased, see Figure 5c. For the group B catalysts, however, the LPG yield was almost constant when the Z/M ratio increased at constant coke, see Figure 5a. In the same way the LPG yield also decreased for this group of catalysts when

235

the matrix surface area increased, see Figure 5c, but as can be seen in Figure 5b the LPG yield decreased slightly when the zeolite surface area increased for the group B catalysts.

25,

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Figure 5a LPG yield (wt%) at constant coke as function of the zeolite to matrix surface area ratio

50

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Figure 5b LPG yield (wt%) at constant coke as a function of the zeolite surface area

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Figure 5c LPG yield (wt%) at constant coke as a function of the matrix surface area (11 = Catalysts group A, ik = Catalysts group B)

A larger matrix surface area often means an increased dehydrogenation activity for the matrix and as a consequence an increased yield of hydrogen. This was also observed in this study when the matrix surface area increased, see Figure 6c. The dehydrogenation effect is of major importance and may be one reason to why the hydrogen yield declined when the Z/M surface area ratio increased, see Figure 6a. The declined hydrogen yield might also depend on an increased catalyst coke selectivity.

236 0.4

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Figure 6a Hydrogen yield (wt%) at constant coke as a function of the zeolite to matrix surface area ratio

0.4 dO

~0.3

"~0.2

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Figure 6c Hydrogen yield (wt%) at constant coke as a function of the matrix surface area (11 = Catalysts group A, A = Catalysts group B)

Parameters other than the Z/M surface area ratio, need to be optimized for optimal catalyst performance. The matrix itself has to be optimized, and the best ratio between the mesoporeand macropore surface areas must be found. For North Sea long residues most of the precracking takes part on the mesopore surface area [ 19,22] but it has been shown that it is also necessary to have enough macropore surface area present in the catalyst [ 19]. To illustrate this requirement the mesopore surface area values are included to show the connection with the naphtha and LCO yields, see Figures 2c and 3c. 4. CONCLUSIONS In order to find an optimum catalyst for the RFCC unit, the results show that it was necessary to optimize each group of catalyst separately with the proper feed. The zeolite to

237 matrix surface area ratio of the catalysts was used for optimization. When cracking paraffinic North Sea long residues, the zeolite to matrix surface area ratio of the catalyst should be as large as possible for maximum naphtha yield. Additional information was also generated by plotting yields as a function of the zeolite surface area and of the matrix surface area separately. Results have shown that the matrix surface area should not be below a minimum value that is determined by the group of catalyst and feed used. Moreover there is a correlation between the catalyst pore volume and the matrix surface area but not between the pore volume and the zeolite surface area. Furthermore no correlation was observed between the catalyst total surface area and cracked products yields. ACKNOWLEDGEMENT The authors are greatful to Statoil for the permission to publish this paper. REFERENCES

10. 11. 12. 13. 14. 15. 16. 17. 18. 19. 20.

L.R. Aalund, Oil Gas J., 88(11) (1990) 33. D. Gledhill, J. Pedersen, "Operating experience with the new VSS riser termination technology", GRACE Davison FCC Technology Conference Lisbon Portugal, September 1-4, 1998. S.-I. Andersson, T. Myrstad, Appl. Catal. A., 159 (1997) 291. S.-I. Andersson, T. Myrstad, Oil Gas Europ. Mag., 23(4) (1997) 19. L.T. Boock, X. Zhao, ACS Symposium Preprints, Div. of Petrol. Chem., 41 (2) (1996) 367. P. O'Connor, A.P. Humphries, ACS Symposium Preprints, Div. of Petrol. Chem., 38(3) (1993) 598. A. Humphries, J.R. Wilcox, Oil Gas J., 87(6) (1989) 45. P. O'Connor, E. van Houtert, Ketjen Catal. Symposium 1986, Scheveningen, The Netherlands, Paper F-8. G.W. Young, J. Creighton, C.C. Wear, R.E. Ritter, NPRA, San Antonio, 29-31 March 1987, AM-87-51. B.A. Lerner, Hydrocarbon Engineering, March 1998, 26. A. Corma, V. Fornes, F. Rey, Zeolites, 13 (1993) 56. S. Alerasool, P.K. Doolin, J.F. Hoffman, in: M.L. Occelli, P. O'Connor (Eds.), Chemical industries, 74 (1998) 99-110. C.C. Wear, R.W. Mott, NPRA, Annual Meeting, March 20-22, 1988, AM-88-73. B.R. Mitchell, Ind.Eng.Chem.Prod.Res.Dev., 19 (1980) 209. M.F.L. Johnson, J. Catal., 52 (1978) 425. ASTM D 4365-85. S.-I. Andersson, J.-E. Otterstedt, "Catalytic Cracking of North Sea Resid", presented at Katalistiks 8th Annual FCC Symposium, Budapest (1987), Paper 21. M.M. Mitchell Jr., J.F. Hoffman, H.F. Moore, in: J.S. Magee, M.M. Mitchell (Eds.), Stud. Surf. Sci. Catal., 76 (1993) 302. S.-I. Andersson, T. Myrstad, AIChE, Spring Meeting, New Orleans, 9-11 March, 1998, Paper 31 d. K. Rajagopalan, E.T. Habib Jr., Hydrocarbon Processing, 71 (9) (1992) 43.

238 21.

22.

H. Haave, P.A. Diddams, "FCC catalyst technology for short contact time applications", GRACE Davison FCC Technology Conference Lisbon Portugal, September 1-4, 1998. J.S. Magee, W.S. Letzsch, ACS Symposium Series, 571 (1994) 349.

Studies in Surface Science and Catalysis 134 M.L. Occelli and P. O'Conner (Editors) 9 2001 Elsevier Science B.V. All rights reserved

239

An experimental protocol to evaluate FCC stripper performance in terms of coke yield and composition Colin E. Snape *a, Youva R. Tyagi a, Miguel Castro Diaz a, Shona C. Martin a, Peter J. Hall, a Ron Hughes b and C.L. (Arthur) Koon b University of Strathclyde, Department of Pure and Applied Chemistry, Glasgow G1 1XL Scotland, U.K.

a

b University of Salford, Chemical Engineering Unit, Salford M5 4WT, U.K. * Current address: University of Nottingham, School of Chemical, Mining and Environmental Engineering, University Park, Nottingham NG7 2RD, UK Tests have been conducted in a microactivity test (MAT) and a fluidised-bed reactor to develop an experimental protocol to determine how the yield and composition of coke and the associated catalyst surface area vary as a function of stripper conditions in fluid catalytic cracking (FCC). In both reactors, the use of rapid quenching has allowed the relatively short stripping times encountered in FCC units to be simulated. Low sulphur vacuum gas oils (VGO) with a low metal equilibrium catalyst (E-cat) were used for stripping periods of up to 20 minutes. Significant variations occur in the structure of both hard and soft coke during stripping. Although the hard coke becomes more highly condensed with prolonged stripping, the surface area reduction by the hard coke remains fairly constant for stripping periods in excess of c a . 5-10 minutes and is small (10 m 2 g-l) in relation to the loss of surface area from the soft coke. The use of about 70 g of catalyst in the fluidised-bed provides sufficient sample for demineralisation to recover the hard coke for 13C NMR analysis after the initial extraction of the soft coke. Indeed, a further pool of soft (chloroform-soluble) coke is physically entrapped within the catalyst pore structure and is only released after demineralisation. In fact, this second soft coke fraction is more highly aromatic than the first and ultimately controls the final coke yield. The structural information obtained has been used to formulate a model for the stripping process where the soft coke II fraction undergoes cracking in competition with coke formation and evaporative removal from the catalyst.

1. INTRODUCTION Coked catalyst in a fluid catalytic cracking (FCC) unit first passes to a steam stripper to remove residual volatiles and then it is transferred to the regenerator vessel where the coke is burned in a stream of air. Since the catalyst acts as a heat-transfer medium with the heat liberated by coke combustion providing the energy for the endothermic cracking reactions, the

240 coke selectivity can markedly affect a unit's profitability. It is now generally accepted that, as well as being formed via the actual cracking reactions, coke also arises from the thermal and metal-mediated (Ni/V) reactions, together with the entrained products which are symptomatic of incomplete stripping and can contribute to the overall level of coke [1]. The entrained products increase the hydrogen content of the coke and the additional air requirement gives rise to excessively high temperatures in the regenerator and additional steam, which in turn contribute significantly to the deactivation of FCC catalysts. The highly dynamic situation within a FCC unit is further complicated by the thermal reactions, which occur in the stripper section and can affect the yield and structure of the insoluble (hard) coke. Although the deactivation of FCC catalysts via coke deposition has been the subject of much investigation since the 1940s [2,3], there is still a lack of knowledge on the contributions of the thermal, catalytic and metal-mediated mechanisms outlined above to the overall level of coke formation. This situation has arisen from the inherent difficulties of characterising the structure of insoluble cokes at the low concentrations encountered in FCC Units. Indeed, to facilitate coke characterisation, fundamental studies thus far on the ultra-stable (US) type Y zeolites have often involved excessively high levels of carbon deposition in relation to normal FCC operation [4]. Moreover, the behaviour with small molecules, where coke can be formed directly (catalytically) from the reactant within the zeolite framework, is quite different to that observed with heavy feedstocks where the yield of coke (typically c a . 1%) is independent of catalyst/oil ratio [5]. To determine how coke yield and composition and catalyst surface area vary as a function of stripper conditions, we have conducted tests recently in a microactivity test (MAT) reactor using a vacuum gas oil (VGO) feed [6]. It was found that significant structural variations occur for both the hard (chloroform-insoluble) and soft (chloroform-soluble) coke as stripping progresses. Although the hard coke became more highly condensed with prolonged stripping, the surface area reduction by the hard coke remains fairly constant for stripping periods over c a . 5-10 minutes. To provide larger and more representative soft and hard coke samples for characterisation, particularly by ~3C NMR, stripping tests have now been conducted in a fluidised bed reactor. Further, the experimental protocol now incorporates rapid quenching that allows the relatively short stripping times encountered in FCC units to be simulated. A nominally low metal equilibrium catalyst (E-cat) has been used here with stripping periods of up to 20 min. Surface areas have been determined before and after removal of the soluble (soft) coke with chloroform. Further, hard coke concentrates have been prepared by demineralisation with hydrofluoric and hydrochloric acids for characterisation by solid state 13C nuclear magnetic resonance (NMR). This approach was successfully demonstrated for FCC refinery catalysts [7,8] where the cokes were found to be highly aromatic in character (carbon aromaticities > 0.95), but differences in feedstock composition were still reflected in the structure of the cokes. The aim of this contribution is to present the latest findings from a unique experimental approach which is being developed to determine how the yield, composition and spatial distribution of FCC coke as a function of stripping conditions. The goal is then to use the experimental data to construct a kinetic model to rationalise the effects of variables, notably catalyst composition in terms of matrix characteristics and gas flow rate, on stripper performance.

24I

2. EXPERIMENTAL 2.1 Stripping Tests The stripping tests were conducted using two low sulfur VGOs (%S = 1.3%, with the second one being more waxy than the first with a higher atomic H/C ratio (1.75 cf. 1.60) with a low metals E-cat whose characteristics are summarised in Table 1. A standard MAT reactor (was used of diameter 12.5 mm, in which the catalyst was in the form of a packed bed [6]. VGO was delivered at a fixed rate via a water-jacketed syringe pump for a pre-determined time, in order to obtain the catalyst to oil ratio required. Following the oil injection, the catalyst bed which had been maintained at 520~ was swept with nitrogen (flow of 30 ml min -1) for the desired stripping period. After the required stripping period, the reactor tube was rapidly disconnected and removed from the furnace. Table 1 Characteristics of the low metals equilibrium catalyst. Parameter Amount Re203 0.82% A120~ 36.0% Ni 826 ppm V 855 ppm MAT yield 65% Hydrogen Factor . 2.0 The fluidised bed reactor was of 52 mm diameter and was capable therefore, of handling a much larger charge of catalyst for subsequent analysis. For each test, 70 g of catalyst (75-150 Dm particles used to prevent elutriation from the bed) was placed in the bed and the reactor was allowed to heat up to 520~ under the flow of the fluidising nitrogen (1.5 dm 3 min-1). A further flow of 1.5 dm 3 min 1 of nitrogen was used to assist feeding the viscous VGOs into the centre of the bed, 15 g being fed over a period of 40 s. The conditions were chosen to match closely as possible those used in the stripping tests conducted in the MAT reactor with the same catalyst to oil mass ratio (5:1) being adopted. In both reactors where stripping periods up to 30 min. were employed, zero stripping time was estimated from the superficial gas velocities through the bed. For each reactor, a test was carried out with the flow rate reduced by hals

2.2 Catalyst work-up and coke characterisation After stripping, the catalysts from the fluidised-bed tests were recovered and extracted with chloroform under reflux using 10 ml of chloroform per gram of sample. The chloroformsolubles (soft coke I) were recovered and characterised by 1H NMR and size exclusion chromatography (SEC). The chloroform-extracted catalysts from the fluidised-bed reactor were then demineralised with HC1/HF to prepare the coke concentrates (50-100 rag) for solidstate 13C NMR analysis [7,8]. After demineralisation, the initial hard coke concentrates were extracted in chloroform to remove any soft coke that had been physically entrapped within the catalyst matrix (soft coke II). Carbon, hydrogen and nitrogen contents of the initial coked catalysts and the soft and hard coke concentrates were determined using a Perkin-Elmer 2400

242 analyser and sulphur contents were measured using the Sulphazo III method. To deduce the proportions of the catalyst carbon accounted for by the second soft coke fractions the yields and C contents of these fractions were used. BET surface area measurements were carried out on the catalysts before and after removal of the first soft coke fraction using a Micromeritics ASAP 2000 apparatus. A Bruker 250 MHz instrument was used to obtain the 1H NMR spectra of the soft cokes in chloroform-d. SEC was carried out to estimate the number and weight average molecular masses (M n and Mw) of the soft coke fractions based on polystyrene standards, a mixed bed PL gel column being employed with RI detection and chloroform as the eluting solvent.

3. RESULTS AND DISCUSSION 3.1 Trends in coke yield and surface area Figure 1 compares the total carbon contents of the stripped catalysts recovered from the fluidised-bed and MAT reactors using the more waxy VGO investigated with the reactors being removed rapidly from their respective furnaces to quench the coke. The agreement between the two systems is reasonable given the differences that exist in scale and gas flow rate. Compared to the previous data reported earlier for the MAT reactor with much slower quenching [6], the carbon contents at short stripping times (< 3min.) are higher. In both reactors, the carbon content decreases to 1.5% after 5 min., but there is a further, albeit small, decrease to 1.3% after 20 minutes stripping.

Figure 1. Comparison of carbon contents with stripping time for the MAT and fluidised-bed reactors using the more waxy V GO feed.

243 Figure 2 presents the variations in carbon contents for the stripped E-cat from the fluidisedbed reactor using the less waxy VGO before and after chloroform extraction. As anticipated, the plot reveals that most of the initial soft coke (that extractable by chloroform) is removed fairly early in the stripping process with the hard coke content reaching a constant level after about 10 minutes. However, the quantities of soft coke recovered for the low metal catalyst after 10 minutes still accounted for 9% of the total carbon and this decreased to 2% after 20 minutes. The trend in Figure 2 suggests that little of the easily extractable coke (soft coke I) is carbonised to form hard coke. The final hard coke content of ca. 0.9% is less than that obtained using the more waxy VGO investigated (Figures 1 and 2).

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Figure 2. Variation of carbon content before and after removal of soft coke I with stripping time for the fluidised-bed reactor using the less waxy VGO. Figure 3 presents the BET surface areas before and after chloroform extraction of the recovered catalysts (removal of soft coke I) from the fluidised-bed tests using the less waxy VGO, together with the surface area of the as-received E-cat. Before stripping commences, ca. 40% of the surface area is lost with the loss being mainly from the micropores with the relatively small proportion of mesopores (ca. 15% of the total surface area) as observed by BET being unaffected. The surface area after chloroform extraction reaches a fairly constant value after stripping for about 10 minutes. Indeed, the chloroform-insoluble coke remaining after 5 minutes is responsible for a relatively small proportion of the initial loss of surface area (only ca. 25%, 15 out of 60 m 2 g-l, Figure 3). The overall distribution of coke carbon between the different fractions as a function of stripping time is presented in Figure 4 for four of the stripping times investigated in the

244 fluidised-bed reactor. Remarkably, at short stripping times, much of the soft coke is physically entrapped within the catalyst (soft coke II) with the hard coke accounting for only

Figure 3. Distribution of coke carbon between soft coke I and II and hard coke as a function of stripping time with stripping time for the less waxy VGO in the fluidised-bed reactor.

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~ 1

2

. . . . . . . . 3 4 5 6 7

8

t

~

,

!

t

~

J

I

t

i

t

t

I

9

10

11

12

13

14

15

16

17

18

19

20

21

Stripping Time (Minutes)

Figure 4. Variation of BET surface area with stripping time using the less waxy VGO in the fluidised-bed reactor.

245 c a . 0.5% C (w/w of catalyst). However, by 3 min., the hard coke content has reached its limiting value of 0.9% (Figure 4). Thus, approximately one third of the entrapped soft coke (II) is carbonised further to form insoluble hard coke. Intuitively, when the gas velocity falls, this fraction is expected to increase, because it is harder for the entrapped soft coke to be volatilised. In fact, increases of c a . 0.2% carbon were also observed for the less waxy VGO with both the MAT and fluidised bed reactors by decreasing the flow rate by a factor of 2 (stripping time of 2 min.). Thus, as a function of both the flow rate of the stripping gas and the catalyst properties, the fate of the entrapped soft coke is probably the major controlling factor on the final coke yield as further exemplified in the following sections.

chlorofo rm-d

SOFT

SOFT

COKE-

II

COKE

- I

PPM

Figure 5. ~H NMR spectra of soft coke I and II fractions from the catalyst stripped for 10 min. in the fluidised-bed reactor (TMS is the internal standard, tetramethylsilane).

3.2 Coke Composition It has been shown previously, that initially the soft coke (ca. 1% carbon- Figures. 2 and 4) is highly aliphatic and contains small aromatic groups [6]. Therefore after relatively short

246 stripping times, soft coke I resembles unconverted feed. Figure 5 shows the 1H NMR spectra of both soft coke fractions from the catalyst stripped for 10 minutes in the fluidised bed reactor. The aromatic hydrogen peaks occur between 6.5 to 8.5 ppm while the dominant aliphatic peak at 1.25 ppm arises from long alkyl chains. The spectra show that soft coke II is considerably more aromatic than soft coke I. Further, two ring and larger systems are present in much higher concentrations in soft coke II as indicated by the more intense bands between 7.2 and 8.5 ppm. However, long alkyl chains do survive in significant proportions. However, these are expected to crack to yield significant quantities of lower alkanes/alkenes as stripping proceeds. The molecular masses of the soft coke I fractions show no systematic variation with stripping time (estimated as weight average Mw in the range 500-600 from SEC measurements) and are similar to those of the initial VGO (Mw of 500). Table 2 Atomic H/C ratios and carbon aromaticities for the hard coke concentrates from the fluidisedbed reactor before and after removal of soft coke II (less waxy VGO). Stripping time Initial concentrate Final concentrate

wc

(,r

WC

10 sec 0.85 0.62 0.63 3 min 0.82 0.63 0.56 20 min 0.60 0.88 0.59 * Value of 0.91 obtained from the more quantitatively reliable technique [7].

250

200

150

100 PPH

50

L_~ 0.86* 0.88 0.90 single pulse excitation

0

-50

247 Figure 6. CP/MAS 13C NMR spectra of hard coke concentrates after removal of soft coke II obtained from the catalysts stripped in the fluidised-bed reactor using the less waxy V GO. Table 2 presents the atomic H/C ratios and aromaticities of the hard coke fractions before and after removal of soft coke II and Figure 6 shows the CP ~3C NMR spectra of the final coke concentrates obtained from the 10s, 3 min. and 20 min. stripping runs in the fluidised-bed reactor. The aromatic carbon peak is centred at 130 ppm on the left-hand side of the spectra while the broader aliphatic peak centred at 20-30 ppm represents the aliphatic fraction. The aromaticities of the initial hard coke concentrates were found to be extremely low at very short stripping times but the aliphatic carbon contents decrease markedly with increasing stripping times. However, after extraction of soft coke II, the values for the final hard coke concentrates are c a . 0.9 and are comparable to those of the two chloroform-insoluble fractions examined from the MAT reactor at longer stripping times where little entrapped soft coke will remain. This level of discrimination against aromatic carbon is similar to that encountered previously with FCC cokes [7,8]. Clearly, the high aromaticity of the final hard coke formed before stripping actually commences (0.5% C at 10 s, Figure 4) suggests that the volatile yields should be extremely low. In contrast to the hard coke, the entrapped soft coke (II) still possesses significant aliphatic character (see ~H NMR spectra of soft coke I and II after 10 min. stripping, Figure 5) that is responsible for the low aromaticities of the initial coke concentrates (Table 2) and which should result in significant quantities of cracked hydrocarbon gases and gasoline from stripping (as for soft coke I). However, the soft coke II fractions do have carbon aromaticities of 0.4-0.5 (estimated from the proportion of aromatic hydrogen determined by 1H NMR and their atomic H/C ratios). Thus, from a structural standpoint, soft coke II will produce vastly more hard coke than the easily extractable soft coke I. Indeed, the fine balance between long chain alkyl and polycyclic aromatic moieties, is consistent with the volatile yields obtained from the soft coke II being the controlling factor on the amount of coke that can be removed by stripping as indicated by the variation in coke yields reported in the previous section as a function of changes in sweep gas velocity. The carbon skeletal parameters derived from the solid-state 13C NMR spectra of the coke concentrates obtained at 5 and 15 minutes previously from the MAT reactor [7] showed that although the aromaticities are similar, the average ring size increases from 2 to 3 rings in going from 5 to 15 minutes stripping. Further, the hard coke obtained from the VGO is considerably less condensed than that found from a vacuum residue feed in an actual FCC unit [6], providing further evidence for differences in feedstock composition still being reflected in the structure of the cokes. The structural changes occurring within the hard coke do not appear to have any influence the surface area of the catalyst, which after removal of the soft coke, remains fairly constant during prolonged stripping (Figure 3).

