studies in Surface Science and Catalysis 106 HYDROTREATMENT AND HYDROCRACKING OF OIL FRACTIONS
Technologisch Instituut
Studies in Surface Science and Catalysis Advisory Editors: B. Delmon and J.T. Yates Vol.106
HYDROTREATMENT AND HYDROCRACKING OF OIL FRACTIONS
Proceedings of the 1st International Symposium/6th European Workshop, Oostende, Belgium, February 17-19,1997 Editors G.R Froment Universiteit Gent, Gent, Belgium B. Delmon Universite Catholique de Louvain, Louvain-La-Neuve, Belgium P. Grange Universite Catholique de Louvain, Louvain-La-Neuve, Belgium
1997 ELSEVIER Amsterdam — Lausanne — New York — Oxford — Shannon — Tokyo
ELSEVIER SCIENCE B.V. Sara Burgerhartstraat25 P.O. Box 2 1 U 0 0 0 AE Amsterdam, The Netherlands
ISBN 0-444-82556-8 © 1997 Elsevier Science B.V. All rights reserved. No part of this publication may be reproduced, stored in a retrieval system or transmitted in any form or by any means, electronic, mechanical, photocopying, recording or otherwise, without the prior written permission of the publisher, Elsevier Science B.V., Copyright & Permissions Department, RO. Box 521,1000 AM Amsterdam, The Netherlands. Special regulations for readers in the U.S.A. - This publication has been registered with the Copyright Clearance Center Inc. (CCC), 222 Rosewood Drive, Danvers, MA 01923. Information can be obtained from the CCC about conditions under which photocopies of parts of this publication may be made in the U.S.A. All other copyright questions, including photocopying outside of the U.S.A., should be referred to the copyright owner, Elsevier Science B.V, unless otherwise specified. No responsibility is assumed by the publisher for any injury and/or damage to persons or property as a matter of products liability, negligence or otherwise, or from any use or operation of any methods, products, instructions or ideas contained in the material herein. This book is printed on acid-free paper. Printed in The Netherlands
COHTBNTS Introduction
xi
SCS7 WO^TM tSCTVWBB Processes and catalysts for hydrocracking of heavy oil and residues F. Morel, S. Kressmann, V. Harle and S. Kasztelan
1
An improved process for the production of environmentally friendly diesel fuels J. Grootjans and C. Olivier
17
Hydroprocessing to produce reformulated gasolines - the ISAL*"" process G.J. Antes, B. Solari and R. Monque
27
Molecules, catalysts and reactors in the hydro-processing of oil fractions W.H.J. Stork
41
Simultaneous HDN/HDS of model compounds over Ni-Mo sulfide catalysts L. Zhang and U.S. Ozkan
69
Kinetics of the catalytic removal of the sulphur components from the light cycle oil of a catalytic cracking unit G.F. Froment, G.A. Depauwand V. Vanrysselbergtie
83
A review of catalytic hydrotreating processes for the upgrading of liquids produced by flash pyrolysis R. Maggi and B. Delmon
99
Dual-functional Ni-Mo sulfide catalysts on zeolite-alumina supports for hydrotreating and hydrocracking of heavy oils H. Shimada, S. Yoshitomi, T. Sato, N. Matsubayashi, M. Imamura, Y. Yoshimura and A. Nishijima
115
VI
ORAL COMMDI^CAlllOHS CATALYTIC ASPECTS Hydrodracking of Cio hydrocarbons over a sulfided NiMo/Y zeolite catalyst J.L Lemberton, A. Baudon, M. Guisnet, N. Marchal and S. Mignard
129
Novel hydrotreating catalysts based on synthetic clay minerals B. Leiiveld, W.C.A. Huyben, A J. van Dillen, J.W. Geus and D.C. Koningsberger
137
Influence of the location of the metal sulfide in NiMo/HY hydrocracking catalysts D. Cornet, M. El Qotbi and J. Leglise
147
Acidity induced by H2S adsorption on unpromoted and promoted sulfided catalysts C. Petit, F. Mauge andJ.-C. Lavalley
157
Organo metallic siloxanes as an active component of hydrotreatment catalysts I.M. Kolesnikov, A.V. Yablonsky, M.M. Sugungun, S.I. Kolesnikov and M. Y. Kilyanov
167
Alumina supported HDS catalysts prepared by impregnation with new heteropoli-compounds A. Griboval, P. Blanchard, E. Payen, M. Fournierand J.L. Dubois
181
Genesis, characterizations and HDS activity of Mo-P-alumina based hydrotreating catalysts prepared by a sol-gel method R. Iwamoto and J. Grimblot
195
Effects of ethylenediamine on the preparation of HDS catalysts : comparison between Ni-Mo and Co-Mo based solids P. Blanchard, E. Payen, J. Grimblot, O. Poulet and R. Loutaty
211
Creation of acidic sites by hydrogen spillover In model hydrocracking systems A.M. Stumbo, P. Grange and B. Delmon
225
Application of ASA supported noble metal catalysts in the deep hydrodesulphurisation of diesel fuel H.R. Reinhoudt, R. Troost, S. van Schalkwijk, A.D. van Langeveld, S.T. Sie, H. Schuiz, D. Chadwick, J. Cambra, V.H.J, de Beer, J.A.R. van Veen, J.L.G. Fierro and J. A. Moulijn
237
Reactor runaway in pyrolysis gasoline hydrogenation E. Goossens, R. Donker and F. van den Brink
245
vu Surface property of alumina-supported Mo carbide and its activity for HDN T. Miyao, K. Oshikawa, S. Omi and M. Nagai
255
The design of base metal catalysts for hydrotreating reactions; temperature programmed sulphidation of NiW/AlaOa catalysts and their activity in the hydrodesulphurisatlon of thiophene and dibenzothiophene H.R. Reinhoudt, A.D. van Langeveld, R. Mariscal, V.H.J, de Beer, J.A.R. van Veen, S.T. Sie and J.A. Moulijn
263
THEORY AND CATALYTIC DEACTIVATION Surface science models of CoMoS hydrodesulfurization catalysts A.M. de Jong, V.H.J, de Beer, J.A.R. van Veen andJ.W. Niemantsverdriet
273
Molecular mechanics modelling of the interactions between M0S2 layers and alumina or silica support Ph. Faye, E. Payen and D. Bougeard
281
In-situ FT-IR study of NO adsorbed on C0-M0/AI2O3 sulfided at high pressure (<5.1 MPa) N. Koizumi, M. lijima, T. Mochizul
293
Compound formation and hydrogen activity at sulfided catalysts : a combined surface science and quantum chemical approach J. Paul, H. Al<pati, P. Nordlander, W.S. Oh, D.W. Goodman and B. Demirel
303
Deactivation studies on NiO-MoO3/Al203 and C0O-M0O3/AI2O3 hydrodesulphurization catalysts R. Marinl
307
Characterization of aged catalyst from hydrotreating petroleum residue M.T.Martinez, J.M. Jimenez, M.A. Callejas, F.J. Gomez, C. Rial and E. Carbd
311
PROCESSES Hydrotreatment of spent lube oil: catalysts and reactor performance 0. Yiol<ari, S. Morphi, E. Siokou, F. Satra, S. Bebelis, 0. Vayenas, C. Karavassilis and G. Deligiorgis
323
Catalytic hydrodesulfurization of petroleum middle distillate and model sulfur compounds over a series of catalysts activity and scheme E. Lecrenay and I. Mochida
333
Vlll
Hydrotreating of compounds and mixtures of compounds having mercapto and hydroxyl groups T.'R. Viljava and A.O.I. Krause
343
Influence of high Mo loading on the HYD/HDS selectivity of alumina supported M0S2 catalysts P. Da Silva, N. Marchal and S. Kasztelan
353
Low temperature hydrocracking of paraffinic hydrocarbons over hybrid catalysts /. Nakamura, K. Sunada and K. Fujimoto
361
Tail-selective hydrocracking of heavy gas oil in diesel production M.V. Landau, LO. Kogan and M. Herskowitz
371
KINETICS Influence of the hydrocarbon chain length on the kinetics of the hydroisomerization and hydrocracking of n-paraffins B. Debrabandere and G.F.Froment
379
Aromatics hydrogenation over supported platinum catalysts : the influence of sulfur on the kinetics of toluene hydrogenation over Pt/Y-zeolite catalysts H. Bergem, E.A. Blekkan and A. Holnfien
391
Kinetic study of the hydrodenitrogenation of pyridine and piperidine on a NiMo catalyst R.Pille and G.F. Froment
403
PO&TBRB KINETICS AND PROCESSES Kinetic modelling of HDN reaction over (Ni)Mo(P)/Al203 catalysts M. Jian and R. Prins
415
A kinetic model for hydrodesulfurisation M. Sau, C.S.L. Narasimhan andR.P.Verma
421
Hydrotreating of gas-oils : a comparison of trickle-bed and upflow fixed bed lab scale reactors R. Myrstad, J. Steinsland Rosvoll, K. Grande and E.A.BIekkan
437
Trickle-bed reactor modeling for middle-distillates hydrotreatment C.G.Dassori, N. Fernandez, R. Arteca, A. Diaz, and S. Buitrago
443
IX
Petroleum residua hydrotreating on Co and/or Ni containing catalysts v. Kogan and N.M. Parfenova
449
Saturation of aromatics in diesel fuels : the catalytic toxicities of sulfur and nitrogen compounds P. Kokayeff and GJAntos
463
Production of high octane gasoline components by hydroprocessing of coalderived aromatic hydrocarbons B. Demirel and W.H. Wiser
469
Process developments in gasoil hydrotreating R.C. Lawrence, D.H. McKinley and M.A. Wood
479
CATALYTIC ASPECTS Benzene hydrogenation over transition metal carbides C. Marguez-Alvarez, J.B. Claridge, A. P. York, J. Sloan and M.LH. Green
485
Hydrodesulfurization of dibenzothiophene in a micro trickle bed reactor D. Letourneur, M. Vrinat and R. Bacaud
491
Use of dispersed catalysts for fossil fuel upgrading A.S. Hirschon and R.B. Wilson
499
Hydrodesulphurization of gas oil using C0-M0/AI2O3 catalyst A.S. Nasution and E. Jasjfi
505
The application of cobalt containing acidic zeolites as catalysts for hydrodesulfurization reactions T.I. Koranyi, N.H. Pham, A. Jentys and H. Vinek
509
Hydrodearomatizatlon of naphtenic base machine oils G. Kons, H.-J. MCiller, M. Vicari, E. Schwab and M. Walter
519
Highly dispersed metal sulfide catalysts for the hydroconversion of vacuum destination residues K. Buker, H. Berndt, B. Lucke and W. Kotowski
523
Hydrogenation of tetralin over a sulfided ruthenium on Y zeolite catalyst: comparison with a sulfided NIMo on alumina catalyst J.L Lemberton, M. Cattenot, V. Kougionas, M. Mhaouer, J.L Portefaix, M. Breysse and G.Perot
529
Catalysts of phosphotungstic or phosphomolybdic acids on different supports from dimethylformamide solutions L.R. Pizzio, P.G. Vazquez, M.G. Gonzalez, M.N. Blanco, C.V. CSceres and H.J. Thomas
535
Infrared study on the acid sites on nitrided molybdena-alumina catalysts M. Nagai, O. Uchino, T. Kusagaya and S. 0ml
541
Synthesis and characterization of zirconia-alumina mixed oxides F.Dumeignil, P. Blanchard, E. Payen, J. Grimblot and O. Poulet
547
Selective synthesis of methylcyclopentane from cyclohexane using Pt-zeolite hybrid catalyst /. Nakamura, A. Zhang and K. Fujimoto
561
Hydrocracking activity of NiMo - USY zeolite hydrotreating catalysts B. Egia, J.F. Cambra, B. Guemez, P.L Arias, B. Pawelec and J.L.G. Fierro
567
Authors index
573
General Information
577
XI
y^l^*y^lyfji l^yYj^j I' *^yf U j ^
The new American, European and Japanese environmental regulations call for advanced hydrotreatment processes for HDS and HDN, the removal of S- and Nicomponents from oil fractions. They will also alter the product slate of the oil refineries and the hydrocarbon composition of these products. Hydrocracking will play an important part in this shift. Adapting the operating conditions will not suffice to reach the desired product specifications and yields. Adequate catalysts will have to be developed. Powerful tools are now available for this, among others surface science techniques, molecular modeling, new types of reactors operated in a non steady mode. Another instrument in the improvement of hydrotreatment and hydrocracking units is the availability of more realistic kinetic models. These are based upon a judicious insight into the reaction mechanism, also provided by the above mentioned tools. Progress in the analytical techniques has allowed to reduce the lumping of components In these kinetic models and first order kinetic equations are gradually replaced by equations accounting for the adsorption of the various components. More detailed and more realistic reactor models are now based upon rigorous hydrodynamic models and their application has become possible through the rapidly Increasing possibilities of computers. A global perspective and an inspection of the state of the art are timely. This is exactly what this symposium on Hydrotreatment and Hydrocracking of Oil Fractions is aiming at. It is a worldwide extension of the 5 European Workshops held from 1981 onwards. We have chosen a prestigious location for the Symposium, which will also favor the personal contacts between the participants. We are grateful to Ms. R. Peys and her staff at Tl-K VIV for the organization of the meeting and for the preparation of these Proceedings.
Prof.Dr. B. Delmon, U.C.L. Prof.Dr.ir. G.F. Froment, R.U.G. Prof.Dr. P. Grange, U.C.L.
The 1st International Symposium on Hydrotreatment and Hydrocracking of Oil Fractions was organized by : The Technological Institute of the Royal Flemish Society of Engineers (TlK VIV). The K VIV is the professional organization of the academically trained Flemish engineers. It represents more than 11.000 members. In 1940 the Society founded the Technological Institute, with the aim of disseminating information on scientific and technological development by means of seminars, lectures, courses, congresses and conferences. Address :
Technological Institute vzw Desgulnlei 214, B - 2018 Antwerpen 8 +32 3 216 09 96 - e +32 3 216 06 89 - e-mail:
[email protected]
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® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
Processes and Catalysts for Hydrocracking of Heavy Oil and Residues F. Morel^, S. Kressmann^, V. Harle^, S. Kasztelan^ Institut Frangais du Petrole, 1 Centre d'Etudes et de Developpement Industriel "Rene Navarre", B.P. 3, 69390 Vernaison, France. 2 Division Cinetique et Catalyse, B.P. 311, 92506 Rueil-Malmaison, France. Atmospheric or vacuum residue can be converted into valuable distillates using high reaction temperature, high hydrogen pressure and low contact time hydroprocessing units. Various residue hydrocracking processes are now commercially employed using fixed bed, moving bed or ebullated bed reactors. The choice of process type depends mainly on the amount of metals and asphaltens in the feed and on the level of conversion required. Various improvements have been introduced in the last decade to increase run length, conversion level, products qualities and stability of the residual fuel. These improvements include on stream catalyst replacement systems, swing reactors, improved feed distribution, guard bed materials limiting pressure drop, coke resistant catalysts, complex association of catalysts using particle size, activity and pore size grading. Further improvement of the resistance of catalysts to deactivation by coke and metal deposits and of the hydrodenitrogenation activity are two major challenges for the development of new residue hydrocracking catalysts and processes. 1. INTRODUCTION Hydrodesulphurisation of atmospheric residue is a well established residuum upgrading process which requires catalysts designed to remove and accumulate metals and to desulphurise the feed [1-3]. However, by increasing the reaction temperature and using appropriate catalysts, the conversion of a substantial fraction of the residue into distillates can also be obtained. The objective of residue u p g r a d i n g is therefore more and more switching from hydrodesulphurisation (HDS) to conversion or hydrocracking (HDC) [1-3]. The HDC of atmospheric or vacuum residue lead to the production of valuable distillates and to the minimisation of the amount of unconverted residual fuel oil which is deeply purified and suitable for low sulphur fuel oil (LSFO).
Residue HDC processes are therefore able to satisfy a number of objectives driven by market tendencies and by more stringent environmental regulations such as a decreasing market for fuel oil, an increasing need for LSFO, an increasing need for clean transportation fuels (gasoline and diesel) and tighter requirements on the use or disposal of the worst refinery residues causing a trend toward "minimum residue" or "zero residue" refinery strategies. Although very capital expensive and in competition with thermal conversion processes, residue HDC processes will develop in the future as they gain in maturity and as more and more synergy will be found with other conversion processes such as distillate hydrocracking, fluid catalytic cracking, coking and residue gasification for hydrogen, steam and power generation. Today, commercially proven residue catalytic HDC processes use fixed bed, cocurrent and counter-current moving bed or ebullated bed reactors [4]. Several slurry reactor processes are also proposed and several demonstration plants have been operated, nevertheless, full scale units have not yet been built [5]. In parallel, many catalysts have been also developed over the years and in particular catalysts packages have been designed for fixed bed units specially adapted to minimise pressure drop and maximise metal retention capacity, HDS activity, run length and conversion level as well as unconverted residue stability. In this paper we briefly review the recent major developments in residue HDC processes and catalysts. 2. RESIDUE HYDROCRACKING PROCESSES 2.1. Hydrocracking operating conditions Atmospheric residue (AR) and vacuum residue (VR) are the most difficult feed to convert catalytically because they contain in a concentrated form most of the impurities contained in the crude oil and among them, asphaltens and metals (essentially nickel and vanadium). Asphaltens are complex aggregates of large molecules containing heteroatoms (S, N, O and metals) which tend to inhibit all of the catalytic functions by coking and metal sulphides deposition during the reaction. In residue HDC one of the challenge is to convert as much asphaltens as possible into lighter compounds. Metals are removed from their host molecules by catalytic hydrodemetallisation (HDM). It is known that metals deactivate the active phase and accumulate on the catalyst surface in the form of sulphide particles during the reaction. Eventually, a large accumulation of metal deposits lead to pore plugging [1,2,4,6-11]. Both catalytic and thermal cracking occur in residue HDC processes and the proportion of each type of reactions depends mainly on the reaction temperature. Catalytic hydrocracking is favoured by temperature and by hydrogen pressure. Thermal cracking involves free radical mechanism to convert large residue molecules to hydrocarbon gases, naphtha, distillates and gas oil. Thermal
conversion is highly favoured by an increase of temperature and is nearly independent from hydrogen pressure. However, a high hydrogen pressure is useful for the hydrogenation of the highly reactive unsaturated compounds generated by thermal cracking preventing polymerisation reactions leading to coke formation. An increase of the hydrogen pressure also have a tremendous impact on the quality of the products by providing high level of desulphurisation and hydrogenation. Hence residue hydrocracking processes are high hydrogen pressure, high temperature and low contact time processes [2,12]. There are two major disadvantages to run at high temperature in residue hydrocracking. The first one is an increased rate of deactivation by coking of the catalysts reducing the cycle length and the second is the formation of highly unsaturated and insoluble compounds which tends to form dry sludge or sediments leading to a poor stability of the residual fuel. These sediments may simply be asphaltens or fragments of asphaltens precipitating as a result of the decrease of the ratio of resins to asphaltens obtained by hydrogenation and conversion [13]. A large number of technology have been developed to overcome the troubles generated by the use of high reaction temperature in the processing of heavy feeds containing metals contaminants like oil residua. Figure 1 indicates schematically the various combination of fluid and catalyst flows that can be found in commercial reactors. Various types of reactors are employed such as fixed bed, co-current and counter-current moving bed, ebullated bed and slurry reactors. Fixed Bed (Trickled)
Moving Bed
HC*M2
MC^Mj
I
A
1*1
MC*M2
Moving Bed Ebullated Bed
(Contercurrcnt)
CATA.
I
CATA.
Figure 1. Residue hydroconversion reactors.
HC • H2 • CATA n^^nj^i^M.M
MC^Mj
CATA.
J*Uy
MC • Mj
Slurry Reactor
(Fluidiscd)
HC*H2
CATA.
itL/
HC*H2
(Cocurrent)
MC • H j
A
CATA.
jtL/
CATA.
A
Jti
HC • H2 • CATA
Table 1 summarises some of the main features of residue HDC processes using these various types of reactors. These processes are in general composed of several large reactors in serie and sometimes several parallel trains of several reactors in serie. For example, the Mina AbduUa unit in Kuwait has two trains of four fixed bed reactors each totalling about 1600 m^ of catalyst [5]. The Amoco Texas City unit possesses three trains of three ebullated bed reactors in series each [2]. Table 1. Main Features of Residue Hydrocracking Processes. Fixed Swing Moving bed fixed bed bed 40 1 3 Number of units (est'd) 500 500-700 120 Ni+V max. in feed (ppm) 100-200 100-200 100-200 Pressure (MPa) 380-420 380-420 380-420 Temperature (°C) 0.1-0.5 0.1-0.5 0.1-0.5 Unit LHSV (h"!) medium good good Unit Operability 60-70 60-70 Max. Conv. to 550°C- (wt%) 50-70 good Residue Stability good good Unit cycle length (month) 6-12 12 CO+ 0.55-0.7 1 1 RCC^ --1.2x3 --1.2x3 Catalyst part, size (mm) --1.2x3 --120 --120 -120 Particles/cm^ [4] Vol. % Cat. in reactor [4] --60 --60 -60 + CO : Continuous operation * RCC : Relative catalyst consumption for same feed for one
EbuUating bed 9 >700 100-200 400-440 0.2-1 medium 70-80 medium CO+ 1.4-2 -0.8x3 -250 -40
Slurry 3 demo >700 100-300 420-480 0.2-1 difficult 80-95 bad CO+ -0.002 -2.4 109 -1
year cycle
All five types of reactors shown in Figure 1 allow to reach the objective of converting a substantial fraction of the residue. Nevertheless as indicated in Table 1, a net advantage is obtained for ebullating bed and slurry processes which can process residues at higher reaction temperatures with no pressure drop. However these two processes produce lower quality products than fixed and moving bed processes as will be seen hereafter. 2.2. Fixed bed processes The most standard residue hydroconversion processes use fixed bed catalytic reactors. The main process licensers are CHEVRON, UOP/UNOCAL, EXXON, SHELL and IFP. Fixed bed hydroprocessing units can be operated in two modes, HDS and HDC by increasing the reaction temperature and using different catalysts [12]. Table 2 shows the products properties obtained from the two modes of operation on the same feed (Kuwait vacuum residue containing 5.51 wt% S) using the HYVAHL-F process from IFP.
In Table 2, the HDS mode allows to reach a conversion of the 550°C+ residue of 35 wt%. The converted product is mainly a desulphurised vacuum gas oil. For the HDC mode. Table 2 shows that the conversion of the 550°C+ residue reach 56 wt% and substantial amounts of gas oil and naphtha are obtained. In both mode of operation, the unconverted atmospheric residue is a desulphurised and stable high quality LSFO with 0.75 wt% or 0.8 wt% sulphur respectively. Results obtained on two other feeds and reported in Table 2 indicate that a conversion level of 63 wt % can be reached with a fixed bed process. The fixed bed process is well adapted for feedstock containing less than 100/120 wt ppm Ni+V for a one year cycle length. This cycle length is determined by the HDM activity and the saturation of the metal retention capacity of the HDM catalyst. The operability of the fixed bed process is good. Table 2. Hydroconversion of various residues in fixed bed HYVAHL-F Process HDC HDC HDC Mode HDS AL/AH AR Safaniya VR Residue Kuwait VR 0.988 1.035 1.031 Specific Gravity 1.031 3.95 5.28 5.51 Sulphur, wt % 5.51 2900 4600 3600 Nitrogen, wt ppm 3600 13.8 23.0 21.8 Conradson Carbon, wt% 21.8 5,7 11.5 9.0 Asphaltens C7, wt % 9.0 104 203 169 Ni+V, wt ppm 169 Yields /conversions MOR 2.2 3.5 Gasoline, wt % 4.6 1.0 19.7 21.5 Gas oil, wt % 19.5 9.5 40.8 32.8 Vacuum gas oil, wt % 29.5 20.5 37.1 29.7 43.2 Vacuum residue, wt % 64.6 91 89 HDS, wt % 90 88 HDM, wt % 93 94 98 98 57 63 35 56 Conv 550°C+, wt % Hydrotreated AR 0.970 0.934 0.965 0.963 Specific Gravity 0.50 0.70 0.8 0.75 Sulphur, wt % Conradson Carbon, wt% 10.0 4.0 10.3 8.8 0.3 0.7 Asphaltens C7, wt % 5.3 1.3 20 4 2.0 11 Ni+V, wt ppm For feedstock containing larger metal content (up to 250-300 wt ppm Ni+V), a new concept of fixed bed reactors including a swing guard reactors system has been developed by IFF (HYVAHL-S process. Figure 2). The process scheme includes swing guard fixed bed reactors which can be switched in operation and several downstream fixed bed reactors in serie. When the HDM catalyst contained in one of the guard reactor in operation is deactivated, a suitable procedure and technology allow to bypass this reactor, replace the catalyst
with a fresh one and put back the reactor on stream. All these operations are done without shutting down the unit, avoiding any production loss. The guard reactor volume and the operating temperature are optimised to achieve a high HDM rate while ensuring satisfactory cycle length in this guard reactor. The swing reactor system has all the benefit of fixed bed reactors, namely no catalyst attrition, no entrainement of fines and allows the unit to achieve a one year operation cycle in the main reactors for most available crude oils. 2.3. Moving bed processes In moving bed reactors the catalyst circulates by gravity, in plug flow, inside the reactor and fresh catalyst is added periodically at the top of the reactor while spent catalyst is withdrawn at its bottom. The main advantage of the moving bed process is its capacity to process, with long cycle length, high metal content feedstock. HDM
Demetallization, Conversion Swing Reactors
u
HDS
Desulfurization Refining
Hydrogen
To Fractionation Section
Figure 2. Fixed bed HYVAHL-S process with swing reactors. Several processes are proposed by various process hcensers such as the cocurrent down flow moving bed in the HYCON process (SHELL), the counter current moving bed OCR (On stream Catalyst Replacement) process (CHEVRON) and HYVAHL-M process (IFF). The best configuration is the counter current mode because the spent catalyst saturated by metals meet the fresh feed at the bottom of the reactor whereas the fresh catalyst reacts with an already demetalHsed feed at the top of the reactor. This configuration results in a lower catalyst consumption than with other processes.
When moving through the reactor, the catalyst is submitted to high mechanical forces that may lead to attrition of the catalyst and therefore to bed plugging problems on the top of the first downstream fixed bed reactor. The catalyst attrition in moving bed can be reduced through optimal hydrodynamic conditions [14] and use of high attrition resistance catalysts. Moving bed processes need however special equipment and procedures for a safe and effective catalyst transfer into and out of the high pressure and high temperature reactor. This involves several low and high pressure vessels as indicated in Figure 3 for the OCR process. The product yields and qualities are similar to fixed bed process for the same operating conditions (Table 3). However, the operability is more difficult due to the operation of continuous catalyst renewal in high P-T conditions and to the entrainment of catalyst fines to the downstream fixed bed reactor.
FrMh Catalytt Bin Vapor/Liquid
Hydrogen and Feed OB
Figure 3. Moving bed OCR reactor.
cj,|a|yrt , WHhdrawal
Recycle OU
Figure 4. Ebullated bed H-Oil reactor.
With these processes the conversion can reach 70 wt % depending on the nature of the feed and the reactivity of the asphaltens. To reach a higher conversion level is difficult because increasing the reaction temperature induce an extensive coking of the catalyst and catalyst bed plugging. 2.4. Ebullated bed process Two ebullated bed technology are currently licensed, namely the LC-Fining (ABB/Lummus Crest) and the H-Oil (IFP/HRI) processes. In an ebullated reactor, the fluids circulate up flow in the reactor. A recirculating pump expands the catalytic bed and maintains the catalyst in suspension. The expanded bed volume is 30 to 50 % larger than the bed volume
at rest as shown in Figure 4. Due to the expansion of the catalyst bed, the ebullated bed is perfectly mixed and isothermal and can work at a much higher temperature than fixed or moving bed reactors. Therefore a higher conversion of the feed can be reached. In addition, bed plugging due to coke build-up between the catalyst grains is prevented. However, the ebullated bed works at higher space velocity, i.e. smaller catalyst volume, than fixed or moving beds reactors with the same reactor volume. The ebullated bed reactor uses also a continuous catalyst renewal system at the top of the reactor and a continuous withdrawal system at its bottom and this allows a continuous operation of the unit. To reach a high level of conversion (> 50-60%) the main problem is to limit the sedimentation of heavy compounds and asphaltens [7] which lead to deposits on the reactor internal parts and down stream vessels as well as on the catalyst and cause operability problems and catalyst deactivation [5, 6]. This increases the catalyst consumption and decreases the stability of the residual fuel. The high temperature used in ebullated bed lead however to lower product qualities than with fixed and moving bed reactors (Table 3). Table 3. Examples of products properties from heavy Safaniya VR hydroconversion using various processes. Slurry reactor Fixed/Moving bed EbuUating bed Naphtha 10-15 5-15 1-5 Yield/Feed, wt % 0.720 0.710-0.720 0.710-0.740 Specific Gravity 0.06 0.01-0.2 <0.01 Sulphur, wt % 200 Nitrogen, wt ppm 50-100 <20 Gas Oil 40-45 20-30 10-25 Yield/Feed, wt % 0.866 0.840-0.860 0.850-0.875 Specific Gravity 0.7 0.1-0.5 <0.1 Sulphur, wt % 1800 >500 300-1200 Nitrogen, wt ppm Vacuum Gas Oil 20-25 Yield/Feed, wt % 25-35 20-35 1.010 0.925-0.970 0.925-0.935 Specific Gravity 2.2 0.5-2.0 0.25-0.5 Sulphur, wt % Nitrogen, wt ppm 4300 1600-4000 1500-2500 <2 <2 <0.3 Conradson Garb, wt % Vacuum Residue 10-20 15-35 Yield/Feed, wt % 30-60 1.160 1.035-1.100 0.990-1.030 Specific Gravity 2.7 1-3 0.7-1.5 Sulphur, wt % 11000 >3300 3000-4000 Nitrogen, wt ppm 26 >20 5-10 Asphaltens C7, wt % <1.1 1.2 Stability, Shell P-value 1.5-1.6
2.5. Slurry reactor process Several slurry reactor processes have been developed at the demonstration stage such as VEBA COMBI Cracking (VEBA Oel), UCAN (CANMET/UOP), Super Oil Cracking (Asahi Chemical) and HDH (INTEVEP) processes. Other slurry processes have been developed at the pilot stage such as MICROCAT RC (EXXON) and MRH (Idemitsu Kosan Co/M.W. Kellogg) processes. All these reactors use small amounts of additive, or catalyst precursors (from 0.1 to 3 wt%). The level of conversion can however reach 95%, but product quality is very poor as indicated in Table 3. 2.6. Products properties Table 3 summarises typical properties of products from the HDC of a Safaniya VR using various residue HDC processes. It can be observed that the properties of the converted products may be out of specification. In particular distillates may contain too much sulphur, aromatics and nitrogen for a direct use in other conversion processes or even for blending. Of particular importance is the high nitrogen content in these distillates indicating a poor hydrodenitrogenation (HDN) performance of these processes. The various distillates obtained from residue HDC processes may therefore require further hydroprocessing to reach cetane number and sulphur content requirements for diesel and low aromatic content for jet fuel for example. This can be obtained by the association of downstream hydrotreating processes at the same pressure than the residue HDC process so-called integrated hydrotreating. The hydrotreated atmospheric residue, on the other hand, is a good feedstock for residue catalytic cracking (RCC). The reduction in Conradson Carbon and metals in the hydrotreated residue allows to maximise the gasoline yield and prevents rapid deactivation of RCC catalysts, thereby limiting its consumption. In addition, the desulphurisation of the RCC feedstock reduces SOx emissions and sulphur content of liquid products. 3. RESID HYDROCRACKING CATALYSTS 3.1. Fixed bed catalysts Fixed bed processes use in general an association of several catalysts and guard bed materials within the guard reactor and the several reactors of the unit. For example, an association of up to 10 catalysts is claimed in the literature [14]. Therefore each catalyst as well as the whole catalytic system must be optimised to reach the unit objectives. This optimisation is closely connected to the choice of :he operating parameters (hydrogen partial pressure, liquid hourly spatial ''elocity, hydrogen/hydrocarbon ratio, temperature) and the nature of the feed. The concepts driving these associations are the particle size grading, activity rading and pore size grading. A recent study reports that the cycle length can be
10 increased by more than 50% by selecting the good grading of catalysts in fixed bed processes [17]. Particle size grading is used to minimise pressure drop and improve fluid distribution at the top of the first reactor. Activity and pore size grading are both needed to improve metal storage capacity, HDS activity, conversion level and run length. One typical association of catalysts is indicated in Table 4. It is a combination of an HDM, a balanced HDM/HDS and an HDS catalysts with a guard bed material on top of the first bed. The guard bed material is mainly used to trap solid particles remaining in the feed such as coke particles coming from heating tubes, iron scales from corrosion, dissolved impurities such as iron, arsenic, calcium containing compounds, sodium chloride, silicon contained in upstream additives, etc. The guard bed material also improve fluid distribution over all of the radial section of the reactor. New guard bed materials have been developed by a number of catalyst manufacturers (AKZO, CRITERION, TOPSOE, PROCATALYSE). These materials have very low surface area with an ultra large macroporosity (Table 4). This porosity increases the void space in the guard bed compared to filled alumina balls. Such a materiel trap iron much more efficiently than a conventional HDM catalyst as shown in Table 5. Table 4. Typical properties of fixed bed catalysts. Guard HDM material catalysts Sph./Extr. Shape Sph./Pellets Size (mm) 10-3 6-1.2 Active Phase Mo/NiMo <1 80-180 Surface Area (m^/g) <0.15 0.7-1.2 Pore Volume (cm^/g) MPD (nm) 20-100 10-100 |Lim Ni+V (wt% fresh cat.) >50 HDM act. High HDS act. Low Catalyst Load (%vol) <5 30-60
HDS/HDM catalysts Extr. 1.6-0.8 NiMo/CoMo 150-220 0.5-0.8 10-20 30-40 Medium Medium 10-40
HDS catalysts Extr. 1.6-0.8 CoMo 180-250 0.4-0.7 8-12 10-20 Low High 30-60
The main goal of the HDM catalysts designed for hydrocracking units is to have a high capacity for metal retention and to be stable as long as possible while carrying out the bulk of the conversion. This is obtained by optimising the HDM activity, pore structure and acidity of the catalyst. The main features of HDM catalyst are therefore a large pore volume to store as much metal deposits as possible and a large amount of macropores to allow resins and asphaltens to diffuse deep inside the grain and be cracked (Table 6). A significant example of HDM catalyst structure is the "chestnut bur" type, shown in Figure 5 [9,18]. This type of catalyst allows to obtain a uniform distribution of metals in the grain. Therefore high metal retention capacity up to lOOg
11 Ni+V/lOOg of fresh catalyst are reached depending on the feed and operating conditions as well as process objectives. Table 5. Analysis of used HDM catalyst and guard bed material Metal Guard HDM (wt %) material catalysts V 3.9 11.7 Fe 2.0 0.1 To improve HDS and HDC performances of fixed bed processes, a balanced activity catalysts can be used after the HDM catalysts. Its main function is to desulphurise the demetallised feed even if some noticeable trace of metals are still present in the feed (30-60 wt ppm). This catalyst has a porosity specially designed to accumulate a significant amount of metals and to provide more HDS activity (Table 4). HDM/HE3S
1 cm = 0,62 jLim
Figure 5. "Chestnut bur" HDM catalyst (SEM analysis). The refining catalysts located after the HDM catalysts, or after the HDS/HDM catalysts, must desulphurise the already deeply demetallised and transformed feed and reduce Conradson Carbon. The refining catalyst has a greater HDS function than the HDM function as indicated in Table 4. These catalysts have a porous texture specially designed for these reactions and a large specific surface area but a low metal retention capacity. The refining catalysts must also have a very low acidity as well as a good hydrogenation activity to limit deactivation by coking.
12 Table 6- Typical HDM catalysts properties for residue hydrocracking Fixed bed Moving bed Ebullated bed 6-1.2 Size (mm) 3-1.2 0.8 Sph./Extr. Sph./Extr. Extr. Shape Active Phase Mo, NiMo Mo, NiMo, NiV CoMo/NiMo 80-180 150-200 150-350 Surface Area (m^/g) 0.7-1.2 0.5-0.8 0.4-0.8 Pore Volume (cm^/g) 3.2. Moving bed HDM catalysts Moving bed HDM catalysts are not very different from fixed bed catalysts. While spheres are better suited than extrudates as they flow better, the use of extrudate is made possible by maintaining a small bed expansion in the reactor [14]. However, care must be taken that the mechanical resistance and particularly the attrition resistance of the catalyst be excellent to avoid too much formation of catalyst fine particles that will create problems for the downstream equipment and reactors. 3.3. Ebullated bed catalysts Ebullated bed reactors require specially designed catalysts. They can be either CoMo or NiMo in the form of small diameter extrudates in order to allow fluidisation. They must be very resistant to attrition to avoid too much catalyst fines formation. These catalysts must also work at high temperature and their acidic properties must be optimised accordingly. Some acidified catalysts, containing for example boron, silica or zeolite have been used to reach high level of conversion. However, this kind of acidified catalysts lead to too large level of sediment formation for industrial uses [19]. The pore size distribution of HDM catalysts is also an important factor to improve conversion level according to literature. Recent reports indicate that the use of bimodal alumina based catalysts with optimised pore size distribution can achieve higher level of conversion with acceptable level of sediment [20]. For the optimised catalyst, presence of macropores in the range 300-800 A is claimed to enable the cracking of a portion of large size asphaltens inhibiting sediment formation whereas the amount of mesopores in the range 100-200 A is claimed to improve HDS, HDC and HDM activities and favour cracking of smaller size asphaltens [20]. 3.4. Slurry reactor catalysts Many different types of materials have been claimed as catalysts, catalyst precursors or additives for slurry processes. Some examples are iron impregnated subituminous coal or lignite, pure iron sulphate, spent FFC catalysts, spent and
13 crushed hydroprocessing catalysts, iron rich clays, group VIA or VIII salts such as molybdophosphoric acid, molybdenum naphthenate etc... (21). Dispersion or particle size appears to be an important parameter to obtain a very high level of conversion. The process of formation of the actual active phase, which can be formed in situ or eventually prepared ex situ, as well as the injection of the catalyst in the reactor are also important parameters. 3.5. Catalysts deactivation Limiting the deactivation of the catalysts is one of the main challenge to develop improved catalysts whatever the type of reactor chosen. Resid hydrotreating catalysts are deactivated simultaneously by coke and metal deposits and it is not always clear which effect is dominating, in which operating conditions, and this at the different stage of the catalyst life [6]. In addition, the main mode of deactivation can be different whether one consider the HDM, HDS/HDM or HDS catalysts in an association of catalysts. The metals are known to deposit on the catalytic surface in the form of spatially dispersed large crystallites (2 to 30 nm) of metal sulphides such as V3S4 or V2S3 and Ni3S2 or Ni2S growing perpendicularly to the alumina surface from fixed nucleation sites [9,22]. This deposition lead to a decrease of the surface area and to the generation of pore restrictions that hinder the diffusion of large asphaltenic molecules inside the catalyst. The effects of metals on the catalytic surface are complex as they can interact with the active phase and be themselves active for instance in HDM [23]. Despite on going deposition of nickel and vanadium sulphides, which can reach up to 10-25 monolayers on the surface depending on the catalyst properties [11], the catalyst remains remarkably active. A number of hypothesis have been proposed to explain this effect such as the migration of deposited Ni and V sulphides to the uncovered alumina surface leaving the initial active phase unperturbed and accessible to the reactants [24] or the migration of the molybdenum sulphide on top of the deposited Ni and V sulphide layers [8]. It has been also proposed that the remaining activity is due to a combination of sulphides of vanadium, nickel and molybdenum, creating therefore new active sites [25]. Coalescence of the M0S2 slabs (crystallites of 30 nm were observed) and destruction of the CoMoS phase are other phenomenon partly responsible for the deactivation [28, 29]. Carbon or coke is recognised to accumulate quickly on the catalyst surface within the first days until the steady state is reached [9]. The initial deposits block smaller pores and may cause up to a 50% loss of surface area [10]. The amount, the rate and the nature of coke are essentially controlled by the operating parameters [25]. It has been shown that a small amount of metal deposits well distributed in the grain (< 1 wt %) can deactivate more efficiently the catalyst than a rather large amount of coke [19]. However, metals are usually not uniformely deposited in the grain and therefore both carbon and metals deactivate the active phase.
14 While for HDM and bimodal catalysts the main deactivation is mainly due to metal deposits, for HDS catalysts, the smaller pore size prevents a large fraction of the inner volume of the catalysts from deactivation by metals deposits. In that case, the HDS catalysts deactivation is likely dominated by coke build up on the catalyst surface [12]. The final rapid deactivation step is pore plugging by metals or carbon deposits. It can be progressive for macroporous catalysts or more sudden for mesoporous monomodal catalysts as shown by the very different deactivation curves reported in Figure 6 which have been obtained in isothermal conditions for two different catalysts. The amount of nickel and vanadium as well as the amount of coke necessary to plug a catalyst is therefore a function of the catalyst pore size distribution, its HDM and HYD activity, its acidity and the operating conditions used. 100
1
----
60
o
4-i
CO
HDS Catalyst
V
V
•> •M
40
o
^
^ ^
OH 20
["
HDM Catalyst
80
0) Q X
1
- - ^
"^-^ N
Time on stream Figure 6. Isothermal deactivation curve of typical HDM and HDS catalysts. I m p r o v e m e n t of the performance and cycle length of residue hydroconversion processes will need well selected association of catalysts each of them having a specific processing objective. While it is obvious that improving the resistance of each of the catalysts and of the whole catalytic system to deactivation by coke and metal deposits is one of the challenge of catalyst research and development, improving the HDN performance remain the most difficult challenge.
15
4. CONCLUSION Commercial residue HDC processes use either fixed bed, moving bed or ebullated bed technologies. The choice of process depends mainly on the amount of contaminants in the feed and on the required conversion levels while maintaining a continuous operation of minimum one year. Among the main developments made in the recent years, on stream catalysts replacement systems, swing reactors, improved flow distribution, uses of guard bed materials, coke resistant catalysts and complex association of catalysts with particle size, pore size distribution and activity grading have been implemented to obtain higher levels of conversion, better products qualities and increased run length while maintaining residual fuel stability. Further improvements will come from better optimisation of each catalyst and of the association of catalysts, from improved resistance of catalysts to deactivation by metals and coke and from improved hydrodenitrogenation activity. Although very capital expensive and in competition with thermal conversion processes, residue HDC will develop in the future as processes gain in maturity and as more and more synergy will be found with other conversion processes. REFERENCES 1. T. Ohtsuka, Catal. Rev. Sci. Eng., 16 (1977) 291. 2. W.I. Beaton, R.J. Bertolacini, Catal. Rev. Sci. Eng., 33 (1991) 281. 3. H. Koyama, E. Nagai, H. Torii, H. Kumagai, Stud. Surf. Sci. Catal., 100 (1996) 147. 4. F. M. Dautzenberg, J.C. Dedeken, Catal. Rev. Sci. Eng., 31 (1985) 121. 5. A. Al-Nasser, S.R. Chaudhuri, S. Battacharya, Stud. Surf. Sci. Catal., 100 (1996) 171. 6. C.H. Bartholomew, Chemistry Industries, Marcel Dekker, New-York, Vol. 58 (1994) 1. 7. R.J. Quann, R.A. Ware, C.H. Hung, J. Wei, Adv. Chem. Eng., 14 (1988) 95. 8. T.H. Heisch, B.L. Meyers, J.B. Hall, G.L. Ott, J. Catal., 86 (1984) 147. 9. H. Toulhoat, R. Szymansky, J.C. Plumail, Catal. Today, 7 (1990) 531. 10. D.S. Thakur, M.G. Thomas, Ind. Eng. Chem. Prod. Res. Dev., 23 (1984) 349. 11. P.W. Tamm, H.F. Harsberger, A.G. Bridge, Ind. Eng. Chem. Proc. Des. Dev., 20 (1981) 262. 12. M. de Wind, Y. Miyauchi, K. Fujita, Stud. Surf. Sci. Catal., 100 (1996) 157. 13. M. Absi-Halabi, A. Stanislaus, D.L. Trimm, Appl. Catal., 72 (1991) 193. 14. G.L. Scheuerman, D.R. Johnson, B.E. Reynolds, R.W. Bechtel, R.S. Threlkel, Fuel Proc. Techn., 35 (1993) 39. 15. LA. Wiehe, Ind. Eng. Chem. Res., 32 (1993) 2447. 16. S. Kamatsu, Y. Hori, S. Shimizu, Hydrocarb. Proc, May (1985) 42. 17. J. Bartholdy, B.H. Cooper, Stud. Surf. Sci. Catal, 100 (1996) 117. 18. C. Maier, J.P. Peries, A. Quignard, NPRA Annual meeting, Los angeles, 1986, paper AM-86-56.
16 19. E.P. Dai, C.N. Campbell, Chemistry Industries, Marcel Dekker, New-York, Vol. 58 (1994) 127. 20. A. Stanislaus, M. Absi-Halabi, Z. Khan, Stud. Surf. Sci. Catal, 100 (1996) 189. 21. A. Del Bianco, N. Panariti, M. Marchionna, Chemtech, Novembre 1995, 35. 22. B. J. Smith, J. Wei, J. Catal, 132 (1991) 21. 23. C. Takaushi, S. Asaoka, S. Nakata, Y. Shiroto, Preprints, ACS Div. Petrol. Chem., 30 (1985) 96. 24. Y.W. Chen, M.C. Tsai, C. Li, Ind. Eng. Chem. Res., 33 (1994) 2040. 25. H.D. Simpson, Stud. Surf. Sci. Catal., 100 (1996) 265. 26. J. Bartholdy, B.H. Cooper, A.C. Jacobsen, Preprints, ACS Div. Petrol. Chem., 38 (1993) 386. 27. G. Gualda, S. Kasztelan, Stud. Surf. Sci. Catal., 88 (1994) 145. 28. J. Wei, X. Zhao, Chem. Engi. Sci., 47 (1992) 2721. 29. X. Zhao, J. Wei, J. Catal., 147 (1994) 429.
® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
17
AN IMPROVED PROCESS FOR THE PRODUCTION OF ENVIRONMENTALLY FRIENDLY DIESEL FUELS J.Grootjans and C.Olivier FINA RESEARCH S.A. Zone Industrielle B-7181 FELUY 1. INTRODUCTION In 1996, the European Commission has issued a draft directive for fuel specifications and passenger car emissions to be implemented by the year 2000. The directive announces a possible further severization of fuel specifications to become effective by the year 2005. The EU member states will be authorized to require special fuels, generically called "city diesel", in specific local areas with extreme air quality problems. For automotive diesel, the proposed specifications for the year 2000 are:
Cetane Number Spec. Grav. T-95 % Poly-aromatics Sulphur
g/ml C wt % ppm
51 Min 0.845 Max 360 Max 11 Max 350 Max
Swedish "Class 1" diesel is a well known example of a local fuel having the following extreme properties: Cetane Number Spec. Grav. T-95% Total Aromatics Sulphur
g/ml C vol% ppm
50 Min 0.800 - 0.820 285 Max 5 Max 10 Max
Considering the uncertainty of future automotive diesel specifications and the need to adapt the refining schemes in order to comply with the year 2000 specifications, the oil refining industry needs to implement refining processes with great flexibility. The mid and longterm anticipated diesel specifications raise several conflicting issues: • The increase of cetane number will tend to further back out cracked stocks from the diesel pool unless more efficient hydrotreating catalysts become available. Another option is to reduce kerosine blending into the diesel, because of its low natural cetane number.
18 • The pressure on density and distillation range implies that less high boiling range material can be tolerated in the raw diesel cut. Low densities will further reduce the possibility to blend cracked stocks . • Further reduction of sulphur and a possible control on aromatics content are a further challenge to catalyst manufacturors and suppliers of refining technology. The most refractory sulphur species are found in cracked stocks and high boiling range material. Since consumption of diesel is still growing all over Europe, the new specifications create a great need to convert heavy atmospheric gasoils into more valuable products. 2. A FLEXIBLE REFINING SCHEME Figure 1 describes schematically a refining scheme that is believed to provide maximum flexibility for coping with future more stringent diesel specifications. The middle distillate is separated on the crude tower into a raw diesel cut and a heavy atmospheric gasoil. In general, the refiner has great flexibility in adjusting the cut point. The diesel stream is desulphurized on a conventional hydrodesulphurization process. Lowering density and T95 of this virgin stream will obviously facilitate deep hydrodesulphurization. The HAGO (heavy atmospheric gasoil) is treated on an AKZO/FINA CFI process. This process developed by Fina Research and AKZO NOBEL combines catalytic dewaxing and hydrotreating such as mild hydrocracking into a single process step. The operating conditions can be adjusted over a wide range according to the feed properties and processing objectives. Of particular interest is the conversion to diesel range material that can be achieved on this process as well as the dramatic improvement of cold flow properties. The stabelized product is separated under cut point control into a diesel stream and bottoms, that can be recycled to the reactor or upgraded as fuel oil component or FCC feedstock. The diesel product is added to the desulphurized virgin diesel and hydrodearomatized on an Arofining process. This process, developed by Fina Research, converts the aromatics into the corresponding naphthenes while further desulphurizing the "city diesel" type product. Arab Light crude oil has been selected for developing this refining scheme. The inspections of the middle distillate streams are summarized in Table 1.
19 CRUOE TOWER
STAESUZER
SIRIFFERS
m I
Naphttia
HDS
ft
J-
ft
t
*—• AIOBIMNG -L-
rT«
I.
Diesd
Naphtha
CH/MHC
SPUTIER
X
-•
BngetoRidQa
Figure 1. Refining Scheme for the Production of "City Diesel"
DIESEL
'
NapUfaa
HAGO
Yield on Crude
wt%
30.0
10.0
Spec.Grav. Sulphur Nitrogen Cloud Point
g/ml wt% ppm C
0.8266 0.901 62 -18
0.8831 1.766 443 23
Pour Point
C
-24
21
51.6
49.6
27.2
36.4
< 350 C
85.0
27.0
<370C
90.0
34.0
Cetane Index HPLC Aromatics
wt%
TBP Distillation
wt%
Table 1. Analyses of Middle Distillate Feedstreams
20
2. DIESEL HDS Figure 2 summarizes the performance of a state of the art HDS catalyst operating at medium pressure. In spite of the dramatic increase in activity of commercial diesel HDS catalysts observed over the last decade, it remains difficult to reduce product sulphur content much below the 500 ppm level, unless the feedstream is cut shorter. At medium operating pressure, an increase in cetane number of 6 to 8 points can however be achieved. The heavy tail boiling above 340 C contains the more difficult substituted aromatic sulphur species, that are the most refractory for conventional Co/Mo catalysts.
DIESEL HDS
WT% S 0.2
0.18+ ^} :W;d^^%M:yl^^ 0.16+''-'--^l^^^^^i^^^^ 0.12 4 * - - ^^ ^; ••;i^-^',:/:: :-^*^^
0.1 +:
0.04 t 0 . 0 2 "p-". -visV/ '.',i^'^
I '
'"'•• '^"' ''''-'^'
Figure 2. Performance of a "State of the Art" Diesel HDS Catalyst 3. HAGO CFI/MHC The "Cold Flow Improvement"/ Mild Hydrocracking process combines in a single process step, hydrotreating and hydrocracking of hydrocarbons and selective cracking of linear or slightly branched paraffins on a zeolitic catalyst of balanced activity.. The catalytic system can be designed and optimized for each particular application. In the proposed refining scheme, conversion of the HAGO into naphtha and deeply desulphurized diesel is the main objective. Figure 3 illustrates the yields of naphtha and diesel product obtained over a wide conversion range in once trough operation. Full conversion of 370+C material can be achieved by recycling the non-converted bottoms.
21
HAGO CFI/MHC 50 45-j40 + 35 t ^ 30 Q 25-f
', '. '~:'Sn:^:i'^^::z,x/t,^S-^;
-I
UJ
>- 20 4'
15-1" 10
5f 15
0
20 25 370+C CONVERSION WT%
30
Figure 3. CFI/MHC Selectivities vs Conversion The stabilized product is separated on a splitter into a diesel stream and bottoms having the following properties: Property TBP Cut Yield on HAGO Specific Gravity Sulphur HPLC Aromatics Mono Poly Cetane Number 1 Pour Point
C Wt% g/ml ppm Wt%
C
DffiSEL
BOTTOMS 1
180-370 43.97 0.8712 485
370+ 29.57 0.8906 670
37.6 13.9 41.7 <-51
-24
Table 3. CFI/MHC Diesel and Bottoms Inspections The diesel product has a medium aromatics content and excellent cold flow properties. The bottom product is an attractive fuel oil component or could be further upgraded as FCC feedstock. The effect of the splitter cut point on diesel sulphur content is illustrated by Figure 4:
22 SULPHUR DISTRIBUTION 500 450 400
(0 i
350
:
300
'
*
^'
*-
Q.
250200 150
300
310
320
340
330
350
360
370
380
390
400
CUT POINT DEG C
Figure 4. CFI/MHC Diesel Sulphur Content vs Cut Point 4.AROFINING Arofining is a hydrodearomatization process developped by Fina Research in the early seventies. Typically, the process is used to produce solvents and burning kerosene. The key feature of the process is the sulphur resistant catalyst, for which recently, a new generation has been developped. The catalyst formulation is based on a noble metal impregnated silico-alumina carrier. The difficulty of silico-alumina carriers is the control on impurities and their formulation into pellets or extrudates when starting from the inorganic precursors. A stepwise increase of activity has been achieved by formulating the carrier from organic precursors, permitting full control on impurities and shaping. A particular property of the catalyst is that sulphur species inhibit the dearomatization activity. This phenomenom is however compensated by increasing the operating temperature and fully reversible, even after large feedstock sulphur spikes have gone through the system. Nitrogen however, adsorbs much more strongly on the active sites and eventually, hot hydrogen stripping is required to recover the initial activity. A rate equation of the form: k.p(A).p(H2) r =
1 + K(H2S).p(H2S) +K(NH3).p(NH3)
can be used to predict the product aromatics content. The activation energy Ea was determined from experimental data using a sulphur and nitrogen free feedstock. The adsorption terms were determined using sulphur and nitrogen containing feedstocks and analyzing the ratios k(feed) k (N=S=0)
1 1 + K(H2S).p(H2S) + K(NH3).p(NH3)
A reasonable fit of the experimental data was found for a heat of adsorption of H2S of 4,810 cal/mole and 11,465 cal/mole for NH3. This indicates that as expected, feedstock nitrogen inhibits the activity more strongly then sulphur.
23 For different feedstocks predicted product aromatics content is compared to the experimental data in Figure 5.
10
12
A (Exp) Wt%
Figure 5. Parity plot for product aromatics content By adjusting cut points on the crude unit and on the CFI/MHC splitter, and by selecting the appropriate operating conditions on the diesel HDS and HAGO CFI/MHC processes, a feedstock to the Arofining process having properties as summarized in Table 5 is obtained. Specific Gravity Sulphur HPLC Aromatics Mono Poly T95 Cetane Number i Cloud Point
ml/g ppm Wt%
c c
0.8273 200 25.5 2.2 340 59.3 -9
Table 5. Combined AroHning Feedstock This feedstock was dearomatized at medium and high severity, yielding the diesel products as summarized in Table 6:
24 Property
Medium Severity
High Severity 1
Yield on feed Wt% Spec. Grav. g/ml Sulphur ppm ASTMD1500 colour HPLC Aromatics Wt% Cetane Number Cloud Point C
97.0 0.819 30 <0.5 10.0 68.6 -10 337
96.0 0.813 5 <0.5 5.0 69.2 -10 335
|T95
C
Table 6. AroHned Diesels 5. HYDROGEN CONSUMPTION Table 7 summarizes the material balances of the hydrotreating processes, the hydrogen consumption refers to the chemical consumption only:
Feed H2
Wt%
Fuel Gas LPG Naphtha Diesel 1 Bottoms
DIESEL HDS
HAGO CFI/MHC
100 0.67
100 0.97
0.95
2.53 9.36 15.54 43.97 29.57
0.60 99.12
AROFINING 1 10 % A 5%A 100 100 0.78 1.16 0.04 2.00 98.74
0.18 3.41 97.57
Table 7. Material Balances On total middle distillate feedstreams, the material balance is summarized as follows: Diesel Wt% Aromatics Middle distillate Hydrogen Fuel Gas LPG Naphtha Diesel Fuel Oil
Wt%
10
5
100.00 1.41
100.00 1.74
1.35 2.37 6.03 84.26 7.40
1.35 2.48 7.25 83.26 7.40
Most noticable is the very substantial hydrogen consumption. The naphtha produced in this refining scheme can be upgraded by catalytic reforming. For the selected crude origin, the N+2A content is of the order of 65 %.
25 6. CONCLUSIONS A flexible refining scheme has been developed for producing environmentally friendly diesel products . The flexibility stems from the use of very effective catalysts and optimal control on the cut points of the feedstreams and process severities. The examples have been developed starting from Arabian Light middle distillates, which is a rather severe case. Obviously, other scenarios can be developed on the basis of less extreme product properties, or starting from lighter crude oils and cracked stocks. Acknowledgement: Part of this work was supported by the European program JOU 2-CT92-0121. References: 1. H.W.Homan Free, T.Schockaert and J.W.M.Sonnemans, Fuel Processing Technology, 35 (1993) 111-117 2. C.Olivier and H.W.Homan Free, HTI Quaterly: Autumn 1995
This Page Intentionally Left Blank
® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
27
Hydroprocessing to Produce Reformulated Gasolines: The ISAL™ Process G. J. Antos UOP Des Plaines, Illinois U.S.A. and B. Solari, R. Monque INTEVEP, S.A. Caracas, Venezuela
As a result of forthcoming environmental regulations, such as those established in the U.S. Clean Air Act, petroleum refiners around the world are searching for low-investment solutions to accomplish hydroprocessing requirements These regulations are expressed in the Complex Model, which brings together sulfur content, olefin content, vapor pressure, and boiling range in an interactive fashion. Processes need to successfiiUy reduce these pollutants and still maintain high yields of high octane-product. Because FCC naphtha accounts for 90% of the sulfijr and olefins in the entire gasoline pool, hydrotreating this material is an attractive process alternative, provided that octane losses are minimized. The ISAL"^"^ process, jointly developed by INTEVEP, S.A., and UOP, is a selective hydrotreating process to improve naphtha quality by reducing sulfiir, nitrogen, and olefins without octane loss. The paper discusses aspects of the process and catalyst chemistries leading to these desirable results. 1. BACKGROUND Refiners who are involved in producing transportation fiiels continue to face challenges that are the result of increasingly stringent legislation covering vehicular emissions and fiiel quality. Environmental regulations covering both gasoline and diesel fuels are appearing in many areas of
28 the world. Regulations covering gasoline are similar to those established within the framework of the U.S. Clean Air Act and subsequent amendments. The time line indicating the timing of critical events in gasoline regulation is shown in Figure 1. Reformulated gasoline to meet Phase I regulations has already been brought to the marketplace. Still ahead are the implementation of the Complex Model for gasoline formulation and the introduction of Phase II reformulated gasoline in the year 2000.
1990
- Clean Air Act Signed (11/90) • Reg-Neg Begins (2/91) -Agreementin Principle (8/91) - OxyFuel Begins (11 /92) • Complex Model & Final Regulations (12/93)
1 9 9 8 T : > - Complex Model Mandatory (1/98) 1999^ 2 0 0 0 ^ — RFG-Phase II
Figure 1. Gasoline Reformulation Time Line To meet the requirements of these regulations with a minimum impact on profitability, new technology is needed. Such a process with a new catalyst has been under development in the laboratories of INTEVEP, S. A., and together UOP and INTEVEP, S. A., have completed the development and offer the ISAL"^"^ process and catalyst technology for license.
29 2. EXPERIMENTAL SECTION 2.1. Feedstocks Naphthas and fluid catalytic cracking (FCC) gasoline samples were obtained from several commercial refinery sources. Pure hydrocarbons were used in microreactor testing. 2.2. Analytical Hydrocarbon analyses were those typically performed in petroleum laboratories, including gas chromatography, mass spectrometry, elemental analyses, and physical properties analyses. Research octane number (RON) and motor octane number (MON) were measured by standard engine techniques. 2.3 Catalyst The proprietary ISAL catalyst is generically a bifunctional catalyst. It contains a specially modified acidity function. Non-noble metals supply the hydrogenation function. ISAL-type catalysts are preparations with variations in the catalyst formulation used to exemplify particular catalytic features. The catalysts were tested in the extruded form. 2.4. Pilot Plant Testing A variety of pilot plant equipment was used in the testing, from typical microreactor equipment for pure component work to larger-scale pilot plants for the experiments with commercial naphthas. Mass balances were obtained as needed to yield data for process characterization. 3. IMPACT OF THE COMPLEX MODEL The components of the gasoline pool in a refinery vary for each specific case. Average contributions from the various gasoline component processes have been computed. ^ Figure 2 shows a typical scenario.
30
Coker Gasoline 1% Light SR 2%
Poly Gasoline 3%
Figure 2. Average U.S. Gasoline Pool The Octane of the gasoline is a result of contributions from significant concentrations of aromatics, olefins, and branched paraffins. The endpoint, benzene concentration, and sulfur level are a function of the feedstocks used in the various contributing processes. To meet the emissions reductions called for in the Clean Air Act of 1995, the Simple Model was used to generate the requirements for Phase I reformulated gasoline. The basis was the refiner's own 1990 product average along with mandated benzene and oxygen levels. The emissions reductions called for in Phase II are based on the Complex Model. The reference basis changes to an industrywide baseline. The refiner may choose any composition as long as oxygen and benzene specifications are met and the emissions reductions are achieved. The year 2000 goals require further reductions in emissions of volatile organic compounds (VOCs), toxics emissions, and also a reduction in NOx emissions. The emissions of VOC, toxics, and NOx depend on many compositional parameters, and this forms the basis for the Complex Model."^ Figure 3 demonstrates the impact of olefin, sulfur, and aromatics concentrations on the reduction of NOx emissions.
31 Effect of Olefins
20 10
1990 U.S. Industry Avg. /~9.2vol-%
10
30
15 20 Olefins, vol-%
Effect of Sulfur
800
400 600 Sulfur, ppm
1,000
Effect of Aromatics
10
20
30
40
50
Aromatics, vol-%
Figure 3. Impact on NOx Emission Reduction Although a refiner may theoretically have the ability to choose whatever composition he wishes, in actuality the pool components may dictate specific courses of treatment that need to be
32 employed to attain compliance. Figure 4 relates these same three variables in an operating network. The refiner must keep in mind his octane needs as well, and these needs are heavily influenced by olefin and aromatic concentrations. Figures 3 and 4 indicate that sulfiir content has the largest impact on NOx emission reduction and that sulfiir reduction in the gasoline pool is a likely scenario for many refiners. Sulfiir level also has a medium-sized effect on toxics emissions reductions as well. Furthermore, in the proposed California Air Research Board (CARB) regulations for gasoline sold in California, sulfiir level is limited to 50 ppm. Clearly, the Complex Model will significantly affect refiners who wish to market in the United States or Europe. Hydrotreating to remove sulfur will be the key to meeting these regulations.
Summer Region 2 RFG - NOx Reduction = 6.8%
20
*^*^"""-.''"-.
15
10
r
^
"-:"••••... ^x ^ \ *^^ •"••••
L
^ N
'\^ '•"..% ^^\ *^ \
5h 01 0
\
Aromatics, vol-7%
\
100
• •.
\
\
^
\
'•'
v
\ vX
\^ \ ^
35 26.3
\
20
200 Sulfur, ppm
Figure 4. Trade Sulfur for Olefins or Aromatics to Meet Phase 11 NOx
\ >
11
300
15
400
33 4. FCC GASOLINE POST-TREATMENT Extensive Hydrotreating of the total worldwide gasoline production capacity of more than 2x10^ barrels per day would be expensive.^ However, inspection of the gasoline process components indicates that approximately 98% of the sulfur in the pool is present in the FCC naphtha (Figure 5). Further, an analysis of one typical sample of FCC gasoline indicated that most of the sulfur is in the heavier cuts of the naphtha (Figure 6). Figure 6 also indicates the nature of the sulfur molecules present in the various cuts. Similar results confirming these observations have been reported elsewhere."^
HeayFOCrsbphtha
(13?-22?Q
U^FOCI^phtha (C^13?Q ,^2%
Oi^^-^
Figure 5. Sulfur Contribution in Gasoline Pool Hydrotreating the entire feedstock to the FCC unit is an attractive choice in terms of the ultimate FCC gasoline product; the tangible benefits are higher overall gasoline yields, lower sulfur in the gasoline, and lower SOx emissions. However, capital investment for feed desulfurization is estimated to be at least $20 to 40 million greater than for posttreatment. Thus, an FCC gasoline posttreatment process becomes the most likely scheme for achieving sulfur reduction in the pool.
34
8000
C2SH CjSH
C3SH
C4SH
C5SH ^
CH3
^C^Hg
^C3H,Q^|
s
6000 B
^ 3
4000
C/3
2000
0
L
FBR 109 138 151 169 194 214 237 271 286 324 347 392
EP
TBP Cutpoint, °F Figure 6. Sulfur Distribution in FCC Gasoline The various hydrotreating reactions that take place in posttreatment processing are: •
Sulfur compounds + H2 -^ H2S + Paraffins (thiophene > benzothiophenes > alkyl substituted benzothiophenes)
•
Olefins + H2 ^ Paraffins
•
Aromatics + H2 ^ Naphthenes
•
Paraffins + H2 -^ Cracked products -^ C4"
•
Paraffins ^ Iso-paraffins
In addition to the beneficial desulfurization reactions, octane loss may be a potential problem because olefins in particular are hydrogenated (Figure 7). If significant cracking to C4- products takes place, the loss of gasoline yield becomes a major expense for posttreating. Taking these factors into account, octane barrel loss during posttreatment can be significant if the process and catalyst are not optimized.
35
u ILU
•
-3
IZI
Q
I
o §
-6 IZI
-9
• IZI
12 1/;
0
1
5
10
15
1
20
25
30
1 35
Absolute Reduction in Olefin Content (vol% feed Olefins - vol% product olefins) Figure 7. FCC Gasoline Hydrofinishing, Olefin Saturation Results in Octane Loss 5. CATALYTIC PERFORMANCE The conversion of pure compounds is instructive in identifying the catalytic traits of the ISAL^^ catalyst. An ISAL-type catalyst was used. Microreactor studies were run with the pure compounds N-octane (NCg); M-xylene (M-X); 1,3 dimethylcyclohexane (1,3 DMCH); and N-octene (C8=). No nitrogen or sulfur compounds were added. Conditions chosen were in the range expected for the process, with constant temperature of 320°C at 400 psig reactor pressure and 1 LHSV. Microreactor test results are on a weight basis, assuming 100 g of feed. The reactivities are those representing the initial state of the catalyst and as such are extremely high. In actual operation with nitrogen- and sulfur-containing feedstocks, activity will be reduced. Cracking and aromatics production are expected to be reduced to a greater extent.
36 Table 1 exemplifies the catalytic features for the ISAL-type catalyst. Paraffins and olefins are the most reactive, and olefins exhibits the highest rate of conversion. Although isomerization appears to be the most significant reaction for all molecules, the catalyst is able to achieve molecularweight reduction, dehydrocyclization, and recombination of hydrocarbon fi-agments. Table 1. Pure Component Reactivities over ISAL-Type Catalyst NC8
M-X
1,3DMCH
C8=
Unreacted
45.5
75.1
91.6
1.2
Isomerized
26.2
21.6
5.9
29.5
Cyclicized
5.7
2.8
-
14.4
Aromatized
4.1
-
0.2
7.6
Cracked (C7")
18.5
0.5
2.3
47.3
Feedstock Products:
Overall, the catalyst exhibits reactivities that are desirable for FCC naphtha posttreatment: strong hydrogenation capability, isomerization capability, and ability to reduce molecular weight without sacrificing aromatics. As a result, the ability to reduce sulfixr level in the FCC gasoline with minimum impact on RON and MON and minimized Cs^ loss is a real expectation with the optimized catalyst functionalities of the ISAL^^ catalyst. Feedstocks from refineries demonstrate these same aspects of the ISAL^"^ catalyst performance. Figure 8 demonstrates the processing of a heavy virgin naphtha, a Ce^ FCC naphtha, and a Cg^ FCC naphtha over the ISAL™ catalyst. The road octane (RON+MON/2) of the product is shown as a flinction of volumetric liquid yield. The losses of road octane observed with hydrotreating of the FCC gasoline in Figure 7 can be offset by employing the ISAL'r'^ catalyst, as demonstrated in Figure 8. Yield loss may be minimized, and in fact overall production of octane barrels may be increased despite this loss.
37 Table 2 relates experimental data comparing the results of hydrotreating a Cg^ FCC naphtha with those obtained by processing the same stream with the I SAL technology. Significant reduction in the loss of road octane is observed with the ISAL process. Other data has been reported elsewhere. Table 2. Data Comparing Hydrotreating and ISAL Technologies Feedstock:
Product:
Cg^ FCC Naphtha
Hydrotreating Technology
ISAL Technology
Sulflir, wppm: 975
<10
<10
Nitrogen, wppm: 102
<1
<1
RON+MON/2 : 83.7
72.5
81.0
Olefins, wt%: 25.1
2.7
2.8
Volume: 100
100.4
97.0
The processing of actual refinery feedstocks with ISAL technology demonstrates that this expectation is met. Figure 8 presents the data from the processing of an C8+ cut of a particular FCC naptha over the ISAL catalyst. The road octane (RON+MON/2) of the gasoline product is shown as afiinctionof the volumetric liquid yield. The product sulfiir is less than 10 wppm in all cases. The road octane of the feed is indicated as is the road octane observed if the naptha were severely hydrotreated to the same sulfiir level. Full recovery of road octane, which is experienced with hydrotreating, is attained with the ISAL process and catalyst. Full recovery is achieved with minimal C5+ liquid yield loss. If reid vapor pressure (RVP) is not limiting, some of the C4 hydrocarbons may be recovered and blended into the gasoline to enhance octane barrel recovery. Figure 9 indicates that MON recovery is the key. In fact, raising MON to values higher than that of the feedstock to the ISAL process can be achieved by selectively choosing processing conditions. Additional data have been published elsewhere.
38 86 84 „^^ CM Z
Fres/? Feed
82
O 80 + z 78
o D^ 0)
76
c 74
4-<
(0
o
o •o
(0
o Q:
72
Severely HDT Feed -
70
— I
1
1
91 92 93
1
1
\
1
1
94 95 96 97 98
1
1
99 100 101
Cs* Yield, LV-%
Figure 8. Recovery of road octane using ISAL catalyst 81 80 79 78 77 76 E 75 z 74 73 c 72 (0 o 71 70
o
o o
-| -
f^Q t
L
^A
*T^^ ^
TTTtVv,,,^^
Fresh Feefi^-^.^
4
• \ •
• QAx/Ar^K/ U r ^ T C ^ A r l
*
1
91
92
oevereiy n u1 i reeu 1 1
— 1 — — 1
93
94
95
96
97
1
98
1 —— 1
99 100
Cs* Yield, LV-%
Figure 9. Recovery of Motor octane utilizing ISAL catalyst
39 6. CONCLUSIONS The ISAL"^"^ technology employs a catalyst that allows the refinery to hydroprocess FCC gasoline to improve its quality. Reductions in sulfur, nitrogen, and olefin content are all achieved. The T90 of the FCC naphtha may also be reduced. Aromatics levels remain relatively constant. The ISAL"^"^ catalyst is therefore able to produce a gasoline with no octane loss, and in fact octane gain may be achieved. FCC naphtha upgraded with the ISAL processing technology is ideally suited for meeting reformulated gasoline specifications in today's global market.
REFERENCES 1.
S.T. Sie, Evaluation of Catalysts for Reforming. In: G.J. Antos, A.M. Aitani, J.M. Parera, eds., Catalytic Naphtha Reforming: Science and Technology, Marcel Dekker, Inc., New York, 1995, p. 182.
2.
The Challenge of Reformulated Gasoline: An Update on the Clean Air Act and the Refining Industry. UOP, Des Plaines, 1994.
3.
L.D. Krenzke, and others, Hydrotreating Technology Improvements for Low-Emissions Fuels, 1996 NPRA Annual Meeting, March 17-19, 1996. Paper AM-96-67.
4.
H.J. Lovink, Naphtha Hydrotreatment. In: G.J. Antos, A.M. Aitani, J.M Parera, eds.. Catalytic Naphtha Reforming: Science and Technology, Marcel Dekker, Inc., New York, 1995., p. 261.
40 5.
T.A. Nguyen, and M. Skripek, Reducing Sulfur in FCC Gasoline via Hydrotreating. 1994 AIChE Spring National Meeting, April 17-21, 1994. Paper 54C.
6.
J. A. Salazar, and others, ISAL Technology; A New Alternative to Producing Reformulated Gasolines. Preprints of Division of Petroleum Chemistry, American Chemistry Society 212th National Meeting, August 25-29, 1996.
® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
41
Molecules, catalysts and reactors in hydroprocessing of oil fractions W.HJ. Stork Shell International Chemicals B.V., P.O. Box 38000,1030 BN Amsterdam, Netherlands Abstract With the tightening product specifications and poor refinery margins the demands on catalyst performance are ever increasing. Obviously the reactor configuration and reaction conditions under which a catalyst operates in the refinery are crucial for its performance. It is illustrated for many examples in hydroprocessing of oil fractions that improved and novel reactor systems/conditions offer gains in performance that form a most welcome addition to those from improved catalysts by themselves. It is expected that such improved reactor systems will increasingly be used. The combined rather than separate development of such new reactor systems and the dedicated catalysts to maximally exploit these systems can offer a large additional synergy.
1. INTRODUCTION Oil fractions are (hydro)processed in the refinery primarily to manufacture the transportation fuels, that are presently essential to our society. The increasingly tight environmental product specifications, such as for sulphur content, put ever stronger demands on the refinery systems. On top of this, the refinery margins are currently very low, which requires that these tightened specifications are met in a very cost-effective fashion. In this area, hydroprocessing can indeed play a major role, for instance by increasing the depth of desulphurization in a hydrodesulphurization unit. Without new or improved technology, however, it has been estimated that to meet the tighter sulphur specifications in Western Europe alone, extremely high investment would be required. Obviously there is a strong incentive to improve the technology and thus reduce the required investments. One very effective way is the development of improved catalysts, since these can generally be used in existing units without capital investment. Thus, the theme of this conference is very well chosen. A sole concentration on improved catalysts would however be too narrow an approach, since this is only one option: also improved reaction conditions and improved reactor systems/configurations can significantly contribute to meeting the above challenge. At the suggestion of the organizing committee we will therefore discuss integrated approaches to the challenges listed earlier Starting from the basic chemical requirements, we will review the diversity of reactor concepts and catalysts used in the various areas of hydrotreating and hydrocracking processes, as expressed in the tide: Molecules, catalysts and reactors in hydroprocessing. In this we take the normal trickle-flow reactor, as shown in Fig. 1 [1], as a familiar starting point. The reactor contains the catalyst bed(s), with above that the liquid and gas distributing systems, a variety of larger particles above and below the catalyst bed (e.g. to reduce fouHng), quench systems using gas or Hquid, etc. The design of these units, internals, dimensioning, bed lengths, distributing systems etc to achieve maximum efficiency
42
SUPPORT BALLS
CATALYST
SUPPORT BALLS •GRID
li'
PRODUCT
Fig. 1. Downflow Reactor [1].
and safety against runaways, minimize fouling etc contains a lot of proprietary knowledge, which is outside the scope of this paper; suffice it to say that with new steel types being developed continuously new possibilities emerge in terms of maximum temperatures and hydrogen pressures. Sometimes external factors also play a crucial role in determining the size of a reactor — in one case it was the maximum size that could be transported over an existing railway! 2. CATALYTIC REFORMING — AN EXAMPLE As a striking example for the interplay between chemistry, catalysts, reactors and process conditions, though outside of hydrotreating, I will briefly discuss the developments in catalytic reforming [2]. In this important process essentially the octane number of a naphtha (oil fraction with boiling point between about 80 and 200°C) is increased and clearly, with the lead phase-out in gasoline, the demands on reforming have increased continuously. The main reactions are aromatization through dehydrogenation of naphthenes and dehydrocyclization of paraffins; hydrocracking is an important undesired side reaction, because it decreases the yield of liquid product and hydrogen. These reactions are sunmiarized in Fig. 2.
Naphthenes —> Aromatics + H2 Paraffins —> Aromatics + H2 Naphthenes, Paraffins -h H2 —> Hydrocracked Products Feedstock + Catalyst —> Coked Catalyst + H2 (Catalyst Deactivation) Fig. 2. Main reactions in catalytic reforming.
43
CATALYST
CENTER PIPEPRODUCT
Fig. 3. Radial flow reactor [1].
The reactions overall are strongly endothermic, and hence a series of reactors is used with interstage heating; these reactors are of the radial flow (Fig. 3) rather than axial flow type to reduce the pressure drop. Several commercial processes have been developed along these lines, such as UOP's Platforming, IFP's reforming, Engelhard's Magnaforming, Exxon's Powerforming, Chevron's Rheniforming and Amoco's Ultraforming. Differences between these are discussed in [3]; one of the variations is the envisaged lifetime of the catalyst, ranging from a week to a month (low. pressure operation) in cyclic, fully regenerative processes (Fig. 4a) where an extra swing reactor allows taking out one reactor of the train for regeneration, to semi-regenerative processes (operating at higher pressure), with a life time of about 1 year (Fig. 4b). For thermodynamic reasons one wants to operate the process at high temperature and low hydrogen partial pressure; this also enhances the rate of the difficult dehydrocyclization reaction. High hydrogen pressures, which reduce catalyst deactivation by coking (Fig. 5 [4]), on the other hand, also increase the rate of hydrocracking. Thefirstreformers using (monometallic) platinum catalysts were therefore run at a hydrogen pressure of about 35 bar [3]. The advent of the more stable bi(multi)metallic (Pt/Re/(S)) catalysts has allowed a significant reduction in operating pressure, significantly contributing to an increased yield in the "semi-regenerative" ("SR") reformers, as shown in Fig. 6 [5]. Recent refinements here are the use of multiple catalyst systems (e.g. UOP's R72/R56 [6], or Criterion's Pt/Sn//Pt/Re system, PS7/PR9 [12]), where the front reactors in which the rapid endothermic naphthenes dehydrogenation occurs on average at low temperatures use a different catalyst (with high selectivity/lower stability) than the back-end reactors, in which the slow dehydrocyclization occurs at high temperatures putting higher demands on stability, leading to overall yield improvements, and the use of skewed Pt/Re catalysts, (high Re), intrinsically very stable but sensitive to sulphur, often in combination with
44 (a)
NET HYDROGEN TO REFINERY ,
RECYCLE GAS
& \
m P. p^ l_^ L,^
FLASH DRUM
TO STABILIZER
NET HYDROGEN TO REFINERY ,
RECYCLE GAS
Pk pL
Krarrn L ^ t^
0
& FLASH DRUM
-REGENERATION PIPING TO STABILIZER
Fig. 4. Schematic flow diagram of semi-regenerative and cyclic reforming processes [1].