3.3 Modelling the stripping process The experimental evidence obtained thus far is used here to develop a simple kinetic and mass transfer model for the stripping process. This model will be improved once information on product yields has been obtained to add to that reported here on the distribution and structure

248 of the coke. Figure 7 illustrates that the overall loss of carbon from the catalyst follows firstorder kinetics reasonably well for the conditions used in the MAT and fluidised bed reactor. In terms of the three forms of coke identified here, namely easily extractable and physicallyentrapped soft coke and chloroform-insoluble hard coke, it is assumed that, due to its highly aromatic character, the hard coke makes little contribution to the carbon lost and the aromatic moieties become more condensed as stripping proceeds. Since soft coke I resembles unconverted feed, the majority should crack readily to yield light alkanes/alkenes and alkylbenzenes as stripping proceeds. However, as indicated by the PAH moieties formed in relatively small quantities at long stripping times, just like the initial feed during cracking, a relatively small proportion of the easily extractable soft coke can form hard coke, particularly at low sweep gas velocities where the PAHs are less likely to escape via evaporation.

//"

4.5

3.5 ~_ 3 _~_

O

2.5 1.5 1

0.5 0

j 0

2

4

6

8

12

10

14

16

Stripping Time (Minutes)

Figure 7. First-order kinetic plot for the loss of the easily extractable coke (soft coke I) in stripping for fluidised-bed reactor using the less waxy VGO, C o being the initial soft coke I carbon.

Heavy decant oil (evaporative loss)

(--

Soft coke I

--~ kl

Light products

Soft coke II

--~ kl'

Light products

,1"k2 Hard coke As indicated earlier, the fate of the entrapped soft coke is probably the major controlling factor in determining the final coke yield. Under the experimental conditions used here for the less waxy VGO, between 30 and 100% of the soft coke II fraction has contributed to the

249 final carbon content. The trend in Figure 4 strongly suggests that at long stripping times, relatively little of the soft coke II will survive as such. Therefore, it is proposed that some of the soft coke II can crack to light products based on the fact that 50-60% of the carbon is aliphatic and, in this respect, resembles soft coke I. However, this global reaction which should have a similar rate constant to soft coke I, proceeds in competition with the formation of hard coke. Sweep gas is clearly of paramount importance and it is proposed that increasing its velocity increases the proportion of soft coke II that can escape mainly by evaporation, i.e. with relatively minimal structural alteration and should contribute to the decant oil (highest boiling fraction) in the stripped product. Overall, the model can be represented qualitatively as follows when relatively high sweep gas velocities are used so that cracking is the dominant reaction undergone by the easily extractable soft coke. Clearly, the relative contributions from evaporative and cracking processes to the removal of the physically-entrapped soft coke have still to be evaluated, but the presence of steam, as opposed to nitrogen, could have a chemical influence based on the extensive literature on the pyrolysis of coals and oil shales [ 11 ].

4. CONCLUSIONS AND F U R T H E R RESEARCH

The tests conducted in the fluidised-bed reactor with rapid quenching have allowed the variation in coke composition at short stripping times to be investigated for the first time and has provided sufficient hard coke for 13C NMR analysis after the initial extraction of the soft coke. The surface area reduction by the hard coke remains fairly constant for stripping periods over c a . 5-10 minutes. Further, this reduction of only c a . 15 m 2 gl for the low metals E-cat investigated is small in relation to the initial loss of surface area from the easily extractable soft coke. It has been found that a further pool of soft (chloroform-soluble) coke is physically entrapped within the catalyst pore structure and is only released after demineralisation. In fact, this second soft coke fraction is much more highly aromatic than the first and ultimately controls the final coke yield. For the combination of E-cat and VGO investigated here, about half of the final hard coke content is derived from this second soft coke fraction. However, this fraction increases as the nitrogen flow rate decreases. Transferring the deactivated catalyst from the fluidised bed after very short stripping times (nominally 10 s) reactor to separate fixed-bed reactors will now provide a means to obtain accurate mass balances for stripping and ascertain whether the use of steam, as opposed to nitrogen, affects the stripping process. SANS measurements in conjunction with contrast matching are in progress to probe the affect of coke on the microporous structure and the accessibility of hydrocarbons. Experiments with deuterated methanol as the matching agent have established that a significant proportion of the mesopores are inaccessible in a partially stripped equilibrium catalyst [10]. We are also conducting stripping tests in the fluidised-bed with refinery E-cats of vastly different composition in terms of metals content and the amount and apparent accessibility of the matrix.

250 ACKNOWLEDGEMENTS The authors thank the Engineering & Physical Sciences Research Council (EPSRC) for financial support (Grant Nos. GR/L57289 and 58743) and Dr. N. Gudde of BP Oil International Limited, Oil Technology Centre, Sunbury-on-Thames, Middlesex TW16 7LN, UK for supplying the equilibrium catalyst and many helpful discussions.

REFERENCES

1. P. O'Connor, P. and A.C. Pouwels, in Catalyst Deactivation, Studies in Surface Science and Catalysis, Catalyst Deactivation 88 (1994) 129 and references therein. 2. E. H. Wolfe, A. Alfani, Catal. Rev. Sci. Eng., 24 (1982) 329. 3. J.B. Butt, Catalyst Deactivation, Adv. Chem. Ser., 109 (1972) 259. 4. W.A. Groten, B.W. Wojciechowski, B.K. Hunter, J. Catal., 125 (1990) 311. 5. P. Turlier, M. Forissier, P. Rivault, I. Pitault, J.R. Bernard, in Fluid Catalytic Cracking III, Am. Chem. Soc. Symp. Ser. No. 571, (eds. M.L. Occelli and P. O'Connor), p 98 (1994). 6. C.L. Koon, R. Hughes, T.R. Tyagi, M. Castro Diaz, S.C. Martin, P.J. Hall, C.E. Snape, Proc. 2 nd International Conference on Refinery Processing, p 523 (1999) held in conjunction with the AIChE 1999 Spring National Meeting, Houston, 14-18 March 1999. 7. C.E. Snape, B.J. McGhee, J. Andresen, R. Hughes, C.L. Koon, G. Hutchings, Appl. Catal. A: General, 129 (1995) 125. 8. A.A.H. Mohammed, B.J. McGhee, J., Andr6sen, C.E. Snape, R. Hughes, in Fluid Catalytic Cracking III, Am. Chem. Soc. Syrup. Ser. No. 571, (eds. M.L. Occelli, and P. O'Connor), p 279 (1997). 9. J.D. Rocha, S.D. Brown, G.D. Love, C.E. Snape, J. of Anal. and Appl. Pyrolysis, 40-41 (1997) 91. 10. P.J. Hall, Y.R. Tyagi, C.E. Snape, S.D. Brown, M. Castro Diaz, R. Hughes, C.L. Koon, J. Calo, Ind. Eng. Chem., submitted. 11. E. Ekinci, A.E. Putun, M. Citiroglu, G.D. Love, C.J. Lafferty, C.E. Snape, Fuel, 71 (1992) 1511 and references therein.

Studies in Surface Science and Catalysis 134 M.U Occelli and P. O'Conner (Editors) 9 2001 Elsevier Science B.V. All rights reserved

251

Use of ~3C-Labelled Compounds to Probe Catalytic Coke Formation In Fluid Catalytic Cracking Carol L. Wallace a, Colin E. Snape a, Nick J. Gudde b, Graham W. Ketley band Anthony E. Fallick c a University of Strathclyde, Department of Pure and Applied Chemistry, Thomas Graham

Building, 295 Cathedral Street, Glasgow G1 1XL, UK b BP, Oil Technology Centre, Chertsey Road, Sunbury-on-Thames, Middlesex TW16 7LN, UK

c Scottish Universities Research & Reactor Centre, East Kilbride, Glasgow G75 0QU, UK In order to demonstrate how the fate of individual species can be followed during fluid catalytic cracking (FCC) and to ascertain their contribution to the coke formed, I3C labelled benzene and toluene were incorporated into two actual FCC feeds, namely a low and a high sulfur vacuum gas oil. Gas chromatography-isotope ratio mass spectrometry was used to determine the distribution of 13C label incorporated into the liquid products and, to determine the enrichment of the 13C label in the coke, conventional sealed-tube combustion was used. Extensive migration of the 13C label was evident in the liquid products, particularly in the case of the toluene-doped feed where much of the 13C label identified occurred in alkylbenzenes other than toluene as a result of transalkylation reactions. This was also confirmed by solution state 13C NMR. In addition, incorporation of approximately 0.2 % of the starting x3C label was found in the cokes for both the ~3C toluene and benzene doped feeds, providing direct evidence that some of the catalytic coke, albeit a small proportion, was derived from benzene and toluene, but the greater contribution was from toluene (labelled at only the methyl carbon).

1. INTRODUCTION Coke formation in fluid catalytic cracking (FCC) can markedly affect a unit's performance and for over fifty years, it has been the subject of much investigation. [1-3]. Coke can arise from a number of sources, namely acid catalysed, metals mediated and pure thermal r e a c t i ~ ~. Therefore, means of quantifying the respective contributions of the three main coke-forming pathways will be of significant benefit. The behaviour of small alkenes and aromatics, where carbonisation can be initiated within the zeolite framework to form catalytic coke is expected to be different to that for heavy species

252 [4], where, intuitively, the formation of metals mediated and thermal coke from the feed is expected to occur in extra-framework mesopores. In this study, 13C labelled benzene and toluene have been doped into high and low sulfur vacuum gas oils to aid in the understanding of how catalytic reaction pathways contribute to coke formation. The distributions of the labels have been monitored using gas chromatography isotope-ratio mass spectrometry (GCIRMS) for the liquid products and conventional sealed-tube combustion for the cokes. Compound specific isotope analysis using GC-IRMS is now a well established technique in the petroleum industry, coal science, environmental and biomedical research [5-10]. More specifically to FCC, Filley et al recently reported the isotopic composition of selected carbonisation products obtained from a FCC decant oil doped with 13C-enriched 4methyldibenzothiophene [11]. However, only selected compounds were identified and no overall mass balance to account for label distribution was attempted. This investigation represents the first attempt to obtain a quantitative audit into the fate of individual constituents within FCC product streams, particularly with respect to the reaction pathways leading to catalytic coke. A preliminary account of this investigation has been presented in the proceedings of the 2nd International Conference on Refinery Processing hosted by American Institute of Chemical Engineers [ 12].

2. EXPERIMENTAL

2.1 Feedstock and Product Preparation The liquid products and cokes investigated were prepared from a low and high sulfur vacuum gas oils (VGO) using a microactivity test (MAT) reactor (ASTM D3907, D5154). The low sulfur VGO was doped with 8% w/w toluene and benzene, and the high sulfur VGO one with 8% w/w toluene, but only 4% w/w benzene. The carbon in the methyl position of the toluene was labelled as opposed to all carbons in the case of the benzene (purity >99% for both the labelled compounds). Therefore, when the 0.8 g of feed used in the MAT experiments was doped with 8% toluene, the 13C enrichment equated to approximately 8 mg and, similarly for benzene, doping gave enrichments of either 30 or 60 mg excess lac. The reactor contained 4.08 g of steamed zeolite making the catalyst to feed (0.8 g) ratio equal to 5.0. The feed was delivered in 18 seconds to the reactor at 510~ before the system was flushed with nitrogen for 14 minutes. The liquid products were trapped in the collection bulb cooled with a beaker of iced water. 2.2 Analysis Compound specific ~13C measurements of the liquid products, generated from the MAT experiments, were carried out on a VG Isochrom II | GC-IRMS instrument as described by McRae et al [ 13]. A gas chromatograph is used to separate the individual components of the sample that then pass into a copper oxide combustion furnace. The resultant CO2, produced for each separated species, is continuously analysed by a single inlet, triple-collector mass spectrometer. Differences in the isotopic composition of carbon-containing substances are expressed in the conventional ~5-notation giving the permil (%0) deviation of the isotope ratio of the sample (sa) relative to that of a standard (st), i.e.

253

J

613 Csa -" (13 C/12 C)sa -- 1 x 10 3 (13 C/12 C)st

(i)

The standard commonly used is Peedee belemnite (PDB), whose 8~3C value defines 0 %0 on the 8-scale. In this initial study isotopic compositions have also been expressed on an atomic percent basis (13C %) which represents the excess of 13C above the natural abundance:

i 13Csal 1 13c%-[ / 13Csa/ ooo + R~t~~-j

+1

1.11

(ii)

Where R~t is the 13C/12C isotopic ratio of PDB (0.011237) and 1.11% is the natural abundance of 13C. Helium was employed as the carrier gas and a temperature program of 40~ (5 mins) to 320~ (5 mins) at 5~ rain -~produced the best possible separation of sample peaks. For convenience compounds were grouped, according to the classes listed in Table 1. However, it was not possible to determine isotopic values for components eluting from the column within two minutes of the injection period. The liquid products were also analysed by normal GC-MS (Varian 3400 GC linked to a Finnigan MAT TSQ70 triple quadruple MS, ionising energy 70eV; ion source temperature 150~ transfer line temperature 300~ to identify the major constituents present. Table 1 Approximate boiling point distributions and compositions of the liquid products. Group Boiling Point Product Distribution Distribution (~ 1 80 - 130 C6-C 8 alkanes, cyclo-alkanes, toluene 2 130 - 180 C9-C10 alkanes, C2-C 4 alkyl benzenes 3 180 - 220 Cll-C~2 alkanes, C4- Cs alkyl benzenes, naphthalene 4 220- 270 C13-C14 alkanes, C~-C2 substituted naphthalenes 5 270- 340 C15-C21 alkanes, substituted naphthalenes 6 270- 340 C~5-C21 alkanes,C2-C4 substituted naphthalenes 7 340 3-membered polycyclics i.e. anthracene, phenanthrene 8 340- 360 Alkyl substituted 3-ring polycyclics 9 340- 360 Alkyl substituted 3-ring polycyclics 10 360 -450 Alkyl substituted 3-4 ring polycyclics

254 Equation (iii) was used to calculate the mass of 13C incorporation into the liquid product groups. The mass of each group was determined from normal GC analysis of the liquid products doped with an internal standard (triacontane, C30H62). A Carlo Erba 4130 instrument was used, equipped with an FID detector and an SGE 25m fused silica capillary column coated with BP-1. The same temperature program used for the GC-IRMS analyses was employed.

13C Mass =

Mass Standard / Respo---~se~ S-i----andard,]x Response Area Group x 13C O~

(iii)

Unfortunately, due to excessive evaporative losses, solution state 13C NMR was not carried out on the liquid products detailed above. Therefore, a second series of samples were generated to obtain ~3C NMR data; the low sulfur feed was doped with 4 % ~3C toluene and 8 % ~3C benzene, respectively. The spectra for the whole liquid products were obtained in deuterochloroform (50% solution) with the free induction decays being collected in 32K data sets over a spectral width of 16 kHz, using a 40 ~ pulse, a pulse delay of 2.5 s and a 0.524 s acquisition time. A known weight of TKS (tetrakistrimethylsilane) standard was added to a predetermined weight of liquid product. In order to quantify the extent of 13C incorporation into the major peaks of the labelled samples, the ~3C NMR spectra of the corresponding unlabelled samples were also obtained. These quantitative experiments were run overnight with a small amount of 0.1M solution of chromium acetylacetonate [Cr(AcAc)3] in deuterochloroform added. All the spectra were processed using an exponential multiplication with a 10 Hz line broadening factor. Sealed-tube combustion of the coked catalysts and ~3C labelled liquid products was conducted using copper oxide in a sealed evacuated tube at a temperature of 850~ The tube was then broken under vacuum, combustion products collected and directed through a series of vacuum lines. Water was removed from the combustion products by a dry ice/acetone slush trap. The remaining carbon dioxide was collected for stable isotope analysis by mass spectrometry. Equation (iv) was used to determine the mass of ~3C present in the coke. To calculate the enrichment of '3C, the mass of ~3C present in the cokes derived from feeds doped with ordinary toluene and benzene was subtracted from the value obtained in equation (iv). ]3C Mass = (Carbon Content Coked Catalyst x Mass Catalyst Used)x ~3C %

(iv)

3. RESULTS AND DISCUSSION

3.1 Liquid Products Figures 1 to 4 present the distributions of the 13C labels, expressed as the excess atomic percent (Equation (ii)), for the different product groups (Table 1). The general similarity between the profiles, implies that the sulfur content of the feed did not markedly affect the product distribution.

255

Figure 1. Distribution of 13C label based on ~3C % values (Equation (ii)) from the low sulfur feed doped with 8 % ~3C toluene.

Figure 2. Distribution of ~3C label based on ~3C % values (Equation (ii)) from the low sulfur feed doped with 8 % 13C benzene. Table 2 contains the 8~3C data on both a permil and an excess atomic percent 13C basis. To illustrate the portion of the starting enrichment being incorporated into each group, the mass of ~3C is expressed as a percentage of the total mass of ~3C added to the feeds. These results are summarised in Table 3. The accumulative errors for the balance are estimated to be at

256 least approximately + 10%. These are due to inherent errors associated with the integrated peaks in normal GC analyses, coupled with the fact that in GC-IRMS, the smaller peaks are below the threshold level for detection. Also, due the transfer times of over 30s from the GC though the combustion furnace to the mass spectrometer, the peaks are considerably broader than in normal GC analysis.

Figure 3. Distribution of ~3C label based on ~3C % values (Equation (ii)) from the high sulfur feed doped with 8 % ~3C toluene.

Figure 4. Distribution of ~3C label based on 13C % values (Equation (ii)) from the high sulfur feed doped with 4 % ~3C benzene.

257

Table 2 Isotopic ratios (813C, Equation (i)) and the concentrations of excess 13C (Equation. (ii)) for the liquid products Feed Low sulfur doped with High sulfur doped with Group 8% 13C C7H8 8% 13C C6H6 8% 13C C7H8 4% 13C C6H6 813C 13C% 813C 13C~ 813C 13C~ 813C 13C% 1 2 3 4 5 6 7 8 9

2490 3110 149 29.8 42.6 44.3 17.1 -3.2 -0.1

2.66 3.30 0.16 0.03 0.05 0.05 0.02 <0.01 <0.01

10

-7.2

<0.01

3780 2220 397 232 203 127 71.5 42.7 34.4

3.98 2.38 0.44 0.26 0.22 0.14 0.08 0.05 0.04

2190 3020 158 23.9 35.3 39.5 33.0 1.9 3.5

2.35 3.21 0.17 0.03 0.04 0.04 0.03 <0.01 <0.01

-

-5.1

<0.01

-

2450 1810 195 111 89.8 49.0 37.1 15.0 0.2

2.62 1.95 0.22 0.12 0.10 0.06 0.04 0.02 <0.01

1.1

<0.01

Table 3 Overall distributions of the 13C labels in the liquid products groups (Table 1). *Estimated error = + 10 % of total 13C. High Sulfur doped with Feed Low sulfur doped with 8% 13C C7H8 4% 13C C6H6 8% 13C C7H8 8% 13C C6H6 Group 1

2 3 4 5 6 7 8 9 10 TOTAL

23

33 2 0.1 0.1 0.1 0.04 3.8xl 0.3 3.2xl 0.3 1.9xl 0-3 58%*

5

4 0.9 0.1 0.1 0.07 0.02 4.8xl 0.3 2.8xl 0.3 10%*

10

20 2 0.1 0.1 0.2 0.03 7.2xl 0.3 0.01 6.0xl 0.3 32%*

5

5 0.8 0.1 0.1 0.1 0.03 5.2xl 0.3 1.9x10 .3 2.2xl 0.3 11%*

The data illustrate that the 13C label has been incorporated significantly into groups 1 and 2. Indeed, the C2-C4 alkylbenzenes which, as indicated by GC-MS are present in greater concentrations than the C6-C8 alkanes, contain the majority of the 13C label accounted for in the liquid products. Thus, transalkylation appears the dominant reaction pathway responsible for migration of the 13C labels. Despite the large experimental error associated with the overall label balances (see following), the results indicate that the labelled toluene has undergone the much greater extent of rearrangement. Relatively small quantities of the 13C label occur in a range of PAHs, which includes naphthalene, phenanthrene, anthracene and chrysene. These species in the decant oil boiling range are probably the active intermediates

258 in coke formation that must form by a combination of transalkylation, cyclisation and dehydrogenation. Table 3 indicates that the recoveries of the label in the liquid products are 10-11% for benzene and 32 and 58% for toluene. These seemingly low recoveries arise from a number of factors. Firstly, hydrocarbon gases were not analysed. Further, due to time delays in the analyses, the label present in the light ends of the liquid products was inevitably lost. Similarly, GC-IRMS analyses did not take account of light end C5-C6 hydrocarbons and, therefore any unconverted benzene was excluded from the balance. Sealed-tube combustion of the ~3C labelled liquid products gave recoveries of between 3 and 20%. These low recoveries were the result of problems with the experimental procedure in which the loss of a significant fraction of the more volatile components could not be avoided. The solution state ~aC NMR spectra obtained for the second set of liquid products obtained from the low sulfur feed doped with normal and labelled benzene and toluene are shown in Figures 5-8. In the case of the products from the ~3C benzene doped feed, the extremely large peak present at 129 ppm in the aromatic region of the spectrum, confirms that a large amount of unconverted benzene is present in the liquid product (Figure 5). Indeed, the 129 ppm peak is still prominent in the spectrum of the product from the feed doped with unlabeled benzene and, from the intensity of this peak, it is estimated that the benzene content is 8 + 2 % w/w, suggesting that only a small fraction of the benzene has been converted and little had been lost by evaporation (Figure 6). As indicated above, unconverted benzene was not taken into account for the overall the isotopic mass balances. For the ~3C toluene spiked product, unconverted toluene is evident and, from the intensity of the methyl peak at 22 ppm (Figure 7), it is estimated that 40% of the13C has undergone rearrangement. This estimate is reasonably consistent with those from the overall isotopic mass balances (Table 3).

220

200

180

S6 0

S4 0

S2 0

1 O0

80

60

Figure 5: ~3C NMR of low sulfur feed doped with 8 % 13C benzene

40

20

0

259

....

~

~. . . .

~

o

~o

Figure 6" '3C NMR low sulfur feed doped with 8 % unlabelled benzene. /

/

" ................

i';'o . . . . . . . . . . . . . . . ;';'o . . . . . . . . . . . . . . .

;'~'o . . . . . . . . . . . .

i';'o . . . . . . . . . . . . . . . ;'go . . . . . . . . . . . . . . . . ;'J . . . . . . . . . . . . . . . . . 6'; . . . . . . . . . . . . . . PpM

20 . . . . . . . . . .

2o

Figure 7: 13CNMR of low sulfur feed doped with 4 % 13Ctoluene. Table 4 Peak ratios from the '3C NMR Spectra Doped low sulfur feed Ratio 8 % '3C Benzene Aliphatic 9TKS 8 % Benzene Aliphatic " TKS 4 % '3C Toluene Aromatic" TKS 4 % Toluene Aromatic" TKS

Ratio Value 5 5 3 2

~

260

................. W .... 2ao

t60

140

laO

!oo.

eo

60

..... 40

20

Figure 8" 13C NMR of low sulphur feed doped with 4 % unlabelled toluene. Comparing the ratio of integrals of the internal standard, TKS to that of the aliphatic region in the case of benzene, and the aromatic region in the case of toluene for the labelled and unlabelled products, provides an indication of whether the transformations undergone by the labelled carbons could have involved entering and leaving phenyl rings through tropylium carbocations (C7H7+). These ratios are listed in Table 4 and they indicate that virtually no label from the ~3C benzene was incorporated into the aliphatic components of the liquid product. However, the higher aromatic to TKS intensity ratio of 3 for the 13C labelled toluene compared to the unlabelled toluene (Table 4), suggests that some of the label in the methyl position has been incorporated into phenyl tings. Table 5 Enrichment of the 13C labels in the cokes for the recovered catalysts. High Sulfur doped with Low Sulfur doped with 8% 8% 13C C7Hg 4% 13C C6H6 13C C7H8 8% 13C C6H6 % Coke on Catalyst 0.64 0.64 0.64 0.65 Weight Coke (mg) 26 26 26 27 ~513C -11.83 44.17 -7.91 75.3 Weight 13C (mg) 0.29 0.30 0.29 0.32 Enrichment 13C (mg) 0.014 0.030 0.016 0.051 O~ 13C of Coke 0.05 0.11 0.06 0.19 % Feed Enrichment 0.18 0.05 0.19 0.18 3.2 Coke Concentrates

Table 5 presents the data obtained from the sealed-tube combustion of the coked catalysts. The excess of 13C in the cokes is expressed on a percent basis of (i) the amount of carbon in the coke and (ii) the amount of the initial ~3C label added. The more positive ~3C values in

261 relation to the unlabelled toluene and benzene doped feeds that gave 813C values of-25.8 %0 and -24.9 %o, indicates that some incorporation of the labelled carbons into the coke has occurred. However, both the percentages of the initial 13C labels incorporated into the cokes, 0.05-0.19% and, the percentages of coke carbon derived from these aromatics are low (maximum of 0.19%). This indicates that benzene and toluene are minor contributors to the overall level of catalytic coke produced. Therefore, it is likely that alkenes and other aromatic species in FCC products, such as indenes are the major precursors of the catalytic coke.

4. FUTURE W O R K This study with 13C labelled benzene and toluene has provided the first direct evidence that light aromatics can contribute to catalytic coke in FCC. In terms of future work, conventional-sealed tube combustion is being used to ascertain if the natural isotopic compositions of cokes derived from acid-catalysed and thermal pathways vary. The premise is that catalytic coke could experience extensive isotopic fractionation since many carboncarbon bonds have to be formed for low molecular mass precursors. Since ~2C-~2Cforms and cleaves faster than 12C-13C, both isotopically lighter products and cokes i.e. with a more negative 8~3C values could arise. In terms of cracked products, the study by Sackett [14] has established that methane produced from both catalytic and thermal cracking of n-octadecane is generally subject to marked isotopic fractions (ca. 25 %0). On the other hand, thermal coke, which is formed from large molecular precursors in the feed, will produce coke that should have a similar isotopic signature as the feed. Table 6 provides some preliminary results for cokes generated in fluidised bed reactors. The data indicate that catalytic coke formed from n-nonene is lighter than the initial feed. In contrast, n-hexadecane-derived coke has retained the same isotopic signature as the starting feed, suggesting that the coke is formed from uncracked species. These cokes were obtained using a fluidised bed reactor and a low metals equilibrium catalyst [15]. The thermal cokes were also produced in the fluidised-bed but the catalyst used was a clay with essentially zero activity. Therefore, the resultant cokes from the VGO and decant oil used were truly thermal in nature. Again, the results are consistent with the hypothesis that coke formed from heavy feeds displays no significant isotopic fractionation effect in relation to the feed. Although values for metal-mediated coke have not yet been obtained, these preliminary results suggest that similarities in the isotopic signatures of feeds and cokes is indicative of thermal coke while lighter isotopic values may be obtained with contributions from catalytic coke. Table 6 Preliminary results from conventional-sealed tube combustion for different coke types Feed 813C Feed 813C Coke n-Nonene -23 %o -26 %0 n-Hexadecane -31%0 -31%0 VGO -28 %o -28 %0 Decant Oil -28 %o -28 %o

262

REFERENCES [1] E. H. Wolfe, A. Alfani, Catal. Rev. Sci. Eng., 24 (1982) 329. [2] R. Hughes, Deactivation of Catalysts, Academic Press, London 1984 [3] M. Guisnet, P. Magnoux, B. Delmon and G. F. Froment (Eds), Studies in Surface Science and Catalysis, Catalyst Deactivation, 88 (1994) 53. [4] P. Turlier, M. Forissier, P. Rivault, I. Pitault and J.R. Bernard, M. L. Occelli and P. O'Connor (Eds), Fluid Catalytic Cracking III, Am. Chem. Soc. Symp. Ser., No. 571 (1994) 98. [5] M. Schoell, J. Brookes and D. H. Welte (Eds), Advances in Petroleum Geochemistry, 1 (1984) 215, Academic Press, London. [6] G. Steer, T. Ohuchi, K. Muehlenbachs, Fuel Process. Technol., 15 (1987) 429. [7] V. P. O'Malley, T. A. Abrajano Jr., J. Hellon, Org. Geochem., 21 (1994) 809. [8] C. McRae, G. D. Love, I. P. Murray, C. E. Snape, A. E. Fallick, Anal. Comm., 33 (1996) 331. [9] Z. K. Guo, A. H. Luke, W. P. Lee, D. Schoeller, Anal. Chem., 65 (1993) 1954 and references therein. [ 10] K. J. Goodman, J. T. Brenna, Anal. Chem., 64 (1992) 1088 and references therein. [11] T. R. Filley, R. M. Filley, S. Eser, K. H. Freeman, Energy and Fuels, 11 (1997) 637. [12] C.L. Wallace, C.E. Snape, N.J. Gudde and G.W. Kettley, Proc. 2"d International Conference on Refinery Processing, pp 523-528. AIChE Spring National Meeting, Houston, 14-18 March 1999. [13] C. McRae, C. E. Snape, A. E. Fallick, The Analyst, 123 (1998) 1519. [14] W.M. Sackett, Geochim. et Cosmochim. Acta, 42 (1978) 571. [15] C. E. Snape, B. J. McGhee, S. C. Martin, J. M. Andresen, Catalysis Today, 37 (1997) 285.

Studies in Surface Scienceand Catalysis 134 M.L. Occelli and P. O'Conner (Editors) 9 2001 Elsevier Science B.V. All rights reserved

BIFUNCTIONALITY

IN CATALYTIC

263

CRACKING

CATALYSIS

W.L. Schuette a and A.E. Schweizerb

aDeceased bRetired from" ExxonMobil Refining and Supply Company, Process Research Laboratories, P.O. Box 2226, Baton Rouge, Louisiana 70821-2226 USA Two chemistries have been found to be operative in catalytic cracking catalysis. One is the conventional carbonium ion chemistry catalyzed largely by the Bronsted acid sites provided by the zeolite framework alumina. Radical catalytic chemistry also occurs. It initiates the formation of olefins by dehydrogenation and/or dealkylation of saturates to form olefins. The radical chemistry is catalyzed by Lewis acid electron accepting sites provided by various forms of amorphous alumina. The olefins, once formed, diffuse to Bronsted acid sites where they pick up protons to become carbenium ions. The carbenium ions readily convert to lower molecular weight carbenium ions. Thus conversion is bifunctional involving initiation by Lewis acid sites and propagation by Bronsted acid sites. The Lewis and Bronsted are synergistic to one another in providing conversion. Zeolite nonframework alumina is a particularly effective radical catalyst. Increased radical initiation leads to lower coke and dry gas selectivity and higher naphtha olefins. Zeolitc nonframework alumina is synergistic to ZSM-5 for the production of light olefins, e.g., propylene and butenes. 1. INTRODUCTION Cat cracking is the major petroleum refining catalytic process. Approximately 35% of all petroleum crude feeds are processed through a cat cracker. The principal products of the cat cracker are gasoline, propylene, butenes, isobutane, mid-distillate, and a high boiling bottoms product. Usually it is desirable to maximize the conversion of bottoms and mid-distillate to lighter products principally gasoline. However, currently the production of light olefins, e.g., propylene, for chemicals is receiving considerable attention. While the initial catalytic cracking catalysts were amorphous silica-aluminas such as acid treated clays, materials have improved markedly over the years. The inclusion of zeolites alone revolutionized catalytic cracking by providing much higher activity, which has enabled refiners to significantly expand their cat cracking capacity. Over the past fifty years capacity increases of more than five fold have been achieved. There have also been several less revolutionary but, nevertheless, very important improvements in zeolitic cat cracking catalysis. One was a shift to lower unit cell size faujasites which increased olefins and octane while reducing the production of coke and dry

264 gas. Another was the incorporation of ZSM-5 additives which recracked naphtha olefins into lighter olefins while improving the octane of the remaining naphtha. Despite all this progress, a comprehensive understanding of the mechanisms of catalytic cracking remained elusive. Most of the attempts to provide an understanding focused solely on the carbonium ion chemistry catalyzed by both the acid sites of the zeolite and amorphous matrix. A few researchers, e.g., McVicker et. al., tried to point out that other chemistries (radical) may also be catalyzed by additional sites on the catalyst not associated with carbonium ion sites. The present paper draws from evidence in the open literature that (a) indeed radical catalysis does occur and that it plays an important role in catalytic cracking catalysis, (b) radical catalysis occurs at electron accepting sites (Lewis acid sites), (c) radical catalysis results in the dehydrogenation of saturates to form olefins, and (d) the olefins once formed diffuse to Bronsted acid sites where they pick up protons to become carbenium ions and subsequently convert. Thus the radical and carbenium ion sites act concertedly as a bifunctional catalyst. It will be shown that recent advancements in cat cracking catalysis are fully consistent with this understanding ofbifunctionality. 2. THE NATURE OF BIFUNCTIONALITY Any discussion of bifunctional catalysis will profit from a review of Weisz's classic work on the subject (1). A synopsis of this work is provided in Figure 1. This figure displays the apparatus and results which establish the role of bifunctionality in the conversion of n-butane. The apparatus consists of two reactors in series. The first reactor contains a Pt catalyst. The second contains a Mordenite catalyst, n-Butane and hydrogen are passed through the two reactors. The temperature of the first reactor (Pt) is varied between 200 ~ and 600 ~ C. The temperature of the second reactor is held constant at 230 ~ C. The conversion of n-butane in the second reactor is monitored as a function of the temperature in the first reactor. It is observed that the conversion of n-butane in the second reactor increases as the temperature of the first reactor is increased. Presumably increasing temperature in the first reactor increases dehydrogenation of butane to butene. The increased production of butene in the first reactor results in increased conversion of C4 species in the second reactor. Thus it appears that the overall conversion of n-butane is bifunctional in character and involves the production of olefinic intermediates. The olefins produced in the first reactor probably become associated with protons from the Mordenite in the second reactor and become carbenium ions. Conversion to other molecular species then ensues.

265


r.-

(Pretreot) Reoqlor I F

~ P i ......... ~:

(Crocking)

,s

Reoctor Lr'I ..

>

-

$omNing loop

[

..... ooo~

:2

g o1_! tD ga 9s ~ 50 s')

_

~a

tD u J..

o

0 200

300

400

500

Temperature in reoctor [ , *C

n-El~tcne Hydrogen

-6

.5 .4

-3

I0~0 10 tO

-~ ~0

.~ )0

Otefin Concenlrahon (relo)ive)o butane )

Fig. 1. Weisz demonstrates bifunctionality in butane cracking

3. RADICAL CATALYSIS IN ISOBUTANE CONVERSION In a seminal study on the conversion of isobutane over solid acids G. B. McVicker et. al. proposed that some solid acids catalyzed radical-like chemistry (2). This was very provocative since radical chemistry was generally believed to be thermal and not catalytic chemistry. McVicker showed that thermal decomposition of isobutane results in equilmolar production of methane and propylene and also butenes and hydrogen. He then proceeded to demonstrate that amorphous solid acids such as alumina, silica-alumina, and halogenated aluminas produced the same products in the same proportions but at lower temperatures indicating that these materials were indeed catalysts for radical-like chemistry (Table 1). The halogenated aluminas were most active in promoting the radical-like catalysis. This demonstrated that not only did radical catalysts exist but that it is conceivable to formulate fairly high activity versions of such catalysts. McVicker went on to postulate that the radical mechanism involved electron acceptor sites on the solid acid catalyst which resulted in the production of radical cations. The radical cations subsequently decompose into methane and propylene or into butenes and hydrogen.

266 Table 1 Various amorphous aluminas act as radical catalysts in isobutane conversion CATALYST

NONE

AL203

SIO2-AL203

0.9% CLAL203

TE,MPERA'D,YRE. K

0.9% FAL303

923

873

873

873

823

1.6 ! .8 2.5 2. ! 9 3.7

1.6 1.4 4.2 2. ! 3 4. I

1.5 ! .7 4.9 2.12 5.5

8. I 7.4 7.6 2.25 6.5

7.3 7.0 9.0 2.22 14.2

-

0.1 0.2

0.1 TR

0.6 0.3

RADICAL PRODUCTS, MOL% METHANE PROPYLENE BUTENES H/C RATIO I-I2 (ESTIMATE) (~ARBONIUM ION PRODUCTS, MOL% PROPANE N-BUTANE PENTANES I_SOBUTANE, MOL% MINOR, MOL%

TR

-

94.0

92.6

91.4

76.0

74.7

0.1

0.2

0.2

0.8

I.!

REFERENCE2

The choice of isobutane as a molecular probe for radical catalysis was inspired. Isobutane is not very susceptible to conversion by carbonium ion chemistry since it cannot undergo beta scission. Therefore it is well suited to probe other catalytic sites which may be present near the carbonium ion sites. 4. C A R B O N I U M ION C A T A L Y S I S IN I S O B U T A N E C O N V E R S I O N

McVicker also studied the conversion of isobutane over faujasite (2). He used a sample of Union Carbide's LZ-Y82 (ultrastable Y, USY) for this purpose. Thermal treatment used in the production of LZ-Y82 results in some of the zeolite framework alumina being driven from the structure to extra framework positions as an amorphous component (zeolite nonframework alumina). Over this zeolite he noted the appearance of methane and propylene but also large quantities of saturated products (see Table 2). He proposed that the saturated products observed with faujasite arose from carbonium ion assisted chemistry. Implicit in McVicker's conclusion is that two catalytic chemistries are operative in the conversion of isobutane over LZ-Y82. One is radical-like in character and the other is carbonium (carbenium) ion chemistry. He further proposed that radical catalysis is prerequisite to carbonium ion chemistry. On the amorphous solid acid catalysts he found very little production of saturates indicating little or no carbonium ion assisted chemistry for the conversion of isobutane over these materials

267 Table 2 Faujasite produces secondary carbonium ion products from isobutane cracking 1,5% F-AL203

FAUJASITE

CATALYST

CONDITIONS TEMPERATURE, K CATALYST, GM _RADICAL PRODUCTS. MOL% .. METHANE PROPYLENE BUTENES H/C RATIO H2 (ESTIMATE)

823 1

773 1

773 1

773 0.3

723 1

9.7 8.2 9.7 2.23 17.0

! .g 2.0 2.8 2.19 4.7

2.5 0.9 ! .5 2.52 0

1.0 0.7 0.8 2.48 0

0.4 0.2 0.7 2.48 0

1.3 0.4 TR

0.3 0.3 TR

I 1.6 7.9 3.9

2.5 2.2 i .0

2.7 4.0 ! .7

68.5

92.4

69.5

91.2

89.8

CARBONIUM ION PRODUCTS, MQL% PROPANE N-BUTANE PENTANES !SOBUTANE. MOIL.*/, REFERENCE 2

It is instructive to examine the faujasite data of Table 2 more closely. At 773 ~ K yield data are provided for two different catalyst charges, 1.0 and 0.3 grams. Isobutane conversions in these experiments were 30.5 and 8.8 mol.% respectively. The magnitude of the various yield changes with changing conversion can provide information about initial and secondary products (3). For instance a reduction in conversion will reduce secondary products more than primary products. Note the large reduction in the saturates (carbonium ion products) as conversion is decreased by a factor of ca. 3. This indicates that the carbonium ion products are secondary products of isobutane conversion. They are not formed directly by some protolytic catalysis but certainly by some further chemistry involving reactive intermediates. Note also that methane changes more nearly in proportion to conversion. This suggests that methane is a primary product of isobutane conversion. If methane is a primary product of conversion then its co-product would be propylene. Methane and propylene have been found to be radical-like products of isobutane conversion over amorphous solid acids. They are likely formed as initial products of radical catalysis by the zeolite nonframework alumina. As mentioned earlier zeolite nonframework alumina is an amorphous alumina. It would be expected to have high Lewis acidity. Methane is a stable initiation product. Propylene is a reactive initiation product which can pick up a proton from a Bronsted site to become a carbenium ion. From this perspective it is more accurate to term the Bronsted acid catalyzed chemistry carbenium rather than carbonium ion chemistry. Propylene can convert to other molecular weight olefins by disproportionation. All olefins can convert to saturates by hydride transfer. Hydride transfer is a form of carbenium ion chemistry (4). As reactive intermediates, olefins are both being formed (radical catalysis) and destroyed (carbenium ion catalysis). Thus they should change

268 less with conversion than either stable primary products (methane) or stable secondary products (saturates). The data of Table 2 show this to be true. 5. RADICAL CATALYSIS BY LEWIS ACID SITES Another study reports on the conversion of isobutane over various amorphous solid acid catalysts (5). This study reported a strong correlation between a rate constant for isobutane conversion and the Lewis acidity of various amorphous solid acid catalysts. The rate constant for isobutane conversion was found to increase in direct proportion to the concentration of Lewis acid sites (Figure 2). The simplest explanation is that the Lewis acid sites catalyze the radical-like conversion of isobutane. An alternative explanation is the Lewis acid sites simply enhance the acidity of the Bronsted sites by charge withdrawal (3, p. 69). However if this were the case one should see evidence in increased production of saturates. This is not seen.

0[

T.s

E.

;c

1.0

o u o N

t

.

0.5

|

) IO4 .x

,,

2 4 6 8. Concenlrolion of Lewis ocid Qroq)s,mn~lm 2

Fig. 2. Lewis acidity drives isobutane conversion on amorphous solid acid catalysts

6. INITIATION/PROPAGATION IN BIFUNCTIONAL ISOBUTANE CRACKING McVicker further reports on isobutane partial pressure effects (2). He observes that the partial pressure dependency over the amorphous solid acids is first order for methane and

269 olefin formation while the appearance of saturates over faujasite is second order. This may be regarded as evidence that initiation chemistry is monomolecular while the propagation chemistry is bimolecular. Initiation by demethylation or dehydrogenation would be expected to be monomolecular. On the other hand the secondary conversion of light olefins has been found to be by dimerization/recracking, a bimolecular reaction (Figure 3)(3). Also the conversion of olefins into paraffins by hydride transfer would involve bimolecular reactions. This is further evidence that initiation of isobutane conversion involves radical-like monomolecular chemistry while propagation is provided by bimolecular carbenium ion chemistry.