45
T(°C)
-
@ = Pressure (bar)
/
10
St 1 RON = C
@)
/
/
^^@^^
5 To
1
1
0
1
.
1
3
.
1
5 Time (arbitrary unit)
Fig. 5. Typical influence of pressure on stability [4]. 100 HYDROGEN Q
_J UJ
90
> § 80
'^^_ v:
0^
'^--/\
"
•^.^ '^-.^
C5+ PLATFORMATE
CL
70
143
0—
C C^ • Co + C3 +^A 1 ^
Q
O a.
n—
285
""^
428
570
PRESSURE,psig
Fig. 6. Effect of pressure on yield structure [5],
a sulphur guard bed. Sulphur guard beds appear imperative in the operation of the new, zeolitic type of aromatization catalysts, Pt/K/L, of the Aromax process, that are very selective to lower aromatics [7], However, the basic issue in reforming is that the conditions that favour a higher liquid product and hydrogen yield also lead to a lower catalyst stability. It was therefore a major step forward when the importance of catalyst life was drastically reduced by UOP's introduction in 1971 of the continuously regenerated reformer, "CCR", shown schematically in Fig. 7 (later also IFP introduced a continuous process). Here the catalyst is continuously replaced and regenerated after only days of operation which allows a drastic reduction in operating pressure, currently down to some 3.5 bar, which favours the desired reactions and which suppresses the undesired hydrocracking reaction. A large overall yield increase as shown in (Fig. 6) is the net result, and clearly the majority of the new units are of the low pressure, CCR type (Fig. 8 [8]). Catalyst improvements for a CCR unit will be targeted towards e.g. higher intrinsic selectivities, better regenerability, and high mechanical strength rather than towards the better stability desirable for a SR unit, and hence will follow a different set of rules: these are generally Pt/Sn catalysts. The above clearly illustrates the impact that a change in reaction
46
FUEL GAS
REGENERATED CATALYST
SPENT CATALYST
LIGHT ENDS TO RECOVERY
NET GAS
CHARGE PLATFORM ATE
Fig. 7. UOP Continuous plarforming process [5].
^^H
6h 0)
o
5
4
/ ^ 6 0 ' s & 70s
U
I
2\-
50S&60S
70 s & 8 0 s 1
1 10
1
15
1 20
1 25
1
30
35
PRESS (bars) Fig. 8. Increasing trends in the severity of naphtha reforming processes (Arrow indicates increasing severity of operation) [8].
conditions and reactor concept has by itself, and also how it affects the direction for catalyst developments. 3. HYDROTREATMENT AND HYDROCRACKING OF OIL FRACTIONS 3.1. Introduction As is well known, the first step in the processing of crude oil is the separation by distillation into a number of boiling fractions. The naphtha, with boiling point between 80 and 200°C is the basis for gasoline, and an important chemical feedstock; kerosine (bp between 150 and 250°C) is
47 used for aviation fuel, and gasoil (bp between 250 and 370°C) is used a.o. for diesel (automotive gasoil). All these fractions already have the right boiling point range and "only" have to be brought to specification, in which hydrotreating processes such as hydrodesulphurization and hydrogenation often play a major role. The crude oil also contains a significant amount of components boiling above about 370°C, the residue in the atmospheric distillation, and these can be converted into the valuable transportation fuels by (hydro)cracking, either by processing the entire residue, or (usually) by processing the vacuum gasoil, a fraction with boiling point between about 370 and 520°C that is recovered by vacuumflashingthe atmospheric residue. We will discuss here the reactor aspects of hydrotreating and hydrocracking by going to progressively heavier feedstocks, and then in the next section inventorize the various reactor issues encountered, and the common reasons behind those choices. 3.2. Hydrotreatment of light gases Light gases are produced in the refinery in various processes, such as catalytic cracking, and are often for used further conversion into e.g. gasoline fractions by e.g. alkylation. In most of these applications it is important to minimize the amount of acetylenes and diolefins by selective hydrogenation, processes for which dedicated catalysts have been developed. Thus one may estimate that the removal of the C4, C5 dienes from the alkylation feedstock yearly would save several millions of dollars in terms of reduced acid consumption in the alkylation unit. For cost-effectiveness, it would be very attractive to achieve this hydrogenation applying catalytic distillation technology [9]. In this operation the catalyst is placed inside the (existing) distillation column, requiring a sufficiently high reaction rate under the distillation conditions (obviously temperature, reactant concentration, depth of conversion, hydrogen partialpressure and reaction inhibition and selectivity are all relevant). If appUcable, however, the technology has intrinsic advantages in terms of reaction control, heat of reaction removal, removal of coke precursors and low capital costs. Indeed this technology, well-known from e.g. MTBE production [10], has now found a number of applications in selective hydrogenation of light gases (Fig. 9). In the commercial selective hydrogenation of C4, it is shown in Fig. 10 [11], that three catalytic
>-Distillate Product Hydrogen
•
Hydrocarbon feed
•
> • Bottoms Product Fig. 9. CD Hydro reaction column (ref. [11]).
48
Oligomer free Reflux
>-Offgas Catalyst Bed
Hydrogen Hydrocarbon Feed
^
Distillate Product
Fig. 10. CD Hydro Desulphurization and anti-fouling mechanism [11]. I 1,3-Butadiene
• Butene-1
50
"r 60
70
80
90
Packing Height (*/o of total bed) Fig. 11. Concentration profiles in CD Hydro Catalyst Bed [11].
functions are performed successively: first methyl mercaptan absorbs strongly and reacts with butadiene to form heavy sulphides going to the bottom of the column, then butadiene absorbs and is selectively hydrogenated and finally double bond isomerization occurs (see Fig. 11). Further appHcations of such technology may be expected.
49 Diolefins —> (Ni/S or Pd catalysts, 50-150°C) —> Monoolefins Monoolefins —> (NiMo/S catalysts, 250°C) —> Saturates Desulphurization —>(NiMo/S catalysts, 330°C) Diolefins —> (200°C or higher) —> Rapid fouling Fig. 12. Chemistry of pyrolysis gasoline desulphurization (schematic).
3.3. Hydrotreating of naphtha With the advent of the bimetaUic reforming catalysts, combined with the higher proportion of cracked feedstocks, the demands on naphtha desulphurization have only increased. Improved catalysts are available (see below), and while the technology itself has not changed much, the cracked feedstocks have favoured operation at higher pressures. New catalytic guard materials are available to remove silicon and iron [12] and protect the HDS catalyst; with the cracked feedstocks the recombination of H2S with olefins to mercaptans can limit the maximum operation temperature [13]. The hydroprocessing of pyrolysis gasoline from ethylene crackers requires a special set-up, since these contain large amounts of olefins and diolefins, defining the following chemistry (Fig. 12). The diolefins are especially nasty since they easily give severe fouling in an HDS unit. Therefore the pyrolysis gasoline is first processed at low temperatures over a selective hydrogenation catalyst using multiple catalyst beds to convert the diolefins into mono olefins, with product recycle of hydrogenated gasoil to limit the temperature rise over the reactor, to dilute the reactants and to dissolve formed polymers. In the next reactor the olefins are hydrogenated in a similar fashion, andfinallythe product is desulphurized (Fig. 13 [14]). 3.4. Hydroprocessing of middle distillates The specifications for automotive gasoil have become much more stringent recently, a trend which may very well continue; the sulphur specification in particular is now set to 500 ppm in many places. Apart from this, the Swedish and the Califomian authorities have promoted the introduction of a special grade of diesel, low in aromatics and sulphur, which is now also available in the UK. The basic chemistry of desulphurization, and aromatics hydrogenation in gasoils is shown in Fig. 14.. As is common in processing of oil fractions, in the ever deeper desulphurization of gasoil the most refractory sulphur species remain, which have been identified as the alkyl substituted dibenzthiophenes (e.g. ref. [15]). As a consequence, in the example quoted in ref. [16], a doubling of the reactor volume/halving of the feed rate would be required to go from e.g. 0.1 %S to 0.05 %S. Alternatively the reactor temperature could be increased by some 20°C, but this certainly reduces catalyst life and may also negatively affect product properties such as aromatics level and/or colour stability. The development of ever more active HDS catalysts has helped significantly, but the activity differences between various generations of Co/Mo desulphurization catalysts are about 15-20% (and quite difficult to realize) (Fig. 15), which is clearly by itself not enough to reach the new specification at otherwise unchanged conditions. A new option though apparently not yet commercial is the use of a dedicated ("alkyl-substituted DBT conversion") catalyst in the bottom part of this reactor [17,18]. Thus, at present, with many existing "single-stage" units, the severest sulphur specifications can only be achieved at a severe penalty to the catalyst runlength, especially if the other product properties also have
50 FIRST STAGE
SECOND STAGE
THIRD STAGE
^
LIGHT GASOLINE AROMATICS CONCENTRATE
^ HEAVY GASOLINE
GAS OIL HYDROGEN PYROLYSIS GASOLINE/GAS OIL
Fig. 13. Pyrolysis gasoline hydrotreating unit [14].
Sulphur compounds —> CoMo/S or NiMo/S catalysts —> Desulphurization -range of reactivities -alkylsubstituted DBT's most refractory compounds -inhibition by H2S Aromatics hydrogenation -equilibrium limited (partial pressure H2, 7) -Ni/Mo/S catalysts —> moderate activity, moderately poisoned -noble metals catalysts —> high activity —> strong poisoning by (organic!) S, N compounds Coke formation —> Deactivates CoMo/S and NiMo/S catalysts —> high T, low pp H2 Fig. 14. Chemistry of desulphurization and aromatics hydrogenation in gasoils.
to be met. Only single-stage units that operate at higher hydrogen pressures (favouring reaction rates, product properties at higher operating temperature and catalyst stability) offer significant scope. This holds for desulphurization, but even stronger for aromatics hydrogenation, where thermodynamics limits the use of the higher reaction temperatures. In this situation a process line-up in which a second reactor, with a fresh supply of hydrogen without H2S or NH3, is added, can be attractive ("two-stage" Hne-up). Even for conventional desulphurization catalysts hydrogen sulphide is a catalyst poison, due to competitive absorption, and the normal H2S partial pressures cause a significant loss in catalyst activity [19]. For noble metal catalysts, the detrimental effect of H2S (and NH3) of course is much stronger. A second-stage reactor, however, might employ dedicated, rather H2S sensitive catalysts to further reduce the sulphur levels [17,18].
51 IMPROVEMENTS IN CATALYST Rel.Vol.Activity
TECHNOLOGY Diesel Sulphur S p e c ^ ^ w t
200
150
100
2000
I Sulphur Spec
QRec.Voi.Act
Fig. 15. Activity of generations of HDS catalysts.
As was indicated above, the larger incentive for the two-stage unit is in aromatics hydrogenation using a noble metal catalyst in the second stage. The temperature in this stage must be low so that sufficiently high aromatics conversions are possible by thermodynamics, which puts special demands on catalyst activity. The favoured noble metal catalysts for this are those systems that have a relatively high tolerance for sulphur and nitrogen compounds in the feed; in this quite some progress has been made [16]. Theflowscheme of the integrated two-stage Shell Middle Distillate Hydrogenation process, running since 1992 in Gothenburg, Sweden, is shown in Fig. 16 [16]. Here the second-stage catalyst, at low pressure, can cope with (organic) S and N levels in thefirst-stageeffluent of some 100 and 50 ppm, respectively, and, on top of the aromatics hydrogenation, also achieves a desirable hydrodecyclization. The line-up implemented in the co-current Synsat process described in ref. [20] is similar, while another realization for this is shown in Fig. 17 [19]. An alternative line-up is based on the use of a counter-current second-stage reactor. In the normal co-current reactor the reaction rate always is highest near the top of the reactor (particularly for reactions of order 2 such as HDS) generating high concentrations of H2S already at the top of the reactor, negatively influencing the activity of the remainder of the catalyst (Fig. 18 [21]). In the counter-current line-up, this H2S (and/or NH3) is quickly removed from the reactor without affecting the activity of the bottom part of the catalyst. Hence, the counter-current operation potentially has the largest advantage when the catalyst is very sensitive to gases generated at the top of the catalyst bed, for noble metal catalysts therefore. The counter-current reactor was originally developed by Lummus in their Aerosat technology and later implemented together with Criterion in the Synsat unit in the Scanraff refinery in Sweden [22]. The detailed line-up is shown in Fig. 19, with the second-stage rector with "pure" hydrogen in counter-current operation. It is reported in ref. [22] that the counter-current operation allows a reduction in catalyst volume of 30-35% for reaching the same aromatics conversion relative to conventional all co-current operation (10-15% for a HDS target). Thus the SMDH and the Synsat processes are two options to produce low aromatics diesel oil; the paper by Grootjans [23] addresses the manufacture of this in more detail. (Ref. [16] gives a good listing of catalysts claimed for this duty.)
52 Fresh Gas rR-1 I XI
r*4m\^ FEED
^
<¥
M.
W
<S)
MBU
DEA
1^
FUEL GAS
CZZZ)
R-1 = HYDRODESULPHURISATION R-2 = HYDROGENATION
Fig. 16. Shell Middle Distillate Hydrogenation process [16].
Fig. 17. Reactor concept according to ref. [19].
PRODUCT
•53 100r
w t «»/o S in oil
°/oHDS
100
^'/o of catalyst bed
100
Fig. 18. H2S partial pressure profile in counter-current and co-current operation during hydrodesulphurization of gas oil [21].
The yield in the refinery of in-spec gasoil can be increased in some cases by a dewaxing step: in some cases the end point (boiUng point of the highest boiling components) of the gasoil is limited by the specification on coldflowproperties: crystallization of the wax molecules in the gasoil at low temperatures clearly is undesirable. This can be circumvented by application of catalytic dewaxing, mainly pioneered by Mobil, using shape selective zeolites such as ZSM-5 (MFI) [24]. In the original Mobil designs the dewaxing was carried out in a separate reactor, with a catalyst lifetime of about one month. More recent developments include the AKZO/Fina Cold Flow Improvement process, which achieves combined hydrotreating and dewaxing of hydrocarbons and can be used in existing units (possibly after a revamp to accommodate the cracked products) [25]. 3.5. Hydrocracking of vacuum gasoil The emphasis in the above sections was on improving the properties of the boiling fractions rather than changing their boiling range, treating rather than cracking. Since on average a significant fraction of the crude oil has a boiling point above about 400°C, it is important to produce transportation fuels from that fraction by cracking. This can be done in several ways, but in this section we limit ourselves to cracking of the vacuum distillate (boiling point roughly between 370 and 550°C) in the presence of hydrogen, hydrocracking. The basic underlying chemistry is shown in Fig. 20. The process is carried out in several configurations, but because the cracking catalyst is easily poisoned by organic N compounds, the first stage of a hydrocracldng unit always ensures denitrogenation rather than cracking, and generally uses dedicated (Ni/Mo) catalysts. The denitrogenated vacuum distillate feed is then cracked, either by processing the entire first-stage effluent ("series-flow"), or after removal of produced gas (H2S, NH3!) and light
wl P
I NTERSTAGE STRIPPER
9 REACT(
r HYDROCARBONS
A
L
ii
-
DIESEL PRODUCT
Fig. 19. Counter current SYNSAT line-up [ 2 2 ] .
FRESH
MAKE-UPWIS COMPRESSOR
ABSORBER
GAS
55 • Bifunctional catalysis, using a balanced (de)hydrogenation and an acidic catalytic function • Mechanism: dehydrogenation yielding reactive species (e.g. olefin from paraffin) that - adsorbs on the acid site - cracking of the reactive species - hydrogenation of the cracked species, and desorption • Poisoning of the acid function by (organic) nitrogen compounds necessitating denitrogenation • Hydrogen partial pressure reduces catalyst coking • Reactivity of feedstock components in decreasing order: aromatics -> naphthenes -^ paraffins • Zeolitic hydrocracking catalysts may exhibit shape/size exclusion effects Fig. 20. Catalytic chemistry of hydrocracking (schematic).
hydrocarbon products ("two-stage"); in both cases unconverted high boihng material can be (partially) recycled. According to the basic chemistry, the aromatics are converted first, followed by the naphthenes, while the paraffins are — not surprisingly — the least reactive. The unconverted product is therefore hydrogen rich, and may advantageously be used as FCC feedstock, ethylene cracker feedstock, and as lubricating base oil (partial conversion operation, or selective hydrocracking). The various process configurations are shown schematically in Fig. 21 [26]; the series-flow configuration is cheaper both in capital and in operating costs. The most modest approach is Mild hydrocracking, in which an existing unit, generally used for desulphurization of vacuum distillate, and hence rather low in hydrogen partial pressure and flow rate, is used to achieve a partial conversion in terms of cracking. Areas of attention are the quench capabilities and the product work-up section; at too high conversion catalyst Hfe will be short due to coke formation [27]. The same configuration, but then at the "proper" hydrogen pressure and flow, is used in single-stage or series-flow hydrocracking, possibly with recycle of unconverted product. Here the cracking catalyst has to work in the presence of the ammonia generated by denitrogenation, and hence at high temperatures. By contrast, in two-stage operation no ammonia is present and low operation temperatures, or less active catalysts can be used in the cracking stage. The process line-up chosen has important consequences for the catalyst development. In two-stage hydrocracking the catalyst activity is generally not a problem, and high conversions to naphtha can be achieved. The catalyst need not be resistant to ammonia or hydrogen sulphide, and for that reason also noble metal hydrocracking catalysts can be attractive. Indeed in general the hydrogenation activity will be limiting in this duty, and open to improvement as shown by Ward [28]. The low operation temperature also leads to more severe diffusional barriers than in series-flow operation, and a clear build-up effect of large, ring shaped molecules can be seen in recycle operation over zeolitic catalysts (Fig. 22 [29]). Since amorphous catalysts, with their larger pores do not exhibit this phenomenon, often composite catalysts are developed for this duty [29]. In series-flow operation, on the other hand, the intrinsic activity of the acidic function has been limited so much by the ammonia that this becomes important [28], and here often zeolitic catalysts are used. The product yield pattern has also shifted to middle distillates at the expense of naphtha as illustrated in Fig. 23, because the higher effective ratio of hydrogenation to cracking activity favours this, and also, because the lighter product are quickly removed from the reactor by evaporation (a form of catalytic distillation that has been used very specifically also in the wax hydrocracking step in Shell's middle distillate synthesis process [32]). Thus a modification of the process line-up (or conditions) is an alternative to a catalyst modification (as shown in Fig. 24 [31]) to influence selectivity. Combinations of series-flow and two-stage
56
(a)
FRESH GAS
:reO-
RECYCLE GAS
r
RECYCLE COMPRESSOR p ^
r^ 1st STAGE
t
2nd STAGE
HP SEP
bp<370" C
r\ FRACTIONATION
t
LP SEP I bp>370'C
^ A FEED
CXD
FRESH GAS
(b)
iree^
1st STAGE
t
f
RECYCLE GAS
r
RECYCLE COMPRESSOR p ^
2nd STAGE
HP SEP
A
FRACTIONATION
t
LP SEP
4
WJ FEED
bp<370'C
j bp>370'C
oo
Fig. 21. Process schemes for (a) series-flow, and (b) two-stage hydrocracking [26].
57
25
25
Carbon number
Carbon number
Fig. 22. FIMS spectra of the unconverted material after hydrocracking treated VGO over (a) all-amorphous, and (b) all-zeoUtic catalysts (two-stage operation), Z indicates the hydrocarbon stoichiometry, CnHm+z^ i-^. polynaphthenes occur for Z < 0, and aromatics are possible at Z < —4 [29].
58 8070-^ 60(D (1)
I/)
2000 ppm N
bO
(U
-M
OJ
— .
40
n TJ
30
•M
U)
2020
40
—T" 60
—r~ 80
100
Fig. 23. Effect of N on middle-distillate selectivity [30]. KEROSINE SELECTIVITY
U N I T CELL SIZE ( A )
Fig. 24. Kerosine selectivity versus zeolite unit cell size [31].
cracking are also being used, in three reactor systems in which the second reactor contains a cracking catalyst in series-flow operation; more recently stacked beds of a denitrogenation catalyst and a cracking catalyst in thefirststage have been introduced to achieve the same effect [29]. Here, as so often in hydrocracking the management of the exotherm reaction, with a high activation energy, puts demands on the quench systems. In hydrocracking units sometimes the products have to be slightly upgraded: for this purpose special catalysts are used at the bottom of the HCU unit to reduce the mercaptan formation by recombination, while in hydrocracking to produce lubricating base oils, a separate, low-temperaturefinishingstep is applied to hydrogenate aromatics [33]. A specific problem in hydrocracking units processing heavy feedstocks is the occurrence of polynuclear aromatics, that build up in recycle, and eventually precipitate on heat exchangers
59 (a)
L
S
L
G
L
G
L
(b)
L
G
POREUZE KATALYSATOR DEELTJES.^^
> I
i ^ ^
'^^P
1 / Jiff
t ^t ^
N
*
^
^
^?
^
^fi>
^
/ ^^^Sx
M M M^ G
TL
G
TL
G
TL
G
G
1 i1 G
Fig. 25. Three level porosity concepts for counter current operation [35].
("red death"). Several options exist to combat this phenomenon, one of which is a special absorption step to remove these species for subsequent disposal [34]. A clear current limitation of the counter-current reactor is the requirement of low pressure drop to avoid flooding of the reactor. To this end large shaped catalyst particles such as 5 mm Raschig rings could be envisaged, possibly also shell type catalysts. Sooner or later however the requirement of low pressure drop is conflicting with a high volumetric catalyst activity and a high catalyst utilization. Sie [35] therefore has proposed alternative solutions, based on the insight that in hydroprocessing the gas/liquid mass transfer is generally not rate limiting. In this way constructions are devised such as shown in Fig. 25 [35], in which the gas/liquid contacting is only periodically intensified. Such constructions, as yet still in the research phase, could allow higher gas rates and extension of the counter-current concept to areas such as hydrocracking, where one needs high gas rates, and also may wish to minimize the poisoning effect of the ammonia generated at the top of the catalyst bed by denitrogenation of the organic nitrogen compounds [21,35]. 3.6. Residue hydroprocessing Instead of recovering the vacuum distillate for conversion into transportation fuels, one can also process the entire residue of the atmospheric distillation, either to obtain a heavy low-sulphur fuel (desulphurization), or, again, to produce more transportation fuels (cracking, also called conversion in the oil industry). The atmospheric residue differs from the vacuum distillate in boiling point distribution (the vacuum residue, with components boiling above 520°C, is included), but even more importantly, in that it contains large amount of contaminant species, such as metals (in particular Ni and V) and asphaltenes (large aromatic compounds), which are very strong catalyst poisons. Furthermore, in the residues the concentration of sulphur and nitrogen containing molecules is higher than in the distillates, while on average these molecules have a lower reactivity. As a consequence of the basic chemical factors, outlined in Fig. 26, residue processing conditions have to be quite severe, high in temperature and in hydrogen pressure.
60 Sulphur compounds HDS —> CoMo/S or NiMo/S catalysts: low reactivity some diffusion effects Ni/V compounds HDM —> CoMo/S, NiMo/S or autocatalytic: low reactivity strong diffusion effects HDM: Ni and V deposition on thie catalyst, poisoning HDS activity. I interstitial deposition (would lead to) —> severe catalyst bed fouling residual compounds cracking —> mainly thermal, not catalytic residual compounds at high T —> severe catalyst coking and fouling hydrogenation of residues —> at a given hydrogenation level asphaltenes precipitate
oo
Fig. 26. Basic chemistry of hydroprocessing of residual oil fractions. PORE SIZE
WIDE
• ASPHALTENE PENETRATION HYDRO CONVERSION ACTIVITY METAL STORA6E
USE
1
TYPE
TOTAL VERY LOW VERY GOOD DEMETALLIZATION CATALYST A
MEDIUM
NARROW
MEDIUM
SHALLOW
MODERATELY HIGH
HIGH
FAIR
POOR
FRONT END TAIL END HYDROCONVERSION HYDROCONVERSION CATALYST CATALYST B
C
Fig. 27. ABC catalyst concept in residue HDS [37].
The reactor concept chosen strongly depends on the processing goal. The emphasis mostly has been on the desulphurization of residues to produce low-sulphur heavy fuel, or on the direct residue hydrocracking. For the desulphurization appHcation one desires a high degree of desulphurization, and therefore high catalyst activities and excellent staging. Fixed bed trickle bed reactors have been an obvious choice, and the technology has been developed by Chevron, Unocal, UOP, Exxon and Shell [36]. Catalyst deactivation was controlled by on the one hand high hydrogen pressures to reduce coke formation, and on the other hand by dedicated catalyst systems favouring demetallization in the front end, and highly stable desulphurization at the back end of the catalyst system (ABC concept. Fig. 27 [37]). In general the catalysts are contained in several large reactors; since the front end, demetallization, catalyst may deactivate more rapidly and the risk of interstitial deposition is to be minimized, sometimes small guard reactors are used that can be alternated, or put off-line; also the introduction of the feedstock at various points lower down the reactor with increasing runlength has been described [38], With the desire to increase cracking in thefixedbed units, by increasing temperature, also the demetallization rate increases
61 Fresh Catalyst Bin
Product to RDS Reactor
High Pressure Catalyst Vessel
OCR Reactor
Feed in
Spend Catalyst Bin Fig. 28. OCR catalyst replacement system [36].
and hence catalyst deactivation increases. The ultimate solutions, therefore, particularly for high metals feedstocks, are continuous HDM catalyst replacement systems, such as developed by Chevron (OCR, on line catalyst replacement, with a counter current operation, the feed moving upflow, Fig. 28 [36]) and by Shell HYCON (co-current downflow in the bunker flow reactors. Fig. 29a, b [37], processing even vacuum residue). They also have the advantage that the HDIVI catalyst leaving the system is fully loaded, contrary to the spent HDM catalyst from afixedbed reactor, where a metals profile will exist over the catalyst bed [36]. In residue hydroprocessing, more than in any other application, guard materials consisting of large, porous, low activity materials are used to catch Fe, scale and salt species before they plug the catalyst bed with its fine particles. Improved catalysts for the trickle bed units should have higher activity for HDM or HDS, higher stability (e.g. in terms of metals uptake capacity), or higher activity for cracking, e.g. by the use of zeolites (some approaches are discussed in refs. [12] and [39]). A recent example from our laboratory is given in Fig. 30 [12]. The main alternative is the use of expanded bed or ebullating bed reactors (see Fig. 31 [40]), which have a larger liquid hold-up, smaller catalyst concentration, with less risk of fouling, and with continuous catalyst replacement. A tight control of the inlet temperature however is critical [41]. The process has been developed in two versions, by ABB Lummus Crest ("LC-Fining") and by HRI ("H-Oil")- The catalyst bed is expanded by the liquid flow that is accomplished by pumping recycle liquid recovered near the top of the reactor to the plenum beneath the bed. These reactors are clearly intended and suited for the high-temperature cracking application rather than for deep desulphurization; important limitations are the lack of staging, including that of different catalysts, inherent to continuous stirred tank reactors. A further development of the concept has been implemented in the Texas City unit of Amoco, where three reactors
62 (a)
1
1
CTS
(CTS= CATALYST TRANSPORT
TO DISTILLATE SECTION
SYSTEM)
(b) FIXED BED REACTOR Fresh Catalyst
MOVING BED REACTOR (BUNKERFLOW)
Liquid > ' Gas
STATIONAR CATALYST BED
CONTINUOUS FLOW OF GAS AND LIQUID
Liquid Gas
CONTINUOUS FLOW OF GAS AND LIQUID
Liquid Gas Spend Catalyst
Fig. 29. Bunker flow system [37].
are used in series, improving staging and in principle allowing the use of optimized catalysts in each separate reactor (Fig. 32 [42]). Improved catalysts in the ebuUating bed units should allow higher conversions (cracldng) without leading to asphaltenesflocculation;some approaches are discussed in ref. [43]. Figure 33 [44] gives an impressive example of an improved catalyst that allows higher conversions without problems as to sediment formation; ref. [45] on the other
63 +50 (+90)
^4 •
^ +40 + ""^ (+72) "d) -30 b (-54) ?
CD
+20
^ (+36) ^^ +10 (+18) Base
RN -400^__^^^..*^—5—-rrW-
-L
^^<^ ^ 4
J^^^^t^k.
-7^^ ^-^-''^^^^^
RN-450
L^x^
1
1
500
1
1000
1500
1
2000
2500
Time on Stream, Hrs Fig. 30. Recent catalyst developments show significant improvement in residue desulphurization (ref. [12], fig. 20).
Catalyst addition Vapor/liquid i^'^'A separator
^Recycle pan
-Normal bed level
m
-Settled bed level
[——Distributor plate
Feed inlet
Recycle pump Catalyst withdrawal Fig. 31. Expanded bed reactor [40].
64 RESID HYDROPROCESSING HS
RHU PS VR
r-CKKh -CK>Q -{XH L-OCHlH
G •D A
GO
H V
C JZ
c
GO
T
AC Fig. 32. Expanded bed reactors in series [42].
0.2
BFDSHDIMENT.WT'/o
0.15 h
STANDARD CATALYST
0.1
0.05 h
NEW CATALYST
( TEX 2710)
_L_
50
55
60
65
70
75
80
85
538 C + CONVERSION,V<»/o Fig. 33. Sediment versus 538C+ conversion in ebuUating bed duty [44].
hand also quotes many examples of more active catalysts (including zeolite containing ones) without this low sediment feature, that are therefore not so suitable for this operation. A further step when going from the catalytic trickleflowreactor to the (catalytic) ebullating bed reactor is the Veba Combi Cracking process and related processes, where the cracking step is carried out at temperatures between 440 and 485°C, and high hydrogen pressures, without a catalyst, using only a finely ground antifouling agent. The product is then separated in a hot separator; the cracked product, which does not contain the worst catalyst poisons any more, is then hydrotreated at high temperatures and hydrogen pressures in fixed bed, while the unconverted product can be used for combustion, gasification or coking (Fig. 34 [46]) Clearly therefore also in residue hydroprocessing several options for the reactor type and conditions exist. Judging from the fact that both new fixed bed units (such as Shell's unit in Yokkaichi refinery) and new ebullating bed units (such as in Tula, IVLexico, and AGIP, Italy) are
65 Vacuum Residue
Hydrogenation Residue Fig. 34. Veba Combi Cracking system [46].
being built, they have about comparable merits. Again, these different technologies all put their own demands on the catalyst developments. 4. ANALYSIS OF THE VARIOUS REACTOR OPTIONS USED IN OIL HYDROPROCESSING From the preceding survey, the following generalizations can be made as to the various reactor types and modifications implemented or being studied to contribute to cost effectively meeting the current demands to oil refining and hydroprocessing. 1. Continuous catalyst replacement. These reactors, in various manifestations, are now used in catalytic reforming ("CCR"), in residue hydroprocessing ("H-Oil" or "LC-Fining" for ebullating beds, at high operation temperatures), and the OCR (Chevron) or HYCON bunker(Shell) for processing high-metals residues. These reactors allow relaxation of the targets for catalyst stability and can therefore use catalysts optimized with respect to activity or selectivity, or alternatively, process more difficult feedstocks economically. 2. Reactor configurations allowing processing at low levels of catalyst poisoning. With the increasing demands on conversion, it becomes more and more attractive to use a catalyst under conditions that poisoning is minimized. This is achieved by the conventional deep desulphurization of a naphtha prior to reforming, by mild diolefins removal in pyrolysis gasoline hydrotreating, or by the denitrogenation prior to hydrocracking, but recently this has been taken a step further. Thus in reforming the use of sulphur guardbeds increases, and even appears crucial for the application of the extremely sulphur sensitive zeolite based catalysts. In distillate hydroprocessing the use of a H2S scrubber of the recycle gas is a first step, but a two-stage system, in co-current operation has been introduced. To this end idle reactors, or specially designed types, can be applied. A further step is counter-current operation. At present this is still limited in terms of applicable gas rates, and hence to special applications, but research in progress in this area opens interesting perspectives to wider applications.
66 3. Reactors with selective product removal Reactor systems in which particular components are preferably removed fast, e.g. by distillation or by membrane transport have the potential to the influence the product selectivity, and these effects used consciously in hydrocracking. Catalytic distillation is rapidly gaining more applications also because it offers the possibility to better use existing equipment. The quoted example shows that also in hydroprocessing this technology is reaching a high level of sophistication. Obviously here clear limitations exist as to process conditions that are feasible, which may in turn be translated again to demands put on the catalysts for such potential applications. The above shows that a diversification from the simple downflow reactor has started, with many niche technologies. As stressed earlier, these bring significant added possibilities to meet the refining challenges of today, often without major investments, and hence form a valuable additional tool to the gains that can be achieved through development of improved catalysts for the conventional reactor systems. A strong interaction exists between catalyst improvements per se, and the reactor technology development, in that the novel technologies imply a different set of performance targets for improved catalysts as compared to improved catalysts for conventional technologies. Ultimately this again confirms our conviction that catalyst development is best done in an environment where process and reactor development are also being studied. Acknowledgement The author is indebted to Professor Dr. Ir. S.T. Sie for the many stimulating discussions, also on the present subject matter, over many years. References [1] [2] [3] [4] [5] [6] [7] [8] [9] [10] [11] [12] [13] [14] [15] [16] [17] [18] [19] [20] [21] [22]
M.D. Edgar, in: Applied Industrial Catalysis, Vol. 1, B.E. Leach, ed„ Academic Press, New York, 1983, p. 123. Catalytic Naphtha Reforming, G.J. Antos, A.M. Aitani and J.M. Parera, eds.. Marcel Dekker, New York, 1995. A.M. Aitani, in ref. [2], p. 409. J. Barbier, E. Churin, P. Marecot and J.C. Menezo, Appl. Catal. 36, 1988, 277. J.A. Weiszmann, UOP Platforming process, in: Handbook of Petroleum Refining Processes, R.A. Meyers, ed., McGraw-Hill, New York, 1986, p. 3-3. R.S. Haizmann, M.D. Moser, and D.Y. Wei, paper AM 95-30 presented at 1995 NPRA Meeting, March 19-21, 1995, San Francisco. P.W. Tamm, D.H. Mohr, and C.R. Wilson, in: Catalysis 1987, Studies in Surfaces Science and Catalysis Series, Vol. 38, J.W. Ward, ed., Elsevier, Amsterdam, 1987, p. 325. S. Sivasanker and P. Ratnasamy, in ref. [2], p. 483. J.L. DeGarmo, V.N. Parulekar and V. Pinjala, Chem. Eng. Progress, March 1992, p. 43-50. M.F. Doherty and G. Buzad, Trans. Inst. Chem. Engineers, part A, 70 (1992) p. 448-458. R. Barchas, R. Samarth, and G. Gildert, Fuels Reformulation, Vol. 3, no 5, September 1993, p. 44-52. D.C. McCulloch, Oil Gas J. Conference, Januuary 1996. D.C. McCulloch, in ref. [1], p. 69; H.J. Lovink, in ref. [2], p. 257. C.T. Adams, C.A. Trevino, in: ACS Symposium Series, 32, 1976, 412. E.g. A. Amorelli, YD. Amos, C.P. Haelsig, J.J. Kosman, R.J. Jonker, M. de Wind, and J. Vrieling, Hydrocarbon Processing, June 1992, p. 93-101. J.R van den Berg, J.R Lucien, G. Germaine, and G.L.B. Thielemans, Fuels Proc. Technol. 35, 1993, 119-136. I. Mochida, K. Sakanishi, X. Ma, S. Nagao, T. Isoda, Catalysis Today, 29, 1996, 185-189. Wo patent 9531518, to Exxon. European Patent application 0553920A1, to Shell. E.W Davies, A.J. Suchanek, and M.C. Baldassari, paper AM-94-62 at the 1994 NPRA Annual Meeting, March 20-22, San Antonio, Texas. R. Krishna and S.T. Sie, Chem. Eng. Science, 49 no 24A, 1994, p. 4029-4065. D. Dave, A. Gupta, K. Karlsson, A.J. Suchanek and H. van Stralen, paper AM-93-24 at the 1993 Spring Annual NPRA Meeting.
67 [23] [24] [25] [26] [27] [28]
[29] [30] [31] [32] [33] [34] [35] [36] [37] [38] [39] [40] [41] [42] [43]
[44] [45] [46]
J. Grootjans, this conference. N.Y. Chen, R.L. Gorring, H.R. Ireland and T.R. Stein, Oil Gas J. 75, 1977, June 6, 165-1701. H.W. Homan Free, T. Schockaert and J.W.M. Sonnemans, Fuel Proc. Technol. 35, 1993, p. 111-117. J.K. Minderhoud and J.A.R. van Veen, Fuel Process. Technol. 35, 1993, 87-110. J.W. Gosselink, A. van der Paverd and W.H.J. Stork, in: Catalysts in petroleum refining, D.L. Trimm, S. Akashah, M. Absi-Halabi, and A. Bishara, eds., Elsevier, Amsterdam, 1990, p. 385. J.W. Ward, Fuel Process. Technol. 35, 1993, 55-85; J.W. Ward, in: Preparation of catalysts ///, G. Poncelet, P. Grange, and P.A. Jacobs, eds., Elsevier, Amsterdam, 1983, p. 587. T. Huizinga, J.M.H. Teunissen, J.K. Minderhoud, J.A.R. van Veen, and H.J. Springett, paper at the AIChE Meeting, San Francisco, November 15, 1994. P.J. Nat, J.W.F.M. Schoonhoven, and F.L. Plantinga, in: Catalysts in Petroleum Refining, D.L. Trimm et al., eds., Elsevier, Amsterdam, 1989, P 399. I.E. Maxwell, T. Huizinga, A.A. Esener, A. Hoek, F. van der Meerakker, W.H.J. Stork, and O. Sy, Oil Gas J. April 22, 1991, 77. S.T. Sie, J. Filers and J.K. Minderhoud, in: Proc. 9th Congress on Catalysis, Vol. 2, Calgary 1988; M.J. Phillips and M. Teman, eds, Chem. Institute Canada, Ottawa, 1988, p. 743. J.F. le Page, Applied Heterogeneous Catalysis. Design, Manufacture, Use of Solid Catalysts, Eds Technip, Paris, 1987. N.M. Abdul Latif, in: Catalysts in petroleum refining, D.L Trimm, S. Akashah, M.Absi-Halabi, and A. Bishara, eds., Elsevier, Amsterdam, 1990, p. 349 S.T. Sie, NPT Procestechnologie, March 1995 p. 9; April 1995, p. 9. G.L. Scheuerman, D.R. Johnson, B.E. Reynolds, R.W. Bachtel, and R.S. Threlkel, Fuel Process. Technol. 35, 1993, 39-54. J.M. Oelderik, S.T. Sie and D. Bode, Applied Catal. 47, 1989, p. 1. OK patent 201405BCB. I.E. Maxwell and W.H.J. Stork, in: Introduction to zeolite science and practice, H. van Bekkum, E.M. Flanigan and J.C. Jansen, eds.. Studies in surface science and catalysis. Vol. 58, Elsevier, Amsterdam, 1991, p. 571 and references therein. R.M. Eccles, Fuel Process. Technol. 35, 1993, 21-38. WI. Beaton and R.J. Bertolacini, Catal. Rev. Sci. Eng. 33 (3&4), 1991, 281-317. J.F Mosby, R.D. Buttke, J.A. Cox, and C. Nikolaides, Chem. Eng. Science, 41, 1986, 989-995. J.D. Carruthers, J.S. Brinen, D.A. Komar, and S. Greenhouse, in: Catalytic Hydroprocessing of Petroleum and Distillates, M.C. Oballa and S.S. Shih, eds., Dekker, New York, 1994, p. 175-199; E.P Dai, D.E. Sherwood, and B.R. Martin, Chem. Eng. Science, 45, 1990, 2625-2631. G. Nongbri, G.A. Clausen, J.R. Huang, D.E. Shelf, C.A. Paul, and A.I. Rodarte, ibid, p. 55-69. E.R Dai and C.N. Campbell, ibid, p. 127-141. K. Niemann and F. Wenzel, Fuel Process. Technol. 35, 1993, 1-20
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® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
69
Simultaneous HDN/HDS of Model Compounds over Ni-Mo Sulfide Catalysts Liping Zhang and limit S. Ozkan Department of Chemical Engineering, The Ohio State University 140 West 19th Avenue, Columbus, OH 43210, USA
Abstract Reaction studies of model compounds are combined with pre- and postreaction characterization of the alumina-supported monometallic Ni, Mo and bimetallic Ni-Mo sulfide catalysts to acquire a better understanding of the catalytic phenomena involved in the hydrodenitrogenation catalysis over Ni-Mo system. The catalytic job distribution of Ni and Mo centers and the form of active sites for hydrogenation and hydrogenolysis steps in HDN reactions are examined. The effect of sulfur compounds on the catalytic sites and on the individual steps of the HDN network are discussed.
1. INTRODUCTION In recent years, hydrodenitrogenation (HDN) catalysis has been gaining importance due to growing environmental concerns and an increasing need for processing heavy oils and low-quality stocks which contain larger quantities of nitrogen compounds than the traditional petroleum stocks. Aluminasupported Co-Mo and Ni-Mo sulfide catalysts have long been used for HDS and HDN of refinery feedstocks. However, unlike the extensive in-situ characterization done for HDS reactions over Co-Mo/AlgOg catalysts, which h a s led to a fairly thorough understanding of the structure and catalytic mechanisms at the atomic and molecular level [1], the fundamental understanding of HDN catalysis over the Ni-Mo/AlaOg system still lags behind its industrial applications. It is partly due to the fact that HDN reactions a r e more complex and more difficult to catalyze than HDS [1-6]. HDN reactions usually consist of many intermediate aromatic ring hydrogenation and C-N bond cleavage steps which require a balanced dual functionality from the catalysts [2-4, 7, 8]. In addition, hydrogenation reactions are sometimes controlled by thermodynamic equilibrium and there is an intricate inter-play between the kinetically controlled regimes and the thermodynamically controlled regimes. Also, sulfur containing compounds in reaction feed have been found to affect the performance of Ni-Mo sulfides in HDN, with the effect
70 varjang significantly, depending on the rate-controlling step(s) [4, 7-10]. Therefore, more work is needed addressing both the structural and the kinetic aspects of the system to acquire a better understanding of the HDN catalysis. The general structure of Ni-Mo/y-AlgOg catalysts is thought to be the widely accepted "Ni-Mo-S" structure proposed by Tops0e and coworkers for HDS reactions [1, 11]. In the "Ni-Mo-S" structure. Mo exists in slabs of a ]V[oS2-like phase with Ni atoms decorating the edges of the IVEoSg platelets. Although this phase structure is well accepted, the exact form and characteristics of the active sites responsible for HDN reactions have not yet been fully determined. Until now, the suggested active sites and the catalyst deactivation models for HDN have not fully accounted for the role of Ni. Also requiring further study is the role of hydrogen sulfide or other sulfur-containing compounds on HDN reactions, including in-situ and post-reaction characterization of the catalysts and kinetic investigations. Polycyclic nitrogen compounds having both carbocycles and five- or sixmembered heterocycles are the major nitrogen-containing species in petroleum feedstocks. Indole and quinoline are typical model compounds representative of the five-membered and six-membered polycyclic nitrogen species, respectively. While five-membered nitrogen heterocycles account for about 2/3 of total nitrogen content in petroleum crude [1], majority of HDN studies have been conducted with quinoline [2-10, and references therein] and reports on indole are fewer and less extensive. In this study, bulk and y-Al203-supported Ni, Mo and bimetallic Ni-Mo catalysts with different ratios of Ni/Mo were prepared to perform individual and simultaneous hydrodenitrogenation and hydrodesulfurization of model compounds such as pyridine/thiophene and indole/benzothiophene. The role of HgS, which appears to affect HDN efficiency significantly in industrial practice, was investigated in some detail. Indole HDN reaction network was examined at temperatures ranging from 200 to 400°C, and pressures ranging from 100 and 1000 psig (14.7 psi = 1 atm), with a wide range of feed compositions of indole, H2S, benzothiophene, and o-ethylaniline. Pre- and post-reaction characterization of the catalysts were performed using various techniques, including X-ray photoelectron spectroscopy (XPS), temperature-programmed desorption (TPD), and temperature-programmed reduction (TPR), laserRaman spectroscopy, X-ray diffraction, SEM, and BET surface area. This paper will review the major results and conclusions from our reaction and characterization studies, some of which have been reported earlier [12-15], and present some of the new findings fi:*om the indole HDN studies.
2. EXPERIMENTAL Alumina-supported catalysts with different Ni and Mo loadings were prepared by wet co-impregnation of Y-AI2O3 with aqueous solutions of ammonium heptamolybdate and nickel nitrate. The preparation procedures and the characterization results of the oxide precursors were presented previously [12, 13]. The catalyst compositions are reported as weight percentages of the oxide precursors, i.e., M0O3, NiO, following the convention
71 commonly used in HDS literature. The BET surface area of the catalysts used in these studies varied between 165 and 195 mVg, with pure alumina giving the highest surface area at 195 vcfilg. The X-ray photoelectron spectrometers used for this work were Physical Electronics/Perkin Elmer (model 550) ESCA/Auger or V.G. ESCALAB Mark II, operated at 15kV and 20mA with Mg Ka radiation (1253.6 eV). Post-sulfidation and post-reaction analyses were done without exposure to the atmosphere, using a controUed-atmosphere chamber. Binding energies were referenced to Al 2p of 74.4 eV. SEM images were collected using a Phillips XL30 Scaning Electroscope. The reaction was carried out in a 4 mm I.D. stainless steel vapor-phase flow reactor. The total surface area of the catalyst in the reactor was kept constant at 50 m^ which gave a catalyst bed length of about 6 mm. Prior to all reaction studies, the catalysts were sulfided in situ at 400°C with 10% H2 in H2S for 10 hours. The feed and product streams were analyzed by an on-line HP 5890 A gas chromatograph with both thermal conductivity (TCD) and flame ionization (FID) detectors. All reaction data were taken after steady-state was established. For pyridine HDN studies, the reaction temperature was varied in the 320 to 400°C range. The pressure was maintained at 100 psig for all reactions. The hydrogen flow rate was at 70 cm3(STP)/min. The inlet thiophene and pyridine concentrations were kept constant at 0.87% each in H2 for all reactions. I n experiments which involved pyridine hydrodenitrogenation in the presence of H2S, the H2S concentration was also kept at 0.87%. The pyridine conversion levels were in the range of 0-50%. For indole HDN studies, reaction temperature was varied between 200-400°C. Pressures were 100 or 1000 psig. The Hg total flow rate was 30, 60, or 300 cm^(STP)/min. Three different reactant concentrations of 0.023, 0.046, and 0.069% were used for indole, benzothiophene, o-ethylaniline with balance being Hg. The mole ratio of HgS to H2 was varied in the range of 0 to 1.63x10"^. The products were identified by GC-MS analysis in conjunction with injection of pure compounds. TPD/TPR experiments were performed in an apparatus built in-house, as described previously [16]. All samples were sulfided in-situ using the s a m e procedure as in reaction studies. The reactor effluent composition w a s continuously monitored as a function of sample temperature by a m a s s spectrometer (Hewlett-Packard, MS Engine 5989A). 3. RESULTS AND DISCUSSIONS 3.1. Post-sulfidation and post-reaction characterization of catalysts X-ray photoelectron spectra of Ni 2p region showed the presence of NiAl204 over the sulfided monometallic Ni/y-AlgOg catalysts at a binding energy of 856.1 eV. The peak corresponding to this species, which is the result of Ni and y-alumina interaction, became very intense after individual and simultaneous thiophene HDS and p3n:"idine HDN reactions. The Ni 2p spectra for the sulfided bimetallic Ni-Mo/y-AlgOg catalysts, however, did not show the
72 presence of NiAl204 either before and after the HDS/HDN reactions, indicating that Ni in the bimetalhc catalysts exists in a different coordination environment than it does in the mono-metaUic Ni catalysts. On the other hand, pre- and post-reaction spectra of Mo 3d and S 2p for both Mo and Ni-Mo sulfide catalysts presented no detectable differences and showed the characteristics of M0S2 phase. These results suggest that the structure of the catalysts prepared in this work can best be represented in terms of the Ni-Mo-S phase structure proposed by Tops0e and coworkers. The intensity ratio of Mo/Al obtained from XPS measurements indicate that the addition of Ni enhanced the dispersion of Mo over the support and this enhancement was especially significant after sulfiding the catalysts. Post-reaction XPS measurements were also performed for catalysts with different on-stream time (up to 1000 h) for indole HDN/benzothiophene HDS reactions. It was found that the mole ratio of Ni/Mo decreased with aging. The loss of Ni on the surface of Ni-Mo catalyst after reaction was particularly signified by the change of Ni auger lines. In addition, the Ni spectra for the 3%NiO-15%Mo03/Y-Al203 catalyst after 1000 h of HDN/HDS reactions showed the presence of NiAl204, a feature observed over the monometallic Ni catalyst. SEM characterization of the aged catalyst also showed the presence of some Nirich crystals. These results suggest that more Ni sites were lost than Mo sites during the course of catalyst deactivation. There are two possible m e c h a n i s m s accounting for the loss of Ni active sites. First, Ni sites are more prone to being covered by carbon deposition since they are located on the edges of the M0S2 slabs. Deactivation studies in the literature have confirmed that coking, that spreads from the alumina support to the active phase, is one of the major causes of the deactivation of sulfide catalysts [1, 17]. Second, migration of NiMo-S slabs may force the segregation of some Ni atoms from the Ni-Mo-S structure, as evidenced by the presence of NiAl204. It is highly likely that both of these mechanisms may be in effect during the course of deactivation. 3^. HDN of pyiidine in the presence of thiophene and H2S Under our reaction conditions, the carbon containing products produced by HDN and HDS/HDN reactions were mainly C5 hydrocarbons and piperidine. Pentylamine, which was suggested as an intermediate in HDN reaction, w a s not detected. The results suggested a simplified, two-step reaction pathway, as proposed by Mcllvried [18]. In the first step, p3nddine is hydrogenated to form piperidine and in the second step, hydrogenolysis of piperidine through C-N bond cleavage takes place to form C5 and ammonia. This simple reaction scheme makes it possible to distinguish the sites responsible for hydrogenation and hydrogenolysis steps in HDN network. From pyridine HDN thermodynamic equilibrium data [19], we find that at the relatively low reaction pressure, the partial pressure ratio of piperidine to pyridine in the product stream is very close to the thermodynamic equilibrium ratio at temperatures above 350°C, indicating tliat this step is thermodynamically controlled. As the reaction temperature gets lower, the Ppipe/Ppyj. ratio gets further away from the thermodynamic equilibrium, indicating a kinetically controlled regime. The Pcs/Ppiper ^^^ Pcs^^pyr ratios in the product stream are far from equilibrium.
73 Figure 1 shows the effect of Ni and Mo loading on the two steps of pyridine reactions at 320°C where pyridine/piperidine thermodynamic equiUbrium is not estabUshed. The piperidine formation rate and the pyridine conversion rate are seen to increase with Ni loading level (Fig. la). Although it is not a sharp maximum there is a slow decline in these two rates with further increases in Ni loading above 5%. The C5 formation rate does not seem to be affected by the nickel loading. These observations suggest that Ni promotes the hydrogenation, but has no effect on the hydrogenolysis reaction. The fact that the promotional effect of Ni on the hydrogenation step goes through a maximum with increasing Ni loading can be explained in terms of the Ni-Mo-S phase structure. When the Ni loading exceeds the amount that corresponds to the available sites on the edges of M0S2 slabs, the excess Ni will form an inactive phase which, in turn, may block some of the Mo active centers. [1].
0.16
0.16
0 1 2 3 4 5 6 7 8
3
NiO, wt% (15% M0O3) Pyridine
- & - Piperidine
5
7
9
11 13 15 17
M0O3, wt% (3% NiO) Q C5
Figure 1. Effect of loading on pyridine conversion, piperidine and C5 production rates at 320°C.
74 0.08 fl o
S'^ -d •
0.06
P a p^ a 0.04
^. ;S o o a a
2 =^ 0.02 ;H
0
>
o
u
0 5
7
9
11 13 15 17
M0O3, wt% (3% NiO) Pyridine
NiO, wt% (15% M0O3) Piperidine
-a-C5
Figure 2. Effect of loading on pyridine conversion, piperidine and C5 production rates at 360°C. When the effect of Mo loading on hydrogenation and hydrogenolysis rates is examined under the same conditions (Figure lb), we see that Mo does not show the same enhancement effect on piperidine production rate as Ni does. Instead, increasing Mo loading increases C5 formation and pyridine conversion rates and decreases piperidine formation rate. The increase in C5 and p5n:'idine conversion rates with increasing Mo loading becomes more pronounced at 360°C as shown in Fig. 2a. On the other hand, the Ni loading has a negligible effect altogether (Fig. 2b). Since at 360°C, piperidine production rate is controlled by thermodynamic equilibrium, this observation reiterates the conclusion that Ni has no effect on hydrogenolysis steps, as it is already seen at 320 °C through the C5 yield which essentially remains constant regardless of Ni loading. The trends seen in these figures are significant in providing the first clues about the role of Ni and Mo-associated sites. It appears that under conditions where the first hydrogenation step is kinetically controlled, the addition of Ni increases the p3n:-idine conversion rate. When the first step approaches thermodynamic equilibrium, the addition of Ni no longer has a promotional effect, suggesting an assignment of the hydrogenation function to Ni-associated sites. When the role of Mo is considered, we see that it has an enhancement effect on C5 production rates at all times by promoting the piperidine hydrogenolysis reaction. The effect of increased Mo loading on the pyridine conversion rate is mainly an indirect one and is facilitated through the
75 consumption of piperidine in the hydrogenolysis reaction, which drives the first step fiirther to the right. The effect of sulfur compounds (HgS and thiophene) on pyridine HDN is found to be strongly dependent on several factors. When hydrogenolysis of piperidine is the rate determining step, pyridine HDN is enhanced by gas phase thiophene and H2S over both Mo and Ni-Mo catalysts. This enhancement effect is most pronounced over the bimetallic catalyst. Figure 3 gives an example of the effect of sulfur compounds on pyridine conversion and C5 formation rates over the monometallic and bimetallic catalysts at 400°C. In the absence of sulfur compounds, significant levels of piperidine were observed in the product stream, making the C5 production rate much lower than the pyridine conversion rate over both Mo and Ni-Mo catalysts. Under these conditions, the pjrridine/piperidine equilibrium is established and the rate-determining step is the hydrogenolysis of piperidine. When thiophene or H2S is present, however, piperidine is never detected in the reactor effluent, making C5 formation rate equals to pyridine conversion rate. An implication of this observation is that the rate determining step is no longer piperidine hydrogenolysis, but pyridine hydrogenation. The enhancement effect of sulfur compounds is facilitated through an increase in the hydrogenolysis sites. The increased HDN activity is much more evident over the bimetallic catalyst due to the strong promotional effect of nickel on the hydrogenation step. It can be seen from Figure 3 that, the enhancement effect of thiophene on pyridine HDN over the Ni-Mo catalyst is more pronounced than that of H2S. Since HgS is the product from thiophene HDS, it seems that thiophene has a n additional enhancement role t h a n HgS and this additional role is linked to Ni promoter.
Mo
S HDN only
H HDN with thiophene
Ni-Mo
HDN with H2S
Figure 3. Effect of sulfur compounds on the pyridine conversion and C5 production rates (400°C, catalysts: 10%NiO/Y-Al2O3, 20%MOO3/Y-A12O3, 3%NiO15%Mo03/Y-Al203)
76 3.3. Temperature-prograinined reduction and temperature-programmed desorption studies TPD experiments over the sulfided Ni-Mo/y-AlgOg and Mo/y-AlgOg catalysts showed two HgS desorption peaks. For both catalysts, the first peak corresponded to the desorption of weakly adsorbed H2S fi^om the a l u m i n a support. The second H2S desorption peaks had the same on-set temperatures of 300°C and maximum temperatures of 480 and 490°C for the 20% MoO^y-Alfi^ and 3%NiO-15%Mo03/Y-Al203 catalysts, respectively. In addition, the shapes of the second HgS TPD peaks for the two catalysts were very similar. The TPD profile of the bare alumina support did not show the high temperature feature. Based on these TPD results, it is conceivable that the second HgS peak represents the desorption of H2S from the catalytic active sites associated with Mo atoms only. Prior to the H2 TPR of the sulfided catalysts, a degassing treatment at 500°C was performed to remove all the adsorbed H2S species so that any HgS that evolves during the TPR experiment is a result of the reduction reaction and is not due to desorption of H2S which is left on the surface from the sulfidation process. The TPR profiles for bimetallic Ni-Mo/Y-Al203 catalysts were quite different from those of the monometallic Ni/y-Al203, and Mo/y-Al203 catalysts. However, bimetallic Ni-Mo/y-Al203 catalysts with different Ni loadings showed very similar patterns. The detailed results are presented elsewhere [20]. There were two low temperatiu-e H2S peaks at ca. 160°C and 220°C, which could be due to the removal of sulfur from Mo and Ni centers, respectively. ^The ratio of H2S peak areas for Mo- and Ni-associated peaks was slightly larger than 1 for the 3%NiO-15%Mo03/y-Al203 catalyst and about 1 for the 5%NiO-15%Mo03/y-Al203 and 7%NiO-15%Mo03/y-Al203 catalysts. It is conceivable that these two peaks correspond to the formation of sulfur vacancies on the edge planes of Ni-Mo-S phase. The 1:1 ratio for the Mo- and Ni- associated S vacancies for catalysts with high Ni loadings is in agreement with the maximum Ni accessibility to the M0S2 plane that is concluded in the literature [1]. 3.4. Active sites and their catalytic functions Combining the results of our kinetic and characterization studies with some findings in the literature, we propose two major types of active sites promoting HDN of nitrogen heterocycles. Type I: these are hydrogenation sites consisting of sulfur vacancies associated with Mo (type la sites) or Ni in Ni-Mo-S phase (type lb sites). T5rpe II: these are hydrogenolysis sites consisting of Bronsted acid centers associated with Mo atoms only. According to Yang and Satterfield [91, the adsorption and dissociation of an H2S molecule can convert a sulfur vacancy to a Bronsted acid site and a sulfliydryl group (SH), but the adsorption is readily reversible if HgS is removed fi^om the reaction system. The results from our pyridine HDN studies can be explained in terms of these active site assignments. The catalytic job distribution of Ni and Mo associated centers in pyridine HDN can be summarized as Ni-associated sulfur vacancies in the Ni-Mo-S phase (type lb sites) being responsible for hydrogenation steps, whereas the primary function of Mo being to promote C-N bond hydrogenolysis reactions through Bronsted acid sites. For Ni-Mo
77 catalysts, Mo-associated S vacancies are not important for hydrogenation reactions due to their much lower intrinsic activity compared to that of Niassociated ones. The role of sulfur compounds in the pyridine HDN catalytic scheme is envisioned to be multi-faceted: 1) HgS in the gas phase helps maintain a certain content of Bronsted acid sites (type II sites). Thiophene promotes hydrogenolysis of piperidine indirectly via HgS formed during HDS reaction. 2) Thiophene in the gas phase also helps the pyridine hydrogenation step over the bimetallic Ni-Mo catalysts by keeping Ni active sites in an effective form for pyridine hydrogenation reaction, probably by converting double S vacancies to single vacancies. 3) Thiophene in the gas phase does not enhance the hydrogenation step over the mono-metallic Mo catalyst, but inhibits it by reducing the number of available hydrogenation sites. 3.4. HDN of indole in the presence of H^S, benzothiophene, and o-ethylaniline The reaction network of indole HDN based on the proposals from the literature [6, 21-29] and this work is depicted in Figure 4, including the acronyms used for the compounds discussed. There is general consensus that the reaction network in indole HDN starts out with hydrogenation of the heterocyclic ring in a reversible step which leads to indoline formation, dictated by thermodynamic equilibrium under most conditions. Several of the previous reports on indole HDN suggest that o-ethylaniline is the exclusive intermediate toward the formation of hydrocarbons [21-27] following indoline formation. One of the more recent studies on indole HDN which was conducted over a sulfided NiMoP/Y-Al203 catalyst [28] proposed the denitrogenation pattern for indole to be analogous to that found for quinoline [2, 3], i.e., the hydrocarbon products being formed predominantly by complete hydrogenation of both the heterocyclic and the benzene rings prior to the cleavage of C-N bond. There are also suggestions that both routes could be playing an important role in the overall network [6,29]. Indole
Indoline
OEA
r OHI
I OECHA " N
EB
t JLXV.
ECH Figure 4. Indole HDN reaction network.
78 A key question regarding the reaction network of indole HDN is the role of two N-containing intermediates, o-ethylcyclohexylamine (OECHA) and octahydroindole (OHI), since they were not detected in many of the studies reported. Another question that has not been addressed very much is the effect of sulfur compounds. This study was designed to determine the important intermediates and to differentiate among the individual steps involved in the indole reaction network in the presence and the absence of sulfur compounds. Benzothiophene was chosen as the model sulfur-compound in this study. Ethylbenzene (EB) was the only major product in the HDS of benzothiophene (BT) over the Ni-Mo catalyst. The hydrocarbon products from BT HDS were about 99% of EB and 1% of ethylcyclohexane (ECH) throughout the temperature range from 200 to 400°C at 100 psig. In contrast to benzothiophene HDS, ECH was always a major product for indole HDN over Ni-Mo catalyst, implying that the hydrogenation of benzene ring is important in the coiu^se of indole HDN. By selecting reaction parameters which favor hydrogenation reactions, OHI and both cis- and trans-OECHA were observed in our reaction experiments. Since the hydrogenation of EB to ECH does not seem to be important under our reaction conditions, as evidenced by the BT HDS results, the hydrogenation of benzene ring should occur prior to the C-N bond cleavage. This, in turn, suggests that OECHA is the intermediate in ECH formation. According to the reaction network proposed, there are two possible routes for OECHA formation, i.e., the hydrogenation of OEA and the hydrogenolysis of OHI. It has been reported in the literature that, in a mixture with other nitrogen-containing compounds, aniline-tjrpe molecules are least reactive [3, 28, 30, 31]. To determine the reactivity of OEA in an indole HDN reaction mixture, we performed a set of experiments keeping the total concentration of nitrogen compounds constant, but replacing one half of indole with OEA. The results from these experiments are summarized in Table 1. The percentage of OEA in the product stream resulting from a co-feed of indole and OEA, although somewhat higher than that resulting from a OEA-free feed, is less than 10 % of the OEA concentration which was in the feed. This result clearly shows that OEA remains highly reactive even in the presence of indole/indoline. It should also be noted that the ECH/EB ratio decreases when half of indole in the feed is replaced by the OEA intermediate. If the conversion of indoline is through the OEA route only, same ECH/EB ratio would be expected in this case. The lower ECH/EB ratio from indole + OEA HDN suggests that both OEA and OHI routes are important for the conversion of indoline. Table 1 HDN of indole and indole+OEA mixture over the Ni-Mo catalyst at 320°C and 100 psig
Feed indole alone indole + OEA
indole mol% inlet outlet inlet outlet
0.046 0.023 0.023 0.017
OEA mol% 0 0.0016 0.023 0.0022
ECH/EB 2.03 1.72
79 It has been reported in the hterature that hydrogenation reactions are enhanced by increasing H2 pressure whereas hydrogenolysis reactions are not very sensitive to Hg pressure [32]. Based on this observation, one would expect that increasing H2 pressure will increase the importance of OHI route in indole HDN reactions and will result in larger OHI and OECHA presence in the reactor effluent if hydrogenolysis reaction rates are relatively lower. The experiments we performed at 1000 psig for indole HDN in the absence of sulfur compounds over the Ni-Mo catalyst showed that it was indeed the case. Figure 4 compares the selectivity of some major products from indole HDN at 100 and 1000 psig H2 pressure. At low pressure, only trace amounts of OECHA and OHI were present, whereas at high pressure, OECHA had the highest selectivity among N-containing species and the relative amount of OHI was also increased. Also at high pressure, OEA, an intermediate from indoline hydrogenolysis, had the lowest selectivity among all N compounds, whereas it had the highest selectivity at low pressure. When hydrocarbon product selectivities are compared, the most pronounced difference is that ethylbenzene selectivity dropped while ECH selectivity increased when H2 pressure increased was from 100 to 1000 psig. This result suggests that due to a much stronger hydrogenation function of the catalyst at high pressure, the route via OEA in indole HDN becomes less important with increasing H2 pressure in the absence of sulfur compounds.
H
0^2 r
Cr 10
20
30
Selectivity %
0
10
20
30
40
50
Selectivity %
Figure 5. Comparison of indole HDN product selectivities at 100 and 1000 psig. (3%NiO-15%Mo03 catalyst, 320°C, feed: 0.046% indole in hydrogen)
80 Two routes are proposed for the ECH formation from OECHA HDN as shown in the network. The route via p-ehmination (OECHA -> ECHE ) and hydrogenation (ECHE -^ ECH) has been well estabhshed [4-6, 33]. The route without ECHE as an intermediate could be envisioned as a hydrogenolysis reaction or, more accurately, as a Hofmann degradation reaction [6, 33, 34] in which the -NHg group of OECHA is replaced by a -SH group whicli goes through hydrogenolysis very easily and rapidly. The results obtained from indole HDN by varying the HgS-to-Hg ratio in the feed from 0 to 1.63 xlO'^ u n d e r ICKX) psig over the bimetallic catalyst are presented in Figure 6. As shown in the figure, with increasing H2S-to-H2 ratio, the indole conversion and ECH production increased steadily up to a H2S/H2 ratio of 0.3x10'^. The ECHE and EB production rates, on the other hand, showed very Uttle change with HgS/Hg ratio. OECHA production rate decreased rapidly with increasing HgS concentration (Figure 6B) suggesting a direct correspondence between the ECH and OECHA production rates. These results suggest the presence of a second OECHA -» ECH route, which does not go through the ECHE intermediate.
0.4 0.8 1.2 1.6 0 0.4 0.8 1.2 l.( Mole Ratio, H2S/H2 xlOO Mole Ratio, H2S/H2 xlOO Figure 6. Effect of H2S-to-H2 ratio on indole HDN over the 3%NiO-15%]V[o03 catalyst (320°C, 1000 psig , feed: 0.046% indole in 300 cm^(STP)/min hydrogen). As shown in Figure 6B, there was a sharp decrease of indoline, OHI, and OECHA production rates with increasing HgS concentration, especially at lower H2S/H2 ratios. Since C-N bond cleavage is involved in the further reaction of each of these species, this observation indicates that the higher activity in indole HDN with increasing H2S/H2 ratios is directly linked to the enhancement effect of HgS on C-N bond cleavage (hydrogenolysis) reactions.
81 The appearance of significant amounts of these intermediate species at h i g h hydrogen pressures in the absence of sulfur compounds (Figure 5) seems to suggest that, at higher pressures, the hydrogenation function of the catalyst is relatively strong and more HgS is needed to balance the hydrogenation a n d hydrogenolysis functions of the Ni-Mo sulfide HDN catalysts. The necessity of replenishing the sulfur sites with HgS is especially relevant to industrial practice since high hydrogen partial pressures tend to deplete the sulfur on the catalyst more readily, leading to a greater loss of hydrogenolysis activity. 4. SUMMARY/CONCLUSIONS Combining the results of our kinetic and characterization studies with some findings in literature, a catalytic job distribution of Ni and Mo centers in HDN reactions is proposed. It is envisioned that, in the HDN of nitrogen heterocycles, Ni-associated sulfur vacancies in the Ni-Mo-S phase a r e responsible for hydrogenation steps, whereas the primary function of Mo is to promote C-N bond hydrogenolysis reactions through Bronsted acid sites, which are generated fi*om the adsorption and dissociation of H2S on Moassociated sulfur vacancies. The conversion of Mo-associated S vacancies to Bronsted acid sites by HgS is a reversible process strongly depending on temperature and HgS-to-Hg ratios. The reaction network of indole HDN is a complex one, consisting of multiple steps of hydrogenation, hydrogenolysis, P-elimination, Hofmann degradation, and dehydrogenation. The reactivity can not be characterized in terms of a single rate-determining step, but is mainly controlled by the relative strengths of the hydrogenation and hydrogenolysis functions of the catalyst. As for the effect of sulfur compounds on the HDN activity, a "universal" effect which accounts for all conditions can not be defined. This effect varies significantly depending on whether the hydrogenation of the heterocyclic-ring is kinetically or thermodynamically controlled and which steps and which sites are dominant in the overall catalytic scheme. At h i g h temperatures and high Hg pressures, more HgS is needed to maintain the proper balance between the hydrogenation and the hydrogenolysis functions of the catalyst. Acknowledgment The financial support provided for this work by the National Science Foundation through the Grant HRD-9023778 is gratefully acknowledged. REIFERENCES 1. 2.
Tops0e, H., Clausen, B.S., and Massoth, F.E., Hydrotreating Catalysis, Springer-Verlag, Berling,1996. C.N. Satterfield, Heterogeneous Catalysis in Industrial Practice, 2nd ed., McGraw-Hill, 1991, p383.
82 3. 4. 5. 6. 7. 8. 9. 10. 11. 12. 13. 14. 15. 16. 17. 18. 19. 20. 21 22. 23. 24. 25. 26. 27. 28. 29. 30. 31. 32. 33. 34.
G. Perot, Catal. Today, 10 (1991) 447. M.J. Girgis and B.C. Gates, Ind. Eng. Chem. Res., 30 (1991) 2021. T.C. Ho, Catal. Rev.-Sci. Eng., 30 (1988) 117. R. Prins, in Knozinger, Ertl and Weitkamp (eds.). Encyclopedia of Catalysis, to be published. C.N. Satterfield, M. Modell, and J.A. Wilkens, Ind. Eng. Chem. Process Des. Develop., 19 (1980) 154. C.N. Satterfield and S. Gultekin, Ind. Eng. Chem. Process Des. Develop., 20 (1981) 62. S.H. Yang and C.N. Satterfield, J. Catal., 81 (1983) 168. S.H. Yang and C.N. Satterfield, Ind. Eng. Chem. Process Des. Develop. 23 (1984)20. H. Tops0e and B.S. Clausen, Catal. Rev.-Sci. Eng. 26 (1984) 395. U.S. Ozkan, S. Ni, L. Zhang, and E. Moctezuma, Energy Fuels, 8 (1994) 249. U.S. Ozkan, L. Zhang, S. Ni and E. Moctezuma, J. Catal., 148 (1994) 181. U.S. Ozkan, L. Zhang, S. Ni and E. Moctezuma, Energy Fuels, 8 (1994) 830. L. Zhang and U.S. Ozkan, in J.W. Hightower et al. eds., 11th Int. Cong. Catal - 40th Anniversary, 1996, 1223. U.S. Ozkan, Y. Cai, M.W. Kumthekar, and L. Zhang, J. Catal., 142 (1993) 182. J.G. Weissman and J.C. Edwards, Appl. Catal. AiGeneral 142 (1996) 289. H. G. Mcllvried, Ind. Eng. Chem. Process Des. Develop., 10 (1971) 125. C.N. Satterfield and J.F. Cocchetto, AIChE J., 21 (1975) 1107. L. Zhang and U.S. Ozkan, to be summited to J.Catal. L.D. RoUmann, J. Catal. 46 (1977) 243. E.W. Stern, J. Catal., 57 (1979) 390. E.O. Odebunmi and D.F. OUis, J. Catal., 80 (1983) 76. J.-L. Olive, S. Biyoko, C. Moulinas and P. Geneste, Appl. Catal., 19 (1985) 165. P. Zeuthen, P. Stolze and U.B. Pedersen, Bull. Soc. Chim. Belg., 96 (1987) 985. J. Shabtai, G. Que, K. Balusami, N.K. Nag and F.E. Massoth, J. Catal., 113 (1988) 206. F.E. Massoth, K. Balusami and J. Shabtai, J. Catal., 122 (1990) 256. M. Callant, P. Grange, K.A. Holder, and B. Delmon, Bull. Soc. Chim. Belg., 104(1995)245. M.V. Bhnde, Ph.D. Dissertation, University of Delawere, Newark, 1979. S. Kasztelan, T. des Courieres, and M. Breysse, Catal. Today, 10 (1991) 433. S.-J. Liaw, A. Raje, K.V.R. Chary, and B.H. Appl. Catal. A:General 123 (1995)251. S.R. Shih, J.R. Katzer, H. Kwart, and A.B. Stiles, Am. Chem. Soc. Div. Petrol. Chem. Prepr., 22 (1977) 919. N. Nelson and R.B. Levy, J. Catal. 58 (1979) 485. J.L. Portefaix, M. Cattenot, M. Guerriche, J. ThivoUe-Cazat and M. Breysse, Catal. Today, 10 (1991) 473.
® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
83
Kinetics of the catalytic removal of the sulphur components from the light cycle oil of a catalytic cracking unit G.F. Froment, G.A. Depauw and V. Vanrysselberghe Laboratorium voor Petrochemische Techniek, Rijksuniversiteit Gent, Krijgslaan 281, B-9000 Gent, Belgium ABSTRACT First order kinetics for the hydrodesulphurization of 40 sulphur components of the benzothiophene and dibenzothiophene families were determined using experimental data obtained with a light cycle oil. Substituents next to the sulphur atom reduce the reactivity with respect to the nonsubstituted component. Hougen Watson rate equations for substituted dibenzothiophenes were related to those for the head of the family, dibenzothiophene, through global multiplication factors f„ s^ny ^"^ frsDUT- These factors are the products of the electronic and steric effects of the substituents on the adsorption equilibrium constants and of the electronic effects on the rate coefficients. They were evaluated for a number of substituted dibenzothiophenes. 1. INTRODUCTION Light cycle oil (LCO) contains various aromatic sulphur components such as benzothiophenes (BTs), dibenzothiophenes (DBTs) and naphthothiophenes [1]. These components have to be removed for both technical and environmental reasons. Hence the sulphur content of diesel is limited to 0.05wt% from October 1996 onwards. Detailed kinetic equations for the sulphur components are required for the modeling of the trickle bed reactors used for hydrodesulphurization (HDS). LCO is an ideal feedstock for the determination of the kinetic data since all interesting refractory sulphur components are present in it and since the content of aliphatic sulphur components as well as the length of the alkyl side chains are limited, thus facilitating the quantification. A kinetic modeling of the hydrodesulphurization reactions based upon structural contributions was developed to limit the number of parameters [2]. The rates for reactions involving substituted components were related to those of a non substituted reference component in terms of the effect of the substituents on the rate coefficients and the adsorption equilibrium constants. Based upon a number of assumptions, multiplying factors were introduced for the electronic and steric effect of the substituents on the adsorption and for the electronic effect on the reaction rate. In that way the number of parameters for the HDS of a set of mono-, di- and trimethyl substituted DBTs was reduced from 1133 to 93. According to the rate equations derived for the HDS of DBT [3] and assuming that the rate expressions
84 for the HDS of the methyl substituted DBTs are of an identical form, this number can be further reduced to 35. A limited set of experiments with complex mixtures, such as LCO, can be used to obtain the global multiplication factors for various components. These global factors are the products of the effects of the substituents on both the rate coefficients and the adsorption equilibrium constants. The determination of the structural contributions as defined by Froment et al. [2] requires additional experimental data obtained from model components such as 4-methyldibenzothiophene (4-MeDBT), 4,6-dimethyldibenzothiophene (4,6-DiMeDBT) and one of the trimethyldibenzothiophenes. 2. EXPERIMENTAL SET-UP The liquid hydrocarbon was fed into the reactor with a high-pressure pump. The hydrogen, the hydrogen sulphide and the nitrogen feed were controlled and metered with a set of electronic mass flow controllers. Hydrogen sulphide and nitrogen were used in the pretreatment of the catalyst. The gases and the liquid feed were preheated and mixed before entering the reactor. The effluent section was also heated to avoid condensation. The reaction was carried out in a multiphase Robinson-Mahoney reactor. The temperature was measured by means of thermocouples and controlled by a PID temperature controller. The pressure was controlled by a back pressure regulator. The effluent of the reactor consisted of gas and liquid phases at high pressure and high temperature. Both phases were separated by means of a cyclone. The liquid was collected in the liquid holder. The cyclone and the liquid holder were kept at the same pressure and temperature as in the reactor, to avoid changes in composition of both phases. The gas phase was cooled, so as to condense heavy fractions, and was then scrubbed by means of a sodium hydroxide solution to remove hydrogen sulphide before venting. The liquid product was cooled and flashed under ambient conditions. The light gases, dissolved in the liquid phase, were partially desorbed and collected in a gas burette. Product samples were taken off-line for GC analysis. 3. ANALYSIS The GC-AED system (Hewlett-Packard 5921 A) was used for the quantification of sulphur components and the determination of the total sulphur content of the LCO mixture. The atomic emission detector (AED) is element specific and has a high dynamic range for C, H, N and S. The total sulphur content of the LCO (1.28wt%) was determined with the AED using hydrogen linearization. Detailed identification and quantification results of the sulphur components in LCO are presented in a separate paper by Depauw and Froment [1]. A quadrupole mass spectrometer, part of the Hewlett-Packard 5989A system was used for quantification as well. Electron impact ionisation mass spectra (m/z 40-400) were obtained at 70eV at a rate of 1.6 scans/second. The quadrupole temperature was lOO^'C, the ion source temperature 250''C and the transfer line temperature 250''C. The GC separation was performed with a Hewlett-Packard 5890 Series 11 instrument on a 50 m x 0.2 mm fused silica capillary column coated with a 0.5 ^m film of cross linked 100% dimethylsiloxane (HPPONA). Helium was used as carrier gas (0.645 ml/min at 35''C). The column was
85 temperature programmed from 35°C(5 min) to SCCCIS min) at a rate of 2.5°C/min and further to 200°C(5 min) at a rate of 2.0°C/min and finally to 250X at a rate of TC/min. The injector temperature was 250°C. One microliter of undiluted sample was injected at a split ratio of 63:1. The conversion of a molecule can be calculated by comparing the feed and effluent ion chromatograms. The effluent chromatograms are scaled using fluorene as an internal standard. This molecule is present in the LCO feed and is not produced nor hydrogenated or evaporated significantly under reaction conditions. It does not co-elute with other components with the same mass. The most important fragment of fiuorene with a m/z ratio of 166 was used for the scaling. An ion chromatogram (m/z =166) of an effluent is multiplied with a scaling factor in order to get the same surface for the fluorene peak as was obtained for this peak in the feed analysis. All other ion chromatograms of the effluent are multiplied with the same factor. The surfaces of the peaks in the effluent and the feed ion chromatograms are then linearly proportional in an identical way to the molar flows at the reactor exit and entrance respectively. The conversions of all components in the feed are derived from the relative decrease of their fragments. In case of co-elution of several molecules an ion fragment which is typical for the molecule considered has to be used for the calculations of the conversion. 4. EXPERIMENTAL PROGRAM Experiments were performed with a light cycle oil (LCO) of a catalytic cracking unit containing 1.28wt% sulphur and with a boiling range of 225-370''C. Experiments were carried out at temperatures between 240 and 320°C. The liquid pumping rate was varied between 15 and 27 ml/hr. The total pressure was 80 bar, and the hydrogen to hydrocarbon ratio 344 Nl/1. The molar hydrogen to methane ratio was 6.4. The number of experiments amounted to 25. The catalyst used was the commercial HDS catalyst AKZO Ketjenfine 742. It was crushed to a size between 710 and 800jitm to avoid diffusional limitations and 2.53 gcat were diluted with nonporous inert alumina. The absence of diffusional limitations was calculated using the Weisz-Prater criterion. The reaction mechanism and the intrinsic kinetic equations for the hydrodesulphurization of dibenzothiophene were derived by Vanrysselberghe and Froment [3]. In order to determine the kinetics for the hydrogenation of naphthalene into tetralin, experiments were carried out using a solution of 2wt% dibenzothiophene and l-5wt% naphthalene in a paraffinic mixture (Cjo-CiJ. The temperature was varied between 240 and 300''C, the molar hydrogen to hydrocarbon ratio between 1.10 and 1.36. The molar feed flow rate of dibenzothiophene was varied between 1.23 10'^ and 3.94 10'^ kmol/hr, the naphthalene molar feed flow rate between 1.71 10'^ and 1.42 lO'** kmol/hr. The total pressure was 80 bar.