0.7-~

I UN~MOt.ECULAqo ~'xJN

_ o.s-

~vt

'~'. 0.5 -

E ~ 0.4 ~,

0.3-

/ X__. /

0.2 -

CRAC,,,,o

0

o.I-

" ,,0 9. ~ ~

0 -- 0 ~ " 9

/

\(~

" ~

~'--~------O

0,----0

I.

I _

I ,

.I

2

3

4

5

I, 6.

I

"I

7

8

J

9

CARBON NUMBER

Fig. 3. Light olefins convert by bimolecular cracking

In fact McVicker goes on to propose kinetics which incorporates both initiation and propagation catalysis in the conversion of isobutane (2). In his formulation the initiation catalysis provided by the electron acceptor sites is found to be multiplicative to the propagation catalysis provided by the faujasite in facilitating total isobutane conversion activity. Thus not only do solid acid catalysts provide catalysis for both radical-like and carbenium ion assisted chemistries, but the chemistries interact synergistically to provide the overall conversion of isobutane. All the above information is assimilated in an understanding in which initiation is associated with the formation of olefins by dehydrogenation of isobutane over Lewis acid

270 electron accepting sites. These olelfins, once formed, diffuse to Bronsted acid sites, pick protons and become carbenium ions. Conversion then proceeds by beta scission. In this schema there are distinct sites involved in the initiation and propagation chemistries. No conversion will occur if the initiation of olefins does not happen. No propagation will occur without the presence of initiated olefins. Conversion involves the formation of olefinic intermediates. The appearance of saturates is due to secondary hydride transfer reactions catalyzed by Bronsted sites. 7. HISTORICAL EVIDENCE FOR BIFUNCTIONAL CAT CRACKING CATALYSIS Bifunctional catalysis is well understood. As noted above, Weisz was prominent in the development of such understanding for metal/acid bifunctional catalysts (1). A key feature of bifunctional catalysis is a synergistic interaction between the two types of sites in providing activity. A catalyst activity synergy between amorphous silica-alumina and faujasite in composite cat cracking catalysts has long been noted (7). The present authors believe that the interaction between radical initiation and carbenium ion propagation is the explanation for this synergy. More recent evidence is provided by the use of amorphous alumina as an activity enhancement in zeolitic cracking catalysts. By itself alumina has been found to provide very low cracking activity (4, p. 46). However when combined with faujasite in a cracking catalyst it much enhances the activity of the composite catalyst (Table 3). For example the inclusion of 20 wt.% alumina in a catalyst containing 20 wt.% USY increased the catalyst activity 78%. This improvement is much more than would be expected from the additive contribution of a very low activity ingredient. It is suggestive of a synergistic contribution as may arise from bifunctional catalysis. The authors contend that the amorphous alumina has very high Lewis acidity but very little Bronsted acidity. Thus it provides excellent initiation but little ability to propagate the conversion. Faujasite, on the other hand, has a much larger ratio of Bronsted to Lewis acidity. Thus it provides much higher propagation to initiation capability. Combining both the amorphous alumina and the faujasite in a common catalyst provides a better balance of initiation and propagation than either ingredient alone; hence improved activity. Table 3 Addition of matrix alumina much increases activity o f zeolite cracking catalyst CATALYSTCOMPOSITION.WT% ULTRASTABLEZEOLITE AL203 SI02-AL203 GEL

20 0 80

20 20 60

50 1

64 1.78

pERFORMANCE 430~ CONVERSION, VOL.%

SECONDORDERACTIVH'Y REFERENCE 8

271 8. ZEOLITE N O N F R A M E W O R K ALUMINA AS A RADICAL CATALYST Zeolite nonframework alumina is derived from zeolite framework alumina which has been driven from the framework by exposure to severe hydrothermal conditions such as exist in a cat cracker regenerator. The academic community has long been aware of a synergistic interaction between zeolite framework and nonframework alumina (9). Prominent explanations for the enhanced activity associated with zeolite nonframework alumina involve chemical interactions between the framework and nonframework aluminas which enhance the activity of the framework sites. The present authors propose that the nonframework alumina is an amorphous alumina that has very high Lewis acidity and thus very high initiation activity but low propagation activity because of its low Bronsted acidity. Its high initiation activity complements very nicely the very Bronsted acidity of the zeolite framework sites. The two activities are multiplicative in providing enhanced bifunctional zeolite activity. 9. GASOLINE COMPOSITION AS EVIDENCE FOR INITIATION Further evidence for initiation is found in gasoline composition data obtained on amorphous and zeolitic cracking catalysts (10). These data (Table 4) show that the amorphous catalysts, which have a greater ratio of Lewis/Bronsted acidity than the zeolitic catalysts provide much higher naphtha olefins and less of the paraffin and aromatic products of secondary hydrogen transfer. Hydrogen (hydride) transfer was noted earlier in this paper to be catalyzed by Bronsted acid sites. Early explanations for the lower naphtha olefins with the zeolitic catalysts focused on the high hydrogen transfer activity of these catalysts.

Table 4 Cat cracked gasolinecomposition reflectsradicalinitiationchemistry FEED CATALYST

CALIFORNIAVIRGIN GAS OIL

CALIFORNIA COKER GAS OIL

GACHSARAN GAS OIL

DURAB~ 5

DURABEAD1

DUP.ABEAD5

DURABEAD1_

.DURABEAD5

D U R A B ~ 1_

21.0 19.3 14.6 45.0

8.7 10.4 43.7 37.3

21.8 13.4 19.0 45.9

12.0 9.5 42.8 35.8

31.9 14.3 16.3 37.4

21.2 15.7 30.2 33.1

GASOLINE, VOL.% PARAFFINS CYCLOPARAFINS OLEFINS AROMATICS

"

DURABEAD 5 IS EARLYGENERATIONZEOLITE(REHX) DURABEAD 1 IS AMORPHOUS SILICA-ALUMINA REFERENCE 10

It is, however, possible to set forth expectations for naphtha olefinicity based on kinetic analyses. One begins by writing out the differential equation for the production and destruction

272 of naphtha olefins. This is accomplished in equation I below. dCN=/dT = klkp(1 -X) z -

k H C N --

-kxCN=

(1)

In this equation CN= is the yield of naphtha olefins, T is the contact time, kl is the initiation rate constraint, kp is the propagation rate constant, X is the fractional conversion of the feed, kH is the hydrogen transfer rate constant, and kx is the rate constant for the recracking of naphtha olefins to light olefins. Thus the equation depicts the formation of naphtha olefins by the second order conversion of feed molecules. Both initiation and propagation are involved in the conversion of feed molecules into naphtha olefins. The equation also depicts the destruction of naphtha olefins by hydrogen transfer and recracking. Both of these destruction steps involve only propagation or Bronsted sites. The maximum in naphtha olefins may be determined by setting the derivative in equation I equal to zero. The maximum in naphtha olefins is then determined by equation 2. CN=Max = klkp(1 -XMax) 2/(kH+kx)

(2)

Note that in this equation k p , kH, and kx are all related solely to propagation or Bronsted acidity in the numerator and denominator of the equation. Thus the Bronsted acidity effectively cancels out and the maximum in naphtha olefins is expected to be a function of initiation or Lewis acidity only. This kinetically determined expectation is in direct opposition to the opinion that naphtha olefins are set by the hydrogen transfer activity of the catalyst. This is very important for catalyst design. One increases naphtha olefins by increasing the initiation activity of the cracking catalyst not by minimizing Bronsted acidity. 10. ZEOLITE N O N F R A M E W O R K ALUMNA AS SUPERIOR RADICAL CATALYST Some convincing corroboration of bifunctionality in cat cracking catalysis is provided by data generated by Lercher (11). In this publication Lercher provides MAT data on a series of commercial cat cracking catalysts. The MAT test is defined by ASTM method 3907 and uses a standard vacuum gas oil (VGO) as feed as opposed to model compounds. Complete yield structures were provided at constant conditions. From this data it is possible to calculate rate constants for the hydrogen transfer which converts isobutene to isobutane (the ratio of isobutane to isobutene at the constant conditions of the MAT test). Summarized data are provided in Table 5. These data provide a six fold variation in hydrogen transfer activity. These hydrogen transfer rate constants were taken to be a measure of the carbenium ion activity of the subject catalysts. In addition second order rate constants for overall conversion at 421~ F were calculated. These rate constants were calculated by excluding the aromatic content of the feed (23.7 wt.%) as unconvertible. The rate constants for overall conversion covered a range of nearly five fold in activity. Thus by these two measurements we have quantified indications of a rate constant for initiation (radical activity) multiplicative to a rate constant for propagation (carbenium. ion activity) in overall conversion and a rate constant for

273

carbenium ion activity as measured by isobutene hydrogen transfer. Simple division of the former by the latter should provide a measure of the initiation (radical) activity. Table 5 Deconvolution of carbenium ion and radical activities f,ATAI,X~

UTE

UTA

UTO

trrc

UTS

UTR

5.68 1.68 1.98 71.5

3.07 2.40 1.22 63.6

3.39 2.23 1.38 61.1

4.76 2.03 1.79 68.6

6.75 0.88 2.97 72.5

4.66 1.79 1.66 64.2

3.42 14.9 4.35

1.28 5.02 3.92

1.52 4.01 2.64

2.35 8.93 3.81

7.67 18.8 2.45

2.61 5.32 2.04

YIELDS.WT.% ISOBUTANE ISOBUTENE COKE 421 F CONV. RATE CONSTANTS IC.4 HYDROGEN TRANSFER 421 F CONV. RADICAL ACWIVITY REFERENCE 11

The results of this computation are provided in Figure 4. This figure is a plot of radical activity versus zeolite nonfrarnework alumina. Notice the strong correlation. This suggests that zeolite nonframework alumina is a very strong radical catalyst. Zeolite nonframework alumina is known to have predominantly Lewis acidity (9). Thus our exercise sought to compute a measure of initiation catalysis and the result identified a prominent Lewis acid material as the source of such catalysis.

I

4.5

I

3.5

/

> v3 o < J 2.5 < _O 2 es

/

-

v

1.5

1 0.5

0

0.5

1

t.5

2

2.5

3

3.5

4

ZEOLITE NONFRAMEWORK ALUMINA, WT.%

Fig. 4. R a d i c a l activity v e r s u s z e o l i t e n o n f r a m e w o r k alumina

4.5

274 Linear regression of the radical catalysis rate constants against catalyst properties showed that matrix alumina may also play a role in the radical catalysis measured with the feed used and at the conditions studied in Lercher's report (11). The regressed equation relating radical contributions to catalyst properties are shown on Figure 4. The coefficient is much smaller for matrix relative to the zeolite, nonfrarmework alumina. However there is much more matrix than zeolite nonfrarmework alumina in these catalysts. In one catalyst (UTR) the matrix provides more than a third of the total initiation activity. The very high activity of the zeolite nonframework alumina may be due a combination of very high intrinsic activity and its location in close proximity to the propagating zeolite framework sites. How well the bifunctional model fits Lercher's MAT data is shown in Figure 5. His figure presents a plot of actual versus predicted 421~ F conversion. The authors believe that the excellent fit is strong corroboration for radical initiation and carbenium ion propagation in cat cracking catalysis.

75 i g w

Y

70

y = 1.00Z 0 u w w IIL w

Al

65

7

u o w Fa w w a.

50

55

60

65

70

75

ACTUAL COKE FREE CONVERSION,WT.'/,

Fig. 5. Actual versus predicted 421 F coke free conversion

11. COKE SELECTIVITY IN BIFUNCTIONAL CAT CRACKING CATALYSIS

Coke selectivity is the ratio of coke yield to the second order conversion activity. It is a very important parameter because most cat crackers operate at a coke burning limitation. At this limitation improved (lower) coke selectivity can provide conversion credits in the reactor. The MAT data of Lercher (11) provide some interesting insights on coke production. In Figure 6 we plot the MAT coke yield from the six catalysts of this study versus the square root of the carbenium ion activity of these catalysts. Recall that isobutene hydrogen transfer was

275 used as a measure of carbenium ion activity. Note the strong correlation. This suggests that most of the coke made from the feed used in this study is associated with carbenium ion activity provided by the Bronsted acid sites. Likely much of this coke is due olefin oligermerization/condensation chemistry. The square root relation likely results from a Voorhies type coke make formulation in which the rate of coke formation is inversely related to the amount of coke on the catalyst. 3.5

2.5

v i.08:

z~

.<

J

w" ~e

J

J

J

0 1.5 o

0.5

0

0.5

1

1.5

2

2.5

SQRT(KHT)

Fig. 6. Carbenium ion catalysis dominant factor in coke production

Coke selectivity is sometimes expressed as its reciprocal, dynamic activity. It is the second order activity for conversion divided by the coke yield. Since activity is driven by the radical activity multiplying the carbenium ion activity and coke is proportional to the square root of the carbenium ion activity, one would expect dynamic activity to correlate with the product of radical activity and the square root of the carbenium ion activity. This expectation is plotted in Figure 7. Note the excellent correlation. The correlation shows that one may increase dynamic activity, improve coke selectivity, by increasing either radical activity or carbenium ion activity or both. However recall that radical activity increases as the square of the zeolite nonframework alumina. This suggests that dynamic activity will be much improved through increased zeolite nonframework alumina. In faujasite zeolite nonframework alumina is increased by decreasing unit cell size. The merits of reduced unit cell size for improved coke selectivity have been previously reported (13). The present bifunctional rationale for this is new.

276

/ >-

6

/

> 5 IU < o_ 4

/

./

--

f-

.<3 >. 0

2

0

1

2

3

4

5

6

7

8

9

RADICAL ACTIVITY X 8QRTIC40/C4,,)

Fig. 7. Dynamic activity driven by product of radical and carbenium ion activities 12. ZEOLITE NONFRAMEWORK ALUM1NA/ZSM-5 SYNERGY It is generally recognized that ZSM-5, used in conjunction with a base cracking catalyst, results in the recracking of naphtha olefins to light C5- olefins (14). It has also been found that ZSM-5 makes greater increments of light olefins from low hydrogen transfer base cracking catalysts (Figure 8). In this figure higher rare earths correspond with higher unit cell sizes and higher amounts of zeolite framework alumina. An hypothesis for the synergy between low zeolite framework alumina and ZSM-5 is that the reduced hydrogen transfer associated with the lower zeolite framework alumina results in less conversion of naphtha olefins to uncrackable naphtha paraffins. 1.8 1.6

i

i

~1'~ 1.4 .J . , 1.2 z_ i1_

1

l

I

- e - DEL:i'A C 3 0 L E F I N S / -II- DELTA C40LEFINS ~-

9 III

"-,,,..

J 0 O.8 l"r"

\.

0 0.6 -1

Ir

\II~

0.4 .J uJ Q 0.2 0

24.21

24.22

24.23

24.24

24.25

24.26

24.27

24.28

24.29

UNIT CELL SIZE, ANGSTROMS

Fig. 8. Delta light olefin yields from ZSM-5 decrease with increasing unit cell size of base catalyst

277 More definitive expectations can be developed by formulating a kinetic expression for the formation and destruction of naphtha olefins over a base catalyst with ZSM-5 present. This is shown as equation 3 which is a minor extension of equation I to include ZSM-5. AII the symbols are the same as those defined for equations I and 2 except for kz dCN-/dT = kHkp(1 - X ) 2 -

kHCN---kxCN =

-kzCN=

(3)

which is the rate constant for ZSM-5 recracking of naphtha olefins. The maximum naphtha olefins may be calculated by setting derivative in equation 3 equal to zero. The result is equation 4. This equation is identical to equation 2 except for the appearance of CN=Max = klkp(1-XMax)Z/(kH+kx+kz)

(4)

kz in the denominator. Thus naphtha olefins are expected to decrease because they are recracked to light olefins by the ZSM-5. The production of light olefins is described by equation 5. In this equation CL= is the yield of light olefins. At the maximum naphtha dCL=/dT = kxCN= + kzCN= = (kx + kz)CN=

(5)

olefin yield equation 5 becomes equation 6. This equation predicts that ZSM-5 will dCL=/dT = (kx + kz)k&p(1 -XMax)z/(kn+kx+kz)

(6)

increase the rate of light olefin production. However kz appears in both the numerator and the denominator. Thus the effectiveness of ZSM-5 addition is expected to diminish as the amount of ZSM-5 is increased. Not so for initiation, kl is present only in the numerator. The rate of production of light olefins will always increase in direct proportion to the amount of initiation catalysis provided, kl. kl is always multiplicative to the ZSM-5 contribution and therefore synergistic for the production of light olefins. Initiation catalysis makes naphtha olefins which feed ZSM-5 for the production of light olefins 13. ARTIFICIAL ZEOLITE N O N F R A M E W O R K ALUMINA Zeolite nonframework alumina has been shown to be a particularly effective catalyst for radical initiation. Zeolite nonframework alumina is formed as zeolite framework alumina is driven from the zeolite framework. Thus the amount of zeolite nonframework alumina is limited to that originally present in the zeolite framework. This limitation could be removed if an artificial phase of zeolite nonframework alumina could be added to the zeolite. This has been accomplished. It was found that the addition of 10 to 15 wt.% alumina to various faujasites by vapor deposition from aluminum acetylacetonate led to remarkably increased activity of the finished cracking catalysts (15). Analytical characterization established that the added alumina resulted in increased Lewis acidity of the treated zeolite (15). The effectiveness of the artificial zeolite nonframework alumina is strong corroboration for the present theory of bifunctionality in cat cracking catalysis.

278 14. CONCLUSIONS Based on a survey of literature, it is concluded that catalytic cracking catalysis is fully bifunctional. The mechanism for conversion involves first the initiation of olefins by electron accepting (e.g. Lewis acid) sites. The newly formed olefins diffuse to the Bronsted sites where they pick up protons to become carbenium ions, which are then easily converted. The Bronsted acid sites are provided predominantly, but not entirely, by the faujasite framework alumina. Zeolite nonframework alumina is a much more effective radical catalyst than is amorphous matrix alumina in the catalyst matrix. The radical and carbenium ion sites are multiplicative in providing conversion. Coke selectivity can be controlled by controlling the balance between radical and carbenium ion sites. Generally a high proportion of radical catalysis favors low coke selectivity. An artificial source of zeolite nonframework alumina can be added to the zeolite to much enhance the activity of the zeolite. Zeolite nonframework alumina is synergistic to ZSM-5 for the production of light olefins, e.g., propylene and butenes and increasing cat naphtha octanes. REFERENCES

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

Weisz, P.B., Chemtech., August 1973. p. 498. McVicker, G.B., G.M. Kramer, and J.J. Ziemiak, J Catal., 83,286. (1983). John, T.M., and B.W. Wojciechowski, J Catal., 37, 240 (1975). Gates, B.C., J R. Katzer, and G.C.A. Schuit, Che..mistry of Catalytic Processes, McGraw-Hill, 1979. Sato M., T. Aonuma, and Y. Shiba, Proc. 3d Int. Con. Catal., p. 396, North-Holland Amsterdam, 1965. Abbot, J., and B.W. Wojciechowski, Canadian Journal of Chemical Engineering, Vol. 63, No. 3, p. 462-469. Plank, C.J., B.H. Davis (ed.) and W.P. Hettinger (ed.), Heterogenious Catalysis, ACS Symposium Series 222, 1983, p. 269 Gladrow, E.M., W.E. Winter, and W. L. Schuette, US Patent 4,259,212 (1981). Beyerlein, R.A., C. Choi-Feng, J.B. Hasll, B.J. Huggins, and G.J. Ray, Topics in Catalysis 4 (1997) 27-42. Eastwood, S.C., C.J. Plank, and P. B. Weisz, Proc. 8th World Pet. Cong., Moscow, 1971. Brait, A., K. Seshab, H. Weinstabl, A. Ecker, and J.A. Lercher, Applied Catalysis A: General 169 (1998) 315-329. Beyerlein, R.A., C. Choi-Feng, J.B. Hasll, B.J. Huggins, and G.J. Ray, Topics in Catalvs~s 4 (1997) 27-42. Blazek, J.J., A New Era for the Coke Selective Fluid Cracking, Cata.lyst, Davison Catyalagram 75 (1987). Miles, A.D., AKZO NOBEL Catalyst Courier, 27, August 1996. Schuette, W.L., and A.E. Schweizer, European Patent Application EP 0 749 781 A2 (1996).

Studies in Surface Scienceand Catalysis 134 M.L. Occelli and P. O'Conner (Editors) 9 2001 ElsevierScienceB.V. All rights reserved

279

Catalytic C r a c k i n g of A l k y l b e n z e n e s . Y-zeolites w i t h D i f f e r e n t C r y s t a l Sizes S. A1-Khattaf and H. de Lasa Chemical Reactor Engineering Centre, Faculty of Engineering, University of Western Ontario, London, Ontario, Canada N6A 5B9 Catalytic cracking of both cumene and 1,3,5-triisopropylbenzene (1,3,5-TIPB) is investigated in a novel CREC Riser Simulator. Reaction testing is developed using different reaction times, temperatures and Y-zeolite sizes (0.4-pm and 0.9-~m). At 350-450~ higher 1,3,5-TIPB conversions were obtained using 0.4-pm Y-zeolites. Differences on 1,3,5-TIPB conversions were, however, less significant at 500-550~ It is speculated that, in the 350-450~ range, this phenomenon is the result of an increased 1,3,5-TIPB limited diffusional transport. Regarding cumene and benzene, main 1,3,5-TIPB cracking products, it is found that their yields decrease and increase respectively with both crystal size and temperature. In addition and for achieving a better understanding of the cracking of 1,3,5-TIPB, the study was complemented with cumene catalytic cracking, a cracking reaction expected to be free of transport limitations in Y-zeolites. 1. I N T R O D U C T I O N Y-zeolites have been used extensively in fluid catalytic cracking (FCC) since 1962. The 60-~m commercial FCC catalysts are manufactured with 1-2-~m dispersed in an amorphous silica-alumina matrix forming the pellets [1]. The majority of the active sites are located within the zeolite pore structure. In order for the reaction to proceed, molecules have to diffuse through the large matrix pores into the zeolite crystals. As a result only certain hydrocarbon species with a molecular size smaller than 10.2 [2] can penetrate the zeolite pore structure. For the larger hydrocarbon molecules, only the active sites situated on the external surface area of the zeolites [3], representing about 3% of the total surface area, are available [4]. Over the past decade, processing heavy oil in FCC units has become more prevalent due to the declining availability and higher price of light crudes. These general market trends have made upgrading resides economically attractive and very essential to the FCC units existing [5]. However, processing heavy feedstocks containing more bulky molecules offer new challenges given the hindered transport of hydrocarbon species inside the zeolite structure. While diffusion in the catalyst matrix belongs to the well-known Knudsen regime, diffusion in zeolites falls into configurational regime [2,6]. This kind of diffusion is a process with an activation energy usually substantially greater than for other kinds of diffusion [2]. Since the molecule size in the pore is nearly the size of the passageway, the diffusion of molecules in zeolites is governed by a continuous interaction between the zeolite crystal and diffusing molecules [2,6].

280 Regarding hydrocarbon cracking in zeolites, it is claimed that catalytic reactivity is strongly influenced by chemical species transport. Hydrocarbon diffusivity can be affected by a variety of factors, including the relative size of pores and of the reactant molecules [7]. Catalysts having pores in a given diameter range, only admit molecules with a diameter smaller than the maximum pore size. Admitted reactant molecules can have full or partial access to the catalyst active sites. Larger molecules, especially those with diameters approaching the cutoff size, are strongly hindered in their diffusion. In this regard, it was suggested that product molecules could also be trapped in the small pores filled with adsorbed reactant molecules [6]. Diffusivity limitations of some hydrocarbon molecules and its effects on catalytic cracking were reported by Nace [8]. When a REHX catalyst was used for cracking, the rate constant increased with molecule size up to a maximum value. Then, the rate constant decreased drastically for three and four ring reactant molecules. This was an experimental confirmation that there are diffusional constrains in the intracrystalline zeolite pore structure with diffusivity becoming more limited with increasing molecular size. On the other hand, Maselli et al [3] reported that small Yzeolite were both more active and produced more liquid products than the large ones. Rajagopalan et al [9] cracked gas oil using small and large Y-zeolite particle. These authors concluded that small size zeohtes produce more gasoline and LCO and less coke and light gases than the large ones. Gianetto et al [10] pointed out that the small crystal Y-zeohtes produce more olefins, iso-paraffins and less aromatics and coke than the larger zeolites. N-hexadecane and 1,3,5-TIPB were used to confirm diffusivity effects on zeolite performance [11]. The extra large pore MCM-41 with 2.5-nm pore size, pure H-Y zeolite with 0.75-nm pore size, and commercial FCC catalyst were used. It was found that accessibility to the internal acid sites influences both catalyst activity and selectivity, with MCM-41 yielding the higher 1,3,5-TIPB conversions. In this respect, A1-Khattaf and de Lasa [12] have recently estimated the effective gas oil diffusivity in Y-zeolites at typical reaction conditions. On this basis effectiveness factors significantly smaller than one were calculated for gas oil cracking on Y-zeolites with crystals having 1 ~m and 0.1-~m size. Moreover, since secondary reactions in catalytic cracking, case of hydrogen transfer, occur between smaller molecules, an effectiveness factor of one was assigned to the secondary reactions. Thus, the accounting for diffusional constrains, using effectiveness factors, it allows demonstrating the major influence of the diffusivity phenomena on the product composition.

2. EXPERIMENTAL PROCEDURES 1,3,5-TIPB (Fluka 92075) and cumene were employed, in the present study, as model reactants compound. Regarding the FCC catalysts, commercially available Yzeolites (small crystal and large) from Tosoh Company were used. Properties of these zeolites are reported in Table 1.

281 Table 1. Properties of the small and large zeolites of this study Small Zeolite Na20 (wt%) Si02/A1203 (moYmol) Unit cell size (A) Crystal size (~m)

4.1 .... 5.6 24.49 0.4

Large zeolite 0.25 5.7 24.51 0.9

These Na-zeolites were ion exchanged with NH4NO3 to replace the sodium cation with NH4§ Following this, NH3 was removed and the H form of the zeolites was spray-dried using kaolin as the filler and silica sol as the binder. The resulting 60-pro catalyst particles had the following composition: 30 wt% zeolite, 50 wt% kaoline, and 20 wt% silica sol. The process of sodium removal was repeated for the pelletized catalyst. Following this the catalyst was calcined during 2 hr at 600~ Finally, the fluidizable catalyst particles (60-pm average size) were treated with 100% steam at 760~ for 5 hr. Table 2 reports the catalyst main properties following catalyst pretreatment. The unit cell size was determined by X-ray diffraction following ASTM D-3942-80. Surface area was measured using the BET method. It can be observed that the initial zeolite unit cell size of 24.5 .~ was reduced, following steaming, for both zeolites, to around 24.28/k. Table 2. Properties of CAT-SC and CAT-LC catalysts used in this study

Unit cell size (/k) BET Surface Area (m2/g) Na20 (wt%) Crystal Size (~m)

CAT'SC 24.28 " 148 Negligible 0.4

CAT-LC 24.28 155 Negligible 0.9

3. R E A C T I O N EVALUATION The catalytic activity of the Y-zeolite catalyst, prepared using the various techniques described above, was measured in a Riser Simulator using cumene and 1,3,5 triisopropylbenzene (1,3,5-TIPB) as the hydrocarbon feedstocks. The Riser Simulator is a novel bench scale with internal recycle unit. This reactor, invented by de Lasa [13], overcomes the technical problems of the standard micro-activity test (MAT).

282

Fig. 1. Schematic description of the CREC Riser Simulator Unit. A schematic diagram of the Riser Simulator is reported in Fig. 1. The Riser Simulator consists of two outer shells, lower section and upper section that permits to load or to unload the catalyst easily. This reactor was designed in such way that an annular space is created between the outer portion of the basket and the inner part of the reactor shell. A metallic gasket seals the two chambers, an impeller located in the upper section. A packing gland assembly and a cooling jacket surrounds the shaft supports the impeller. Upon rotation of the shaft, gas is forced outward from the center of the impeller towards the walls. This creates a lower pressure in the center region of the impeller thus inducing flow of gas upward through the catalyst chamber from the bottom of the reactor annular region where the pressure is slightly higher. The impeller provides a fluidized bed of catalyst particles as well as intense gas mixing inside the reactor. The Riser Simulator operates in conjunction with a series of sampling valves that allow, following a predetermined sequence, to inject hydrocarbons and withdraw products in short periods of time. Reaction products are measured by a Hewlett Packard 5890A GC with an FID detector and a capillary column HP-1, 25 m crosslinked methyl silicone with an outer diameter of 0.22 mm and an internal diameter of 0.33 microns. Detailed description of various Riser Simulator components and sequence of injection and sampling can be found in Pruski [14]. 4. R E S U L T S AND DISCUSSION Figs. 2a and 2b report the 1,3,5-TIPB conversion for a C/O=5 at various temperatures (350~ 400~ 450~ 500~ 525~ 550~ and reaction times (3s, 5s, 7s, 10s). This covers experiments developed with both the CAT-LC and the CAT-SC catalysts (CAT-SL: catalyst with large crystal, CAT-SC: catalyst with small crystal).

283

70 605040-

o

30-

~

20-

3~

100 300

,

,

350

400

,,,

450

,

500

,

,

550

10

600

Temperature(C)

Fig. 2a. Change of the 1,3,5-TIPB conversion With temperature and reaction time for CAT-LC, C/O=5. (x) 10 sec. (~k) 7 sec. (~ 5 sec. (O) 3 sec.

Fig. 2b. Change of 1,3,5-TIPB conversion with temperature and reaction time for CAT-SC, C/O=5. (x) 10 sec. (•) 7 sec. (-) 5 sec. (~) 3 sec.

For both catalysts and for at all the temperatures studied, the 1,3,5-TIPB conversion increases with reaction time (3-10 sec.). This is an expected behavior given that at longer residence times there is an increased opportunity of the 1,3,5-TIPB hydrocarbon molecules to be cracked. This trend is however, moderated at longer residence times due to coke formation. The effect of the temperature on the 1,3,5-TIPB conversion is also reported in Figs. 3a and 3b for a 3- and 5-sec. reaction time and C/O=5. It can be observed that increasing the temperature, for a set reaction time strongly augments the conversion up to a given thermal level. From this temperature and on, there is a much less significant influence of the temperature on the 1,3,5-TIPB conversion. For instance, for the CAT-LC catalyst the 1,3,5-TIPB conversion increases with the temperature until it almost stabilizes at 525~ Additional increases of the temperature have a weak influence on the 1,3,5-TIPB conversion. On the other hand the CAT-SC catalyst displays a similar behavior at 500~ For instance, at 5-sec. reaction time (Fig. 3b), the conversion is around 47% for 500~ and 49% for 550~ It has to be stressed that this type of influence of temperature on the 1,3,5-TIPB conversion was observed systematically for 3-, 5-, 7- and 10-sec. reaction times. Regarding the above-described phenomena, it is postulated that this is the result of the competition between 1,3,5-TIPB diffusion and catalytic cracking. At lowtemperatures, the 1,3,5-TIPB diffusion is the controlling step while at higher temperatures the intrinsic 1,3,5-TIPB reaction becomes the controlling rate mechanism. In this respect, given that the 1,3,5-TIPB molecule has a critical diameter of 9.5 A [15], it can be expected that diffusional constrained molecules evolve in zeolites with 7.4 A cage openings. This type of configurational constrained diffusion leads to molecules evolving with high interaction with zeolite crystal walls [2].

284

50. A

40

etm

r-

50

o I~ 3 0 .

40

o

o

tO o

"0 u

m 20

IZl

0. l--

eL 20 i.-

~

1 tt

e-

30

c'~

10

0

0 300

350

400

450

500

550

300

600

350

400

450

500

550

600

Temperature ( C )

Temperature (C)

Fig. 3a. Change of 1,3,5-TIPB conversion with temperature for CAT-SC and CAT-LC. Reaction time=3 sec. C/O=5. (.) CAT-SC, (m) CAT-LC.

Fig. 3b. Change of 1,3,5-TIPB conversion with temperature for CAT-SC and CAT-LC. Reaction time=5 sec. C/O=5. (.) CAT-SC, ( I ) CAT-LC.

In agreement with this, in the 350-450~ range and for 3-, 5-, 7- and 10 sec. contact times the small Y-zeolite crystal gives higher 1,3,5-TIPB conversions than the larger crystals. These results are consistent with those published by Aguiar et al [16] who carried out catalytic cracking experiments at 380~ While Aguiar et al [16] interpreted these results on the unlikely basis of a larger external surface area available for the smaller zeolite crystals, this is in our view the consequence of differences in diffusional constrains. Diffusional transport affecting Y-zeolite performance is further supported by the high activation energy for the 1,3,5-TIPB diffusion, in the 70-80 KJ/mole range [17], and the much lower activation energy for alkyl-aromatic cracking, 15-20 KJ/mole [18]. Thus, the activation energy for diffusion is expected to be much larger than the one for catalytic cracking. As a result, at lower temperatures the overall reaction is diffusionally limited and it is strongly affected by changes in the thermal levels. Increasing the temperature, however, makes the rate of 1,3,5-TIPB diffusion to increase proportionally more than the intrinsic catalytic cracking rate. The overall reaction becomes then controlled by the intrinsic kinetics. Consistent with this there is, in the lower temperature, a pronounced effect of crystal size on the 1,3,5-TIPB conversion with this effect becoming less significant at higher temperatures. This observation is consistent with the potential changes of the modified Thiele modulus [19]. For catalytic cracking of model compounds, a modified Thiele modulus ~ can be defined as,

h'- 1 Ikin~intP:

1 /kin,og)intPce[_(Zl_E2)/2RT]

aext ------~~eff=aext~

(1)

Deft,~

with a e x t (l/m) being the external specific surface area of the zeolite crystal, k i n the intrinsic kinetic constant (m3/Kgcat.s), Deft the effective diffusivity (m2/s), ~b the

285 catalytic activity decay function, E1 the activation energy for the intrinsic catalytic cracking reaction and E2 the activation energy for the effective hydrocarbon diffusivity. In fact, under conditions where El<E2 the very important result for FCC catalysts of raising temperatures reducing the modified Thiele modulus and consequently increasing the effectiveness factor, ~1= tan h(h')/h' can be supported. Thus, trends on h' and 11 allow to predict that by augmenting the reaction temperature the cracking reaction evolves from a diffusionally controlled regime to a chemical controlled regime. As well, given the dependence of the h' with the external surface area, it can be expected that, as it was observed in the present study, transition between regimes will take place, for the smaller crystals (smaller h') at lower temperatures. Figs. 4a and 4b present a description of the changes of various product lumps with reaction time at 400~ This allows to postulate that the catalytic cracking of 1,3,5TIPB is a reaction network with three prevailing in series reactions (Fig. 5): a) dealkylation of 1,3,5-TIPB to form 1,3-DIPB and propylene, b) de-alkylation of 1,3-DIPB to give cumene and propene, c) de-alkylation of cumene to form benzene and propene. While the above mentioned steps are the dominant ones, there are other reactions such as disproportionation, isomerization and condensation [20,21] that may affect gas phase product distribution and coke formation. The extent of these steps may depend on both temperature and catalyst properties [20].

14. 12-

24 -

10-

~. 6 4_ 2. Oq 0

9

-

e

~

18"o ~. 12-

~

6 0 2

4

6

8

10

Time (sec)

Fig. 4a. Change of product yields with reaction time at 400~ ([]) propylene (A) benzene, (-)cumene, ([7) 1,3-IPB.

0

2

4

6

8

10

Time (sec)

Fig. 4b. Change of product yields with reaction time at 550~ ([]) propylene (A) benzene, (-) cumene, ([3) 1,3-DIPB.

286 1,3,5 TIPB

Iv"

ram"

1,3 DIPB

~

I I I I I I !

~

cumene --------+ benzene

~,~

~

~

~

I

~,

I I I

~

I

I

I I I I I I I I I I I

~

%%%

~

I

%%%

I I

"--24+ -I~ coke +

"

propylene

t.

Fig. 5. Schematic description of the catalytic cracking of 1,3,5-TIPB.

Thus, given the nature of the 1,3,5-TIPB cracking reaction it is important to examine the influence of operating conditions and crystal size on product selectivity. For instance, it was observed that the 1,3-DIPB selectivity (first de-alkylation step) increases with reaction time and decreases with both temperature and crystal size. As well, Figs. 4a and 4b show that cumene selectivities were always much higher than 1,3-DIPB selectivities (almost 5 times). This finding was assigned to the higher crackability of the 1,3-DIPB molecule versus the cumene molecule. Regarding the changes of 1,3-DIPB selectivity with 1,3,5-TIPB conversion it was found that 1,3-DIPB selectivity increases first and decreases later with 1,3,5-TIPB conversion (or reaction time) yielding a maximum selectivity at 7 sec. (Figs. 6a and 65).

6

12

~5

0 0

:_~

s

-~

6

9~

4

o

[]

3 m 2 a

9

u)

m n

4

m

9

2

[]

TO

0

20 1,3,5

40

60

TIPB conve rs ion(%)

Fig.6a. Change of 1,3-DIPB selectivity with 1,3,5-TIPB conversion. T=400~ C/O=5. (-) CAT-SC, ([]) CAT-LC.

0 o

~o

,o

60

~o

1,3,5 TIPB conversion (%)

Fig.6b. Change of 1,3-DIPB selectivity with 1,3,5-TIPB conversion. T=450~ C/0=5. (*) CAT-SC, ([]) CAT-LC.

Fig. 7 reports the effect of both temperature and crystal size on 1,3-DIPB selectivity with 1,3-DIPB selectivity diminishing with temperature and zeolite crystal size. It should be noticed that CAT-LC consistently yielded lower 1,3-DIPB selectivities than CAT-SC with these differences decreasing while temperature

287 increased. This is in agreement with a significant reduction of transport limitations in the CAT-SC catalyst and as a consequence of an important increase of the dealkylation of 1,3,5-TIPB rate constant with the rate constants for the other steps remaining essentially unchanged.

30

12

9 9

25 > o

.u

e Ell

20

[]

[] 15-

w

m

4

~

2

10-

E o

0

50

350

400

450

500

Temperature

550

600

(~

Fig. 7. Change of 1,3-DIPB selectivity with temperature.Reaction time-5 sec, C/O=5, (~ CAT-SC, ([]) CAT-LC.

-

o

,

-

.'o

6'0

80

1,3,5 TIPI3 Conversion(%)

Fig. 8. Change of cumene selectivity with 1,3,5-TIPB conversion for CAT-SC and CAT-LC. Temperature =450~ C/O=5.(~ CAT-SC ([]) CAT-LC.

Cumene is formed through double 1,3,5-TIPB de-alkylation with further cumene de-alkylation yielding benzene and propylene [21]. Since both n-propylbenzene and ethyltoluene were detected in the reaction product, it is highly probable that there is some contribution of cumene isomerization. In addition, cumene molecules can disproportionate to form DIPB or TIPB [20,21] with the observed cumene selectivity being a balance between cumene formation and cumene consumption reactions. Regarding the changes in cumene selectivity, this study shows that cumene selectivity increases with 1,3,5-TIPB conversion and contact time. However, cumene selectivity reaches a maximum level at around 7-sec. time and then decreases (Fig. 8). Fig. 8 reports results for 450~ and for both CAT-LC and CAT-SC, showing that cumene selectivity is affected by crystal size. Regarding changes of cumene selectivity with temperature, Fig. 9 reports that CAT-LC produces in the 400-500~ lower cumene than CAT-SC. Benzene selectivity changes with temperature and reaction time are reported in Figs. 4a and 4b. As contact time reaches 10-sec, cumene undergoes further dealkylation giving more benzene and propylene and increasing consequently benzene selectivity. Benzene selectivity and its change with conversion and crystal size are reported in Fig. 10.

288

35 ,,.-,.

3o

[]

20

[]

O

[]

0 350

400

450

500

550

600

Temperature (~

Fig. 9. Change of cumene selectivity with temperature for CAT-SC and CAT-LC. Reaction time=5 sec., C/O=5. (~ CAT-SC, ( 9 CAT-LC. Fig. 11 describes the effect of both temperature and crystal size on benzene formation. For both CAT-SC and CAT-LC benzene selectivity remains low in the lower temperature range, at 350-450~ the 1,3,5-TIPB cracking reaction is in good progress but still incomplete yielding mainly cumene and propylene. However, benzene selectivity augments significantly at 450-550~ Regarding the effects of crystal sizes, it can be observed that CAT-SC displays at 400-450~ much lower benzene selectivity than the CAT-LC (Fig. 11) with both crystals yield similar benzene levels at 550~ Thus, it can be concluded that catalytic cracking of 1,3,5-TIPB at lower temperatures and using CAT-SC reduces the benzene selectivity from 16% to 12%.

35.

[]

30.

|

mmm O []

,~

[] []

25-

> o o o r o

.o

1,3,5 TIPB conversion (%)

Fig. 10. Change of benzene selectivity with the 1,3,5-TIPB conversion. Temperature =450~ C/O=5. (o) CATSC, ([]) CAT-LC.

[]

20.

[]

15.

o N C

10.

m

5.

0

0

tl

[]

[]

[] ....

350

400

450

,

,,

500

550

600

Temperature ( ~ )

Fig. 11. Change of benzene selectivity with temperature. Reaction time =5 sec, C/O=5. (o) CAT-SC, ( 9 ) CAT-LC.

289 This intrinsic ability of the smaller crystals of allowing primary reactions, in this case the cracking of 1,3,5-TIPB, to occur more readily can be exploited as a new approach for the engineering of catalysts and catalytic cracking processes yielding environmentally friendly gasoline (lower benzene content). Regarding propylene/benzene, a theoretical molar ratio of 3 (1.62 weight ratio) is the one t h a t corresponds to the complete conversion of 1,3,5-TIPB into benzene and propylene. Deviation of this stoichiometric ratio may indicate the influence of other reactions. For instance, a propylene/benzene ratio less than 1.62 at 550~ can be considered an indication of an increased importance of the disproportionation reactions, with perhaps two or even three cumene molecules producing one or two molecules of benzene and DIPB or TIPB [20,21]. In this respect, Frilette et al [22] reported a deficiency in propylene at T=470~ when comparing propylene/benzene formed in cumene cracking. On the other hand if this ratio exceeds 1.62, with cumene being a main product, this shows the contribution of partial de-alkylation of 1,3,5TIPB. The CAT-SC showed at 500~ and 5-sec. propylene/benzene ratios higher t h a n the ones for the CAT-LC: 1.9 versus 1.7. This difference was even more pronounced at 350-450~ with the CAT-SC almost doubling at 400~ and 5-sec. the propylene/benzene ratios observed for the CAT-LC: 7.1 versus 3.8. This indicates that as crystal size increases the extent of de-alkylation of alkyl-aromatics tends to be more pronounced yielding a higher fraction of benzene and propylene. Disproportionation and trans-alkylation are important reactions for interconversion of aromatics, especially for di-alkyl-benzens and benzene production [20,21]. Thus, cumene consumed by disproportionation reaction gives more benzene and different TIPB isomers. In this respect, the CAT-LC showed to be more active for this kind of reaction. For example, at 350-450~ the CAT-LC produced higher levels of disproportionation products than the CAT-SC with these fractions being respectively 6 wt% and 4.7 wt% of the converted 1,3,5-TIPB. These fractions decreased, however, with temperature and this is considered an indication of a declining influence of disproportionation reactions. Regarding the complementary studies developed with cumene, it is interesting to note t h a t cumene catalytic cracking is a test reaction frequently used to evaluate catalyst performance. Cumene cracks forming benzene and propene [21] with propylene/benzene ratio being smaller than one and this given the contribution of disproportionation and isomerization [20]. Cumene cracking was developed in the Riser Simulator at 3-10 sec., 400-550~ and a set C/O ratio of 5. It was found t h a t cumene conversion increased with both contact time and temperature (Figs. 12a and 12b). For example, at 3-sec. contact time, cumene conversions were 20 and 33% for 400 and 550 ~ respectively. Cumene conversions at 10-sec. were 25 and 50% for 400 and 550 ~ respectively. Given t h a t cumene critical diameter is 6.8A and the Y-zeolite opening is 7.4A, the cumene cracking proceeds with little influence of transport constrains. In this respect, propene and benzene conversions found in the present studies were very close for both CAT-LC and CAT-SC and conditions in the 3-10 sec and 400-550~ range (Fig. 13). Thus, no significant effect was found in terms of effects of zeolite crystal size on cumene conversion.