86 5. RESULTS AND DISCUSSION 5.1. First order kinetics In order to compare the kinetic data obtained in this work with literature data and to provide quantitative results on the reactivity, first order rate coefficients were determined from the experimental conversions. The conversion X; of component i is defined as: pj. in
X.=
T;
^
out
(1)
'
The conversion Xj of component i is directly obtained from the experimental data. The continuity equation for this component in the perfectly mixed reactor is given by: F/"-r.W=Fr
(2)
Substitution of (2) into (1) results in: r.W
with
Fi'"=V'"Ci'"
(3)
Using the first order approach, the rate of removal is assumed to have the form: r=k.C. I I
(4)
The ratio of the concentration of i at reactor conditions and in the inlet CJC-^ can be written as (l-Xi)pV"'/p"'V°"^ The rate coefficient is finally written as a function of measured quantities: k.=—L-l-^ '
1-Xj W
p
(5)
The activation energy and preexponential factor are determined from the Arrhenius relationship: krAiexp(-JfL)
(6)
87 Results for the BT-family The results for 27 sulphur components of the BT-family are presented in Table 1. The conversions used for the determination of these properties were obtained with the GC/MS system, which shows less peak overlapping than the GC-AED analysis. The Arrhenius plots for a set of substituted benzothiophenes are shown in Figure 1. The correlation coefficients R^, the ratio of the regression sum of squares to the total sum of squares, varied from 0.986 for 7-EtBT to 0.999 for 2567-TeMeBT. Table 1 Preexponential factors, activation energies and first order rate coefficients at T=280°C for various members of the BT-family Sulphur component
A, [mVkg,,yh]
E,i [kJ/kmol]
k, [mVkg,,,/h] atT=280°C
BT 54-6-MeBT 3+4-MeBT 7-MeBT 27-DiMeBT 23-DiMeBT 2-EtBT 7-EtBT 24-DiMeBT 56-DiMeBT 45-DiMeBT 34-DiMeBT 35-h36-DiMeBT 257-TriMeBT 357-TriMeBT 267-TriMeBT 356-TriMeBT 234-TriMeBT 235+236-TriMeBT 7-PrBT 2357-TeMeBT 2367-TeMeBT 2567-TeMeBT
6.537 10^^' 4.253 10' 8.687 10' 2.025 10' 2.717 10^' 5.511 lO''* 2.173 10' 2.394 10^« 4.412 10'" 7.423 10' 9.943 10' 6.025 10*^ 3.717 10^2 5.102 10'" 1.212 10^" 4.371 10'^ 1.528 10" 9.660 10'' 1.816 10'2 8.112 10'^ 2.026 10'^ 5.629 10*2 5.297 10'"
121100 107900 85900 113300 187000 179100 88200 125300 172400 112200 115900 87170 142100 178400 123500 147500 130100 210000 150200 171100 142600 159700 134400
2.587 10' 3.109 10' 8.737 10-2 5.059 10-2 6.591 10-^ 9.990 10-' 1.016 10' 4.645 10-2 2.852 10-2 1.885 10' 1.087 10' 3.345 10-2 1.921 10' 8.197 10-' 3.180 10-2 6.515 10-' 8.258 10-2 1.363 10-2 1.382 10-2 5.840 10-2 7.185 10-' 4.906 10-' 1.147 10-2
88 1/T(*1000)(K'^) 1.78 1.83 H
1.
1.93 h
• 27-DMBT 1 B 7-MBT A 2367-TMBT xBT X 356-TMBT • 2567-TMBT + 7-EBT - 235+236-TMBT - 56-DMBT o 45-DMBT D 357.TMBT 1 A 257-TMBT
Figure 1
Arrhenius plots for a set of components of the BT-family.
Components with substituents in position 2 show an important decline in reactivity with respect to benzothiophene. This is less pronounced for substituents in position 7 and position 3. The decline in reactivity due to substitution in position 2 and 7 is caused by the steric hindrance of these substituents on the vertical adsorption through the sulphur atom. A substituent in position 3 causes steric hindrance on the adsorption through the double bond of the thiophenic ring. Substituents in the positions 4, 5 and 6 hardly affect the reactivity of the component. The first order rate coefficients of the methyl substituted benzothiophenes can be related to the rate coefficient of benzothiophene as follows: 1. AAAJAAA^ Kj-r2l7 131415 16 Kjj
(7)
where f4=f5=f6. The power 6j equals one if the substituted BT contains a methyl substituent in position j , otherwise 5j is zero. The product f2'f7Y3'f4'fJ*f6' 's a global multiplication factor if the removal of the benzothiophenes occurs mainly on one type of catalytic site. The parameters fj were obtained by minimization of the objective function:
89
A 2:[ln(iL)-ln(^)]^ >=1
'^BT
^^^
BT
where n=23. The parameter estimates for T=240X, T=260°C, T=280°C and T=300X are given in Table 2. The parameters f, could not be determined at T=320°C because benzothiophene was completely converted at this temperature for the flow rates that could be applied in the equipment. Table 2 Values of the fj parameters of the rate coefficient prediction model for four temperatures. Parameter
T=240X
T=260°C
T=280°C
T=300°C
f2
0.0479
0.0882
0.125
0.183
h
0.334
0.295
0.298
0.318
f3
0.395
0.429
0.481
0.609
f4 = f5 = f6
1.078
0.871
0.886
0.807
The calculated versus experimental first order rate coefficients for the substituted BTs are shown in Figure 2 for all temperatures. The correlation coefficient R^ is 0.973. From the parameter values of Table 2 it is seen that a methyl substituent in position 2 reduces the reactivity at T=240''C with 95% with respect to BT. For methyl groups in positions 7 and 3 the reduction at T=240°C is respectively 67% and 61 %. The effect predicted by the model of the 4, 5 and 6 methyl substituents is a small increase at T=240°C and a reduction in reactivity at the other temperatures. The l, parameter values increase with temperature for substituents on the thiophene moiety and tend to decrease for substituents on the benzene moiety. Results for the DBT-family These results are given in Table 3. Houalla et al. [4] determined first order rate coefficients for DBT, 4-MeDBT and 4,6-DiMeDBT at 300''C and 102 bar on a C0M0/AI2O3 catalyst. The ratio between the rate coefficient of the substituted DBT and DBT obtained by Houalla et al. [4] and in the present work is given in Table 4. The ratios for 4-MeDBT and 4,6-DiMeDBT are higher than those found by Houalla et al. [4] The reactivity obtained for 4-MeDBT with respect to 4,6-DiMeDBT is almost 3 in this work and 1.35 according to the results of Houalla et al. [4]
90
Calculated ki [m^ /kg cat hr] 1
0.001
0.003
0.01
0.03
0.1
0.3
Experimental ki [m^ /kgcathr] Figure 2 Calculated versus experimental first order rate coefficients at all temperatures for components of the BT-family. Table 3 Preexponential factors, activation energies and first order rate coefficients at T=320°C for various members of the DBT-family. Sulphur component
A; [mVkg^Jh]
E,i [kJ/kmol]
ki [m'/kg^/h] atT=320°C
DBT 1-MeDBT 2+3-MeDBT 4-MeDBT 46-DiMeDBT 24-DiMeDBT 13-DiMeDBT 23-DiMeDBT 4-PrDBT 4Et6MeDBT 146-TriMeDBT 346-TriMeDBT
3.974 10' 1.717 10' 7.579 la' 1.181 10" 5.161 10' 3.371 10^ 5.421 10' 4.714 10" 5.220 10' 5.499 10' 2.364 10' 9.252 10"
64700 57800 91400 76850 54100 59400 96300 72500 48700 63900 59300 82700
1.064 10-' 1.892 10' 1.035 10' 2.115 10-' 8.152 10-" 3.457 10-' 2.253 10-' 2.566 10' 3.060 10' 1.174 10-' 1.821 10-' 4.733 10'
91 Table 4 Comparison of the reactivities found by Houalla et al. [4] and the present work at T = 3 0 0 X for 4-MeDBT and 4,6-DiMeDBT with respect to DBT llaetal. [4] '^-MeDBT'l^DBT
Present work
0.090
0.35
0.067
0.12
5.2. Hougen Watson kinetics Introduction Aromatic sulphur components are converted on a sites by hydrogenolysis and on r sites by hydrogenation. Several authors [3,5,6] found for the HDS of benzothiophene and dibenzothiophene that the surface reaction between adsorbed species is the rate determining step on both sites. The resulting Hougen Watson rate equation is generalised here for all aromatic sulphur components. The rate equation for the disappearance of an aromatic sulphur component i in a complex mixture becomes: K^\o^Ha^i
^H,
r.=' DEN„(T,C„...,C„)
^r^ir^Hr^i
^H,
DEN^(T,C„...,C„)
(9)
where the first term relates to the hydrogenolysis and the second to hydrogenation. The adsorption characteristics on both catalytic sites differ, as reflected by the denominators DEN„ and DEN^. These contain the concentrations of all adsorbing species and their temperature dependence. The functions represented by DEN„ and DEN^ are identical for all rate equations. Hydrogenation reactions of aromatics occur only on the r sites. The surface reaction was observed to be the rate determining step for biphenyl [3,6] and naphthalene hydrogenation. Since the hydrogenation reactions are irreversible their rate can be written: _
k..K.^K„^C. C^
^^^^
' DEN^(T,Cp...,C„) The rate equations for the hydrogenation of sulphur components and aromatics contain the same denominator DEN^. Determination of the denominators DEN,, and DEN, In complex mixtures the denominators DEN,, and DEN^ cannot be calculated a priori, since not all adsorbing species and their corresponding adsorption equilibrium constants are
92 known. Relating the rates of substituted sulphur components in complex mixtures to those of the unsubstituted heads of the families, requires the knowledge of both denominators DEN„ and DEN^ for each LCO experiment. For the complex mixture LCO these can be calculated considering model components for which kj^Kj^KH^ and kj^Kj^Hr ^^^ known, since these products are invariant of the mixture composition. The product k,)BT„K[)BT^KH„ was determined for the hydrogenolysis of dibenzothiophene [3]. The products k^mTT^Dm^Hr^ kupH^BPHr^HT and kN^Nr^Hr were determined for the hydrogenation of dibenzothiophene [3], biphenyl [3] and naphthalene in the present work. These products can be substituted in the rate equations of the reactions of dibenzothiophene and naphthalene in the LCO mixture. Now the denominators of the rate equations of these model components can be calculated. These denominators DEN„ and DEN^ are identical for the rate equations of all the sulphur compounds in the LCO. The expressions for the conversion of DBT, the conversion of DBT into biphenyl and cyclohexylbenzene, the conversion of naphthalene and the conversion of naphthalene into tetralin in a LCO in the completely mixed reactor can be written:
" " ^ " " C ^ V ^ DEN„(T,C„...,C„) _
W
DEN/T,C„...,C„)
^ H , J- ^ D B T o ^ D B T o ^ H o ^ D B T _ * ^ B P H T ^ B 1 » H T ' ^ H r ^ B P H -I
'"'" C ^ V^ DEN„(T,C„...,C„) " DEN,(T,C„...,C„)^ " " ' c £ v ^ DEN,(T,C, . . - ^ % NT
"
c^
NT
C.) DEN,(T,C
V^'DEN^(T,C„...,C„)^
V„„ =2.44 10'» exp[. -122.8 10^ ] kmol/kgjh
ko„,
=3.36 10-" exp[ilM_!2!] m'/kmol
KJ
mVkmol
=2.87 10'* exp[Zl^^4^] k'1101/kg.a/h
KJ
k„p„^ =3.41 10" exp[-^^Jiy^] kmol/kgjh
KJ
(13)
Cj'
(15)
with
K„„T„ =7.57 10'
(12)
(14)
HT
C ; V-^DEN/T,C„...,C„)^
KH„
(11)
93 K„,
=1.40 10-" e x p [ i i | i L l ^ ]
K^^^ =2.50 10-' exp[
76.8 10^
mVkmol
]
m'/kmol
K „ _ =4.97 10-" exp[ 37.9 10^ ]
m'/kmol
-185.9 10^
k^,
=1.57 10" exp[-
K^,
=5.06 10-' e x p [ ^ 4 ^ i ^ ] R,a.T
] kmol/kgjh mVkmol
The values of the unknowns, DEN„ and DEN^ were estimated for each LCO experiment by means of regression. They are shown in Figures 3 and 4 at T=320°C for various liquid compositions expressed in terms of a molar averaged conversion defined as follows: X
1
=
E ^i y*
(16)
Eyi with Xj the conversions of a set of selected components (BT, DBT, Naphtho[2,/-Z7]thiophene, 4-MeDBT, 4,6-DiMeDBT, Naphthalene and Phenanthrene) and yj the corresponding mole fractions in the LCO feed. The adsorption is weaker at higher temperatures on both the r and a sites and the denominators decrease. For a given temperature, the coverage of the r sites increases and that of the a sites decreases with molar averaged conversion. The reacting species in the LCO mixture adsorb to a higher extent on the a sites than the reaction products. 220
200
27
30
33
Molar averaged conversion [%]
Figure 3: Denominator corresponding to the c7-sites at T=320°C as a function of the molar averaged conversion.
27
30
33
Molar averaged conversion [%]
Figure 4: Denominator corresponding to the T-sites at T=320°C as a Function of the molar averaged conversion.
94 Determination of multiplication factors In a second step the numerical values of both denominators can be used in the rate equations of substituted aromatic sulphur components. The rate equations for these substituted sulphur components, sDBT e.g., can be related to that of the unsubstituted head of the family, DBT [2]: f
]c
K
K
f
fsDBT-^sDBT ^",1 pEN (T,C„...,C„)
\c
K
K
DEN,(T,C„...,C„) ^
The global multiplication factors f„ ^DBT and f^soBT are the products of the electronic and steric influences of the substituents on the adsorption equilibrium constant and of the electronic effects on the rate coefficient. These factors depend on the temperature. The conversion of a sDBT in a completely mixed reactor can then be written: _ ^
^sDBT
Hj |. *asDBT^DBTo ^ D B T o
Ho
^ ° ' ' ^ ~ C ^ V^ DEN„(T,C„...,C„)
TSDBT^DBTT ^ D B T T ' ^ H T - .
/1 0\
DEN/T,C„...,C„)^
The global multiplication factors f^ SDBT and frsDBT are the unknowns in this equation. As mentioned ^mja^mTcf^Ha and kj)j,T^t)j,TrKHr were already derived by Vanrysselberghe and Froment [3]. The denominators were determined for the compositions reached in each LCO experiment as explained above. The multiplication factors at a given temperature are obtained by minimization of the objective function:
/ ^
(^SDBT"^.S[)BT)
^
^
where X',DBT are the experimental conversions. In Table 6 the global multiplication factors t,i)BT and t,[)BT at T=320X are given for 4-MeDBT, for 4,6-DiMeDBT and for 2 and 3MeDBT which were not separated by GC-MS. 6. CONCLUSIONS The reactivity of various sulphur components of the benzothiophene and dibenzothiophene family has been demonstrated using first order rate coefficients. Substituents next to the sulphur atom reduce the reactivity with respect to the unsubstituted component. Substituents in other positions can either increase or decrease the reactivity. The electronic and steric influences of the substituents on the rate coefficient and the adsorption equilibrium
95 Table 6 Global multiplication factors f„,i,BT and t,^,^ at T=320°C for 2+3-MeDBT, 4-MeDBT and 4,6-DiMeDBT. Component
K sDHT
I i sDBT
2+3-MeDBT
1.3
1.1
4-MeDBT
0.15
3.3
4,6-DiMeDBT
0.037
1.8
constant in the more refined approach of Froment et al. [2] were combined into global multiplication factors f„ .DBT and f^ ^DBT i^ the Hougen Watson rate equations. The denominators corresponding to the a and r sites in the LCO can be obtained by using the Hougen Watson rate equations of the hydrodesulphurization of dibenzothiophene and the hydrogenation of naphthalene into tetralin. The a sites were found to be more occupied than the T sites. The global multiplication factors can be determined in a next step using the numerical values for the denominators. The next level of refinement involves the determination of the structural contributions as defined by Froment et al. [2]. This approach requires experimental data with 4-MeDBT, 4,6-DiMeDBT and one of the trimethyldibenzothiophenes as model components. The contributions kELo''''''(m;0;0), k^J'''''(n\;0;0), Ki,L+sTr''''''(m;0;0) and the product KsTa^^'^(4;0;0)KELo^"'^(m;0;0) can be obtained from experiments with 4-MeDBT. Combining these results with the global multiplication factors obtained with the LCO experiments, KsTa^*^^(4;0;0) can be calculated by dividing the global multiplication factor fs^DBx for 4-MeDBT by that for the 1, 2 or 3-MeDBT. An identical approach has to be followed to determine the structural contributions for the di- and trimethylDBTs. ACKNOWLEDGEMENT This work was funded by the European Commission under the Joule program contract no. JOU2-0121. V. Vanrysselberghe and G.A. Depauw are also grateful for a contribution from the Center of Excellence Grant awarded to the Laboratorium voor Petrochemische Techniek by the Belgian Ministry of Science. We wish to thank R. Le Gall for his cooperation. REFERENCES L G.A. Depauw and G.F. Froment, Journal of Chromatography, to be published. 2. G.F. Froment, G.A. Depauw and V. Vanrysselberghe, Ind. Eng. Chem. Res., 33 (1994) 2975.
96 3. V. Vanrysselberghe and G.F. Froment, Ind. Eng. Chem. Res., to be published. 4. M. Houalla, D.H. Broderick, A.V. Sapre, N.K. Nag, V.H.J. De Beer, B.C. Gates, and H. Kwart, J. of Catal., 61 (1980) 523. 5. I.A. Van Parys and G.F. Froment, Ind. Eng. Chem. Prod. Res. Dev., 25 (1986) 431. 6. R. Edvinsson and S. Irandoust, Ind. Eng. Chem. Res., 32 (1993) 391. NOMENCLATURE Ai
c> Eai pin pout
kE.."»T(m;0;0)
W,™^(m;0;0)
K,J«'T(m;0;0) KE,,.sT.™^(ni;0;0)
K.sT„""(4;0;0)
•^gas
T yin yout
w XfiPH ^CHB ^DBT ^sDBT
preexponential factor liquid concentration of component i inlet liquid concentration of component i activation energy inlet molar liquid flow rate of component i outlet molar liquid flow rate of component i electronic effect of one methyl group on the rate coefficient for the hydrogenolysis of sDBT electronic effect of one methyl group on the rate coefficient for the hydrogenation of sDBT first order rate coefficient of component i rate coefficient of component i on s sites electronic effect of one methyl group on the adsorption equilibrium constant of sDBT on the a sites electronic and steric effect of one methyl group on the adsorption equilibrium constant of sDBT on the r sites steric effect of a methyl group in position 4 on the adsorption equilibrium constant of sDBT on the a sites adsorption equilibrium constant of component i on s sites gas constant total rate of disappearance of component i absolute temperature total inlet volumetric liquid flow rate total outlet volumetric liquid flow rate total catalyst mass conversion of dibenzothiophene into biphenyl conversion of dibenzothiophene into cyclohexylbenzene conversion of dibenzothiophene conversion of substituted dibenzothiophene conversion of component i conversion of naphthalene conversion of naphthalene into tetralin calculated conversion of substituted dibenzothiophene
m-Vkg,,yh kmol/mL^ kj/kmol kmol/h kmol/h
mVkgJh kmol/(kg,,ih)
mL^/kmol kJ/kmol/K kmol/(kg,,ih) K mVh mVh kgcat
97 GREEK SYMBOLS p p^ a r
liquid density at reactor conditions inlet liquid density hydrogenolysis site hydrogenation site
SUBSCRIPTS BPH CHB DBT H H2 a r
biphenyl cyclohexylbenzene dibenzothiophene atomic hydrogen molecular hydrogen with respect to the hydrogenolysis function with respect to the hydrogenation function
SUPERSCRIPTS ^ in out
calculated inlet conditions outlet conditions
kg/m^ kg/m^
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® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P, Grange, editors
99
A review of catalytic hydrotreating processes for the upgrading of Uqmds producedfayflash pyrolysis R. Maggi and B. Delmon Unit6 de Catalyse et Chimie des Mat^riaux Divises, University Catholique de Louvain, Place Croix du Sud 2/17,1348 Louvain-la-Neuve, Belgium ABSTRACT Liquids produced by flash pjnrolysis of biomass and solid wastes are intended to be used in direct combustion but they contain a high quantity of oxygenated molecules which causes unwanted characteristics such as thermal instability, tendency to polymerise, corrosion and low heating value. These properties can be improved by partial or total elimination of oxygen atoms by catalytic hydrotreating with production of hydrocarbons and water. This paper reviews the development of this process from the first empirical tests using real oils and industrial sulphided cobalt-molybdenum and nickel-molybdenum supported on alumina catalysts to the development of a specially dedicated catfidjrtic system. Special attention is given to the catalytic aspects: the reaction schemes, the identification of the different catalytic functions and the control of the reaction. Finally, we discuss the utilisation of activated carbon as a new support. 1. WHY UPGRADE PYROLYSIS OILS? All biomass available as energy feedstock consists of chemically complex ligno-cellulosic materials. Their thermochemical and/or catalytic conversion produces gases, liquids and charcoal in various proportions. The liquids offer advantages in transport, storage, handling, retrofitting and flexibility of use. In addition they have a much Idgher energy density than the starting biomass. Fundamental studies such as those carried out by Shafizaded (1) indicate that high liquid yields from biomass can be obtained by pyrolysis, which is the thermal decomposition of the ligno-cellulosic matter either in the complete absence of oxidising agent, or with a limited supply in order to avoid gasification. In the last 10 years, liquid jdelds as high as 85% have been obtained by fast or flash pyrolysis (2-6) which involves extremely short residence times and extremely fast heat exchanges followed by rapid quenching. The valuable intermediary products are thus preserved before further repolymerisation.
100 These liqiiids, generally called oils, bio-oils, fast or flash p3n:olysis oils or biocrudes, rarely meet the standards required for fuels, but are nevertheless intended to be used as such in direct combustion in boilers, internal combustion engines or turbines. They have the aspect of a tar, they are viscous and not completely volatile and they do not mix with fossil fuels. In addition they are thermally imstable and tend to polymerise with time, temperature and light. These unwanted characteristics are related to the high oxygen content (up to 52%) present under the form of almost all oxygenated fimctions and as dissolved water (15-30%). Thus it is generally agreed that, to be used as fuels, bio-oils must be processed to remove oxygen. Two deoxygenation methods are currently proposed in the literature. One method proposes deoxygenation with simultaneous synthesis of gasoline-type compounds at atmospheric pressure through simultaneous dehydration-decarboxylation reactions over acidic zeolite catalysts and without reducing gases. This zeolite ZSM-5 is well-known for the production of gasoline from methanol. Its activity for the deoxygenation of other small oxygenated compoimds such as aldehydes and ketones has also been proved several times. At a typical temperature of 450 °C, oxygen is rejected as H2O, CO2 and CO (7). The maximal theoretical yield in hydrocarbons is 42% for flash p3rrolysis oils, but the literature indicates yields as low as 17-18% (8) because of the loss of carbon due to the high coke deposition (up to 15% of coke deposition on the catalyst and an extra 15% as suspended carbon) and because of the low conversion of the phenolic compounds (9). Moreover, the large molecules present in p3nrolysis vapours do not have access to the pores of the catalyst and, consequently, are not selectively converted. The other method proposed is hydrodeoxygenation (HDO) a t high temperature under hydrogen pressure in the presence of a catalyst (10). The reactions occurring are elimination of oxygen as water, elimination of nitrogen as ammonia, and hydrogenation-hydrocracking of large molecules. The reaction conditions and the catalysts (sulphided nickel-molybdenum or cobaltmolybdeniim supported on alumina) are similar to those used in the petrolexmi refining processes. The degree of deoxygenation can be easily modulated from simple stabilisation (elimination of more reactive functions such as carbonyl, olefins and carboxyles) to complete refining with a maximal theoretical hydrocarbon yield of 50%. This paper reviews the process of upgrading of bio-oils by catalytic hydrotreating from the first empirical tests to the recent development of new adequate catalytic systems. The reaction schemes, the potential inhibitors and the parameters enabling the control of the HDO reactions are discussed, as well as the catalyst deactivation by coke deposition. Finally, we discuss the utilisation of alternative neutral supports such as activated carbon. 2. THE EARLY YEARS The first studies concerning the upgrading by catalytic hydrotreating of vegetable oils were carried out in the early 80s' after the oil crisis. The main objective of these studies was to prove the feasibility of the hydrodeoxygenation process to produce hydrocarbonaceous fuels. Very little attention was paid to the setting up of the reaction conditions nor to the selection of adequate
101 catalysts, and most of the results were qualitative or obtained in inaccurate experimental conditions. The first quantitative and reproducible tests were performed by Elliott and Baker (11) in a flow reactor at 390°C and 13,5 MPa with hydroliquefaction oils. These oils contain more aromatics, polyaromatics and phenols than those produced by pyrolysis. This chemical composition is reflected by a low oxygen content (20%) compared to the high oxygen content of pyrolysis oils (52%). Different catalysts were tested: Ni, CoMo oxide, sulphided CoMo and NiMo, all supported on alumina. The best results were obtained with sulphided CoMo. The hydrogen consumption was 600 1 H2/1 fed oil and the yield in hydrocarbons was 75% of the fed oil. Later, on the basis of experiences with hydroliquefaction oils, the same authors hydrotreated oils obtained by p)rrolysis. However, the first experiences were relatively unsuccessful because of the extensive polymerisation of pyrolysis oils occurring at usual hydrotreating temperatures (12). The solution adopted by Elliott and Baker to avoid this thermal degradation was a pretreatment at lower temperature (270-280°C) aiming at eliminating more chemical functions such as aldehydes, ketones, carboxylic acids and esters (13). The yield in liquid hydrocarbons was 35% of the fed oil (dry basis), the maximal theoretical jdeld being 60%. This difference is due to the coke deposition leading to the quick deactivation of the catalyst. Elliott and Baker contributed greatly to the further development of hydrodeoxygenation process by setting up this two-step treatment, but very little attention was paid by these authors to the catalyst activation procedure and to keeping the sulphide state of the catalysts during the reaction. In fact, one important difference between pyrolysis oils and fossil fuels is that bio-oils do not contain any sulphur, thus H2S must be added during the process in order to preserve the sulphide state of the catalysts. Churin et al (14) hydrotreated bio-oils in a mechanically stirred batch reactor. The catalysts, industrial sulphided CoMo or NiMo supported on alumina, were directly suspended in the oil and CS2, which decomposes to H2S during the reaction, was added in order to maintain the sulphide state of the catalysts. This type of reactor does not allow a good contact between the liquid and the catalyst and, in addition, the concentration of the catalyst is low compared to fixed beds. In order to offset these problems, he used a hydrogen donor solvent (tetraline) which favours the hydrogen transfer, thus 40% yields in hydrocarbons were obtained. Later, it was demonstrated that tetraline could be replaced by a fraction of treated bio-oil or by a petroleum fraction such as diesel. Churin demonstrated that the hydrogen consumption begins at reaction temperatures of 200°C when a diluted bio-oil is treated. Other hydrodeoxygenation experiences were carried out in a batch reactor by Gagnon et al. (15). They treated bio-oils produced by vacuum pyrolysis which easily polymerise due to the high quantity of carbonyls. They developed a pretreatment at very low temperatures (60-100°C) using Ru/alumina catalysts. The stabilisation of the oil is explained by the hydrogenation of aldehydic functions in sugars contained in the oils. Gevert et al. (16) demonstrated that the alkalis contained in the bio-oils (ash) have a negative influence on the catalytic activity. They obtained good hydrocarbon yields (80%) when the oils were pre-treated by neutralisation.
102 Other groups also hydrotreated both Uquefaction and pyrolysis oils (17,18) in batch reactors but did not study the reaction parameters nor the catalysts or the process development. All these authors proved the feasibility of the deoxygenation by catalytic hydrotreating of oils produced by pjnrolysis. Important contributions such as the two-step treatment and the use of a solvent donor of hydrogen were carried out d u r i n g t h e s e early y e a r s . In p a r a l l e l , a n i n - d e p t h chemical characterisation of the oils was undertaken by different groups (19,20) in order to study the extremely complex composition of these oils containing hundreds of different molecules. Nevertheless, no optimisation of the reaction conditions nor of the catalytic system was done and, after this preliminary period, systematic studies leading to the optimisation and the scaling-up of the process became necessary. 3. THE MODEL COMPOUNDS APPROACH In 1989, Laurent (21) imdertook more fundamental studies concerning the hydrodeoxygenation of bio-oils in order to elucidate t h e main reaction pathways, the influence of the most important reaction parameters, the competition between different molecules and the possible inhibitors or poisons. This meant a complete kinetic and chemical study of the different reactions occurring during the upgrading process. For the whole study, he used industrial sulphided CoMo and NiMo supported on aliunina catalysts which had appeared to be the most adequate in the preliminary tests with real oils. Such a task would be extremely difficult, even impossible, with real oils because of the complexity of these liquids which are mixtures of hundreds of organic molecules: phenols, poly-substituted phenols, guaiacyls, aldehydes, ketones, carboxylic acids and esters, linear and aromatic ethers, sugars and others. In this case, model compounds representative of the real oil facilitate the analytical follow-up and the understanding of the different parameters. In addition, a lower concentration of the reactive chemical functions limits the polymerisation caused by the thermal reactions. It is very important that the tests carried out with model compound be as representative as possible of the real situation. This is why Laurent used model mixtures instead of isolated molecules, so as to establish the interactions and the competitions between molecules. Two different mixtures were defined, the first one contained the most reactive molecules: an aromatic ketone (4-methylacetophenone), a carboxylic ester (diethyl-decanedioate) and an aromatic ether (guaiacol). This "low temperature" solution mimicked the stabilisation step (< 300°C) prior to further hydrotreating. The second mixture, called "high temperature" solution, contained substituted phenols (4-methylphenol, 2ethylphenol) and dibenzofuran. It mimicked the total deoxygenation process requiring temperatures above 340°C. These molecules were selected on the basis of an in-depth chemical characterisation by fractionation into different chemical families followed by deep spectroscopic characterisation of each family (19).
103 4. FIRST STEP: REACTION SCHEMES The first part of the work performed by Laurent (21) in the frame of his PhD thesis was the elucidation of the hydrodeoxygenation reaction pathways of the four most abundant chemical functions in the bio-oils: phenol, carbonyl (ketone), carboxyl and methoxy. Special attention was given to the study of the catalytic activity and selectivity, as well as the possible competition between molecules. 4.1. Phenols The hydrodeoxygenation of phenols is a key reaction of the hydroprocessing of bio-oils when the production of highly refined hydrocarbon products such as transportation fuel is considered. Phenols and their derivatives represent an important portion of the organic part of these oils and, in addition, they are among the most difficult molecules to deoxygenate due to their high chemical stability. The hydrodeoxygenation process (HDO) has not been studied as much as hydrodenitrogenation (HDN) and hydrodesulfurisation (HDS), because oxygenated molecules are present in very low concentrations in petroleimi and because these molecules are not as harmful for the environment and for the catalysts as other heteroatom-containing molecules. Nevertheless, the HDO of phenolic compounds has been studied by several authors in the last years (2224). Two different catalytic functions of hydrotreating catalysts play a role in their conversion. One reaction scheme is the hydrogenation of the aromatic ring immediately followed by the deoxygenation of the intermediary cyclohexanol, which is generally not detected (25). The second pathway is the direct elimination of oxygen by hydrogenolysis of the Carom-0 bond. These two paths involve structurally different active sites (26,27).
Figure 1. Reaction scheme of 4-methylphenol (28) Laurent (28) studied particularly the hydrodeoxygenation of 4-methylphenol in solution with 2-ethylphenol and dibenzofuran in a batch reactor at 340°C
104 using sulphided CoMo and NiMo supported on alumina. CS2, which decomposes to H2S in the reaction conditions, was added in order to keep the sulphide state of the catalysts. 4-methylphenol is converted at a reasonable rate at 340°C, in typical runs its end conversion was 80-100%. 2-ethylphenol conversion rate was always 5 times lower than that of 4-methylphenol. Dibenzofiiran practically did not react. Figure 1 shows the hydrodeoxygenation reaction scheme for 4-methylphenol which is converted in toluene by hydrogenolysis and in methyl-cyclohexane by hydrogenation of the aromatic ring. 1 and 4-methyl-cyclohexene and 1,2cyclopentane are also produced by hydrogenation in minor amounts. In the expression of results, these minor products are grouped with methylcyclohexane. The intermediary 4-methyl-cyclohexanol has never been detected. Since these two path begin with strictly parallel reactions (fig. 2), a linear relation between toluene and methyl-cyclohexane is observed at moderate conversion rates. The deviation from the straight line at conversions higher than 50% is due to the hydrogenation of toluene to methylcyclohexane. This does not affect the measurement and expression of the selectivity methylcyclohexane/toluene, since Laurent performed it using the first conversion points.
T
0,0
0,1
0,2
0,3
'
I
0,4
0,5
CTOL
Co 4MP Figure 2. Relation between hydrogenation products and toluene (28) In the competition between molecules, the increase of 2-ethylphenol concentration causes a decrease of all reaction rates for both catalysts and results obtained by Laurent (28) indicate that this inhibition is more pronounced on the hydrogenolysis than on the hydrogenation. The influence of dibenzofuran was not considered by Laurent but, a priori y it has no specific influence in the conversion of 4-methylphenol since its concentration is almost constant and identical in all nins. 4.2. Carbonyl, carboxylic and guaiacyl grouiis The study of the hydrodeoxygenation of these groups is very important in the context of the upgrading of bio-oils, since they are the main cause of instability
105 and polymerisation because of their high chemical reactivity. The deoxygenation of these compounds leads to the stabilisation of the oils. This stabUisation could be the first step of a full refining process in order to avoid pol3rmerisation at the standard hydrotreating temperatures or, another interesting possibility, the reaction could be stopped at this stage leading to the production of a partially deoxygenated oil useful for the electricity production in turbines or diesel engines (10). However, the literature concerning the hydrodeoxygenation of these groups is very scarce. The reduction of carbonyls has been intensively studied in organic chemistry, as well as the transformation of ketones and aldehydes into alcohol over platinum group metal catalysts in very mild conditions (29). Maier et al studied the transformation of ketones in a methylene group over a metallic nickel catalysts (30). Weisser et al reported the hydrogenation of ketones over single metal sulphides (25). Concerning typical bimetallic hydrotreating catalysts, only Durand et al (31) worked on the hydrogenation of ketones over sulphided NiMo/alimiina. Laurent (32) studied the catalytic reaction schemes for these three groups using a model mixture containing 4-methyl-acetophenone, diethyldecanedioate and guaiacol. The tests were carried out in a batch reactor at different temperatures (260-300°C). The catalysts were industrial sulphided CoMo and NiMo supported on alumina. The sulphide state of the catalysts was maintained by addition of H2S. 4-Methyl-acetophenone: the conversion of the 4-methyl-acetophenone is very fast at 260°C, reaching 100% in less than 2 hours. The carbonyl is reduced to a CH2 with a very high selectivity. The only product observed is ethylmethylbenzene, the carbonyl group being hydrogenated to the alcohol which is quickly dehydrated under the reaction conditions. In addition, it is well known that yalumina catalyses dehydration reactions. The C=0 double bond hydrogenation is then the rate-limiting step (32). Figure 3 shows the reaction pathway.
,CHo
Figure 3. Hydrodeoxygenation reaction scheme for 4-methyl-acetophenone (32) Diethyl-decanedioate: according to the results obtained by Laurent (32), the hydrodeoxygenation of diethyl-decanedioate requires a temperature of aroimd 300°C over both NiMo and CoMo to be converted at a substantial rate. In fact, carboxyls are more refractory to deoxygenation t h a n carbonyls. Figure 4 shows
the reaction scheme leading to three linear alkanes: octane, nonane and decane. The intermediary products are Cs and C9 ethyl esters and their corresponding acids. It is reported in the literature (25) that carboxylic groups
106 are hydrogenated to CH3 groups with hemiacetal, aldehyde and alcohol as intermediary products which are quickly converted above 250^C on sulphide catalysts. These product have never been observed by Laurent. On the basis of these observations, Laurent proposes two mechanisms: one is the hydrogenation of the carboxylic group leading to a CH3 (reaction 1), the other is the rupture of the C-C bond leading to the complete decarboxylation (reaction 2). Octane is produced when reaction 2 occurs at both ends and decane si produced with reaction 1 at both ends. Nonane is produced by a combination of the two mechanisms. He defined the selectivity as the ratio between octane and decane. Laurent reports that a third reaction mechanism accounts for the production of carboxylic acids as intermediary. Using the corresponding pure acid (decanoic acid), he demonstrated that it is less reactive than the corresponding ester, giving slightly higher yields of decarboxylated products. Concerning the behaviour of the two catalysts, NiMo has a higher decarboxylating activity than CoMo, which could be due to the difference of acidity and the ensuing cracking activity between the two catalysts. Finally, he reported that the octane/decane selectivity slightly decreases with the conversion of the reactant.
^^ ( 1 ) 0 (2) O / H5C2-0fc^^(CH2)8 - C-0-C2H5^
O ( 1 ) ^ C10H22 H3C -(CH2)8 - C-O-C2H5 ^ ( > ^ ^ d ^ ^"20
(2) XH3C -(CH2)7 - C-O-C2H5 ^CT ( 2 ) ^ QH 18 Figure 4. Hydrodeoxygenation reaction schemes for carboxylic esters (32) Guaiacol: figure 5 presents the hydrodeoxygenation mechanism for guaiacol. The first step in the conversion of this molecule is the rupture of the C-CH3 bond leading to the formation of catechol which is then converted into phenol by elimination of one hydroxyl group. This phenol is subsequently transformed into benzene and cyclohexane. Under the reaction conditions used by Laurent, the conversion of guaiacol was limited to catechol and phenol (32). The total amoxmt of reaction products never accoimted for the converted guaiacol. Laurent reported 15-30% default in the molar balance at a conversion rate of 60%. Other authors reported 20% default at moderate conversion over CoMo at 250°C (33) and 10 to 50% conversion with the same catalysts at 300°C (34). These defaults in the molar balance can be related to the formation of heavy products or coke by analogy with the high tendency to form char during P)n:olysis of guaiacyls and hydroxyphenols (35).
107
Figure 5. Hydrodeoxygenation mechanism of guaiacol (32) Laurent (32) reported a decrease of the conversion rate of guaiacol over sxilphide catalysts. This decrease can be explained by the formation of coke and the ensuing blockage of active sites. But the production of catechol which adsorb strongly on y-alxxmina can also explain this deactivation (36). 5, SECOND STEP: ESIFLUENCE OF H2O, NHg AND HaS The refining or the stabilisation of highly oxygenated bio-oils require the total or partial elimination of oxygenated fimctions contained in molecules such as those above mentioned. The feasibility of this deoxygenation process has been proved with model compoimds at temperatures varying between 200-350°C over bimetallic sulphide catalysts. However, this reactivity can be different with real oils due to the presence of poisons or inhibitors: water which is dissolved in biooils (up to 30%), nitrogen which may be present in quantities as high as 3% in bio-oils and sulphur which is added during the hydrodeoxygenation reaction to keep the sulphide state of the catalysts (bio-oils contain extremely low quantities or even no sulphur). The influence of these three compounds, which are potential inhibitors or poisons but could also promote certain reactions, has been extensively studied in all hydrotreating reactions (24,37-41). Laurent (21,28,42) specifically studied their influence on hydrodeoxygenation reactions of model compounds representing bio-oils. Table 1 summarises the influence of ammonia, hydrogen sulphide and water on the different catalytic reactions involving the above mentioned model compoiinds. A scale going from — for a very strong inhibition to +++ for a very strong promoting effect has been used in the table in order to facilitate the comparison, 0 is used when there is no influence. As indicated in the table, water has no influence or very little inhibiting effect for all reactions except the hydrolysis of carboxylic esters which is promoted. Ammonia appears as a strong inhibitor of almost all reactions and this for both CoMo and NiMo. Surprisingly, it does not affect the hydrogenation of the ketonic group. Hydrogen sulphide has very little influence on the hydrogenation of ketones over sulphided CoMo, while it depresses the same reaction with NiMo. It has a promoting effect on the decarboxylation of carboxylic esters as well as on the hydrogenation of phenols. Demethylation of guaiacol is not affected.
108 Table 1 Inhibiting or promoting effects of NH3. H2S and H2O on the HDO reactions (21) Compound NH3 HgS HgO Reaction 4-Methyl-phenol hydrogenation
—
-0+
0
4-Methyl-phenol hydrogenolysis
—
~
0
4-Methyl-acetophenone hydrogenation
0
~
0
Carboxyl ester hydrolysis
—
0
++
Decarboxylation
—
+
0
Carboxyl ester hydrogenation
—
-
0
Guaiacol demethylation
~
0
0
These results suggest that both ammonia and hydrogen sulphide could be used for the control of the hydrodeoxygenation reactions: carbonyl groups could be selectively eliminated from complex feeds under a pressure of ammonia. On the other hand, carboxylic groups could be selectively eliminated by direct decarboxylation controlling the hydrogen sulphide pressure.
a CO
a o
0,05 0,10 0,15 0,20 H2S cone, (mole/1) Figure 6. Evolution of the hydrogenation (kj^cg • ) and hydrogenolysis (k^oL ^^ rate constants of the CoMo catalyst as a function of H2S pseudo in concentration (28). Another important effect of H2S partial pressure is the control of the hydrodeoxygenation of phenols, since the hydrogenation/hydrogenolysis selectivity is strongly influenced, specially over CoMo. This control of the reaction could allow the control of the hydrogen consumption which is extremely important for the upgrading of real bio-oils. In fact, four molecules
109 of hydrogen are consumed via the hydrogenation pathway against only one via hydrogenolysis, the hydrogen consumption being multiplied by 2.3 in the H2S concentration range investigated (28). Figures 6 illustrates the influence of H2S partial pressure on the hydrogenation/hydrogenolysis selectivity over CoMo (28). Similar observations could be made for the conversion of carboxylic ester since t h e decarboxylation pathway consumes more hydrogen t h a n the hydrogenation one (42). & PROBLEM: DEACTIVATION BY COKE DEPOSITION Experiences performed in the early years with real bio-oils indicated a low stability of the catalytic system. Experiments could not r\in over a few days, nor even a few hours: the catalyst, embedded in coke, plugged the reactor. This could be explained by the still intense polymerisation occurring at the stabilisation temperatures (250-300°C), but also by the deactivation of the catalyst. It was t h u s necessary to understand this phenomenon. Laurent (21,28,32) calculated molar balances for each of the molecules he studied, and concluded t h a t these balances reached or were very close to 100% for the hydrodeoxygenation of the ketonic, carboxyl and hydroxyl groups.
100
150
200
250
300
Reaction time (min) Fig. 7. Evolution of the molar fraction of guaiacol (•), catechol (A), phenol (O) and their siim (H) as a function of the reaction time (CoMo, 280 °C), (32). The situation was completely different for guaiacol and catechol. As mentioned in section 4.2., up to 30% default in molar balance can be observed at a conversion rate of 60% of guaiacol. Figure 7 presents the molar balance for a typical run over CoMo, where a 20% default can be observed. The balance is still poorer with NiMo. This tendency to coke formation is tj^ical for phenols containing two or more oxygenated substitutes such as guaiacyls and hydroxyphenols. Klein et al. (33,34) gave valuable information concerning the
110 HDO reactions of anisole and guaiacol. They reported t h a t guaiacol reacts faster t h a n anisole but t h a t anisole is quantitatively converted, whereas for guaiacol it is not possible to close mass balances. Laurent and Centeno thus undertook a systematic study of this phenomenon of coke formation (43). They carried out experiences with guaiacol in the standard model mixture, guaiacol alone, hexadecane (solvent) alone, catechol, phenol and methylanisol over the traditional sulphided CoMo/y-alumina but also over y-alumina. Results obtained by these authors are presented in table 2. Two important conclusions can be drawn: 1) the coke deposition is very similar with guaiacol in mixture or alone, confirming that the other model molecules used by Laurent do not form coke and, 2) the mixture reacts over pure alumina with a poor rate constant but a high coke deposition, meaning t h a t the coke formation must be attributed to the aliunina support. Catechol has a reactivity similar to t h a t of guaiacol and also leads to a high coke formation. The catalysts used for phenol and methylanisol present a much lower coke content, indicating t h a t benzenic molecules containing only one oxygen have a lower propensity to this phenomenon. This coke would result from the condensation of guaiacol and catechol themselves. Table 2 Initial rate constant and quantities of coke deposited (43) Rate constant Reactant Catalyst (min-1. g.-l).103
Carbon content (%p) 1.8
Hexadecane (solvent) Guaiacol in mixture
C0M0/Y-AI2O3
9.4
8.9
Guaiacol in mixture
Y-AI2O3
2.9
10.3
Guaiacol
C0M0/Y-AI2O3
9.2
7.8 (85% conv.)
Catechol
C0M0/Y-AI2O3
1L7
5.5 (90% conv.)
Phenol
C0M0/Y-AI2O3
0
3.5 (0% conv.)
Methylanisol
C0M0/Y-AI2O3
19.7
2.8 (95% conv.)
C0M0/Y-AI2O3
7. THE FUTURE: A NEW CATALYTIC SYSTEM The conversion of guaiacol in experiences performed by L a u r e n t and Centeno can be attributed to the alumina support since, without active phase, it has a certain activity for the conversion of the reactant. Unfortunately, it also has a high activity for the formation of coke and heavy products. On this basis, neutral supports such as activated carbon or silica or even non supported catalysts could be a good alternative to avoid coke formation. Centeno et al.
Ill explored these possibilities (44). They prepared different catalysts: CoMo supported on activated carbon and on silica, and non supported. Table 3 presents results obtained with these catalysts compared to the traditional alumina.
Tables
Rate constants and phenol/catechol selectivity obtained with diflFerent cat. (44) Catalyst Rate constant Phenol/catechol (min'\ g. cat.'\cm^) (%) -^0 rAl203 0.35 12.6 1.30 COMO/Y-A1203 2.0 0.28 CoMo/Si 0.22 89.3 CoMo/C 0.39 8.0 CoMoS This table shows that the catalyst supported on activated carbon seems very promising because, even if its activity is very poor compared to alumina, it has a very high phenol/catechol selectivity. The other proposed catalysts do not show interesting possibilities since both activity and selectivity are low. Concerning the coke formation, the catalyst supported on carbon is also very interesting. Evidently, no values of carbon content are available for activated carbon-supported catalysts, but Centeno reported that the catalyst is not embedded by coke after reaction and remains active. In a recent congress (45), he presented the molar balance for the hydrodeoxygenation of guaiacol over CoMo/C compared to that obtained over CoMo/alumina (figure 8).
^
0,028 1
1
0,026 <
2 0,024-
o o
•*• 0,022-
1 ^ o
3
o
O 0,020S
0,0180,016 1 ()
1
20
O CoMo/Al-2| • CoMo/C A NiMo/C
•
1
40
o •
o 1
60
•
ft
o 1
80
Time (min)
'
0 1
100
•
1
120
Figure 8. Molar balances for guaiacol conversion over CoMo/C and NiMi/C and CoMo/alumina
112 The figure shows clearly that the mass balances reach 100% for carbon supported catalysts. This point, in addition to the fact that these catalysts are highly selective in phenol, confirm the first tests. Nevertheless, both CoMo and NiMo supported on carbon still have a poor activity. This activity must be increased. This could be achieved by improvement of the preparation procedure, better dispersion of active phases and modification of the micro structure of carbon. In fact, activated carbons have very high BET area (>1000mVg) with a high percentage of micropores. In parallel, other catalysts such as noble metals supported on carbon are being studied and tested. These t3rpes of highly hydrogenating catalysts are particularly interesting in the actual context of the stabilisation of bio-oils leading to the production of a fuel to be fed in a diesel engine for the production of electricity. This fuel would be only stabilised by saturation of double bonds limiting the deoxygenation and, of course, the hydrogen consimiption.
REFERENCES 1. Shafizaded, F., Industrial P3n:olysis and Cellulosic Materials, Applied Pol. Symp., 28 (1975) 153. 2. Scott D. S. and Piskorz J., Can. J. Chem. Eng, 60 (1982) 666. 3. Medina E. and Cuevas A., in 6th European Conference on Biomass for Energy, Industry and Environment, Grassi, G. et al. (eds.), (1991) 1200. 4. Diebold, J. and Scahill, J., in ACS S3maposium on Production, Analysis and Upgrading of Pyrolysis Oils from Biomass, (1987) 21. 5. Graham, R., Freel, B., Huffman, D. and Bergougnou, M., in Advances in Thermochemical Biomass Conversion, Bridgwater, A. V. (eds.), (1992) 1275. 6. Roy, C, de Caumia, B., Brouillard D. and Menard, H., in Fimdamentals of Thermochemical Biomass Conversion, Overend, R. P., Milne, T. A. and Mudge, L.K. (eds.), (1985) 237. 7. Chang, C. D., Silvestri, A. J., Journal of Catalysis, 47 (1977) 249. 8. Diebold, J. P., Scahill, J. W., Evans, R. I., in Biomass Thermochemical Conversion Contractors' Meeting, Minneapolis, USA, (1985) 31. 9. Renaud, M., Grandmaison, J., Roy, C. and Kaliaguine, S., in Pyrolysis Oils from Biomass: Producing, Analysing and Upgrading, Soltes, E. and Milne, T. (eds.), (1988) 290. 10. Laurent, E. and Delmon B., in 7th EC Conference on Biomass for Energy, Industry and Environment, Grassi, G. et al. (eds.), 1992. 11. Elhott, D. and Baker, E., SAE paper n° 859096,1985. 12. Elliott, D. and Baker E., in Energy from Biomass and Wastes X, Klass, D. L. eds., (1987) 765 13. Elliott, D. and Baker E., in Thermochemical Biomass Conversion, Bridgwater, A. V. and Kuester, J. L. eds., (1988) 883. 14. Churin, E. et al., in Thermochemical Biomass Conversion, Bridgwater, A. V. and Kuester, J. L. eds., (1988) 878. 15. Gagnon, J. and Kaliaguine, S., Ind. Eng. Chem. Res., 27 (1988) 1783.
113 16. Gevert, B., in Advances in Thermochemical Biomass Conversion, Bridgwater, A. V. (eds.), 1992. 17. Nelte, A. and Meier zu Kocker, H., in International Congress Eurofonim New Energies, Ferrero, G. and Grassi, G. eds., 3 (1988) 673. 18. Elamin, A., Capart, R. and G61us, M., in 6th European Conference on Biomass for Energy, Industry and Environment, Grassi, G. et al. (eds.), (1991)693. 19. Maggi, R. and Delmon, B., Fuel, 73 (1994) 671. 20. Pakdel, H. and Roy, C , Am. Chem. Soc. Div. Fuel Chem. Prep., 32(2) (1987) 203. 21. Laurent, E., Etude et controle des reactions d'hydrod^soxygenation lors de rhydrorafifinage des huiles de pyrolyse de la biomasse, PhD thesis, UCL, Belgiimi, 1993. 22. Satterfield, C. and Yang S., J. Catalysis, 80 (1983) 56. 23. Odebunmi, E. and Ollis, D., J. Catalysis, 80 (1983) 56. 24. Gevert, B., Otterstedt, J. and Massoth, F., Applied Catalysis, 31 (1987) 119. 25. Weiser, O. and Landa, S., Sulphide catalysts: Their Properties and Applications, Pergamon, 1973. 26. Stuchly, V. and Beranek, L., AppUed Catalysis, 35 (1987) 35. 27. Okamoto, Y., Maezawa, A. and Imanaka, T., J. Catalysis, 120 (1989) 29. 28. Laurent, E. and Delmon, B., Ind. Eng. Chem. Res., 32(11) (1993) 2516. 29. Rylander, P., Catalytic hydrogenation over platinum metal. Academic Press, (1967) 229. 30. Maier, W., Gergman, K , Bleicher, W. and Schleyer, R., Tetrahedron Letters, 22 (1981) 4227. 31. Durand, R., Geneste, P., Moreau, C. and Pirat, J., J. Catalysis, 90 (1984) 147. 32. Laurent, E. and Delmon, B., AppHed Catalysis A, 109 (1994) 77. 33. Hurff, S. and Klein, M., Ind. Eng. Chem. Fimdam., 22 (1983) 426. 34. Petrocelli, F. and Klein, M., Fuel Sci. Tech., 5 (1987) 63. 35. Bredemberg, J. and Ceylan, R., Fuel, 62 (1983) 343. 36. Bredemberg, J. and Sarbak, Z., J. Chem. Tech. Biotechnol., 42 (1988) 221. 37. Lemberton, J., Touzeyidio, M. and Guisnet, M., Applied Catal., 54 (1989) 91. 38. Satterfield, C. and Smith, C , Ind. Eng. Chem. Process Des., 25 (1986) 942. 39. Gultekin, S., Ali, S. and Satterfield, C , Ind. Eng. Chem. Process Des., 25 (1986)431. 40. La Vopa, V. and Satterfield, C , J. Catal., 110 (1988) 375. 41. Broderick, D. and Gates, B., AIChEJ, 27 (1981) 663. 42. Laurent, E. and Delmon, B., AppUed Catalysis A, 109 (1994) 97. 43. Laurent, E., Centeno, A. and Delmon, B., Catalyst Deactivation, Delmon, B. and Froment, G., eds., (1994), 573. 44. Centeno, A., Laurent, E. and Delmon, B., J. Catal., 154 (1995) 288. 45. Centeno. A., David, 0., Vanbellinghen, C , Maggi, R. and Delmon, B., in Developments in Thermochemical Biomass Conversion, Bridgwater, A. V. (eds.), in press.
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® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
115
Dual-functional Ni-Mo sulfide catalysts on zeolite-alumina supports for hydrotreating and hydrocracking of heavy oils H. Shimada^, S. Yoshitomi, T. Sato^, N. Matsubayashi^, M. Imamura^, Y. Yoshimura^ and A. Nishijima^ ^Surface Chemistry Department, National Institute of Materials and Chemical Research, Tsukuba, Ibaraki 305, Japan 'faculty of Engineering, Shibaura Institute of Technology, Shibaura, Minato-ku, Tokyo 108, Japan Ni-Mo sulfide catalysts supported on the mixtures of 7-AI2O3 and HY zeolite with various ratios were prepared and tested in the hydroprocessing reactions of heavy oils to reveal the roles of the dual-fimctions of hydroprocessing catalysts. Hydrogenation activity solely functioned for the hydrodesulfiirization and hydrodearomatization of heavy feedstocks with high nitrogen contents, typically oil sand bitumen and coal-derived oils. For other feedstocks, the dual-fiinctionality played important roles not only in hydrocracking but also in hydrodenitrogenation. The optimum balance of the hydrogenation/hydrocracking activities greatly depend on the feedstock properties. Pre-hydrotreatment which reduced the concentration of heavy nitrogen-containing materials was effective for maximizing the dualfunctionality. 1. mXRODUCTION To achieve high-quality petroleum products with desired distribution from crude oils, most of the recent petroleum refineries consist of distillation, cracking, and hydrotreatment as the major unit processes. Light and middle fractions obtained by distillation can be relatively easy to convert into valuable final products through hydrotreating processes, mainly hydrodesulfurization (HDS). On the other hand, heavierfi-actionsthan atmospheric gas oil cannot be converted into valuable products merely through the hydrotreating processes, particularly due to the recent decreasing importance of fuel oil. This results in the demand to convert heavier fractions into light components by either hydrocracking (HCK) or fluid catalytic cracking (FCC) processes. Since late 1980s, environmental concerns have been placing increasing emphasis on the production of clean transportation fuels, for example, diesel fuels with low sulfur and aromaticity. There is an increasing demand for middle distillates such as kerosene, turbine and diesel fuels particularly in developed countries. All these circumstances have been making the HCK process more attractive one and resulting in the developments of processes and catalysts [1-6]. The technology of the present HCK processes were established long years
116 ago, however, there still exists much room to improve the efficiency and flexibility, which are deeply related to the catalyst development. The catalysts applied to the HCK processes in general possess two catalytic fiinctions, namely hydrogenation (HYD) and HCK fiinctions. For the HCK of heavy feedstocks in the presence of H2S, metal sulfides such as Ni-Mo, Co-Mo or Ni-W sulfides are used to provide the catalyst with HYD function. Amorphous mixed oxide or zeolite-containing AI2O3 has been applied to the HCK catalyst supports for long years to provide the catalyst with HCK function arising from the solid acidity. At present, the use of zeolite-based supports with high activity and stability is prevailing because of the recent rapid development of the zeolite technology. The combination of the Ni-Mo, Co-Mo and Ni-W sulfides is same as that of hydrotreating catalysts, that have been extensively studied regarding the genesis and structure of the active sites, the reaction mechanisms, the catalyst deactivation and other respects by many groups [7]. Lots of investigations to elucidate the structure and properties of zeolite have been carried out to improve the catalytic performance of zeolitic materials, however, most of the studies have been dedicated to the development of FCC catalysts [8]. In this context, when compared with the hydrotreating or FCC catalysts, much smaller number of fiindamental studies have been reported on the HCK catalysts with dual-functionality in spite of large numbers of patents and process-oriented studies. The present study has been conducted aimed at the understanding of the roles of the catalytic dual-functionality in the HCK reactions. Commercially available catalysts, which contain many kinds of additives to improve the catalytic activities, possess HCK activities even though the catalog claims 7-AI2O3 supported catalyst. Thus, we have prepared several kinds of catalysts with different composition of supports. Afi:er loading the Ni-Mo sulfide, the catalytic activities of the catalysts with different balances of HCK and HYD activities have been examined for hydroprocessing heavy oils. To discuss the relationships between the catalytic functionality and the feedstock properties, some kinds of synfliels have been used in addition to petroleum vacuum gas oil (VGO) and residual crude (RC). Taking into account that the HYC catalysts are employed both at the first and second-stage reactors, prehydrotreated feedstocks have been tested in addition to raw feedstocks. In the last of the paper,fiatureaspects for the improvement of the HCK catalysts are presented. 2. EXPERIMENTAL 2.1 Catalysts Zeolite-containing supports were prepared by extruding mixtures of H-type high Si02/Al203 Y-zeolite (Si/Al~10) and 7-AI2O3 powders into pellets with a diameter of 1/16 inch. For high HY zeolite-containing supports, non-porous AI2O3 was added as a binder. Mordenite and Ti02 supports were prepared with the same procedures. All the supports were calcined at 500 °C prior to the following catalyst preparation. NiO (1.7 wt%) and M0O3 (6.7 wt%) were loaded on the supports by the incipient wetness method using an aqueous solution of Ni(N03)2-6H20 and (NH4)6Mo7024.4H20. The catalysts were then dried at 110 °C for 10 h and calcined at 500 °C for 3 h. In addition to the above laboratory-prepared catalysts, a commercially available 7-AI2O3 supported catalyst with 4.0 wt% of NiO and 14 wt% of M0O3 was employed as a reference catalyst. All the catalysts were presulfided with a gas flow of 5
117 Table 1 List of catalysts Catalyst NiO (wt%) NM/Al'' 4.0 NM/HYo 1.7 NM/HY2 1.7 NM/HY7 1.7 NM/HYio 1.7 1.7 NM/HM 1.7 NM/Ti
M0O3 (wt%) 14.0 6.7 6.7 6.7 6.7 6.7 6.7
Support Composition AI2O3 AI2O3 HY (20wt%)- AI2O3 HY (70wt%)- AI2O3 HY zeolite*^ H-type Mordenite Ti02^
Physical Properties S.A. P.V. (ml/g) (m'/g) 141 0.47 254 0.76 312 0.74 495 0.48 579 0.39 397 0.32 151 0.39
Acidity APD (mmol/g)
(A)
130 120 95 38 27 32 100
*5
1.1 1.8 4.2 5.5 3.4 0.27
*1 Commercially available catalyst. *2 Non-porous AI2O3 was used as a binder. Measured by N2 adsorption. S.A.(surface area), P.V.(pore volume), A.P.D. (apparent pore diameter). *4- Estimated by temperature programmed desorption (TPD) of NH3. The numbers show the amount of NH3 disorbed over 150 ^C. Not measured. Table 2 Properties of feedstocks Feedstock VGO RC CL-VGO CL SAND SHALE H-VGO*^ H-RC'' H-CL*^ H-SAND H-SHALE
Distil.'' {%rc) -350 350-548 8 85 9 49 46 53 68 11 30 18 41 48 15 22 65 38 -
80 54 26 47 -
5487 41 1 21 52 11
Asp.*^ (wt%) 0 3.3 0 10.1 9.6 0.9
5 24 9 15 -
0 1.0 4.7 1.6 0
Elemental composition H/C (-)*' S (%) 1.68 2.96 1.67 3.23 1.10 0.048 1.12 0.50 1.57 4.57 1.66 0.53 1.81 1.67 1.16 1.57 1.85
0.078 0.33 0.068 0.35 0.15
N (%) 0.083 0.18 0.68 0.67 0.38 0.99 0.021 0.13 0.34 0.33 0.15
Hydrotreated over a Ni-Mo/Al203 catalyst at 425 C for 3 h at a constant pressure of 14.7 MP with a H2flowof 1 1/min. *2 Simulated distillation. *3 Amount of asphaltene measured as a Soxhlet extraction residue by hexane. *4 Atomic ratio.
118 vol.% H2S-95 vol.% H2 for 2 h at 400 °C before reaction. Table 1 summarizes the chemical compositions and physical properties of the catalysts used. 2.2 Feedstocks The petroleum feedstocks used in the present study were VGO and RC from Middle East crude oil. To examine the catalytic activities for heavier hydrocarbons, Australian Morwell coal-derived crude oil (CL) and vacuum gas oil fraction (CL-VGO), oil sand bitumen (SAND) from Athabasca in Canada, shale oil (SHALE) from Mona in China, were also employed. Properties and elemental compositions of the feedstocks are illustrated in Table 2. The SHALE and CL contained large amounts of light fractions (350 °C"), since they were not distilled. Hydrotreated feedstocks were prepared over a commercially available NiM0/AI2O3 catalyst. 2.3 Reaction procedures The HYD and HCK activities of the catalysts were evaluated as the basic functions by model test reactions using 1-methyltetralin for HYD and dimethylnaphthalene for HCK. Detailed procedures of the model test reactions were described in a previous paper [9]. All the hydroprocessing reactions were carried out in laboratory batch reactors with an inner volume of 50 cm^. The initial charge to the reactor was 10 cm"^ of feedstock, 0.5 g of catalyst, and 9.8 MPa of H2. The reactions were performed at 425 °C for 2 h. Other details were already described in a previous paper [10]. 3. RESULTS AND DISCUSSION 3.1 Basic functions of catalysts Basic functions of the catalysts evaluated by the model test reactions are summarized in Table 3. The HYD activity of the NM/HYx catalysts decreases with increasing zeolite content. The present catalysts except the NM/Al catalyst contain comparatively low concentrations of NiO and M0O3 to minimize the catalyst agglomeration on the supports. However, the low HYD activities of the catalysts with high zeolite contents are presumably due to the poor dispersion of Ni and Mo sulfides on the external surface of zeolite. It should be noted that the numbers in Table 3 present the results performed in batch reactors, thus not being proportional to the intrinsic catalytic activities particularly for high conversion regions. In contrast to the HYD activity, the HCK activity of the catalyst increases with increasing zeolite content. This is due to the high acidity of zeolite, as indicated in Table 1. It has been confirmed that the AI2O3 support alone does not crack diphenylmethane in the given reaction conditions, while large amounts of benzen have been produced over the zeolite-containing supports. This indicates that the cracking of diphenylmethane is catalyzed on the Br0nsted acid sites but not on the Lewis acid sites. It should be noted that the TPD method employed in the present study cannot distinguish the Lewis acidity of AI2O3 from the Br0nsted acidity of zeolite. In the reaction over the NM/HYQ catalyst, diphenylmethane is cracked on the Br0nsted acid sites of Ni-Mo sulfides [11, 12]. The results in Table 3 demonstrates that the balance of HCK/HYD activities can to a large extent be controlled by the zeolite content in the support, while catalysts possessing both high HYD and high HCK activities are not available by simply changing the zeolite content.
119 Table 3 Basic functions of catalysts Catalyst N M / H Y Q NM/HY2 N]VraY7 NM/HYio 41 38 HYD*^ 62 53 86 HCK*^ 6 57 87
NM/HM NM/Al NM/Ti 35 60 57 56 11 20
* 1: Yields of 1- and 5-methyltetralin in the hydrogenation of 1-methylnaphthalene. *2: Yields of benzene and toluene in the hydrocracking of diphenylmethane.
Table 4 Hydroprocessing of WGO Catalyst*^ TiofC)'^ TsoCcr AH/C^ H2 cons, (mg/g)*"^ HDN (%)*^ HDS (%)*^
NM^Yo 318 433 0.05 3.3 28 46
NM/HY2 NM/HY7 NM/HYio 130 81 266 383 354 414 0.18 0.14 0.18 7.3 6.8 5.5 87 91 70 85 80 82
NM/Al NM/Ti 275 283 414 418 0.15 0.15 5.0 5.9 68 83 83 89
* 1: NM/HM was not applied to the reaction. *2: Temperature at which 10 or 50 vol.% oil distills over. Tio and T50 for the feedstock are respectively 345 and 447 ""C. *3: Change in H/C ratio. *4: Hydrogen consumption during the reaction. *5: Nitrogen removal. *6: Sulfur removal.
Hydroprocessing ofRC NMy^Yo NM/HY2 NMyHYv NMyllYio NM/HM NM/Al NM/Ti Catalyst 52 25 16 53 36 9 29 HDA (%)*~^' 24 27 23 30 24 10 HDN (%) 19 47 23 58 34 33 54 50 HDS (%) 0.02 0.05 0.02 -0.01 0.00 -0.01 0.09 AH/C 5.4 7.5 3.7 4.6 6.3 5.0 4.6 H2 cons, (mg/g)
^ 1: Asphaltene removal.
120 The commercially available NM/Al catalyst demonstrates lower HYD but higher HCK activities than the NM/HYQ catalyst. The Ni-Mo active component of the NM/HYQ catalyst is highly dispersed on the pure 7-AI2O3 support, yielding higher HYD activity in spite of the half Ni and Mo loading of the NM/Al catalyst. On the other hand, some additives in the NM/Al catalyst improve the HCK activity of the AI2O3 supported catalyst, either directly by the solid acidity or through the enhancement of the Br0nsted acidity of nickel-molybdenum sulfide. For instance it is known that the addition of phosphorus increases the HCK activity of Ni-Mo/A^Os catalysts [13]. The lowest HYD activity observed for the NM/HM catalyst is probably due to poor dispersion of Ni and Mo sulfide on the external surface of H-Mordenite crystals. The NM/HM catalyst with strong acidity does not exhibit high HCK activity either, compared with the NM/HYio and NM/HY7 catalysts. This suggests that cracking of diphenylmethane does not require strong acidity but favors large numbers of weak or moderate Br0nsted acid sites. The NM/Ti catalyst gives higher HCK activity than the AI2O3 supported NM/HYQ and NM/Al catalysts in spite of the very low acidity (Table 1). In a previous paper [9], we reported that a Ti02 supported molybdenum sulfide catalyst gave significantly high HCK activity without decreasing the HYD activity and discussed that electron transfer from the Ti02 support to molybdenum sulfide played an important role in the enhancement of the Br0nsted acidity induced under the reaction conditions. The Ni promoted catalyst in the present study show the same trend but the enhancement of the HCK activity by the support is depressed. The dispersion of molybdenum and nickel sulfides on Ti02 is as high as that on AI2O3, which makes the HYD activity of the NM/Ti catalyst higher than the zeolite supported catalysts. 3.2 Hydrocracking and hydrotreating of petroleum VGO and RC Table 4 summarizes the reaction results of petroleum VGO over the catalysts. The index Tio obtained by simulated distillation represents the conversion to light fractions; a low Tio value indicates much production of gasses. The T50 index correlates the total conversion of the heavy fraction into the light one. Among the NM/HYx catalysts, Tio and T50 decrease with increasing zeolite content. It is also shown that larger amounts of H2 are consumed over the zeolite supported catalysts, which yield liquid products with higher H/C atomic ratios. These indicate that hydrocracking is the dominant upgrading reaction and that the heavy fractions in VGO are readily hydrocracked over the acid sites of zeolite with the production of gaseous fractions. It should be noted that the present batch-type reaction system enhances the gas production compared with practical flow-type reactors. As regard to heteroatom removals, the NM/HY7 and NM/HYio catalysts demonstrate higher hydrodenitrogenation (HDN) activities than the other catalysts, while the NM/Al catalyst shows a superior HDS activity to the high zeolite-content catalysts. The Br0nsted acidity of the zeolite support functions effective for HDN, but not very effective for HDS of VGO. High dispersion of the sulfide catalysts leading to a large number of the catalytically active sites is a more important factor than the dual-functionality for the HDS of VGO fractions. The NM/HYo catalyst with the highest HYD activity in the model test reactions exhibits the lowest performances for all the items. The NM/Ti catalyst does not show high activities in spite of the high HYD and HCK performances evaluated by the model test reactions. The induced Br0nsted acidity of the NM/Ti catalyst, which cannot be measured by the TPD
121 method, does not function in the of the real feedstocks. These suggest that some kind of acidity enhancement, probably solid acidity, is essential for the hydrotreating catalysts to achieve high heteroatom removal activities from VGO fractions. In fact, some previous studies evidenced advantages of zeolite incorporation into the support of hydrotreating catalysts in HDS and HDN of heavy feedstocks [14] or model compounds [15], though modified 7-AI2O3 support is still prevailing in commercial use. The acidity enhancement by the additives may function enough effectively for the heteroatom removals from VGO fractions. The catalytic activities in the hydroprocessing of petroleum RC are summarized in Table 5. In contrast to the processing of VGO, the superiority of the zeolite supported catalysts is much depressed. As indicated by the largest amount of H2 consumption over the NM/HYio catalyst, the acid sites of the zeolite function as the HCK active sites. However, as evidenced by the low asphaltene and sulfur removal among the HYx supported catalysts, the HYD active sites function more effective to the upgrading of RC than the HCK active sites. The negative values observed for the AH/C in the reaction over the NM/HY7 and NM/HY[o catalysts suggest that the acid sites of the zeolite promote retrogressive reactions with reducing the HYD activity. These low activities of the zeolite supported catalysts may be partly due to the pore diffusional limitation of large molecules into the micropores of zeolite. Careful observation of Table 5 shows that the optimum balance of HYD/HCK differs among each reaction. For nitrogen removal in which the NM/HY7 catalyst gives the highest performance, the acidity of zeolite contributes to the reaction to a certain extent. Asphaltene removal (HDA) also to some extent requires the dual-functionality, as evidenced by the highest activity by the NM/HY2 catalyst. On the other hand, the HDS activity is controlled dominantly by the HYD function as observed in the order of catalytic activity, NM/HYQ > NM/HY2 > NM/HY7 > NM/HYio. The NM/Al catalyst presents high performances for overall reactions, again probably due to the enhancement of the acidity by unknown additives in the commercial catalyst. Neither of the NM/Ti nor NM/HM catalyst showed high performance for the upgrading of RC. The active sites of NM/Ti does not function effective to heavy feedstocks, although the superiority of a Ti02 supported Ni-Mo catalyst for HDS was reported in a recent paper [16]. Mordenite with strong acidity and smaller pore mouth (7 A) than Y-type faujasite (8 A) has often been used for hydroisomerization of light hydrocarbons [3]. Minja and Ternan [17] reported that the incorporation of mordenite into AI2O3 was effective for hydrodemetalization but not for hydrocracking of heavy oil. The strong acidity of the external surface of mordenite is presumably not appropriate for hydrocracking of polycyclic compounds. Table 6 illustrates the reaction results obtained for the hydrotreated feedstocks. Since the first-stage treatment was carried out in a semi-batch autoclave, light fractions produced during the hydrotreatment were removed with the hydrogen flow. This has resulted in the feedstocks for the second stage with relatively small portions of light fractions as shown in Table 2, while the asphalten and heteroatom contents are much reduced. Table 6 evidently displays that the HCK active sites function much more effectively in the reaction of H-VGO than in the reaction of VGO. Further, the M0/HY70 catalyst, HY (70wt%)-Al2O3 supported molybdenum sulfide catalyst without Ni promoter, exhibits higher HDS and HDN activities than the NM/Al and NM/HYQ catalysts. The most refractory nitrogen- or sulfide-containing
122 Table 6 Hydroprocessing of H-VGO and H-RC Catalyst NM/HYo NMmY2 NM/HY7 NM/HYio NM/Al Mo/HY7*^ H-VGO 254 138 277 70 59 59 Tio CC) 362 112 144 400 113 409 T50 ('C) 0.10 0.08 0.07 0.27 0.27 AH/C 4.5 13.5 5.7 3.7 H2 cons, (mg/g) 12.5 12.7 64 86 90 HDN (%) 87 91 91 94 81 74 98 98 79 HDS (%) H-RC 27 51 39 HDN (%) 0.11 0.05 0.08 AH/C 4.2 3.6 2.4 H2 cons. (mg/g)_ * 1: Mo (10 wt%)/HY(70%)-Al2O3 catalyst.
40 50 60 HY activity (%)
52 0.21 4.9
51 0.11 3.7
40 50 60 HY activity (%)
Fig. 1. H2Consumption during the reaction as a function of HYD activity a) O VGO A CL-VGO • SHALE (b) O RC ^ CL • SAND (a) 0 A over NM/Al) (
-
123 compounds are alkyl-substituted dibenzothiophenes or aery dines with steric hindrance [18, 19]. To break the steric hindrance, hydrogenation of the aromatic rings or removal of the alkyl groups is needed [20]. The above results suggest that the dual-functionality with high HYD activities are more suitable for deep HDS. The remarkable reactivity improvements of RC after the first stage treatment are evidenced by the larger AH/C values over the NM/HY7 and NM/HYio catalysts than the other catalysts. Also, HDN and AH/C during the reaction of H-RC increase with the increase of the zeolite content in the support. The H2 consumption in the reaction of RC does not result in the nitrogen removal or the increase in the H/C ratios of the product, while H2 in the reaction of H-RC is consumed to upgrade the liquid product properties. The HCK active sites of the zeolite thus function effectively in the reaction of H-RC. This is in good agreement with the fact that HCK is catalyzed prior to the complete hydrogenation of the aromatic rings over the acid sites [21] and metal sulfides [22 ]. The poor activities of the zeolite supported catalysts for the heavy feedstocks were often attributed to the pore diffusional limitation [3]. However, the molecular sizes of H-RC are not very different from those of RC, since the pre-hydrotreatment does not heavily crack the molecules. The zeolite-supported catalysts gave larger amounts of H2 consumption in the reactions of RC than in the reactions of H-RC in spite of much less total upgrading reactions. This indicates that the HCK of RC is catalyzed in the micropores, while other reactions take place on the external surface of zeolite which is poisoned during the reaction of RC. When the feedstock properties of RC and H-RC are compared, significant differences are observed in the sulfur and asphaltene contents. It is very unlikely that sulfur-containing compounds in RC poison the HCK active sites. The reactivity improvements of the RC by hydrotreatment is presumably due to the reduction of the asphaltene content; the polar nitrogen-containing compounds which are strongly adsorbed on the HCK active sites have been removed in the first-stage hydrotreatment. In the case of VGO, partial hydrogenation of the polyaromatic rings evidenced by the increase of the H/C ratio during the hydrotreatment probably results in a high degree of HCK. These results refer to the importance of hydrotreating reactions prior to the use of solid-acid catalysts. To further discuss the relationship between the feedstock properties and reactivity for heavy feedstocks, the hydroprocessing reactions of synfuels are discussed in the following session. 3.3 Hydrocracking and hydrotreatment of synfuels Fig. 1 shows the relationship between the HYD activity of the catalysts and hydrogen consumption during the reaction which is assumed to be an index for total upgrading reactions. For SHALE, CL-VGO and CL, H2 consumption increases with increasing HYD activity, whereas opposite relationships are observed for RC and VGO. An intermediate trend is observed for SAND. Comparison of the feedstock properties in Table 2 indicates that the reactivity differences among the feedstocks are very likely attributed to the nitrogen concentrations in the feedstock. In the reaction of VGO and RC with low nitrogen concentrations, the acid sites of zeolite catalyze HCK reactions with H2 consumption. In contrast, HYD is the major upgrading reaction for SHALE, CL-VGO and CL with high nitrogen concentrations. Figs. 2-5 show the HDS, HDN, AH/C and HDA activities over the catalysts as a function of HYD activity. It is evident that HDS (Fig. 2) and AH/C (Fig. 4) increase with increasing catalytic HYD activity except for the reaction of VGO over NM/HYQ. These reactions are
124
40 50 60 "^ 40 50 60 HY activity (%) HY activity (%) Fig. 2. HDS during the reaction as a function of HYD activity (a) O VGO • SHALE (b) O RC ^ CL • SAND ( ® A a overNM/Al)
40 50 60 HY activity (%)
40
50 60 HY activity (%)
Fig. 3. HDN during the reaction as a function of HYDactivity (a) O VGO A CL-VGO • SHALE (b) O RC ^ CL • SAND
40
50 60 HY activity (%)
40
50 60 HY activity (%)
Fig. 4. AH/C of feedstocks during the reaction as afiinctionof HYD activity (a) O VGO ^ CL-VGO • SHALE (b) O RC ^ CL • SAND
125
40 50 60 HY activity (%) Fig. 5. HDA during the reaction as a function of HYD activity O RC A CL • SAND
1
H f &
1
^o ^<^c 4h o \ ^ L
A
1 A
CL
^ ^ y^ - • • ^ s n
-
^ • SI
.
SAND\^ ^ ^ 2 ^ "~" 1
1
1
1
_ 1 ^
40 50 60 40 50 60 HY activity (%) HY activity (%) Fig. 6. H2 consumption during the reaction of the hydrotreated feedstocks as a function of HYD activity (a) O VGO A CL-VGO • SHALE (b) O RC ^ CL • SAND
ft j
^^•^Tr~
•
\
SAND " 1
1
1
40 50 60 40 50 60 HY activity (%) HY activity (%) Fig. 7. HDN of the hydrotreated feedstocks as a fiinction of HYD activity (a) O VGO A CL-VGO • SHALE (b) O RC ^ CL • SAND
126 catalyzed solely on the HYD active sites. In other words, the dispersion of the nickelmolybdenum sulfide is of more importance than the dual functionality, as discussed in the previous section. The profiles of HDN activities (Fig. 3) are almost same as the corresponding patterns for H2 consumption except for SAND. This indicates that the total upgrading reaction proceeds in parallel with HDN reactions. The HCK active sites are presumably occupied by nitrogen-containing heavy molecules in the initial stage of the reaction. With the progress of HDN from the heavy molecules, the HCK active sites get recovered andfiinctionfor other upgrading reactions. The HDA of SAND and CL increase with increasing catalytic HYD activity, while that of RC is the highest over the NM/HY70 catalyst (Fig. 5). This discrepancy is attributed to the structural difference of the asphaltenes contained in CL, SAND and RC. The asphaltenes in CL and SAND are converted into hexane-soluble fractions by the saturation of aromatic rings, however, the cleavage of C-C, C-N or C-S bonding is needed, to some extent, to convert the RC asphaltenes. As a general trend observed in Figs. 2-5, high HYD/HCK activity ratios are suitable for the upgrading reactions of heavy feedstocks. Nevertheless, the NM/HYQ catalyst exhibits very low performances in most of the upgrading reactions except for the saturation of aromatic rings. In contrast to the low activities of the NM/HYQ catalyst, the NM/Al catalyst gives much better performances than expected from the HYD function, particularly for HDN reactions (Fig. 3). The dual-fiinctionality of the catalysts, which does not necessarily arise from soUd acidity of zeolite, play very important roles in hydrotreatment as well as HCK, even when the feedstocks are quite heavy. Slightly different trends are observed in the upgrading reaction of the hydrotreated feedstocks as shown in Figs. 6 and 7. The NM/HY7 and NM/HYio catalysts show better performance than the NM/HY2 and NM/HYQ catalysts in the reaction of H-SAND in addition to H-RC and H-VGO, indicative of the increased role of HCK activities in the second-stage reaction. This is attributed to the reduced concentrations of asphaltenes and nitrogen by the hydrotreatment prior to the reaction over the zeolite containing catalysts. The HDN of the HCL and H-CL-VGO increases with decreasing HYD activity, in spite of the decrease in the total H2 consumption. In the reaction of coal-derived liquid with high aromaticity, hydrogen is consumed as a major part for the saturation of the aromatic rings over the HYD active sites, while the HCK active sites participate in the HDN reactions. The pre-hydrotreatment greatly enhances the activities of the HCK sites, however, the NM/Al still exhibits almost same or even higher HDN performance than the NM/HY7 and NM/HYIQ catalysts. This is partly attributed to the present batch reaction system, where the gaseous nitrogen-containing materials such as NH3 heavily inhibits the upgrading reactions over solid acid sites. In the practical flow-type reactors, light basic compounds are not adsorbed on the acid sites because of the low partial pressure and adsorption parameters [23]. 3.4 Future aspects for improvements of dual-functionality As has been described, the dual-fiinctionality of the catalysts play important roles in both HCK and hydrotreatment reactions. The improvements of the dual-fiinctionality is therefore one of the key issues for the developments of better HCK catalysts. In this context, control of the solid acidity of zeolite by modification is of great importance to exclude unfavorable reactions. In the field of modern zeolite technology, dealumination of zeolite has been attracting much attention [15], although many of the studies have been dedicated to the development of FCC catalysts [8]. In addition to the development of high Si/Al zeolite, these
127 studies have pointed out that the dealumination results in the emergence of mesoporosities [24] as well as the formation of amorphous AI2O3 [25], both of which greatly influence the catalytic performance in the hydroprocessing reactions. The mesoporores ranging from 1 to 50 nm reduce the disadvantage of zeolitic microporosoties. Thus formed amorphous AI2O3 provides the zeolite support with favorable sites for high dispersion of nickel molybdenum sulfide [26], while it may cause the misleading relationship between the chemical Si/Al ratio and catalytic performances of the zeolite [27]. Another important issue lies in the elucidation of the interaction between the HYD and HCK active sites. Evidently, they are not independent each other. Some reports have been published to study the interaction of Ni and Mo sulfide and zeolite framework [28, 29]. These indicate that more precise understanding of zeolite structures and properties is essential, tofiartherdevelop the zeolite technology. The present results indicates that the pore dififiision in the zeolitic micropores is not a critical factor in the hydroprocessing reaction of heavy feedstocks. Consequently, the external surface of zeolite particles presumably play more important roles than the inside of the micropores. This indicates that the elucidation of the external surface properties of zeolite is quite important for understanding the zeolite fiinctions. From this point of view, the present authors have recently developed a new method to analyze the depth profile of zeolite particles by the use of X-ray photoelectron spectroscopy with an excitation energy tunable synchrotron radiation X-ray source [30]. Combining the bulk information from, for instance, solid state nmr, detailed understanding of the catalytic properties of zeolite is expected. 4. CONCLUSIONS The main purpose of the present study is to elucidate the roles of dual-fianctionality in the hydroprocessing reactions. Based on the resuhs on the different ratios of HY/AI2O3 supported catalysts, the following conclusions are given. 1. The optimum balance of the HYD/HCK ratio depends on the feedstock properties, particularly nitrogen and asphaltene concentrations. For the upgrading of very heavy feedstocks with the high concentrations of nitrogen and asphaltene, HYD active sites Sanction for HDS and aromatic ring saturation but HCK active sites do not fianction properly. 2. For other feedstocks, dual-functionality play important roles in both HCK and hydrotreatment reactions. Thus, HCK activity which does not necessarily arise from the zeolitic acid sites is essential for most of the upgrading catalysts. 3. Pre-hydrotreatment enhances the performances of the dual-fiinctionality. This is due to a decrease of heavy nitrogen-containing compounds and partial hydrogenation of aromatic rings. 4. In the reaction of heavy feedstocks, the external surface of zeolite functions as the HCK active sites. Thus, limitation by the micropore diffusion is rather less important. REFERENCES 1. 2. 3. 4.
T. Yan, Ind. Eng. Chem. Proc. Des. Dev., 22 (1983) 154. H. Sue, M. Fujita, Oil & Gas J., May 26 (1986) 51. I. E. Maxwell, Catal. Today, 1 (1987) 385. A. Hoek, T. Huizinga, I. E. Maxwell, W. Stork, F. J. van de Meerakker, O. Sy, Oil & Gas J., April 22 (1991) 77.
128 5. 6. 7. 8. 9. 10. 11. 12. 13. 14. 15. 16. 17. 18. 19. 20. 21. 22. 23. 24. 25. 26. 27. 28. 29. 30.
J. W. Ward, Fuel Processing Techno!., 35 (1993) 55. J. K. Minderhoud and J. A. R. van Veen, Fuel Processing TechnoL, 35 (1993) 87. For example, H. Tops0e, B. S. Clausen, F. E. Massoth, Hydrotreating Catalysts, Springer-Verlag, Berlin Heidelberg, 1996. For example, J. Scherzer, Catal. Rev. Sci. Eng., 31 (1989) 215. H. Shimada, T. Sato, Y. Yoshimura, J. Hiraishi, A. Nishijima, J. Catal., 110 (1980) 275. H. Shimada, T. Sato, Y. Yoshimura, A. Hinata, S. Yoshitomi, A. C. Mares, A. Nishijima, Fuel Processing TechnoL, 25 (1990) 153. S. H. Yang and C. N. Satterfield, J. Catal., 81 (1983) 168. N-Y. Tops0e and H. Tops0e, J. Catal, 139 (1993) 641. J. M. Levis, R. A. Kydd, P. M. Boorman, P. H. van Rhyn, Applied Catal., A84 (1992) 103. R. S. Mann, I. S. Sambi, K. C. Khulbe, Ind. Eng. Chem. Res., 27 (1988) 1788. P. L. Arias, J. F. Cambra, M. B. Guemez, J. A. Legarreta, B. Pawelec, J. L. G. Fierro, Bull. Soc. Chim. Belg., 104 (1995) 197. K. Segawa, M. Katsuta, F. Kameda, Catal. Today, 29 (1996) 215. R. J. A. Minja and M. Ternan, Energy & Fuels, 5 (1991) 117. M. J. Girgis, B. C. Gates, Ind. Eng. Chem. Res., 30 (1991) 2021. T. Kabe, A. Ishihara, H. Tajima, Ind. Eng. Chem. Res., 31 (1992) 1577. X. Ma, K. Sakanishi, I. Mochida, Ind. Eng. Chem. Res., 33 (1994) 218. A. T. Lapinas, M. T. Klein, B. C. Gates, A. Macris, J. E. Lyons, Ind. Eng. Chem. Res., 26 (1987) 1026. Y. Miki, Y. Sugimoto, Fuel Processing TechnoL, 43 (1995) 137. V. LaVopa and C. N. Satterfield, J. Catal., 110 (1988) 375. L. E. Maier, P. H. Bigeard, A. Billon, p. Dufi-esne, NPRA Annual meeting (1988). R. Fang, G. Harvey, H. W. Kouwenhoven, R. Prins, 130 (1995) 67.1 J. LegUse, A Janin, J. C. Lavalley, D. Cornet, J. Catal., 114 (1988) 388. R. D. Bezman, Catal. Today, 13 (1992) 143. R. Cid, F. J. Gil Llambias, M. Gonzalez, A. L. Agudo, Catal. Lett., 24 (1994) 147. J. Leglise, J. M. Manoli, C.Potvin, G. Djega-Mariadassou, D. Comet, J. Catal., 152 (1995)275. H. Shimada, N. Matsubayashi, M. Imamura, T. Sato, A. Nishijima, Catal., Lett., 39 (1996) 125.
® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
129
Hydrocracking o f C i o hydrocarbons over a sulfided N i M o / Y zeolite catalyst.
J.L. Lemberton^, A. Baudon^, M. Guisnet^ N. Marchal^ and S. Mignard^ ^ Laboratoire de Catalyse en Chimie Organique, URA CNRS 350, Universite de Poitiers, 40 avenue du Recteur Pineau, 86022 Poitiers Cedex, France b Institut Fran^ais du Petrole, 1 et 4 avenue de Bois-Preau, BP 311, 92506 Rueil-Malmaison Cedex, France
Abstract The transformation of different CJQ model molecules (n-decane, n-butyl cyclohexane, tbutyl cyclohexane, n-butyl benzene and tetralin) was studied on a sulfided NiMoA^ zeolite catalyst under industrial conditions : 380°C, 5.7 MPa hydrogen pressure, presence of hydrogen sulfide and ammonia. Depending on the molecule, different reactions were observed, catalyzed either by the acid sites, or by the hydrogenating sites, or by both sites (bifixnctional mechanism). Isomerization could concern the chain or the cycle, and occurred always through a bifimctional mechanism. The formation of the light products was mainly consecutive to isomerization, and consequently occurred also through the bifimctional mechanism. However, some light products were formed through condensation-cracking reactions. Purely acid-catalyzed dealkylation of the alkyl aromatics was also observed, whereas the hydro-dehydrogenating fiinction of the catalyst was responsible ft)r the hydrogenation of the aromatics and the aromatization of the naphthenes.
1. BVTRODUCTION In a modem refinery, hydrocracking is complementary to catalytic cracking : polyaromatic compounds can be transformed through hydrocracking, which cannot be through catalytic cracking [1,2]. Moreover, the great versatility of hydrocracking makes it possible to equilibrate the supply (gasoline, gasoil,...) and the demand. Hydrocracking feeds contain various hydrocarbons, such as paraffins, naphthenes or aromatics, as well as sulfiir and nitrogen-containing compounds. Catalysts are bifimctional, associating a hydrodehydrogenating fiinction (sulfides of groups VI and VIII metals) with an acidic one (zeolite) [1-3]. On these catalysts, the transformation of hydrocarbons involves reaction steps both on hydrogenating sites (hydrogenation and dehydrogenation), and on acid sites (rearrangement, cracking). Consequently, the hydrocracking scheme can most likely depend on the nature and on the structure of the hydrocarbon. The aim of this work was to establish under industrial conditions the scheme of transformation of several CJQ hydrocarbons, representative for the hydrocracking feed families.
130 namely n-decane (paraffins), n-butyl cyclohexane and t-butyl cyclohexane (naphthenes), n-butyl benzene (aromatics) and tetralin (naphtheno-aromatics).
2. EXPERIMENTAL The hydrocracking of the C\Q hydrocarbons was carried out in a flow reactor at 380°C under a 6 MPa total pressure. Dimethyl disulfide and aniline were added to the hydrocarbon in order to generate respectively H2S and NH3. The standard reaction conditions were : catalyst weight 1.75 g, hydrocarbon contact time 0.18 mn, hydrogen/hydrocarbon molar ratio 20, p hydrocarbon = 0.285 MPa, p H2 = 5.7 MPa, p NH3 = 5 kPa, p H2S = 6.1 kPa. The catalyst was first sulfided in the reactor by a dimethyldisulfide/aniline/n-heptane mixture, under the same flow rates and the same pressures as those used for the C|o hydrocarbon hydrocracking reaction. The sulfiding feed was injected starting at 150°C, then the temperature was raised to 285°C (1-hour stage), 325°C (1-hour stage), 350°C (2-hour stage) and finally to 380°C (reaction temperature). After about 2 hours at 380°C, the catalyst exhibited a stable activity for n-heptane hydrocracking. The sulfiding feed was then replaced by the reaction one, containing the choosen C^Q hydrocarbon. All the reaction products were analyzed on-line by gas-liquid chromatography (Varian 3400) on a 50-m CP-Sil5 capillary column (Chrompack) with a temperature programming firom 40to70°C(5°C.mn-l). Special experiments were carried out to measure the amount of carbon deposited on the spent catalyst: after sulfidation and stabilization of the catalyst, the Cjo hydrocracking reaction was performed during 6 hours, then the feed injection was stopped and the catalyst was cooled down rapidly under a small hydrogen flow rate (3.5 l.h-1) in order to eliminate the feed remaining in the reactor. The carbon content of the catalyst was then measured by the "Service Central d'Analyses CNRS".
3. RESULTS AND DISCUSSION
Figures 1 to 7 show the distribution of the reaction products as a fijnction of the conversion of each hydrocarbon. The different conversion values were obtained by changing the contact time on the stabilized catalyst. n-Decane transforms into isomerization products and light products (figure 1). Isomerization products are monobranched isomers M (methyl nonanes, plus some ethyl octanes and propyl heptane) and multibranched isomers B (mainly dimethyl octanes). Figure 1 shows that both M and B appear as primary reaction products. Light products C (propane, butanes, pentanes, hexanes and heptanes) are secondary products, and consequently are formed fi-om the isomers. Moreover, these products are highly branched (the iso/normal ratio i/n is about 2 for the C5 and C5 molecules, and 1 for the C4), which indicates that they come mainly fi-om the multibranched isomers. Consequently, a parallel-consecutive bifimctional reaction can explain the distribution of n-decane transformation products (scheme 1). The isomerization reaction proceeds through protonated cyclopropanes [4,5], whereas cracking involves the cleavage of a C-C bond in B position of a carbocation [6-8]. The fact that multibranched isomers B appear as primary products means that the add function of the catalyst is not well balanced by the
131 hydrogenating one : several acid sites may be encountered by the olefinic intermediates during their difiiision between two hydrogenating sites [9-11]. On the other hand, the C3/C7 and C4/C6 molar ratios are lower than unity. This excess in the amounts of C7 and C5 molecules indicates that condensation-cracking reactions occur, for example between an olefin produced by cracking and a carbocation : €3"^ + Cio"*" ~> C13 -^ C7 + C^. This type of reaction can also participate in the formation of the isomerized products, but only to a small extent since isomerization is a primary reaction, but not the formation of the light products.
40 1
.c
30
C3 • C10
\/s/\/\/
Ns 20 -
/
^*""~^M \y\/sy\^\
10 m— •
—H
•
6
0 -
()
60
40
20
C4 • C
^
•
C3 C7
X(n-Cw)
Figure 1 Distribution of n-decane transformation products
Scheme 1
Isomers I are also the primary products of n-butyl cyclohexane transformation (figure 2). They are mainly dialkyl cyclopentanes such as methyl butyl cyclopentane, as well as dialkyl cyclohexanes (methyl propyl cyclohexane), these latter being most likely formed through isomerization of the former [12]. Small amounts of aromatisation products A (alkyl benzenes) are also observed as primary reaction products (figure 2).
80
J
/C
60 •^ 40:
\
^^,^—^"B
"SI
• ^
20
^^r'^Tn^n-f^— •
0 4 ' *••
0
^ V 6 - o -'
20
-'—r—=^^^ 40
60
*
•
—
—
-^A.
1— -^-=—1 80 100
X (n-BCH)
Figure 2. Distribution of n-butylcyclohexane transformation products
^
o Scheme 2
\
132 Cracking (C) is a secondary reaction, yielding mainly butanes and cyclic C5 molecules : cyclohexane and methyl cyclopentane (no significant ring opening is observed), which indicates that a breaking of the alkyl chain occurs. Since cracking is a secondary reaction, the direct breaking of this chain does not occur, and the chain must isomerize through a protonated cyclopropane before cracking (scheme 3) :
o
^
o
Scheme 3
K
0
The isomerization of the n-butyi carbocation into a i-butyl one is most likely the fastest reaction, since it transforms a secondary carbocation into a tertiary one (the isomerization of the secondary i-butyl carbocation into the tertiary one occurs through a very rapid hydride shift). Moreover [6-8], the cracking of the i-butyl carbocation (tertiary -> secondary) is also faster than the cracking of the s-butyl carbocation (secondary -> secondary). This accounts for the fact that n-butane is formed in smaller amounts than i-butane : the iC4/nC4 ratio is about 4, whatever the n-butyl cyclohexane conversion. Then, n-butyl cyclohexane transforms through two parallel reactions (scheme 2) : isomerization of the cycle or isomerization of the chain followed by cracking. On the other hand, the C4/C5 molar ratio is lower than unity, which indicates there are condensation-cracking reactions, as was the case with n-decane. The products formed from t-butyl cyclohexane (figure 3) are isomers I (dialkyl cyclopentanes and some traces of dialkyl cyclohexanes), cracking products C, practically only C4 (mainly i-butane, iC/^nC4 = 20) and C6 (mainly methyl cyclopentane), and traces of t-butyl benzene A. Figure 3 indicates that isomerization is again a primary reaction, and that cracking is a secondary one. Consequently, as was the case with n-butyl cyclohexane, there is no direct breaking of the alkyl chain. Since the cracking products are almost entirely i-butane and methyl cyclopentane, the scission of the t-butyl chain occurs most likely after the isomerization of the cycle into a methyl cyclopentane (scheme 4). This can be explained by the nature of the intermediate carbocations. Direct cracking would transform a secondary carbocation into a tertiary one, whereas cracking after isomerization would transform a tertiary carbocation into a tertiary one (scheme 5). The latter is reported to be the fastest cracking reaction [6-8]. Contrary to what was observed with the two preceeding molecules, the C4/C5 molar ratio is
133 close to unity, which shows that condensation-cracking reactions, if they exist, do not play here an important role.
^ Figure 3 Distribution of t-butylcyclohexane transformation products
cr— (2f-. Scheme 4
cr^
o Scheme 5
n-Butyl benzene transformation products are quite different. Indeed one can observe the n-butyl benzene hydrogenation product H (n-butyl cyclohexane), the isomers of the hydrogenated product (HI, dialkyl cyclopentanes and dialkyl cyclohexanes), crackmg products C, and isomerization products I (dialkyl benzenes). Figure 4 and 5 show that H, HI and I are apparently primary products, whereas cracking products are secondary reaction products. However, it is clear that the isomers HI result from the isomerization of n-butyl cyclohexane, and are consequently consecutive to the hydrogenation of the aromatic ring. The n-butyl cyclohexane produced (scheme 6) transforms in turn according to scheme 2, this reaction being very rapid. The cracking products are mainly cyclic C5 (benzene, methyl cyclopentane and cyclohexane) and butanes, n-butane being now formal in greater amounts than i-butane (iC4/nC4 = 0.6). This increase in the n-butane formation can be related to the formation of significant amounts of benzene. Indeed (scheme 7), n-butane and benzene can be produced by the dealkylation of n-butyl benzene, this reaction occurring through a purely acid mechanism. However, a direct dealkydation of n-butyl benzene would yield a primary butyl carbocation. The dealkylation probably would occur after the isomerization of the chain through a protonated cyclopropane (by the same mechanism as in scheme 3). However, one can suppose than dealkylation occurs only with s-butyl benzene, since the i-butyl isomer would also produce a primary carbocation (sciieme 7). Lasdy, the C4/C5 molar ratio is about 1, which indicates that the participation of condensation-cracking reactions to the formation of the cracking products seems negligible.
134 40
20-1
'HI
A C
30 •
X
^ 10
^ 20
10 J •9i-^^
0 -1
()
20
X
c- H
1 1— 40 60 X (n-BBz)
j^^^P^ 100
80
0
20
•
*
*
1 1 40 60 X(ii-BBz)
:
1 80
1 100
Figure 4 Figure 5 Distribution of n-butylbenzene transformation products
+
nC4
(SCHEME 2)
Scheme 6
o
— Co:
^<
Scheme 7
Finally, the transformation of tetralin yields many different products (figures 6 and 7) : isomers I (methyl indans), hydrogenation products H (decalins), aromatics A (naphthalene, methyl or dimethyl naphthalenes), isomers of the hydrogenated product HI (formed through isomerization of one or two cycles of the decalins into methyl cyclopentanes), alkyl tetralins Alk (methyl or dimethyl tetrdins), and cracking products C with either 10 carbon atoms, resulting from the opening of a cycle (n-butyl benzene and n-butyl cyclohexane), or less than 10 carbon atoms (mainly cyclohexane, methyl cyclopentane and butanes, the iC4/nC4 ratio being about 1.5). Isomerization products, hydrogenation products and naphthalene are primary reaction products of tetralin, whereas the hydrogenated isomers and the cracking products are secondary products, which suggests the reaction scheme 8. The dehydrogenation of tetralin into naphthalene is initially much favoured when compared with the hydrogenation reaction (figure 6). The naphthalene/tetralin ratio is about 0.2, i.e. close to the value at the thermodynamical equilibrium [13]. The hydrogenated isomers may be produced either through isomerization of the hydrogenation products, or through hydrogenation of the isomers. n-Butyl benzene probably is not formed through the opening of the saturated ring of tetralin, since this reaction would produce an unstable primary carbocation. It is most likely formed from 2-methyl indan through a secondary carijocation (1-methyl indan would also produce a primary carijocation). n-Butyl cyclohexane can be obtained through hydrogenation of n-butyl benzene (as in scheme 6), but also through the opening of one saturated ring in the decalins. n-Butyl cyclohexane
135 transforms in turn as shown in scheme 2. The C4/C5 molar ratio is very small (0.1-0.3 in the 20-80% tetralin conversion range), which indicates that condensation/cracking reactions also participate in the formation of the cracking products. Lastly, the methyl (or dimethyl) tetralins (or naphthalenes) result probably from a reaction between tetralin or naphthalene and the methyl fragments produced by dimethyl disulfide decomposition [14]. •HI
40
30
N? 20-
V H
^KJ^
10-
»^7~-^
>->.' A
0-
<}
20
40
60
80
100
X(THN)
Figure 6 Figure 7 Distribution of tetralin transformation products
\ —
Scheme 8
0:5 - 0 0
-
(SCHEME 2)
(SCHEME 6)
4. CONCLUSION Several reactions were observed during the hydrocracking of the model Cjo hydrocarbons, namely the isomerization or the cleavage of C-C bonds, the hydrogenation of aromatic rings or the dehydrogenation of saturated rings. Isomerization is always a btfunctional reaction, occuring through protonated cyclopropanes, and concerns the cycle or the chain, depending on the nature of the hydrocarbon : isomerization of the cycle of t-butyl cyclohexane, isomerization of the chain of n-butyl benzene, both reactions with n-butyl cyclohexane. The cleavage of the C-C bonds can also occur through a bifunctional mechanism. In this case, cracking is consecutive to isomerization, which means that the acid function of the catalyst is not balanced by the hydrogenating one. Bifimctional cracking concerns mainly the chain, since no significant ring opening is observed. Light molecules can also be formed through condensation-cracking reactions, but also through an acid-catalyzed dealkylation (n-butyl
136 benzene). The last type of reaction observed is catalyzed only by the hydro-dehydrogenating sites of the catalyst : hydrogenation of the aromatic ring of aromatic molecules and dehydrogenation of the saturated ring of the naphthenes. It is clear that the structure of the hydrocarbon modifies very much the rates of the dififerent reactions, and consequently the selectivity of the hydrocracking reaction. Precise kinetic measurements would allow to estimate each reaction rate, and to establish the relationship between the structure of the hydrocarbon and the selectivity for hydrocrackmg.
REFERENCES 1.
J.H. Gary and G.E. Handwerk, in << Petroleum Refining, Technology and Economics », p. 149, Dekker, New-York, 1994
2.
A.P. Bolton, in « Zeolite Chemistry and Catalysis », J.A. Rabo (ed.), p.714, American Chemical Society, Washington, DC, 1976
3.
M. Guisnet, F. Alvarez, G. Gianetto and G. Perot, Catal. Today, 1 (1987) 415
4.
F. Chevalier, M. Guisnet and R. Maurel, in « Proceedings 6th International Congress on Catalysis », G C Bond et al. (eds), I, p.478. The Chemical Society, London, 1977
5.
D.M. Brouwer, Rec. Trav. Chim., 87 (1968) 1435.
6.
J. Weitkamp, J.A Martens and P.A. Jacobs, Applied Catal., 8 (1983) 123
7.
F. Alvarez, F.R. Ribeiro, G. Perot, C. Thomazeau and M. Guisnet, J. Catal., in press.
8.
P.A. Jacobs and J.A Martens, in «Introduction to Zeolite Science and Practice », H. Van Bekkum et al. (eds.), Elsevier, Studies in Surface Science and Catalysis, 58 (1991)445.
9.
G. Gianetto, G. Perot and M. Guisnet, Ind. Eng. Chem. Prod. Res. Dev., 25 (1986) 481
10.
G. Gianetto, F. Alvarez, F.R. Ribeiro, G. Perot and M. Guisnet, in « Guidelines for Mastering the Properties of Molecular Sieves », D. Barthomeuf et al. (eds.), NATO ASI Series, Plenum, New-Yoric, 221, 1990, 355
11.
M. Guisnet, C. Thomazeau, J.L. Lemberton and S.Mignard, J. Catal., 151 (1995) 102
12.
J. Weitkamp, P A Jacobs and S. Ernst, in « Structure and Reactivity of Modified Zeolites », P.A. Jacobs et al. Eds., Elsevier, 18 (1984) 279
13.
C.G. Frye, J. Chem. Eng. Data, 14 (1962) 372
14.
J. Leglise, A Janin, J.C. Uvalley and D.Comet, J. Catal., 114 (1988) 388
1997 Elsevier Science B. V. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
137
Novel Hydrotreating Catalysts Based on Synthetic Clay Minerals R.G. Leliveld, W.C.A. Huyben, AJ. van Dillen, J.W. Geus and D.C. Koningsberger Department of Inorganic Chemistry, Debye Institute, University of Utrecht, P.O. Box 80083, 3508 TB Utrecht, The Netherlands Saponites with Co, Ni, Zn, Mg and combinations of Co and Mg in the octahedral layer were synthesised. Their applicability as catalysts and supports for hydrotreating catalysts was investigated. The stability of these clays in a sulfidic environment was studied using Temperature Programmed Sulfidation. The catalytic performance of the saponites both bare and impregnated with Co/Mo was tested in the hydrodesulfurisation of thiophene. 1. Introduction In modem refineries hydrocracking of heavy oil feedstocks to high-value transportation fuels is a flexible process dependent on feed and desired product selectivity. Most catalysts currently used are bifunctional consisting of a hydrogenating component, such as noble metals or transition metal sulfides, together with an acidic function usually incorporated in the support [1-2]. One of the first hydrocracking catalysts to be commercially used was a supported Fe montmorillonite catalyst developed by ICI in the late 1930's [3]. Nowadays almost all hydrocracking catalysts are based upon zeolites due to their favorable acidic and textural properties [4], However, the zeolite structure can lead to pore exclusion of the large molecules present in heavy crudes. The limited size of zeolite pores has led to a renewed interest in acidic clay minerals as an alternative support since the platelike structure of clays might avoid transport limitations [5-6]. Clays belong to the group of phyllosilicates: layered or twodimensional silicates build of octahedral and tetrahedral sheets as represented in Figure 1.
Figure 1. Structure smectite Hx-Mg3Si4.xAlxOio(OH)2: T, tetrahedral and O, octahedral sheet The varying and difficult to control texture and composition of natural clays and the presence of impurities are severe drawbacks to the use as catalysts [7-9]. Recently, our group reported on a novel route to synthesize a class of 2:1 trioctahedral clays called saponites [10] at
138 nonhydrothermal conditions. In saponites tetrahedral Si atoms are partly substituted by Al atoms, while the octahedral layer contains divalent cations. The new synthesis route allows careful control of the clay's texture and acidity. This study is directed towards the aptness of synthetic saponites with differing octahedral ions as catalysts and as support for hydrocracking catalysts with emphasis on their resistance against sulfur. For this reason, saponites with Co, Ni, Zn, Mg and combinations of Co and Mg incorporated in the octahedral sheet were synthesised and studied with Temperature Programmed Sulfidation and XRD/TEM. The catalytic activity of the bare saponites and of Co/Mo impregnated samples were tested in the hydrodesulfurisation of thiophene. A good hydrocracking catalysts requires a well dispersed hydrogenating function to prevent excess cracking on the acidic sites. The performance in Thiophene hydrodesulfurisation (HDS) can offer insight in the dispersion of the metal sulfides on the surface of the clay support [11]. 2. Experimental Catalyst preparation Saponites with Mg, Zn, Co and Ni as octahedral ions and mixtures of Mg and Co were synthesised at 363 K and 1 atmosphere from a Si/Al gel (Si/Al ratio 12.3) and a solution containing urea and the M ^-nitrate [10]. After ion-exchange with ammonium (IM NH4CI solution) the saponites were dried at 393 K and calcined at 723 K. The samples are denoted by the octahedral ions followed by their mutual ratio e.g. Co:Mg 1:11 saponite. Mo and Co impregnated catalysts were prepared by incipient wetness impregnation of the saponite supports with aqueous solutions containing the required amounts of (NH4)6Mo7024-6 H2O and Co(N03)3- 6 H2O. The impregnated saponites were subsequently dried in an air flow at 298 K, dried in stagnant air at 393 K and calcined at 723 K. Characterisation Catalyst samples were examined with X-ray powder Diffraction (XRD) with an Enraf Nonius FR590 diffractometer using CoKa radiation. Transmission Electron Microscopy (TEM) was performed with a Philips EM-420 instrument operated at an accelerating voltage of 120 kV. Diffuse Reflectance Infrared Fourier Transform (DRIFT) spectra were recorded on diluted samples (10% in KBr) using a Perkin Elmer 1600 series spectrometer. Extended X-ray Absorption Fine Structure (EXAFS) measurements were performed at EXAFS station 9.2 of the SRS at Daresbury (U.K.). The saponite samples were pressed into self-supporting wafers and mounted in an in-situ EXAFS cell [12]. The spectra were recorded for the Co-K edge at liquid nitrogen temperature. Details of the data analysis have been described elsewhere [1314]. Phase shifts and backscattering amplitudes from reference compounds were used to calculate the EXAFS contributions: CoO for the Co-Co and Co(OH)2 for the Co-0 scattererbackscatterer pair. The Co-Mg and Co-Si references were theoretically calculated with the FEFF-3.1code[15]. Temperature-Programmed Sulfidation and Hydrodesulfurisation of Thiophene Temperature Programmed Sulfidation profiles were recorded in an automated microflow apparatus. The samples (1.0 ml, sieve fraction 150-425 |im) were sulfided in a 100 ml/min flow of H2S/H2/Ar (10/40/50). After 30 min at room temperature the temperature was linearly raised (5K/min) to 673 K. The H2S concentration before and after the reactor was monitored with a Varian UV/Vis spectrophotometer at X=232 nm. After 30 min at 673 K the flow was
139 switched to 51.2 ml/min thiophene/H2/Ar (2.4/87.9/9.6). Analysis of the reactor effluent was done with a gaschromatograph using a Chrompack CP-sil-5 CB column and FID. The catalytic activity in the conversion of thiophene was measured at temperatures between 673 and 423 K after 6 hr on stream at 673 K (atmospheric pressure). 3. Results 3.1 Characterization After synthesis for 20 hours and calcination at 723 K all saponites were characterised with XRD, TEM and nitrogen physisorption. XRD confirmed that in all cases a 2:1 clay triooctahedral clay was formed as indicated by the presence of the characteristic d(060) reflection at about 1.54 A [16]. TEM showed the saponites to consist of small platelets with an open 'house of cards' structure, as shown in Figure 2a for a Co saponite. The size of the platelets varied from 300 nm for the Zn-saponite to 25-40 nm for the Mg, Ni and Co-clays. The B.E.T. surface areas reflected the size of the platelets varying from 105 mig (0.14 ml/g) for the Zn-saponite to 570 m /g (0.46 ml/g) for the Mg sample. The surface areas of the mixed Co-Mg saponites were intermediate between the pure Mg and Co saponite (410 m^/g). The structure of the mixed Co-Mg saponite platelets was quite uniform as revealed by TEM. There was no indication of separate Co and Mg platelets, as EDX confirmed the presence of both Co and Mg within the same clay platelet.To study the distribution of the Co and Mg cations within the octahedral layer the samples were characterised with DRIFT and EXAFS.
•290 nm • • B i t ^ ^ ^
130 nm
Figure 2. TEM photograph of a fresh (a) and sulfided (b) Co-saponite Si/Al=12.3 DRIFT and EXAFS The hydroxy 1 stretching range between 3800 and 3100 cm"^ of the pure and mixed Co-Mg saponites is represented in Figure 2. All spectra were recorded at 673 K under an Ar flow to minimise the amount of water present within the saponite samples. The spectra contained two bands referred to as Vi(OH) and V2(0H). The position of the V2(0H) band at 3723 cm"^ was independent of the Co-Mg ratio except for the pure Co saponite. The broad Vi(OH) band varied from 3762 to 3615 cm' for the Mg and Co saponite, respectively. The Co saponite and Co-Mg saponites with Co:Mg=l:29 and 1:1 were characterised with EXAFS of the Co-K edge. The absolute part of the k weighed Fourier Transform of the three samples,as shown in Figure 3, exhibited two shells. The first shell between 1 and 2 A (phase
140 uncorrected) is due to the nearest six oxygen neighbours of the Co ions in the octahedral layer. The second shell between 2.2 and 3.4 A is made up from neighbouring octahedral cations and the Si/Al atoms of the tetrahedral sheet. When the amount of Mg in the octahedral layer increased the amplitude of the second shell drasticly reduced. This reduction is due to a destructive interference of the Co-Co and Co-Mg scatterer-backscatter contributions [17]. The results of the data analysis results are presented in Table 1. 12
-Co •Co-Mg 1:11 Co-Mg 1:29 J
10
;:3
§ u
H
t-H
o
(U
4
tin
2
^^ 3625
3450
3275
6
*§ o
3800
8
0
3100
Wavenumber (cm") Figure 3: DRIFT spectra of Co-Mg saponites Figure 4. k^ weighted Fourier Transform The number of neighbours and the interatomic distances for pure Co-saponite were in accordance with the known crystal structure [17]. As Si and Al are neighbours in the periodic table the backscattering phase and amplitude are similar justifying a fit of the k weighted Fourier Transform with 4 Si neighbours at 3.27 A. The fit of the Co-Mg samples incorporated the expected Mg neighbours at a slightly shorter distance than the Co ions. Table 1 Data Analysis Results Co-K EXAFS of Co-Mg saponites, R-space fit, k Fourier Transform scatterer Co-saponite O Co Si Co-Mg 1:1 saponite 0 Mg Co Si Co-Mg 1:29 saponite 0 Mg Co Si
N
Aa'(10-^ A^)
R(A)
AEo(eV)'
6.1 6.2 4.1
-3 2 33
2.09 3.11 3.27
-0.4 1.4 5.3
6.0 2.5 2.5 4.1
-7 104 168 -2
2.08 3.02 3.14 3.27
0.4 1.0 1.5 2.6
6.0 5.8 0.4 4.5
17 6 33 -13
2.05 3.12 3.15 3.23
2.3 1.5 2.3 8.8
^^ The estimated accuracy of the fit parameters is 20 % for N and Aa^ and 1% for R and AE,
141 3.2 Temperature Programmed Sulfidation Figure 5 shows the H2S absorption per unit weight for the pure Mg, Zn, Co and Ni saponites.
273
373
473
573
673
Temperature (K)
Figure 5. TPS of Co, Ni, Zn and Mg saponite
273
373
473
573
Temperature (K)
673
Figure 6. TPS of Co and Co-Mg 1:1 saponite
Except for Zn, all saponites showed a desorption of H2S between 298 and 343 K due to release of physisorbed H2S. When the temperature was raised both the Co, Ni and Zn saponites consumed a large amount of H2S. The consumption of H2S was not completed at 673 K. When the H2S absorption was monitored isothermally at 673 K, the consumption quickly decreased. Only when the temperature was further raised above 673 K an extra H2S uptake was observed, indicating the extent of sulfidation to be predominantly temperature controlled. The Mg-saponite does not show any uptake or release of H2S at temperatures higher than 353 K pointing to a stable oxidic structure. Figure 6 presents the TPS profiles of the Co and Co-Mg 1:1 saponite. The H2S uptake per Co atom for the mixed Co-Mg saponite was substantially less than for the Co-saponite. This points to a stabilisation of the Co atoms in the clay lattice of the Co-Mg saponites relative to the pure Co saponite. 1
d(060)
d(060)
9"
^ •S
I
sS
_
D
k 1 11
v^—^
R
L-^.,,.^ A
50
20
100
150
Figure 7. XRD of fresh and sulfided Mg (A and B ) and Co saponite (C and D)
50
100
150
20
Figure 8. XRD of fresh (A) and sulfided (B) Co-Mg 1:1 saponite
142 The XRD spectra of a fresh Mg and Co saponite and those after sulfidation are shown in Figure 7. The XRD pattern of the Mg-saponite after sulfidation did not change with respect to the fresh saponite whereas for Co it was quite different. The characteristic clay reflections at 1.54 A (Tl"* 20) and 2.63 A (40'' 29) practically disappeared and two reflections at 1.74 A (62° 29) and 2.96 A (35° 29) characteristic of Co9Sg evolved. In agreement with the TPS results, XRD patterns of sulfided Ni and Zn saponites similarly showed the presence of crystalline Ni3S2 and P-ZnS, respectively, and a loss of the clay structure. Figure lb contains the TEM photograph of the sulfided Co-saponite. The small clay platelets disappeared and instead large crystalline particles were formed together with amorphous silica-alumina. In contrast, XRD spectra of the fresh and sulfided Co-Mg 1:1 saponite are shown in Figure 8. After sulfidation the 2:1 clay structure was still present as indicated by the d(060) reflection and confirmed by TEM. The sample suffered some loss of crystallinity as indicated by the broadening of the clay reflections leading to a decrease in intensity. 3.3 Thiophene Hydrodesulfurisation The thiophene conversion of the bare Co,Ni, Zn and Mg saponites and those impregnated with 7.5 wt% M0O3 is presented in Figure 9a and 9b, respectively. The usual reaction products were observed, viz., 1-butene, cis- and trans-butene, butane and H2S. Additionally the Co and Ni-catalysts produced substantial amounts of iso-butane (up to 8 %). As shown in Figure 9a Zn and Mg saponites did not exhibit any HDS activity. The Ni and Co-saponite, however, were quite active with conversions of 77 and 42 %, respectively. When Mo was impregnated the Mg-saponite became active exhibiting a conversion of 21 % at 673 K. Addition of Mo to the Co-saponite lowered the activity from 42% to 20%. The performance of the Ni-saponite remained more or less unchanged when promoted with Mo (conversions of 77 VS.82 % ) .
: -^Nj
100
J
e^ 50
.2
_ -^e-Mg
j /
]
^
0
0
u 0 373
473
573
Temperature (K)
673
373
473
573
673
Temperature (K)
Figure 9. Performance of bare Co, Ni, Zn and Mg saponites (a) and impregnated with Mo (b) The conversion curves of the mixed Co-Mg saponites are shown in Figure 10a. As could be expected the activity of the saponites increased with the amount of Co in the octahedral layer. The drastic increase in activity going from the Co-Mg 1:11 to the 1:1 saponite is explained by the amount of Co rising from 4 wt% to 21 wt%. Figure 10b compares the HDS activity of a 4 wt % Co impregnated Mg-saponite with that of the Co-Mg 1:11 saponite.
143 Hardly no difference in activity is observed with both catalysts. Clearly, the addition of Mo substantially increased the activity, the co-impregnated CoMo/Mg catalyst being slightly more active than the Mo impregnated Co-Mg saponite. 100 ^ 75 o
50
-Co-Mg Co/Mg -Mo/Co-Mg -CoMo/Mg - CoMo-stev
o 25 O
(A)
373
473
573
Temperature (K)
673
373
(B)
- alumina
473
573
673
Temperature (K)
Figure 10. Performance of mixed Co-Mg saponites (A) and saponites impregnated with CoMo (B). Impregnated catalysts are denoted by the impregnated metal(s) / saponite support. During the 6 hr on stream at 673 k the conversion of the saponites stabilised from conversions of 100% to the values presented here. Simultaneously the selectivity towards ibutane and C1-C3 products decreased. After oxidative regeneration and subsequent sulfiding the activity could be restored to the initial 100%. For reasons of comparison a non acidic CoMo impregnated Mg stevensite was prepared. The Mg-stevensite (no Si/Al exchange, some vacancies in the octahedral layer but no acidity) indeed showed no deactivation and the performance is comparable to that of a conventional CoMo/ y-Al203 catalyst (13 wt %> M0O3, 4.3 wt % C03O4). 4. Discussion 4.1 Distribution Co-Mg in mixed Co-Mg saponites The Vi(OH) and V2(0H) bands in the DRIFT spectra of the saponites can be attributed to the hydroxyl group of the M^^-(OH) units in the octahedral sheet and the Si-OH groups present at the edges of the tetrahedral sheets, respectively [18-19]. The position of the Vi(OH) band depends on the nature of the three coordinating octahedral ions. With two differing octahedral cations as with Co and Mg the Vi(OH) band is expected to split in four separate bands: Mg3-0H, Mg2Co-OH, MgCo2-OH and C03-OH. A random distribution of the octahedral ions with a ratio of 1:1 should result in a relative intensity of the bands of 1:2:2:1 [20]. Due to line broadening our spectra only show one broad Vi(OH) band. In the spectrum of the Co-Mg 1:1 saponite the frequency of this band is fairly between those of the pure Mg and Co-saponite. Together with the symmetric shape of the band covering the region of the individual Vi(OH) bands this points to a homogeneous distribution of the octahedral ions. More detailed insight can be obtained from the EXAFS data. As can be seen from Table 1 the coordination numbers are also in accordance with a homogeneous distribution of the Mg and Co ions in the octahedral sheet. The coordination numbers Nco-co ^^^ ^Co- Mg of ^-4 and
144 5.8 for the Co-Mg 1:29 saponite are close to the expected values of 0.2 and 5.8, respectively. Moreover, in the Co-Mg 1:1 saponite the coordination numbers Nco-co and Nco-Mg are equal showing that each Co ion is surrounded by an equal number of both Mg and Co ions. If domain segregation had taken place during synthesis the ratio Nco-co/Nco-Mg would be higher than the ratio of 1.0 calculated from the stoichiometric amounts. The sum of both coordination numbers is not exact 6, but within the limits of accuracy. AUtogether, based on the DRIFT and EXAFS results, it can be concluded that in the mixed Co-Mg saponites the Co and Mg ions are uniformly distributed in the octahedral sheet. 4.2 Sulfur Resistance To apply synthetic saponites as hydrocracking catalysts the clay lattice should be stable in an environment where rather large quantities of sulfur are present. The XRD/TEM and TPS profiles of the different saponites clearly indicate that when the octahedral layer only consists of Co, Ni or Zn the clay structure collapses upon exposure to sulfiir and the bulk sulfides CogSg, Ni3S2 and P-ZnS are formed. The Mg-saponite, however, is stable and no sulfide formation takes place. These results can be discussed in terms of the thermodynamic data of these compounds. Sulfidation of saponites with H2S proceeds through: M3Si40io(OH)2 + 3 H2S
•
3MS + 4Si02 + 4H20
(1)
The calculated changes of Gibbs free energy of the reaction are given in Table 2. The tabulated data are those of talc minerals, a structural analogue of saponites with no Si/Al substitution in the tetrahedral layer. As less than 10 % percent of the Si is replaced by Al these data can be considered to be representative for the equivalent saponites. Data for Co talc were not available but can be expected to follow the trend of the metal oxides. Table 2 Gibbs free energy of formation and calculated Gibbs free energy per mol metal sulfide formed for sulfidation of talcs with H2S [21-23] Compound Mg3Si40io(OH)2 Ni3Si40io(OH)2 Zn3Si40io(OH)2
MgO NiO CoO ZnO
Temperature (K) 298 298 298 700 700 700 700
AG" (kJ/mol) -1357 -1071 -1146 -150 -66 -68 -93
AGsuif(kJ/mol) 56 -91 -97 33 -70 -59 -74
From Table 2 it is clear that sulfiding of a Mg saponite is thermodynamically unfavorable, whereas Ni and Zn saponites are liable to react with sulfiir. The same trend is observed for the simple metal oxides including CoO. This low reactivity of Mg towards sulfur is due to the enhanced stability of Mg in an oxidic lattice compared to the other metals as indicated by the values of Gibbs free energy of formation. The enhanced stability of Co in a mixed Co-Mg saponite can thus be understood in terms of thermodynamics. Mixing Co and Mg will result in a AGsuif intermediate between those of the pure Mg and Co saponite. Apparently, according to
145 our XRD observations, AGsuif must be still postive for a Co-Mg 1:1 saponite preventing a complete collapse of the clay structure upon sulfidation at 673 K. However, the TPS pattern shows the presence of a readily sulfidable species, indicated by the low temperature consumption and production of H2S. Most likely this species can be identified as Co atoms exposed at the edges of the clay platelets. These Co atoms are not stabilized by the clay lattice and can be considered to sulfide relatively easy. Only at temperatures above 523 K, when an increasing consumption of H2S is observed, the more hidden Co is sulfided and extracted from the octahedral layer. It is clear, however, that at 673 K a part of Co is not sulfided and remains in the clay's octahedral layer. 4.3 HDS Activity The HDS activity of the saponites has to be discussed both in terms of sulfur resistance and active metal sulfides. As discussed above the bare Zn, Co and Ni saponite detoriate upon sulfidation and P-ZnS, Co9Sg, and Ni3S2 are formed. The latter two are known to be active in HDS, especially when these phases are highly dispersed. This explains the observed activity of the collapsed saponites. ZnS clearly is not active in HDS just as the bare Mg-saponite. The activity of the Mo impregnated Mg-saponite can entirely be ascribed to the presence of molydenum sulfide slabs. For the Zn-saponite the collapse of the clay structure causes the M0S2 to sinter and causes subsequent loss of activity. Probably the sintering of CoQSg is promoted by Mo decreasing the activity relative to the non-impregnated saponite. It is also worthwile to compare the HDS activity of the bare Co-saponite (Fig. 9a) with the bare Co-Mg saponites (Fig 10a). This demonstrates that the specific activity of COQSS is low in comparison with the dispersed cobalt sulfide species formed after sulfiding of the mixed CoMg saponite. Although for the mixed Co-Mg 1:29 saponite the amount of Co is almost a factor 30 lower, the conversion is at about the same level as for the pure Co saponite. This indicates that the turnover rate for the active Co-S species is very high, indicating a high dispersion. The sulfided Co species at the edges of the clay platelets is probably very well dispersed and might even be stabilised by the clay lattice through remaining oxygen bonds. Such a species can account for the HDS activity observed. Additionally dispersed cobalt sulfide will be present supported on the saponite platelets after extraction from the clay lattice. Remarkably, the HDS performance of the mixed Co-Mg saponite and the Co impregnated Mg support are very close. As we can fairly assume that all of the impregnated Co has been sulfided, in contrast to the mixed saponite, one might expect a larger activity of the impregnated sample. However, it seems that the lower amount of cobalt sulfide is compensated by its higher dispersion. Only when Mo is impregnated, which significantly increases the activity, the impregnated sample is slightly more active, probably as a result of a better CoMo dispersion. As noted earlier, the saponites suffer from deactivation during HDS. This deactivation can be attributed to the acidic nature of the supports, as also indicated by the initially formed ibutane and C1-C3 products. The stevensite with no Si/Al substitution, which is therefore far less acidic, indeed does not suffer from deactivation and the performance is as good as that of a conventional CoMo/alumina catalyst. For this reason, the dispersion of the metal-sulfide phase on the stevensite, which is comparable to those on the saponite carriers, can be considered to be quite high. Further research will be directed towards the tuning of the acidic properties of the saponite and the hydrogenation properties of the metal-sulfide. This balance
146 plays a major role in hydrocracking, where both catalyst properties are equally important in obtaining a good activity and selectivity towards the desired products. S.Conclusions This study shows that saponites with only Co, Ni or Zn as octahedral ions are not stable upon exposition to H2S and the clay structure collapses under formation of the bulk metal sulfides. Saponites with only Mg or with combinations of Mg and metals as Co or Ni incorporated in the octahedral layer are stable upon sulfidation. Using a mixed Co-Mg saponite the Co atoms at the edges of the platelets are sulfided as part of the lattice Co. The formed Co sulfide species are highly active in the thiophene HDS and can be compared to a coventional alumina catalyst. Similar results can be obtained by impregnating bare Mg-saponites with active HDS metals such as Co/Mo. Acknowledgements Akzo-Nobel Chemicals B.V. is acknowledged for financial support of this research and Marianne Smolenaars and Tijmen Ros for parts of the experimental work. References 1. J.W. Ward, Fuel Processing Technology, 35 (1993) 55. 2. J.K. Minderhoud, J.A.R. van Veen, Fuel Processing Technology, 35 (1993) 87 3. H.E . Swift in "Advanced Materials in Catalysis", eds. J.T. Burton, R.L. Garten, Academic Press, New York, (1977) 209 4. I.E. Maxwell, W.H.J. Stork ,Stud. Surf Sci. Catal., 58 (1991) 571 5. M.L. Ocelli, R.J. Bernard, Catalysis Today, 2 (1988) 309 6. M.F. Rosa-Brussin, Catal. Rev. -Sci.Eng, 37(1) (1995) 1 7. Y. Sakata, C.H. Hamrin Jr., Ind. Chem.Prod. Res.Dev., 22 (1983) 250 8. J.T. Kloprogge, W.J.J. Welters, E. Booy, V.H.J, de Beer, R.A. van Santen, J.W. Geus, J.B.H. Jansen, App. Cat. A, 97 (1993) 77 9. R.K. Sharma, E.S. Olson, Prepr. Pap.-Am. Chem. Soc. Dive. Fuel Chem., 39 (1994) 702 10. R.J.M.J. Vogels, M.J.H.V. Kerkhoffs, J.W. Geus, Prep.Cat. VI, G. Poncelet (ed), (1995) 1153. 11. H. Topsoe, B.S. Clausen, App. Cat., 25 (1986) 273 12. M. Vaarkamp, B.L. Mojet, M.J. Kappers, J.T. Miller, D.C. Koningsberger, J. Phys Chem., 99 (1995)16067. 13. J.B.A. Van Zon, D.C. Koningsberger, H. Van Blik, D.E. Sayers, J. Chem. Phys., 82 (1985) 5742. 14. F.B.M. Duivenoorden, D.C. Koningsberger, Y.S. Uh, B.C. Gates, J. Am. Chem. Soc, 108 (1986) 6524. 15. J. Mustre de Leon, J.J. Rehr, S.I. Zabinsky, Phys. Rev. B, 44 (1991) 4146 16. H. Suquet, C. de la Calle, H. Pezerat, Clays and Clay Min., 23 (1975) 1 17. A. Manceau, Canadian Mineralogist, 28 (1990) 321 18. R.J.M.J. Vogels, Ph.D. Thesis, Utrecht, 1996 19. V.C. Farmer, Mineral Mag., 31 (1958) 829 20. R.W.T. Wilkins, J. Ito, Am. Mineral., 52 (1967) 1649 21.1.Barin, O. Knacke, " Thermochemical Properties of Inorg. Substances", Springer-Verlag ,1973 22. Y. Tardy, Amer. J. Sci., 279 (1979) 217 23. Y. Tardy, R.M. Garrels, Geochim. Cosmochim. Acta, 38 (1974) 1101
® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
147
Influence of the location of the metal sulfide in NiMo/HY hydrocracking catalysts D. Comet, M. El Qotbi, and J. Leglise Catalyse et Spectrochimie, URA-CNRS-0414, ISMRA, Universite de Caen, 6 Bd du Marechal Juin, 14050 Caen Cedex, France - Fax : (+33) 31 45 28 22
Two sulfided NiMo/HY zeolites were tested as bifiinctional hydroconversion catalysts. The supports had framework Si/Al = 5 or 17, and had thus widely different acidity. Mechanical mixtures of a non-acidic NiMo/Al203 and the NiMo zeolites or the bare HY were similarly tested. The simultaneous hydroconversion of benzene and n-heptane was performed at 250320°C under 8 MPa overall pressure with a constant supply of dimethyldisulfide. The influence of the location of the NiMo has been appraised by comparing the activities in both reactions and the selectivity into isomeric heptanes. All catalysts had comparable hydrogenation capacity for benzene. For the conversion of heptane, higher activity and selectivity were obtained when the NiMo sulfide is inside the zeolite, i.e. close to the protonic sites. Adding external NiMo did not improve the catalytic performance of the NiMoAf zeolites. The importance of the concentration of the heptene intermediates and of their transport in the zeolite is outlined.
1. INTRODUCTION Bifunctional catalysts combining a hydro-dehydrogenation capacity with a surface acidity are widely used in alkane hydroconversion (1). Zeolites doped with Pd or Pt convert low sulfur feeds at near atmospheric pressure (2). Alkanes are selectively isomerized at temperatures below 300°C (3, 4). Naphthenes also react via ring opening while aromatization occurs at higher temperature (5). With such noble metal catalysts, the hydrogenation function is strong enough to balance the acidic function so that cracking is maintained low. Zeolite catalysts where the hydrogenation function is brought about by a CoMo-type sulfide are effective in converting S- and N-containing feeds. However, being less hydrogenating than the noble metals, they operate at higher temperatures and under a high hydrogen pressure (2). They favor hence cracking more than isomerization (6-9). In this study, we tested some catalysts made of HY zeolites and NiMo sulfide. More precisely, we determined how the activity and selectivity for hydrocracking could depend upon the location of the sulfide. This sulfide was either built-up in the zeolite grains, or brought about as a separate NiMo/Al203 component admixed with the HY zeolite. The catalyst performance was evaluated through the simultaneous conversion of n-heptane and benzene.
148 2. EXPERIMENTAL 2.1. Catalysts Support HY-5 was a stabilized HY (Linde, LZY 82) with a lattice Si/Al atom ratio equal to 5; it contained some extra-lattice material. Support HY-17 (Si/Al = 17) was obtained by steaming a NH4Y at 740°C followed by washing with HCl. Both supports were treated with solutions containing Ni then Mo ions (10). The final Ni and Mo loadings were respectively 2.8 and 9.1 wt.% for NiMoA^-S; 2.2 and 9.8% for NiMoA^-17. The commercial NiMo/Al203 catalyst (Procatalyse, HR 346) contained 2.8% Ni and 9.3% Mo. Mechanical mixtures of the HY zeolites with the NiMo/Al203 were realized by grinding the two powders in a mortar, pelletizing and grinding again to 0.16-0.2 mm size. Mixtures of NiMo/HY and NiMo/Al203 were obtained similarly. 2.2. Catalytic measurements Benzene (Be) and n-heptane (Hp) were simultaneously converted at 250-320°C in a flow reactor operated at 8 MPa hydrogen pressure. Prior to the reaction, the catalysts were sulfided at 320°C. The same Hp - Be - DMDS (dimethyldisulfide) mbrture with mass composition 7820-2 was used as the sulfiding agent then as the reagent. DMDS was completely converted into H2S and methane. For the reaction, the feed rate of hydrocarbons was 0.16 g/h, so that partial pressure were 37 kPa for Hp and 12 kPa for Be. n-Heptane was converted into isomeric heptanes and cracked light alkanes. Benzene was hydrogenated into cyclohexane and methylcyclopentane. When no further products appeared, the fractional conversion of heptane, XJ^Q and x^^^, and that of benzene xj^y^j were deduced from GC analysis. With catalyst NiMoA^-5, hexanes and lighter alkanes appeared as secondary products in the transformation of benzene (10). They had virtually no influence on x^^ but had to be included in deriving xj^y^j. 3. RESULTS AND INTERPRETATION 3.1. Hydrogenation of benzene In the range 260-320°C, the conversion of benzene was rather low. The extent of benzene hydrogenation recorded at constant mass velocity for the two sulfided NiMoAf and the NiMo/Al203 catalyst is plotted on figure 1 as a function of temperature. The rates of hydrogenation measured at 280°C are shown in table 1; they are very close for NiMoA^-5 and NiMo/Y-17 amounting to about 2/3 of the rate measured for NiMo/Al203 (0.075 mol/hkg). Hence, the dispersion of the NiMo sulfide appears nearly as good on zeolites as on alumina. Yet, zeolite HY-5 has a much higher ion-exchange capacity than HY-17, while the mesoporous volume is more developed in HY-17 as a result of acid leaching (11). Accordingly, the location of the NiMo phases in the zeolites was ascertained by electron microscopy coupled with local analysis by EDX (11). Some results are listed in table 1. In the calcined state, the major part of Ni was accommodated inside the grains for both HY supports. Similarly, Mo was found essentially at the interior of HY-17, but only half of the Mo could penetrate the grains of HY-5, the remaining part being detected as external M0O3 microcrystals alone or associated with amorphous alumina.
149 14 12 10 8 6 o/Y-17
4 2 -J
240
260
NiMo/Y-5
L-
280 300 T e m p e r a t u r e (°C)
320
340
Figure 1. Benzene hydrogenation over NiMoA^ and NiMo/Al203 catalysts (benzene flow rate 0.37 10-3 mol/h; catalyst mass 0.17 g). Table 1 NiMoAf catalysts: proportion of internal Ni and Mo deduced from EDX; rates of reaction measured at 280°C over the sulfided catalysts Benzene
Internal metal (%) Sulfided
Oxidic
Heptane
Ni
Mo
Ni
Mo
Rate mol/hkg
NiMoA^-5
>95
53
43
46
0.048
2.4
1.24
NiMoA^-17
>95
>95
50
83
0.056
0.55
1.75
NiMo/Al203
-
-
-
-
0.075
<0.01
Catalyst
Rate mol/hkg
TOF*
h-1
-
* TOF referred to framework aluminum After sulfidation and catalytic reaction, about half of the Ni had moved outward the zeolites. Moreover, the solid NiMoAf-17 had lost a large part of its internal Mo. In the TEM images, many M0S2 microparticles with a slab structure were found inside the grains. These sulfide particles are preferably found in the mesopores created by the dealumination treatment; they are slightly larger on HY-17 than on HY-5. Welters et al. (12) reported lesser proportions of internal Ni or Mo (15% each) in a CaY, but such a support is probably devoid of mesopores. Furthermore EDX analysis of the sulfided NiMoAf proves that Ni and Mo remained associated everywhere in the grains, so that internal mixed NiMo sulfide is likely to occur. By contrast, the large M0O3 particles outside the grains were not fiilly transformed; only their outer parts were sulfided.
150 The hydrogenation capacity measured over the NiMoAT zeolites may be attributed largely to the M0S2 microparticles promoted by Ni. These are at the interior of the zeolite grains, and display a larger specific area than the bulky particles. From the reactivity data, the dispersion of the active sulfide appears equivalent in both zeolites and slightly lower than on AI2O3. 3.2. Hydrocracking of n-heptane The conversion of n-heptane measured at a constant flow rate is plotted on figure 2a as a fimction of temperature. Both NiMoAf catalysts appeared active in converting heptane. However, no appreciable conversion was observed with the NiMo/Al203 catalyst even at 320°C. The rates of hydroconversion measured at 280°C are gathered in table 1. Catalyst NiMo/Y-5 is clearly more active than NiMo/Y-17. This is in line with a higher density of acid sites on the Al-rich support. A relationship between the rate of hydrocracking and acidity was then sought. The density of acid sites was taken as that of.fi'amework aluminum N^i- IR measurements had shown that the actual N ^ were lower on the sulfided NiMoAf than on the corresponding HY supports (11). With these figures, the turnover fi-equencies do not appear very different on both NiMoA^ catalysts (Table 1); the higher turnover obtained for NiMo/Y-17 may reveal stronger Bronsted sites (13). b - Selectivity for heptane isomers
a - Conversion of n-heptane 100 I
20
xjz
NiMoA'-S
3^80^
NiMoA'-S
15 h
r (°)
22 60
2 > 10 L
NiMoA^-l7
> c o
0
/%
NiMo/Y-17 (3xW) (»a)
^
1
89
u
40
o
*
fiMolY-nia)
ex.
20 f
240
260
280 300 320 Temperature (°C)
340
1
20
1
40 60 Cracking Yield
1 ^*'^3o
80
100
Figure 2. Hydrocracking of n-heptane over NiMoA" and NiMo/Al203 (heptane flow rate 1.12 10-3 mol/h; catalyst mass: NiMoA^-5, 0.17 g (O); NiMo/Y-17, 0.17 g (D), 0.51 g (H)). Figure 2b displays the conversion into isomers versus that into cracked products. Measurements performed with equal masses of each catalyst (0.17 g) gave a maximum of 13% isomers over NiMoAf-5, but only 5% over NiMoAf-17. Such amounts are considerably under those reported over Pt zeolites (3, 4; 14-17). Therefore, the hydrogenation fimction in the NiMoAf catalyst is much below their acidic fimction, and these solids are far fi-om featuring
151 ideal hydrocracking (3). The selectivity for isomerization could be somewhat improved upon addition of nitrogen compounds to the feed; but at the expense of activity (6, 7). Although the above results suggest that NiMo/Y-17 is more cracking-oriented than NiMoAf-5, their intrinsic selectivities are not necessarily different. In fact, the activation energy is higher for cracking than for isomerization. Thus, catalyst NiMo/Y-17, which was the least active, required a higher temperature to achieve the same conversion as NiMo/Y-5. Consequently, cracking was enhanced with respect to isomerization. However, measurements carried out with a triple mass of NiMoAf-n afforded similar selectivities for both catalysts (Figure 2b). The product distributions listed in tables 2 and 3 confirm that similar reaction mechanisms operate. The conversion of heptane increased with temperature. However, very little variations were observed in the distribution of heptane isomers (8). Monobranched (MB) as well as dibranched isomers (DB) appeared as initial products; the ratio MB/DB remained close to 3. Higher values, up to 10-20, have been noticed at low conversion over Pt/Y catalyst (15). In addition, the gem-dimethylpentanes amounted to less than 10% of the DB isomers, and only traces of 2,2,3-trimethylpentane could be detected. Table 2 Hydroconversion of heptane over NiMoAf catalysts at 280^C. Distribution of heptane isomers Catalyst
Conversion (%) Overall
Iso.
Distribution (mol%)
3MH
2MH
3EP 24DMP 23DMP 22DMP + 33DMP
NiMoA^-5
34.2
10.4
40.5
30.5
2.9
9.3
14.4
2.4
NiMo/Y-17
35.4
10.2
44.5
29.2
2.5
8.9
12.9
2.0
NiMoA^-5 + NiMo/Al203
41.2
10.7
40.8
31.9
2.9
8.7
13.3
2.4
NiMoA^-17 4NiMo/Al203
9.3
5.7
46.9
30.4
1.7
7.8
11.8
1.4
Such distributions of isomers are readily explained by the bifunctional mechanism. Dehydrogenation of heptane over the NiMo yields n-heptene which migrates toward a protonic site where it is adsorbed as a Cj"*" carbenium ion (1, 14). n-CyHu + H"^ <^ n-CiH^s
(stepp)
The linear carbenium ion is then isomerized into a monobranched species. n-C7Hj5 <=> MBCTH^S
(stepi)
This MB carbenium ion is either desorbed as isomeric heptene MBrCiUts + C7H14 <=> MB-C7H14 + C7HJ5
(step d)
or further isomerized into a dibranched species. MB-C7Hjt5
o
DB-C7Ht5
(step ii)
152 Step d which yields isoheptanes is faster than steps i and ii whenever the concentration of alkenes is maintained high enough, Le, when dehydrogenation and alkene migration are faster than further isomerization and cracking. This is obviously not the case with the NiMoA^ zeolites, where DB isomers appeared initially. The acidic function is then too high, but the two NiMoA^ catalysts cannot be differentiated in this respect. Furthermore, cracked products appeared at very low conversion levels. Therefore on the NiMo/Y, P-splitting of the carbenium ions as for instance DB-C7Ht5
<=> C3H6 + i-C4H^
(stepc)
occurred at a rate comparable with steps i and ii. Cracking patterns are reported in table 3 at similar conversion levels. Once more, no difference arose between the two NiMo/Y. The most salient feature resides in the values of Ncra» '•^- the number of light alkanes molecules resulting from the cracking of 1.00 heptane molecules. According to classical monomolecular hydrocracking (MC), the number N^a shouW be 200. The lower values reported here indicate that in addition to the MC mechanism, a dimerization cracking process (DC) also occurred (8, 18). This DC mechanism accounts for the excess of C4 with respect to C3, and also to the formation of C5 and C5. From these data, the participation of the DC path to the overall cracking process was evaluated at 14% for both NiMo/Y-5 and NiMo/Y-17 (8). Table 3 Hydroconversion of heptane over NiMoA^ catalysts at 280°C. Distribution of cracked products Catalyst
Distribution (mol/100 mol nCy cracked)
Conversion (%) Overall
Cra.
Nc3
NiC4
NnC4
NiMoA^-5
34.2
23.8
89.4
88.2
7.0
5.0
NiMoA^-17
35.4
25.2
92.3
86.8
8.0
NiMoA^-S + NiMo/AljOs
41.2
30.5
91.7
88.0
NiMoA^-17 + NiMo/Al203
9.3
3.6
92.0
86.8
NiC5 N^cs
Nc6
^cra
0.3
1.9
191.8
3.5
0.3
2.0
192.9
6.6
3.9
0.4
1.7
192.3
8.7
3.2
0.3
1.2
192.2
Another feature in the cracking pattern of heptane (Table 3) is the high amount of isobutane among the C4. With heptane, ideal hydrocracking would predict 50% isobutane at low conversion (3). Such a value was actually reported over a near-ideal catalyst at low temperature (3, 15), but the degree of branching increased with conversion, up to 88% at high temperature (3). Higher values above 95% were obtained with non-ideal Pt/HY catalysts at 200-250°C (15, 16). For both of our sulfided NiMoAT catalysts, the degree of C4 branching reached 94% at 260°C, and decreased to 86% when increasing the temperature at 320°C. Again, this is due to a low hydrogenation function of the NiMo catalysts.
153 3.3. Mixtures of zeolites and NiMo/Al203 The NiMo zeolites were prepared with the purpose of building the hydrogenating sites close to the acidic sites, and thus minimizing the distance between them. Conversely, the distance between the hydrogenating and the acidic sites may be increased by putting them on two different solids: NiMo/Al203 and the bare HY zeolite. A first mixed catalyst (A') containing 0.17 g of HY-5 and 0.34 g of NiMo/Al203 was compared with NiMo/Y-5 (sample A: 0.17 g) so that both catalysts had similar amounts of acidic sites. The extent of benzene hydrogenation over mixture A was found twice as high as over catalyst A in agreement with the NiMo content. However, the conversion of heptane over A was much lower than over A (Figure 3a). Thus, over catalyst A having the NiMo outside the zeolite, the rate of hydroconversion wasfifteentimes lower than for the catalyst with NiMo inside. Furthermore, the proportion of heptane isomers observed with the mixture A was much lower than with the NiMo zeolite (Figure 3b), although the distributions of isomers and cracked products were not appreciably changed (Tables 2 and 3). b - Selectivity for heptane isomers
a - Conversion of n-heptane 20
15 h-
,NiMoA^-5 pi/
s 8
/ hj
HY-5 + NiMo/Al203 240
260 280 300 320 Temperature (°C)
340
0
NiMoA^-5 + /^NiMo/Al203 >sp
HY-5 + NiMo/Al203
r
1
0
20
. L —
1
1 ^^Sfa
40 60 80 Cracking Yield
100
Figure 3. Hydrocracking of n-heptane over NiMo catalysts containing zeolite Y-5. A second mixed catalyst (A") containing equal masses (0.17 g) of NiMo/Y-5 and NiMo/Al203 was elaborated with the hope of improving the hydrogenating capacity of catalyst NiMoA^-5 while keeping the same amount of acid sites. Accordingly, benzene conversion was higher with catalyst A" than with A. However the conversion of heptane was nearly the same on both catalysts (Figure 3a). Furthermore, they exhibited comparable yields of heptane isomers (Figure 3b). Similarly, two mechanical mixtures were realized with catalysts containing zeolite Y-17. The mixture (B') containing 0.17 g of the neat HY-17 and twice as much NiMo/Al203 converted heptane at a lower rate than 0.17 g of NiMoAf-17 (Figure 4a). Internal NiMo
154 appeared once more preferable, the rate over NiMoAf-17 being 5 times that of the mixture B'. The product distributions in tables 2 and 3 show that the reaction proceeds through similar reaction mechanisms. The selectivity for isomers observed with B' was consistently low, but not significantly under that of NiMoA^-17 (Figure 4b). Finally, mixture B" was made by admixing 0.17 g of NiMoA^-17 with a two-fold excess of NiMo/Al203. As for NiMoA^-5, the conversion of heptane over NiMoA^-17 was not improved by adding external NiMo (Figure 4a). However, a slight improvement of the selectivity was observed (Figure 4b). b - Selectivity for heptane isomers
a - Conversion of n-heptane 20
100 80 o
60 h
15 NiMoA^-17 + NiMo/Al203
NiMoA^-17 + NiMo/AhCb NiMoA'-n
NiMoA'-n
c o U
o 40 hHY-17 + NiMo/Al203
HY-17 + NiMo/Al203|
ao 20 0 240
260
280
300
Temperature (°C)
320
340
20
40
60
80
100
Cracking Yield
Figure 4. Hydrocracking of n-heptane over NiMo catalysts containing zeolite Y-17. 3.4. Distance between hydrogenating and acidic sites The conversion of heptane over mechanical mixtures A' and B' of the HY zeolites and the NiMo/Al203 proves some cooperation between the acidic and hydrogenation centers, since the level of activity was much higher than on either component. The NiMo zeolites are therefore bifunctional catalysts, and the hydroconversion proceeds through heptene intermediates. Good hydrocracking catalysts should possess a high hydro-dehydrogenation activity, together with a short distance between the metallic and the acidic sites. The first condition is clearly not fiilfiUed with the NiMoY catalysts where cracking prevails over isomerization. Thus, the steady concentration CQ of linear heptenes at the acid sites is likely well under the limit C^q at equilibrium with n-heptane. Therefore, the desorption of isoheptenes (step d) is rather slow. As mixtures A' and B' exhibited lower activities and selectivities than the NiMoY, it may be inferred that the migration of heptene toward the inside of the zeolite is rate limiting. In this case, the longer intersite distance results in CQ values lower than in the corresponding NiMoY. Some quantitative confirmation of the transport limitation in those mixtures may be sought fi"om the Weisz's criterion for polystep reaction systems (19), which is
155
dt Ceq D where dN/dt is the rate per unit volume, R the radius of the zeolite particles, and D the effective difRisivity of linear heptenes. If diffusion is slow, O will be much greater than 1. O was evaluated for the conversion of heptane at 320 °C over the mixed catalyst A' containing support HY-5. Measurements by light scattering afforded a surface-averaged radius R = 1.15 ^m. Diffusivity D was estimated at 10"^ m^ s'^ from NMR measurements on NaX (20). An equilibrium concentration Cgq = 1.7 10"^ mol m"^ was assessed from thermodynamic data (21). With the observed value dN/dt = 0.65 mol s"^ m"^, we deduced O = 0.5. This not clearly in favor of diffusional limitation. Ruckenstein for instance reported a value O = 0.23 for the conversion of methylcyclopentane at 470'^C and 1 atm over a p zeolite (22) and concluded to the absence of limitation. The same synergism between Pt/Al203 and various zeolites was reported for light alkane isomerization (23). However, with our A' and B' mixtures, some limitation is obvious from the rate and selectivity data. The low O value may be due to a poor estimate of D. Moreover C^q in the above formula really stands for the actual concentration CQ outside the zeolite grains; this may be well under Cgq since dehydrogenation over NiMo/Al203 is not fast enough. Therefore, in the NiMo-HY system, the intersite distance appears as the main factor governing the activity. In both NiMo zeolites, the NiMo sulfide seems rather well dispersed and the two kind of sites cooperate efficiently. There, the alkene concentration CQ is higher than with mixtures A' and B', although substantially below the equilibrium value. The internal Co is hardly enhanced when NiMo/Ai203 is admixed with the NiMoY zeolites because then the distance between the acidic and external NiMo is too high and transport limitation occurs. Such transport limitation however is less severe with catalysts containing HY-17 simply because they are less active. Accordingly, hydroisomerization on mixture B" is slightly improved over that of NiMoA^-17. The above conclusions were deduced from the heptane test performed at low temperature, and will not necessarily hold under more severe conditions. For instance, the homogeneous distribution of the NiMo inside zeolite Y appeared favorable for the conversion of heptane. A different behavior was noted when methylcyclopentane reacted over a Pt mordenite (24). There, because of coking, a distribution of Pt at the periphery of the grains was preferred. Besides, external NiMo proved useful in industrial hydrocracking catalyst (6, 25) probably because it improves the diffusivity of large molecules and the resistance to coking. REFERENCES 1. H.L. Coonradt and W.E. Garwood, Ind. Eng. Chem. Prod. Res. Dev., 3 (1964) 38. 2. J.F. Le Page, J. Cosyns, P. Courty, E. Freund, J.P. Franck, Y. Jacquin, B. Juguin, C. Marcilly, G. Martino, J. Miquel, R. Montamal, A. Sugier and H. van Landeghem, Applied Heterogeneous Catalysis, Technip, Paris ,1987. 3. J. Weitkamp, Hydrocracking and Hydrotreating, J.W. Ward and G.A. Quader (eds.), ACS Symp. Ser., Vol. 20, p. 1, Amer. Chem. Soc, Washington DC, 1975. 4. J.P. Franck, M. El Malki and R. Montamal, Rev. Inst. Fr. Petr., 36 (1981) 211.