290

45

35

40

II

30-

~ 35 g

~" ao

25-

II

20-

8 20 ~ 10.

g,

5-

5 0 350

0

400 450 500 550 Temperature (~

600

Fig. 12a. Change of cumene conversion with temperature Reaction time = 3 sec., C/O=5. (~ CAT-SC, ( 9 CAT-LC.

350

,oo

~5o

500

5~o

60o

Temperature (~

Fig. 12b. Change of cumene conversion with temperature. Reaction time = 5 sec., C/O=5. (~ CAT-SC, ( e ) CAT-LC.

Cracking of cumene yields, in addition to propene and benzene, other products such as cymene (methylcumene), toluene, ethylbenzene, n-propylbenzene and 1,3-DIPB and its isomers. Regarding cymene, its formation is reported to take place through disproportionation between two cumene molecules forming cymene and ethylbenzene. At 400~ cymene selectivity with CAT-LC was 16% and 9.3% for 3 sec and 10 sec respectively. At 550~ cymene selectivity was 8.2% and 3.13% for 3 sec and 10 sec respectively. It was observed, however, that the cymene/ethylbenzene ratio was much larger t h a n 1 [21] and this was the result of further ethylbenzene reaction with formation of toluene or ethyltoluene or simply ethylbenzene de-alkylation to benzene and ethylene. DIPB was also an identified product with the m-DIPB/p-DIPB ratio being, at 400~ in the 2.48-2.68 range for 3 sec-10 sec reaction time. However, it was noticed t h a t this ratio decreased with temperature. For example, at 500~ these ratios were 1.25 and 1.5 for 3-sec. and 10-sec. respectively, while at 550~ they were 0.333 and 1.28 for 3-sec. and 10-sec. respectively. Isomerization can also be an important reaction during cumene cracking [21] having n-propylbenzene formed in appreciable amounts. Selectivity for n-propylbenzene was found to oscillate between 1 and 1.5 and this for various reaction times, different temperatures and for both catalysts.

291

7O =~" 60 el O

~

40

o 1-

| N

30

C O

m

20 0

lo

20

3'0

40

Cumene conversion(%)

Fig. 13. Change of benzene selectivity with cumene conversion. Reaction time - 5 sec, C/O=5. (.) CAT-SC, ([]) CAT-LC. In summary, the reaction of cumene cracking on CAT-LC and CAT-SC catalysts demonstrate that changes in Y-zeolite crystal size do no affect the cumene cracking reaction yielding the same overall conversion and cracking products.

5. C O N C L U S I O N On the basis of the results of the present study the following can be concluded: a. Investigation of catalytic cracking of 1,3,5-TIPB and cumene is developed in a novel CREC Riser Simulator under different reaction times and various temperatures using Y-zeolites of different sizes: CAT-SC (0.4-~m) and CAT-LC (0.9-~m). b. Catalytic cracking of cumene illustrates that operation of Y-zeolites, free of diffusional constraints, yields for both CAT-SC and CAT-LC essentially identical conversions and selectivities. c. It is found that for both zeolites the 1,3,5-TIPB conversion increases, at a set C/O ratio, with reaction time. Moreover, at low temperatures (350-450~ the smaller size zeolite crystals give higher 1,3,5-TIPB conversions. Differences in the 1,3,5TIPB conversions are, however, less significant at 500-550~ It is speculated that this phenomenon is the result of 1,3,5-TIPB limited diffusional transport in the Yzeolites. d. Regarding cumene and benzene, main 1,3,5-TIPB cracking products, it is found that their yields decrease and increase respectively with both crystal size and temperature. This intrinsic ability of the smaller crystals of allowing primary reactions, in this case the cracking of 1, 3, 5-TIPB, to occur more readily can be exploited as a new approach for the engineering of catalysts and catalytic cracking processes yielding environmentally friendly gasoline (lower benzene content).

ACKNOWLEDGMENTS The authors would like to express our appreciation to the Saudi Arabian Government that provided a postgraduate scholarship to Mr. S. A1-Khattaf. We are very grateful for the financial support of the Natural Sciences and Engineering Research

292 Council of Canada. We also would like to acknowledge Dr A. H u m p h r i e s from Akzo Nobel and J. Macaoay from Engelhard companies, who helped with the catalyst preparation used in the present study.

NOTATION aext

dr Deft h' kin

E1 E2 R T (Pint

pc q

specific external surface area of the zeohte crystals (l/m) zeolite crystal equivalent diameter (m) effective diffusivity for 1,3,5-TIPB (m2/s) modified Thiele modulus (-) intrinsic kinetic constant for cracking of 1,3,5-TIPB (m3/Kgcat.s). energy of activation for the cracking reaction (kJ/mole). energy of activation for the hydrocarbon diffusion (kJ/mole). universal gas constant (kJ/mol.K). t e m p e r a t u r e (K). intrinsic deactivation function (-). zeolite density (kg/m3). effectiveness factor (-).

REFERENCES 1. P.B. Venuto, E.T. Habib, Fluid Catalytic Cracking with Zeolite Catalysts, Marcel Dekker, New York, 1979. 2. J. Karger, D.M. Ruthven, John Wiley & Sons, Inc, 1992. 3. J. Maselli, A. Peters, Catal. Rev.-Sci. Eng., 26(3&4) (1984) 525. 4. A. Humphries, J. Wilcox, Oil & Gas Journal, Feb.6 1989, 45-51. 5. R.M. Billimoria, P.K. Ladwig, T.A. Cavanaugh, and A.Y. Hu, Presented at King Fahd University of Petroleum and Minerals, Dhahran, Saudi Arabia, Dec. 4-5, 1994. 6. J. Tsikoyiannis, J. Wei, Chem. Eng. Sci, 46 (1). (1991) 233. 7. P. O'Connor, A.P. Humphries, Am. Chem. Soc. Div. Petrol. Chem. Preprints. 38(3) (1993), 598. 8. D.M. Nace, Ind. Eng. Chem. Prod. Res. Dev., 9 (1970) 203. 9. K. Rajagopalan, A.W. Peters, G.C. Edwards, Applied Catalysis, 23 (1986) 69. 10. A. Gianetto, H.I. Farag, P.B. Alberto, H.I. de Lasa, Ind. Eng. Chem. Res. 33(1994) 3053. 11. K. Roos, A. Liepold, H. Koch, W. Reschetilowski, Chem. Eng. Technol. 20 (1997) 326. 12. S. Al-Khattaf, H. I. de Lasa, Ind. Eng. Chem. Res. 38 (1999) 1350. 13. H.I. de Lasa, USA Patent 5, 102 (1992) 628. 14. J. Pruski, MESc Thesis, Univ. Western Ontario, London, Ontario, Canada, 1995. 15. S. Batia, Zeolite Catalysis: Principles and Applications, CRC Press, Inc, Boca Raton, Florida, 1989. 16. E.F. Aguiar, M. L. Murta Valle, M.P. Silva, D.F. Silva, Zeolites, 15 (1995) 620. 17. D.M. Ruthven, B. Kaul, Ind. Eng. Chem. Res. 32 (1993) 2053. 18. D.W. Kraemer, Ph.D Dissertation, University of Western Ontario, London, Ontario, Canada, 1991. 19. J. Smith, Third Edition, McGraw-Hill, Inc, 1981. 20. T. Tsai, S. Liu, I.Wang, Applied Catalysis. 181 (1999) 355. 21. A. Corma, B. W. Wojciechowski, Catal. Rev.-Sci. Eng., 24 (1982) 1. 22. V.J. Frillette, P.B. Weisz, R.L. Golden, J. CataI., 1(1962) 301.

Studies in Surface Science and Catalysis 134 M.L. Occelli and P. O'Conner (Editors) 9 2001 Elsevier Science B.V. All rights reserved

293

On the m e c h a n i s m o f f o r m a t i o n o f o r g a n i z e d m e s o p o r o u s silica that m a y be used as catalysts for F C C R. Zana a, J. Frasch b, M. Soulardb, B. Lebeau b, and J. Patarin b a Institut Charles Sadron (CNRS-ULP), 6 rue Boussingault, F-67000 Strasbourg, France # b Laboratoire de Mat6riaux Min6raux (CNRS-ENSCMu), 3 rue Alfred Werner, F-68093 Mulhouse, France

Organized mesoporous silica (OMS) constitute new solids with many potential applications, particularly in the field of catalysis. These materials with very well organized pores can be obtained upon lowering the pH of highly basic sodium silicate solutions in the presence of cationic surfactants such as cetyltrimethylammonium bromide (CTAB) or chloride (CTAC). The first steps of formation (prior to precipitation) of these materials have been investigated by in situ probing fluorescence techniques. The results for the CTAB systems indicate that only a small fraction of micelle-bound bromide ions is exchange by hydroxyl and silicate ions. Moreover, the presence of these additives increases very little the micelle aggregation number in CTAB and CTAC-containing systems, indicating that there is hardly any micelle growth under the experimental conditions used. Keywords: Organized mesoporous silica, surfactants, spectrofluorometry, time-resolved fluorescence quenching

1. INTRODUCTION A new family of mesoporous molecular sieves with regular and constant pore diameters in the range 1.5-10 nm, designated as M41S, was synthesized in 1992 (1,2). These organized mesoporous silicate-based (OMS) materials constitute new solids with many potential applications, particularly in the field of catalysis. These exciting materials were prepared from silicate-based mixtures in the presence of surfactants as structure-directing agents. Extensive investigations have been carried out to study the mechanism of formation of these materials (2,3-8). Most of the mechanisms proposed thus far are based on the anionic exchange between surfactant counterions and silicate anions. In this study, we have investigated by means of fluorescence probing techniques systems made up of CTAB and CTAC (cetyltrimethylammonium bromide and chloride) and large excesses of sodium silicate and sodium hydroxide, "R. Zana: Phone number 33 3 88 41 40 42

294 with overall compositions similar to those used when preparing OMS. Our purpose was to test whether the surfactant counterions are effectively exchanged by silicate anions and if, under these conditions, the presence of silica induced micelle growth.

2. EXPERIMENTAL SECTION The experiments were performed at a fixed concentration of 0.1 M CTAB or CTAC in pure water and upon addition of first NaOH and then SiO2 (water glass, Na2Si307). The pH was decreased from 13.6 to 11.6 (clear solutions) in the CTAC system by adding HC1 or replacing a part of NaOH by NaC1. These experimental conditions (decrease of the pH) are very close to those used for the synthesis of OMS materials. Experiments were also performed using octameric silicate species (Si8020)8 as silica source. The latter species (D4R units) were stabilized by adding tetramethylammonium hydroxide (TMAOH) and methanol (MeOH) to the silicate solution. The presence of the D4R units was evidenced by 29 Si liquid-state NMR spectroscopy. Two different fluorescent molecules were used in the fluorescence probing experiments: pyrene alone (spectrofluorometry) or in conjunction with the fluorescence quencher cetylpiridinium chloride (time-resolved fluorescence quenching), and dipyrenylpropane (spectrofluorometry). The following parameters were obtained from these fluorescence studies: (i) the micelle micropolarity and microviscosity from the emission spectra of pyrene (I1/I3 intensity ratio) and dipyrenylpropane (IM. . . . . JIExcime r intensity ratio, referred to as IM/IE), respectively, (ii) the micelle aggregation number (N) and the pyrene fluorescence lifetime ('0 from the fluorescence decay curves of micelle-solubilized pyrene.

3. RESULTS AND DISCUSSION The results from experiments using water glass as silica source are listed in Table 1. From the Iz/I3 and IM/IE intensity ratios and for systems free of methanol, it appears that no significant change of the micropolarity and only a slight increase of the microviscosity are observed upon lowering the pH. Concerning the micelle aggregation number N, the values seem to be significantly larger for CTAB than for CTAC-containing systems in the absence of any additive (see experiments 1 and 5). However, for these systems (flee of methanol) the variations of N upon addition of NaOH and of NaOH + silica are qualitatively similar for the CTAB and CTAC solutions and in all cases correspond to quasi-spherical micelles. Thus, the addition of 0.5 M NaOH brings about a very small increase of N which is likely the result of two antagonistic effects: (i) a decrease of N due to the exchange of a small part of micellebound bromide ions by added hydroxyl ions; (ii) an increase of N associated with the increased ionic strength of the system due to the added NaOH.

295 Table 1 Systems CTAB/SiOJNaOH/(TMAOH/MeOH), CTAC/SiO2/NaOH/NaC1, and CTAB/SiO2/ TMAOH(/MeOH): Values of the pyrene fluorescence intensity ratio I1/I3 (micropolarity), pyrene fluorescence lifetime x, micelle aggregation number N, ZEIM/IE (microviscosity), and pH. N~ systems (stoechiometry in M) I1/I3 z (ns) N XEIM/IEb( pH ns) Systems with water glass as silica source 8.3 (1) a 0.1 CTAB 1.35 165 145 572 13.6 (2) a 0.1 CTAB : 0.5 NaOH 1.35 177 149 569 12.9 (3) a 0.1 CTAB : 0.5 SiO2 : 0.9 NaOH 1.35 180 161 650 (4) a 0.1 CTAB :0.5 SiO2 : 0.3 NaOH : 0.7 260 100 TMAOH : 21% MeOH 7.0 372 (5) 0.1 CTAC 1.38 155(342a) 103(101a) 13.6 373 (6) 0.1 CTAC : 0.5 NaOH 1.36 165 106 13.2 414 (7) 0.1 CTAC :0.5 SiO2:0.9 NaOH 1.36 190(342 a) 113(116a) 12.6 476 (8) 0.1 CTAC :0.5 S i O 2 : 0.6 NaOH : 0.3 NaC1 1.35 175 154 12.1 (9) 0.1 CTAC : 0.5 SiO2 : 0.5 NaOH : 0.4 NaC1 1.36 177 157 11.6 (10) 0.1CTAC:0.5SiO2:0.4NaOH:0.5NaC1 1.36 172 152 12.5 (11) 0.1 CTAC : 0.5 SiO2 : 0.9 NaOH : 0.3 HC1 1.36 175 143 11.6 (12) 0.1 CTAC : 0.5 S i O 2 : 0.9 NaOH : 0.5 HC1 1.34 171 148 Systems with D4R as silica source (13) a 0.1 CTAB :21% MeOH c 1.29 195 81 1.21 196 127 (14) a 0.1 CTAB : 1.0 TMAOH (15) a 0.1 CTAB : 1.0 TMAOH : 21% MeOH c 1.23 247 77 272 96 (16) a 0.1CTAB :0.5 SiO2:1.0TMAOH :21%MeOI-F 1.13 a deaerated systems. b the product "I:EIM/IE ('1~ E - - lifetime of the dipyrenylpropane excimer) is proportional to the viscosity of the environment of the probe molecule. c value in volume percent.

The values of the pyrene fluorescence lifetime x in deaerated solutions of CTAB and CTAC (systems 1 and 5 in Table 1) are very different, 165 vs 342 ns, owing to the efficient quenching of the pyrene fluorescence by the micelle-bound bromide ions, micelle-bound chloride ions having no quenching effect (9,10). Micelle-bound hydroxyl ions also have no quenching effect on the pyrene fluorescence since x has nearly the same value in deaerated solutions of CTAC and of cetyltrimethylammonium hydroxide (CTAOH) (11). The 9 value for system 7 in Table 1 shows that silicate anions also do not quench the pyrene fluorescence. Indeed, the x values in deaerated solutions of CTAC and of CTAC + NaOH + sodium silicate (systems 5 and 7 in Table 1) are identical (342 ns) whereas, any quenching by silicate anions would have resulted in a lower value of 9 in system 7. The quenching of the pyrene fluorescence by the bromide ions can thus be used to detect a possible exchange of micellebound bromide ions by added hydroxyl or silicate anions (or by any non-quenching counterion for that matter). Indeed, when such an exchange occurs, the concentration of

296 micelle-bound bromide ion decreases and the pyrene lifetime increases from the value z(Br) = 165 ns in the absence of exchange to ~(C1) = 342 ns in the case of complete exchange upon addition of hydroxyl, silicate, or chloride ions. As an example, from the z value of the experiment 3 (without methanol) and according to reference (12) the fraction of exchanged bromide ions is around 16%. The second set of experiments used octameric silicate species (D4R units) instead of water glass (see Table 1, systems 13-16). The results show that the presence of TMAOH and MeOH results in lower values of I1/I3, i.e. of micelle polarity. The presence of MeOH also causes a decrease of the value of the aggregation number (compare systems 1 and 13, or 14 and 15), a well known behavior (13). Last the values of the lifetimes (~) in these experiments are all larger than for the comparable experiments in the absence of MeOH and TMAOH. However, this is essentially due to the presence of MeOH and TMAOH and not to the nature of the silicate species. Indeed, a similarly large increase of lifetime was measured in system 4 which contains MeOH and TMAOH but where the silica is not under the form of octamers, as indicated by the 29 Si liquid-state NMR experiment performed on this system. In the latter case, the absence of D4R units is due to the presence of sodium cations.

4. CONCLUSION These results conclusively show that the introduction of hydroxyl and silicate anions into micellar solutions of CTAB results in the exchange of only a small fraction of micelle-bound bromide ions by the added ions. This is far from the assumption made in the current models of formation of organized mesoporous silica which imply that silicate anions substitute nearly all surfactant counterions. The question which then arises in view of our results concerns the step where the surfactant/silica interactions take place, resulting in the organization of the system and in the formation of organized mesoporous silica. We feel that the most important step in the process is the formation of prepolymers of silica in regions of the system where, upon addition of HC1, there is a large local excess of HC1 which induces the polymerization of silicate anions. The interaction of these prepolymers with surfactant micelles maybe responsible for micelle growth and subsequent reorganization of the silica/micelle complexes. The coupling of these processes to further silica polymerization in the complexes may ultimately lead to the mesoporous silica. At this stage of the study we cannot exclude that interactions may occur between free (as opposed to micellized) surfactant ions and silicate anions, with the micelles acting as reservoirs of surfactant ions, as suggested by Huo et al. (5).

REFERENCES 1. C.T. Kresge, M.E. Leonowicz, W.J. Roth, J.C.Vartuli, and J.S. Beck, Nature, 359 (1992) 710. 2. J.S. Beck, J.C. Vartuli, W.J. Roth, M.E. Leonowicz, C.T. Kresge, K.D. Schmitt, C.T.W. Chu, D.H. Olson, E.W. Sheppard, S.B. McCullen, J.B. Higgins, and J.L. Schlenker, J. Am.

297 Chem. Soc., 114 (1992) 10834. 3. A. Monnier, F. Scht~th, Q. Huo, D. Kumar, D. Margolese, R.S. Maxwell, G.D. Stucky, M. Krishnamurty, P. Petroff, A. Firouzi, M. Janicke, and B.F. Chmelka, Science, 261 (1993) 1299. 4. C.Y. Chen, S.L. Burkett, H.X. Li, and M.E. Davis, Microporous Mater., 2 (1993) 27. 5. Q. Huo, D. Margolese, U. Ciesla, D.G. Demuth, P. Feng, T.E. Gier, P. Sieger, A. Firouzi, B.F. Chmelka, F. Schath, and G.D. Stucky, Chem. Mater., 6 (1994) 1176. 6. J.C. Vartuli, C.T. Kresge, M.E. Leonowicz, A.S. Chu, S.B. McCullen, I.D. Johnson, and E.W. Sheppard, Chem. Mater., 6 (1994) 2070. 7. J.S. Beck, J.C. Vartuli, G.J. Kennedy, C.T. Kresge, W.J. Roth, and S.E. Schramm, Chem. Mater., 6 (1994) 1816. 8. A. Firouzi, D. Kumar, L.M. Bull, T. Besier, P. Sieger, Q. Huo, S.A. Walker, J.A. Zasadzinski, C. Glinka, J. Nicol, D. Margolese, G.D. Stucky, and B.F. Chmelka, Science, 267 (1995) 1138. 9. E. Abuin, E. Lissi, N. Bianchi, L. Miola, and F.H. Quina, J. Phys. Chem., 97 (1983) 5166. 10. E. Abuin and E. Lissi, J. Colloid Interface Sci., 143 (1991) 97. 11. P. Lianos and R. Zana, J. Phys. Chem., 87 (1983) 1289. 12. R. Zana, J. Frasch, M. Soulard, B. Lebeau, and J. Patarin, Langmuir, 15 (1999) 2603. 13. R. Zana, Adv. Colloid Interface Sci., 57 (1995) 1.