156 5. P. G. Smimiotis and E. Ruckenstein, Catal. Lett., 17 (1993) 341. 6. P. Dufresne, A. Quesada and S. Mignard, Catalysis in petroleum refining 1989, Kuwait, 1989, D.L. Trimm, S. Akashah, M. Absi-Halabi and A. Bishara (eds.). Stud. Surf. Sci. Catal, Vol. 53, p. 301, Elsevier, Amsterdam, 1990. 7. M. Guisnet, J.L. Lemberton, C. Thomazeau and S. Mignard, Proc. 9^ Int. Zeol. Conf, Montreal, 1992, R. von Ballmoos, J.B. Higgins and M.M.J. Treacy (eds.). Vol. 2, p. 413, Butterworth/Heinemann, Boston/London, 1993. 8. J. Leglise, M. El Qotbi and D. Comet, Coll. Czech. Chem. Commun., 57 (1992) 882. 9. W.J.J. Welters, O.H. van der Waerden, V.H.J. de Beer and R.A. van Santen, Ind. Eng. Chem. Res., 34 (1995) 1166. 10. J. Leglise, M. El Qotbi, J.M. Goupil and D. Comet, Catal. Lett., 10 (1991) 103. 11. J. Leglise, J.M. Manoli, C. Potvin, G. Djega-Mariadassou and D. Comet, J. Catal., 152 (1995) 275. 12. W.J.J. Welters, O.H. van der Waerden, W. Zandgergen, V.H.J, de Beer and R.A. van Santen, Ind. Eng. Chem. Res., 34 (1995) 1156. 13. R. Beaumont and D. Barthomeuf, J. Catal., 26 (1972) 218. 14. J. Weitkamp, Ind. Eng. Chem. Prod. Res. Dev., 21 (1982) 550. 15. G.E. Giannetto, G.R. Perot and M. R Guisnet, Ind. Eng. Chem. Prod. Res. Dev., 25 (1986)481. 16. J.A. Martens, P.A. Jacobs and J. Weitkamp, Appl. Catal., 20 (1986) 239. 17. G.F. Froment, Catal. today, 1 (1987) 455. 18. A Lopez Agudo, A. Asensio and A. Corma, J. Catal., 69 (1981) 274. 19. P.B. Weisz, Adv. Catal., 13 (1962) 137. 20. M.F.M. Fost, Introduction to zeolite science and practice, H. van Bekkum, EM. Flanigen and J.C. Jansen (eds.). Stud. Surf. Sci. Catal., Vol. 58, Ch. 11, p. 391, Elsevier, Amsterdam, 1991. 21. RA. Alberty and C. A. Gehrig, J. Phys. Chem. Ref Data, 14 (1985) 803. 22. P. G. Smimiotis and E. Ruckenstein, Ind. Eng. Chem. Res., 33 (1994) 493. 23. G. Perot, P. Hilaireau and M. Guisnet, Proc. 6^ Int. Zeol. Conf, Reno, 1983, D. Olson and A. Bisio (eds.), p. 427, Butterworth, Guildford, 1984. 24. B.A Lemer, B.T. CaviU and W. M. H. Sachtler, Appl. Catal., 21 (1994) 23. 25. A. Nishijima, S. Yoshitomi, H. Shimada, Y. Yoshimura, T. Sato and N. Matsubayashi, Proc. 9* Int. Cong. Catal., Calgary, 1988, MJ. Phillips and M. Teman (eds.). Vol. 1, p. 174, Chemical Institute Canada, Ottawa, 1988.
® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
157
ACIDITY INDUCED BY H2S ADSORPTION ON UNPROMOTED AND PROMOTED SULHDED CATALYSTS Cyrille Petit, Frangoise Mauge* and Jean - Claude Lavalley Laboratoire SPECTROCAT - URA CNRS 414 - ISMRA - Universite de Caen 6, Bd du Marshal Juin -14050 CAEN cedex (France)
The effect of H2S adsorption on the nature and the number of acid sites of alumina supported sulfided catalysts was characterized by IR spectroscopy. For monitoring the variation of the surface properties, dimethyl-2,6-pyridine, a probe molecule more convenient to detect Bronsted acidity than pyridine, was adsorbed at room temperature (r.t). Before H2S introduction, a few number of Bronsted acid sites was detected, particularly on the promoted catalysts. Whatever the catalysts, H2S adsorption at r.t led to the creation of Br0nsted acid sites at the expense of Lewis acid sites. The concentration of the created Bronsted acid sites depended on the presence and on the nature of the promoter. The highest number of Bronsted acid sites was detected on CoMo/AI2O3. 1 - Introduction The activity of hydrotreating catalysts was generally related to the number of coordinatively unsaturated sites (1). In particular, IR spectroscopy of probe molecules like CO allowed us to differentiate unpromoted and promoted coordinatively unsaturated sites and good correlations between absorption bands characteristic of promoted sites and HDS activity were previously obtained (2-4). Nevertheless, studies of the catalytic properties of sulfided catalysts indicate the importance of Bronsted acidity (5). Moreover, the concentration of sulfur vacancies can be modified under H2S which is present during the reaction since the formation of Bronsted acid SH groups is then expected. The direct detection of SH groups is difficult by IR spectroscopy and the adsorption of a probe molecule able to be protonated by such groups is a way to assess their presence. Pyridine is generally used and it was shown that at temperatures typical of the reaction, pyridinium species were formed on M0/AI2O3 and C0M0/AI2O3 (6,7). Previous studies pointed out that dimethyl-2,6-pyridine (DMP) was more sensitive to the presence of Bronsted acidity than pyridine (8,9). For example, we showed that no Bronsted acid sites were detected by pyridine adsorption on Ti02-Zr02 mixed oxides whereas, in the same conditions, the
158 formation of protonated species was observed by adsorption of dimethyl-2,6pyridine (10). This can be related to the higher basicity of this probe molecule (pyridine, pKa = 5.2 - DMP, pKa = 6.7). Moreover, it has been proposed that the presence of methyl groups adjacent to the nitrogen atom is unfavorable for coordination of DMP. The aim of the present work is to determine if dimethyl-2,6-pyridine is a suitable probe molecule to characterize the weak Bronsted acidity of sulfided catalysts and to establish if the presence of H2S increases the Bronsted acidity of these catalysts. 2. Experimental The catalysts were either prepared in our laboratory, by pore filling impregnation of a y-alumina : M0/AI2O3 (9.3% Mo), C0/AI2O3 (3.1% Co), NiW/Al203 (2.5% Ni - 16.6% W) or provided by Procatalyse CoMo/AI2O3 (HR 306) and NiMo/Al203 (HR 346). The samples were pressed into self-supporting wafers (disc = 2 cm^). All the studied samples were sulfided in situ in the IR cell. Firstly, the catalysts were evacuated at increasing temperature from 298 to 473 K (ramp : 0.15 K.s-^). Then, 80 torrs of a mixture of H2 + H2S (15% vol.) was introduced. The temperature was raised from 473 to 673 K (ramp : 0.15 K.s"l) with a plateau of one hour at 673 K. Subsequently, the catalysts were evacuated at 673 K for 10 min. Two other sulfidations for 2 hours and 15 hours were performed at 673 K. The ultimate sulfidation was followed by an evacuation at 673 K for one hour to 3 10"^ torr before cooling at room temperature under evacuation. Dimethyl-2,6-pyridine (3 torrs) was contacted with the catalyst and then evacuated at room temperature (r.t.) for 15 minutes (P = 210"^ torr). Calibrated small doses of H2S were successively introduced and finally a pressure of 15 torrs of H2S was contacted with the disc. This was followed by an evacuation at r.t. for 15 minutes (P= 8.10-^ torr). Spectra were recorded on a Fourier transform infrared spectrometer Nicolet 60 SX. For sulfided catalysts, the spectra reported below correspond to a mass of 10 mg whereas for alumina, they correspond to a mass of 8 mg which is almost the mass of the support present in a disc of 10 mg of a sulfided catalyst. 3. Results 3.1. Effect of H2S adsorption on various metal oxides Previous infrared studies on the adsorption of dimethyl-2,6-pyridine (DMP) reported that the wavenumbers of the bands of this probe molecule depended on the adsorption site, in particular, those corresponding to the vga and V8b vibrations occurring in the 1660-1580 cmr^ range (8-12). On pure alumina or
159 ^ 1580 A
b s o r b a n c e
1650
1600
1550
Wavenumbers (cm"^)
B
1602 ; 1613, /\ !
0.04
1
1654
A
J
•
1
1
'
t
<
>
"
^ 'A
/^'
\ /
1
by /aT.'<
'"">
1650
/
\» xL***
1600
1550
Wavenumbers (cm"^)
Figure 1 : Effect of the introduction of H2S on the spectra of DMP adsorbed on various metal oxides : A - silica; B - alumina. a - Before H2S adsorption ; b - After introduction of 15 torrs of H2S followed by an evacuation at r.t. fluorinated alumina and on zeolites, previous works showed that bands between 1655 and 1625 cm"^ characterized the formation of protonated species whereas bands between 1618 and 1580 cm"^ were due to DMP interaction with Lewis acid sites or to weakly adsorbed DMP. Spectra of the DMP species formed have been compared before and after H2S adsorption. We characterized first the adsorption of this molecule on silica, a quite inert metal oxide (figure lA). In the 1700-1500 cm"^ zone, two bands are detected at 1602 and 1580 cm'^ which can be attributed to H-bonded species, in agreement with the concomitant perturbation of the Si-OH band, and to physisbrbed species. Introduction of H2S followed by an evacuation at r.t. induces a decrease of the intensity of the two bands but no new^ species appears. On alumina, DMP adsorption followed by an evacuation at r.t. mainly leads to the appearance of bands at 1613, 1602 and 1580 cm'^ (figure IB). As on silica, the band at 1602 cm"^ indicates the presence of DMP in weak interaction with the surface (hydrogen bonded and/or physisorbed species) whereas the band at 1613 cm"^ characterized DMP adsorption on strong Lewis acid sites. The band at 1580 cm'^ is common to these two types of interaction. The weak band at 1654 cm'^ indicates the presence of a small number of Bronsted acid sites, hitroduction of small doses of H2S on alumina followed by an evacuation at r.t leads to a strong decrease of the number of DMP species coordinated and/or in weak interaction with the surface (figure IB). In parallel, we note an enhancement of the
160 1625
0.04
B
164( 11613 ^^^
JL 1602 A
"-i^"^^
1650 1600 1550 Waveninrbers (cnr')
I
1
•
" 1
•
1
'i^' 1
j
1650 1600 1550 Wavenurribers (crtr^)
Figure 2 : Effect of the introduction of H2S on the spectra of DMP adsorbed on the various components of a Co-Mo catalyst : A - Before H2S adsorption ; B - After introduction of 15 torrs of H2S followed by an evacuation at r.t (*) AI2O3, ( ) M0/AI2O3, {- -* ) C0/AI2O3, ( ) C0M0/AI2O3
C0M0/AI2O3 M0/AI2O3 C0/AI2O3
1650
1600 Wavenumbers (cm-^)
1550
Figure 3 : Effect of H2S on DMP adsorption for the various components of a sulfided C0M0/AI2O3 catalyst. Subtracted spectra after H2S adsorption (15 torrs of H2S followed by an evacuation at r.t.) minus before H2S introduction.
161 intensity of the 1654 cm-^ band and the appearance of a shoulder at 1629 cm-^ indicating an increase of the number of the protonated species. 3.2 Effect of H2S adsorption on sulfided Mo, Co and C0M0/AI2O3 We compared the effect of the H2S adsorption on the various components of a sulfided CoMo catalyst i.e. M0/AI2O3, C0/AI2O3 and Co-Mo/AI2O3. In a first step, we studied the consequence of the H2S adsorption followed by an evacuation at r.t. (figures 2 and 3). In a second step, we analyzed more quantitatively the effect of the introduction of small doses of H2S(figure 4). On sulfided Mo/AI2O3, the adsorption of DMP leads to the appearance of the same species as on pure alumina, i.e. DMP in weak interaction with the surface, coordinated species and a few number of protonated species. No band characteristic of DMP adsorption on molybdenum sites is detected. H2S introduction leads to the same features as on alumina, i.e. the decrease of the number of weakly adsorbed species and coordinated species and the creation of Bronsted acid sites. Nevertheless, the number of Bronsted sites created is higher than on alumina. On sulfided C0/AI2O3, the bands characteristic of DMP adsorption are close to those observed on alumina except the band at 1602 cm"^ which is less intense. After H2S adsorption, the number of Bronsted sites detected is close to that observed on alumina in the same conditions (Fig. 2). Before H2S adsorption, the sulfided C0M0/AI2O3 catalyst presents the highest amount of Bronsted acid sites and the smallest number of Lewis acid sites and DMP species in weak interaction with the surface (Fig. 2A). The H2S adsorption leads to a marked increase of Bronsted acid sites whereas the other species are almost no more detected (Fig. 2B). In order to clearly characterize the creation and the poisoning of sites due to H2S adsorption on all these catalysts, we present in figure 3 the difference between spectra scanned after and before H2S introduction. Positive and negative bands correspond to an increase or a decrease of the number of sites, respectively, due to H2S adsorption. Figure 3 provides evidence for the creation of Bronsted acid sites on all the catalysts (intensity increase of the vga band at 1645-1655 cm"^ and the V8b band at 1625 cm"^). It appears that the number of sites so created follows the order: C0M0/AI2O3 > M0/AI2O3 > C0/AI2O3 > AI2O3 Negative bands observed at 1613,1602 and 1580 cm'^ correspond to a decrease of the number of Lewis acid sites and DMP weakly adsorbed. This decrease is smaller on Mo/AI2O3 and C0/AI2O3 than on alumina and maximum for CoMo/AI2O3. The effect of introduction of small doses of H2S on the C0M0/AI2O3 catalyst is presented in figure 4A. It shows a continuous creation of Bronsted acid sites at the expense of the bands characterizing Lewis acid sites and DMP species in weak interaction with the surface. In order to follow more quantitatively this effect, we studied the intensity variation, for each dose introduced, from spectra resulting from the subtraction of those obtained after and before H2S introduction, as previously described. The area of the bands corresponding to the creation of
162
1650 1600 1550 Wavenumbers (cm"^)
A1203
•
C0/AI2O3
M0/AI2O3
•
C0M0/AI2O3
H2S introduced (nmoles)
Figure 4 : A - Effect of the introduction of small doses of H2S on the spectra of DMP adsorbed on a C0M0/AI2O3 catalyst. B - Variation of the area of the bands characterizing the creation of Brjlnsted acid sites versus the amount of H2S introduced for the various components of a CoMo/ AI2O3 catalyst. Bronsted acid sites is measured between 1680 to 1615 cm'^ and its variation with the amount of H2S introduced is reported in figure 4B. For all the catalysts, the variation of the number of Bronsted acid sites versus the number of micromoles of H2S introduced is close; it quickly reaches saturation. Nevertheless, it reaches saturation for higher H2S doses on Mo and C0M0/AI2O3 (~10 jimoles) than on alumina and Co/AI2O3 (2 jimoles). This is in agreement with the higher amount of Bronsted acid sites created on M0/AI2O3 and C0M0/AI2O3. We remark, for these two catalysts, that the number of Bronsted sites created is similar. Nevertheless, figure 4B shows that, after evacuation of H2S at r.t, the number of protonated species decreases on Mo/AI2O3 whereas it does not change after evacuation on C0M0/AI2O3. This could indicate that the created Bronsted acid sites are stronger on the promoted catalyst than on Mo/AI2O3. 3.3 Effect of H2S adsorption on various sulfided catalysts This study was extended to the sulfided NiMo/Al203 and NiW/Al203 catalysts (figure 5). DMP adsorption leads to the same bands as on sulfided C0M0/AI2O3. The intensity of the massif of bands at 1613, 1602 and 1580 cm-^ is higher on NiMo/Al203 than on CoMo/AI2O3 whereas on NiW/AI2O3, it is weaker. By contrast, the band near 1650 cm"^ which characterizes Bronsted acidity, presents the same intensity on the three catalysts before H2S introduction at r.t.
163
Ai
1625 1646/\1613 1580 1602 r
1650
1600
1550
1650
1600
1550
Wavenumbers (cmO
Wavenumbers (cm-^
Figure 5 : Effect of the introduction of H2S on the spectra of DMP adsorbed on various sulfided catalysts: ( ) Before H2S adsorption; ( ) After introduction of 15 torrs of H2S followed by an evacuation at r.t.
Figure 6 : Effect of H2S or CH3SH on the spectra of DMP adsorbed on M0/AI2O3 : a - DMP species present after evacuation at r.t. ; b - after H2S adsorption followed by an evacuation at r.t ; c : after CH3SH adsorption followed by an evacuation at r.t.
The effect of H2S adsorption is qualitatively the same. Nevertheless, figure 5 indicates that the highest amount of Bronsted acid sites is created on CoMo/AI2O3 whereas the number of sites present on NiMo/Al203 and NiW/AI2O3 is similar. 3.4 Effect of CH3SH adsorption on sulfided M0/AI2O3 The previous experiments clearly show that H2S addition creates Bronsted acid sites on sulfided catalysts. However, these experiments do not allow us to specify if H2S adsorption occurs dissociatively or not. Indeed, the intensity of the v(SH) band due to adsorbed H2S is so weak that no information can be drawn concerning the H2S adsorption mode. The v(SH) band of CH3SH seems more intense, in particular when it is coordinatively chemisorbed. On pure alumina, O. Saur et al. (13) showed that the first doses of CH3SH adsorbed dissociatively since the v(SH) band was not detected, while for higher amounts of CH3SH introduced, its adsorption became undissociative since the v(SH) band was then detected at 2560 cm'^. Therefore, instead of adsorbing H2S, we introduced CH3SH in order to study the v(SH) band. Adsorption of CH3SH at r.t. leads to the same results as for H2S concerning the DMP adsorption (figure 6) : the intensity of bands at 1613, 1602 and 1580 cm-^ decreases while that of the band near 1650 cm-^ increases. The number of Br0nsted acid sites created by H2S and CH3SH introduction is close. We note, from
164 the intensity of the v(CH3) bands, that CH3SH adsorption resists to evacuation at r.t. whereas no band around 2560 cm"^ can be detected. This suggests that CH3SH adsorbs dissociatively on this sulfided catalyst. This result could be extended to H2S adsorption, at least for the first doses introduced. 3.5 Effect of H2 adsorption on sulfided Mo and C0M0/AI2O3 Following the same procedure as that used for H2S adsorption, hydrogen has been introduced at r.t on sulfided M0/AI2O3 and C0M0/AI2O3. Contrarily to H2S, H2 introduction does not induce any modification of DMP adsorption : no new species are formed and the number of coordinated species and weakly adsorbed species stays constant This absence of interaction of hydrogen at r.t with the sulfided catalyst surface validates the method used. Indeed, it points out that DMP species interact sufficiently strongly with alumina or the sulfided phases to resist to the introduction of a neutral gas whereas the adsorption is sufficiently weak to be sensitive to H2S introduction.
4. Discussion Sulfidation of alumina at 673 K, followed by an evacuation at the same temperature, does not strongly modify its acidic properties since DMP leads to similar species when adsorbed on sulfided or pure alumina (10). This is in agreement with Ziolek et al. (14) who reported close catalytic properties of alumina before and after sulfidation as well as a very low sulfur content Nevertheless, H2S adsorption at r.t induces some modifications on the alumina acidic properties since it slightly increases the number of Bronsted acid sites and decreases that of Lewis acid sites. On sulfided catalysts, DMP adsorption at r.t provides evidence for the presence of some Bronsted acid sites. These sites are maximum on sulfided C0M0/AI2O3. This shows that such sites are formed on the sulfided phase and it indicates the presence of residual SH groups. Further adsorption of H2S at r.t followed by an evacuation at r.t. leads to the formation of supplementary Bransted acid sites. The number of the sites created varies in the following order : C0M0/AI2O3 > M0/AI2O3 > C0/AI2O3 > AI2O3. These results show that, although H2S adsorption also induces the creation of Bronsted acidity on alumina, the presence of M0S2 slabs favors the creation of Br0nsted acid sites particularly when they are promoted by cobalt atoms. Therefore, such sites are involved in the apparition of acidic SH groups. Introduction of small doses of H2S reveals that the same niunber of sites is created on M0/AI2O3 and C0M0/AI2O3. However on C0M0/AI2O3, the protonated species are more stable. This could indicate that the SH sites present on C0M0/AI2O3 are stronger than on M0/AI2O3. By analogy with results obtained with CH3SH, a dissociative adsorption of H2S on sulfided M0/AI2O3 is suggested. It can occur on couples of coordinatively unsaturated (cus) Mo and S sites. We note that the number
165 of SH sites present after evacuation varies in accordance with the catalytic activity measured in thiophene hydrodesulfurization of these catalysts (15). The introduction of small doses of H2S provides also evidence for a simultaneous decrease of the number of Lewis acid sites and weakly adsorbed species. Taking into account the results obtained on silica, we can proposed that H2S displaces the weakly linked species without creation of Brensted acidity, while the poisoning of cus sites present on alumina and on the sulfided phases leads to the creation of Bronsted acid sites. Comparison with results obtained from H2S adsorption on NiMo/Al203 and NiW/Al203 shows that the number of Bronsted sites created is maximum on C0M0/AI2O3. This result can be related to a study of Van Gestel et al. (16) on the sensitivity to the H2S amount of NiMo/Al203 and C0M0/AI2O3 in HDS of thiophene. At high H2S levels, these authors provide evidence for a higher efficiency of sulfur sites created on CoMo than on NiMo catalysts. In our conditions, hydrogen introduction does not induce any acidity. This does not mean that hydrogen does not generate acidity in conditions closer to those of the reaction. 5. Conclusion This study shows the presence of Bronsted acid sites on the support and more specifically on the sulfided phases, promoted or not. Adsorption of H2S at r.t. increases the number of such sites at the expense of Lewis acid sites. Our results suggest that H2S adsorbed dissociatively, likely on cus Mo and S couples. The number of Bronsted acid sites created by H2S adsorption depends on the catalyst. For the Co-Mo catalysts, their variation is correlated to their activity for HDS of thiophene : C0M0/AI2O3 > M0/AI2O3 > C0/AI2O3 > AI2O3. Finally, on NiMo/Al203 and NiW/Al203, H2S adsorption induces a number of Bronsted acid sites lower than that created on CoMo/AI2O3. Acknoledgements The authors thaiJc J. Van Gestel for stimulating suggestions. REFERENCES 1 - R. R. Chianelli, Catal. Rev.-Sci. Eng., 26 (1984) 361. 2 - F. Mauge, A. Vallet, J. Bachelier, J.C. Duchet and J.C. Lavalley, J. Catal., in press. 3 - F. Mauge, J.C. Duchet, J.C. Lavalley, S. Houssenbay, E. Payen, J. Grimblot and S. Kasztelan, Catal. Today, 10 (1991) 561. 4 - J.A. De Los Reyes J.A., M. Vrinat, M. Breysse, F. Mauge and J.C. Lavalley, Catal. Lett., 13 (1992) 213.
166 5678910 11 1213 14 15 16 -
F.E. Massoth and G. Muralidhar in Fourth International Conference on Chemistry and Uses of Molybdenum (H.F. Barry and P.C.H. Mitchell, Eds.) p.343. Climax Molybdenum Co., Ann Arbor, MI, 1982. N.Y. T0ps0e, H. Topsee and F.E. Masoth, J. Catal., 119 (1989) 252. N.Y. T0ps0e and H. T0ps0e, J. Catal., 139 (1993) 641. P.A. Jacobs and C.F. Heylen, J. Catal., 34 (1974) 267. E.R.A. Matulewicz, F.P.J.M. Kerkhof, L.A. Mouljin and H.J. Reistma, J. Colloid Interface Chem., T7 (1980) 110. C. Lahousse, A. Aboulayt, F. Maug^ , J. Bachelier and J.C. Lavalley, J. Mol. Cat, 84 (1993) 283. A. Corma, C. Rodellas and V. Fomest, J. Catal., 88 (1984) 374. S. Jolly, J. Saussey, J.C. Lavalley, N. Zanier, E. Benazzi and J.F. Joly, Ber. Bunsenges. Phys. Chem., 97 (1993) 313. O. Saur, T. Chevreau, J. Lamotte, J. Travert and J.C. Lavalley, J. Chem. Soc. Farad. Trans. 1, Tl (1981) 427. M. Ziolek, J. Kujawa, O. Saur and J.C. Lavalley, J. Mol. Catal., 97 (1995) 49. J. Bachelier, M.J. Tilliette, M. Cornac, J.C. Duchet, J.C. Lavalley and D. Cornet, Bull. Soc. Chem. Belg., 93 (1984) 743. J. Van Gestel, L. Finot, J. LegHse and J.C. Duchet, Bull. Soc. Chim. Belg. 4-5 (1995) 189.
® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F, Froment, B. Delmon and P. Grange, editors
ORGANO METALLIC SILOXANES AS AN ACTIVE COMPONENTS OF HYDROTREATING CATALYSTS. Kolesnikov, I. M., Yablonsky, A.V., Sugungun, M. M., Kolesnikov, S. L, Kilyanov, M.Y State Gubkin Academy of Oil and Gas 117296 Moscow, Leninsky prospect 65, RUSSIA.
ABSTRACT Structural and contents optimization of hydrotreating catalysts and their subsequent activation with organo metallic siloxane based on the theory of catalysis by polyhedra was discussed.
INTRODUCTION Industrially, for the removal of sulfur and or sulfur compounds from oil fractions under an increased hydrogen pressure and temperature a process known generally as hydrotreatment, various catalysts are used which may contain a mixture of Al-Mo-0 or Al-Ni-Mo-0, and even catalysts containing zeolite additives and many others [1,4 ]. Metal sulfides of Ni-W, Cu-W, Co-Mo, Ni-Mo and others are used for hyrodesulfurization of sulfurous feed [ 5,6 ]. The activity of the oxides catalysts is determined by the ratio of metal oxides in them, type of the metal and operating regimes of the hydrotreatment process in general. The optimal content of the catalysts can be determined by the theory of catalysis by polyhedra put forward by one of us in the 1960s [7,9 ]. Active centers of the hydrotreating catalysts based on one of the postulates of the theory were considered to be the ensembles of the following tetrahedras: [ NiO4.MoO4.AlO4 ] , [ M0S4AIO4 ], [ NiS4A104], [ WO4.M0O4.AIO4 ] and others. Tetrahedras of the types : [ C0O6.M0O4.AIO6 ], [ Ni06.Mo06.A106], [ M0S6AIO6 ] are less active [7,8]. Catalysts are prepared by impregnating Co, Mo and Ni salts to y- AI2O3, that allows a sequential shaping of the polyhedral structures. Uncontrolled interaction of the impregnated salts with y- AI2O3 lattice will results to an alternate and chaotic distribution of both the active and non-active polyhedras at the surface of the carrier and makes the activity regulation much more tedious. Thermodynamic method for optimization of hydrotreating catalysts and the application of organo metallic siloxane as both a catalyst and activator are discussed in this paper.
167
168 EXPERIMENTAL Synthesis of organo metallic siloxanes: Synthesis of the organo metallic siloxanes is carried under laboratory conditions using the following technique:Solutions of metal chlorides ( C0CI2, M0CI5, NiCb or FeCh ) are prepared and dissolved in a pure absolute acetone. Powdered aluminophenylsiloxane (APS ) is dissolved in a pure ethyl spirit. Both solutions are then mixed. The reaction mixture is heated for 1 hour and the solvents (acetone and ethyl spirit) were distilled out. A mixture of aluminophenylsiloxane (APS ) and chloride compounds at the following molar ratios were synthesized. APS :CoCl2= 1.0: 0.15 APS : MoCl = 1.0 : 0.3 ; 1.0 : 0.5 ; 1.0 : 1.0 APS : C0CI2: WCl6 = 1.0 : 0.003 : 0.04 Infra-red spectra within thefrequencyinterval of 400-3600 cm"^ was carried out. Molecular weight of each of the sample compound was determined. Ferro-Molybdenum catalysts were synthesized by crafting FeCls and M0CI5 salts to the surface of the Y-AI2O3 carrier. Surface areas were determined on "Sorptomatic". Activities of the organo metallic siloxanes were determined in the homogenous thiophene desulfiirization from toluene. Catalysts activities in hydrothermal decomposition of sulfiir compounds in micro flow reactors in the stream of helium gas were determined. Chromatographic analysis of the products was carried out off-line. THERMODYNAMICS OF POLYHEDRA TRANSFORMATION Activity of the catalysts (types: C0O.M0O3.AI2O3, NiO.MoGB.AbOB ) is defined by the ability of ensemble of theirs tetrahedras to redistribute electrons as in the scheme: electron
[iMe^^04] + [2Me^^04] <^ [iMe^'"04] + [2Me^'"04]
(1)
The electron distribution happens under the effects of temperature or inter reactions of the reacting molecules. These characteristic catalyst's behaviors allows the interchange of CoO or NiO in them by FeO for the production of similar catalysts of type FeO.MoOs.AbOs. Equilibrium reaction constant in agreement with equation (1 ) above can in that case be written as : Kc =
rFe^^04l.rMo^'^0.1 [Fe'^04].[Mo'^04]
(2)
169 Solid catalyst's equilibrium concentrations in the ensembles are determined as follows: [Fe^'^04]= - ^ - x
; [Mo^'^04] =
(3)
r-x
[ Fe^'^04 ] = [ Mo^''04 ] = X Active centers concentration in catalysts are less than the sum total number of polyhedras as such we can write: a --X a+b
»
b --x a-\-b
X and
»
x
(4)
Putting (4) and ( 3 ) in (2 ), we can derive
K = a/b/
(5)
Ax + lY
and
^=V^
(6)
a
Differentiating ( 6) with respect to a/b and equating to zero we get,
4a/by
faV
='
(7)
From where, a=b
(8)
Subsequently, an optimal catalyst content can be considered to be: 99.9 % wt. FeO , 20.1 % wt. M0O3, 70 % wt. AI2O3. Catalysts of various contents were synthesized in the laboratory and theirs respective activities in thiophene hydrotreatment were determined, results are shovm in table 1 below.
170 Table 1. Optimal contents of the hydrotreating catalyst (FeO.MoOs.AI2O3). Catalyst M0O3 (% wt.) FeO (% wt.) AI2O3 (% wt.) I 18.51 11.49 70 II 20.01 9.99 70 III 21.51 8.49 70
Conversion 92.3 99.8 91.6
It can be noticed from the table that, the experimental results corresponds to the predicted optimal values calculated earlier before above with only a very slight deviation. Optimizing the y-AhOs contents in the catalysts will also make it possible to attain its maximal activity. Thermodynamic computations gives the optimal contents of C0O.M0O.AI2O3 as [CoO] 35 wt. % , [ M0O4 ] - 65 wt. %. The experimental results are shown in table 2 below. Table 2. Optimization of Al-Co-Mo-O catalyst. Catalyst CoO ( % wt.) M0O3 (% wt.) I 3.32 1.75 II 3.50 6.50 III 5.8 10.70 IV 7.80 5.0
AI2O3 ( % wt.) 95 90 83.5 77.1
Conversion 11 45 77 56
Catalyst III was successfully produced in a commercial scale. ACTIVATION OF SOLID CATALYSTS. Catalyst III in table 2 above was used as a sample for activation. Cobalt Phenylsiloxane, Ferro Phenlysiloxane and Tungsten Phenyl Siloxane were applied as the activators. The catalyst was calcined at 550 °C for 3 hours, after which it was mixed with solution of Methyl Phenyl Siloxane in Toluene prepared earlier. The mixture was left overnight at 20°C and the solvent (toluene) was distilled out. The modified catalysts was tested in the a microreactor in the stream of helium and hydrogen at 450°C in the hydrodesulftirazation of thiophene. Activity data for Al-Co-Mo-0 , modified Co-Ph-Si, Fe-Ph-Si and W-Ph-Si in the hydrodesulftirazion of thiophen is shown in figure 1 , tables 3 and 4 below.
171 Fig.l. Effect of Concenration of CoPhSi on conversion at various tenperatures.
-^593K -*-623K - - 653K
— I —
0.25
0.5
— I —
1
1.5
— I
2.5
Mass % of CoPhSi
Table 3. Effect of the content of Fe-Ph-Si on the surface of Al-Co-Mo-O catalyst in hyrodesulfurization of thiophene. Tiophene conversion wt. % at temperature Fe-Ph-Si % v^. 653 673 593 523 69 72 36.0 7.1 0 88 84 66 26 0.5 71 85 96 16 1.0 95 94 83 35 2.0 72 96 95 34 3.0 Tungsten Phenyl Siloxane 43 63 3 0.25 68 69 6 0.50 1.0 6 56 58 It can be noticed from the data that, modified catalyst's activity changes non linearly through a maximum. In all cases the modified catalyst's activity is greater than the original initial unmodified. Clear explanations of these data can be received from analysis of the catalyst's IR-spectra. INFRA-RED SPECTRA. Infra-red spectra of the Al-Co-Mo-0 and the W-Ph-Si, Co-Ph-Si catalysts were obtained and the results are shown in figure 3 below. Spectral absorption bands for the Al-Co-Mo-0 catalyst at 500-900 cm"^ can be explained as due to the presence of condensed polyhedras in it. These bands are precise in the spectra of Co-Ph-Si and W-Ph-Si.
172 Figure 2. IR-spectra: 1-AlCoMoO, 2- WPhSi, 3- AlCoMoO + 0.25 wt. % WPhSi, 4- CoPhSi, 5-AlCoMoO + CoPhSi.
3iOD
3H0O
3000
2600 i9oo
ISiO
i^OO S^oo
yoo Srfff
In accordance with literature data absorption bands at the frequency interval of 730-810 cm"^ corresponds to the oscillatory motion of Co-0 and W-0 bonds in the "isolated " [ C0O4 ] and [ WO4 ] tetrahedras. It can then be concluded that. Cobalt - Tungsten Siloxane compounds increases the number of active centers on the activated catalyst's surface. Al-Co-Mo-0 catalyst was modified with Phenyl Silocsane of the following metals: Cu, Fe, Co, Ni, Mo and W, histogram of these modified catalysts is shown in figure 3 below.
173 Fig. 3. Effect of the nature of metal in the modified Al-Co-Mo-O catalyst on the conversion of thiophene ( T = 673K, Me-Ph-Si = 1 % wt.).
^
100 n
n
80-
0
V
e
60-
' s i
40-
0
n
20 n-
^^
1
I I I
1II 1 I1 I1I1 "
I "
Cu
1
Fe
1
Co
1
Ni
1
1
Mo W
Metal Phenyl Silocsane It can be noticed from the histogram that, the most active among the Silocsane are Fe-Ph-Si and Co-Ph-Si. ACTIVATION OF COMMERCIAL HYDROTREATING SPENT CATALYSTS Hydrotreating catalysts, e.g. NiO.MoOa.AliOs used in the commercial units for purification of petroleum fractions deactivates after a prolonged period of exploitation in the units. This deactivation among other reasons was mainly due to structural reorientation of the active components. Usual regeneration of such deactivated catalysts with aim of retaining the activity is in most cases impossible, reactivation in that case is possible only by employing chemical methods. We here recommends activation of hydrotreating commercially spent catalysts ( Al-Ni- Mo-0 ) with Cobalt Phenyl Siloxane and Ferro Phenyl Siloxane. The activity of Al-Ni-Mo-0 commercially spent catalyst in hyrodesulfiirization of vacuum gas oil is in the range of 35 - 40 wt. %. Hydrodesufrilrization of thiophene with commercially spent Al-Ni-Mo-0 catalyst and Metallorganic Siloxane modified catalyst was investigated in a catalytic microreactor. Data of these experiments is shown in table 4 below.
174 Table 4. Thiophene conversion on spent and modified catalysts Conversion, at temperatures, K Catalyst 638 653 563 623 30.4 Spent ( A ) 47.7 63.6 65.9 A + 0.30 % 47.1 70.6 77.1 82.3 wt. CoPhSi 81.5 A + 0.36 % 49.3 70.7 87.6 wt. CoPhSi A + 0.25 % 36.5 45.3 66.8 wt. FePhSi A + 0.50 % 42.5 50.7 67.5 wt. FePhSi A+1.00% 41.9 49.7 66.9 wt. FePhSi A + 2.00 % 30 46.5 61.5 wt. FePhSi
673 68.2 91.2 93.5 68.7 70.6 70.0 65.7
Following conclusion can be made from the above data: 1. The activity of the siloxanes are in order CoPhSi > FePhSi 2. Modification of the spent catalyst with CoPhSi greatly increase the initial spent catalyst conversion, i.e. from 68.2 wt. % to 93 wt. % 3. Ferro Phenyl Siloxane can also serves as a modifier of Al-Ni-Mo-0. 4. Conversion of catalysts modified with Ferro Phenyl Siloxanes follows a non linear relation that passes through a maximum. Organometallic siloxanes are capable of increasing and / or retaining the activity of both fresh and spent hydrotreating catalysts at high temperatures. Solid catalysts to the surface of which 0.25 - 0.5 wt. % of siloxanes is crafted exhibits a maximum conversion. CATALYTIC HOMOGENOUS CONVERSION OF SULFUR COMPOUNDS Organometallic Siloxane can exhibits a considerable catalytic activity in homogenous conversion of sulfiirous compounds. Catalytic activity of organometallic siloxane in a thermal conversion of a model mixture of 6 % solution of thiophene in toluene to which 2-20 wt. % of catalyst is added was investigated. Experiment was carried out in an autoclave at an initial hydrogen pressure of 12 atmosphere. The autoclave was heated to 423 K for 1-5 hours. AlCoWPhSi, AlMoPhSi, AlCoPhSi and AlPhSi were applied as catalysts. Results of the experiments are shown in table 5 below.
175 Table 5. Thiopene conversion on AlCoWPhSi ( P = 12 atm.., T = 423K, catalyst concentration = 20 wt. % ). Time, T (hours ) Thiophene conversion, x y (number experimental (wt. %) repetitions) 3.3 6 1 8.1 8.8 2 17 20 3 5 28 19 It can be seen that, thiophene conversion increase exponentially on increasing the reaction time. On regeneration of the catalyst sample taken from the reaction mixture the initial catalyst activity is almost fully regained. Effect of the concentration of added AlCoPhSi on the thiophene conversion is shown in figure 4. (Note: A = AlCoPhSi, B=AlMoPhSi) Fig. 4 Effects of catalyst concentration and time on homogenous conversion of thiopene in toluene. 40 n 35 3025x,%20 1510-
•
•
^—g
—
-4-A2%wt. Hli-A10%wt. -'-A20%wt. -Hie-B10%wt - • - B 20%wt.
5U 1W'
()
1
1
1
2
1
1
1
3
4
5
Time, t ( hours )
Thiophene conversion as can be noticed gives an S-shaped curves. Homogenous catalytic conversion of thiophene increases with increase in the concentration of the catalysts. The activity of the metallorganic siloxane as was concluded from experimental data increase in the following order: AlMoPhSi >AlCoPhSi >AlCoWPhSi. This allows us to noticed also that, in the hyrodesulfurization of thiophene the activity of the ensemble of tetrahedras are in order : [ AIO4. Mo04.Si04 ] > [ AlO4.CoO4.SiO4 ] > [ AlO4.CoO4.WO4.SiO4 ] .
176 EFFECT OF AlPhSi: MeCln MOLAR RATIO ON CATALYST ACTIVITY The organometallic siloxanes were synthesized by inter reacting AlPhSi with metal chlorides (e.g. M0CI5 or C0CI2). Catalytic activities of compounds with different AIO4 : Me04 ratio in desulfurization of thiopenes were investigated. Results of such investigation are shown in table 6. Table 6. Activity of metallorganic siloxane with different AlPhSi: MeCln ratio. Catalyst AlPhSi: MeCln Parameters Temperature K Time, hours X,( %) 1.0:0.15 AlMoPhSi 150 3 0 1.0:0.25 AlMoPhSi 150 3 20.6 38.4 1.0:0.30 AlMoPhSi 150 3 AlMoPhSi 1.0: 1.0 150 3 52.3 1.0:0.15 150 1 AlCoPhSi 9.3 1 AlCoPhSi 1.0:0.5 150 10.7 AlCoPhSi 1.0: 1.0 150 1 31.6 Hyperbolic change in activity of the compounds on varying the AlPhSi: MeCln ratio in homogenous thiophen hyrodesulfurization can be noticed from the above table. This can possibly be connected to a numerical increase of the ensemble of [AlO4.MoO4.SiO4] tetrahedras per unit weight of the organo metallic siloxane. TETRAHEDRAL ENERGY STATE AND ACTIVITY Activity of a given ensemble of tetrahedras increases with increase in energy of electrostatic field caused by the presence of tetrahedras of types [ Mo^^04.Si'^"^04 ] or [ Fe^^04.Si'^'^04 ]. In these ensemble positive fields were developed by cations while the negative electrostatic fields are due to the anions and or oxygen ions. Tetrahedras in a solid catalysts lattice may be oriented at the surface by theirs apexes F, edges C, or facets S. Tetrahedras of different orientations at the phase surface will develop electrostatic fields of different configurations. The value of such electrostatic fields can be numerically determined. For computation of fields developed by the ensemble of MeO[A104]3 tetrahedras following assumptions were made: 1. All tetrahedras are ideal and symmetrical. 2. Tetrahedras are interconnected by a bridge of oxygen ions. 3. All positive charges are discretely distributed among the cations and negative charges on oxygen ions. 4. Charges are considered as point charges. Electrostatic potential developed at any point of the tetrahedras by cations and ions can be calculated by the formula:
177 N
where (Ji is a point charge, ^p J ^i radius vector along the axis of the nucleus, N number of ions in the tetrahedras. Potential fields developed by FeO[A104]3 and MoO[A104]3 are calculated in table 7 below. Table 7. Potential field developed by ensemble of tetrahedras with different orientation. Distance FeO[A104]3 MoO[A104]3 from the center, A S C F S C F -3.32 -4.12 27.8 0.76 37.4 1.20 0.1 7.87 -2.82 0.90 -2.79 0.50 10.7 0.3 -2.82 -2.04 3.86 0.30 5.42 0.66 0.5 2.12 0.14 -1.39 0.50 -1.59 0.7 3.20 0.98 -0.04 0.25 -1.15 -1.08 1.0 1.65 Electrostatic field developed in the ensemble of tetrahedras by Fe^"^ ions is greater than those developed by Mo^"^ ions which explains the difference in activities of the two in a solid catalyst as can be noticed in figure 2. Ensemble of Co^'^0[A104]3 and Mo^"'0[A104] were as well investigated and the difference in electrostatic fields confirm the high activity of AlMoPhSi catalyst over AlCoPhSi (Fig. 4 ). PROPOSED REACTION MECHANISM FOR HYDRODESULFURIZATION. Catalysts active centers for the hydrodesulfurization of sulfur compounds from oil fractions are in accordance with theory of catalysis by polyhedra considered to be the ensemble of Mo[A104], CoO[A104], FeO[A104] and other similar tetrahedras. Sulfiir compounds under the effect of electromagnetic and electrostatic fields developed by those tetrahedras are absorbed and excited at the surface of the ensembles [ 9 ]. Simultaneous excitation and absorption of the sulfiir compounds and hydrogen molecules is however required. Successive stages of hydrodesulfiirization of thiophene is sketched in figure 5 below.
178 Fig. 5 Stages of thiophene conversion over the ensemble of [ AIO4.C0O4.M0O4 ] tetrahedras.
1
^E C4H4S* V UOICARAS] C4H4S AI.C0.M0.O <>ut. Al-Co-Mo oAs^. Al-Co-Mo Catalvst A ^ Al-Co-Mo ^'A^-c^//g
^ ^ J H2*.C4H4S1 e.%ct.
«A?»-y •H
Cs-^^^"
^^A<)
C4H*6.H 2
Al-Co-Mo
^^-tr. Ci,«^^H* C4H5SH Al-Co-Mo
C4H8 Al-Co-Mo
C4H 8 Al-Co-Mo
Al-Co-Mo
C4H5SH /H2 Al-Co-Mo AX^-tr^-
^c-
C)<-i^-
Cs
C4H 6:H2 ^ C4H 6 ^ lAl-Co-Mo ^'^^ Al-Co-Mo
C4H 6.H2S Al-Co-Mo
C4H5SH H 2 Al-Co-Mo
Following stages are outlined in the scheme above :• Absorption of thiophene molecules • Excitement of thiophene molecules • Adsorption and excitement of H2 molecules • Bonds distribution between the two excited molecules with formation of semihydrated exited molecules • Adsorption and excitement of a second H2 molecule • Inter action of exited C4H5SH and H 2 molecules with formation of butadiene and H2S • Desorption of the H2S molecules from the ensemble of the tetrahedras due to the destruction of the orbital symmetry • Adsorption of H2 molecule on the exited C4H 5 • Excitement of H2 molecules • Bonds distribution between the exited butadiene molecule and exited H2 molecule with formation of an exited 1-butadiene molecule ( C4H*8) • Energy emission by the exited 1-butadiene molecule
179
Q^8
^ Q^8
Desorption of C4H8 molecule at ground state from the ensemble of the tetrahedras due to misbalance of orbital symmetry with a subsequent regeneration of [ AIO4.C0O4.M0O4 ] tetrahedras that can enter into the catalytic circle again.
All the proceedings stages can fully be explained by methods such as molecular dynamics, electrostatic theory, quanta-chemical theory, kinetics and thermodynamics of elementary processes, thermodynamic of transition states, irreversible processes and coordinated systems. REFERENCES 1. M. I. Loctev, A. A. Slikin, The Structure and Physico-chemical properties of AlCoMo catalysts, (in Russian), Moscow, VINITI, (1980). 2. F. E. Massoth, J. Catalysis, (1973), V.30, p.204. 3. P. Cajotrdo, A. Mathieux, P. Grange, B. Delmon, Appl. Catal, (1982), V.3, N.4, p.347. 4. B. K. Nefedov, E.D. Radchenko, R. R. Aliev, Catalysts for processes of intensive treatment of petroleum. (in Russian), Moscow, Chemistry, (1992 ). 5. A. N. Starsev, Catal. Rev.-Sci. Eng., (1995 ) , V.37, p.353 6. W. Nieman, B. S. Clausen, H. Popsoe, Catal. Lett., (1990), V.41, p.355 7. I. M. Kolesnikov, Introduction to the Theory of Catalysis by Polyhedra, Cherkassy, NIITEKHIM.Zh, (1977), N1455/77, 81p , ibid., (1989 ), N 343-hp-89, 195p. 8. I. M. Kolesnikov, S. I. Kolesnikov, Oxidation Commun., (1995 ), V.18, Nl, p. 1-20. 9. I. M. Kolesnikov, The Generalized Quantum Chemical Principle and Mechanism of Catalytic Processes, Cherkassy, ONITECHIM, Moscow, (1981 ), 198p.
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® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
181
Alumina supported HDS catalysts prepared by impregnation with new heteropolycompounds A. Griboval\ P. Blanchard\ E. Payen\ M. Foumier\ J.L. Dubois^ ^ Laboratoire de Catalyse Homogene et Heterogene, URA CNRS n° 402, Universite des Sciences et Technologies de Lille, Bat. C3, 59650 Villeneuve d'Ascq, FRANCE. Tel: (33) 20.43.49.47, Fax : (33) 20.43.65.61. ^ Centre de Recherche Elf Solaize, Elf Antar France, BP 22, 69360 Solaize, FRANCE Tel: (33) 78.02.60.93, Fax : (33) 78.02.60.16. Novel heteropolyanions, with various Co/Mo ratio have been synthesized and impregnated on an alumina carrier using the standard incipient wetness method. These heteropolycompounds have the advantage over the conventional impregnating solutions to have the metals in the same compound. The efficiency of these catalysts for the hydrotreatment has been compared to conventional catalysts with the same metal loading using thiophene HDS as the reaction test. It appears that these heteropoly based catalysts, in which the Co/Mo atomic ratio is easily monitored are very effective for hydrodesulphurization.
1. INTRODUCTION The use of Mo based catalysts supported on alumina and promoted by Co or Ni is well established to hydrotreat petroleum fractions. Lowering the sulphur content of refined products by hydrotreatment is required to respond to the new regulations. In order to achieve this goal, new formulations or new preparation procedures, to produce more active catalysts, are preferred instead of using more drastic conditions in refineries. Additives are generally used to improve the activity of classical C0M0/AI2O3 catalysts, and phosphorus is a common one. Although many studies have been devoted to explain the role of this additive (in the active phase as well as in the preparation process of the oxidic precursor), its precise role remains still unclear. The presence of P2Mo5023^' anion in the impregnating solution of MoP based catalysts has been proposed [1-2]. For the CoMoP catalyst it has been shown that both active metals Co and Mo should be in close contact in the impregnating solution [3]. However a well defined heteropolycompound including the promotor has never been clearly identified in the impregnating solution nor on the oxidic precursor. Molybdenum is a versatile element and may create various structures, the Keggin and the Anderson heteropolyanions being well known. In the last decade solid heteropolyacids and their salts have been extensively used as catalysts for the mild oxidation [4]. In hydrodesulphurization (HDS) interesting results were reported in the literature [5] with alumina supported Anderson molybdenum heteropolycompounds of the general formula (H6XrY6024)''' (Y = Mo, W;
182 X = Co^^, Co^^, N P ^ . , . ) , However their use induces a low Co content as promotor and thereby a relatively low activity. The 12-molybdodophosphoric acid was also proposed as precursor for the preparation of HDS catalysts [6]. The aim of the present work was to prepare HD^ catalysts with heteropolyanions of the Keggin structure as starting material. Novel heteropolyanions, with various Co/Mo ratios have been synthesized. A reduction step was necessary to increase the Co/Mo ratio and various procedures have been used. These new heteropolystructures were impregnated on an alumina carrier using the standard incipient wetness method and physical characterizations were performed at various steps of the synthesis of the oxidic precursor. The efficiency of these catalysts for the hydrotreatment is compared with that of catalysts with the same metal loading prepared by conventional methods using thiophene HDS as a reaction test.
2. EXPERIMENTAL 2.1. Sample preparation All chemical reagents used in this study (Na2Mo04, H3PO4) were purissim Fluka Chemicals. Raw heteropolyacid The pure heteropolyacid H3PM012O40 (H3PM0) or H4SiMoi204o (H4SiMo) were prepared according to literature data [7]. The C03/2PM012O40 (C03/2PM0) as well as the Ni3/2PMoi204o (Ni3/2PMo) salts were obtained through the procedure detailed in Figure 1. H3PM012O40, 13H2O solution^3H\PMoi204o^
3/2 Ba(0H)2
Ba3/2PMOi2O40 3/2 C0SO4
or 3/2 NiS04 C03/2PM012O40 orNi3/2PMoi204o + 3/2 BaS04
Evaporation
> bulk compound
Figure 1. Preparation of C03/2PM012O40 or Ni3/2PMoi204o The Co2SiMoi204o (Co2SiMo) was prepared through the same procedure and instead of the sulphate salt, nickel or cobalt carbonate could be used. Reduced heteropolyacid The starting material, PMo 12040^" (H7PM0) was prepared by reaction under inert atmosphere, of a mixture of sodium molybdate, phosphoric acid and MoOCls^' obtained by hydrolysis of M0CI5 in a 3M HCl solution according to the following equation [8]: H3PO4 + 8 Mo04^" + 4 MoOCls^" + ¥C-^ H4PMoi204o^" + 20 CI"
183 The procedure is detailed in Figure 2. Preparation of an out gassed solution H20-H3P04( 1 M)-Na2Mo04(2M) addition of MoCl5-HCl(3M) solution heating 5 hr at 353 K addition of HCl crystallisation at 273 K filtration drying Figure 2. Preparation of H4PM012040^' acid The C07/2PM012O40 (C07/2PM0*) was obtained by adding to a solution of the purified H7PM0 acid, cobalt sulphate, after neutralisation with barium hydroxide, or cobalt carbonate. Preparation could also be performed by adding Co metal to a solution of the purified H3PM0 acid followed by a substitution of protons by cobalt ions. The following solids were thus obtained: H3C0PM012O40 (H3C0PM0), H3C02PM012O40 (H3C02PM0), H4Co2SiMoi204o (H4Co2SiMo), C05/2PM012O40 (C05/2PM0), Co7/2PMo,204o° (Co7/2PMo°), Co4SiMoi204o(Co4SiMo), Co7/2SiMoi204o(Co7/2SiMo). In order to compare the effect of the sequence of the reaction, two procedures were used i.e. reduction before or after the substitution of the H"" ions. In order to maintain the pH at a constant value during the reduction process, citric acid, the PKA of which being 3.5, was used as a buffer. 2.2. Catalyst preparation The catalysts were prepared by incipient wetness impregnation of a y AI2O3 (pore volume: 0.8 cm^.g'*; specific surface area: 250 m^.g'^) with solutions containing the appropriate amounts of molybdenum, i.e. 14 wt % as M0O3. All the well defined heteropolycompounds were dissolved in the appropriate volume of water and the solution used for the impregnation. For comparison purposes, catalysts (CoMoPl, CoMoP2) with the same Mo loading have been prepared by the standard impregnation method using a solution containing ammonium heptamolybdate, phosphoric acid and cobalt nitrate in the appropriate amounts to prepare two reference catalysts: -CoMoPI: 2 wt % P2O5, 3 wt % CoO, 14 wt % M0O3 (Cc/Mo = 0.41, P/Mo = 0.3) -CoMoP2: 1 wt % P2O5, 2 wt % CoO, 14 wt % M0O3 (Co/Mo = 0.27, P/Mo = 0.15, i.e. the stoichiometry of the C07/2PM0 /AI2O3 catalyst) The impregnated extrudates were dried at 353 K overnight and then calcined at 673 K for 4 h, under N2 or air as stated. The nomenclature of the supported catalysts recalls the molybdenum precursor used. The Mo content was determined using X-ray fluorescence by the Central Analytical Division of the CNRS (Vernaison-France).
184 2.3. Characterizations X-ray diffraction (X.R.D.) was carried out with a Siemens D5000 difFractometer equipped with a monochromator and a Cu X-ray tube. Fourier Transform infrared (FTIR) spectra were recorded in the 400-4000 cm"* spectral range with a Nicollet 510 spectrometer using pellets with spectroscopic grade potassium bromide. Laser Raman spectroscopy (LRS) was performed using a Raman microprobe (XY from Dilor), equipped with a photodiode array. The exciting light source was an Ar"*" laser emitting the 488 nm line with a power at the sample of 10 mW. UV-Vis diffuse reflectance spectra (DRS) were recorded, in the 250-900 nm spectral range, with a diffuse reflectance accessory. Samples were placed in quartz cells and were scanned against a pure alumina background. The X-ray photoelectron spectra (XPS) were obtained by using an AEI 200 EB spectrometer equipped with an aluminium X-ray source working at 300 W. Binding energies were measured by reference to Cls peak at 285 eV. The ^*P NMR measurements of the impregnating solutions were carried out on a Bruker AC300 spectrometer. Chemical shifts are negative towards higher field and are referenced to external 85 % H3PO4 as a standard. The solvent used was H2O and all the solutions were prepared at 0.14 M i.e. the concentration of the impregnating solutions. The ^*P measurements of the bulk compounds and the supported catalysts were carried out on a Bruker ASXIOO, operating at 40.53 MHz in MAS experiments. The pulse length and the relaxation time were respectively equal to 1.8 |is and 20 s. Catalytic activities for thiophene HDS were measured at atmospheric pressure in a flowtype reactor packed with 200 mg of catalyst. The solids were first sulphided with a H2S/H2 (10/90) mixture at a flow rate of 100 cmVmin at 673 K for 2 hr and then cooled down to 573 K. After purification by vacuum distillation, thiophene was introduced in the reactor at constant pressure (50 torr) in a flow of purified hydrogen (20 cmVmin.). The reaction products were analyzed by gas chromatography.
3. RESULTS AND DISCUSSION 3.1. Bulk Compounds Chemical analysis confirms the stoichiometry of the various bulk heteropolycompounds which were synthesized. Figure 3 shows the XRD pattern of the phosphomolybdate samples. The broad peaks observed on the high 20 range characteristic of the oxygen lattice of the H3PM0 acid [9] are observed on the XRD pattern of the C03/2PM0 (Fig 3b) and Ni3/2PMo (not reported here) solids. This means that the Keggin structure is preserved upon substitution of the HT by Co or Ni. However a shift of the peaks in the low 20 range as well as a variation of their relative intensities are observed, suggesting that the introduction of Co induces a modification of the lattice parameters. The XRD pattern of the reduced entity H7PM0 (Fig 3c) presents large similarities with the previous one. This means that upon reduction the Keggin structure is also preserved. Whatever the preparation procedure, the XRD features of the C07/2PM0* (Fig 3e) and Co7/2PMo° (not reported here) are similar but differ from those observed for H7PM0. This suggests that the substitution would induce a change of the crystal structure even though the Keggin structure itself is preserved.
185 It should also be pointed out that the X R D pattern o f the H3C02PM0 (Fig 3d) is similar to the C07/2PM0* one. This confirms the reduction o f the H3PM0 by C o metal and the aforementioned effect o f the Co on the crystal structure. Similar evolution o f the X R D features were observed for silicomolybdates compounds, the X R D pattern o f which are not reported in this work.
(a)
''••^V-T"\
15
,
.
,
35
•
55
(b)
75
29
J#T
15
• • .
35
55
j[jM^ m^MJ^^K^ 15
35
I 26 ^ 75
55
.'1.
75
29
(d)
(c)
,I!U
.. •
|illi'V^-->--.^ IST
35
55
I 20
75
(e)
W'^-''^*'^^*^"
'"'•^S-NW..
15
35
55
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29
Figure 3. X R D pattern of: (a) H3PM0, (b) C03/2PM0, (c) H7PM0, (d) H3C02PM0, (e) C07/2PM0*
The IR spectra o f these samples are shown in Figure 4. The main lines characteristics of the Keggin structure at 1065 and 965 c n \ \ assigned respectively t o the PO and MoOt (the subscript t designs a terminal M o = 0 bond) stretching modes as well as the 870 (br) and
186 790 cm"* characteristic of the MoOMo bridge are observed. No drastic modification is observed upon substitution of the PT by the Co atoms (Fig 4a, 4b, 4c) which indicates that the structural Keggin unit is maintained.
WAVENUMBER (cm"*) Figure 4. FTIR spectra of (a) H3PM0, (b) C03/2PM0, (c) Ni3/2PMo, (d) H3C02PM0, (e) C05/2PM0, (f) Co7/2PMo°, (g) H7PM0, (h) C07/2PM0* Upon reduction the intensities of the PO and MoOMo bands (Fig 4g) decrease while that of the MoOt band remains almost unchanged except the splitting. These peculiar changes were well described by M. Fournier & al. and ascribed to a bipolaron effect [10]. Upon substitution of all the H^ ions the IR spectrum remains unchanged (Fig 4h), except a small shift of the band due to terminal MoOt bond. The spectra of the H3C02PM0 and C07/2PM0 (Fig 4d, 4f) present some similarities with the previous one which confirms that the reduction takes place through
187 addition of Co metal. But two bands at 1065 and 1045 cm'\ even more intense for the C05/2PM0 solid (Fig 4e), are observed. The former one should be ascribed to unreduced H3PM0 whereas the latter one would likely correspond to PC0M011O40 ion [11]. However whatever the sequence of this direct preparation with Co metal we may consider t^^t the reduced entity is the major component of the sample. LRS confirms the results obtained by IR and XRD and the Raman spectra are reported in Figure 5. The Raman spectrum of the cobalt salts i.e. C03/2PM0 (Fig 5b) or Co2SiMo (Fig 5d) are similar to the spectrum of the parent heteropolyacids (Fig 5a, 5c). However a shift of the MoOt stretching mode at 974 and 993 cm"^ is respectively observed for the Co2SiMo and C03/2PM0 samples. This shift is characteristic of the substitution of the H" ions by a polarizing cation which induces a modification of the anion-anion interaction [8]. This is in agreement v^th the variation of the lattice parameter suggested by the XRD features.
(e)
993
(0
(b)
J
970 902
251 603
551
987
^^^
J
v(cm"*)
(c)
r904
^^278
251 I \
Mi W)
955
«°
r
f,
v(cm')
Figure 5. Raman spectra of (a) H3PM0, (b) C03/2PM0, (c) HjSiMo, (d) CozSiMo, (e) H3C02PM0, (0 Co7/2PMo°
188 The spectra of the reduced heteropolycompounds H3C02PM0 and Co7/2PMo° (Fig 5e, 5f) show a line at around 920-930 cm"^ However the lines of bulk M0O3 (820 and 990 cm"^) are also evidenced which suggests that a decomposition has occurred which could be due to the analysis method (heating by the laser beam). X-ray photoelectron spectra of the Mo3d electrons state are reported in Figure 6. The spectrum of the C03/2PM0 (Fig 6a) exhibits the MoVI 3d^^ and 3d^^ characteristic of the Keggin unit at 233.1 and 236.3 eV [12]. For H7PM0 (Fig 6c) this doublet is broadened on the low binding energy site which induces a small shift in agreement with the four-electron reduction of this heteropoly compound. Similar spectra (Fig 6d, 6e) are observed for the H3C02PM0 and C07/2PM0 samples whatever their preparation method. The Mo 3d XPS features of C05/2PM0 solid (Fig 6b) shows an intermediate behaviour between unreduced and four-electron reduced compounds which suggests a two-electron reduction.
El(eV)
^
248
243
238
Figure 6. Spectra of the Mo3d electrons of (a) C03/2PM0, (b) C05/2PM0, (c) H7PM0, (d) H3C02PM0, (e) Co7/2PMo° All the aforementionned physical characterizations show that the Keggin unit is preserved upon reduction and/or substitution of the H"*" ions.
189 3.2. Solutions The impregnating solutions were characterized by ^^P NMR. The spectrum of 12-molybdophosphoric heteropolyacid (not reported here) is in agreement with the literature data (main line at 5 = - 3.14 ppm) [2,13-15]. This spectrum also shows two lines of lower intensity at 6 = - 0.93 ppm and 5 = + 0.17 ppm which are ascribed respectively to H6PMo9034^" and to H2PO4" moiesties. The C03/2PM0 exhibits the same NMR features (Fig 7b) but all the spectrum is shifted toward the low fields. For comparison purpose the spectrum of a solution of H3PM0 in which cobalt nitrate has been introduced (Co/Mo = 0,12) is also reported in Figure 7a. The same three lines are observed but at a lower shift. Such a shift could be ascribed to the paramagnetic effect of the cobalt atom and these results. indicate that a strong interaction between the Co and the heteropolystructure should exist in the C03/2PM0 solid. (a)
l«WV^>>A»VHMiy»J»l»»V'>M*>/v'»W^>**vl^'
15
v * ^ j » 'y>;vY^^.» f^^Aj^^^^mtti-^A
-10
10
(b)
6(ppm)
15
-10
10
Figure 7. ^^P-RMN Spectra of: (a) H3PM012O40 +l,5Co(N03)2, (b) C03/2PM012O40
In Figure 8 is reported the NMR spectrum of the H7PM0 which shows a main line at - 12.3 ppm, characteristic of the existence in solution of the P-form of the four-electron reduced heteropolyacid [14].
5(ppm)
- 2 - 4 - 6 - 8
-10
-12
-14
Figure 8. '^P-RMN spectra of H7PM0
190 It is well known that the reduction of 12-molybdophosphoric acid results in the transformation of the a-form (Keggin structure) into the P-form because the reduced P4-form (i.e. four-electron reduced)) is more stable than the a-form (a4) [14]. This result confirms the presence of a four-electron reduced molybdophosphoric acid and indicates that this compound remains reduced when it is dissolved in water. A line is also observed at - 5.32 ppm which correspond to the a2-form i.e. reduced by two electron [14] (with a molar ratio a4/ P4 of about .2) It should be noticed that the solutions which contain a reduced heteropolycompound are impregnated right after the dissolution of the polyanion in water. The spectrum of the C07/2PM0 (not reported here) is similar to the spectrum of H7PM0 The only difference results in the shift (A5 « 18 ppm) and a broadening of the lines which are due to the presence of cobalt. However, the shift is greater than it would be expected if we only consider the magnetic susceptibility. The broadening of the line, which is in relation with the relaxation time, is ascribed to the strong interaction of the Co with the heteropolystructure. So it can be deduced that the reduced heteropoly compounds are stable in solution. These results also show that a strong interaction exists between the Co^^ ion and the heteromolybdate entities. 3.3 Alumina supported heteropolyanions The alumina supported heteropolycompounds were characterized by DRS and ^^P MASNMR. The spectra are respectively reported in Figure 9 and Figure 10. The UV spectrum of alumina supported C03/2PM0 (Fig 9a) shows a main band at 550 nm representative of a cobalt aquocomplex [16] whereas the dried alumina supported H7PM0 (Fig 9b) exhibits the band characteristic of the p4 form at 685 nm [17], This shows that upon impregnation the nature of the heteropolycompound is not modified. A band is also observed at 344 nm which is not identified with the available data. Upon calcination in N2 of the C07/2PM0 a broad band is observed between 500 and 700 nm whereas the bands usually observed for cobalt supported on alumina are observed after calcination in air [18]. The ^^P NMR measurement of bulk H7PM0 (Fig 10a) shows a line at - 12.41 ppm, which corresponds to the p4-form.The presence of cobalt in the bulk Co7/2PMo° (Fig 10b) induces a shift (5 = -10.25 ppm) and a broadening of this line but the variation is lower than the one observed for the solution, so this suggests a lower interaction of the cobalt with the heteropolyanion in the bulk compound. After impregnation on the alumina and drying in air, the line of Co7/2PMo° (Fig 10c) is observed at - 8.25 ppm. The calcination in air (Fig lOd) does not influence the results (5 = - 8.21ppm). The presence of a single line in the ^^P NMR spectra as well as its position are indicative of the preservation of the heteropolyanion structure after the impregnation on alumina and calcination, but the greater shift suggests a stronger interaction HPA-Co in the catalyst than in bulk compound.
191 ^nf (a)
M^
700
300 400
(d)
1-^ X(nm) 800 900
Figure 9. UV spectra of supported catalysts: (a) C03/2PM0 /AI2O3 dried in air, (b) H7PM0/AI2O3 dried in N2, (c) Co7/2PMo°/Al203 dried and calcined in air, (d) Co7/2PMo°/Al203 dried and calcined in N2 -12.14
100
5 (ppin)
Figure 10. ^^P MAS-NMR Spectra (7kHz)of: (a) H7PM0, (b) Co7/2PMo°, (c) Co7/2PMo°/Al203 dried in air, (d) Co7/2PMo°/Al203 dried and calcined in air
192 3.4. Activity It should be mentionned that the HDS conversions are identical if the impregnation is performed directly on the extrudates or on grinded alumina (size of the particle .5 mm). This means that the impregnating solution penetrates inside the pore of the carrier and that we have no mass transfer limitation during the impregnation of the extrudates. The conversions in HDS of thiophene are reported in table 1. It shows that the catalysts prepared with unreduced heteropolyanions are effective for hydrodesulphurization (N"" 3 to 9) and that the values obtained with phosphomolybdate compounds are similar to those obtained with silicomolybdate ones. However the promoting factor is lower than the one currently observed for solids prepared by a classical impregnation with ammonium heptamolybdate (see N° 3; 4 and 7, 8, 9). This is due to the low Co content of these solids. Table 1 Activity in HDS of thiophene (% conversion of thiophene) N° Compound Thermic Treatment 1 CoMoPl drying and calcination: N2 2 C0M0P2 drying and calcination: N2 3 drying : air H3PM0 4 drying: air C03/2PM0 5 drying and calcination: air 6 drying: air Ni3/2PMo 7 drying: air H^SiMo 8 Co2SiMo drying and calcination: air 9 Co3/2SiMo drying and calcination: air 10 C05/2PM0 drying: N2 11 drying : N2 C07/2PM0° 12 drying and calcination: N2 13 drying and calcination: air 14 drying and calcination: N2 C07/2PM0* 15 Co7/2SiMo drying and calcination: N2 16 Co4SiMo drying and calcination: N2 17 H3C0PM0 drying: N2 18 drying and calcination: N2 19 H3C02PM0 drying: N2 20 drying and calcination: N2 21 drying and calcination: N2 H4C62SiMo
Conversion (%) 27 22 4 11 15 21 4.9 19 14 20 19 32 30 29 29.5 30 12 12.5 18 18.5 21.5
Higher conversions are observed for solids with higher Co loadings (see N° 10 to 16) which can be obtained through a reduction process of the heteropolyanion. Moreover it should be pointed out that calcination in air or N2is favourable to the activity (N° 11, 12, 13). The results presented lines 10 to 21, show that the substitution of the H^ ion is needed to increase the Co loading and consequently the thiophene HDS conversion. However the calcination allows an increase of conversion only for catalysts in which the K" ions are
193 exchanged (lines 17-21 compared to lines 10-16). Further experiments are now in progress to explain these differences. Up to now, from the results obtained in this work, it appears that the most active catalyst is the C07/2PM0, the conversion of which is higher than the one observed for the reference CoMoPl catalyst. It is even more active than the CoMoP2 one which has the same stoichiometry as the C07/2PM0 catalyst but its Co/Co+Mo ratio is lower than .28, the optimum value currently admitted. This could be correlated to the aforementionned strong interaction between the promotor and the heteropolyanion identified in the impregnating solution as well as on the oxidic precursor. This interaction could decrease the fraction of Co atoms involved in the formation of the well known surface "C0AI2O4" species [19], which is not available for the decoration of the M0S2 crystallites.
4. CONCLUSION The main findings of this work can be summarized as follows: i) The reduction allows us to increase the Co/Mo atomic ratio of the bulk phosphomolybdate and the silicomolybdate saUs of the Keggin structure, ii) The nature of these heteropolyanions is not modified after solubilization in water and a strong interaction between Co^^ ion and this heteropolyanion in solution has been evidenced, iii) Phosphomolybdate anion is preserved after impregnation and drying or calcination, iv) The increase of the Co/Mo ratio improves the thiophene HDS conversion, the value of which is higher than the one observed with a catalyst prepared by conventional dry impregnation with AHM, although with a lower Co/Mo atomic ratio. In conclusion, this work has shown that Keggin heteropolyanions containing promotor and base metals are convenient precursors for the preparation of HDS catalysts.
REFERENCES 1. J.A.R. van Veen, P.A.J.M. Hendriks, R.R. Andrea, E.J.G.M. Romers, and A.E. Wilson, J. Phys. Chem., 94 (1990), 5282. 2. W.C. Cheng, N.P. Luthra, J. Catal., 109 (1988), 163 3. J.A.R. van Veen, E. Gerkema, A.M. van der Kraan, A. Knoester, J. Chem. Soc. Chem. Commun., 1684(1987). 4. R.J.J. Jansen, H.M. Van Veldhuizen, M.A. Schwegler and H. Van Bekkum, Rec. Trav. Chim., Pays-Bas, 113, 115 (1994). 5. A.M. Maitra, N.W. Cant and D.L. Trimm, Appl. Catal, 48 (1989), 187. 6. Y. Okamoto, T. Gomi, Y. mori, T. Imanaka, S. Teraniski, React. Kinet. Catal. Lett., 22, 3-4 (1983), 417. 7. C. Sanchez, J. Livage, J.P. Launay, M. Fournier, Y. Jeannin, J. am. Chem. Soc, 104 (1982), 3194.
194 8. C. RocchiccioIi-DetchefF, M. Foumier, R. Franck, R. Thouvenot, Inorg. Chem., 22 (1983), 207. 9. C. RocchiccioIi-DetchefF, M. Amirouche, M. Fournier, J. catal., (48) 138 (1992), 445. 10. M. Foumier, C. Rocchiccioli-Detcheff, L.P. Kazansky, Chem. Phys. Lett., 123 (1994), 294. 11. C. RocchiccioIi-DetchefF and R. Thouvenot, J. Chem. Res., Synop., 46 (1977), miniprint 549 (1976) 12. L.P. Kazansky, Contribution From the institute oF Physical Chemistry, Moscow 117071, USSR, (1979), 70. 13. R.I. Maksimovskaya, V.M. Bondareva, Russian Journal oF inorganic Chemistry, 39, 8 (1994), 1238. 14. D.Z. Herranz, Contribution a I'etude des heteropolyanions molybdo-tungsto phosphoriques etude par RMN de ^*P, These, Paris, 1981. 15. R. Massart, R. Contant, J.M. Fruchart, J.P. Ciabrini and M. Fournier, Inorg. Chem., 16 (1977), 2916. 16. L.G. Roberts, F.H. Field, J. Am. Soc, 72 (1950), 4232. 17. R. Massart, Ann. Chim., t.4. (1969), 365. 18. M. Lo Jacono, A. Cimino, G.C.A. Schuit, Gaz. Chim. Ita., 103 (1973), 1281. 19. H. Topsoe, B.S. Clausen, Appl. Catal. 25 (1986), 273.
© 79P7 Elsevier Science B. V, All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
195
Genesis, Characterizations and HDS Activity of Mo-P-Alumina Based Hydrotreating Catalysts Prepared by a Sol-Gel Method R. Iwamoto ^^ and J. Grimblot« «Laboratoire de Catalyse Heterogene et Homogene, URA CNRS D402, Universite des Sciences et Technologies de Lille, 59655 Villeneuve D*Ascq Cedex, France ^ Central Research Laboratories, Idemitsu Kosan Co., Ltd., 1280 Kami-izumi Sodegaura, Chiba, Japan ABSTRACT Mo oxide - P oxide - Aluminum catalysts with a wide range of P loading (014 wt%) were prepared by a sol-gel method to elucidate the role of phosphorous on the textural, structural and catalytic properties of Mo based catalysts. Two different Mo loadings (expected ~20 and ~30wt%Mo) and two kinds of P precursors (phosphoric acid, phosphorus pentoxide) were examined. The structural properties of dried and calcined forms were studied by means of various characterization techniques. Specific surface area (S.S.A.) of catalysts were decreased proportional to the P loading in every series. Especially, the S.S.A. in the series of P2O5 precursor decreased drastically above 7.7wt%P loading. XRD measurements revealed that excess loading of Mo and P within the alumina framework provokes aggregation of bulk M0O3 (above 6.8wt%P in the series of 30wt%Mo for H3PO4 precursor and above 5.5wt%P in the series of P2O5 precursor). From IR measurements, it was found that P and Mo atoms interact with equivalent sites of alumina. From NMR measurements, predominant formation of Mo-P heteropoly complex were observed in the drying step. P interacted strongly not only with alumina framework but also with P itself. P2O5 prefers to polymerize by calcination. It was also found that Mo enhanced the interaction of P with alumina through the formation of P-Mo heteropoly complex. Water extraction tests revealed that Mo and P interacts strongly with the alumina framework. The HDS activity was not promoted by P while excess P decreased HDS activity with the formation of bulk M0O3. 1. INTRODUCTION The active phase of hydrotreating catalysts generally consists of Mo sulfide deposited on y-alumina which was produced by calcination of alumina hydroxides precursors. The Mo precursor is usually introduced to alumina by conventional dry or wet impregnation methods. However, only up to 10-12 wt% Mo can be dispersed by these methods. In previous works, new preparation methods of welldispersed Mo precursor based on a sol-gel method were proposed [1][2]. In this sol-
196 gel method, alumina is obtained by hydrolysis of aluminium tert-butylate or aluminium sec-butylate. Mo is incorporated homogeneously with the alumina precursor during the support preparation. This advanced method can give at least 30 wt% of weU dispersed Mo and higher HDS activity than conventional catalysts. Indeed, the physico-chemical properties of resulting soUds depend on the reactions involved in the sol-gel process (hydrolysis, condensation through alcooxolation, oxolation or olation steps), on the nature of the metal or the associated alkoxide and finally on the reaction condition (temperature, the ratio between the solvent and alkoxide). Furthermore, the sol-gel preparation method is very convenient not only for obtaining active catalysts but also for investigating what happens on the surface of catalysts because of their unique high S.S.A.. To achieve higher activity with the sol-gel catalyst, it is useful to investigate the effect of promoter and additives such as Co, Ni and P on the Mobased sol-gel catalysts. The role of P on the HDS activity for Mo based hydrotreating catalysts has been studied by many researchers, while the precise effect has not been well understood yet [3-5]. Eijsbouts et al. reported that P had no effect on the HDS activity for M0O3/AI2O3 [3]. On the other hand, Lewis et al. and Kim et al. reported a positive effect for HDS reaction in the region of low P loading [4] [5]. In this work, we wished to elucidate the role of P on the Mo-P-alumina sol-gel catalysts which contain high loading of Mo and a wide range of P amounts. Their main structural and textural properties will be compared as well as their performance in thiophene HDS. 2. EXPERIMENTAL 2.1 Catalyst preparation Mo-P-Alumina catalysts were prepared on the basis of a sol-gel method according to the procedures in Figure 1. Alumina was prepared by the hydrolysis of aluminium sec-butylate (ASB) dissolved in 2-butanol (2BN) and 1,3-butanediol (13BD). Mo and P were incorporated with the alumina precursor during the gel preparation. Mo was added to the aluminium alkoxide before hydrolysis as a dispersion of ammonium heptamolybdate (AHM) in 13BD. P was introduced by different ways depending on the nature of precursor. P2O5 precursor was introduced in ASB solution after dissolving in 2BN (Route A). On the other hand, 99% of ortho-phosphoric acid (H3PO4 precursor) was dissolved in 13BD simultaneously with AHM (Route B). The catalysts obtained at each stage are noted to MPDOC-Y)H, MPC(X-Y)P where the MPD, MPC means dried and calcined sample, X,Y means expected loadings in wt% of Mo and P respectively. H, P refer to the nature of P precursor such as H3PO4 or P2O5 respectively. The ratio of H2O/ASB was usually kept at 10. It was noted however as * if H2O/ASB was increased to 100. 2.2 C a t a l y s t s c h a r a c t e r i z a t i o n The chemical compositions were provided by "Service Central d'analyse du CNRS" (Vernaison, France). The obtained powders were characterized by BET
197 Preparation of Catalyst
Preparation of support ASB in 2-butanol at 85*C with stirring for 10 min. molar ratio butanol/ASB: 3
/
I
P205
\
Addition of 1,3-butanedioi
Addition of (NH4)6Mo7024 and 1,3-butanediol molar ratio butanediol/ASB : 2
molar ratio butanediol/ASB: 2
^4 Addition of H2O molar ratio H2O/ASB: 10
Stirring at 8S'C for 1h
Holding at room temperature for 1h without stirring
Drying under reduced pressure at 40«C for 1h and 60*C for 1h T I
MPD(X-Y)H or P*
Drying at 100*C overnight
|
Calcining at 500*C for 3h heating rate 40*C/min.
/
MPC(O-O) Support AI203
\
MPC(X-Y)H or P*
Catalyst Mo03-P205-Al203
|
Figure 1. Procedure for preparation of Mo-P-Al sol-gel catalysts specific surface area (QUANTASORB Jr., Quantachrome; pretreated at 200°C for 30 min.), X-ray powder diffraction(XRD) (Siemens D5000 Diffractometer equipped with a goniometer, a monochromator and a Cu X-ray tube). Infrared Spectroscopy (FTIR, NicoUetSlO Spectrometer, sample was pelleted with KBr), 27A1-NMR (BRUKER ASX400; resonance frequency 104.26MHz, recycling time 3 sec, pulse length 1 ^isec, spinning frequency 15kHz and reference A1(H20)6^"^) and ^iP-NMR (BRUKER ASXIOO; resonance frequency 40.53MHz, recycling time 40 sec, pulse length 2 jisec, spinning frequency 7kHz and reference H3PO4). 2.3 Catalytic activity (HDS) Hydrodesulfurization of thiophene was carried out at atmospheric pressure in a flow type reactor packed with 0.2g of catalyst. The catalyst was sulfided at 400*^0 for 2h with a H2/H2S (90/10) mixture gas at flow rate of 50 ml/min. After cooling down to 300°C, thiophene purified by vacuum distillation was introduced in the reactor at constant pressure (50 torr) with a flow of dried hydrogen (lOml/min.). The reaction products were analyzed by gas chromatography.
198 3. RESULTS AND DISCUSSION 3.1 Specific surface area and chemical composition Table 1 shows chemical composition and specific surface area per gram of calcined catalyst (S.S.A.) of all prepared samples. Amounts of Mo and P were almost those expected except for MPC(30-14)P*. It is suggested that excess water enhanced dissolution of AHM and prevents incorporation with alumina. As already reported in a previous report [1][2], the bare alumina MPC(O-O) and Mo oxide alumina such as MPC(20-0), MPC(30-0) have higher S.S.A. compared with conventional ones [6]. With introducing P, the S.S.A decreased proportional to the amount of P loading in every series. Especially, the S.S.A. decreased drastically above 7.7wt%P loading in the series of P2O5 precursor. It is presumed that P and cracked alcohol residues may block the porosity of sample because large amount of carbon (5.9wt%) was found on the MPC(30-13)P. However the corresponding MPC(30-14)P* which was prepared with increasing H2O/ABS ratio showed less carbon residue but stiU low S.S.A. (13m2/g). Table 1. Chemical composition and S.S.A. of prepared catalysts Catalysts Mo (wt%) P (wt%) Carbon(wt%) MPC(O-O) 0 0.5 0 MPC(20-0)H 0.2 17.5 0 MPC(20-1)H 0.3 17.9 1.6 MPC(20-2)H 17.3 2.2 0.2 MPC(20-3)H 17.2 0.2 3.1 MPC(20-4)H 17.9 4.4 0.2 MPC(20-7)H 6.6 0.3 16.8 MPC(20-11)H 16.4 11.3 0.3 MPC(30-0)H 26.0 0 MPC(30-1)H 0.1 25.9 1.1 MPC(30-2)H 2.2 0.1 26.5 MPC(30-5)H 0.1 4.6 25.3 MPC(30-7)H 25.8 6.8 MPC(30-11)H 0.1 25.3 11.1 MPC(30-1)P 28.3 1.7 0.3 MPC(30-3)P 0.6 27.7 2.7 0.2 MPC(30-6)P 5.5 25.5 0.2 MPC(30-8)P 26.7 7.7 5.9 MPC(30-13)P 12.7 25.3 MPC(30-14)P* 16.5 13.7 0.3 MPC(0-11)H 0 0.3 10.6 MPC(0-1Q)P 0 -_ (10)
S.S.A.(m2/g) 503 586 570 560 559 525 505 428 523 505 443 411 405 263 508 490 439 238 4 13 474 486
3.2 X-ray powder diffraction (XRD) (a)H3P04 precursor Figure 2 shows the XRD patterns of sol-gel catalysts obtained from the H3PO4 precursor. The bare alumina MPC(O-O) can be identified as poorly
199 crystalline y-Al203. For P oxide-alumina MPC(0-11)H and Mo oxide-alumina MPC(30-0), no peak corresponding to P or Mo oxo-compounds can be detected. It is suggested therefore that Mo and P exist as a weU dispersed species. It was also observed that incorporation of Mo and/or P with alumina prevents the formation of structured y-Al203. They seem to be present in an amorphous matrix. For Mo-PAl, the weU dispersed state of Mo oxide is kept up to MPC(20-11)H in the series of 20wt%Mo (not shown here) and up to at least MPC(30-5)H in the series of 30wt%Mo. However, bulk M0O3 can be identified above MPC(30-7)H. This result means that high loading of Mo and P within the alumina framework provokes aggregation of bulk Mo oxide.
40 2theta/«
Figure 2. XRD patterns of Mo-P-Al sol-gel catalysts prepared from the H3PO4 precursor. (a)MPC(O-O), (b)MPC(0-ll)H, (c)MPC(30-0), (d)MPC(30-5)H, (e)MPC(30-7)H, (f)MPC(30-ll)H
Figure 3. XRD patterns of Mo-P-Al sol-gel catalysts prepared from the P2O5 precursor (a)MPC(30-0), (b)MPC(30.3)P, (c)MPC(30-6)P, (d)MPC(30-13)P, (e)MPC(30-14)P*
200 (b)P205 precursor Figure 3 shows the X-ray powder dififraction patterns of sol-gel catalysts obtained from the P2O5 precursor. The addition of P2O5 precursor showed similar effect as the H3PO4 precursor. However, the intensity of bulk M0O3 in MPC(30-6)P which contains 5.5wt%P was almost the same as that observed in MPC(30-7)H which contains 6.8wt%P. It was concluded therefore that the P2O5 precursor enhanced the formation of bulk M0O3 compared with H3PO4 precursor at the same loading of Mo. 3.3 Infrared spectroscopy (a)H3P04 precursor The assignment of IR bands in Mo-P-Alumina based catalysts have been already reported by many researchers [7-13]. Figure 4 shows IR spectra of dried and calcined catalysts obtained from the H3PO4 precursor. For all dried catalysts, a broad band at ~750 cm^ which is assigned to Al-0 stretching was observed. Furthermore, many small bands and shoulders were observed (i.e. at 1458, 1370, 1135 and 1055 cm-i etc.), though it is sometimes difficult to identify. These bands could be assigned to residual alcoholate incorporated in the alumina framework or supported metal complexes, because they are well corresponding to IR spectra of the solvents (2BN,13BD). This fact indicates that the hydrolysis reaction of ASB does not proceed completely in this preparation condition. Bands at 1070 cm-i for MPD(O-O) and MPD(0-11)H is considered as sol-gel boehmite [7]. P containing catalysts such as MPD(30-11)H, MPD(0-11)H have a broad band at ^1100 cm-^ This band can be decomposed into three bands at 1115, 1080 and 1055 cm-i which is assigned to stretching vibration of P=Ot, P-0 and P-O-Mo of heteropoly acid [8] [9]. This result assumes that P-Mo heteropoly compound was formed during the gel precipitation. Specific bands at ~1404, 900 and 845 cm-i which can be assigned to AHM were observed in MPD(30-11)H. It is considered that P prevents the incorporation of Mo within the alumina framework even at the drying step since the intensity of these bands are weU correlated to the P content. For the calcined samples, all the spectra are rather broad. With increasing loading of P, a large broad band appeared again at about 1100 cm^ which can be decomposed into two bands at 1125 and 1090 cm-i. They are assigned to the P=Ot and P-0 respectively [9][10]. In MPC(30-11)H, bands at 1000, 880 and 823 cm-i which are assigned to bulk M0O3 were also detected. This result is in well agreement with the results of XRD. It is suggested that the main part of M0O3 derived from decomposition of bulk AHM with calcination. In MPC(O-O), characteristic three bands at 1640, 1503 and 1425 cm-i which might be assigned to physically or coordinately adsorbed H2O were observed. However, the two of three bands were disappeared with the introduction of Mo and/or P [11]. This result indicates that Mo and P are interacting with equivalent sites of AI2O3. This is a reason why a part of Mo cannot interact with AI2O3 when P content increases.
201
MPD(O-O)
MPC(O-O)
1500
1000
WAVENUMBER(cm-l)
1500
1000
500
WAVENUMBER(cm-l)
Figure 4. IR spectra of Mo-P-Al sol-gel catalysts prepared from the H3PO4 precursor. (A: after drjdng 100°C , B: after calcination at 500°C). (b)P206 precursor Figure 5 shows IR spectra of dried and calcined catalysts obtained from the P2O5 precursor. In MPD(0-10)P, a characteristic band which might be assigned to monomeric P species observed at 1000 cm-i. In MPC(30-13)P, intensity of bands for the residual alcoholate was much more higher than those in MPD(30-11)H. It is suggested that P2O5 prevents the hydrolysis of ASB and eventually, it remains more alcoholate in the final compound. It is also assumed that part of these bands are attributed to organic P complex formed with the alcohol solvent, since the same IR spectra was obtained from dried P2O5 after dissolving it in 2BN. From the literature, the formula of these complexes are P0(0H)2(0but), P0(0H)(0but)2 or P0(0but)3 [12]. These complexes are considered to be formed by the following reactions. P2O5 + 2but-0H+ H2O -> 2PO(OH)2(Obut)
(1)
P2O5 + 4but-0H ^ 2PO(OH)(Obut)2 + H2O
(2)
P2O5 + 6but-0H ^ 2PO(Obut)3 + 3H2O
(3)
202
For calcined samples, MPC(0-10)P showed broad bands between 1000 and 1330 cm-i which are assigned to highly polymerized P oxo-compounds [13][14]. For MPC(30-13)P, the intensity of bands at 1090 and 1125 cm-i which are assigned to P-0 and P=Ot vibration decreased comparing with those in MPC(30-11)H. On the contrary, the intensity of band at 1200 cm-i which is assigned to polymeric P oxocompounds increased. It is suggested that MPC(30-13)P contains also more polymerized P oxo-species than MPC(30-11)H. If the ratio of H2O/ASB increases from 10 to 100, the bands at 1330 cm-i increased significantly. This means that the excess of water during the gel preparation provokes the aggregation of P. It is assumed that P has less interaction with alumina in the drying stage because a large part of P is involved in complexes with the alcohol solvent. In such a case, P prefers to polymerize than to interact with the alumina framework. The bands for bulk M0O3 were also observed at 1000, 880, and 823 cm-i in MPC(30-13)P and MPC(30-14)P*. From the IR measurements, it was found that the P precursor affects significantly on the physicochemical properties of resulting catalysts.
MPC(30-0)
MPD(30-0)
lij
o z < CO
2000
1500 1000 WAVENUMBER(cm-l)
500
2000
1500
1000
500
WAVENUMBER (cm-1)
Figure 5. IR spectra of Mo-P-Al sol-gel catalysts prepared from the P2O5 precursor. A: after drying 100*^0 , B: after calcination at 500°C.
203 3.4 27A1.NMR (a)H3P04 precursor Top peak value of ^^Al-NMR spectra are listed in Table 2. The assignment of 2'^Al-NMR spectra in this region have already been reported by many researchers [15-22]. MPD(O-O) has a single broad signal at 7.2 ppm. This signal is assigned to octahedral alumina [15]. For all the Mo and P containing catalysts, tailing of spectra between 0 and -30 ppm or even presence of a shoulder at -5 ppm were observed depending on the content of Mo or P. This tailing should correspond to octahedral surface aluminium sites shell in which P or Mo are located in a second coordination [16]. These signals seem to be characteristic for sol-gel catalysts since surface informations are emphasized by the extremely large S.S.A.. In addition, P containing catalysts showed another weak signal at ~41 ppm which is assigned to AIPO4. This result indicates that P interacts strongly with alumina framework even in the drying step. It was also revealed that the degree of interaction between P and alumina increased in the presence of Mo because the intensity of AIPO4 in MPD(30-11)H was more stronger than that in MPD(0-11)H. Table 2. Results of ^^Al -NMR obtained from the Mo-P-Al sol-gel catalysts Catalysts MPD(0-0) MPD(0-11)H MPD(20-0) MPD(30-0) MPD(20-11)H MPD(30-11)H MPD(30-13)P MPD(30-14)P* MPD(0-10)P
Before calcination 7.2 40.2 7.1 6.5 6.2 41.2 6.0 41.2 5.9
61.0
. 1.1 13.7 13.6 -0.4 5.2 33.0
(ppm)
-5.0 -5.0 -5.0
After calcination (ppm) 65.5 33.0 6.9 -12.0 65.5 39.0 6.6 5.2 62.2 30.0 -13.6 54.7 27.3 5.3 -13.6 55.0 37.8 4.8 -13.6 36.7 6.6 36.7 37.9 55.0
27.0 13 -2.9 -14.1 26.8 -2.9 -12.4 26.8 6.3
On the calcined bare alumina MPC(O-O), another new signal of tetrahedral aluminium site was observed at ~65 ppm [15]. Furthermore, a broad shoulder appeared at 33 ppm which might be attributed to 5-fold coordinated aluminium sites [17]. This signal is characteristic for the sol-gel alumina since it possesses a highly disordered and poorly crystalline structure as shown by XRD. The signal at ^30 ppm is also observed in the Mo loaded catalysts such as MPC(20-0) or MPC(30-0). The intensity of this signal increased with the increasing amount of Mo. This could be assigned to 5-fold coordinated aluminium sites since introduction of Mo prevents the crystallization of alumina and leads to much more distortion as already shown by XRD. However, another explanation as being due to the presence of a surface tetrahedral Al(OMo)4 cannot be neglected.
204 Furthermore, MPC(30-0) gave a weak shoulder spectra at -13.6 ppm which is assigned to Al2(Mo04)3 [18]. This compound is supposed to be derived from a following equation. 3 MoOaCbulk) + 2 Al203(surface) ^ Al2(Mo04)3
(4)
The formation of Al2(Mo04)3 is more apparent in MPC(30-11)H, because the high P loading favors the formation of bulk M0O3. However, MPC(0-11)H which contains only P also showed the shoulder signal at ~-13 ppm. This signal can be assigned to A1(0P)6 in this case [19]. Hence, the signal of ~-13, 14 ppm might be considered as multiple states of octahedral surface alumina in which terminal OH are exchanged by Mo or P. The spectra for AIPO4 was observed at 37 to 39 ppm for all the P containing catalysts. Though the AIPO4 already existed in the drying step, the main part of AIPO4 forms during calcination. The intensity of AIPO4 for MPC(20-11)H and MPC(30-11)H were much more stronger than that for MPC(0-11)H catalyst. This result indicates again that Mo provokes the formation of AIPO4. Concerning the chemical shift, the top peak values of 6, 5 and 4-fold aluminium sites tend to decrease with the increasing amount of Mo. This might be caused by the increase in distortion of the alumina framework or by the decreases of the aluminium density in the shell of each aluminium sites when the Mo loading increases [20]. (b)P205 precursor MPD(0-10)P had three signals at 61 ppm(weak), 33 ppm(weak) and 5.2 ppm(strong) which correspond to tetrahedral, 5-fold and octahedral aluminium sites respectively. MPD(30-13)P showed a large broad signal at -1.1 ppm and a small sharp signal at 13.7 ppm. The former signal could be assigned to less condensed octahedral aluminium sites [15]. The later signal might be assigned to (Al(OH)„(H20)6.n)(Mo04) or A1(0P)5 [19][21]. This result suggests that the P2O5 precursor prevents drastically the hydrolysis and condensation of the Al-alkoxide. Zaharescu et al. also reported that the rate of hydrolysis and condensation of Sialkoxide is strongly influenced by PO(OR)x complexes in the P-TEOS system [22]. P complexes might affect on the accessibility of the metal alkoxide to water molecules or to other alkoxides for condensation. In calcined catalysts, MPC(0-10)P showed also three signals at 55 ppm, 26.8 ppm(strong) and 6.3 ppm(strong) which correspond to tetrahedral, 5-fold and octahedral aluminium sites respectively similar to MPD(0-10)P. However, their respective populations were strongly modified. It is remarkable that no signal for the AIPO4 was observed in MPC(0-10)P even after calcination. MPC(30-13)P gave a large broad signal at 36.7 ppm and a smaU sharp signal at -14.1 ppm which are assigned to AIPO4 and Al2(Mo04)3 or A1(0P)6 respectively.
205 3.5 31P-NMR (a)H3P04 precursor Table 3 shows the top peak value of ^^P-NMR spectra. The assignment of ^T-NMR spectra in this region has been also reported by several researchers [2327]. For the dried MPD(0-11)H, a broad signal which could be decomposed into 2 signals at -11 and -21 ppm was obtained. They are assigned to monomeric and pol5mieric P oxo-species respectively [23]. On the other hand, the Mo and P containing catalysts such as MPD(20-11)H, MPD(30-5)H and MPD(30-11)H showed another signal at about -15 ppm. This signal could be assigned to a P-Mo heteropoly compound in agreement with the IR observation. All the calcined catalysts showed broad overlapping signals at about -18 and ~-24 ppm which corresponds to polymeric P oxo-species and AIPO4 respectively. Mo containing catalysts such as MPC(20-11)H and MPC(30-11)H gave less polymeric P than MPC(0-11)H. It is suggested that Mo is effective for dispersing P on the alumina through the formation of a P-Mo heteropoly complex. It was found that the top peak value of AIPO4 signal for MPC(30-5)H and MPC(30-11)H shifts 2 ppm to the lower value. Table 3. Results of 3iP-NMR obtained from the Mo-P-Al sol-gel catalysts. Catalysts MPD(0-11)H MPD(20-11)H MPD(30-5)H MPD(30-11)H MPD(30-13)P MPD(30-14)P* MPD(0-10)P
Befor calcination (ppm) -11 -21 -11 -15.9 -21 -11 -14.3 -11 -15.9 -21 -4 to-11 -14.6 -21 .4 to-11 -14.6 -21 -4 to-11 ^21__
After calcination (ppm) -18 -24 -24 -18 -26 -18 -26 -18 -19 -18 -19
-25 -23
b)P206 precursor MPD(0-10)P showed several overlapping signals between -4 and -11 ppm which might be assigned to multiple states of monomeric P such as P0(0H)(0But)2 and P0(0But)(0H)2 including their isomer structures. Zaharescu et al. reported that these complexes are not hydrolyzed by water [22]. MPD(3013)P showed another signal at -14.6 ppm which is assigned to a Mo-P heteropoly compounds. This result means that a part of the organic P can form complexes with Mo as weU as the H3PO4 precursor. Considering from the equilibrium studies by Jian et al. and Cheng et al. [24][25], it is supposed, that the P-Mo heteropoly compound was formed, for example, by the following reactions. 2H^ + 14PO(OH)2(OBut) + 5M07O24 6- ^
7H2P2M05O23 ^- + 14But-0H+ H2O
(5)
206 or M07024 6- + 4H20 ->7Mo04 2- + 8H+ and 6H+ + 2PO(OH)2(OBut) + 5M0O4 2-
(6) ->
H2P2M05O23 ^' + 2But-0H + 3H2O
(7)
After calcination, it was found that MPC(0-10)P showed only polymeric P at '--IQ ppm. No signal for the AIPO4 was obtained at -24 ppm in agreement with 2'7A1-NMR. It is considered that the P2O5 precursor tends to polymerize rather than to interact with the alumina framework because the alcoholate P complexes prevent the interaction with alumina. On the other hand, MPC(30-13)P showed two overlapping signals at -19 and -25 ppm which are attributed to the polymeric P oxo-species and the AIPO4 respectively. These data indicate that the P-Mo heteropoly compounds induce the formation of AIPO4. In the "P2Mo5023"structure, the two P atoms are located at top and bottom of the cluster respectively [25]. Therefore, it is thought to be easier for P to be in contact with the alumina framework. From the above investigation, it was found that the hydrolysis and condensation reaction of the Al-alkoxide are extremely prevented by the P2O5 precursor. Therefore, another MPC(30-14)P* was prepared with increasing the ratio of H2O/ASB=100 to accelerate the hydrolysis and condensation reactions. Figure 5 showed, however, that MPC(30-14)P* gave much more polymeric P oxospecies than MPC(30-13)P. It is considered that excess water shifted the equilibrium equation (5) and (7) to the left hand and consequently prevented the hydrolysis of alcoholate P precursor. Since increasing the H2O/ASB ratio did not improve the hydrolysis reaction, P2O5 might prevent the access of alkoxide molecules to each other. The amounts of Mo and P remaining after water extraction were also investigated in Table 4. In general, M0O3, H3PO4 and heteropoly compounds are easily extracted by water while monolayer molybdate and AIPO4 are hardly extracted [26][27]. The extent of extraction depends strongly on the degree of Table 4. Effect of water extraction on the atomic ratio Mo/Al and P/Al Before water extraction Catalysts MPC(20-0) MPC(20-11)H MPC(30-0) MPC(30-11)H
Mo/AI 0.13 0.18 0.23 0.36
After water extraction
P/Al
Mo/Al
0.37
0.14 0.14 0.25 0.21
0.50
P/Al
0.40 0.55
207
interaction between those compounds and support. It was found that the amount of Mo and P for MPC(20-0), MPC(20-11)H, MPC(30-0) catalysts after the water extraction gave ahnost same value as that of the initial catalysts. Therefore, all the Mo and P oxo-species in these catalysts have strong interaction with the alumina surface. On the other hand, the amount of Mo for MPC(30-11)H apparently decreases with the water extraction. Therefore, it is considered that a part of Mo cannot interact with alumina and leads to the formation of bulk M0O3. As the main conclusions from the characterizations, it appears that reactivity of P with Mo and alumina depends strongly on the nature of the P precursor and on the preparation conditions. The scheme of interaction between P and other component is shown in Figure 6. P0(0H)x(0R)3.x +ROH polymeric ^ P oxo-species +Mo
P-Mo heteropoly species
>
AIPO4
Figure 6. Schematic diagram for P transformation 3.6. Thiophene HDS activity (a)H3P04 precursor Figure 7 shows the thiophene HDS activity and selectivity for the sol-gel Mo-P-Al catalysts prepared from the H3PO4 precursor as a function of P content. It was found that no effect was detected in the series of 20wt%Mo. On the other hand, a negative effect was obtained above 4wt%P in the series of 30wt%Mo. The decrease in the HDS activity should be attributed to the formation of bulk MoOs since bulk MoOs possesses less activity than dispersed Mo. Selectivity of C4 products did not changed significantly whUe the selectivity of hydrogenated compound (butane) decreased slightly with the formation of bulk M0O3
208
/-
I 6.56^- r 1 0
•
t
•
5.5 - [ •
• Mo20wt% • Mo30wt%
\#
54.5 - [
•
43.5 3-
•
•
•
•
1 '^
H
4
6
\
1
10
15
8
P content (wt%) Figure 7. Thiophene HDS activity and selectivity products on Mo-P-Al sol-gel catalysts prepared from H3PO4 precursor. In B, • , • mean the series of 20wt%Mo and n , 0 mean the series of 30wt%Mo. b)P205 precursor Figure 8 shows the thiophene HDS activity and selectivity for the Mo-P-Al sol-gel catalysts prepared from the P2O5 precursor. The almost same trend was obtained as the H3PO4 precursor, though the activity started to decrease above -'4wt%P. This limit is lower than that found in H3PO4 series because the P2O5 precursor favors the formation of bulk MoOs and carbon in the preparation procedure (see Table 1). It is considered that only dispersion of Mo affects on the thiophene HDS activity in Mo-P-Alumina sol-gel catalysts.
209 7 6
5
0<
2
2 !:^
4
d a
3
•^ > d
2
«M
OB V4
0
o
•
' * X.
^
0 Q
0
<
^
1
A
0
1
1
•
10
'
15
lUU -
90 - ^ -^80 ^ 7 0 '1 O60-
butene+butadiene
M
• —•
™^
^^
S.50|40 |30; ^2010 0
butane } • —
^
^
^
\ 4
\ 6
\ 10
P content (wt%) Figure 8. Thiophene HDS activity and selectivity of C4 products on Mo-P-Al sol-gel
catalysts prepared from P2O5 precursor.
4. CONCLUSION Mo-P-Al sol-gel catalysts with a wide range of P loading were prepared to elucidate the role of phosphorous on the textural, structural and catalytic properties of Mo based catalysts. It was found that the amount of P, the nature of P precursors and the preparation conditions affect significantly on the physicochemical properties and HDS activity. In drying step, predominant formation of a P-Mo heteropoly complex was observed. With calcination, P interacts strongly not only with the alumina framework but also with P itself. The interaction between P and alumina was enhanced in the presence of Mo via the formation of a P-Mo heteropoly complex. The P2O5 precursor prevents strongly the hydrolysis and condensation of the Al-alkoxide. The P2O5 has less interaction with alumina in the drying step and tend to polymerize by calcination. The HDS activity was not promoted by P and decrease with the formation of bulk M0O3.
210 REFERENCE I] E. Etienne, E. Ponthieu, E. Payen, and J. Grimblot, J. Non-Cryst. Solids 147 &148, 764 (1992) 2] L. Lebihan, C. Mauchausse, L.Duhamel, J.Grimbot and E.Payen J.Sol-Gel Sden. Tech., 2,837 (1994) 3] S. Eijsbouts, J. van Gestel, J. A. R. van Veen, V. H. J. de Beer and R.Prins, J.Catal., 131, 412 (1991) 4] J. M. Lewis and R. A. Kydd, J. Catal., 136; 478 (1992) 5] S. I. Kim and S. I. Woo, J.Catal., 133, 124 (1992) 6] S. M. A. M. Bouwens, J. P. R. Vissers, V. H. J. de Beer and R. Prins, J. Catal., 112,401(1988) 7] E. Ponthieu, E. Payen, G. M. Pajonk and J. Grimblot, in Proc. 8th Inter. Workshop on Glasses and Ceramics from gels (1995) (in press) 8] A. Spojakina and S. Damyanova, React. Kinet. Catal. Lett., 53, 2, 405 (1994) 9] P. Atanasova and T. Halachev, Appl. Catal., 48, 295 (1989) 10] J. M. Lewis and R. A. Kydd, J. Catal., 132, 465 (1991) II] P. M. Boorman, R. A. Kydd, T. S. Sorensen, K. Chong, J.M.Lewis and W. S. BeU, Fuel, 71, 87, (1992) 12] Wenjian Weng, in Proc. 8th Inter. Workshop on Glasses and Ceramics from gels (1995), (in press) 13] A. Spoja]dna and S. Damyanova and L. Petrov, Appl. Catal., 56, 163, (1989) 14] C. Morterra, G. Magnacca and P. P. de Maestri, J. Catal., 152,384 (1995) 15] Y. Kurokawa,Y. Kobayashi and S. Nakata, Heterogeneous Chem. Rev., 1, 309 (1994) 16] F. M. Bautista, J. M. Campelo, A. Garcia, D. Luna, J. M. Marina and A. A. Romero, Appl. Catal., 96, 175 (1993) 17] S. H. Risbud, R. J. Kirkpatrick, A. P. Taglialavore and B. Montez, J. Am. Ceram. Soc, 70, 1, ClO (1987) 18] I. H. Cho, S. B. Park and J. H. Kwak, J. Mol. Catal., A, 104, 285, (1996) 19] R. K. Brow, R. J. Kirkpatrick and G. L. Turner, J. Am. Ceram. Soc.,73 (8), 2293 (1990) 20] S. Rezgui, B. C. Gates, Chem. Mater., 6, 2386 (1994) 21] J. C. Edwards and N.P. Luthra, J. Catal., 109, 163 (1988) 22] M. Zaharescu, A. Vasilescu, V. Badescu and M. Radu, in Proc. 8th. Inter. Work. Glass Ceram. from Gels (1995), (in press) 23] E.C.Decanio, J. C. Edward, T. R. Scalzo, D. A. Storm and J. W. Bruno, J. Catal., 132,498, (1991) 24] M. Jian and R. Prins, BuU. Soc. Chim. Belg., 104, n4-5, 104 (1995) 25] C. Cheng andN. P. Luthra, J. Catal., 109, 163, (1988) 26] N. R. Gazimzyanov, V. I. Mikhailov and V. V. Volod'ko, Kinet. Catal., 36,5,694, (1995) [27] P. Atanasova, J. Uchytil, M. Kraus and T. Halachev, Appl. Catal., 65, 53 (1990)
® 1997Elsevier Science B.V, All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
211
Effects of ethylenediamine on the preparation of HDS catalysts: Comparison between Ni-Mo and Co-Mo based solids. P. Blanchard*, E. Payen*, J. Grimblot*, O. Poulet** and R. Loutaty^ *Laboratoire de Catalyse Heterogene et Homogene, URA CNRS 402, Universite des Sciences et Technologies de Lille, Bat. C3, 59655 Villeneuve d'Ascq, Cedex, France ^Centre de Recherches Total-France, Gonfreville I'Orcher, 76700 Harfleur, France
In the present paper, we report a new method of HDS catalysts preparation which allows the deposition of high metal loadings on alumina by using ethylenediamine (En) in the impregnating solution. We have shown that the use of En improves the dispersion of Mo as well dispersed oxomolybdenum entities in the oxidic precursor. The promoter (Co or Ni) can be introduced in the Mo impregnating solution. A complete study of the mechanism of adsorption on alumina in the presence of this complexing agent has been undertaken and we report the comparison between Co-Mo-alumina and Ni-Mo-alumina catalysts from physical characterizations as well as from thiophene HDS data.
l.INTRODUCTION Innovation in the development of Co-Mo-AI2O3 and Ni-Mo-AI2O3 catalysts has been driven by the need to produce clean fuels, based on the pressing requirement for environmental protection. Extensive studies on these hydrotreating catalysts revealed that a high dispersion of M0S2 nanocrystallites on alumina and a precise location of Co or Ni on these crystallites to play their promoting role is required for a good HDS activity. These catalysts are obtained by sulphidation of oxidic Co(Ni)-Mo-Al203 precursors. Most of them have been prepared by an incipient wetness impregnation like for the indutrial catalysts. From an industrial point of view, an increase of the molybdenum loading for a given alumina carrier appears as a challenge to enhance the activity per volume of catalyst. Inorganic salts such as ammonium heptamolybdate (AHM), cobalt or nickel nitrates are commonly used as metal precursors in aqueous solution. Ammonia or nitric acid is generally used as a pH monitoring agent to control the nature of the oxomolybdenum species in the impregnating solution. Organic acids [1,2] have sometimes been introduced in the impregnating solution but they have a very limited use in industrial catalyst preparations. In a previous paper [3] we reported the use of ethylenediamine (En) in the impregnating solution and we have shown that it improves the dispersion of the Mo 0x0species in the oxidic precursor with the presence of isolated molybdate entities in the impregnating solution. The promotor (Co or Ni) is generally introduced in the Mo impregnating solution and the use of En gives Co-Mo oxidic precursors where Co and Mo are well dispersed without formation of cobalt molybdate even at high metal loading. In the present report, a more complete study of the mechanism of adsorption in the presence of this
212 complexing agent has been undertaken and we report the comparison between Co-Mo-alumina and Ni-Mo-alumina based from physical characterizations and from thiophene HDS data.
2.EXPERIMENTAL 2.1.Preparation of the impregnating solution and catalysts
The catalysts were prepared by the incipient wetness impregnation method of a commercial Y-AI2O3 (pore volume: 1 cm^g~l; specific surface area: 350 m^g'^) with solutions containing the appropriate amounts of the active metals. The impregnated extrudates were dried at 393K overnight and then calcined at 773K for 3h. The nomenclature of the studied samples indicates the main preparation parameters. A sample will be designated as a CoMoX(En) or aNiMoX(En), where a is the atomic ratio nco(Niy(^Co(Ni)"^"Mo)» ^ ^^® ^ ^ loading in wt% of M0O3 and En the complexing agent (i.e. ethylenediamine). Co (Ni) or En are omitted respectively for samples prepared without cobalt (nickel) or without the complexing agent. Three main series of samples were prepared as follows: - Mo and Co(Ni)Mo series: The impregnating solution was an aqueous solution of AHM at its natural pH (5.5). A one step impregnation at 323K was used for the preparation of the 30Mo sample in order to avoid precipitation. For the Co-Mo and Ni-Mo samples, cobalt or nickel nitrate was dissolved in the AHM solution. - aCo(Ni)MoX(En): The impregnating solution contained AHM (with or without cobalt (nickel) nitrate) dissolved in a 3 M aqueous En solution. For purposes of comparison, some catalysts were prepared by successive impregnation steps: - MoX(En)/aCo(En) and aCo(En)/MoX(En) (or with Ni). The sequence of introduction of the Co (Ni) was varied i.e. impregnation by a Co (Ni) solution with or without En was performed prior to or after the Mo impregnation. An intermediate calcination at 773K during 4h was performed between each impregnation step. Some Co(Ni)-alumina solids, without Mo, were prepared with aqueous cobalt (nickel) nitrate solution containing or not the complexing agent. They will be referred hereafter as CoX(En) or NiX(En) where X is the Co (Ni) loading in wt% of CoO (NiO). The calcined extrudates were sulphided with a 10% vol. H2S in H2 at 673K for 2h. The Mo content was determined X-Ray fluorescence by the « Service Central d'analyses du CNRS »(Vemaison, France). 2.2.Characterizations Laser Raman Spectroscopy (LRS) was performed using a Raman microprobe (XY fi-om Dilor), equipped with a photodiode array. The exciting light source was an Ar"*" laser emitting the 488 nm line with a power at the sample of 10 mW except when otherwise specified. X-ray Photoelectron Spectra (XPS) were obtained by using an AEI ES200B spectrometer equipped with an aluminium X-ray source working at 300W. The binding energies BE have been measured by reference to Al 2p peak of the support at 74.8 eV., a value generally encountered for y AI2O3. Variations of the integrated intensity ratio of typical core levels (I Mo'M^i^_^ijlAl 2p, I Co(Ni) 2^^^_^Jl Al 2p) provides information on the surface composition and the surface atomic ratio Mo/Al, Co/Al or Ni/Al have been deduced from the XPS intensity ratio (integration of the classical differential quantitative XPS equation to infinite
213 depth). Since different oxidation states may exist on the sulphided catalysts, the experimental spectra are rather complicated and, in order to obtain reliable quantitative informations, a decomposition procedure to calculate the extent of molybdenum sulphidation has been used. Ultra-Violet Diffuse Reflectance Spectra (DRS) were recorded in the 250-1100 nm spectral range with a diffuse reflectance accessory. Samples were placed in quartz cuvettes and were scanned against a pure alumina background. High Resolution Electron Microscopy (HREM) analysis was performed on a Philips EM30 electron microscope. The powder, after sulphidation, was dispersed in alcohol for the preparation of the electron microscope grid. Based on the detection of more than 400 crystallites on several micrographs the distribution of lengths L of the elemental layers as well as their stacking N can be obtained. Catalytic activities for thiophene HDS were measured at atmospheric pressure in a flow-type reactor packed with 0.2g of catalyst. The solids were first sulphided with a H2S/H2 (10/90) mixture at a flow rate of 100 ml/min at 673 K for 2h and then cooled down to 573 K. After purification by vacuum distillation, thiophene was introduced in the reactor at constant pressure (50 torr) in a flow of purified hydrogen (20 ml/min). The reaction products were analyzed by gas chromatography.
3.RESULTS 3.I.C0 based oxide precursors Figure 1 shows the UV spectra of the Co solutions. Besides the strong band at 300 nm due to NO3" ions, the broad feature at 475-500 nm characterizes the presence of Co(H20)6^^
300
400
500
600
A.(nm)
700
300
400
500
600
700
800
X (nm)
Figure 1. UV spectra of Co and Ni aqueous solutions, (a) Co nitrate; (b) Co nitrate + En; (c) Ni nitrate; (d) Ni nitrate + En. ions in the cobalt nitrate solution (Fig.la). Upon addition of En, the spectrum is changed (Fig. lb) with a main band at 360 nm and a tail up to -470 nm. A band at 1070 nm (spectral range not presented on Fig.l) was not detected. According to the literature data reported in
214 table 1, presence of such a band is characteristic of the Co(En)3^^ complex ion whereas Co(En)3^'^ does not exhibit such a band. Therefore, contrary to what reported in [3], the impregnating Co solution with En certainly contains Co(En)3^^ ions. Obtention of Co(En)3^^ needs manipulation under inert atmosphere [4,6,7]. It should be noticed that introduction of molybdate does not modify the UV spectra (not shown) and therefore the Co(En)3^^ ions remain in solution in presence of Mo species. Table 1 Main UV bands reported in the literature for Co-En complexes in solution Complexes Bands References 356-480 Co(En)3'' (4) 461 - 1070 (5) 475 - 534 - 1042 (6) 338 - 465 - 730 Co(En)3' (5) 345 - 470 Figure 2 shows the DRS spectra of y-Al203 impregnated by Co and Co+En solutions with different cobalt contents. These samples have been characterized in their wet, dried and calcined states. h
700
800
Figure 2. UV spectra of wet (w), dried (d) or calcined (c) Co-alumina samples, (a) (b) (c): Col w, d, c; (d): Co2 c; (e) (f) (g) Co4 w, d, c; (h) (i) Q): ColO w, d, c; (k) (1) (m): Co4(En) w, d, c.
215 When Co(H20)6^^ is impregnated on alumina and whatever the Co loading, the spectrum of the wet sample presents a broad band at about 500 nm and a weak absorption shoulder at 650 nm (Fig.2a, e, h). This implies that the Co(H20)6^'^ entity is not greatly modified by impregnation. Upon drying or calcination the spectrum of the Co2 sample shows a triplet at 530, 570 and 620 nm (Fig.2d) which evidences the localisation of cobalt in tetrahedral sites of alumina giving a "pseudo-CoAl204" entity as previously discussed by Topsee and al.[8]. At this stage, the initial impregnating entity is completely modified by interaction with the support. The UV spectrum of the Co4 sample (Fig.2g) also shows the features of this surface entity with a strong broadening from 500 to 700 nm. After calcination, the spectrum of the ColO sample (Fig.2j) characterizes the presence of C03O4 (band at 450 nm and a broad shoulder at 700 nm) whereas C03O4 and C0AI2O4 oxides are both evidenced on the Co4 solid. It is however difficult to quantify the proportions of Co involved in each species. Some controlled impregnations of Co at pH 6 or 11 showed that the amount of Co irreversibly adsorbed on the alumina is about .6 to 1.6 wt% (Table 2), a loading which corresponds to the maximum amount of Co which can be deposited on alumina without formation of bulk C03O4 as evidenced by UV spectroscopy. Table 2 Amount of Co irreversibly adsorbed on alumina after impregnation with excess of solution Impregnating solution pH Washing solution Amount of Co (wt%) irreversibly adsorbed H2O + Co 6 H2O 0.65 H2O + Co + En H2O + En 11 1.60 H2O + Co + En 11 H2O lj60 If the impregnating solution contains En, the UV spectrum of the wet sample is different from the one observed for the solution (Fig.2k) The band at 1070 nm is not present. The band at 470 nm should be considered as characteristic of Co(En)3^^ in interaction with the support. After drying, no evolution occurs. This is in agreement with TGA measurements which show a weight loss at about 220°C corresponding to the loss of 3 En molecules per Co atom. In the drying conditions, the supported Co species are not decomposed. Upon calcination, the spectra of Co4(En) (Fig.2m) only characterizes C0AI2O4. It is clear that the use of En allows the preparation of C0-AI2O3 solids without any formation of bulk C03O4 in accordance with the UV characterizations reported in figure 2. The Co 2p3/2 XPS BE (Table 3) are in agreement with the presence of surface C0AI2O4. Formation of CoO (Co 2p3/2 BE = 781.7 eV) or of Co^^ species is not observed. However the ncJnM ratios deduced from XPS of the Co4 and Co4(En) solids are similar. This shows that when C03O4 is formed at the surface (Co4 sample), it is well dispersed probably as nanocristallites. 3.2.Ni based oxide precursors The UV spectra of the impregnating solutions reported in figure l(c,d) shows the features of the Ni(H20)6^^ and Ni(En)3'' complexes [5] respectively. The DRS of the wet and dried samples prepared without En (Fig.3a,b) are similar to the Ni(0H)2 one (Fig.3g). If En is used in the impregnating solution, the main evolution of the spectrum is evidenced after drying; the shape of the band at 550 nm is affected. After calcination and whatever the method of impregnation (with or without En, Fig.3c,f), the UV spectra are similar and present a band at 405 nm which characterizes Ni ions in octahedral sites [9]. This attribution correlates well with the XPS results reported in table 3 (Ni Ip^n BE = 856.6 eV). The shape of the peak as well as
216 the UV spectra do not indicate the formation of bulk NiO (main Ni 2p3/2 peak at BE = 854 eV [10]). From these results it appears that En has no effect on the nature and the dispersion of the Ni^^ species; however the presence of well dispersed NiO nanocrystallites could not be rejected [9]
Figure 3. UV spectra of wet (w), dried (d) or calcined (c) Ni(Mo)-alumina samples. (a) (b) (c): Ni4 w, d, c; (d) (e) (f): Ni4(En) w, d, c; (g) Ni(0H)2; (h): 0.28NiMol4 c; (i): 0.28NiMol4(En) c; (j):0.28Ni/Mo20(En) c; (k): 0.28Ni(En)/Mo20(En). Tables XPS characterization of Co(Ni)-alumina oxide precursors Catalysts nco(Ni)nm' nco(Ni/i^Ai(xlO0 XPS BE Co(Ni) 2p3/2 (eV) Co4 0.9 4.3 782.1 Co4(En) 0.9 4.0 782.2 Ni4 0.9 4.0 856.6 Ni4(En) 0.9 4.0 856.4 3.3.C0-M0 based oxide catalysts. Upon impregnation, drying and calcination, the evolution of the UV features of the CoMo(En) solids reported in figure 4 is quite similar to the one previously described for the Co(En) based solids. The spectrum of the Mo(En)/Co catalyst shows the features of C03O4.
217 We have shown that the tetrahedral MoO/" entity was present in the CoMo(En) solution (basic pH) and that the dispersion of Mo was not affected by the presence of the promotor [3]. The Raman spectrum of the CoMo20(En) calcined sample characterizes well described oxomolybdate entities (line at 950 cm'\ Figure 5a). In contrast, the Raman spectrum of 0.28CoMo30(En) shows a line at 940 cm"^ (Fig.5c) on the broad underlying band of the oxomolybdate entities. This former one should be ascribed by reference to literature data [11], to C0M0O4 not detected by XRD. In view of the very high diffusion cross section of CoMooI compared to supported oxoanions, this line should not be interpreted as the presence of large amount of this compound. This implies that if it exists, it could only be amorphous, microcrystalline or present in a non detectable amount. Whatever the loading, upon increasing the power of the laser beam (Fig. 5b, d), the modification of the Raman spectrum strongly suggests the existence of an interaction between Co and the oxomolybdate entities in the oxidic precursor. The use of En allows the one step impregnation method of Co and Mo at high Mo loadings (5.2 at.nm'^) without formation of bulk C0M0O4 whereas it is always identified if the preparation is performed without En in the impregnating solution as soon as the Mo loading exceeds around 3 Mo atoms.nm"^.
X (nm) Figure 4. UV spectra of wet (w), dried (d) or calcined (c ) Co-Mo-alumina solids, (a) (b) (c): 0.28CoMo20(En) w, d, c; (d): Mo20(En) c; (e): 0.28Co(En)/Mo20(En) c; (f): 0.28Co/Mo20(En) c.
Figure 5. Raman spectra of calcined Co-Moalumina catalysts, (a) (b): 0.28CoMo20(En) with Piaser=100 or 600W; (c) (d): 0.28 CoMo30(En) with Piaser=100 or 600W. (e): 0.28NiMo20(En) with Piaser=100 W.
218
3.4. Ni-Mo based oxide catalysts. The Raman spectrum of 0.28NiMo20(En) (Fig.6e) shows a broad line at 970 cm"* characteristic of a well dispersed oxomolybdate entities. The features of (a) or (b) NiMo04 phases [11] or of an heteropolyanionic Ni-Mo species are not evidenced [12]. Table 4 shows the XPS results of the oxidic precursors. The dispersion of molybdenum is not affected by the presence of the promotor and the binding energy of Ni Ipm characterizes Ni^^ atoms in octahedral sites which could be in the alumina lattice or present as a well dispersed surface entities as evidenced by the quantitative XPS analysis (see the comparison between nNi/nAiXPS and nNi/nAi theoretical). The UV spectra of NiMol4(En) and NiMol4 calcined are reported in fig.3h, i. They characterize the same evolution of the Ni entities compared respectively to the Ni4(En) and Ni4 solids. As for the Ni-based solids, no difference is observed for the final state of NiMoH and NiMol4(En) catalysts. Table 4 XPS characterization of Ni-Mo-alumina oxide precursors Catalysts nNi/nAi (xlO^) n^JriM (xlO^) BE Mo 3d5/2 theoretical XPS (eV) 0.28NiMol4 2.5 3 233.4 0.28NiMol4(En) 2.5 3.3 233.2 0.28NiMo20(En) 3.6 3J 232.8
BE Ni 2pi/2 (eV) 856.6 856.5 856.2
3.5. Sulphided catalysts.
The Mo 3d5/2 BE of all samples characterizes the presence of molybdenum disulphide. The dispersion of the Mo is not modified upon sulphidation the rate of which is about the same for all the samples studied (-- 80%).
Figure 6. HREM picture of the sulphided 0.28 CoMo30(En) catalyst. As this preparation method allows the deposition of higher Mo loading as well dispersed oxomolybdate entities, it would be interesting to compare the morphology of the disulphide crystallites with those obtained on conventionally prepared systems. Therefore some of these samples have been observed by HREM (Fig.6). For all the samples studied in this work, single layer crystallites are the dominant species. Upon increasing the Mo loading from 14 to 30 wt % as molybdenum oxide (with or without En), the stacking increases slightly (N=1.4 to 1.6) and the length L of the nanocrystallites is about 33 A. At high Mo loading some long and wrapped crystallites are observed. The effect of the carrier on the morphology of the
219 disulphide crystallites is observed in this work by comparison with the results obtained on a classical alumina carrier with a specific surface area of about 200 mV^ [13]. The more divided nature of the alumina induces a dispersion of the size of the disulphide crystallites as evidenced by the standard deviation (1.3nm). This has to be correlated to smaller size and the polydispersion of the alumina platelets. Upon introduction of the promotor (Co or Ni), no drastic change of the morphology is observed whereas the mean stacking (N = 2.4) is generally observed for promoted Mo based catalysts prepared by coimpregnation [13]. Table 5 XPS characterization of Ni-Mo and Co-Mo-alumina oxide and sulphided catalysts. Catalysts BE FWHM(eV) BE difference nNKCo/nAixlO^ nNi(Co/nAi xlO^ Ni(Co) Ni2p3/2 Ni(Co)2p3/2 (XPS) (XPS) 2p3/2 (eV) (Co 2pm) oxide sulphide Mo 3ds/2 Phase NiMoS [14] 625 1.8 854.2 0.28NiMol4 854.0 624.6 2.8 3.0 3.1 854.1 624.9 2.5 0.28NiMol4(En) 2.9 3.3 0.28NiMo20(En) 854.2 625.2 2.5 3.8 3.3 625.5 0.28Ni(En)/Mo20(En) 854.2 2.6 0.28Ni/Mo20(En) 854.3 625.2 2.6 549.9 0.28CoMol4 779.0 2.6 3.0 2.5 0.28CoMol4(En) 549.8 3.6 4.0 779.0 2.5 550.0 0.28CoMo20(En) 779.0 2.6 4.9 5.2 779.1 0.28CoMo30(En) 550.1 2.4 6.4 8.3 0.28Co/Mo20(En) 779.2 550.2 2.7 4.2 4.0 0.28Co(En)/Mo20(En) 779.4 550.5 2.4 5.0 5.0
In sulphided CoMo catalysts, three different types of Co have already been evidenced, i.e. cobalt in the alumina lattice, CogSg and Co in the so-called CoMoS phase. It has been shown that XPS could identify the nature of the Ni [14] or the Co species [15]. The main XPS characteristics of the Co and Ni in these sulphided catalysts are reported in table 5. Whatever the Mo loading or the sequence of impregnation the BE difference BECO-BEMO characterizes a promotor atom in a decoration position. The FWHM (full width at half maximum) of the Co2p3/2 peak increases when the catalyst has not been prepared with En, showing a higher amount of Co involved in CopSg formation during the sulphidation. Moreover, the atomic ratios nco/nAi of the Co-Mo solids deduced from the XPS intensity ratio decrease upon sulphidation, confirming that CopSg is formed (due to the presence of C0M0O4). In counterpart, this ratio increases when the solid is prepared with En. This has been correlated to the large amount of well dispersed Co associated with the M0S2 crystallites. Unfortunatly, for the Ni-Mo catalysts the value of the FWHM does not allow to conclude on the presence or not of a Ni sulphide entity. However, the evolution of nNi/nAi upon sulphidation is exactly the opposite of the one observed for Co-promoted catalysts. This ratio decreases when the catalysts are prepared with En and increases for the NiMol4. One can postulate that this decrease is due to the formation of Ni sulphide independant of M0S2; it is formed by the sulphidation of a well dispersed NiO phase not evidenced on the oxidic precursor with the techniques used in this study.
220
3.6. Activity measurements. Table 6 Thiophene HDS Activity of selected Ni-Mo and Co-Mo alumina catalysts. Catalysts Thiophene HDS activity (% conversion) 0.28NiMol4 44 39 0.28NiMol4(En) 0.28NiMo20(En) 49 0.28Ni/Mo20(En) 41 0,28Ni(En)/Mo20(En) 41 0.28CoMol4 25 0.28CoMol4(En) 33 0.28CoMo20(En) 38 0.28CoMo30(En) 42 0.28Co/Mo20(En) 31 0.8Co(En)/Mo20(En) 36 Q.28Co(En)/Mo30(En) 44 Table 6 shows the results of thiophene HDS activity. Whatever the mode of impregnation, successive or simultaneous, a higher conversion is always observed if Co has been introduced as Co-En complex by comparison with solids prepared with the same sequence and the same Mo loading. With the Ni based catalysts, whatever the method of preparation (one step or two steps with Ni after Mo) the use of En does not induce a higher conversion of thiophene. Even, a decrease of the activity is observed for the NiMol4(En) sample. The comparison is not possible for a 20 wt % M0O3 loading as the NiMo20 solid presents some heterogeneity already observed with Co-Mo catalyst. When the preparation is carried out by successive impregnation, the effect of En is not clear. Indeed, no evolution of activity is observed. Moreover the 0.28Ni/Mo20(En) and 0.28Ni(En)/Mo(En) solids have lower activities than the 0.28NiMo20(En). Upon increasing the Mo loading, the activity per gram of catalyst increases but the promoting factor (conversion ratio between promoted and unpromoted catalysts at the same Mo loading) remains constant for a simultaneous impregnation with En: about 6 for Co promoted catalysts and about 8 for Ni promoted catalysts [3].
4. DISCUSSION 4.1. Co, Ni based solids. It has been well described that, in a classical impregnation of Co or Ni on alumina (without a complexing agent), the formation of bulk C03O4 starts at a coverage of 0.7 Co at. nm'^ [16] whereas bulk NiO is not observed until 3 Ni at. nm'^ [17, 18]. Our results on solids prepared without a complexing agent confirm these results. Nevertheless, this study shows that segregation begins during the drying step for the C0/AI2O3 samples whereas no evolution occurs at this step for the Ni/Al203 ones. The formation of C03O4 proceeds through Co atoms involved in an entity characterized in the wet and dried states by UV bands at 500 and 650 nm which should correspond to an hydroxyde precipitate. At 4 wt %, the well dispersed Co atoms
221 on the calcined sample are already in strong interaction with the support after the drying as evidenced by UV results. It is indeed very interesting to note that the amount of Co adsorbed during impregnation with a large excess of solution corresponds to the amount of Co which can be deposited by dry impregnation without formation of C03O4. The formation of this well dispersed entity should proceed through interaction of the Co(H20)6^'^ entities with the Al-OH groups of the carrier. Upon calcination this entity is transformed into the well dispersed surface "C0AI2O4" phase whereas the hydroxyde precipitate is transformed into bulk C03O4. Considering that cobalt hydroxyde is not so well defined than Ni(0H)2, this latter is clearly identified by DRS on the wet and dried Ni/Al203 solids. However it should be well dispersed after calcination. This difference has been attributed to a better occupation by Ni of the tetrahedral and octahedral sites of AI2O3. Addition of En during the preparation of C0/AI2O3 and Ni/Al203 solids respectively inhibites the formation of bulk C03O4, but does not change the Ni repartition on the support. The Ni(En)3^^ and Co(En)3^^ ions interact electrostatically with the surface of the support during impregnation and drying as evidenced by DRS; this is in agreement with the data of Clause and al. [19,20]. The high stability of these complexes, as shown by TGA, allows to avoid the precipitation of cobalt oxohydroxyde. Isolated entities well dispersed and interacting with the carrier are obtained for temperatures below 200°C. These entities are characterized by UV bands at 350 and 470 nm for Co species and 350 and 550 nm for the Ni ones. Upon calcination, degradation of these surface complexes induces a location of the Ni^^ or Co^^ ions in lacunar sites of the carrier. Then, the differences observed for the calcined Co and Ni/Al203 solids could be ascribed to a better occupation by Ni of tetrahedral and octahedral sites of alumina as it has been often proposed [21] but also to a different chemical stability of the aquocomplexes presents in the impregnating solution in the pores of alumina. The lower stability of the Co entities in presence of the carrier induces a precipitation due to an increase of the pH of the solution in the pores. These precipitates should transform upon calcination into C03O4. The use of a more stable complex of Co implies a better dispersion of the element in interaction with the alumina carrier. However the dispersion of the Ni is good and its preparation does not require the use of a complexing agent. We have to take also into account the reduction of the Co^^ complexes which occurs during its decomposition. Hathaway and al. have already observed this phenomenon [6] for silica supported Co(En)3^^ and attributed it to the presence of O^" groups and residual H2O molecules. However in this work the reduction of Co^"^ to Co^^ which occurs upon calcination should rather be due to the decomposition products of En. So a better knowledge of the role of the reduction of Co^^ to Co^^ on the localisation of this metal is necessary. 4.2. CoMo (NiMoyAliOa catalysts. The location of Co in the sulphided catalysts depends on its location in the oxide precursor i.e. i) the existence or not of bulk oxides such as C03O4 or C0M0O4 which forms upon sulphidation bulk C09S8 and ii) the repartition of Co in octahedral and tetrahedral sites of the alumina lattice; they can migrate upon sulphidation to give the so-called CoMoS phase [22]. The use of En allows the preparation of Co-Mo catalysts at high Mo loadings without formation of unwanted oxides such as C03O4 or C0M0O4. This permits the obtention of an optimal sulphide phase without C09S8 and induces a better « decoration » of the M0S2
222 cristallites. Formation of C0M0O4 is due to surface precipitation of Mo and Co during the impregnation with the conventionnal method (natural pH) [23]. The use of En inhibites such a precipitation in stabilizing the solution in the pores of the alumina. Van Veen and al. [1] have used nitrilotriacetic acid (NTA) to achieve this goal. They attributed the formation of a pure CoMoS II type structure to the formation of a Co-Mo-NTA complex during the preparation of the oxide precursor. Unfortunatly, the existence of such a Co-Mo-En entity in the impregnating solution has not been characterized in this study. But, after drying, isolated Co^^ complexes and polymolybdates species in mutual interaction are probably adsorbed on the surface of the carrier and upon calcination will lead to the formation of a well dispersed Co-Mo oxide phase in which the cobalt atoms and the polymolybdate phase are in close contact as evidenced in LRS. Moreover, at a loading of 20 wt % M0O3, as the thiophene conversion of solids prepared by sequencial impregnation with En is lower than the one obtained for a solid prepared by simuhaneous impregnation, we may consider that such entities could be present in the dried catalyst. At 30 %, the one step impregnated catalyst is less active that the two steps one, which shows that this method of preparation can not be applied for very high loading due to the limited solubility of all the entities present in solution. It can be concluded that up to 30 wt % M0O3, the use of En allows a good stability of the impregnating solution in the alumina texture which induces a good dispersion of the Co-Mo based species in the oxidic precursor. In counterpart, the results show that En has no effect for Ni-Mo based systems. As the formation of nickel molybdate as well as of bulk nickel hydroxyde are never evidenced on solids prepared with the classical simultaneous impregnation, the use of En is not necessary. The use of En induces a strong interaction between the Ni complex and the support which inhibites the formation of an oxide Ni-Mo phase considered by some authors as being a heteropolyanionic system. Therefore, the Ni atoms in the decorating position are in lower quantity and this should explain the decrease of activity observed when En is used in the preparation and the decrease of the ni^JriM ratio after sulphidation. This is in agreement with results obtained with a solid prepared by a sequential impregnation. Indeed the activities are lower if Ni is introduced after the Mo with or without En. This confirm that Ni and Mo should be associated at each step of the preparation for obtaining a high activity.
5. CONCLUSION In a previous paper [3] we showed that the use of En improves the dispersion of molybdenum up to 5.2 at. Mo/nm^ as well dispersed polymolybdate entities whereas isolated molybdates were present in the impregnating solution. An increase of the thiophene conversion per mass unit of catalyst was observed upon increasing the Mo loading with En in the impregnating solution. For Co-Mo-En and Ni-Mo-En samples, the same evolution of activity is observed in this work. The promoting factor when Co or Ni are added to the catalysts is the same whatever the Mo loading (about 6 for Co-based catalysts and 8 for Ni-based catalysts). The present results clearly suggest that the use of En improves the dispersion of the Mo on the alumina. Moreover, it induces a better dispersion of the Co atoms in a decoration position of the M0S2 crystallites. Such an effect is not observed with the nickel as a promotor. This has been attributed to a different stability of the aquo-complexes or ethylenediamine complexes present in the pores of the alumina support.
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[2] [3] [4] [5] [6] [7] [8] [9] [10] [II] [12] [13] [14] [15] [16] [17] [18] [19] [20] [21] [22] [23]
(a): J.A.R. van Veen, E. Gerkema, A.M. van der Kraan, A. Knoester, J. Chem. Soc. Chem. Commun. 1684, 1987 and (b): S.M.A.M. Bouwens, F.B.M. van Zon, M.P. van Dijk, A.M. van der Kraan, V.H.J, de Beer, J.A.R. van Veen, D.C. Koningsberger, J. Catal. 146, 375, 1994. Yoshimura, N. Matsubayashi, T. Sato, H. Shimada, A. Nishijima, Appl. Catal. A, 79, 145, 1991. P. Blanchard, C. Mauchaussee, E. Payen, J. Grimblot, O. Poulet, N. Boisdron, R Loutaty, Stud. Surf. Sci. Catal., Elsevier, Amsterdam, 91, 1037, 1995. L.G. Roberts, F.H. Field, J. Am. Chem. Soc. 72, 4232, 1950. C.K. Jorgensen, Adv. Chem. Phys. 5, 33, 1963. B.J. Hattaway, C.E. Lewis, J. Chem. Soc. A, 1183, 1969. M.C-L. Yang, R A Palmer, J. Am. Chem. Soc. 97, 5390, 1975. N.Y. Tops0e, H. Tops0e, J. Catal. 75, 354, 1982. J. Abart, E. Delgado, G. Ertl, H. Jeziorowski, H. Knozinger, N. Thiele, X. ZH. Wang, Appl. Catal. 2, 155, 1982. P. Dufresne, E. Payen, J. Grimblot, J.P. Bonnelle, J. Phys. Chem. 85, 2345, 1981. E. Payen, Doctoral Thesis, Lille, France, 1983. H. Jeziorowski, H. Knozinger, Appl. Surf. Sci. 5, 335, 1980. E. Payen, R. Hubaut, S. Kasztelan, O. Poulet, J. Grimblot, J. Catal. 147, 123, 1994. S. Houssenbay, S. Kasztelan, H. Toulhoat, J.P. Bonnelle, J. Grimblot, J. Phys. Chem. 93, 7176, 1989. L Alstrup, L Chorkendorff, R. Candia, B.S. Clausen, H. Topsoe, J. Catal. 77, 397, 1982. C. Wivel, B.S. Clausen, R. Candia, S. Morup, H. Topsoe, J. Catal. 87, 497, 1984. M. Wu, D.M. Hercules, J. Phys. Chem. 83, 2003, 1979. J.M. Rynkowski, T. Paryjczak, M. Lenik, Appl. Catal. A : Gen. 106, 73, 1993. L. Bonneviot, O. Clause, M. Che, A. Manceau, H. Dexpert, Catal. Today, 6, 39, 1989. O. Clause, M. Kermarec, L. Bonneviot, F. Villain, M. Che, J. Am. Chem. Soc. 114, 4709, 1992. P Dufresne, E. Payen, J. Grimblot, J.P. Bonnelle, J. Phys. Chem. 85, 2344, 1981. H. Topsoe, B.S. Clausen, Appl. Catal. 25, 273, 1986. J.A.R. van Veen, E. Gerkema, A.M. van der Kraan, P.A.J.M. Hendriks, H. Beens, J. Catal. 133, 112, 1992..
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® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
225
Creation of acidic sites by hydrogen spillover in model hydrocracking systems. Alexandre M. STUMBO, Paul GRANGE, Bernard DELMON University catholique de Louvain, Unite de catalyse et chimie des matdriaux divis6s. Place Croix du Sud, 2/17, B-1348 Louvain-la-Neuve (Belgium) ABSTRACT The results concern mechanical mixtures of sulfided CoMo (14% M0O3, 3% CoO) deposited on a non-acidic support, i.e. Si02, with silica-aluminas (6.5,12 or 60 wt.% AI2O3). A strong synergy between the two components is observed: (i) in the extremely selective cracking of diphenyl methane (DPM) to benzene and toluene under hydrogen pressure (5 MPa; 673 K; feed (wt.%): 29.5 DPM, 70 n-dodecane and 0.5 benzothiophene added for keeping the catalyst sulfided); (ii) FTIR band at 1540 cm-^ of pyridine relatively strongly adsorbed on Bronsted acid sites; and (iii) extend of D-H exchange as detected by FTIR OD bands. No hydrogenation of the benzenic rings is observed. Spillover hydrogen increases the acidity, the rate of cracking of DPM and the rate of D-H exchange. A new mechanism for explaining the cracking of DPM in the presence of hydrogen is proposed. In addition to creating Bronsted sites, spillover hydrogen would act as a dual hydrogenizing agent, with spillover H+ creating first a carbonium ion, which would decompose to carbenium, this carbenium reacting with a spillover hydride ion H- to form toluene. 1 . INTRODUCTION The objective of this work was to investigate the possible role of spillover hydrogen (Hso) in mild hydrocracking. More precisely we investigated whether the acidity of silica-aluminas could be enhanced by Hso- There are strong indications that an enhancement of the acidity of zeolites could be due to spillover (1-5). Similar effects seem to occur with other acidic materials (1,6). However, the catalysts used in these last studies were obtained by impregnation of acidic supports with noble metals. The effects could be extremely sensitive to the impregnation procedure and to modifications due to the impregnation process. With the objective to clarify the situation with respect to non-zeolitic materials and to avoid the pitfall of impregnation, we used silica-aluminas and mixed them mechanically with the Hso donor. On the other hand, the sources of Hso in previous studies were metals. This is not relevant to usual hydrocracking in the petroleum refining industry, where catalysts work in the presence of sulfur containing gases, and are very often constituted of sulfided CoMo or NiMo mixtures. We therefore decided to work with a sulfided CoMo catalyst supported on a non-acidic support, mechanically mixed with silica-aluminas. It was thus hoped to detect changes of acidity due only to the phase in contact with the source of Hso, ^^ the exclusion of those due to this source. We used the cracking of diphenylmethane (DPM) as test reaction. Using sulfide catalysts, DPM splits to benzene and toluene without hydrogenation of the unsaturated ring (5). The reaction is simple, thus facilitating the interpretation of results. The activity data, showing a synergy between CoMo/Si02 and silica-alumina, have been presented previously (7). We report here data concerning the enhancement of Bronsted acidity
226 and H-D exchange. A tentative cracking mechanism, involving spillover species, will be suggested. 2 . EXPERIMENTAL 2 . 1 . Materials Three commercial amorphous silica-aluminas were used as acidic phases. They contained 6.5, 12 and 60 wt.% AI2O3 and will be called SA6, SA12 and SA60, respectively. Their surface area was 500 m^g-^ and their pore volume and pore size were between 0.9 and 1.3 ml.g-1 and 6.6 and 111 A, respectively. The precursor of the Hso donor, a CoMo/Si02 catalyst (14 wt% M0O3 and 3% CoO) in its oxide form, was prepared by impregnation of silica (130 m^.g-^) with aqueous solutions of successively cobalt acetate and ammonium heptamolybdate (Merck, ultra pure). After each impregnation step, the sample was dried overnight at 393K and calcined at 673K for 2 hours, under a stream of air (Air Liquide, S). In order to make the mechanical mixtures, the pure phase was first ground and sieved, to obtain particles of sizes under 40 jim. The powders, previously dried ovemight at 393K, in the desired proportions, and n-pentane (15ml/g solid) were mixed to form a suspension, which was immersed in an ultrasonic bath for 5 minutes. The suspension was submitted to vigorous mechanical agitation (30(X) rpm) with an Ultra-Turrax T-50, for 10 minutes; n-pentane was then evaporated at room temperature, under a stream of argon (Air Liquide, N46) and under continuous magnetic stirring. After drying at 393K ovemight, the powder was pressed (10 ton.cm-^), gently ground and sieved. Particles of sizes between 0.315 and 0.5 mm were obtained. The pure phases were submitted to exactly the same procedure. The samples will be identified by their relative weight content of silica-alumina, named Rm, defined as: P wt.% Si02 - AI2O3 ... ^m- ^^^^ CoMo/Si02 + wt.% Si02 - AI2O3 ^ ^ An extensive quantitative XPS study did not detect any contamination of the silica-aluminas with Co or Mo. The sulfidation was realized in situ. A flow (100 ml.min-^ of argon (Air Liquide, N46) was first established, the temperature raised to 423K (heating rate: 10 K.min-^) and maintained at this value for 30 min. The gas was then changed to a mixture of 15% (vol.) H2S (Air Liquide, N28) in H2 (Air Liquide, N30) at the same flow rate. The temperature was raised to 673 K, at lOK.min-^ and kept at that level for 2 hours. 2 . 2 . Catalytic tests The reaction was carried out in a continuous-flow tubular reactor. The catalytic bed (1 g catalyst diluted with enough carborundum to reach a volume of 4 cm^, i.e., a height of 7 cm in the reactor) was placed between two plugs of glass wool. The rest of the reactor was filled with 1 mm glass spheres. It was verified that glass and carborundum are completely inert. A thermocouple was placed in a well along the axis of the reactor. The liquid feed was mixed with pure H2 and vaporized before reaching the catalytic bed. The effluents of the reactor were condensed under high pressure. At the end of the 2 h sulfidation period (see above), the pumping in of the model liquid feed was initiated. After a delay long enough to allow the charge to reach the catalytic bed, the gas was switched to pure H2 (Air Liquide, N30) and the pressure was progressively raised until its final value. The test temperature was 673 K and the total pressure 5 MPa. The liquid feed flow rate was 16.5 ml.h.i and the hydrogen flow rate 24 l.h"^ (STP).
227 The model liquid feed was composed of 29.5 wt.% of DPM (Fluka, 99+%), 70% of ndodecane (Aldrich, 99+%) and 0.5% of benzothiophene (Aldrich, 95%). The role of benzothiophene was to produce H2S and keep the catalyst sulfided. Liquid samples taken at regular intervals were analyzed by gas chromatography, using an Intersmat IGC 120 FL apparatus, equipped with a flame ionization detector (PH2 = 1«4 bar; Pair = 1.0 bar). The products wre separated by a capillary column, AUtech CP-Sil-8CB (50 m x 0.25 mm, film thickness of 0.4 |im). He (Air Liquide, N45) was the carrier gas. The estimated precision of the chromatographic analysis is about 5%. The results of the catalytic tests were expressed as DPM total conversion, calculated as follows: % conversion = % ^ x l O O
[2]
where Co and Cf are, respectively, the concentrations of DPM at the inlet and the outlet of the reactor. The experimental results, expressed as DPM total conversions, were compared to a theoretical sum of the individual contributions of the pure phases, calculated assuming, as an approximation, a zero-order reaction and the absence of interactions of any kind between these phases: Ct = ^ X C R „ = I + ( 1 - ^ ) C R ^ = O
[3]
where CRin=l and CRm=0 are, respectively, the experimental conversions corresponding to pure silica-alumina and pure CoMo/Si02. Differences between these calculated values Ct and the experimental resuhs CRm constitute an evidence of a synergy between the phases. The effect can be quantified using a quantity we call "intensity of synergy" (Isyn): Isyn = ^ ^ ^ ^ 1 0 0 [4] 2 . 3 . FTIR of adsorbed pyridine The samples were pressed (2 ton.cm"^, for 15s) in the form of wafers (about 13 mm diameter and weighing between 3 and 5 mg). These wafers were placed in a specially designed Pyrex cell, that allowed the heating of the sample under vacuum or controlled atmosphere. IR spectra could be taken through NaCl windows. The samples were submitted to exactly the same sulfidation procedure as the samples used for catalytic tests, followed by 2 h of heating at 673 K under vacuum (about 2xlO-^Pa). After cooling under vacuum, pyridine was adsorbed at room temperature for 30 minutes. The samples were then outgassed in four steps of Ih: the first one at room temperature and the others at 423 K and 523 K. FTIR spectra were taken before pyridine adsorption and after each outgassing step, with a Fourier Transform Infrared Spectrometer Bruker IFS-88 (spectral resolution set at 1 cm-^). Each spectrum represented the average of at least 50 scans (analysis time about 5 min). Corrections have been made to take into account the differences in weight and surface of the wafers. After each test, each wafer was weighed and its cross section measured with a planimeter. The results were corrected Ac) to represent those of a "standard wafer" of 5 mg and 25 units of area, according to the following expression: 25 5 Ac = AeXo-x—[5] where Ae is the experimental integral absorbance of the band considered, Sw is the cross section of the wafer (in arbitrary units) and mw is the weight of the wafer (mg).
228 2 . 4 . H-D Exchange The samples were prepared and pretreated and the IR spectra were taken and normalized exactly as in the case of FTIR of adsorbed pyridine. A small refinement was however added. Since the IR beam does not pass through the center of the wafer, each sample was analyzed when in 2 different positions. The differences are inferior to 5%. The reported data correspond to the average of the 2 measurements. In order to avoid possible interferences, the cells used, although of exactly the same design, were different. For measurement of the H-D exchange, the temperature was raised to 423 K, at a rate of 10 K.min-i. After temperature stabilization, deuterium (Air Liquide N28; 80 kPa) was admitted in the cell. Several spectra were taken at regular intervals during approximately 24 h. Each one represented the average of at least 50 scans (analysis time about 5 min.). Before each analysis, the sample was cooled to room temperature. Preliminary tests had shown that no exchange took place at that temperature, even after several hours. The amount of deuterium exchanged was calculated from the total area of the OD bands situated between 2800 and 2100 cm-i, where no peaks are detected in the absence of deuterium. The surface was measured using the software supplied by the FTIR manufacturer. The experimental results are compared to theoretical values assuming no interaction between the phases, using a formula identical to that used for the catalytic activity. 3 . RESULTS 3 . 1 . Catalytic tests The products of diphenylmethane (DPM) cracking were benzene and toluene. Small amounts of polymerized products were formed. No cyclohexane, nor other totally or partially hydrogenated products (as, for example, cyclohexylphenyhnethane) have been detected. Figure 1 shows, as an example, results obtained with SA60, as a function of the composition of the mechanical mixtures, one hour after the beginning of the reaction and at the steady-state. The dashed lines represent the sum of the individual contributions of the pure phases, calculated according to assumptions made in the experimental part. ^ 100
SA60
Similar results are obtained with the other silica-aluminas. Each series presents a maximum of activity, but at 50 different Rm values (8). SA6 series has a maximum between Rm values of 50 and 75, whereas SA12 series has a maximum around Rm = 50, and the present SA60 1 >ri series near Rm = 75. A very important synergetic effect is observed in all series, 100 i.e., the cracking activity of the SABojfCoMojSiO * SA60) fwt %J mechanical mixtures is considerably Figure 1. Diphenylmethane conversion as a higher than the sum of the individual function of the composition of mixtures CoMo/Si02 contributions of the pure phases. Figure 2 shows the intensity of the synergy Igyn + SA60: o after 1 h; • at steady state (24 h). in each series, as defined in tne experimental part, as a function of the composition of the mechanical mixtures. The cracking activity of the mixtures can be as much as 8.5 times higher than the properly averaged sum of the activities of its components.
" N\;
229 NO
800
SA12 c
1.00
Figure 3 shows the evolution of the activity with time on stream. There is a loss of activity between the beginning of the sampling (time on stream of Ih) and the steady state.
jg' SA60
r
The three series display similar deactivation patterns. Pure silicaaluminas are strongly deactivated at the beginning of the reaction. They lose about 80% of their activity before 50 100 reaching the steady state. The loss of activity of the pure CoMo/SiC)2 catalyst SAJiCoMojSiO^*SA)[wt%J is slower and much less pronounced Figure 2. Intensity of the synergy in DPM cracking, (about 15%). Mechanical mixtures Isyn> as a function of the composition of the represent an intermediate case between mechanical mixtures. the situations described above. They lose some activity in the first hours of reaction, but the loss is not so pronounced (between 35 and 50%) and the decrease of activity is slower than in pure silica-aluminas. The deactivation is accompanied by coke deposition. Table 1 shows the carbon contents after catalytic test on the pure phases and mechanical mixtures of CoMo/Si02 with S A60. The experimental results are compared to the theoretical values obtained by adding the contributions of the individual pure phases, account taken of their proportions in the mixture. The carlx)n contents in the other mechanical mixtures are also very significandy lower than the calculated values.
100
SA6
'^ 50
*
^
100
%t 50
SA12
4-
D
i i 100
SA6 0
) 50
h
* 10
••
20 Time [h]
Figure 3. Evolution of DPM conversion with time on stream. • CoMo/SiOa; • Rm= 25; A R^ = 50; O R^ = 75; O SA
230 Table 1. Carbon content (wt%) in mixtures of C0M0/AI2O3 with SA60 after catalytic test (24h). Rm 0 0.25 0.50 0.75 1.00
experimental 0.45 2.3 4.5 8.4 25.4
if C content same as in pure components . 6.7 12.9 19.0 -
3 . 2 . FTIR of adsorbed pyridine The CoMo/Si02 catalyst did not present any Bronsted acidity. Figure 4 indicates the amounts of Bronsted sites, as measured by the surface of the IR peak at approximately 1540 cm-i, as a function of the composition of the mechanical mixtures made with SA60. A very clear effect is observed at all outgassing temperatures. After treatment at 523 K, the amount of Bronsted sites increases by 300 to 600%. In the SA6 series (9), the amount of Bronsted sites measured after outgassing at 523 K increases slightiy 0 (by about 15% and 30% for R^ = 50 A and 75, respectively), when compared 0 to the sum of the individual 0.19 contributions of the pure phases. The ) \ ^ effect is smaller after degassing at room temperature or 423K. In the SA12 series, a clear increase is observed after heating at 423 K. The 0.00 50 100 effect becomes more evident after the SAeojCoMolSiO^ *SA60) [wt %} treatment at 523 K, when the pyridine in interaction with the weaker sites has Figure 4. Amount of Bronsted sites, as measured been eliminated. In this case, all the by the absorbance of adsorbed pyridine at 1540 cm-i, mechanical mixtures present more after degassing at: • room temperature; • 423 K acidic sites (63%, 187% and 214% for and • 523 K. The curve corresponds to the trend Rm = 75, 25 and 50, respectively) than observed after desorption at 523 K. the added contributions of the pure phases. 0.38
SA60
3 . 3 . Isotopic exchange using D2 All the samples showed two kinds of OD bands: a relatively sharp one around 2760 cm-^ and a broad one at lower frequencies, in a range whose limits were always between 2700 and 2100 cm-i. The former is assigned to isolated OD species and the latter to OD species in interaction (hydrogen-bonded species (10-12). The pure silica-aluminas exchanged tiie lowest amounts of deuterium among all the samples. Their OD bands were not very intense and increased very slowly with time. The pure CoMo/Si02 catalyst, where a potential source of spillover hydrogen is present, exchanged a higher amount of deuterium tiian the pure silica-aluminas. The peak of isolated deuteroxyls represented the major part of the deuterium exchanged. The band of interacting deuteroxyls was much weaker tiian the other one. The three series of mechanical mixtures had bands located exactiy at die same positions as those of the pure phases. However, their intensities were
231 considerably higher. The growth in intensity of the interacting deuteroxyls bands was more important than that of the isolated deuteroxyls. SA60
§
8 •
65
\
^/# -Q
^Q ^QQ SA6o/(CoMo/SiO, *SA60I fwt%] / ^ Figure 5. Total area of the OD bands after 20h of exchange. is expressed, as previously, by Isyn, as defined in a general way in the experimental part. The enhancement of the D-H exchange due to the contact between CoMo/Si02 and the silicaaluminas is considerable: it corresponds to a multiplication by a factor of about 23 1000 when 1 g of SA60 is mixed with 0.33 g of CoMo/Si02. Qualitatively similar results, often with still more intense synergies, are observed in the presence ofHaS, and at other temperatures (9). ^
4.
DISCUSSION
Figure 5 shows the variation of the total area of the OD bands after 20 h of exchange, as a function of the composition of the mechanical mixtures of CoMo/Si02 with SA60. The maximum of the OD bands intensity corresponds to the maximum of catalytic activity (fig. 1). The same observation can be made for the series of mixtures containing SA6 and SA12. Figure 6 summarizes the data Concerning the intensity of the synergy between the mixed phases with respect to enhancement of the OD bands. This
4SA60
j
/ ^SA12
^ ^^
^oo
SAfCoMo/SiO^ * SA) [wt %J
Figure 6. Intensity of the synergy in OD exchange, Isyn, as a function of the composition of the mechanical mixtures.
Taken together, the three groups of results, namely activity in DPM cracking, intensity of the 1540 cm-^ band of adsorbed pyridine and D-H exchange, strongly suggest that spillover hydrogen emitted by the sulfided CoMo/Si02 catalysts promotes the acidity of silica-aluminas. Before formulating a more precise conclusion, however, a few points should be discussed. On the other hand, our introduction indicated that the cracking of DPM did not correspond to an ideal hydrocracking. We shall therefore be led to discuss a possible mechanism for this cracking. In parallel with similar new proposals in literature, the end of the discussion will suggest the possibility that spillover hydrogen could have two interrelated roles, that of increasing the Bronsted acidity of silica-aluminas and that of acting as real reactant in a new type of hydrocracking. 4.1.