Studies in Surface Science and Catalysis 134 M.L. Occelli and P. O'Conner (Editors) 9 2001 Elsevier Science B.V. All rights reserved

299

Catalyst Assembly Technology in FCC. Part I: A review of the concept, history and developments P. O'Connor a, p. Imhof b and S.J. Yanik c aAkzo Nobel Catalysts, Stationsplein 4, P.O.Box 247, 3800AE, Amersfoort, The Netherlands bAkzo Nobel Research Center Catalysts, Amsterdam, The Netherlands c Akzo Nobel Catalysts, Marketing Center Singapore 1. SHORT HISTORY OF FCC CATALYST TECHNOLOGY Revisiting the past from the perspective of the present is an ideal tool to inseminate the rethinking process. Unfortunately because of the frequently proprietary nature of the FCC catalyst manufacturing process, which is often not sufficiently protected by patents, the open literature in this field is relatively limited, and is mainly dominated by the development and application of the zeolites applied in FCC. We will digress from this well known path and try to focus more on the non-zeolitic aspects of FCC Catalyst technology. A.D. Reichle [1] gives an excellent review of the various types and forms of Cracking Catalysts which were developed and used in the early days (Table 1). The first catalyst, the Super Fitrol [1], was produced by Filtrol in their plant by activating various clays with acid, producing large pore dealuminated materials containing acid sites. These materials were originally used for bleaching edible oils and decolorizing hydrocarbons. Reichle describes the testing of many types of synthetic mixed oxide catalysts, some of which were 2 to 3 times more active than the activated clay based types. However, the stability was rather poor. Only the SiO2eAI203 and SiO2oMgO survived through the development stage into commercial testing. The optimum Al203 content and SiO2eAI203 interaction using impregnated dry SiO2 gels defined a broad maximum in activity covering the 10 to 25% Al203 range. The high AI203 catalysts (HA, 25% AI203) clearly exceeded the low AI203 catalysts (LA) and Super Filtrol steamed activity level. Co-currently Micro spheroidal (MS) catalysts were developed, recognizing the advantages for a higher alumina content catalyst and the improvements in impregnation efficiency with small (<100 microns) particles compared to the traditional lumps of silica hydro-gel. The original process of preparing MS catalysts was quite burdensome and involved the dispersion of silica sol in oil, using the oil viscosity, degree of agitation and the use of emulsifiers to produce the correct particle size and morphology. Special drying conditions without

300

Table 1. Early days of Cracking Catalysts Year

Process

Reactor System

Catalyst Type

Form

1920 1939

McAfee Houdry

Batch Fixed bed

AICI3 Clay

1940

Suspensoid Liquid phase Clay

1942

FCC

fluid bed

Clay

1945

Moving bed

Clay

1942

TCC, Houdryflow FCC

Granulated Acid treated Granulated Ex-luboil decolorizing. Powdered Super Filtrol Acid treated Powdered\ Acid treated Pellets

Fluid bed

1946

FCC

Fluid bed

Synthetic SiO2eAI203 Granulated Synthetic SiO2eAI203 Microspheres

damaging the catalysts were developed. The advantages especially in reduction of catalyst losses were however great. Eventually the emulsion process was replaced by spray drying of the impregnated gel. Spray drying still is today the way all FCC manufacturers compound and form their MS catalysts. The development of improved synthetic catalysts continued in the following years: Planck and Drake [2], for instance studied the effect of preparation conditions on the porous structure and physical properties of SIO2.AI203 gels. Braithwaite [3] describes the addition of Clay in the preparation of synthetic SiO2eAI203 catalysts to provide additional macro porosity. Magee and Blazek [4] explain how most of the first generation zeolite catalysts were based predominately on SIO2.AI203 gels ("Matrix") in which the zeolite is added at some point prior to spray drying: "When properly done (not as in Fig. 1), an extremely homogeneous dispersion of zeolite/matrix co-catalyst results; which contains an extensive micro- and macro pore structure". In the "ln-situ" crystallization method, as patented by Haden et al [5] and applied by Engelhard kaolin based microspheres are prepared and calcined, where after zeolites are crystallized in the microsphere, leaving zeolite in an "AI203-enriched matrix".

301

Fig. 1. Non-ideal incorporation of zeolite in SiOz*AIzO3 gel Secor [6] describes how Filtrol was producing a zeolite type catalyst based on reacting the acid activated Halloysite with silica, exchanging with rare-earth or magnesia, while using the AI203 gel leached out from the clay as a binder. Magee and Blazek [4] and Rajagopalan and Habib [7] state that additional macroporosity introduced in clay based or semi-synthetic clay containing catalysts does not significantly alter the catalytic performance. Elliot [8] and Ostermaier [9] introduce the use of a Silica hydro-sol based binder for the incorporation of zeolites. Silica hydro sol is a polymerized silica dispersed in water to form a clear continuos phase, which is a binder giving dramatic improvements in attrition resistance and density. The result is the creation of particles which are encased in a hard resilient shell of a vitreous material, as illustrated by the following artist impression (Fig. 2):

Fig. 2. Incorporation of zeolite in SiO2 Hydro-Sol The low activity of the silica hydro-sol relative to the SIO2.AI203 based systems enabled the selectivities of nearly pure zeolite cracking, resulting in significant improvements in gasoline and coke yields in Fluid bed and MAT test units, as reported by Rajagopalan and Habib [7].

302

Alafandi and Stamires [10] and Lim and Stamires [11] nevertheless continue to pursue the "Filtrol" route using AI203 gel and/or SiO2eAI203 gels as binder and as catalytically functional materials, stressing the significance of porosity and permeability of the microsphere. Young and Rajagopolan [12] fail to confirm the presence of diffusion restrictions in FCC. O'Connor and Van Houtert [13] however question the models used to set criteria for mass transfer limitations in the non-steady state FCC, and conclude from their short Contact time pilot riser experiments that combining zeolites with a "diffusion enhancing" matrix can result in significant product selectivity and product property improvements. The industrial benefits of non-zeolitic "matrix" on bottoms cracking in Heavy VGO and Resid FCC are confirmed by Godinho [14] and Hettinger [15]. Hettinger gives the example of the more porous semi-synthetics containing kaolin: The kaolin has little or no cracking activity, and in the laboratory (MAT test) catalyst activity was directly related to the silica-alumina gel content. However, the catalysts performed much better in commercial tests than anticipated. Occelli [16] explains the effect of kaolin using atomic force microscopy (AFM). AFM reveals that the FCC surface is formed by the aggregation of kaolin particles held together by the binder. Voids between clay platelets generate slit like openings that are mainly responsible for FCC microporosity. The occurrence of missing plates, misaligned plates and the occlusion of clay platelets aggregates yields much larger pores in the mesoporosity range.

2. CATALYST ASSEMBLY, ACCESSIBILITY AND STRENGTH 2.1 Catalyst Assembly Technology (CAT) Recently De Jong [17] introduced the term "Catalyst Assembly" for the actions (processes) to control the composition, structure and location of the active phases in three dimensions in a catalyst particle. In this sense the production of a solid catalyst (like FCC, see Fig. 3) resembles more the assembling of a mechanical watch or an electronic chip, rather than the making (mixing) of a homogeneous chemical compound. The interactions (processes) which occur during the catalyst particle preparation procedure will determine both the catalytic and physical properties of the catalyst:

Components

+ Process

Interaction v

C o m p o s i t e s + Structure

Mann et al [18,19] also raise the analogy between the assembly of engineered particle surfaces like electronic chips and FCC catalyst particle design. They again

303

Fig. 3. Assembly of FCC catalyst from the components. demonstrate that some of the present commercial FCC catalyst are very inhomogeneous and do not correspond in morphology or internal pore architecture to the ideal models we sometimes have in mind. Rijnten [20], Mann et a1118], De Jong [17], and Davies et al [21] demonstrate several aspects of what can be more generally described as "Particle Surface Engineering Technology". This field has progressed significantly in the last years with the advent of new analytical techniques that allow for a better understanding of the particle morphology and assembly at the micrometer and the nanometer scales. One of these novel techniques, the Low Melting Point Alloy (LMPA) penetration, method as described by Mann [22] can be effectively used to illustrate the inhomogeneous character of many FCC catalyst particles (Fig. 4). The foregoing visualization is achieved by penetrating the catalyst meso and macro pores with low melting point alloys, which are made visible by SEM.

2.2 Catalyst Accessibility O'Connor and Humphries [23] define the term Accessibility in FCC, by asserting that catalyst sites are accessible if they can be reached by the compounds which are supposed to interact with these sites within a certain time limitation as set by the catalytic process. They portray the relevance of accessibility for FCC, and highlight the importance of issues like contact time in testing, feedstock type, catalyst fouling by coke, metals and the effect of reverse-accessibility or "egressibility" on hydrocarbon entrainment and stripping efficiency.

304

Fig 4. Model of Alloy penetration of an "average" FCC particle. Falabella et al [24] apply the cracking of 1,3,5,triisopropylbenzene (TIPB) to study the activity and accessibility of zeolites, and conclude that both accessibility as well as acidity seem to control the rate of reaction. O'Connor et al [25] emphasize that, as a consequence of the trend towards reduced contact times in FCC between catalyst particles and hydrocarbons; the mass transfer rate of the hydrocarbons into the catalyst particle and the mass transfer rate of the product out of the catalyst particles will start to play a bigger role. Realistic short contact time tests are needed in order to evaluate the catalysts as the mass transfer characteristics are not adequately exposed in existing often relatively long injection time laboratory tests. O'Connor and Humphries [23] and Boot et al [26] claim the existence of new proprietary methods to characterize catalyst accessibility based on the non-steady state diffusion of hydrocarbons into FCC particles. Hodgson et al [27] demonstrate how these new techniques can be used to quantify catalyst accessibility and optimize catalyst effectiveness in an industrial FCC operation. They observe a significant loss of accessibility as quantified in an Accessibility Index (AAI) as a function of catalyst age and metals content. The rapid loss of AAI in the equilibrium catalyst with metals and also coke laydown was investigated, and although all catalysts are effected by this phenomenon, only the catalysts with a high fresh AAI, retained an acceptable value versus the age distribution. These catalysts are produced using the LA ("Filtror') Technology.

305

2.3. Catalyst Strength While the benefits of the improved permeability or accessibility of the LA catalysts are hard to establish in most of the used laboratory testing and deactivation protocols, the shortcomings in terms of attrition resistance are evident from any traditional attrition test. It was therefore extremely puzzling to see that our refiners, were using these type of catalysts in various of their industrial units with the clear benefits of improved bottoms cracking, but without any penalty in catalyst losses. O'Connor and Pouwels [28] describe how by accident they discovered that in the non-ambient tests units like Cyclic deactivation units and Pilot risers there was a distinct different behavior in losses and fines generation versus attrition for the LA catalysts compared to other catalysts. Based on these observations a Hot Attrition Test (HAT) was developed to measure the attrition mechanisms at non-ambient and FCC like temperatures. As was suspected, the HAT confirmed the different behavior of LA catalysts. Many substances and certainly most porous materials will show a decrease of strength with increase of temperature. The collapse of a building in a fire is often caused by the decrease of the strength of the concrete structure and not by the burning of the structure. The same holds for FCC: As can be seen in Table 2 the attrition increases with increasing temperature. There is a difference in the degree of this increase in attrition for LA and other catalyst systems. For LA the change is small, for the other systems it is considerable. The result is that the relative high attrition of LA at room temperature is hardly increased at 700~ while the attrition of most other systems increases by about threefold.

Table 2

Classic (ambient) versus Hot Attrition results

CAT

AI-KTM Classic 20~

HAT Hot 700~

Comment

A1 A2 B C D(LA)

5 15 6 5 15

15 4O 18 15 15

Average Off spec product Average Average Average

306 3. RECENT DEVELOPMENTS IN CATALYST ASSEMBLY

Notwithstanding the often misleading feedback of conventional laboratory tests, and thanks to the puzzling industrial results with LA catalyst, we were able to focus catalyst development towards the real relevant catalyst goals in FCC. This is not to say that laboratory testing is not of importance, it surely is, but considering the complexity of simulating FCC we have to be modest in the interpretation realizing that each deactivation and/or test protocol only gives a partial indication of one of the catalyst characteristics. Last but not least laboratory results should be constantly scrutinized and compared to the actual industrial results. Some new concepts in Catalyst Assembly technology were conceived and developed. A new technology to be commercialized in is code-named "Jade". "Jade" is a distinct new Catalyst Assembly Technology which opens the flexibility to improve Catalyst Accessibility and/or Catalyst Strength, while enabling the incorporation of certain novel catalytic materials, which are not very compatible with the present FCC technologies. The first "Jade" prototype materials have been produced while the first industrial trials confirm the expectations for improvements in FCC catalyst performance. Co-currently we are presently performing the first industrial trials with a second new Catalyst Assembly Technology, code-named "Opal" which opens the prospect to introduce also new catalytic functionalities into the existing high Accessibility (LA) catalyst system. Some examples of performance test results are given, making use of various aspects of developments in Catalyst Assembly Technology. Example 1: Resid FCC Cracking.

The example in Table 3 shows how improvements in catalyst performance, which can be exploited either for converting more of the heavy resid bottoms, or by reducing the catalyst "severity" (= consumption) and hence costs. Table 3

Improved Resid Conversion

CAT: Zeolite type: Matrix type:

SiO2-sol ADZ55 ADM40

"LA . . ADZ70 ADM20

.

.

OpaI-A" ADZ70 ADM2060

Yields in %wt on feed @ constant sevedty (CT0=5.5) Conversion Gasoline

64.27 43.56

70.26 46.88

73.44 47.98

307

Light Cycle Oil Bottoms Coke

13.66 22.07 4.58

14.64 15.10 5.26

14.58 11.98 5.25

Zeolite and Matrix types as described by Humphries et al [29]. Catalyst: Deactivated in redox cycles with ~4000 ppm. Ni, ~2000 ppm V. Catalyst tested in a Short Contact (Fluid bed) Time Resid Test (SCTRT). Feedstock: Atmospheric Resid, 20%wt boiling above 530~ by HTGC.

Example 2: Higher VGO Cracking Activity. The example in Table 4 shows test results indicating the potential for increasing FCC conversion and/or reducing catalyst addition rates:

Table 4

Higher VGO Cracking Activity

CAT: Zeolite type: Matrix type:

SiO2-sol ADZ55 ADM40

"Jade-A . . . . ADZ55 ADM4060

Jade-B" ADZ55 ADM4060

Yields in %wt on fresh feed @ constant severity (CTO=4.0) Conversion Gasoline Light Cycle Oil Bottoms Coke

75.67 45.36 14.10 10.23 4.58

77.99 45.91 13.52 8.50 5.26

78.69 45.78 13.06 8.25 5.25

Catalyst: Deactivated in redox cycles with ~1000 ppm. Ni, ~1000 ppm V. Catalyst tested in a (long contact time) Fluid Bed Test (FST). Feedstock: Vacuum Gasoil, ~2%wt Sulfur and ~1000ppm total nitrogen.

Example 3: High Metal Resid FCC Cracking. In Table 5 similar improvements as in the previous examplkes are demonstrated, but now processing a high metal Resid feed, with 10 000 ppm metals on catalyst.

308

Table 5

Improved Resid Cracking

CAT: Zeolite type: Matrix type:

SiO2-sol ADZ50 ADM50

SiO2-sol ADZ50 ADM40.60

"Jade-E" ADZ50 ADM40.60

Yields in %wt @ constant sevedty (CT0=5.5) Conversion Gasoline Light Cycle Oil Bottoms Coke

69.02 44.74 16.27 14.71 6.26

74.12 46.02 14.69 11.19 7.99

NBC* 79.03 80.82 *) Net Bottoms Conversion = l O0-Bottoms-Coke

76.90 46.85 14.20 8.9 8.2 82.9

Catalyst: Deactivated in redox cycles with ~5000 ppm. Ni, ~5000 ppm V. Catalyst tested in a Short Contact (Fluid bed) Time Resid Test (SCTRT). Feedstock: Atmospheric Resid.

4. SUMMARY AND CONCLUSIONS The way in which FCC catalysts are built up from the separate components, being the Catalyst Assembly Technology (CAT) has had an important effect on the developments in FCC Catalysis. Advances in assembly and manufacturing technology will contribute further to improvements in catalyst performance as well as cost-effectiveness.

5. ACKNOWLEDGEMENTS H.N.J. Wijngaards, J.A. van Aken, F. Olthof, E. Brevoord, R. Smeink, R. Jonker, E.M. Berends, W. Veerman, E. Rautianen and many more of our colleagues of Akzo Nobel's Research Center Catalysts. Furthermore we wish to acknowledge the stimulating and fruitful contributions of our colleagues in Brazil at Petrobras and FCC SA.

309

6. REFERENCES

1. A.D. Reichle, "Early Days of Catcracking at Exxon" Paper G-2, Ketjen Catalyst Symposium 1988, May 29m, The Netherlands. H.J. Lovink editor. 2. C.J. Planck and L.C. Drake, J Colloid Science, 1947, vol. 2, p. 399. 3. D.G. Braithwaite, "Silica-Alumina Hydrogel- Kaolinite Catalysts and Processes for Preparation Thereof'. U.S.Patent 3,034,995 (1962). 4. J.S. Magee and J.J. Blazek, "Preparation and Performance of Zeolite Cracking Catalysts." ACS Monograph 171, ACS, Washington, DC, 1976, p. 615. 5. W.L. Haden, "Microspherical Zeolite Cracking Catalyst". U.S. Patents 3,657,154 and 3,663,165 (1972). 6. R.B. Secor, "Method for producing catalysts". U.S. Patents 2,935,463 (1960) and 3,446,727 (1969). 7. K. Rajagopalan and E.T. Habib, Hydrocarbon Processing, September 1992, p. 43. 8. C.H. Elliott, Jr., "Process for preparing a cracking catalyst". U.S. Patent 3,867,308 (1975). 9. J.J. Ostermaier and C.H. Elliott, Jr., "Process for preparing an attrition resistant zeolite hydrocarbon conversion catalyst". U.S. Patent 3,957,689 (1976). 10.H. Alafandi and D. Stamires, "Method of producing zeolitic catalysts with Silica Alumina Matrix". U.S. Patent 4,142,995 (1979). 11.J. Lim and D. Stamires, "Attrition resistant zeolitic catalysts". U.S. Patent 4,086,187 (1978). 12. G.W. Young and K. Rajagopalan, Ind. Eng. Chem. Process. Dev. 24 (1985) 995. 13. P. O'Connor and F.W. van Houtert, Paper F-8, Ketjen Catalyst Symposium 1986, May 25th, The Netherlands. H.Th. Rijnten and H.J. Lovink, editors. 14.M. Godinho Tavares and P. O'Connor, "Bottoms conversion in FCC Operations". First South American Ketjen Catalysts Seminar, September 22-24, 1985. H.J. Lovink, editor. 15.W.P. Hettinger, "Development of RCC Catalyst". ACS Symposium Series 375. ACS, Washington, DC, 1988. M.L. Occelli, editor. 16.M.L. Occelli, S.A.C. Gould, F. Baldiraghi and S. Leoncini, "AFM Imaging of a FCC Catalyst Surface before and after aging", in Fluid Cracking Catalysts, Marcel Dekker, 1998. M.L. Occelli and P. O'Connor, editors. 17. K.P. de Jong, Cattech, June 1998, p. 87. 18.R. Mann, A. AI-Lamy and A. Holt, Trans. I. Chem, 73 (A) 1995, p. 147. 19. R. Mann, K. Khalaf and A. AI-Lamy, "Evaluating pore structure and morphology of hydrocarbon conversion catalysts". ACS Symposium Series 634. ACS, Washington, DC, 1996. P. O'Connor, T. Takatsuka and G.L. Woolery, editors. 20.H.Th. Rijnten, "New Developments in Catalysis Research". Hand-out at Akzo Catalyst Symposium, June 1994, The Netherlands. 21.R. Davies, G.A. Schurr, P. Meenan, R.D. Nelson, H.E. Bergna, C.S.A. Brevett and R.H. Goldbaum, Adv. Matter 1998, 10, no. 15, p. 1264. 22. R. Mann, Trans. I. Chem. E, 71 (A) 1993, p. 551. 23. P. O'Connor and A.P. Humphries, "Accessibility of functional sites in FCC". ACS preprints Vol. 38, no. 3, p. 598, 1993.

310

24. E. Falabella S-Aguiar, M.L. Murta-Valle, E.V. Sobrinho and D. Cardoso, "Cracking of 1,3,5-triisopropylbenzene over deeply dealuminated Y zeolite" in Studies in Surface Science and Catalysis, Elsevier Science, Amsterdam, Vol. 97, 1995, p. 417. L. Bonneviot and S. Kaliaguine, editors. 25.P. O'Connor, F.P. Olthof and R. Smeink, "Accessible catalysts for short contact time cracking", in Fluid Cracking Catalysts, Marcel Dekker, 1998. M.L. Occelli and P. O'Connor, editors. 26.L.A. Boot, M. de Boer, P.H. Desai and P. O'Connor, Hydrocarbon Engineering, September 1998, p. 59. 27. M.C.J. Hodgson, C.K. Looi and S.J. Yanik, "Avoid excessive RFCCU deactivation, improve catalyst accessibility". Akzo Nobel Catalysts Symposium, June 1998, The Netherlands. S. Docter, editor. 28.P. O'Connor and A.C. Pouwels, "Realistic commercial catalyst testing in the Laboratory". Proceedings of 8th International Symposium on Large Chemical Plants, Royal Flemish Society of Engineers, Antwerp, October 1992. 29.A.P. Humphries, R.P. Fletcher and J.R. Pearce, AM-99-63, NPRA 1999 Annual Meeting, March 1999, San Antonio, Texas.