Critical discussion of the results The most striking and convincing proof of a spillover of hydrogen from sulfided CoMo/Si02 to silica-alumina is the considerable enhancement of the D-H exchange rate on the hydroxyls. It is therefore very important to discuss whether the synergy observed in deuterium exchange under our experimental conditions is really a consequence of spillover. Species like
232 HDO or D2O may be responsible for a non-spillover exchange mechanism (13). These species could also be formed by the reaction between oxygen traces and deuterium (13). Some authors also mentioned that water and the corresponding deuterated species were responsible for a direct exchange between the gas phase and the surface of the solid (13,14). The possibility could therefore exist in our case that water, or HDO and/or D2O, that could be present in the deuterium as impurities, would be responsible for the observed exchange. If we accepted this explanation, we should have observed a significant exchange in the pure silica-aluminas, where there was no Hso source. This was not the case. We must therefore conclude that there is no or only a negligible direct exchange through the gas phase. Consequentiy, we must conclude that our results represent a direct evidence for the existence of spillover phenomena. D2 is adsorbed and dissociated on the metal sulfides. The species created in this way can then migrate (spillover) to the surface of the silica (primary spillover) and of the silica-alumina (secondary spillover), making the exchange with the hydroxyls of the latter possible. It is not excluded that water might have another role, namely that of accelerating the exchange between spillover deuterium and -OH groups, but the only point relevant to the present work is the existence of a spillover of deuterium and hydrogen from CoMo/SiOi to silica-alumina. A possible second role of water restricted to the surface of the silica-aluminas does not change our conclusion concerning this existence. Another point only needs a short mention. This is the fact that spillover oxygen diminishes the amount of deposited coke. This coke most probably poisons acidic sites. TTie consequence of this effect is to depress the rate of DPM conversion. lliis is sufficient to explain why, on the whole, the effect of spillover on DPM cracking is less important than on the D-H exchange of -OH groups, where no such depression by coking takes place. A last point in this section concems the relation between the activity (DPM conversion) and the amount of Bronsted sites, as measured by the intensity of the 1540 cm-^ band of adsorbed pyridine. There is some correlation, but no proportionality (fig. 7). Actually, a strict proportionality could not be expected. These reasons will be mentioned hereafter. First of ail, there is a striking difference of activity between the pure siUca-aluminas and the mixtures: for the same acidity, they have a much lower activity. This discrepancy can be easily explained. Pyridine adsorption was performed on freshly sulfided samples, whereas conversion data came from catalysts submitted to reaction conditions. The apparent inconsistency is due to the ability of the Hso to moderate the deactivation by coke deposition on acidic sites. The activity of the pure silica-aluminas is certainly relatively high in the first moments of the reaction but, due to the absence of a Hso source, the acidic sites (whose presence is demonstrated by the FTIR ^ so I I results) are quickly deactivated. On the ' ' ' other hand, mechanical mixtures benefit from the action of Hso^ existing sites m • are partially protected from coke A formation and new sites can be created. mechanical Other remarks concern the mixtures 30 mechanical mixtures and the absence of proportionality between activity and number of Bronsted sites (fig. 7). 1 CoMo Former results of our group (15,16), k SA12 • SA6 • SA60 together with numerous other data of literature, show that the repartition in 0.2 OM strength of Bronsted (and Lewis) sites Bronsted sites (15^0 cm'^ band)is very different according to silicaalumina composition. As the discovery Figure 7. Correlation between the catalytic activity at that spillover hydrogen can create acid the steady state and the amount of Bronsted sites as sites on amorphous silica-aluminas is measured after outgassing at 523K (IR band at 1540 new, there is no surprise that no study cm-i of adsorbed pyridine).
233 has been conducted yet to evaluate the capacity of various sorts of silica-aluminas to generate new Bronsted sites. These could be very different according to the silica-alumina composition. Another reason for not obtaining a linearity is coke deposition which, according to our results (table 1) varies with silica-alumina composition. A last reason is that, as shown recently in our laboratory (17), silica-aluminas exhibit very largefluctuationsof the Si-Al ratios, from spot to spot, this necessarily bringing about fluctuations in nature and strength of the acid sites. There are many reasons to believe that these fluctuations do not have the same intensity at all silicaalumina compositions. 4 . 2 . Cracking mechanism of diphenylmethane The ability of unsupported or supported transition metal sulfides to dissociate molecular hydrogen (or deuterium) is well known (18-21). There is still some uncertainty as to the nature of the hydrogen species formed: H* due to homolytic scission of H2, or H'*"+H", by heterolytic scission. In zeolites, the results of Roland et al. (22) using the Hall effect, demonstrate that H spillover species (Hso) are charged, thus suggesting that the second hypothesis can be retained. The existence of Hgo in the form of H+ could easily explain the D-H exchange and the enhancement of Bronsted acidity. But it is more difficult to explain how Bronsted sites could catalyse the cracking of DPM. An "ideal hydrocracking", through a classical bifunctional mechanism, is impossible because DPM cannot be dehydrogenated. We propose hereafter a mechanism inspired on the one hand, by that proposed for the dealkylation of isopropyl benzene (scheme I), which corresponds to a classical cracking mechanism via a carbonium ion (23) and, on the other hand, by a new mechanism proposed by K. Fujimoto (4,24). The essential idea of K. Fujimoto is that two mobile species are formed, namely H+ and H", and that the first one promotes the creation of the carbenium ion, this reacting with the hydride ion H" after isomerization (case of paraffins) or -H"disproportionation (case of toluene), giving the saturated product. It is possible that the model proposed by Roessner et al., where the second mobile species is an electron (25), is just a CHq - C/7 - CHq variant of Fujimoto's mechanism. Our proposal is represented in scheme H. The first part of this scheme is identical to that of scheme I. The difference with Fujimoto's mechanism for the second part is that we assume that CHq - CH - CHy H+ and H- both migrate to the silicaalumina. This mechanism thus supposes Scheme I. that, in addition to promoting Bronsted acidity, spillover hydrogen, as proton and hydride, is a reactant, namely is consumed in the reaction. The cracking of DPM would then be a real hydrocracking, namely using hydrogen, but this is the form of spillover species.
234 As the effect of spillover hydrogen with DPM is very strong, one may reasonably speculate that the effect described in the present article could also play a very significant role in other cracking reactions taking place in the presence of hydrogen, together with "ideal hydrocracking".
5 . CONCLUSIONS
Scheme 11.
Spillover hydrogen produced by sulfided CoMo/SiOi creates relatively strong acidic sites on silica-aluminas. Although our experiments are limited to the cracking of diphenylmethane, we speculate that this effect could play a role in most cracking reactions made in the presence of hydrogen. We suggest that, in the case of diphenylmethane, the spillover species reacts with reaction intermediates. This would correspond to a new hydrocrackmg mechanism. ACKNOWLEDGMENTS The silica-alumina samples were kindly supplied by Akzo Chemicals bv. One of the authors (AMS) thanks the CNPq (National Council for Scientific and Technological Development, Brazil) for financial support. This work is part of a program concerning the role of spillover hydrogen in hydrotreating catalysts supported by the Federal Science Policy Office of Belgium, through the Interuniversitary Pole of Attraction (PAI program) "Katalyse". REFERENCES 1. H. Hattori, in "New Aspects of Spillover Effect in Catalysis" (T. Inui et al., eds.), p. 69, Elsevier, Amsterdam, 1993. 2. F. Roessner, U. Roland, T. Braunschweig, J. Chem. Soc. Far. Trans., 91 (1995), 1539. 3. I. Nakamura, K. Aimoto, K. Fujimoto, AIChE Symp. Ser., 85 (1989), 15. 4. I. Nakamura, R. Iwamoto, A.I-ino, in "New Aspects of Spillover Effect in Catalysis" (T. Inui et al., eds., Elsevier, Amsterdam, 1993, p. 77. 5. H. Hattori, K. Yamashita, T. Tanabe, K. Tanabe, Proc. 9th Int. Congr. CataL, p. 27. 6. M. Lacroix, G.M. Pajonk, S.J. Teichner, Bull. Soc. Chim. Fr., 7-8 (1981) 265. 7. A.M. Stumbo, P. Grange, B. Delmon, Catal. Lett., 31 (1995) 173. 8. A.M. Stumbo, P. Grange, B. Delmon, 11th Int. Congr. Catalysis, Baltimore, June 30July 5,1996, communication A-3. 9. A.M. Stumbo, PhD thesis. University catholique de Louvain, 1996. 10. R.R. Cavanagh, J.T. Yates, Jr., J. Catal., 68 (1981) 22. 11. E. Baumgarten, E. Denecke, J. Catal., 95 (1985) 296. 12. E. Baumgarten, E. Denecke, J. Catal., 100 (1986) 377.
235 13. D. Bianchi, D. Maret, G.M. Pajonk, SJ. Teichner, in "Spillover of Adsorbed Species" (G.M. Pajonk, SJ. Teichner, J.F. Germain, eds.), Elsevier, Anoisterdam, 1983, pp. 4552. 14. W.C. Conner, Jr., J.F. Cevalos-Candau, N. Shak, V. Haensel, in "Spillover of Adsorbed Species" (G.M. Pajonk, S.J. Teichner, J.F. Germain, eds.), Elsevier, Amsterdam, 1983, pp. 31-43. 15. P.O. Scokart, F.D. Declerck, R. Sempels, P. Rouxhet, J. Chem. Soc, Faraday Trans. I, 73 (1977), 359. 16. J.P. Damon, B. Delmon, J.M. Bonnier, J. Chem. Soc, Faraday Trans. I., 73, (1977) 372. 17. C. Sarbu, M. Ruwet, B. Delmon, submitted. 18. M. Karroua, H. Matralis, P. Grange, B. Delmon, J. Catal., 139 (1993) 371. 19. S. Giraldo, P. Grange, B. Delmon, in "New Aspects of Spillover in Catalysis" (T. Inui et al., eds.), Elsevier, Amsterdam, 1993, pp. 345-348. 20. X. Chu, L.D. Schmidt, J. Catal., 144 (1993) 77. 21. N.M. Rodriguez, R.T.K. Baker, J. Catal., 140 (1993) 287. 22. U. Roland, H. Winkler, H. Bauch, K.H. Steinberg, J. Chem. Soc. Far. Trans., 87 (1991) 3921. 23. P.A. Jacobs, in "Carboniogenic Activity of Zeolites", Elsevier, Amsterdam, 1977, pp. 113-120. 24. K. Fujimoto, in "New Aspects of Spillover Effect in Catalysis" (T. Inui et al., eds., Elsevier, Amsterdam, 1993, pp. 9-16. 25. F. Roessner, U. Roland, T. Braunschweig, J. Chem. Soc. Far. Trans., 91 (1995) 1539.
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® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delinon and P. Grange, editors
231
Application of ASA supported noble metal catalysts in the deep hydrodesulphurisation of diesel fuel. H.R. Reinhoudt^ R. Troost^ S. van Schalkwijk^ A.D. van Langeveld^ S.T. Sie^ H. Schulz^ D. Chadwick', J. Cambra^ V.HJ. de Beer", JA.R. van Veen', J.L.G Fierro^and J A. Moulijn^ ^ Delft University of Technology, 2628 BL Delft, The Netherlands Universitat Karlsruhe, Engler Bunte Institut, 76128 Karlsruhe, Germany ^ Imperial College of Science, Technology and Medicine, London, S W7 2B Y, U.K. ^ Escuela de Ingenierios, 48013 Bilbao, Spain ' Eindhoven University of Technology, 5600 MB Eindhoven, The Netherlands ^C.S.I.C, E-28006 Madrid, Spain ABSTRACT The potential of Amorphous Silica Alumina (ASA) supported Pt and Pd catalysts for deep hydrodesulphurisation (HDS) of diesel ftiels was investigated. It appeared that the ASA supported catalysts exhibit an excellent activity for the conversion of 4-Ethyl, 6-Methyl Dibenzothiophene (4-E,6-M DBT) under model conditions as compared to conventional HDS catalysts and Y-AI2O3 supported noble metal catalysts. Pt/ASA was also tested under practical conditions using a diesel ftiel feed. The Pt/ASA catalyst showed a comparable activity to the NiW/y-Al203 catalyst which was higher than that of the conventional CoMo/y-Al203 catalyst. The main difference of the catalyst was the better hydroconversion of the 4,6 di-alkylated DBT's. The better performance of PVASA in the testing under model conditions as compared to the diesel ftiel HDS can be attributed to poisoning of part of the active phase by basic nitrogen compounds like quinoline. It is concluded that ASA supported noble metal catalysts have a promising potential for deep HDS processing. 1. INTRODUCTION Diesel engines play an important role in transportation and local power generation. The demand for diesel ftiel has increased by about 10% each year for the last decade and therefore diesel ftiel production has become a more important segment in oil refineries. However, diesel engines have an important drawback compared to Otto engines with regard to their emission of particulates. Since these particulates are suspected to cause serious health problems, a lot of research effort is currently put in the reduction of these emissions. In general, three possible ways to reduce particulate emission can be thought of, i) improvement of the combustion process (diesel engine), //) end of pipe catalytic soot oxidation and in) diesel ftiel quality improvement. Although the debate on the effects of the various diesel ftiel components on particulate formation is still going on, aromatics and sulphur content are currently restricted by
238 legislation. Since October 1993, the maximum sulphur content of diesel fuel in California (USA) has been reduced to 0.05 wt% while the maximum allowed aromatics content has been decreased to 10 vol%. The European Union has imposed the same maximum sulphur content starting from October 1996, but has not yet restricted the aromatics content. It is expected that these restrictions will be tightened in the near future. Because a high sulphur level hampers the de-aromatisation of diesel fuel, it can be concluded that deep hydrodesulphurisation (HDS) of diesel fuel is a key process in the upgrading of diesel fuel properties. The remaining sulphur components present in diesel fuel after conventional HDS are mainly dibenzothiophenes. It has been well established [1,2] that especially DBT's with alkyl substituents on the 4 and 6 position have a very low reactivity over conventional hydrotreating catalysts like CoMo/y-Al203. This is attributed to the steric effect of the alkyl groups which hinder desulphurisation through direct hydrogenolysis. It has been demonstrated that hydrogenation of one of the benzene rings of DBT practically lifts the steric hindrance [1], thus enabling subsequent hydrogenolysis. However, the hydrogenation potential of the conventional C0M0/Y-AI2O3 is limited for this hydrogenation step. Other routes to achieve a decrease of the steric hindrance are isomerization or de-alkylation of the alkyl groups, but especially the latter is undesirable in industrial applications. Therefore, it is challenging to develop a catalyst with a high hydrogenation activity under industrial (deep) HDS conditions. Although one could imagine that the H2S partial pressure in deep HDS is substantially lower than in conventional HDS, it is still relatively high for the application of noble metal catalysts supported on a conventional support as Y-AI2O3. It is known from the literature that the use of acidic supports, like zeolites, can increase the sulphur resistance of noble metals as platinum and palladium [3,4]. Since zeolites are expensive and, furthermore, have limited accessibility for large molecules as alkylated DBT's, the purpose of this study was to explore the potential of amorphous silica alumina (ASA) as a support for platinum and palladium for application in the deep HDS of diesel fuel. The performance of ASA supported catalysts under model and practical deep HDS conditions will be shown. Moreover, the influence of nitrogen and aromatic compounds on the activity and stability of the most promising catalysts will be addressed to. 2. EXPERIMENTAL 2.1. Catalyst screening Potentially interesting catalysts for deep HDS of diesel fuel have been screened in a batch autoclave reactor. The batch autoclave reactor set-up used in these experiments was described elsewhere [5]. The tested catalysts were obtained from Shell Research and Technology Centre Amsterdam (SRTCA). The properties of the tested catalysts are given in Table 1. The particle size applied in the tests was between 125 and 250 jam. The noble metal catalysts were reduced in-situ with pure H2 in an integrated reduction cell. This cell was heated at 10 K min" up to 623 K and kept isothermal for 1 h. Next, the catalysts were transferred into the HDS batch reactor, which was pressurised to 6.0 MPa with H2. The NiW/y"Al203 and the CoMo/y-Al203 catalysts were sulfided in-situ in the HDS reactor with a sulfiding mixture of 10% H2S in H2 at 0.2 MPa. The reactor was heated with 10 K min" up to 648 K and kept isothermal for 1 h. A model feed, consisting of 1000 ppm 4-ethyl, 6-methyl dibenzothiophene (4-E,6-M DBT) in n-hexadecane (Aldrich, 99%+), was used to test the
239 activity of the catalysts at 633 K and 6.0 MPa. No initial H2S was added to the reactor, the maximum H2S/H2 ratio at complete conversion was 0.3 mol%. Table 1 Catalysts investigated Metal loading [wt%] Pt/ASA Pd/ASA Pt/Y-Al203 PtPd/y-Al203 CoMo/y-Al203 NiW/Y-Al203
0.7 (Ft) 1.0 (Pd) 1.0 (Pt) l.O(Pt), l.O(Pd) 3.1 (Co), 9.0 (Mo) 1.2 (Ni), 15.6 (W)
The influence of nitrogen containing compounds and aromatics on the conversion of 4-E,6-M DBT was also evaluated in the batch autoclave reactor using similar procedures as described above. The experiments were carried out with 1000 ppm 4-E,6-M DBT in nhexadecane and 0.5 mol% H2S in H2. Quinoline and anthracene were used as model compounds for nitrogen containing compounds and aromatic compounds in diesel fuel respectively. The influence of these compounds was determined both separately and simultaneously. 2.2. Catalyst activity in deep HDS of diesel fuel These tests were carried out in a continuous flow micro-reactor. The reactor has an internal diameter of 10 mm and is equipped with a thermowell (5 mm o.d.). Catalysts selected from the screening experiments were used. Also, the pre-treatment was analogous. For these experiments, a particle size fraction between 180 and 250 jim was used. The catalyst bed consisted of 4 g catalyst diluted with 2 g 100 jam SiC particles to optimise hydrodynamics. Table 2 Composition and physical properties of hydrotreated gasoil feed sulphur content nitrogen content aromatics monodipoly-
[ppm] [ppm]
760 60
[mmol g" ] [mmol g" ] [mmol g' ]
101 9 3
boiling point range density
[K] [kgm-^]
493-653 840
240 The remaining space in the reactor was filled with Si particles (1000 |Lim). A hydrotreated straight-run gasoil (Shell Pemis) was used as a representative feed for deep HDS. The initial sulphur content of the gasoil is 760 ppm S which is mainly present as dibenzothiophenes. Composition and physical properties of the gasoil are given in Table 2. The performance of the catalysts was determined under conditions relevant for industrial HDS, that is, 633 K, 6.0 MPa and WHSV = 4.2 gou gcat"^ h'\ The hydrogen to oil ratio was set at 416 1 kg"\ At complete conversion a maximum H2S/H2 ratio of 0.15 mol% was reached. No additional H2S was added to the feed. After pre-treatment, a stabilisation period of at least 50 h time on stream was allowed before the first sample was taken. The reactor effluent samples were analysed off-line with a Hewlett Packard gas chromatograph (HP 5890 series II) equipped with a 60 m CPSIL-8 CB column (Chrompack, 0.25 mm i.d. and 0.25 |am film thickness) and a sulphur specific detector (Sievers, Sulphur Chemiluminescence Detector (SCD), SCD 355). 3. RESULTS 3.1. Catalyst screening This set of experiments intended to be a screening of potentially interesting catalysts and to determine the performance of the reference catalysts. Hence, the model compound concentration was determined as fimction of reaction time. It appeared that the conversion of the model compound could be described by first order kinetics for all catalysts. On the time scale of a single experiment poisoning nor deactivation effects were observed. The first order rate constants for the catalysts are shown in Figure 1.
2.53
CO
o
1.02
o
0.35
ASA
Pt AI203
1.03
Ml PtPd AI203
0.90
CoMo AI203
NiW AI203
Figure 1. The first order overall reaction rate constant for the conversion of 4-E,6-M DBT over different catalysts (633 K, 6.0 MPa, H2S/H2 = 0.15 mol%, batch reactor).
241 Clearly, Pt/ASA is the most active catalyst for the HDS conversion of 4-E,6-M DBT under the conditions applied. Also, Pd/ASA shows a high activity. Both PtPd/Y-Al203 and NiW/y-Al203 show an activity comparable to that of the CoMo/y-Al203 reference catalyst. Pt/y-Al203 has a considerably lower activity than its ASA counterpart. The ASA support does not show any activity at all.
CHj
CHj CHj
I
C2H5CH3
I 760 ppm S
A.JU
V-AI2O3 180 ppm
IL.A,^'U\AjlAN.AA^,/w».Xu..J>t/^
-Ai,.M.vxovy^.>s..^_^._CoMo/Y-Al203, 140 ppm Kj^—^^-^-^^'..X^-,,..,^^
80 ppm S . NiW/y-AljOs, 70 ppm
Retention time Figure 2. GC/SCD profiles of the diesel feed and desulphurised products after deep HDS at 633 K, 6.0 MPa and WHSV = 4.2 g^i, g^at'^ h"^ over the catalysts indicated in the figure. 3.2. Catalyst activity in deep HDS of diesel fuel To establish the activity for deep HDS of diesel fiiel, some interesting catalysts from the screening experiment, i.e., CoMo/y-Al203, NiW/y-Al203, Pt/ASA and Pt/y-Al203, were selected. The GC/SCD profiles of the sulphur compounds present in the feed are shown in the topmost part of Figure 2; the most important compounds are identified. Also, in Figure 2 the GC/SCD profiles of the desulphurised products are presented for the catalysts tested. From this figure it appears that Pt-ASA and NiW/y-Al203 are the most active catalysts under the reaction conditions applied. The total remaining sulphur concentration in the reactor effluent is slightly lower in case of NiW/y-Al203 (70 ppm) as compared to Pt/ASA (80 ppm). Pt/y-Al203 and CoMo/y-Al203 show lower activities, leading to 180 and 140 ppm S sulphur in the reactor effluent, respectively.
242 For each catalyst the possible occurrence of deactivation has been checked by duplicating the measurement after 50 h on stream. It appeared that on this time scale no significant deactivation of the catalysts was observed. 4. DISCUSSION From the model compound experiments it appears that the ASA supported catalysts exhibit a high activity for 4-E,6-M DBT conversion. This high activity cannot be attributed to the support only, since ASA itself does not have any activity for this reaction. According to literature [3], electron deficient Pt or Pd clusters may be formed on acidic supports resulting in a higher sulphur tolerance of these noble metals and, hence, a higher intrinsic activity for certain reactions. When the performance of the Pt/ASA catalyst is compared to that of the Pt/y-Al203 catalyst, a difference of about one order of magnitude is observed. Clearly, the ASA support has a highly beneficial effect on the noble metal active phase. A possible explanation is a decreased stabilisation of sulphur on Pt clusters induced by the acidic sites of the ASA support. This leads to a larger number of active sites available under steady state reaction conditions, thus yielding a higher hydrogenation activity under the conditions applied. However, an explanation based on differences in dispersion between the alumina and ASA supported catalysts can not be ruled out at present. Based on the diesel fuel activity tests it can be concluded that both NiW/y-Al203 and Pt-ASA show high activity for the deep HDS of diesel fuel. However, one could argue that the difference in performance between NiW/Y-Al203 and Pt/ASA and, for example, a standard CoMo/y-Al203 HDS catalyst, is not very large when expressed in terms of total conversion. In Table 2, the total conversion is given for the four catalysts. It must be taken into account, however, that the feed although pre-desulphurised still contains compounds like DBT and 4-M DBT. These compounds have a reasonable reactivity over conventional HDS catalysts since they can be desulphurised without preceding hydrogenation step, which seems to be essential in the HDS of 4,6 di-alkylated DBT's over conventional catalysts. Hence, to assess the actual performance in deep HDS, it is more relevant to compare the conversion of the refractory sulphur compounds in the feed. As can be observed in Figure 2, compounds like 4-M DBT are almost completely converted over all four catalyst at 633 K and WHSV = 4.2 gon g^at"^ h'\ The most important difference between the NiW/Y-Al203 and Pt-ASA on one hand and C0M0/Y-AI2O3 and Pt/Y-Al203 on the other hand the conversion of the 4,6 di-alkylated DBT's like 4-M,6-M DBT and 4-E,6-M DBT When the conversion is quantified in terms of these refractory sulphur compounds a much larger difference is found (Table 3). Hence, it can be concluded that when very low sulphur levels (< 100 ppm S) are a prerequisite, NiW/Y-Al203 and Pt/ASA are potentially interesting catalysts. By comparing the performance of Pt/ASA and Pt/Y-Al203 it can be concluded that the application of ASA carrier strongly enhances the hydrogenation activity of Pt in a sulphur containing environment. When the performance of NiW/Y-Al203 and Pt/ASA under more practical conditions is compared to that under model conditions, it can be concluded that Pt/ASA behaves slightly less in terms of conversion of refiractory sulphur compounds. As the pre-treatment in both series of experiments was identical and that the H2S level was higher under model conditions,
243 the difference has to be explained in tenns of competitive adsorption and/or poisoning effects induced by other components in the gasoil. Possibly, the performance of Pt/ASA is influenced by strongly adsorbing basic nitrogen compounds or (poly)aromatic compounds. We verified the possible effects of basic nitrogen compounds and aromatics in the batch autoclave reactor under similar reaction conditions as used for the catalyst screening experiments. It appeared that the activity of Pt-ASA for 4-E,6-M DBT was indeed strongly Table 3 Performance in deep HDS of tested catalysts (T = 633 K, WHSV = 4.2 gon g^at"^ h'^), Perform; refractory sulphur is defined as the total amount sulphur minus (mono-substituted) DBT's. Catalyst
Remaining sulphur [ppm]
Conversion [%]
Conversion refractory sulphur [%]
Pt/y-Al^Oa
180
77
59
C0M0/Y-AI2O3
140
82
68
Pt-ASA
80
90
82
NiW/y-Al203
70
91
84
influenced by the presence of 200 ppm quinoline. Quinoline was rapidly converted over Pt/ASA, but 4-E,6-M DBT only reacted very slowly while quinoline was present. After complete conversion of quinoline, 4-E,6-M DBT was converted faster, however with only half the reaction rate as observed in the tests without exposure to quinoline. This observation indicates an irreversible poisoning of part of the active phase on Pt/ASA by quinoline, or its reaction products. Remarkably, when quinoline was added in a higher concentration (1000 ppm), after complete quinoline conversion, the same reaction rate for 4-E,6-M DBT was observed as in the 200 ppm case. Clearly, this indicates that part of the sites active in conversion of 4-E,6-M DBT remains unaffected by quinoline. Tests were also done with both quinoline and anthracene present in the reaction mixture. It appeared that the conversion of 4-E,6-M DBT was not influenced at all by anthracene. It can therefore be concluded that the relatively poor performance of Pt/ASA in diesel fuel HDS as compared to model compound testing, can be explained by poisoning of part of the active sites by basic nitrogen compounds. It seems however, that the nitrogen compound poisoning only affects part of the active sites and that the unaffected sites exhibit a high tolerance towards nitrogen compounds. However, to study these effects in more detail, more extensive research under model conditions has to be done. It should be noted that it is complex to compare NiW/Y-Al203 and Pt/ASA directly because of the different carrier and metal loading. However, considering the possibilities to optimise the present catalysts further, it can be stated that ASA supported Pt and Pd catalysts are very promising catalysts for application in deep HDS processing.
244 5. CONCLUSIONS Amorphous Silica Alumina supported Pt and Pd have excellent HDS performance in the conversion of 4-E,6-M DBT. It is concluded that this can be attributed to ASA-noble metal interaction. Differences in the relative performance between diesel fuel and model compound HDS can be explained by poisoning of part of the active phase by basic nitrogen compounds. Finally, ASA supported Pt and Pd catalysts are promising for the application in deep HDS processing. ACKNOWLEDGEMENT This research was in part supported by the European Union (JOU2-0904). Stimulating discussions with W.R.A.M. Robinson (Eindhoven University), F. Ousmanov and P. Waller (Universitat Karlsruhe), P.L. Arias and J.A. Legaretta (Escuela de Ingenierios) are gratefully acknowledged. REFERENCES 1. T. Kabe, A. Ishihara, Q. Zhang, Appl. Catal. 97 (1993) LI. 2. A. Amorelli, Y.D. Amos, C.P. Halsig, J.J. Kosman, J. Jonker, M. de Wind, J. Vrieling, Hydrocarbon Processing 71 (1992) 93. 3. J.A. Rabo, V. Schomaker, P.E. Pickert, Proc. 3'"^ Int. Congr. Catal., Amsterdam 1964, North Holland, Amsterdam, 2, (1965) 1264. 4. S.T. Homeyer, Z. Karpinski, W.M.H. Sachtler, J. Catal. 123 (1990), 60. 5. J.P. Janssens, Ph.D. Thesis, Delft University of Technology, Delft (1996).
® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F, Froment, B. Delmon and P. Grange, editors
245
Reactor Runaway in Pyrolysis Gasoline Hydrogenation E. Goossens^ R. Donker^ and F. van den Brink^ * DSM Research, Industrial Catalysis Section, P.O. Box 18,6160 MD Geleen, The Netherlands ^ DSM Hydrocarbons, P.O. Box 606, 6160 AP Geleen, The Netherlands
December 23, 1994 the wall of the first stage pyrolysis gasoline hydrogenation reactor in DSM's NAK3 steam cracker ruptured, requiring an emergency shutdown of the plant. A mixture of gasoline, hydrogen and nickel catalyst escaped through the crack and immediately caught fire. There were no personal injuries, but property damage was substantial. There had been no domino effects, nor had there been any danger to the surrounding area. One week after this severe fire NAK3 was put back into operation except for the first stage pygas hydrogenation. Thermodynamic calculations, analysis of the process data and of the catalyst and reactor wall indicated that the bulge and subsequent rupture of the reactor wall were caused by a brief local temperature excursion to 700-750 "C at an operating pressure of 30 bar. This temperature excursion can be explained by maldistribution of the liquid and thus poor heat transfer caused by local excessive carbon formation in the reactor in the period preceding the incident. Disruption of the heat removal can cause a chain of exothermic reactions, viz. hydrogenation of olefins, aromatics, as well as hydrocracking. Results of the investigation as well as the improvements made in the operation and procedures are presented in this paper.
1. INTRODUCTION 1.1 Process description and reactor design The C5+ fraction from a steam cracker contains alkadienes and alkenyl aromatics that will easily form gums, which are detrimental for end uses like automotive gasoline and aromatics production. These compounds need to be removed in order to stabilise the C5+. In the first stage pyrolysis gasoline (pygas) hydrogenation they are selectively hydrogenated over Ni/Al203 catalyst. Mono-olefins and aromatics are allowed to pass the reactor. The reactor is followed by a gas/liquid separator and a distillation unit (figure 1).
246 hydrogen
• fresh hydrogen
{XI—.fresh Cc
j quench
Um ^uenchf
,R601 A,
I*
Hi V601
-M-
start up diluent
—
I
JdistillationJ
I
rQ
MPV6101 to storage
Figure 1.
Schematic representation of pygas hydrogenation unit
The hydrogenation reactor is a trickle-phase unit with two catalyst beds. The system comprises of three phases: hydrogen, liquid C5+ and catalyst. A mixture of hydrogen, C5+ and, optionally, hydrogenated product is fed to the reactor and passed over distributors to the first catalyst bed (3 m). Prior to entering the second catalyst bed (10 m), the feed is mixed with recycled hydrogenated product (quench) and redistributed. Quench is used to control the temperatures in the catalyst beds. The inlet temperature varies from approximately 40 °C (Start Of Run) to 95 °C (End Of Run). The temperature gradient is approximately 60 ° per bed. The temperatures in the catalyst bed are monitored by 18 thermocouples equally divided over 3 thermowells (figure 2). plan view
"TZI" Figure 2.
Arrangement of thermowells and thermocouples in first stage pygas hydrogenation reactor
From the gas/liquid separator, after the reactor, hydrogen is recycled to the inlet of the reactor and made up with fresh hydrogen. The liquid, i.e. hydrogenated product, is partly used
247 for quench, the excess amount being pumped to the distillation section. The distillation section is not in operation during startup of the reactor. In the first phase of start-up, hydrogenated product is recirculated through the reactor without charging raw, i.e. non-hydrogenated, C5+. In this phase, an exothermic effect occurs due to the heat of adsorption of hydrocarbons on the catalyst surface. As the start-up progresses an increasing amount of raw C5+ is charged while a balancing amount of hydrogenated product is sent to storage via MPV 6101. 1.2. The incident December 23, 1994, a few minutes before 18.00 h, the wall of the first-stage pygas hydrogenation reactor ruptured during operation. The incident took place during the start-up of the reactor. Due to the operating pressure of 30 bar a mixture of gasoline, hydrogen and nickel catalyst was blown out of the reactor. The mixture immediately caught fire, resulting in a jet of flames of 40 m. The plant was shut down. There were no personal injuries, no danger to the surroundings, no domino effects but the hydrogenation unit sustained substantial damage. One week after the incident NAK3 was on stream again except for the first-stage pygas hydrogenation unit. 2. RESULTS AND DISCUSSION The investigation was led by a multidisciplinary team of experts from DSM. The regulatory authorities were regularly informed of the progress of the investigation. In the phase of gathering information, production staff were interviewed and various companies operating similar processes were visited. All companies were quite open in discussing their experiences. The problems described here did not seem to be unique. The investigation focused on reconstruction of process data, catalyst characterization, material characterization and thermodynamic calculations in order to establish the cause of the crack and more importantly, to avoid similar problems in the future. 2.1 Process data Reactor R601A had been recharged with fresh catalyst one month prior to the incident. Catalyst was activated and the reactor started up in accordance with the operating instructions. Fully hydrogenated Cg-fraction, free of olefins and aromatics, was used for the startup. The catalyst used had not been presulphided for historical reasons and fear of odour nuisance. Although fresh, unsulphided, nickel catalysts are known for their poor initial selectivity, the process was very stable at the time of the rupture (figure 3). As the plot clearly indicates, temperatures are in the normal operating window and gradually increase from the top to the bottom of the reactor. By the end of the startup, the feed had almost completely been replaced by raw C5+, when the reactor wall ruptured for no apparent reason.
248
^%^ TS6022 ...w.^. TS6019 ys^
TS6016
^ ^ . TS6013 .V
TS6010
time (Dec. 23, 94)
Figure 3.
Process conditions on the day of the incident
From analysis of the process data it became clear that a runaway must have occurred a week before the rupture. This runaway was initiated by a sudden increase in the amount of hydrogenated product being sent to storage via MPV 6101, thereby simultaneously increasing the amount of raw C5+ being charged to the reactor (figure 4).
Figure 4.
Process conditions one week before the incident (hourly averages)
Despite the fact that production staff acted according to operating procedures, temperatures in the reactor rose to 320 °C in 1.5 hours. After another 3.5 hours, temperatures had returned to their normal levels and the hydrogen compressor was restarted. As temperatures immediately started to rise it was decided to shut down, depressurize and purge the reactor with nitrogen. All thermocouples except two indicated a decrease in temperature to normal level: thermocouples TS 6019 and TS 6022 remained at a high level for two days (figure 5).
249
o
Dec. 16, 94
date/time
Figure 5.
Process conditions one week before the incident (hourly averages)
Two days after the initial startup, the reactor was quenched with fully hydrogenated C5, fraction which resulted in a normal temperature level for all thermocouples. After assessment of the situation and the sequence of events, it was concluded that the reactor had not been outside its design specifications (T, P). The restart made on December 21 went very smooth up until to the moment of the rupture. 2.2 Catalyst characterization Characterization by means of TPR, elemental analysis, XRD and SEM proved that the fresh catalyst used was identical to the catalyst previously used. The relative difference in dispersion was 12%. Because of the low absolute metal surface area of approximately 4 m^.g"^ and a high metal loading of 10%, this difference in dispersion is too small to account for any difference in activity. Samples were taken from the catalyst left in and that blown out of the reactor. Excessive carbon formation was observed in many places in the reactor (figure 6): up to 50 wt% C was found in various samples.
1
crack second bed hard coal with pulverized catalyst catalyst with some coal
Figure 6.
Carbon formation throughout the reactor after the incident
250 In the first catalyst bed, carbon had been deposited along the reactor wall and in the centre of the bed on the support grid. Only little carbon had formed near the three thermowells. In what was left of the second bed, carbon deposits were observed along the wall. Carbon deposits were harder and thicker in the vicinity of the crack. Scanning and transmission electron microscopy revealed three types of carbon: amorphous, graphitic and whiskers (figures 7 and 8).
Figure 7.
SEM photograph of whisker-type carbon formed during pygas hydrogenation
1^
Figure 8.
SEM photograph of graphite carbon formed during pygas hydrogenation
251 It is well known that the type of carbon formed strongly depends on the reaction conditions. Atomic carbon and precursors formed by dissociative adsorption of hydrocarbons are transformed to polymeric coke, filamentous coke and graphitic carbon at progressively higher temperatures [1-4]. Polymeric hydrocarbons and graphitic films encapsulate the metal surface and deactivate the catalyst. Filamentous carbon (whiskers) generally do not deactivate the metal surface because the metal particle is lifted from the surface. When present in large amounts, however, the catalyst bed will plug and pellets will break up. The relation between the structure of the precursor and the kinetics and morphology of the coke formed is not well understood. An important parameter in the formation and structure of coke deposited is the presence of hydrogen. In the presence of hydrogen, carbon precursors are gasified which keeps the catalytic surface clean and thus active. Moreover, in a hydrogen-lean atmosphere, the catalyst will take its hydrogen from any hydrocarbons present, thereby enhancing the deposition of coke. In the sequence of events the hydrogen flow was stopped according to operating procedures prior to the incident (i.e. stop the feed of one of the reactants). This slows down the hydrogenation rate but the dehydrogenation and thus the rate of coke formation, will strongly increase. Near the crack, large amounts of whisker type carbon were characterized. It is well known that filamentary carbon has a low tensile strength but a high compressive strength [5]. Locally, these carbon deposits had led to pulverization of the catalyst. Lab tests confirmed that the catalyst tends to grow whisker in a CH4/H2 atmosphere. Initially, the investigation committee was put on the wrong track by the potential compressive forces exerted by filamentary carbon. Later on it became clear from the material characterization (section 2.3) that this compressive force on the wall was not necessary for the crack to occur. No a-Al203 was detected in any of the samples taken inside of the reactor, which would have indicated temperatures higher than 1000 °C. Some samples from outside of the reactor were found to contain a-Al203. Whether this phase was formed prior to the rupture or during the fire that followed caimot be decided. It was clear from elemental analysis that there had been a sulphur gradient present across the reactor: the concentration decreased concurrently in axial direction. This indicates that the feed contained little sulphur. Sulphur is known to enhance catalyst selectivity by selectively poisoning the hydrogenation sites that are most active. 2.3 Material characterization The reactor wall is made of 13CrMo44 steel. The crack, which was approximately 30 cm long and 2 cm wide, was in the centre of a large bulge that had a height of 10 cm. For further characterization a section of 1.5*1.5 m was cut out of the reactor wall, half of which was sent to the regulatory authorities. It was manifest that the wall within a radius of 0.3 m around the crack had briefly been at temperatures of 700-750 "C. At the operating pressure of 30 bar, a wall made of this material will fail at this temperature range. The geometry of the bulge and crack found on the reactor wall are as expected under these conditions. Other potential causes like corrosion, hydrogen embrittlement (low T), Nelson hydrogen attack (high T), material defect at the time of fabrication, creep or low cycle fatigue could be ruled out after an extensive study.
252 2.4 Thermodynamics In order to evaluate the conditions during various stages of operation, thermodynamic calculations were performed. Complete hydrogenation of fresh C5+ results in an adiabatic temperature rise of approximately 600 °C. In the presence of quench, i.e. hydrogenated product containing only olefins and aromatics, this temperature will be approximately 520 "C. If, however, hydrocracking occurs the adiabatic temperature can rise to about 1000 °C. In the case of maldistribution and thus bad heat transfer, temperatures higher than 700 "C can easily be reached. The fact that such temperatures were not detected by the thermocouples indicates that there had been a hot spot near the reactor wall. 3. CONCLUSIONS 3.1 Investigation On the basis of the results discussed here it can be concluded that the rupture of the reactor on December 23, 1994 was caused by a brief local temperature excursion to 700750 °C at an operating pressure of 30 bar. The geometry of the bulge and crack are consistent with this scenario. The hot spot can be explained in terms of liquid maldistribution. Maldistribution hampers heat transfer, leading to unwanted side reactions like hydrogenation of mono-olefins, aromatics and even hydrocracking of hydrocarbons. These reactions are highly exothermic and thus amplify the hot spot, i.e. a high temperature enhances vaporization of the feed, which in its turn results in a decrease of heat transfer due to a lower heat capacity of the vapour (positive feedback, figure 9).
Figure 9.
Formation of hot spot during pygas hydrogenation
The run-up to this situation started with the runaway that took place one week before the rupture. The strong increase in the concentration of fresh C5+ during startup, in combination with a decrease in hydrogen flow, caused the temperature excursion that was
253 observed. This in turn caused excessive carbon formation throughout the reactor, upsetting liquid distribution in the next startup. This eventually resulted in the, undetected, temperature excursion to 700-750 "C on December 23. Catalyst characterization confirmed that large amounts of carbon deposits, having various morphologies, were formed. 3.2 Lessons learned and actions taken The incident demonstrates that a runaway reaction can occur in a trickle-phase reactor loaded with a supported nickel catalyst. Presulphiding the catalyst prior to start up will minimize the possibility of a nmaway and improves the intrinsic safety of the process. Presulphiding can be performed in situ as well as ex situ. In situ sulphiding, by spiking the feed with a sulphur-containing compound like dimethyl sulphide, will result in a sulphur front moving through the reactor bed. Ex situ sulphiding is to be preferred in that it gives a uniform concentration of sulphur throughout the bed right from the start. A critical stage in the operation of a pygas hydrogenation reactor is the startup. The pygas hydrogenation reactor is started up using a specific Cs/Cg stream with a predetermined bromine number, diene number and aromatics content. The new operating procedures ensure that the quench, i.e. hydrogenated product, and hydrogen recirculation used are maximized when wall temperatures exceed 200 °C. They also ensure that a minimum linear velocity of the liquid and a minimum hydrogen:liquid ratio are maintained during operation. These improvements will minimize the chances of excessive carbon formation. As an extra source of information 28 thermocouples have been installed on the outside of the reactor wall. The alarm threshold for these thermocouples has been set at 200 °C. Furthermore, a research program has been set up to investigate the kinetics of runaway reactions during operation of the pygas hydrogenation. Experimental data will be gathered by using model compounds under various conditions (P, T, contact time) and studying their relation with the morphology of the carbon deposits. This incident has once again shown the importance of adequate response to near misses, in this case the runaway which was observed one week before the rupture. A thorough check of the catalyst bed after any runaway is called for in order to prevent future problems. In many cases this will imply that the reactor needs to be opened and the catalyst discharged. The sequence of events that eventually led to the incident had not been identified in the HAZard and OPerability study (HAZOP), although the individual events like runaway and carbon formation were known [6,7]. A thorough Process Safety Analysis (PSA) in an early stage of the process design could have revealed the possibility of this sequence of events. The PSA should include an extensive study of incidents in similar units. This illustrates the importance of a PSA prior to the HAZOP study. Finally, the incident and subsequent analysis illustrate the importance of an open discussion within the industry and academia in topics concerning safety, health and environment.
254 4. ACKNOWLEDGEMENTS Parts of this paper have been reproduced with permission of the American Institute of Chemical Engineers. Copyright © 1996 AIChE. All rights reserved [8]. The authors are indebted to prof. J. Geus and Dr. M. Hoogenraad of Utrecht University for performing the TEM analysis and the catalytic test in the formation of carbon whiskers.
REFERENCES 1. 2. 3. 4. 5. 6. 7. 8.
J. Rostrup-Nielsen and D.L. Trimm, J. Catal., 1977, 48, 155-165. C.H. Bartholomew, Catal. Rev. - Sci. Eng., 1982, 24(1), 67-112. R.T.K. Baker, M.A. Barbier, P.S. Harris, F.S. Feates and R.J. Waite, J. Catal, 1972, 26, 51-62. J. Rostrup-Nielsen, Catalysis - Science and Technology (Eds. J.R. Anderson and M. Boudart), Springer-Verlag, New York, 1-117. M. Hoogenraad, Ph.D. Thesis, Utrecht University, 1996. J.L. Figueiredo and J.J.M. Orfao, Sprechsaal, 1986,112(12), 1139-1142. D.L. Trimm, Progress in Catalyst Deactivation - Proc. NATO Adv. Study Inst. Catal. Deact. (Ed. J.L. Figueiredo), Algarve Portugal, May 18-29, 1981, 31-43. R.A. Donker, Proceedings of the 8th Ethylene Producers' Conference & 5th World Congress of Chemical Engineering, 1996, 5, in press.
® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
255
Surface property of alumina-supported Mo carbide and its activity for HDN Toshihiro Miyao*, Katsuhiko Oshikawa, Shinzo Omi, and Masatoshi Nagai Department of Advanced Materials, Graduate School of Bio-applications and Systems Engineering, Tokyo University of Agriculture and Technology, Koganei, Tokyo 184, Japan The activity and selectivity of the alumina-supported molybdenum carbide for the HDN of carbazole were studied. The effects of catalyst pretreatment and surface composition of the supported Mo carbides on the HDN activity of tiie catalyst were also considered. Catalyst characterization was done using TPR, BET, nitrogen element, XRD, and XPS analyses. It was revealed that the catalyst carburized at 700**C was 2.2 times more active than was the catalyst carburized at SOOT in the HDN of carbazole. The TPR analysis showed that neither transformation of the Mo nitrides to the Mo carbides nor free-carbon deposition occurs below 500T for the alumina-supported catalysts, although above 900''C, a large amount of free carbon is deposited on the surface of the Mo carbide catalysts. For the unsupported Mo catalysts, tiie XRD analysis showed tiiat MOjN was converted to a -MOjC at TOOT, and above 900"C the a-MOjC crystal was transformed to T]-MO3C2 crystal. 1. INTRODUCTION Molybdenum carbide is an active catalyst for various reactions that emulate the catalytic properties of noble metals. Ledoux and his coworkers (1) have reported that in reactions involving n-hexane, Mo carbide catalysts activated by Pt show greater reactivity than does Pt by itself. They have also reported that clean surfaces of M02C exhibit greater reactivity towards reactions such as isomerization and cracking than do conventional Pt catalysts. The isomerization of n-hexane has been studied, revealing that the carbide catalyst can be twentyfive times more active than Pt catalysts, if high specific surface area is obtained (2). Otiier isomerization reactions, such as those for 2- and 3-methylpentane, have been tested by PhamHuu et al. (3). For reactions involving hydrodenitrogenation (HDN), Choi et al. (4) and Schlatter et al. (5) studied tiie HDN of pyridine and quinoline on unsupported MOjC catalysts at 400*C and atmospheric pressure, and reported that M02C were more active than y-MOjN and reduced catalysts. Altiiough a number of studies have focused on unsupported Mo carbides, litfle attention has been paid to the reactions involving the Mo carbides supported on alumina rather than unsupported MOjC powder. Lee et al. (6) have reported the methods of preparation for alumina-supported Mo carbides, such as reduction of M0O3/AI2O3 with Hj followed by carburization, direct carburization with CH^/Hj, and carburization ^ter nitriding with NH3. In this study, the activity and selectivity of the alumina-supported Mo carbides for tiie HDN of carbazole are studied. The effects of catalyst pretreatment and surface composition of the alumina-supported Mo carbides on the HDN activity of the catalysts are also discussed. The Present address: Central Research Institute, Nissan Chemical Co., Tsuboi 722, Funabashi, Chiba 274, Japan.
256 formation of Mo carbides and polymeric carbon on the surface of the catalysts were determined using temperature-programmed reduction (TPR), X-ray diffraction (XRD), and X^-ay photoelectron spectroscopy (XPS). 2. EXPERIMENTAL 2.1. Catalyst preparation 12.5 wt% M0O3/AI2O3 (Nikki Chemical) and 100% M0O3 (Aldrich, 99.99%) were treated in dry air for 1 h at 5 0 0 ^ and either directly carburized or nitrided before carburization: the catalyst was directly carburized at 700T in flowing 20% CH4/H2 (99.9995%) for 3 h, or carburized at 500T, 700T, and 900T in flowing 20% CH4/H2 for 3 h, after nitriding in flowing NH3 gas (99.999%) at 4 1/h from 500T to 700T at a rate of TC /min and holding at 700T for 3 h. The catalysts were
Table 1 Nomenclature of carbide catalysts Temperature ( C) Nomenclature loading Nitriding Carburizing (wt%) 100N7C5 100 700 500 100N7C7 100 700 700 100 700 100N7C9 900 100C7 100 700 12N7C5 12.5 700 500 12N7C7 12.5 700 700 12N7C9 12.5 700 900 12C7 12.5 700
passivated using 1% Oj/He at room temperature for 24 h before activity measurement The 12.5% M0/AI2O3, reduced in flowing H^ B^ ^^ 400**C for 3 h, was also used for comparison. Abbreviated notations of the catalysts are used throughout this paper as shown in Table 1. For example, 12N9C7 denotes 12.5 wt% M0/AI2O3 nitrided at 900'C and followed by carburiding at 700'C, while 100C9 denotes 100% M0O3 carburized at 900T without nitriding. 2.2. Characterization For TPR experiment, the carbided catalyst (0.2 g) was heated in situ at 100"C in flowing He for 1 h after carburization and then cooled to room temperature. The catalyst was linearly heated to 900'*C at a rate of TC /min at a flow of 15 ml/min Hj. The gases desorbed from the catalyst were continuously monitored using an on-line quadrupole mass spectrometer (QMS, UL^^C, MSQ-150A) equipped with a variable-leak valve. The amounts of gases desorbed were calculated by calibration curves. The spectra of the gases were obtained by curve-fitting (Origin, Microcal Co.) the data transferred from the QMS. The specific surface area (Nj-BET) of the catalysts was measured by nitrogen adsorption using a standard BET apparatus after evacuation at 2 0 0 ^ and 1 Pa for 2 h. Analysis of nitrogen element was carried out with a Perkin-Hmer CHN elemental analyzer, while molybdenum analysis was performed using atomic absorption spectroscopy. The compositions of the samples were measured using XRD, with Ni-filtered CuK^ radiation of a Rigdcu X-ray diffractometer operating with a scanning speed of 2° min"^ from 2 6 =15° to 120''. The crystallite size of the sample was estimated from the broadening of the more intense diffraction lines of phase (111) using the Scherrer equation. The surface composition of the catalysts was measured using a Shimazu ESCA 850 spectrometer with monochromatic MgK^ exciting radiation (8 kV, 30 mA). Analyses were carried out at a pressure of 5 X10^ Pa. Binding energy was checked against the Ag 3d5/2 line at 368.0 d/. Overlapping peaks were resolved using a nonlinear least-squares fitting routine and GaussLorents curves.
257 2.3. HDN activity The H D N activity for the H D N of carbazole was measured using a fixedbed microreactor at 3 0 0 T and 10.1 MPa total pressure. The carbided catalyst ( 2 . 0 g, granule) was packed in the middle of the reactor, which was connected upward to a hydrogen gas cylinder and highfeeder pump and downward to a highpressure separator. The reactor feed, consisting of 0.25 wt% carbazole in xylene (nitrogen content: 0.02 wt%), w a s introduced into the reactor at a rate of 2 0 ml/h with a H j flow of 61/h. The H D N rate is calculated from the percentage of conversion calculated on the basis of the number of moles of carbazole converted divided by the number of moles of carbazole in the reactant feed.
3. RESULTS AND DISCUSSION
•
•
^'"""•^^
1
—o—BCH 1 0.15 ~ r
*/
—o—cnn
• /"^
—o-CHB —D—HIIC
•/
-HS—OHC —«—PHC
•/
—•—THCA - — C A «••
(J
200
400
600
1200
Figure 1. Product distribution for the H D N of carbazole on the 12N7C7 catalyst at 300*C and 10.1 MPa total pressure: (BCH) bicyclohexyl, (CHH) cyclohexylcyclohexene, (CHB) cyclohexylbenzene, (HHC) hexahydrocarbazole, (OHC) octahydrocarbazole, (PHC) perhydrocarbazole, (THCA) tetrahydrocarbazole, (CA) carbazDle.
3.1. HDN activity The product distribution in the HDN of carbazole on the 12N7C7 catalyst at 300'C and 10.1 MPa total pressure is shown in Figure 1. The main denitrogenated products were ( ^ = ^ @ ^ (=^(§90 = C g O 1^3,4,4a,9 1,23,4-Tetnihydrobicyclohexyl and cyclohexylcyclohexane. Carbazole Hfcxahydrocarbazoie Decahydrocarbazole carbaxole Cyclohexylbenzene was hardly formed in the reaction. The major hydrogenated product was tetrahydrocarbazole. Other CgO) cSrO CnO Bicyclohexyl Cyckthexylbeozene Cyclohexykyclohexw hydrogenated compounds, hexahydro-, Perhydrocarbazole octahydro-, and perhydrocarbazole were Figure 2. Reaction scheme of the HDN formed in minute amounts during the of carbazole. reaction. The product distribution for 12N7C7 was similar to that for the 12N7 . , . , catalyst, the 12.5% Mo/Al,03 nitrided at 500"C (7) and reduced at 400 C (8). Bicyclohexyl was produced from the C-N bond scission of perhydrocarbazole through tetrahydrocarbazole in a successive hydrogenation of carbazole, as shown in Figure 2. A direct breakage of the C-N bond m carbazole did not take place before a successive hydrogenation of carbazole on the carbided M0/AI2O3 catalyst as well as the nitrided and reduced catalysts. Figure 3 showed that the Mo carbide catalyst (12N7C7) was 2.0 and 1.1 times more active than the catalysts reduced at 500T (12R5) and nitrided at 700T (12N7) for carbazole HDN at 300T, respectively. Carburization enhanced the activity of the Mo/Al^Og catalyst for the HDN of carbazole, compared with the reducing and nitriding methods. The activity of Mo carbide catalyst was extremely high at the first stage but decreased with time on stream to that of the Mo nitride catalyst and reached a steady state in 8 h. JHJ\^
258 3.2. Change of phase of Mo carbide bulk with carbiding temperature The XRD spectra of the unsupported catalysts are shown in Figure 4. M0O3 treated with 20% CH4/H2 at SOOT (100C5) showed M0O2 in the XRD pattern but showed a-MojC at TOOT (100C7). M0O3 was reduced to M0O2 below 500*'C but not transformed to Mo carbides by a direct carburization of M0O3 in a flow of 20% CH4/H2. The carbide prepared directly with a 20% CH4/H2 at 700T (100C7) showed the phase of hexagonal close-packed a-Mo2C, while the carbide carburized at 700**C through the nitride intermediate (100N7C7) showed the phase of facecentered cubic a-Mo2C. Moreover, fee a-Mo2C was transformed to Y1-M03C2 at 900°C (100N7C9). The treatment of yM02N with 20% CH4/H2 at 700°C and 900°C produced a-MOjC and r] -MO3C2, respectively. For the carbided 12.5% M0/AI2O3 catalysts, no carbides were observed in the XRD patterns. 3.3. XPS analysis of bulk Mo carbides Although unsupported Mo carbides were analyzed using XRD, Mo carbides for alumina-supported Mo catalysts could not be fully understood by XRD. In order to obtain more information about the formation of Mo carbides on the surface of the alumina-supported Mo catalysts, the detection of the carbidic elements was focused on by XPS and TPR. Mo3d5/2 Mo3d3/2 and C^^ XPS spectra for the supported and unsupported Mo carbides are shown in Figure 5. 100% MOjC sample has the Mo3d5/2 peak at 227.9 6/ and Cls peak at 283.2 eV. 12N7C7 before reaction showed broad Mo3d peaks at 283.0 and 230.0 d/ but sharp Cls peak at 284.7 d/. The binding energy of XPS Mo3d at 230.0 eV was broad for 12N7C7 but was not as broad as that for the M0/AI2O3 reduced at 400"C. The Mo3d5/2 peak for a mixture of 12.5% M02C and 87.5% AI2O3 was shifted to
0
2
4
6
8
T i m e o n S t r e a m [h]
10
Figure 3. Carbazole conversion versus time on stream at 300°C and 10.1 MPa total pressure.
(a)
M0O2
_A
L^^
eL_
(b)
a-MoiC
(hep)
JL_JA__LJS_ a -M02C
(c)
(fee)
JJL i i A.A.. (d)
r\r^^ M 20.00 40.00
5.00
T1-M03C,
26
A L. 80.00
60.00
Figure 4. The X-ray diffraction patterns of various bulk Mo carbides: (a) C5, (b) C7, (c) N7C7, (d) N7C9
100.00
259 229.2 eV, and a broad Mo3d5/2 peak was observed at a binding energy between 227.9 eV and 230.0 eV, but the Cls peak was broad with a very feeble peak at 283.2 eV. The binding energy of Cls at 283.2 ^ is extremely weak for 12.5% M0/AI2O3 catalyst compared to 100% M02C because of 12.5 wt% loading Mo on alumina. Therefore, the charge effect occurs on the 12.5% M0/AI2O3 carbided at 700"C, and Mo^C was formed on alumina. The M02C crystal had two Cls peaks, one at 283.2 eV (assigned to Mo carbide) and another at 284.7 d/ (assigned to graphite) (1,2). The BET surface area (S ) and the ratio of IMO/IAI in the XPS Mo3d/A12p of the
Mo carbide catalysts are shown in Table 2. These values of IMO/IM ratio were used to evaluate the degree of Mo species dispersed on the surface alumina of the catalysts. The l^H^ ratio was not affected by the surface area of the catalysts during either nitridation or carburization. High surface area of the catalysts were maintained during both treatments. These results indicated that Mo species on the surface of the Mo carbides and nitrides were dispersed on the surface alumina and more stable thermally. Furthermore, the particle size (d) of the unsupported Mo carbide crystal (100N7C7) was determined from the BET surface area (S, 88 m^g"^). d= 6/Sp Since particle density (p) is 9.4 //gm'^, d = 7.3 nm. Thus, the particle size of (fee) M02C crystal is 7.3 nm, which is between 7 and 9 nm, in accordance with previously reported results (9). 3.4. TPR-Confirmation of Mo carbides and carbon species formed by carburization TPR profiles of the alumina-supported molybdenum catalysts are shown in Figure 6. The main gases formed were CH4 and N2, but other gases containing carbon, such as CO and CO2 could not be confirmed. Therefore, the carbidic and
222
292
Binding energy [eV]
280
Figure 5. M03, and C,^ spectra of the carburized M0/AI2O3. (a) 100%Mo,C,(b) 100N7C7, (c) 100% Mo,C + A I A , (d) 12N7C7 before reaction (e) 12N7C7 after reaction, (f) 12R4
Table 2 BET surface area and Mo/Al ratio of catalysts Catalyst Surface area Mc^/Aljp ratio " (mVg) (-) Fresh 245 0.36 12R4 259 0.40 12N5 269 0.37 12N7 195 0.37 12C7 213 0.33 ^' The value determined by XPS analysis
260 polymeric carbon formed through carburization reacts with hydrogen to become CH^. By examining the CH4 formation temperatures during TPR, it is possible to evaluate the carbidic species formed during carburization. For the 12N7C5 catalyst, N2 gas desorbed at 780**C but no methane peak was observed. Because this profile of Nj desorption was similar to that of the Mo nitride (12N7), the Mo nitride was not transformed to Mo carbide below 500'*C. For the 12N7C7 catalyst, the peak of N2 desorption was observed up to 780*'C while the desorption peak of methane at 670T with a small peak at 270'C was observed. The peak at 270*'C is hkely to be assigned to methane adsorbed on the surface of AI2O3 because methane desorption at 270'*C was not observed for the unsupported Mo carbide (100N7C7). The CH4 desorption for the 100N7C7 catalyst showed the peak at 494*'C with a broad peak above 600**C, as shown in Figure 6. In Table 3, the C/Mo ratio decreased to 0.44 from 0.49 when the catalyst was heated to 600'*C during TPR. Lee et al. (9) reported that methane gas was desorbed at 462°C during TPR for unsupported MOjC without carbon deposition on the surface. Furthermore, they pointed out that the desorption of CH4 at 647''C was due to the decomposition of accumulated carbon with the release of methane gas. Therefore, the desorption peaks of methane at 494'C and 670"C are attributed to the decomposition of Mo carbide to form methane and high molecular weight carbon compounds such as graphite, respectively. For 12N7C9 carburized at 900T, a large peak of methane desorption existed at 870T, indicating a shift towards higher temperatures when comparing this peak to that formed at 690'C for the 700Tcarburized sample (12N7C7). This result revealed that the polymeric carbon probably covered the surface of the Mo carbides. If Tj-MOgCj was formed on the surface of the 12.5% M0/AI2O3 as the
0.002
(a) 12N7C5
N,
0.0015 0.001
cii.
200
0
(b) 12N7C7
400
600
Temperature ['CJ
800
200 400 600 800 1000 Temperature f C ] Figure 6. CH^ desorption profiles during TPR in hydrogen over carbide catalysts carburized at 700"C. (a) 12N7C5, (b) 12N7C7, (c) 12N7C9, (d) 100N7C7, (e) 12C7
261 same manner as that of the 100% M0O3 in the carburization in flowing of 20% CH4/H2, ii-MOgCj is underneath a layer of polymeric carbon on the 12N7C9. Since Al^Og was also covered with polymeric carbon, the desorption of metiiane was not observed at 270^. The true activity of the catalyst cannot be determined prior to treatment Furthermore, the desorption peaks of methane at 12C7 were shifted to higher temperatures than those for 12N7C7. Nitriding the unsupported catalyst has been shown to help reduce the production of polymeric carbon in a previous paper (9). From these results, the Mo nitrides were transformed to Mo carbides on the surface for the alumina-supported Mo catalysts above 700**C. However, above 900*'C, polymeric carbon was deposited on the surface of the Mo carbides during carburization treatment of the Mo nitrides.
4
6
Time on stream [hr]
Figure 7. Effect of carburization temperature of pre-nitrided M0/AI2O3 catalysts on carbazole conversion at 300X and 10.1 MPa total pressure.
(•) 12N7C5, (O) 12N7C7, (D) 12N7C9 Table 3 Carbon contents of carburized samples Surface Before TPR After TPR Sample C/Mo area (m^g ') C/Mo 0.02 0.02 N.D.^ 100N7C5 0.44 0.49 88 fee 100N7C7 4.52 N.D.'^ 5.10 100N7C9 0.52 7 0.49 hcpl00C7 0.08 0.10 197 12N7C5 0.37 0.10 12N7C7 230 8.66 9.11 12N7C9 198 12C7 0.13 0.48 213 ' not detected.
3.5. HDN activity and Mo species The change in HDN activity for carbazole HDN with carburizing temperature of the catalysts is shown in Figure 7. The 12N7C7 catalyst was 2.2 times more active than the 12N7C5 and 12N7C9 catalysts. The XPS and TPR studies with complement of XRD analysis of unsupported Mo carbides showed that a-MOjC is formed in 12N7C7, and that 12N7C5 is mainly composed from Mo nitrides. Because 12N7C7 showed the highest activity, it has been shown that Mo carbides have higher activity than do Mo nitrides. Comparing the maximum activities of the a-Mo2C carburized at 700"C and the rj-MOgCj carburized at 900"C, it was foimd that the former had about twice greater activity than the latter. a-Mo^C was more active than ri-MogCj for the HDN of carbazole, although polymeric carbon accumulated on the surface of 12N7C9 causes lower activity tiian 12N7C7. No deactivation was observed more tiian 10 h during tiie HDN reaction at 300**C. Thus, the catalytic activity for tiie HDN activity of carbazole decreased in the following sequence: 12N7C7 > 12N7 > 12N7C5 ^ 12R4 ^ 12N7C9 as shown in Figures 1 and 7. 3.6. Surface Model of Mo Carbide Figure 8 depicts tiie active sites on tiie surface Mo carbide as evaluated from the differences in activity resititing from differing amounts of carbon. Half of the tetrahedral structure of Mo metal in a-MOjC is filled witii carbon atoms, while a third of the structure for T1-MO3C2 is
262 composed of carbon. Since a-MOjC shows greater activity than does T1-MO3C2 on AI2O3 it can be deduced that the presence of carbon depleted sites contributes to the activity of the catalyst. 4. CONCLUSIONS (1) The catalytic activities for the HDN of carbazole decreased in the following sequence.
12N7C7 > 12N7 > 12N7C5 ^ 12R4 ^ 12N7C9
Carburization enhanced the activity of the M0/AI2O3 catalyst for the HDN of carbazole, compared with the reducing and nitriding. The catalyst carburized at 700**C was 2.2 times more active than was the catalyst carburized at 500T in the HDN of carbazole. (2) The TPR analysis show that neither transformation of the Mo nitrides to the Mo carbides nor free-carbon deposition occurred below SOO^C for the dumina• vacancy supported catalysts, although above # carbon atom 900*'C large amount of free carbon is deposited on the surface of the Mo Figure 8. The structure of Mo carbide carbides catalysts. catalysts. (3) From elemental analysis, the 100N7C7 was composed two phases: Mo nitrides and Mo carbides. (4) For the unsupported Mo catalysts, the XRD analysis showed that MOjN was converted to a-Mo2C at 700*C, and above 900**C the a-MOjC crystal was transformed to ri-MOgCj crystal. REFERENCES 1. M. J. Ledoux, C. Pham-Huu, J. Guille, and H. Dunlop, J. Catal., 134 (1992) 383. 2. M J. Ledoux, C. Pham-Huu, H. Dunlop, and J. Guille, "Proceedings, 10th International Congress on Catalysis, Budapest, 1992" (L. Guczi, F. Solymosi, and P. Terenyi, Eds.), p. 955. Elsevier, 1993. 3. C. Pham-Huu, M. J. Ledoux, and J. Guille, J. Catal., 143 (1993) 249. 4. J.-G. Choi, J. R. Brenner, and L. T. Thompson, J. Catal., 154 (1995) 33. 5. J. C. Schlatter, S. T. Oyama, J. E. Metcalfe, III, and J. M. Lambert, Jr., Ind. Eng. Chem. Res. 27 (1988) 1648. 6. J. S. Lee, M. H. Yeom, K. Y. Park, I. Nam, J.S. Chung, Y. G. Kim, and S. H. Moon, J. Catal., 128 (1991) 126. 7. M. Nagai, T. Miyao, and S. Omi, "Hydrotreating Technology for Pollution Control" (M. L. Occdli and R. Chianelli, Eds.), Chap. 18, Marcel Dekker, New York, 1996. 8. M. Nagai, T. Masunaga, and N. Hanaoka, Energy and Fuels, 2 (1988) 645. 9. J. S. Lee, S. T. Oyama, and M. Boudart, J. Catal., 106 (1987) 125.
® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
263
The design of base metal catalysts for hydrotreating reactions; Temperature programmed sulphidation of NiW/Al203 catalysts and their activity in the hydrodesulphurisation of thiophene and dibenzothiophene' H.R. Reinhoudf, A.D. van Langeveld', R. Mariscar, V.H.J. de Beer^ J. A.R. van Veen\ S.T. Sie' and J.A. Moulijn". ' Delft University of Technology, 2628 BL Delft, The Netherlands ^ Eindhoven University of Technology, 5600 MB Eindhoven, The Netherlands ABSTRACT NiW based hydrotreating catalysts have a good performance in hydrodesulphurisation reactions. It appeared that the their performance in the hydrodesulphurisation of dibenzothiophene and thiophene strongly depends on their sulphiding degree, which can be controlled by both the calcining and sulphiding temperature. By combining temperature programmed sulphiding with quasi in-situ XPS and activity measurements, it was concluded that the active phase for the HDS of DBT consist of either micro-crystalline sulphided M on a WO3 substrate, or of Ni^^ dissolved in the in NiW04. In contrast, the active phase for the HDS of thiophene seems to consist of Ni-promoted WS2 structures. 1. INTRODUCTION Hydrotreating processes play a central role in refineries, amongst others to upgrade transport fiiels. A continuous growth in the demand for transport fliels and changes in the relative importance of different crude oil fractions have put more pressure on the performance of conventional hydrotreating catalysts. More active catalysts with higher selectivity for specific hydrotreating reactions are needed to meet legislation for fuel quality in an economically attractive way. For example, in western Europe nowadays an important part of automotive transport is powered by Diesel engines, thus making dieselfiielproduction a growing segment in oil refining. However, the most important drawback of Diesel engines compared to emission controlled Otto engines is the emission of particulates which are suspected to cause serious health problems. Although the debate on the effects of the various diesel fuel components on particulate formation is still going on, aromatics and sulphur content (which are supposed to play a role in the particulate formation) are currently restricted by legislation. Apartfromthe fact that reaching low sulphur levels is a problem of its own, the presence of small amounts of sulphur also hampers the de-aromatisation of diesel fiiel. Therefore, deep hydrodesulphurisation (HDS) of dieselfiielis a key process in the upgrading of diesel fuel properties. The sulphur components which remain in diesel fiiel after conventional HDS are maidy dibenzothiophenes (DBT). It has been well established [1] that especially DBT's with alkyl substituents on the 4 and 6 position have a low reactivity over conventional hydrotreating
264 catalysts like C0M0/Y-AI2O3. This is due to the steric hindrance of the alkyl groups which block direct hydrogenolysis. It has been demonstrated that hydrogenation of one of the benzene rings of DBT lifts the steric hindrance [2], resulting in a higher reactivity in accord with molecular modelling. This observation implies the search for new catalysts with a high activity for reactions like isomeriation and hydrogenation which lift steric hindrance. Recent work on the development of catalysts for deep hydrodesulphurisation [3] revealed that NiW/y-Al203 catalysts are very promising for this application. Despite the importance of hydrotreating catalysts, the production of these catalysts is still mainly based on experience and empirical knowledge. Considering the increasing relevance for dedicated hydrotreating catalysts, the need for a thorough description and understanding of preparationpretreatment-activity relationships is clear. In the past a lot of work was focused on CoMo/y-Al203 and NiMo/y-Al203 catalysts. However, with the expected change to more specific hydrotreating reactions, and possibly also different reaction conditions, other catalysts than CoMo/y-Al203 and NiMo/y-Al203 might be more suitable. In the light of the promising activity in the deep HDS of gasoil, it was concluded to investigate the genesis of the active phase in NiW/y-Al203 catalysts in a detailed and systematic way. Scheffer et al. [4] have shown that especially NiW/y-Al203 is an interesting sytem, since it allows to steer the formation of different phases and morphologies by applying different pretreatment conditions. In this paper we will demonstrate the importance of pretreatment conditions on the activity of NiW/y-Al203 catalysts for different model reactions. The differently pretreated catalysts were characterised with Temperature Programmed Sulphiding (TPS) and quasi insitu X-ray Photoelectron Spectroscopy (XPS). The final aim of the work is the assessment of catalyst design rules for the NiW catalysts based on the knowledge of the correlation between pretreatment conditions, development of the active phase and the catalytic performance in different hydrotreating reactions.
2. EXPERIMENTAL 2.1 Catalyst Preparation The NiW/Al203 catalyst was prepared by pore volume co-impregnation of y-Al203 (Ketjen 0001.5E CK300, high purity, SBEf=190 m^.g'\ pore volume 0.6 ml/g). The aqueous solution contained (NH4)6.Wi2039.xH20.(Aldrich) and Ni(N03)2.6H20 (Aldrich), resulting in a catalyst with 1.2 wt% Ni and 15.2 wt% W, corresponding to 2.5 W/nm^ and 0.6 Ni/nm^. Table 1 Pretreatments of the NiW based catalysts investigated Tsulph [K]
613 673 823
1
923
1
Tcalc [K]
393
673
* •
• • *
823 * * * *
923 * t
1 1
265 Next, the catalyst precursor was dried overnight in air at 393 K, followed by calcining at various temperatures during 1 hour. For the activity measurements, temperature programmed sulphidation and the XPS analysis, the extrudates were ground and sieved for the 125 and 250 jam particle size fraction. The various catalyst investigated are collected in Table 1. 2.2 Catalyst characterisation Temperature programmed sulphiding was performed in an atmospheric plug flow reactor, more details on the equipment can be found in [5]. About 100 mg of the catalyst was diluted with a same amount of SiC. After purging at room temperature with Ar, the catalyst was exposed to the sulphiding mixture containing H2S, H2 and Ar (3, 25 and 72 vol %, respectively) at a totalflowrate of 33 [j,mol/s. After 30 minutes the temperature program, with a linear heating rate of 0.167 K/s, was started. Upon reaching the highest temperature of sulphiding, the sample was cooled in the sulphiding mixture. The signals of all TPS profiles have been normalised to the amount of catalyst. The sulphiding degree of the samples is based on a quantitative sulphiding of the nickeloxide into K13S2, and of the tungstenoxide into WS2. Clearly, the sulphiding degreefromthe TPS also includes chemisorbed S on the active phase of the catalyst, which is not taken into account m the reference point for complete sulphiding. For subsequent XPS analysis, the samples were purged with Ar at room temperature in order to remove residual traces of H2S. Then, the reactor was closed, disconnected from the TPS equipment and transferred into a glove box where the catalyst sample was transferred into the quasi insitu transfer facility for the XPS machine. XPS analysis was performed on a Perkin-Elmer PHI 5400 ESCA system equipped with a hemispherical analyser. Sample excitation was done by AlKa X-rays (1486.6 eV). The catalyst was pressed into an In foil attached to the sample holder under protective conditions in a glove box. Partial pressures of oxygen and water are lower than 0.5 . 10"^ mbar, typically. Transfer of the samples form the glove box into the XPS machine was performed by use of a commercially available transfer chamber. Peak shifts due to charging of the samples was corrected for by taking the Al 2p line of the AI2O3 at 74.2 eV as reference [6]. 2.2 Catalyst performance testing The batch autoclave reactor set-up used in the experiments for the DBT hydrodesulphurisation was described elsewhere [5]. About 200 mg of the catalysts was sulphided quasi insitu in an integrated reactor with 15% H2S in H2 with a flow of 40 [xmol/s at 1.2 MPa. The reactor was heated at 0.167 K/s up to the highest sulpiding temperature and kept isothermal for 1 h. Upon cooling down in the sulphiding mixture, the sulphided catalyst is transferred into the batch autoclave, where is submerged in the model feed, thus avoiding exposure to air. The model feed,which consisted of 2 g dibenzothiophene (Aldrich, 98%) in 100 g n-hexadecane (Aldrich, 99%+), was used to test the activity of the catalysts at 633 K and 100.0 MPa. No initial H2S was added to the reactor, the H2S/H2 ratio at 50% conversion was 2 mol%. Thiophene HDS was performed at atmospheric pressure in a flow reactor. Sulphidation of the catalyst in a mixture of H2S and H2 (50 vol% of both gases) at a total flow rate of 16 lamol/s. The catalyst was heated at 0.167 K/s up to 543 K, kept isothermal at this temperature during 30 minutes, followed by heating at 0.167 K/s up to the desired sulphiding temperature, i.e. 673 K, 823 K or 923 K, where it was kept isothermal during 2 hours. The
266 total flow was 39.5 famol/s, the thiophene content being about 6%. Thiophene conversions were determined at 623 K after 4 hours stabilising the catalysts. From the conversion the first order rate constant for the HDS of thiophene was evaluated. 3. RESULTS 3.1. Temperature programmed sulphidation In Figure la and lb the temperature programmed sulphidation profiles are shown for the catalysts calcined at 673 K and 823 K, respectively. For both catalysts the isothermal room temperature uptake is not shown in thefigures,however, the amount has been taken into account for the quantification of the sulphiding degree. Directly after the start of the heating program, a small amount of H2S is produced without accompanying H2 uptake, followed by a H2S uptake which starting at about 375 K. For the catalyst calcined at 673 K, this H2S uptake is accelerated above 490 K, while simultaneously a minor H2 consumption can be seen. Next, a sharp increase in the H2S concentration is observed which effectively results in a production maximum at about 615 K. Simultaneously, a similar H2 uptake occurs.
H,S
H2 T — I — I — 1 — I — I — r
300
500
700
900
300
1 — I — \ — I — r
500
700
900
-^T[K] Figure 1. Temperature programmed sulphidation profiles of the catalyst calcined at 673 K (left) and 823 K (right). Signals are normalised in micromol of H2S and H2 per 100 mg of catalyst. Note that a negative deflection of the upper profile corresponds to a H2S consumption, whereas for H2 (lower profile) a positive deflection corresponds to a H2 consumption. Quantification of the integrated sulphur uptake at 615 K, yields a sulphidation degree of the active phase of 63 %. The H2S production peaking at 615 K is followed by a second uptake with a maximum consumption at about 700 K, followed by a slow progressive sulphiding of the sample up to 923 K. At the highest temperature (923 K) the sulphiding degree is 87 %. For the catalyst calcined at 823, the H2S and H2 profiles exhibit two major uptakes. However, the amounts of H2S and H2 consumed are much less than in case of the catalyst calcined at 673 K. At 615 K, the total sulphur uptake corresponds to a sulphiding degree of 32 %, while at 923 K the total sulphiding degree is 64 %. Quite remarkably, the H2S production and H2 consumption peaking at 615 K are virtually absent.
267 3.2. Quasi in-situ XPS of sulphided catalysts The XPS spectra of the Ni 2p and the W 4f emission line regions of the catalyst calcined at 823 K in the various stages of sulphidation are collected in Figure 2. Note, that the reference spectra of the oxidic precursors are not shown. The quantitative data, that is, the peak position of the oxidic and sulphidic contribution of the Ni 2p3/2 and W 4fia emission lines and sulphiding degree of both elements upon the various sulphiding steps are collected in Table 2. Upon isothermal sulphiding of the catalyst at 293 K, no significant shift of the emission lines of Ni and W could be observed. However, the Ni 2p3/2 line broadens by about 10% at half height. The peak position of the oxidic M 2p3/2 was nearly constant at 856.6-856.7 eV, for the sulphidic contribution, the peak position was found at 853.6-853.8 eV. At 540 K, 39 % of the Ni was sulphided. At 613 K, the amount of sulphided Ni was increased up to 49 - 58 %, dependant on the time of isothermal sulphiding. After sulphiding at 823 K the amount of sulphidic Ni was about 70 %, and almost independent of the time of isothermal sulphiding, leaving 29 % of the Ni in the oxidic state at the highest temperature of sulphiding. The peak position of the oxidic W 4f^a line was found at 34.4 - 34.8 eV, while that of the sulphidic contribution was found at 31.6 - 32 eV. Sulphiding of the W (9 %) could only be observed at 540 K. After sulphidation at 613 K, the relative amount of sulphidic W increased up to 21 % after 2 hours isothermal sulphiding. Pronounced sulphiding of the W up to 75 % only occurred at 823 K. Note, that about 25 % of the W is not sulphided at the most severe sulphiding conditions. Table 2 Quantification of the XPS spectra shown in Figure 2. 1
A sulph
(K)
Ni 2p3/2 (eV)
tsulph
(min)
W4f7/2
(eV)
1
oxid
sulph
% Wsulph 1
oxid
sulph
0
856.7
-
0
-
-
30
856.2
853.2
3
35.7
-
1
856.5
853.8
39
35.4
-
1 613
5
1 856.4
853.6
49
35.5
32.2
1 613 1 613
30
856.6
853.6
49
35.3
31.8
120
1 856.6
853.6
58
35.4
31.9
60
856.6
853.9
70
35.7
120
856.9
853.9
71
35.6
-
1 298
1 ^"^^ 1 ^23 1 823
% Nisuiph
1
0 0 0 9
1 1 1 1
32.2
11 21 74
1 1 1
32.2
75
1
268 1
1
1
1
1
1
'^' ^ P
1
\
1 1—
854.0 eV
—1—I—I—I—r
W4f
3/2: s
5/:
1/2
32.0 eV 7/2
. y ^ ^ (0 c (D
a
J
1
1
880
1
1
L
I
860
'
'
J
-^ BE [eV]
40
1
1
L
30
Figure 2. Quasi in-situ XPS spectra of the Ni 2p (left) and the W 4f lines of the catalyst calcined at 823 K in its various stages of sulphidation, a) 30 min. at 298 K, b) 1 min. at 540 K, c), d) and e) 5, 30 and 120 min, respectively at 613 K, f) and g) 60 and 120 min. at 773 K, respectively. 3.2. Catalyst activity in the HDS of thiophene and dibenzothiophene Figure 3 gives thefirstorder reaction rate constant for the thiophene hydrodesulphurisation over the various catalysts. Clearly, a progressively increasing activity is observed at higher sulphidation temperatures, both for the catalyst sulphided at 623 K and 823 K. Note, that under mild sulphiding conditions the catalysts calcined at 673 K seems to perform better in the thiophene HDS. However, this difference levels off at a sulphiding temperature of 823 K
269
673 5.7
r
2.8
613
5.3
3.4
673
823
613
923
673
823
923
Tsulf [ K ]
Figure 3. The reaction rate constants for thiophene HDS of the various catalysts calcined at 673 K and 823 K, followed by sulphiding at 613 K, 673 K, 823 K and 923 K. Thefirstorder reaction rate constant for the hydrodesulphurisation of dibenzothiophene over the various catalysts is shown in Figure 4. For both catalysts calcined at 673 K and 823 K the same trend, i.e., a decreasing DBT HDS activity is observed at increasing sulphidation temperatures. The highest activity for the HDS of DBT is found upon 'low temperature sulphiding' of the catalyst calcined at 823 K. Note, that at the highest sulphiding temperature both catalysts have the same performance in the conversion of dibenzothiophene. Quite remarkably, the catalyst calcined at 673 K already reaches the low level of activity upon sulphidation at 823 K, whereas the catalyst calcined at 823 K still has an enhanced activity for the HDS of dibenzothiophene. 12 £.