Studies in Surface Science and Catalysis 134 M.L. Occelli and P. O'Conner (Editors) 9 2001 Elsevier Science B.V. All rights reserved

311

Catalyst Assembly Technology in FCC. Part I1: The influence of fresh and contaminant-affected catalyst structure on FCC performance C. W. Kuehler

a, R. Jonker b, p. Imhof b, S.J. Yanik c and P. O'Connor u

aAkzo Nobel Catalysts, Houston, Texas, USA bAkzo Nobel Research Center Catalysts, Amsterdam, The Netherlands CAkzo Nobel Catalysts, Marketing Center Singapore dAkzo Nobel Catalysts, Amersfoort, The Netherlands

SUMMARY Historically, great emphasis has been placed on the influence of catalyst components (zeolite, matrix, contaminant traps) on the performance of FCC. Although physical qualities such as mesoporosity have been recognized as important, no significant attention has been paid to the influence of mass transfer barriers caused by various manufacturing processes and those induced by certain contaminants on the performance of FCC catalysts. Such barriers will become increasingly important in short contact time FCC. This paper provides examples of both manufacturing and contaminant induced mass transfer barriers and their influence on process performance. Results from a mathematical model illustrating the influence of such barriers are presented. Finally, catalyst architectures that better tolerate the effects of contaminant induced barriers are discussed.

1. INTRODUCTION Fluid catalytic cracking has evolved into a process in which hot, virtually carbon-free catalyst is intimately mixed with a petroleum-derived feedstock, then transported through a turbulent reaction zone with a reactant residence time of 1-4 seconds. The vapor products from the cracking reaction are efficiently separated from the catalyst utilizing a variety of ingenious devices, then cooled and fractionated in a distillation column. The catalyst, partially deactivated by coke formed during the cracking reaction, is first steam-stripped to remove occluded hydrocarbons on its surface and then regenerated by combustion in a separate reactor.

312

This relatively fast, efficient contacting of feed and catalyst contrasts with early reaction systems which featured poor feed/catalyst contacting, long reaction times, and poor catalyst/product separation. The benefits of modern reaction systems are improved selectivity for desirable products. However, because of the reduced reaction time and the elimination of non-selective conversion, an increase in reaction severity is required to maintain adequate conversion. This is achieved by increasing reaction temperature, catalyst-to-oil ratio, or catalyst activity. Typically, higher catalyst activity is achieved by increasing the content of active ingredients in the catalyst. In general, it has been assumed that reaction at the catalytic sites has been the controlling step in the overall reaction system. However, the trends in reactor and catalyst design described above suggest a reexamination of this assumption. Diffusion of the larger reactant molecules through the porous structure of the catalyst particle to the reaction sites could now be a more important factor in the determination of the overall reaction rate. In addition, there must be sufficient time for the reaction products to diffuse out of the catalyst particle. The diffusion must be as efficient as possible to prevent overcracking to undesirable light products. The term accessibility [1] has been used to describe these phenomena. If diffusion to and from the catalytic sites is important in determining the overall reaction rate, then the nature of the porous structure of the catalyst particle should be examined closely. The materials and method of manufacture of the catalyst can dramatically affect the porous structure. In addition to the distribution of pore sizes in the catalyst, the shape of the pores can also be important [1,2]. For the reactants to reach the interior of the catalyst particle, they must pass through the surface, so it is very important that diffusional restrictions be minimized there. In addition to diffusional restrictions created at the surface and interior of the asmanufactured catalyst, additional accessibility problems may develop as a result of contaminants deposited on the catalyst by the feed. These may naturally occur in the petroleum or be introduced during its recovery or refinery processing prior to the FCC. Common contaminants that have been studied extensively are Ni, V, and Na. In general, the focus of these studies has been on their effect on zeolite activity or undesirable reactions such as dehydrogenation catalyzed by the contaminant itself. There has been some discussion regarding the possible effects of V on accessibility. There has been significant study of how these contaminants distribute throughout a catalyst particle and their mobility in the reaction system. There are a wide variety of contaminants that are less widely recognized and studied regarding their effects on the FCC process. Fe and Ca probably have little inherent catalytic activity, but have been associated with surface deposits that clearly could have a significant impact on the access to the interior of a FCC catalyst particle [3].

313

2. FCC CATALYST MORPHOLOGY 2.1 Fresh Catalyst

The design and manufacture of FCC catalyst is complex and there are many proprietary techniques involved [1] (see Part l on page 299 of this volume). In general, however, the manu-facturing routes fall into four categories distinguished primarily by the technique used to bind the catalyst ingredients (zeolite, alumina, kaolin, etc.) into a microsphere suitable for use in an FCC: 9

9 9 9 9

Silica sol Modified silica sol (containing alumina) Alumina sol Alumina gel, and Clay based microspheres

Figure 1. SEM of fresh FCC catalyst with a modified silica sol binder.

314

There are numerous variations on these routes, but virtually any commercially produced FCC catalyst would fall into one of these categories. The first three routes are similar in the sense that the ingredients are mixed with the binder and spray dried into microspheres. In many cases, the microspheres undergo further thermal or chemical treatment. In the last case, microspheres are manufactured from treated clay and given a severe thermal treatment. Zeolite is subsequently grown inside the pores of the catalyst (in situ). The catalyst is finished with further chemical and thermal treatment. Although the active ingredients of all of these catalyst types are essentially the same, the different manufacturing routes create microspheres with remarkably different physical characteristics as evidenced by scanning electron micrographs (SEM) shown in Figures 1-4.

Figure 2. SEM of catalyst made with an alumina sol binder.

315

Figure 3. SEM of catalyst made with an alumina gel binder.

For the silica sol (Figure 1) and alumina sol (Figure 2) routes, there is definite evidence of a "skin" on the surface of the catalyst. For the alumina gel sample (Figure 3), there is no skin. The clay-based "in situ" technology (Figure 4) shows some minor indications of a skin. Although the porous character of the skins is not known, they appear to be denser than the interior and could offer mass transfer resistance. In the case of the silica sol binder, the skin appears to be enriched in silica. The presence of a hard skin can offer advantages in terms of catalyst hardness and hence attrition resistance [1].

316

0

0 0 c

ffl

0 e--

E

0 0

L_

O~

U.I

~

317

2.2 Contaminated Catalyst Over the past few years, an increasing awareness of catalyst morphology has developed in response to performance problems [3]. In many cases, a skin has been observed on the surface of the catalyst with a glazed appearance. Some typical examples follow. Figures 5 and 6 show the surface of a catalyst particle that has experienced only mild contamination. Notice in the high magnification case (Figure 5) that the surface appears rough and there is evidence of access to the particle interior.

Figure 5. SEM of"normal" E-cat. Figures 7-11 illustrate catalyst particles with significant contamination. The morphology of the surfaces varies greatly, but all of the catalysts exhibit a glazed appearance. The surface can be smooth (Figures 7 and 8), slightly dimpled (Figure 9), severely dimpled with nodules (Figure 10), or severely studded with nodules (Figure 11). This glazed surface appearance is associated with a skin or shell that has formed on the surface of the catalyst particle. It is typically a few microns thick and can vary in composition. The most common element present in the skin, besides the typical FCC

318

Figure 6. High magnification SEM of"normal" E-cat.

Si, AI and Re, is Fe. However, fluxing agents like Mg, Na, Ca, etc are also found in varying amount and type. The smooth, glazed appearance with no apparent cracks or pores suggests a low permeability.

3. MASS TRANSFER CHARACTERIZATION 3.1 Accessibility Index An analytical technique has been developed at Akzo Nobel that measures the rate of penetration of appropriately sized probe molecules into FCC catalyst particles. An index quantifying this rate can be defined. This AAI Index does not involve reaction of the probe molecule but simply measures resistance to mass transfer.

319

Figure 7. Contaminated E-cat with smooth glazed surface.

One of the most interesting outcomes of the application of the AAI relates to artificial catalyst deactivation. While increasing deactivation and contamination of a catalyst in a commercial unit decreases the AAI, the same artificial deactivation through steaming increases the AAI. Deactivation using cyclic protocols reduces the AAI, but not to the extent observed in commercial samples. None of these traditional deactivation techniques exhibits the glazed surface observed on commercial samples.

3.2 TiPB Technique Another tool that has proved useful in probing catalyst structural effects is the cracking of 1,3,5-triisopropylbenzene (TiPB). This molecule seems to be of the appropriate size to distinguish the effect of phenomena such as zeolite crystal size from the intrinsic cracking rate at the actual acid site.

320

b4 t~ e,, O O

e,,.

E

t~

e,-

l=

o ,..,.

e-O O

r

ili CO co ~

t~

E e-. "1"

aSai

321

Figure 9. SEM of contaminated catalyst exhibiting mild dimpling.

4. EFFECT OF MASS TRANSFER ON PERFORMANCE

There have been numerous occasions when accessibility has been thought to deleteriously affect catalyst performance. Although some of these cases have been documented elsewhere [3], we would like to focus on a number of recent examples which support the relationship of catalyst performance and accessibility:

322

r

i

t,-

t.__

i

E

i

e.-

"10 0 t,-.

.

.

.

t--X

. , m

>,,

l=

i

t---

t~ 0 "0

.

e-o o 0 LLI 69

.

i

0

d,--: v--

~1::: tT.~

0

t.___

"0 0 e-.

e--

..i

r X

0 i.!.! O9

t___.

II

.,....

323

324

Example 1. This data comes from an RFCC in which a commonly used commercial FCC process model is used in an attempt to predict performance of the unit. The process model is tuned to the unit and then used to predict future performance taking into account feed, catalyst, and operating variables. One of the variables the model does not use is the Accessibility Index. Figure 12 illustrates the difference between what the model predicted and the actual unit conversion. The prediction was not particularly good with deviations as high as 5 vol% in both the positive and negative directions. If we superimpose the AAI on the plot, we see that it tracks exceptionally well with the difference indicating that inclusion of this parameter in the model would improve its prediction significantly. It also demonstrates the significant effect that accessibility (mass transfer) can have on unit performance. Figure 13 illustrates the same phenomenon for gasoline yield. The same correlation turns out to be valid for bottoms make and LPG.

Figure 12. AAI improves RFCC conversion prediction.

325

Figure 13. AAI improves RFCC gasoline prediction.

Figure 14. MAT, At, I, and Fe correlate well after Fe spike

326

Example 2. In this case, we have another FCC cracking resid. The unit was subjected to a spike of contaminant Fe from a source other than the crude oil. The equilibrium catalyst was thoroughly analyzed. Contaminant metals other than Fe were essentially constant during this period. MAT and AAI are plotted against Fe in Figure 14. The correlation of both parameters with Fe is good and suggests a good correlation with MAT and AAI.

Example 3. In another resid case, Fe contamination of the feed occurred with no increase in Ni or V. The source of Fe is unknown. AAI dropped precipitously resulting in MAT and unit conversion losses given in the following table:

Samples Na, wt% Fe, wt% V, ppm Ni, ppm ABD, g/cc Activity AAI SA, m2/g RxT, C RgT, C Avg cat age Feed Effect Fe Unit conv FST FST AAI

Base Sample 1 0.48 0.47 0.46 0.61 936 932 2400 2352 0.83 0.82 69 66 6.3 1.9 125 115 500 700 (partial burn) 44 days African, < 25% resid

Sample 2 0.46 0.64 969 2334 0.84 64 1.9 115

-4% -22/wt% Fe -16/wt% (Fe + Na) -17/wt% Fe

These conversion losses are typical of other cases of AAI loss and Fe contamination. It is interesting to note that although Fe, Na, V, Ca, and Mg have been identified on the surface of highly contaminated catalyst, correlation with catalyst performance has been primarily with Fe concentration.

327

5. MODELING MASS TRANSFER EFFECTS There has been some debate in the literature as to the role mass transfer plays in the FCC process. Young and Rajagopalan [4] conclude that the cracking process is free of both vapor and liquid phase diffusion limitations based on experiments with varying catalyst particle size. This is contrary to their theoretical predictions for liquid phase diffusion. They explain this discrepancy by the inadequacy of the diffusivity prediction for heavier molecules. It is interesting to note that their experiments were performed on steam-deactivated catalysts that have shown anomalous accessibility behavior and no indication of performance penalty as a function of accessibility. O'Connor and Van Houtert [5] analyzed the non-steady state diffusion problem showing that for large molecules only partial penetration of the catalyst particle was possible for typical riser residence times. Their work was focused on the conversion of higher molecular weight petroleum components. The observation of skins on the surface of catalyst particles and associated performance problems as reviewed in the previous section suggests that a skin could be responsible for mass transfer limitations. Certainly, the glazed appearance of these skins would indicate a lower permeability. A mathematical model has been developed to help understand what the implications of a low permeability skin would be on catalyst utilization. The catalyst volume was divided into two zones, a thin surface shell and the remaining interior of the particle. A transient solution was obtained for the following equation ac _ 1 a ac) at r 2 ar (B(r) r2--~ subject to the conditions: 1. there is zero flux at r = 0 (the center of the sphere) 2. the concentration at the outer surface of the sphere (r = R) is at all times t _>0 equal to the concentration in the bulk phase times the ratio between the saturation concentration inside and outside the sphere. 3. the initial (t = 0) concentration within the spheres is identically zero. 4. The concentration and flux are equal at the interface between the shell and interior. Reaction rate, per se, was not considered, but was instead represented by a Freundlich adsorption isotherm: 1

C,d, = re'C"

Two separate isotherm cases were examined:

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1) Linear Isotherm Here, n = 1, and a continuous reaction with no time dependency is represented. There is no deactivation and a continuous coke gradient results. 2) Shrinking Core Here, where n >> 1 (50), a very fast reaction rate relative to diffusion is represented. Deactivation is rapid and a "black and white" coke gradient is observed. Practice will lie somewhere between these two models. Assumptions in this model include: 9 9 9 9 9

No interparticle mass transfer Only adsorption No reaction kinetics No desorption of reaction products No influence of flux of reaction products on intraparticle mass transfer

The general conclusions are represented in Figures 15 and 16 for the two isotherm cases. The utilization of the particle, quantified by the effectiveness factor Cp(t) Eff ectivenness

F a c t o r = c p (equilibrium)

is plotted against the Fourier Number Ff--w

D.t dp2

in which D t dp

= = =

Diffusion coefficient time Particle diameter

The skin was assumed to occupy 2% of the particle radius. Fourier numbers were calculated for various carbon number molecules by estimating the diffusivity and assuming a particle size of 60~ and a time of 3 sec. Effectiveness factors were calculated for a number of cases representing lower relative diffusivity in the shell versus the bulk. Comparing Figures 15 and 16 we see that there is not much effect of the isotherm assumption on the results. As expected, as contact time is lowered, the effect of a

329

less permeable shell is magnified. Obviously, the lower the permeability of the shell, the more dramatic the effect. For a molecule with 40 carbon number, modest reductions of 10 to 20% in catalyst utilization are calculated. For an 80 carbon number, the penalty calculated for the lowest shell permeability is around 40%. It should be realized that the diffusivities are simply estimates and the calculations only serve to illustrate the plausibility of the penalty of a low permeability on catalyst performance. However, the relative carbon number calculations suggest that the penalty will be most severe on bottoms cracking.

6. MASS TRANSFER AND HIGH ACCESSIBILITY CATALYSTS

If, in fact, a low permeability shell is responsible for loss in performance in highly contaminated FCC catalysts, two remedies are possible: A catalyst architecture that renders the shell more porous, or A catalyst architecture that significantly increases the bulk mass transfer capabilities.

Figure 15. Catalyst Particle Utilization Assuming Linear Isotherm

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Figure 16. Catalyst Particle Utilization Assuming Shrinking Core Model

Figure 17. Higher Accessibility Catalyst Retains Advantage as an Fe is Added

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It is somewhat difficult to conceive of how to achieve the first remedy. Possibly the surface composition could be engineered to resist glazing, but there is no evidence to date of this being the case. Certainly, if the permeability of the shell were low enough, increasing the mass transfer in the interior of the particle would never help. However, if the mass transfer rates in the two regions were not too disparate, the second remedy could be effective. This seems to be the probable case based on experimental data to date. This is illustrated in Figure 17 in which AAI data for catalysts that have been contaminated with Fe are plotted. It appears that catalysts with high accessibility structures maintain their advantage over less accessible structures contaminated with incremental Fe that either forms a shell or further defines an already present shell. From a modeling point of view, this can be visualized by starting with the low skin permeability curves of Figures 14 and 15 and observing what happens if the diffusivity ratio is increased. In the following table, some cases in which high accessibility catalysts were utilized to improve performance in high contaminant environments are summarized:

Case A1 A2 B1 B2

Fe (wt%) 0.8 0.8 ....

Delta Conversion LA to HA (wt%) 5 4 2.1 2.0

Delta Bottoms LA to HA (wt%)

-3.4 -2.5

In all of these cases, neither the zeolite character nor content was changed appreciably. Improved performance can be ascribed almost totally to improved accessibility.

7. CONCLUSIONS

It has been shown that shells exist in both fresh and contaminated FCC catalysts that have the potential to negatively affect performance. Based on microscopic examination, these shells appear to be quite dense with the probability of a significantly lower permeability. A mathematical model representing diffusion into a catalyst particle with such a shell indicates the plausibility of this argument. Depending on parameters assumed, catalyst utilization penalties of up to 40% were predicted for large petroleum molecules.

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The solution to resisting this problem commercially appears to be utilization of a catalyst with a high initial accessibility. These catalysts appear to maintain their accessibility advantage over less accessibility catalysts as a contaminant shell is formed or better defined. 8. REFERENCES

1. P. O'Connor, P. Imhof and S.J. Yanik, Catalyst assembly technology in FCC. Part I: A review of the concept, history and developments. This volume, p. 299. 2. M.O. Coppens, Catalysis Today 53 (1999) 225. 3. M.C.J. Hodgson, C.K. Looi and S.J. Yanik, Akzo Nobel Catalysts Symposium 1998, Paper F-4. 4. G.W. Young and K. Rajagopalan, Ind. Eng. Chem. Process Dev., 24 (1985) 995. 5. P. O'Connor and F.W. van Houtert, Ketjen Catalyst Symposium 1986, Paper F-8.

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Keyword Index

Accessibility, 71, 279, 299, 311 Accessibility measurement, 209 Acidity, 41, 263 Activity, 219, 263 Alumina, 41 Basic metal oxide, 3 Bifunctionality, 263 Butylenes, 111 13CNMR, 251 Catalyst, 71, 263 Catalyst activity, 209 Catalyst aging, 87 Catalyst assembly, 299 Catalyst deactivation, 187, 201, 219, 227 Catalyst preparation, 293, 299 Catalyst structure, 311 Catalytic coke, 239 Coke, 187, 239, 251 Coke analysis, 251 Coke formation, 167 Contact time, 153 Crystal size, 279 Deactivation, 87, 133, 311 Dealumination, 3 Entrainment coke, 239 Equilibrium catalysts, 201 Extraction of coke, 239 FCC contaminants, 133 Feedstock, 71 Fluid bed, 111 Fluid cracking catalyst preparation, 59 Gasoline composition, 153

Hard coke, 239, 251 High rare-earth on Y, 141 High silica alumina ratio, 107 History, 299 Hydrogen transfer, 133, 141, 153 Hydrothermal treatment, 59, 87, 187, 219 Industrial units, 87 Kinetic models, 187 Kinetics, 153, 167, 209, 251 Lab scale pilot plant, 167 Labelling, 251 Light olefins optimization, 111 Low-alumina Y (LAY) zeolites, 107 Mass transfer and diffusion, 311 MAT testing, 209 Mathematical model, 87 Mathematics, 187 Matrix, 263 Matrix properties, 71, 227 Mechanisms, 153 Metal oxides, 107 Metals, 201, 219, 227, 263, 263, 311 Micro-riser, 167 Modelling, 311 Models, 219 New test unit, 153 Nickel resistance, 201 Non-framework alumina, 3, 41, 59 Olefins cracking, 141 Olefins in gasoline, 141 Olefins optimization, 133 Organized mesoporous silica, 293

334 Particle binders, 293 Particle formation, 299 Pilot plant, 111 Pilot riser testing, 227 Pore size distribution, 299 Probe molecules, 279 Propylene, 111 Reactor, 111 Resid cracking, 227 Riser testing, 71 Short contact time, 71, 111, 167, 311 Silica, 293 Silica-alumina ratio, 3 Silica structures, 293 Sintering, 311 Soft coke, 239, 251 Solid state aluminum 27 NMR, 41, 59 Solid state silicon 29 NMR, 41, 59 Strippability, 209, 239 Structural design, 299

Sulfur in gasoline, 153 Sulfur species, 153 Surface area, 219, 279 TCC, 107 Total olefins management, 141 Tri-iso-propyl-benzene, 279 Ultrastable zeolites, 3 USY, 3 Vanadium, 133 Vanadium tolerance, 201 Zeolite aging, 59 Zeolite deactivation, 59 Zeolite properties, 227 Zeolite synthesis, 41 Zeolites, 3,263, 279 Zeolites, HY, 41, 59 ZSM-5, 141