'
673
"
12 ] 1 0.5
r 8f
8 4
6.1
c(D
5.5
•
3.0
3.3
^
^
% o
4
823
4.7
J.
B 613
673
823
923
(0
613
673
823
n
3.4
1
923
-^ T^uif [ K ] Figure 4. The reaction rate constants for dibenzothiophene HDS of the various catalysts calcined at 673 K and 823 K, followed by sulphiding at 613 K, 673 K, 823 K and 923 K. 4. DISCUSSION Upon sulphiding the catalyst, three different regimes can be discerned in the TPS profiles of the catalyst calcined at 673 K. In the first H2S uptake, mainly micro-crystalline [4, 8] Ni is sulphided. However, because of the high sulphiding degree also at least part of the W must have exchanged its oxygen by sulphur. This is supported by Raman spectroscopic analysis of the catalyst sulphided under identical conditions which clearly identifies the presence of WS vibrations [9]. The nature of the sharp H2S production at 613 is notfijllyclear at present, the shape of the peak suggests the process to be kinetically limited rather than thermodynamically determined. It could
270 possibly be caused by the nickel catalysed reduction of a tungsten oxysulphide, analogous to C0M0/AI2O3 [7], although the hydrogenation of excess sulphur could also explain the phenomenon. In the second stage, starting at about 615 K, the tungsten is progressively sulphided under simultaneous reductionfromW^^ into W^^ [4]. Following calcination at 823 K, the sulphiding degree of the active phase is strongly reduced (32 % at 613 K) as compared to the catalyst calcined at 673 K (63 % at 613 K). This is also reflected in the XPS spectra of the Ni 2p emission line. At 613 K about 60 % of the Ni is sulphided after calcining at 673 K, while upon calcining at 823 K only 40 % on the Ni can be sulphided under mild conditions. High resolution electron microscopy on catalysts sulphided at 673 K shows the presence of species with dimensions in the range of 0.4 - 0.6 nm [lOJ.Possibly these species represent sulphided micro-crystalline nickel. Clearly, the amount of micro-crystalline nickeloxide is decreased in catalysts calcined at higher temperatures. As can be inferred from Table 2, about 20 % of the Ni is slowly sulphided at 613 K in the isothermal stage. Following Scheffer et al. [4], in this temperature regime both surface nickelaluminates and nickeltungstate are sulphided. Since the presence of nickelaluminate surface compounds can be excluded by FTIR of adsorbed CO [11], the nickel must be stabilised in a tungstate. This corroborates the simultaneous slow sulphiding and reduction of the tungsten from 9 % up to 21 % at 613 K, as can be observed from the XPS data. At 823 K 70 % of the nickel is in the sulphided state, leading to the conclusion that in the temperature interval between 540 K and 823 K about 30 % of the nickel has been sulphided. The binding energy of the oxidic contribution of the Ni 2p3/2 line in this temperature regime is found at 856.4-856.7 eV, being 1 eV lower than the value reported by Ng and Hercules [6]. At 823 about 30 % of the nickel could not be sulphided, which can be attributed to bulk nickelaluminate [4] formed during the calcining at 823 K. This is also supported by the Ni Ip^a line position of this residual oxidic nickel (857.0 eV), which nicely corresponds to the value (857.2 eV) reported by Ng and Hercules for NiAl204 [6]. The oxidic contribution of the W 4f7/2 emission line in all catalysts is found at higher values (35.4 - 35.8 eV) than those reported by Ng and Hercules [6]. This may be due to the polarisation induced by the alumina as a result of the strong interaction between the active phase and the support [12]. The values of the sulphidic W Aa/i contribution are shifted slightly to a higher binding energy (by 0.2- 0.5 eV) as compared to that reported by Ng and Hercules. Possibly this is also due to polarisation by the support. The binding energy of the refractory tungsten, which can not be sulphided at 823 K reasonably corresponds to that reported by Ng and Hercules [4] for Al2(W04)3 To summarise, of the catalyst calcined at 823 K about 40 % of the Ni consists of microcrystalline nickeloxide which can be sulphided up to 540 K, about 30 % is stabilised as nickeltungstate which can be sulphided in the temperature range of 540 K to 823 K, while 30 % of the nickel is bound in bulk nickelaluminate which does not sulphide up to 823 K. Returning to the activity patterns for the hydrodesulphurisation of thiophene in Figure 3 it is clear that a thorough sulphiding of the catalyst is beneficial for an effective desulphurisation of thiophene. Both higher sulphiding temperatures and lower calcination temperatures result in higher sulphiding degrees as can be inferred from the TPS and the XPS data. Obviously, the catalyst for thiophene HDS should preferably be an analogue of the so-called 'CoMoS' phase [13,14]. In contrast, for the HDS of dibenzothiophene sulphiding of the catalyst should be rather limited. A low sulphiding temperature combined with a calcination temperature of about 823 K applied to the NiW based catalyst are beneficial for a high activity in this reaction. In the present stage it is not M y clear if the active phase consists of either micro-crystalline nickelsulphide, or of nickel enclosed in the nickeltungstate. In case the micro-crystalline nickelsulphide is the active phase, it must be
271 supported by a the WO3 layer on the alumina, since nickelsulphide on alumina does exhibit only a very low activity in hydrodesulphurisation reactions. Whatever the case may be, sulphiding of the surface of the tungstenoxide phase does not seem to play a key role in this respect. However, a strong decrease in the HDS activity for dibenzothiophene is observed as soon as the tungsten is being reduced to W^^. The reducibility of the tungsten can be solved at least partially by applying a sufficiently high calcination temperature. 5. CONCLUSIONS The present work clearly shows that catalyst performance testing and optimisation should be performed in using the reactants of relevance. Thiophene HDS should preferably be performed over NiW catalysts completely sulphided. Dibenzothiophene HDS, in contrast, should preferably be done over NiW catalysts sulphided under mild conidtions, thus avoiding the reduction of the W^^. ACKNOWLEDGEMENT This research was in part supported by the European Union (JOU2-0904) and the Netherlands Foundation for Applied Research (STW). Stimulating discussions with dr. W.R.A.M. Robinson (Eindhoven University), dr. D. Chadwick, (Imperial College, London), prof dr. H. Schultz (Universitat Karlsruhe), dr. P.L. Arias, dr. J. Cambra and prof dr. J.A. Legaretta (Escuela de Ingenierios, Bilbao) and prof dr. J.L.G. Fierro (C.S.I.C, Madrid) are gratefully acknowledged. REFERENCES 1) A. Amorelli, Y.D.Amos, C.P. Halsig, J.J. Kosman, R.J. Jonker, M. de Wind and J. Vrieling Hydrocarbon Process., Int.Ed. 71 (1992) 93. 2) T. [Kabe], A. Ishihara and Q. Zang, Appl. Catal.97 (1993) LI. 3) H.R. Reinhoudt, R. Troost, S. van Schalkwijk, A.D. van Langeveld, S.T. Sie, H. Schulz, D. Chadwick, J. Cambra, V.H.J, de Beer, J.A.R. van Veen, J.L.G Fierro and J.A. Moulijn, Proc. 1^ Int. Symp.on Hydrotreatment and Hydrocracking of Oil Fractions, Oostende, Feb. 1997, eds. B. Delmon and G. Froment, Elsevier Amsterdam, 1997 . 4) B. Scheflfer, P.J. Mangnus and J.A. MouHjn, J. Catal. 121 (1990) 18. 5 J.P. Janssens, Ph.D. Thesis, Delft University of Technology (1996). 6)) K.T. Ng and D.M. Hercules, J.Phys.Chem.80 (1976) 2095. 7) P.J. Mangnus, E.K. Poels and J.A. Moulijn, Ind & Eng. Chem. Research 32 (1993) 1818. 8) C.H. Kim, W.L. Yoon, I.C. Lee, S.I. Woo, Appl. Catal. A144 (1996) 159. 9) H.R. Reinhoudt, E. Crezee, J A R Van Veen, AD. Van Langeveld, V.H.J. De Beer, T.S. Sie and J. .A Moulijn, to be published. 10) RM. Stockmann, H.R. Reinhoudt, RPrins, P.J. Kooyman, H.W. Zandbergen, A.D. van Langeveld and J.A. Moulijn, to be published 11) H.R.. Reinhoudt, L. Mulder, E. Crezee, J.A.R. Van Veen, AD. Van Langeveld, V.H.J. De Beer, T.S. Sie, J.L.G. Fierro and J.A. Moulijn, to be published. 12) B. Scheffer, P. Molhoek and J.A. Moulijn, Appl. Catal. 46 (1989) 11 13) H. Topsoe, B. Clausen and F.E. Massoth, 'Hydrotreating catalysis', in: 'Catalysis Science and Technology, Vol. 11, Editors J.R Anderson and M. Boudart, Springer Berlin (1996). 14) RPrins,V.H. J. De Beer and G. Somorjai ,Cat.Rev.-Sci.Eng 31 (1989) 1.
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® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
273
Surface Science Models of CoMoS Hydrodesulfurisation Catalysts A.M. de Jong, V.H.J, de Beer, J.A.R. van Veen and J.W. Niemantsverdriet Schuit Institute of Catalysis, Eindhoven University of Technology, 5600 MB Eindhoven, The Netherlands Characterization of supported catalysts with surface spectroscopic techniques is often limited due to restraints imposed by the support material. The use of flat conducting substrates as a model support offers a way to apply these techniques to their fiill potential. Such surface science models of silica and alumina supported CoMoS catalysts have been made by impregnating thin Si02 and AI2O3 films with a solution of nitrilotriacetic acid (NTA) complexes of cobalt and molybdenum. X-ray Photoelectron Spectroscopy (XPS) spectra indicate that the order in which cobalt and molybdenum transfer to the sulfided state is reversed with respect to oxidic Co and Mo systems prepared by conventional methods, implying that NTA complexation retards the sulfidation of cobalt to temperatures where M0S2 is already formed. Catalytic tests show that the CoMoS model catalysts exhibit activities for thiophene desulflirisation and product distributions similar to those of their high surface area counterparts. 1. INTRODUCTION The last decades have seen a tremendous evolution in surface sensitive methods which provide atomic level information on the physicochemical state of the surface. Application of these methods on catalysts is often seriously hindered by the presence of a porous, electrically insulating support, which causes serious loss of intensity and resolution in the electron and ion spectroscopies. This implies that much, in principle extremely relevant information on the surface state of a catalyst can not be obtained. A solution to this problem is the use of model catalysts, consisting of a conducting substrate with a thin Si02 or AI2O3 layer on top of which the active catalytic phase is deposited [1]. An example of a model catalyst is schematically depicted in Figure 1. It consists of a silicon wafer as a substrate, which is flat and electrically conducting, with a thin layer of the appropriate oxide on top of it. The catalytically active phase is deposited on this model support either by a reaction of a suitable metal complex with the hydroxyl groups of the oxide surface or by impregnation with a solution of the catalyst precursor. In order to make realistic model catalysts it is important to use the same preparation routes as used for porous supported catalysts. Since these model supports are flat, catalyst preparation by pore volume impregnation is not straightforward. However, impregnation can be mimicked by spincoating, according to a
274 active phase
Figure 1. Schematic representation of a model catalyst, consisting of a Si substrate with a thin oxide layer on top of which the catalytically active phase is deposited.
procedure described by Kuipers and coworkers [2]. The principle of spincoating and the resemblance with pore volume impregnation is schematically shown in Figure 2. First the model support is cleaned and then covered with a film of the solution to be spincoated. Afterwards the wafer is placed on the spincoater and spun at the required rotation speed. During spinning a part of the solution is slung from the wafer and the rest of the solvent evaporates, thus depositing the solute on the surface. In this paper we apply this approach to supported CoMoS hydrodesulfurisation (HDS)
^
Pore volume impregnation
•H^
• •
•
Spin-coating impregnation Figure 2. The principle of pore volume impregnation (top) compared with spincoating (bottom). The porous oxidic support is simulated by a flat conducting silicon wafer covered with a thin oxide layer on top. Pore volume impregnation is mimicked by spincoating the solution to be impregnated, which subsequently evaporates. When the solvent evaporates the catalytically active phase is deposited on the support. catalysts. Van Veen et al. [3] have demonstrated that CoMoS II (Tops0e's terminology [4]), the highly active cobalt promoted M0S2 in which cobalt is thought to decorate the edges of M0S2 slabs [5-9], can be synthesized by sulfiding cobalt and molybdenum complexes of
275 nitrilotriacetic acid (NTA). In order to check if the preparation of the CoMoS phase on these model catalysts was successful, we performed thiophene HDS experiments to determine the activity of the model catalysts. In the following we will show that the model catalysts prepared from cobalt and molybdenum NTA complexes yield significant activities comparable to those of practical high surface area CoMoS catalysts. 2. EXPERIMENTAL A silica model support is prepared by oxidizing a silicon wafer with (100) surface orientation and a diameter of 75 mm in air at 825 K for 24 h. After oxidation the wafer is cleaned in a solution of ammonia and hydrogen peroxide at 340 K for 10 min. The surface is rehydroxylated by boiling in water. Angle dependent XPS experiments done as explained in [10, 11] indicate that the Si02 layer is about 5 nm thick after oxidation. During cleaning some Si02 dissolves resulting in an eventual oxide layer thickness of about 3 nm. The alumina model support is prepared by evaporating aluminum oxide on a silicon wafer with a native oxide layer. Prior to evaporation the wafer is cleaned in an HF solution. The evaporated alumina layer is approximately 5 nm thick. The highly active CoMoS phase can be prepared by sulfiding NTA complexes of Co and Mo [3]. The active material is applied to the model supports by spincoating an aqueous solution of these complexes, with a Co/Mo ratio of 1/3, at 2800 rpm. A model describing the spincoating of dilute solutions [12] predicts that eventually afilmof solution of 1.7 jum in height forms on the model support. With a concentration of Mo in the solution of 0.0034 M, this results in a loading of 3.6 Mo atoms per nm^, which is the equivalent of a loading of 12 wt% Mo on a 250 m^/g silica support. The model catalysts are sulfided in a glass tube reactor, with a mixture of 10% H2S in H2 at a flow rate of 60 ml/min while the catalyst is heated at a rate of 2 K/min to the desired temperature, and kept there for half an hour. After sulfidation, the catalyst is cooled under helium flow to room temperature. Next the reactor is closed and introduced into a nitrogen-filled glove box,fromwhere the sample can be transferred to a vessel for transport under nitrogen to the XPS spectrometer. Each sulfidation experiment is done with a fresh piece of model catalyst. XPS is used to study the state of the CoMo model catalysts. XPS spectra are obtained with a VG Escalab 200 spectrometer equipped with an Al Ka source and a hemispherical analyzer equipped with a five channel detector. Measurements are done at 20 eV pass energy. Energy correction is performed by using the CIs peak at 284.7 eV as a reference. The model catalysts are tested in a thiophene HDS reaction to monitor the activity and product distribution. First the model catalysts are sulfided at 675 K for two hours as described above. Then a mixture of 4% thiophene in H2 is passed through the reactor at a rate of 50 ml/min at 675 K. After 3 minutes the reactor is closed and operated as a batch reactor. After the desired reaction time a sample is taken from the reactor with a valved syringe for GC analysis of the reaction products. When a sample is taken the reactor isflushedwith the thiophene/H2 mixture for 3 minutes and closed again for the next analysis. Blank runs of the empty reactor and bare silicon wafers are also performed.
276
245
240
235
230
225
220
Binding energy (eV)
Figure 3. Mo 3d XPS spectra of the CoMo/Al2O3/Si(100) model catalysts as a function of sulfidation temperature. 3. RESULTS AND DISCUSSION The Mo 3d XPS spectra of the alumina supported CoMo model catalysts for different sulfidation temperatures are shown in Figure 3. The spectra of the freshly prepared model catalysts reveal only Mo 3d peaks at a binding energy of 232.5 eV, corresponding to a Mo^^ species. At 350 K a shoulder develops at a binding energy of 230.6 eV, indicating that Mo is already partially reduced. At 450 K the majority of the Mo is transformed to a species with a binding energy of 229.2 eV, which indicates that Mo is mostly in an (oxy)sulfidic form. For a detailed description of the changes in Mo XPS spectra during sulfidation we refer to earlier XPS studies of MoOx/SiO2/Si(100) model catalysts [13, 14] and of crystalline M0O3 powders [15]. The important point to note from Figure 3 is that sulfidation of Mo takes place between 350 and 450 K, while at 400 K the larger part of the Mo has already been converted to the sulfided state. The XPS spectra of Figure 4 indicate that sulfidation of cobalt proceeds slower than that of molybdenum. The Co 2p spectra of the freshly impregnated catalyst as well of these after treatment in H2S/H2 at 300 - 400 K show the pattern characteristic of oxidic cobalt, with a main peak at 781.9 eV and a shake up feature at higher binding energies [16]. The spectra after sulfidation at 500 K and higher have the Co 2p at 779.0 eV and are fully characteristic of sulfided cobalt [17], while the catalyst sulfided at 450 K is in the transition between the two states. Thus, complete sulfidation of the cobalt in the NT A complex takes temperatures between 450 and 500 K. This is in remarkable contrast to measurements on single component cobalt oxide reference
277 A Sulf. 1 temp. (K)
Co2p
/ lWvv^75
d
^
AjJuikAA^
.. A \ 1
(0
c
CO Q.
^ 815
805
h795
^
785
\
775
765
Binding energy (eV)
Figure 4. Co 2p XPS spectra of the CoMo/Al2O3/Si(100) model catalysts as a function of sulfidation temperature. catalysts, where complete sulfidation is observed around 400 K already [9, 18], Unfortunately, Co 2p binding energies give little information on the precise state of the Co and do not allow for further distinction between different sulfidic surroundings of the cobalt [17]. XPS spectra of the SiO2/Si(100) supported model catalysts [19] gave similar results. Again sulfidation of molybdenum is completed before that of cobah. This is in agreement with the fact that these catalysts are prepared from Co and Mo NTA complexes, which are beUeved to have negligible interaction with the support [3], so there are no large effects from the support to be expected. Apparently, the sulfidation of Co is retarded by the NTA ligand to temperatures at which molybdenum is ah-eady in a fully sulfided state. Thesefindingsare in agreement with recent work of Medici and Prins [ 20, 21] on NTA derived NiMoS catalysts and fit well in the picture of CoMoS and NiMoS in which the promoter atoms decorate the edges of M0S2 particles [8, 22] Since the presence of the CoMoS phase cannot be proven definitively by XPS measurements, we have determined the catalytic activity of the model systems for thiophene desulfiirisation and compared it with the catalytic performance of their conventional counterparts, to obtain indirect evidence for the CoMoS phase. To this end, batch thiophene HDS activity tests have been performed with the model catalysts for different reaction times. The only products formed are butenes and C2 and C3 hydrocarbons. The total conversion of thiophene is plotted in Figure 5 for both the model catalysts and for a blank experiment. In the blank experiment, the reactor is loaded with an oxidized and cleaned silicon wafer; this is a model support that was treated in the
278 2,5
2,0
1
1 ^ CoMo/SiO 2/Si
1.5 •CoMo/AI jOg/Si 1,0 {
0,5
i •
1
1 •
i^—: 10
i
i •
•blank
...-^•--•"
J. 1—
20
'^'"
1
1
30
40
h-
50
1
1
1
1
60
70
80
90
1
100
time (min)
Figure 5. Conversion of thiophene in HDS batch reactions at 675 K over the thin film silica (triangles) and alumina (diamonds) supported model catalysts as a function of the batch reaction time. The dots represent conversion rates measured in the blank experiment. The silica and alumina supported model catalysts exhibit similar conversion rates. same way as the model supports used to prepare the model catalysts. Figure 5 reveals that in the blank experiment also some conversion of thiophene occurs. This is thought to be due to thermal decomposition of the thiophene, which may or may not be assisted by the reactor wall. However, the conversion of the Co and Mo loaded model catalysts is significantly higher, thus proving that these model catalysts are indeed active in HDS reactions. The loading of Mo (3.6x10^'^ atoms/cm^ and the total thiophene conversion (about 1.5% after 1 hour) result in an activity for thiophene desulfurization of 2.10"^ molecule s"^ per Mo atom at 675 K. This value agrees very well with activities reported on alumina and silica supported CoMoS catalysts [23, 24], which fall in the range 1.10'^ to 1.10"^ s'^ The fact that the thiophene HDS activity is similar for CoMoS on both silica and aluminafiillyagrees with the earlierfinding[3] that catalysts prepared via the NT A route exhibit minimal influence of the support on the activity. The product distributions for both the silica and alumina supported model catalysts after 1 hour of batch reaction are displayed in Figure 6. The product distribution is independent of the reaction time and is similar on both supports. The ratios of the formed 1-butene, and trans and cis 2-butene are in excellent agreement with those measured at low conversions on conventional, high surface area CoMoS catalysts prepared via the NTA route. Since n-butane is formed through
279
•CoMo/Si02/Si 0,8 DCoMo/Al203/Si g o O
0,6
0,4
0,2
UI C1
C2
C3
1-butene
n-butane
t-2-butene
c-2-butene
Figure 6. Product distribution after 1 hour of batch reaction for the siUca and alumina supported model catalysts. hydrogenation of butenes [25], only a small amount is formed at low conversions. Summarizing, the results of this study show that we can prepare model catalysts for hydrodesulfurisation reactions showing activities and product distributions that are in good agreement with activities and products yielded by these catalysts supported on porous silica and alumina carriers. These model catalysts have the advantage that surface sensitive spectroscopies are applicable to their full potential to study properties of the catalytic surface. XPS spectra of CoMoS model catalysts prepared from NTA complexes as a precursor reveal that molybdenum is the first to form sulfidic species upon treatment in H2/H2S, whereas the sulfidation of cobalt is selectively retarded. Apparently the role of the NTA-complex is here to form an environment for the cobalt that remains stable up to temperatures where M0S2 has already formed. ACKNOWLEDGEMENTS This work was supported by the Netherlands Foundation for Chemical Research (SON) with financial aid from the Netherlands Technology Foundation (STW). REFERENCES 1. P.L.J. Gunter, J.W. Niemantsverdriet, F.H. Ribeiro and G.A. Somorjai, Catal. Rev.-Sci.Eng. submitted. 2. E.W. Kuipers, C. Laszlo, W. Wieldraaijer, Catal. Lett., 17 (1993) 71.
280 3. J A.R. van Veen, E. Gerkema, A.M. van der Kraan, A. Knoester, J. Chem. Soc, Chem. Commun. 22 (1987) 1684. 4. R. Candia, O. S0rensen, J. Villadsen, N. Tops0e, B.S. Clausen, H.Topsoe, Bull. Soc. Chim. Belg. 93(1984)763. 5. R. Prins, V.H.J, de Beer, G.A. Somorjai, Catal. Rev. - Sci. Eng. 31 (1989) 1. 6. S.M.AM. Bouwens, J.AR. van Veen, D.C. Koningsberger, V.H.J, de Beer, R. Prins, J. Phys Chem. 95 (1991) 123. 7. S.M.AM. Bouwens, F.B.M. van Zon, MP. van Dijk, A.M. van der Kraan, V.H.J de Beer, JAR. van Veen, D.C. Koningsberger, J. Catal. 146 (1994) 375. 8. H. Tops0e, B.S. Clausen, Catal. Rev. - Sci. Eng. 26 (1984) 395. 9. M.W.J. Craje, V.H.J, de Beer, JAR. van Veen, A.M. van der Kraan, J. Catal. 143 (1993) 601. 10. A.M. de Jong, L.M. Eshelman, L.J. van IJzendoorn, J.W. Niemantsverdriet, Surf. Interf. Anal. 18(1992)412. 11. P.L.J. Gunter, A.M. de Jong, J.W. Niemantsverdriet, H.J.H. Rheiter, Surf. Interf Anal. 19 (1992) 161. 12. R.M. van Hardeveld, P.L.J. Gunter, L.J. van IJzendoorn, W. Wieldraaijer, E.W. Kuipers, J.W. Niemantsverdriet, Appl. Surf. Sci. 84 (1995) 339. 13. AM. de Jong, H.J. Borg, L.J. van IJzendoorn, V.G.F.M. Soudant, V.H.J, de Beer, JAR. van Veen, J.W. Niemantsverdriet, J. Phys. Chem. 97 (1993) 6477. 14. J.C. Muijsers, Th. Weber, R.M. van Hardeveld, H.W. Zandbergen, J.W. Niemantsverdriet, J. Catal. 157(1995)698. 15. Th. Weber, J.C. Muijsers, J.H.M.C. van Wolput, C.P.J. Verhagen, J.W. Niemantsverdriet, J. Phys. Chem. 100 (1996) 14144. 16. J.F. Moulder, W.F. Stickle, P.E. Sobol, K.D. Bomben, Handbook of XPS, Perkin Elmer Corporation, Eden Prairie MN, USA, 1992. 17. I. Alstrup, I. Chorkendorfif, R Candia, B.S. Clausen, H. Tops0e, J. Catal. 77 (1982) 397. 18. M.W.J. Craje, E. Gerkema, V.H.J, de Beer, A.M. van der Kraan, in M. Occelli, R. Anthony, eds. Advances in hydrotreating catalysts: preparation, characterization and performance, studies in surface science and catalysis, vol. 50; Elsevier: Amsterdam, 1989; p. 165. 19. AM. de Jong, V.H.J, de Beer, J.A.R. van Veen and J.W. Niemantsverdriet, J. Phys. Chem. Lett., accepted for publication. 20. L. Medici, R. Prins, presented at the 2""* European Congres on Catalysis, Europacat II, 1995, Book of abstracts, p. 45. 21. L. Medici, R. Prins, J. Catal. submitted. 22. JAR. van Veen, E. Gerkema. A.M. van der Kraan, P. A. J.M. Hendriks, H. Beens, J. Catal. 133(1992)112. 23. H. Tops0e, B.S. Clausen, F.E. Massoth, Hydrotreating catalysis. Springer-Verlag, Berlin 1996. 24. LA. Medici, Thesis: Influence of chelating ligands on the structure and activity of NiMo hydrotreating catalysts, ETH Zurich, Switzerland 1995. 25. C.N. Satterfield, GW. Roberts, AIChE J. 14 (1968) 159.
® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
281
Molecular mechanics modelling of the interactions between M0S2 layers and alumina or silica support Philippe Faye *'^, Edmond Payen * and Daniel Bougeard ^ * Laboratoire de Catalyse Heterogene et Homogene, U.R.A.-CNRS 402, Bat. C3, Universite des Sciences et Technologies de Lille, 59655 Villeneuve d'Ascq Cedex, France. ^ Laboratoire de Spectrochimie Infrarouge et Raman, U.P.R.-CNRS 2631, Bat. C5, Universite des Sciences et Technologies de Lille, 59655 Villeneuve d'Ascq Cedex, France. The ability of the support to stabilize M0S2 layers has a marked effect on the activity of the active phase of this hydrotreatment catalyst. Thus the non-bonded interactions of M0S2 slabs supported on alumina and silica have been modelled by molecular mechanics. Whatever the carrier surface, a single M0S2 slab lying flat on the support is more stable than a slab perpendicular to the support, but the energy difference between these two configurations decreases as the number of stacked M0S2 sheets increases. Whereas single sheets lying flat on the [111] plane of the y-Al203 are favoured, the [110] plane induces a stacking of several slabs. Whatever the exposed plane of the silica support successively modelled by the P-quartz and the high-cristobalite, the stacking of several M0S2 sheets is favoured and stacked M0S2 cluster in a parallel configuration is more stable on y-Al203 planes than on Si02 surfaces. The intercalation of one layer between a slab and the support or between two contiguous M0S2 slabs leads to a screen effect characterized by a decrease of about two orders of magnitude of the interaction energy of the separated surfaces. This effect increases with the number of inserted sheets. 1. EVTRODUCTION Hydrodesulfurisation (HDS) of petroleum feedstocks is one of the most important industrial processes worldwide, used in the removal of sulfur compounds from crude petroleum. Among several HDS catalysts, molybdenum disulfide (M0S2) associated with a promoter (Co or Ni) or single and dispersed over the surface of a carrier with high specific surface area ( gamma-alumina or silica) is widely used. In future further reduction in the emission of sulphur compounds will require more severe HDS of petroleum fractions. Thus modelling techniques, combining computer graphics and theoretical calculations can be very helpful to improve the understanding of the microscopic steps of the HDS reaction. As far as the hydroprocessing catalyst is concerned, it is obtained from a reductive sulfiding of an oxidic precursor.^' ^ The structure and the genesis of this oxidic precursor as well as the sulfidation process have been widely studied for two decades.^'^ Today the presence of nanocrystals of M0S2 in the sulfided form, which can be visualized by transmission electron microscopy (TEM)^'^^, is well established and it is now generally accepted that the
282
active phase for HDS consists of M0S2 slabs dispersed over a carrier surface. ^^' ^^ In a previous paper^^ the interaction of M0S2 layers with y-Al203 was investigated by means of molecular mechanics. This study led to the elucidation of some open questions concerning the interpretation of high-resolution electron microscopy (HREM) data.^' ^^ In the present work the interaction of M0S2 slabs supported on different silica surfaces is investigated in order to compare the ability of silica and alumina to stabilize the M0S2 active phase. All over this paper the edge M0S2 plane and the support surface are ideally modelled leaving only geometric factors to be analysed. These include the slab orientation on the silica carrier and the distances between the support and the M0S2 slabs. Increasing the interface spacing step by step, an M0S2 slab is tilted from a perpendicular configuration to a parallel configuration and the interaction energy is evaluated by molecular mechanics calculations. Reproducing such experiments with single, double and triple sheets, the influence of the stacking on different silica surfaces has been studied. The energy of the interactions between two contiguous layers or between a sheet and the support surfaces can be calculated from the most stable configuration. Thus the stacking or the dispersion of M0S2 as single sheets on different support surfaces could be deduced. 2. COMPUTER MODELLEVG The method of molecular mechanics currently employed in molecular modelling is a conceptually simple method which relies on embedded empirical parameters. Molecular mechanics calculations, less expensive in term of computational resources, allow the investigation of large systems such as those studied in the present work. The molecular mechanics results of the present work were generated using the program Cerius^ ^^ developed by Molecular Simulations Incorporated running on a Silicon Graphic Indigo workstation. 3. CRYSTAL STRUCTURES Hexagonal M0S2 (PGs/mmc) is the most common form with two units M0S2 in the hexagonal unit cell, the dimensions ao and Co are 3.16 A and 12.295 A, respectively.^^'^^ Bulk M0S2 has a layered structure in which the metal atom is surrounded by a trigonal prismatic array of sulfur atoms. These prisms are joined by their edges to form sheets which are held together by non-bonded forces. Molecular mechanics calculations have been previously carried out for M0S2 slabs truncated along the [1100] and the [21 10] plane, respectively and dispersed over alumina plane surfaces. Only small energy differences have been noticed between both configurations, thus the [1100] edge plane was chosen among the two stable surfaces in order to model the M0S2 slabs. Throughout this work the M0S2 layered cluster is modelled by the stacking of 19 A long by 13.7 A wide rectangular sheets, each of them containing 39 molybdenum atoms. Such a slab containing 117 atoms is of the size of the supported slabs studied by HREM. Notwithstanding the very simple chemical formula of silicon dioxide silica exists under several variety of forms. Each form of silica is characterized by a structure (crystalline or amorphous) and specific surface properties. ^^ Two silica structural forms were ideally modelled in this work, i.e. P-quartz and high-cristobalite. The former one crystallizes in a hexagonal unit
283 cell (P6222) where the dimensions Oo and Co are 5.01 and 5.47 A, respectively. The second one is built according to a square lattice (Fd-3m) with the parameter ao equal to 7.16 A.^^ Whatever the crystal structure of silica, the silicon atoms are tetrahedrally coordinated by oxygen anions. The support was modelled as a slice which consists of five contiguous planes. The surface plane composed of oxygen anions closely packed in rows is 47.4 A long by 33.5 A wide. Different surface planes can be expected, consequently the [1010], [1100] and [1101] surface planes of the P-quartz as well as the [100] and [111] exposed planes of the highcristobalite were investigated. With such dimensions the M0S2 basal area of each slab is smaller than the support surface on which it is deposited. In this way edge effects between catalytic phase and support were reduced and calculations with one to three M0S2 slabs supported on silica surfaces could be performed. 4. FORCE FIELD In molecular mechanics calculations a model structure and a force field are necessary to find the optimized geometry by minimizing the total energy. This energy is written as the sum of bond-stretching (Eb), angle-bending (Ea), torsional (Et) and non-bonding (Enb) interactions. Each interaction is described by a simple analytical expression. The conjugate gradient method was used as minimization algorithm to find the optimized geometry corresponding to the lowest total energy. The Spline switching method was applied to effectively minimize the boundary forces at the cutoff region. Table 1 Force constants used in the M0S2 force field Force Potential function Mo-S stretch
S-Mo-S bend
Eb = Vi kb (R-Ro)^
Ea = C [1 - B (-if cos N0]
Parameters kb = 1125.4 kJ mol'^ A"^ Ro = 2.418A B=l N=4 C = 99.1kJmor'rad'^
Mo-S-Mo bend
Ea = K (Co + Ci cos 9 + C2 cos 20)
Co = C2 (2 cos^ 0o +1) Ci = - 4 C2 cos Go C2 = 'A sin^ 0o K=1010.3kJmor^rad"^ 00=81.64°
Using previous calculations carried out by Mitchell et al.^^' ^^ a force field was developed for MoS2^^ which is summarized in table 1. As the structural constraints are adequately included in the stretching and bending terms, Mitchell et al.^^ assumed that torsional
284 forces are not necessary for molecular mechanics calculations on inorganic extended solids. Therefore in all calculations presented in this work, the torsional force constant was set to zero. Whatever the initial relative position of two rectangular layers, each of them containing 39 molybdenum atoms, the optimized configuration corresponds to parallel stacked layers (Fig. 1). The distance between two contiguous sheets does not differ by more than 2.4% fi-om the values in bulk M0S2. In the previous paper ^^ concerning the interaction between M0S2 and gamma-alumina, preUminary attempts using the Dreiding force field ^^ and allowing the relaxation of the y-Al203 support, showed only minor structural variations at the interface. The small amplitude of these changes is probably due to the possibility to relax the forces over the whole volume of the model support. Thus, in order to limit the use of computational resources and as the influence of the support on the disulfide active phase is the most interesting question, all atoms of the support structure were considered in fixed positions. Non-bonded forces were only applied between the exposed planes of both supports y-Al203 and Si02 and the active phase.
C
Figure 1. Bulk M0S2 layered structure represented by two rectangular slabs truncated along the [1100] edge plane, each of them containing 117 atoms, 39 of which are molybdenum atoms. Two potential fimctions are used to describe the long range electrostatic and the Van der Waals interactions treated as non-bonded forces (Enb). These potential fiinctions and their parameters are summarized in table 2 and table 3. Referring to quantum calculations on the hypothetical Mo-S molecule ^ the charge (Qi) expressed in electron units (e) was set to +0.18e on the molybdenum atoms; the electroneutrality of the M0S2 clusters imposes a charge of-0.09^ on the sulphur atoms. Partial charges on alumina and silica were obtained using the charge equilibration scheme (QEq).^"* They were averaged to +1.2^, +1.6^ and -O.Se on aluminium, silicon and oxygen atoms, respectively. The assignment of the Van der Waals diameters (pii) of each atom is made according to the results published on MoS2^^ and Dreiding force field parameters on the carrier.^^
285 Table 2 Non bonded interactions Forces Coulomb
Van der Waals
Potential function
-t^coul
^o
parameters Qi, Qj charges Rij distance between charges (A) dielectric constant 8 = 1 Co: conversion factor (IMTCSO)
eRij
Evaw ^ Sij [-2 (B^/p^ + (K^p^
]
the well depth e^ = -Jsisi Pij = y2 (pii + pjj)
Table 3 Atomic parameters used in the non-bonded potential functions 8i / kJ mol'^ Qi / electron units (e) atom Mo 0.844 0.18 1.877 S -0.09 1.296 Al 1.2 1.296 Si 1.6 0.401 0 -0.8
Pii/A 5.388 3.694 4.390 4.270 3.405
5. RESULTS AND DISCUSSION In a previous paper^^ the energy of an isolated M0S2 sheet which does not interact with a support surface plane has been evaluated as 711 kJ mor\ This value will be considered as a reference value throughout this study. The energy (E1.2) of the interactions between two contiguous sheets buih in the configuration of bulk M0S2 has been deduced to -927 kJ mol'V It has also been shown that the intercalation of several layers can be considered as a screen effect which increases with the number of inserted sheets. For gamma-alumina supported molybdenum disulfide the existence of covalent bonds between the carrier and the active phase have been suggested earlier.^^' ^^ However, previous molecular mechanics studies carried out on an M0S2 slab docked to the gamma-alumina ^^ have only led to qualitative indications in spite of the geometrical fit and the major simplifications made concerning the Al-O-Mo linkages. Another molecular modelling method like the density functional could be more relevant to the investigation of small M0S2 clusters docked to different support exposed planes. Nevertheless, the molecular mechanical study of non-bonded interactions between the M0S2 slabs, each of them containing 39 molybdenum atoms, and the support (alumina or silica) is the alternative approach. 5.1. M0S2 cluster supported on ^-quartz The HDS active phase is modelled as rectangular layers of M0S2 held together by nonbonded forces. Each sheet of this cluster lying on a silica support is truncated along the [1100] planes. The surface plane of silica is successively assumed to be the [1010], [1100] and [1101]
286 plane of P-quartz where oxygen anions are closely packed in rows. The parallel stacked M0S2 cluster is held to the preferentially exposed plane of P-quartz by non-bonded interactions. The geometrical parameters considered in this model calculation are presented in (Fig. 2). 19 A
<:><>6666^. 54.7 A
45.1 A
Figure 2.Triple-layered rectangular M0S2 cluster truncated along the [1100] edge plane and perpendicularly supported on the [1010] plane of P-quartz. The small circle localises the rotation axis parallel to the Y axis. The X distance perpendicular to the exposed plane of P-quartz is measured between the lowest molybdenum row and the highest surface oxygen plane. It can be considered as the interface measurement between the exposed plane of silica and the [1100] edge surface of M0S2. Both Y and Z slidings describe shifts parallel to the support surface and 9 is the angle between the silica exposed plane and the basal plane of a M0S2 sheet. As already observed for an M0S2 sheet oriented perpendicularly to the [110] Y-AI2O3 plane ^^ a small amplitude in energy is also monitored when a similar layer slides on P-quartz surfaces. A sinusoidal function with small amplitude compared to the influence of the interface parameter X is obtained upon variation of the Y or Z positions. The periodicity of this curve corresponds to the distances between anions rows which compose the exposed plane of the silica surface. Therefore a single M0S2 sheet oriented perpendicularly to the silica support has its lowest energy when it is located between two contiguous oxygen anions rows at an interface distance of 3.4 A (Fig. 3).
2.00 2.60 3.20 3.50 4.00 4.60 5.20 5.80 interface/A Figure 3. Influence on the energy of the interface spacing (X) of one M0S2 slab perpendicular to the [1010] plane of the P-quartz support.
287 Alternatively, a single M0S2 sheet set parallel to the P-quartz surface plane presents its lowest energy when the X distance is close to 4.4 A, which corresponds to a distance of 2.8 A between sulfur atoms of the M0S2 slab and the oxygen anions of the P-quartz surface. As a difference in energy appears between these two configurations, a single sheet was tilted step by step, from 90° to 0°. The study was carried out with a rigid slab in steps between two X values (X=3.4 and 5.4A) successively supported on three P-quartz surface planes ([1010], [1100] and [1101] ). A three4,2 interface/A dimensional graph can be sketched (Fig. 4) for each model and only one minimum is found on these angle/degrees representations (X=4.4A and 9=0°). Whatever the exposed plane of the 1 Si02 carrier surfaces considered, the gain in energy between the most stable perpendicular slab (X=3.4A and e=90°) and the most stable parallel Figure 4. Influence of the interface spacing and the sheet which corresponds to the global angle between the slab and the support surface on the minimum for one sheet supported on energy of one single M0S2 sheet supported on the silica surface is the same order of magnitude. This gain in energy is close [1010] plane of p-quartz to 420 kJ mor\ However, a difference of ca. 25 kJ mol'^ appears according to the surface planes of the P-quartz. As regards this result it is unlikely to find an M0S2 single sheet perpendicularly oriented on the [1010], [1100] or [1101] surface plane of the P-quartz. The total energy (E) of an M0S2 slab lying flat on the P-quartz support is written as a sum of two terms^^^ where Eo is the energy of one single sheet and Ei.sup represents the interaction between the M0S2 slab and the support surface. -t-'
i-
(1)
This calculated energy is of the order of magnitude of the E1.2 attractive energy beween two contiguous M0S2 layers, but depends on the exposed plane of the P-quartz surface. These differences are summarised in table 4 where the energy of the interactions of an M0S2 slab lying flat on a P-quartz surface can be noticed for each exposed plane of the support.
288 Table 4 Energies of interactions (Ei-sup) of a M0S2 slab on different planes of P-quartz Exposed plane of the P-quartz Ei.sup (kJ mol") -587.5 [1010] -537.2 [1100] -557.6 [1101] Whatever the exposed plane of P-quartz, an M0S2 sheet lying flat is stabilized by the support. Therefore the energy (Ei-sup) between the slab and the p-quartz exposed surface plane is higher than the energy which represents the non-bonded forces between two contiguous M0S2 sheets (E1.2 = -927 kJ mol"^). Consequently each surface plane of P-quartz is preferentially covered by stacked M0S2 layers lying flat on the support surface. Calculations identical to those previously reported for one layer, are extrapolated to a double-layered and a triple-layered cluster in order to estimate the interaction energy between the second stacked layer and the exposed surface plane of P-quartz. These two or three slabs each containing 39 molybdenum atoms are built according to the bulk M0S2 structure (Fig. 1), and the clusters are successively supported on the [1010], [1100] and [1101] surface plane of P-quartz. The two variables (9, X) which have been previously defined (Fig. 2) are used to sketch a threedimensional graph showing differences in energy (Fig. 5a, 5b).
interface/A
Figure 5a. Influence of the interface spacing and the angle between the slab and the support surface on the energy of a double M0S2 layered structure supported on the [1010] plane of P-quartz
The [1100] edge row containing only molybdenum atoms belonging to the sheet which is nearing the surface is considered as the rotation axis. Whatever the surface plane considered a similar local minimum of energy appears in a location of stacked layered cluster nearly perpendicular to the support surface, but two global minima are observed which correspond to stacked layers lying flat on the support. Although the gain in energy between the local minimum and the two global minima decreases with the number of stacked layers in the M0S2
289 cluster, a barrier has to be jumped. This series of one to three layers (Table 5) shows that the gain in energy (AE) between the most stable configurations decreases, whereas the energy barrier (Ea) increases, when the number of stacked layers increases. Consequently and as shown with identical M0S2 clusters supported on y-Al203 surfaces, stacked layers of M0S2 lying flat on a silica support is the most stable configuration but the existence of metastable nearly perpendiculariy oriented slabs cannot be excluded, particularly at low temperature. The energy (E) of a double layered cluster of M0S2 lying flat on the siUca exposed plane is evaluated as a sum of the interactions forces and the energy of a single M0S2 sheet ^^^ where Eo, E1.2 and Ei.sup have been eariier evaluated. E2.sup represents the interactions between the support surface and the second sheet. E — 2 Eo + E1.2 + El-sup + E2-SUP (2)
2
3,6 5,2 6,8 8,4 interface/A
10 11,6 angle/degree
Figure 5b. Influence of the interface spacing and the angle between the slab and the support surface on the energy of a triple M0S2 layered structure supported on the [1010] plane of P-quartz
The energy E2.sup is close to zero in agreement with our previous results about the intercalation screen effect. Whatever the nature of the interacting support (M0S2, Y-AI2O3 or Si02) the same screen effect is observed on the intercalation of one single sheet between two contiguous layers or between a sheet and the support surface. Moreover, this screen effect independent of the support surface increases with the number of inserted sheets, as expected. Table 5 Gain in energy and energy barrier for three M0S2 clusters supported on three different exposed planes of P-quartz Exposed Triple-layered Single-layered Double-layered plane of the AE Ea AE Ea AE Ea support (kJ mor^) (kJ mor^) (kJ mor^) (kJ mor^) (kJ mof^) (kJ mol"^) 446.0 0 323.8 77.7 317.6 82.9 [lOiO] 392.8 0 224.3 53.5 [1100] 183.46 87.4 436.5 0 273.2 60.6 266.7 68.6 liiPIL
290 5.2. M0S2 cluster supported on high-cristobalite All previous molecular calculations have been carried out for M0S2 slabs supported on different surface planes of P-quartz. Based on the interpretation of spectroscopic data some authors^^ assumed that the [111] or [100] exposed planes of the high-cristobalite could be found. Similar M0S2 clusters supported on an alternative silica model like high-cristobalite will now be considered. The number of molybdenum atoms of each M0S2 slab was kept constant at 39 and each slab was always truncated along the [1100] edge plane. The [100] and [111] surface plane of the high-cristobalite were ideally modelled by oxygen anions closely packed in rows and similar calculations were performed in order to estimate the influence of the silica structure. The most stable configuration is still the M0S2 single sheet parallel to the support at a distance of 2.8 A. Whatever the exposed surface plane of the silica models considered the study of a series of one to three M0S2 stacked layers shows that the gain in energy between the most stable configurations decreases, whereas the energy barrier increases, when the number of stacked layers increases. According to the most stable configuration which corresponds to an M0S2 layered cluster lying flat on a surface plane of the high-cristobalite, the energy Ei.sup can be deduced. This value representing the interaction of the M0S2 slab contiguous to the support surface is evaluated for each surface plane of the high-cristobalite, the results are gathered in table 6. In spite of the ability of the both model structures of silica to stabilize M0S2 layers lying flat on the surface planes, it appears that the stacking of several sheets is more favoured on the highcristobalite surface planes than on the P-quartz exposed planes. Table 6 Energies of interactions (Ei-sup) of an M0S2 slab on high-cristobalite Exposed plane of high-cristobalite Ei.sup (kJ mol" ) [100] -509.11 [111] -454.6 5.3. Comparison of Si02 and Y-AI2O3 support Using mechanics calculations results which have been carried out for M0S2 clusters supported on y-Al203 planes ^^ and the present results obtained for similar active phase clusters supported on silica, a comparison can be established. Whatever the nature of the support (y-Al203 or Si02) the most stable configuration is always an M0S2 single sheet lying flat on the support at a distance of 2.8 A. Nevertheless, in view of the interacting energy reported in tables 5,6,7 the active phase is more stabilized by Y-AI2O3 surface planes than by silica exposed planes. Whatever the surface plane of the silica models, the energy Ei.sup which represents the forces between the active phase and the support is higher than the energy E1.2 of the interactions between two contiguous M0S2 layers ; whereas the energy Ei.sup between the slab and the [111] Y-AI2O3 support surface becomes smaller than the energy E1.2 (Table 7). Thus the silica support like the [110] Y-AI2O3 surface plane seems to be preferentially covered by stacked M0S2 layers whereas the [111] surface plane of Y-AI2O3 can be covered by several single sheets lying flat on the support. This result is in agreement with the well known ability of the alumina carrier to disperse disulfide crystallites in contrast to the silica support.^^
291 Table 7 Energies of interactions (Ei-sup) between an M0S2 slab and a plane of Y-AI2O3 or (E1.2) between two contiguous parallel M0S2 sheets Nature of the interactions Energy of interactions (kJ mol"^) M0S2 [110] plane of Y-AI2O3 Ei.sup = -649 M0S2 M0S2 E1.2 = -927 M0S2 [1111 plane of Y-AI2O3 Ei.sup = -1020 6. CONCLUSION A suitable force field has been developed in order to accurately model free M0S2 clusters or the non-bonded interactions between the active phase and the surface of the carrier. Such M0S2 layered clusters are stabilized in energy by the stacking of several sheets and whatever the nature of the support (y-Al203 or Si02), the most stable configuration is a single sheet lying flat on the surface ideally modelled by oxygen anions arranged in rows. Nevertheless, the non-bonded interactions between an M0S2 sheet and a plane of y-Al203 are stronger than similar interactions between the same M0S2 slab and whatever plane of the silica support considered in this work. The intercalation of one layer between two contiguous parallel sheets or between an M0S2 slab and the support exposed plane can be considered as a screen effect which increases with the number of inserted sheets. As deduced for M0S2 clusters supported by y-Al203, the energy difference between the metastable perpendicular configuration of M0S2 clusters supported by Si02 planes and the parallel one decreases with the number of stacked sheets whereas the activation energy increases. However the interaction between M0S2 and silica, whatever the surface planes of two models, are smaller than the forces between an M0S2 slab and gamma-alumina. Taking into account the energy (Ei.sup) for each modelled surface planes of the support, it appears that the stacking of several M0S2 layers is more favoured on silica support than on [110] plane of y-Al203 whereas the dispersion of single sheets lying flat on the [111] plane of y-Al203 support is more probable. These results are in agreement with the experimental measures which award to gamma-alumina support a better ability to stabilize well dispersed M0S2 clusters. REFERENCES 1. 2. 3. 4. 5. 6. 7. 8. 9.
F. Massoth, Adv. Catal, 1978, 25, 265 P. T. Vasudevan and F. Zhang, Appl. Catal. A, 1994, 112, 161 P. Amoldy, J.A.M. van den Heijkant, G.D. de Bok and J.A Mouljn, J. Catal, 1985, 92, 35 E. Payen, S. Kasztelan, S. Houssenbay, R. Szymanski and J. Grimblot, J. Phys. Chem, 1989, 93, 6501 E. Payen, S. Kasztelan, J. Phys. Chem., 1994, 4, 363 H. Tops0e, B.S. Clausen, N. Y. Topsee and P. Zeuthen, Catalysis in Petroleum Refining (D.L. Trimm et al. Eds), 1990, 77, Elsevier Amsterdam E. Payen, R. Hubaut, S. Kasztelan, O. Poulet and J. Grimblot, J. Catal, 1994, 147, 123 F. Delannay, Appl Catal, 1985, 16, 135 J.V. Sanders and K.C. Pratt, Micron, 1980, 11, 303
292 10. J.V. Sanders and K.C. Pratt, Micron, 1981, 67, 331 11. R. R. Chianelli, A. F. Ruppert, S. K. Behal, B. H. Kear, A. Wold and R. Kershaw, J. Catal., 1985, 92, 56 12. S. Eijsbouts, J. J. L. Heinerman and H. J. W. Elzerman, Appl. Catal. A, 1993,105, 53 13. P. Faye, E. Payen, D. Bougeard, J. Chem. Soc, Faraday Trans., 1996, 92(13), 2437 14. S. Srinivasan, A K. Datye and C. H. F. Peden, J. Catal, 1992,137, 513 15. Cerius^ version 1.6, Molecular Simulations Incorporated, Burlington, Massachusetts (USA) 16. K. D. Bronsema, J. L. De Boer and F. Jellinek, Z. anorg. allg. Chem., 1986, 540/541, 15 17. R. W. G. Wyckoff, Crystal Structure, 2nd ed., Interscience, New York, 1963,1, 280 18. A.P. Legrand, H. Hommel, A. Tuel, H. Balard, E. Papirer, P. Levitz, M. Czemichowski, R. Erre, H. Van Damme, J.P. Gallas, J.F. Hemid, J.C. Lavalley, O. Barres, A. Burneau, Y. Grillet, Adv. Col. Int. Sci., 1990, 33, 91 19. R. W. G. Wyckoff, Crystal Structure, 2nd ed., Interscience, New York, 1963, 3, 318 20. T. M. Brunier, M. G. B. Drew and P. C. H. Mitchell, Mol. Simulation, 1992, 9, 143 21. T. M. Brunier, M. G. B. Drew and P. C. H. Mitchell, J. Chem. Soc, Faraday Trans., 1992, 88(21), 3225 22. S. L. Mayo, B. D. Olafson and W. A. Goddard III, J. Phys. Chem., 1990, 94, 8897 23. M. A. Zhong-Xin and D. A. I. Shu-Shan, Acta Chim. Sinica, 1989, 3, 201 24. A. K. Rappe and W. A. Goddard III, J. Phys. Chem., 1991, 95, 3358 25. E. Diemann, Th. Weber and A. Miiller, J. Catal., 1994,148,288 26. T. F. Hayden and J. A. Dumesic, J. Catal., 1987,103, 366 27. C. Mauchausse, H. Mozzanega, P. Turlier and J. A. Delmon in "Proceedings, 9th International congress on Catalysis, Calgary, 1988" (M.J. Phillips and M. Ternan, Eds.), p. 775. Chem. Institute of Canada, Ottawa, 1988
© 7PP7 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G,F. Froment, B. Delmon and P. Grange, editors
293
In-situ FT-IR study of NO adsorbed on C0-M0/AI2O3 sulfided at high pressure(<5.1 MPa) Naoto Koizumi, Minoru lijima, Takeo Mochizuki, and Muneyoshi Yamada Department of Applied Chemistry, Faculty of Engineering, Tohoku University, Aoba, Aramaki, Aoba-ku, Sendai 980-77, JAPAN
The surface structure of C0-M0/AI2O3 sulfided at high pressure(^5.1 MPa) was investigated in-situ by FT-IR/DRA study of adsorbed NO. The ratio of the band intensity of NO adsorbed on Co site to that of Mo site was much higher when Co-Mo/A 2 O3 was sulfided at the pressure above 1.1 MPa than when it was sulfided at atmospheric pressure. Combining the results with those found on catalysts varying the Co/Mo molar ratio and Mo K-edge EXAFS, it was suggested that "Co-Mo-S"-like structure is effectively formed on C0-M0/A2O3 by higher pressure sulfiding.
1. INTRODUCTION Sulfided C0-M0/A2O3 plays very important role in hydrotreatment of petroleum fractions, and many efforts have been made so as to improve the catalytic performance. It is well known that the addtion of small amount of Co to M0/A2O3 drastically improve the HDS activity. This is one aspect of synergy effect between Mo and Co. Therefore, many studies have aimed to make clear the active sites in relation to this synergy effect. The latest developments of surface analysis techniques also contributes to evaluate the surface structure of C0-M0/A2O3, and some models of active sites for HDS reaction are proposed. That is, Tops0e et al. have proposed "Co-Mo-S" model by means of Mossbauer emission spectroscopy(MES) as the active site for HDS of thiophenes on C0-M0/A2O3 [1]. Furthermore, local structure of "Co-Mo-S" on Co-Mo/active-C has been proposed by meams of Mo and Co K-edge EXAFS [2,3]. However, some problems are pointed out about "Co-Mo-S" model. One is that there is a complexity with signal assignment in MES. That is, it has been reported that "Co-Mo-S" like Mossbauer signal appears even in Co/active-C [4]. This might make the meaning of the fact ambiguous that the intensity of "Co-Mo-S" signal in MES correlates with thiophene HDS activity. The other is concerning the stability of "Co-Mo-S" in reaction condition. That is, Breysse et al. have reported the results of MES and XPS of unsupported Co-Mo sulfides using HDS of dibenzothiophene under pressurized condition [5-7]. Based on these results, they speculated that Co in "Co-Mo-S" on unsupported Co-Mo sulfide was segregate in their reaction condition. This vSpeculation dose not necessarily exclude the existence of "Co-Mo-S" on CoM0/A2O3, but might limit the role of "Co-Mo-S" in industrial HDS reaction. A s o in relation to this problem, "Remote control model" has been proposed by Delmon et al. [8]
294 Concerning to these problems, some points should be reminded about experimental techniques. That is, i) the characterization techniques adopted so far are mainly MES and EXAFS, whose infonnation is predominated by the over-all structure of the catalysts, not by the surface structure. Therefore, combining these with surface sensitive techniques is necessary to evaluate the surface structure of C0-M0/AI2O3 more clearly. And ii) the catalyst pretreating (sulfiding) conditions reported so far are limited to an atmospheric pressure, which is far from working state of C0-M0/AI2O3. It is possible that the surface structure of C0-M0/AI2O3 sulfided at atmospheric pressure is different from that of C0-M0/AI2O3 sulfided at higher pressure. In order to clarify the active sites for industrial HDS reaction on C0-M0/AI2O3, it is necessary to evaluate the surface structure of C0-M0/AI2O3 under its working state, at least to evaluate in-situ (without exposing the catalyst to the air) the surface structure of C()-Mo/Al203 sulfided at high pressure. In relation to the problem of i), the authors have characterized the surface structure of CoM0/AI2O3 by combining NO probe method(NO uptake measurement and FT-IR/DRA study of adsorbed NO) as surface sensitive technique and Mo K-edge EXAFS, and found the new aspects of active site on C0-M0/AI2O3. That is, we have investigated the benzothiophene HDS activity and the surface structure of C0-M0/AI2O3 which is pretreated by combination of sulfiding and reducing, and reported that the formation ratio of Co site to Mo site correlates with the benzothiophene HDS efficiency [9]. In addition, comparing the surface structure of C0-M0/AI2O3 with that of Ni-Mo/Al203 and Fe-Mo/Al203 in detail, it was found that the growth of MoS2-like structure in the lateral direction on Fe-Mo/Al203, in which synergy effects for HDS of benzothiophene do not appear, is suppressed compared to that on C0-M0/AI2O3 and Ni-Mo/Al203 [10,11]. However, in these studies, the problem of ii) remains to be solved. The present work aims to evaluate the surface structure of C0-M0/AI2O3 under its working state. That is, Co and Mo site distribution on the surface of C0-M0/AI2O3 sulifided at high pressure is investigated by FT-IR/DRA study of adsorbed NO using high temperature( ^ 773 K) and pressure( ^ 5 . 1 MPa) in-situ IR cell. Furthermore, fine structure around Mo of CoM0/AI2O3 is also investigated by Mo K-edge EXAFS using in-situ EXAFS cell which can pretreat the catalyst at the same pressurized condition as FT-IR/DRA study of adsorbed NO.
2. EXPERIMENTAL 2 . 1 . Catalyst preparation Home-made C0-M0/AI2O3, M0/AI2O3 and C0/AI2O3 were used. Commercial Co-Mo /AI2O3(supplied by Nippon Ketjen Co.) was also used. In the following text, commercial CoM0/AI2O3 is denoted by Co-Mo/Al203(NK) so as to distinguish it from home-made CoM0/AI2O3. Home-made C0-M0/AI2O3, M0/AI2O3 and C0/AI2O3 was prepared by incipient wetness method. 7 -Al203(vSupplied by Nippon Ketjen Co., 333 m^/g, 0.76 cm-^/g) was impregnated by (NH4)6Mo7024'4H20(supplied by Wako Pure Chemical Industries Ltd.) solution followed by drying(393 K, 16 h) and calcination(793 K, 12 h), then Mo/Al203(Mo03:15.8 wt%) was obtained. This was impregnated by Co(N03)2*6H20(supplied by Wako Pure Chemical Industries Ltd.) solution followed by drying(393 K, 16 h) and calcination(793 K, 12 h), finally Co-M()/Al203 was obtained. Co/Mo molar ratio was fixed at that of Co-Mo/Al203(NK)(().56
295 mol/mol). Co/Al203(CoO:5.2 wt%) was prepared as mentioned above. 2 . 2 . FT-IR/DRA of adsorbed NO Catalyst sulfiding procedure is as follows. First, finely powdered catalyst was placed in the IR cell. 5%H2S/H2 was introduced into the IR cell at the pressure of 0.1 to 5.1 MPa, at room temperature. Under flowing 5%H2S/H2(30 ml/min(STP)), the catalyst was heated to 673 K at the rate of 20 K/min. This temperature was kept for the time of 1 to 10 h. After that, the catalyst was cooled to room temperature, then He flow was introduced into the IR cell to purge 5%H2S/H2. Unless otherwise stated, sulfidation was carried out for 1 h. 10%NO/He was introduced to the IR cell by pulse method. After flushing the gas-phase NO by He flow, FT-IR/DRA spectrum of NO adsorbed on the catalyst was measured by means of diffuse reflectance method using FYS6000(Bio-Rad) spectrometer equipped with mercurycadmium-telluride detector. 2 . 3 . Mo K-edge EXAFS Catalyst sulfiding procedure is as follows. First, the catalyst was pressed into pellets, and was set in the EXAFS cell. The EXAFS cell was made reffered to that reported by Boudart et al. [12] 5%H2S/H2 was introduced into the EXAFS cell at the pressure of 0.1 to 5.1 MPa, at room temperature. Under flowing 5%H2S/H2(150 ml/min(STP)), the catalyst was held at room temperature for 1 h, and then was heated to 673 K at the rate of 10 K/min. This temperature was kept for 2 h. After that, the catalyst was cooled to room temperature, and then He was introduced into the IR cell to flush 5%H2S/H2. After these procedures, X-ray absorption spectra were measured using Laboratory System equipped with double-crystal monochromator [13,14]. The interatomic distances and the coordination numbers were determined by curve fitting techniques [15]. The backscattering amplitudes and phase shifts reported by McKale et al. [16] were used. M0S2 powder(supplied by Wako Pure Chemical Industries Ltd.) was used as reference for Debye-Waller factor and photoelectron mean free path.
3 . RESULTS AND DISCUSSION 3 . 1 . Effect of sulfiding pressure on Co and Mo site distribution on CoM0/AI2O3 Figure la) shows FT-IR/DRA spectra of NO adsorbed on C0-M0/AI2O3, M0/AI2O3 and C0/AI2O3 sulfided at 3.1 MPa. Two bands appear in the spectrum on C0-M0/AI2O3 at 1850 and 1800 cm'^. Much weaker band also appears around 1700 cm"^. From comparison with the spectra on M0/AI2O3 and C0/AI2O3, the lower frequency(1800 and 1700 cm"^) bands and higher frequency(1850 and 1800 cm"^) bands in C0-M0/AI2O3 can be assigned to Mo site-band and Co site-band respectively. This assignment is consistent with that in our preceding papers using C0-M0/AI2O3 sulfided at 0.1 MPa [9,17,18]. As shown in Figure la), it is noticed that the spectrum of NO adsorbed on C0-M0/AI2O3 is not simple superposition of those on M0/AI2O3 and C0/AI2O3. That is, the intensity of Co site-band is very strong compared to that of Mo site-band in C0-M0/AI2O3 sulfided at 3.1 MPa. For comparison, the catalysts sulfided at 0.1 MPa were also supplied for FT-IR study, as is shown in Figure lb). Comparing Figure la)
296 with Figure lb), the relative intensity of Co site-band of C0-M0/AI2O3 sulfided at 3.1MPa is stronger than that of C0-M0/AI2O3 sulfided at 0.1 MPa. Figure 1 shows that Co and Mo site distribution on the surface of C0-M0/AI2O3 is clearly affected by sulfiding pressure.
2100
1900
1700
Wavenumber /cm "^
1500
2100
1900
1700
1500
Wavenumber /cm "^
Figure 1. FT-IR/DRA spectra of NO adsorbed on Co-Mo, Mo and Co. Sulfiding pressure: a)3.1 MPa, b)0.1 MPa
The effect of sulfiding pressure was further investigated. Figure 2 shows FT-IR/DRA spectra of NO adsorbed on C0-M0/AI2O3 sulfided at the pressure of 0.1 to 5.1 MPa. It is noticed that the spectra of NO adsorbed on C0-M0/AI2O3 sulfided over 1.1 MPa is different from that on C0-M0/AI2O3 sulfided at 0.1 MPa. That is, the intensity ratio of Co site-band to Mo site-band is much higher when C0-M0/AI2O3 is sulfided over 1.1 MPa than when it is sulfided at 0.1 MPa. It is also noticed that no drastic change appears with increasing sulfiding pressure above 1.1 MPa. For comparison, FT-IR/DRA spectra of NO adsorbed on C0-M0/AI2O3 sulfided at 0.1 MPa for the time of 1 to 10 h are shown in Figure 3. The relative intensity of Co site-band slightly increases with increasing sulfiding time. However, comparing Figure 2 with Figure 3, the relative intensity of Co site-band of C0-M0/AI2O3 sulfided at 1.1 MPa(for 1 h) is stronger than that of C0-M0/AI2O3 sulfided at 0.1 MPa for 10 h. These results indicate that the surface structural change induced by 1.1 MPa sulfiding can not be attained by simple extension of sulfiding duration at 0.1 MPa.
297
^
2100
1900
^
1700
Sulfiding pressure
1500
Wavenumber/cm''' Figure 2. FT-IR/DRA spectra of NO adsorbed on Co-Mo sulfided at high pressure.
2100
1900
1700
1500
Wavenumber/cm'^ Figure 3. FT-IR/DRA spectra of NO adsorbed on Go-Mo with various sulfiding time. Sulfiding pressure: 0.1 MPa
3.2. Surface structure of C0-M0/AI2O3 formed by high pressure sulfiding As mentioned above, FT-IR spectrum of NO adsorbed on C0-M0/AI2O3 is not a simple superposition of those on M0/AI2O3 and C0/AI2O3. In order to understand this point with respect to Figure 2, a physical mixture of M0/AI2O3 and C0/AI2O3 was also supplied for FTIR/DRA study of adsorbed NO. Co/Mo molar ratio of the physical mixture was fixed at 0.56, which is equal to that of C0-M0/AI2O3. Figure 4 shows the vSpectra of NO adsorbed on the physical mixture sulfided at the prssure of 0.1 to 5.1 MPa. Comparing Figure 4 with Figure 1, it is clear that the spectra of NO adsorbed on the physical mixture are simple superposition of those on M0/AI2O3 and C0/AI2O3. This supports that Co and Mo site are not interacted each other in the physical mitxure. The intensity ratio of Co site-band to Mo site-band of CoM0/AI2O3 is compared with that of the physical mixture, as shown in Figure 5. Comparing CoM0/AI2O3 with the physical mixture, following two points are noted. (A)The relative intensity of Co site-band of C0-M0/AI2O3 is much stronger than that of the physical mixture in the range of 0.1 to 5.1 MPa. This suggests that Co site is selectively formed on C0-M0/AI2O3 compared to the physical mixture of M0/AI2O3 and C0/AI2O3. And (B)it is also noted that the difference of the relative intensity of Co site-band between them is remarkable when the catalysts are
298 sulfided at the pressure above 1.1 MPa. This suggests that Co site is more profoundly formed when C0-M0/AI2O3 is sulfided at the pressure above 1.1 MPa. These two phenomena (A) and (B) suggest that a kind of surface interaction species is effectively formed between Co and Mo when C0-M0/AI2O3 is sulfided at high pressure.
0.2
Co ^Q
Sulfiding pressure
Co-Mo
5
> s ^ -t—'
if)
c 0) 4 - 1 — ^^ _c
•(/)
c
CD
•a
0.4
1.1 MPa
^ ? cc ^
_Q (D
^2
•
I
CD
^1 •
1900
1700
1500
Wavenumber/cm'^ Figure 4. FT-IR/DRA spectra of NO adsorbed on physical mixture sulfided at high pressure.
•
Physical mixture • •
n 0
2100
•
-
CO
0.1 MPa
•
1 2 3 4 5 Sulfiding pressure /MPa
1 1
Figure 5. Effect of sulfiding pressure on the relative band intensity.
The number of coordinatively unsaturated site(Co and Mo site) is quantified by NO uptake measurements(Figure 6). As shown in Figure 6, the amount of NO adsorbed on Co-Mo /Al203(NK) decreases with increasing sulfiding pressure, which indicates that the number of Co and Mo site decrease with increasing sulfiding pressure within the range examined here. Furthermore, comparing C0-M0/AI2O3 with M0/AI2O3 and C0/AI2O3, following points are noted. That is, when the catalysts are sulfided at 3.1 MPa, the amount of NO adsorbed on C0/AI2O3 is very small, and the amount of NO adsorbed on C0-M0/AI2O3 is much larger than the sum of NO adsorbed on M0/AI2O3 and C0/AI2O3. These results also suggest that the interaction between Co and Mo is effectively formed on C0-M0/AI2O3 when it is sulfided at high pressure. In order to examine the interaction between Co and Mo from a different point of view, the catalysts varying the Co/Mo molar ratio were also supplied for FT-IR/DRA study of adsorbed NO. Figure 7 shows FT-IR/DRA spectra of NO adsorbed on the catalysts sulfided at 3.1 MPa.
299 With increasing the ratio of Co/Mo, the relative intensity of Co site-band increases, whereas that of Mo site-band decreases steaply and almost disappears. The similar phenomenon has been reported for the catalysts sulfided at atmospheric pressure, and interpreted that coordinatively unsaturated Mo site is capped by Co on C0-M0/AI2O3 [19]. This interpretation has brought about "Co-Mo-S" model. Therefore, Figure 7 suggests that "Co-Mo-S"-like structure is formed on C0-M0/AI2O3 sulfided at 3.1 MPa as well as atmospheric pressure. Based on this interpretation, (B) suggests that "Co-Mo-S"-like structure is effectively formed when CoMo/Ah O3 is sulfided at the pressure above 1.1 MPa. Formation of such a specific Co site on C0-M0/AI2O3 can also explain the difference of the relative intensity of Co site-band betvs^een C0-M0/AI2O3 and the physical mixture shown in Figure 5.
'
i
6 0 5
p
Co-Mo
• B
.
E4 i
CD -^ 0 03 0 Q=3
Co-Mo(NK)
_
^
Mo •
k
0 2 2
1
Go
n
1
0
1
1
li
1
1
I—I—1
1 2 3 4 5 Sulfiding pressure /MPa
1
6
Figure 6. Effect of sulfiding pressure on the amount of NO adsorbed on CoMo, Mo and Co.
2100
1900
1700
1500
Wavenumber/cm"^ Figure 7. FT-IR/DRA spectra of NO adsorbed on Co-Mo with various Co/Mo ratio. Sulfiding pressure: 3.1 MPa
3 . 3 . Effect of sulfiding pressure on the morphology of MoSz-Hke structure As mentioned above, the sulfiding pressure dependency of FT-IR/DRA spectra of adsorbed NO can be interpreted that "Co-Mo-S"-like structure is effectively formed when C0-M0/AI2O3 is sulfided at high pressure. In order to make clear the limit of this interpretation, another possibility was further investigated. It is possible that the number of exposed Mo site correlates with the morphology of MoSo-like structure, that is, the growth of MoS2-llike structure in the
300
lateral direction. Therefore, another interpretation for (A) and (B) might be that the growth of MoS2-like structure in the lateral direction is promoted on C0-M0/AI2O3 sulfided at high pressure. This point was examined by Mo K-edge EXAFS.
4 6 Distance /A Figure 8. Fourier transforms of EXAFS spectra of Co-Mo sulfided at high pressure.
4 6 8 Distance /A Figure 9. Fourier transforms of EXAFS spectra of physical mixture sulfided at high pressure.
Figure 8 shows the Fourier transforms of Mo K-edge EXAFS spectra(radial structure function) of C0-M0/AI2O3 sulfided at the pressure of 0.1 to 5.1 MPa. Figure 8 also shows that of M0S2 powder for reference. As shown in Figure 8, only Mo-S and Mo-Mo shells clearly appear in the radial structure function of the catalysts sulfided at the pressure of 0.1 to 5.1 MPa. These two peaks suggest that fine structure around Mo in C0-M0/AI2O3 can be approximated by MoS2-like structure when C0-M0/AI2O3 is sulfided at the preSvSure of 0.1 to 5.1 MPa. It is noticed that the peak intensities of Mo-S and Mo-Mo shells are hardly affected by sulfiding pressure. Coordination numbers calculated about Mo-S and Mo-Mo shells of C0-M0/AI2O3 are shown in Figure 10. As can be seen from Figure 10, both Mo-S and Mo-Mo coordination numbers are almost constant in the range of 0.1 to 5.1 MPa. These results suggest that the
301 morphology of MoS2-like vStructure formed in C0-M0/AI2O3 is hardly affected by higher pressure sulfiding. Figure 9 shows the radial structure function of the physical mixture sulfided at 0.1, 1.1 and 5.1 MPa. The physical mixture examined here is the same as that for the FT-IR/DRA measurement already mentioned. Mo-S and Mo-Mo coordination numbers of the physical mixture are shown in Figure 10. Comparing C0-M0/AI2O3 with the physical mixture, Mo-Mo coordination number of C0-M0/AI2O3 is constant, and comparable with that of the physical mixture in the range of 0.1 to 5.1 MPa. This suggests that the growth of MoS2-like structure in the lateral direction on C0-M0/AI2O3 is comparable with that on the physical mixture. It can be said that the presence of Co does not affect the morphology of MoS2-like structure of sulfided C0-M0/AI2O3. The present Mo K-edge EXAFS studies show that the morphology of MoS2-like structure of C0-M0/AI2O3 is hardly affected by high pressure sulfiding and the presence of Co. It is interesting to compare this finding with that of FT-IR/DRA of adsorbed NO. The present FTIR/DRA study of adsorbed NO shows that Mo site is extinguished by high pressure sulfiding. Therefore, it can be said that the surface structural change induced on C0-M0/AI2O3 by high pressure sulfiding is not accompanied by the growth of MoS2-like structure in the lateral direction. In other words, "Co-Mo-S"-like structure is effectively formed when C0-M0/AI2O3 is sulfided at the pressure above 1.1 MPa.
5
Mo-S
•
§4 3 C C
.9 3
1
•
•
a
D J D
0 •
Mo-Mo •
0
E2 » 0 0
0 , 1 n
1
0 1
1
1
J
1
1
1
L_-1
1
1
2 3 4 5 Sulfiding pressure /MPa
Figure 10. Effect of sulfiding pressure on Mo-S and Mo-Mo coordination numbers. • 9 : Co-Mo, n O- physical mixture
1
302 5. CONCLUSION In order to evaluate the surface structure of C()-Mo/Al203 under its working state, the surface structure of C0-M0/AI2O3 sulfided at high pressure(^5.1 MPa) was investigated in-situ by FTIR/DRA study of adsorbed NO and Mo K-edge EXAFS. Based on these results, it was suggested that "Co-Mo-S"-like structure is effectively formed by higher pressure sulfiding. The present work shows more straightforward results based on which one can discuss about the surface structure of C0-M0/AI2O3 under its working state with respect to Co and Mo site distribution.
ACKNOWLEDGEMENT We thank Professor Yasuo Udagawa for usefull discussion concered with EXAFS measurements and use of EXAFS spectrometer. This work was supported by a Grant-in-Aid for Scientific Research from the Ministry of Education, Science and Culture of Japan.
REFERENCES 1. H. Tops0e and B. S. Clausen, Catal. Rev. Sci. Eng., 26 (1984) 395. 2. S. M. A. Bouwens, R. Prins, V. H. J. de Beer and D. C. Koningsberger, J. Phys. Chem., 94(1990)3711. 3. S. M. A. Bouwens, J. A. R. Van Veen, D. C. Koningsberger, V. H. J. de Beer and R. Prins, J. Phys. Chem., 95 (1991) 123. 4. A. M. van der Kraan, M. W. J. Craje, E. Gerkema and W. L. T. M. Ranselaar, Appl. Catal., 39 (1988) L7. 5. M. Breysse, R. Frety, B. Benaichouba and P. Bussiere, Radiochem. Radioanal. Lett., 59 (1983) 265. 6. M. Vrinat, M. Breysse and R. Frety, Appl. Catal., 12 (1984) 151. 7. M. Breysse, R. Frety and M. Vrinat, Appl. Catal., 12 (1984) 165. 8. B. Delmon, in T. Inui, K. Fujimoto, T. Uchijima and M. Masai, (Editors), New Aspects of Spillover Effect in Catalysis, Elsevier, Amsterdam, 1993, p. 1. 9. S. Kasahara, S. Miyabe, T. Shimizu, H. Takase and M. Yamada, Sekiyu Gakkaishi, 38 (1995) 81. 10. S. Kasahara, N. Koizumi, J. Iwahashi and M. Yamada, Sekiyu Gakkaishi, 38 (1995) 345. 11. S. Kasahara, N. Koizumi, M. Yamada and Y. Udagawa, Sekiyu Gakkaishi, 38 (1995) 439. 12. R. A. Dalla Betta, M. Boudart, K. Foger, D. G. Loffler and Sanchez-Arrieta, Rev. Sci. Instrum., 55(12) (1984) 1910. 13. K. Tohji, Y. Udagawa, T. Kawasaki and K. Mieno, Rev. Sci. Instrum., 59(7) (1988) 1127. 14. Y. Udagawa, The Rigaku Journal, 6 (1989) 20. 15. K. Tohji, Y. Udagawa, S. Tanabe and A. Ueno, J. Am. Chem. S o c , 106 (1984) 612. 16. A.G. McKale, B.W. Veal, A.P. Paulikas, S.-K. Chan and G.S. Knapp, J. Am. Chem. S o c , 110(1988)3763. 17. T. Obara, M. Yamada and A. Amano, Chem. Lett., (1986), 2003. 18. M. Yamada, S. Kasahara and K. Kawahara, in T. Inui, K. Fujimoto, T. Uchijima and M. Masai, (Editors), New Aspects of Spillover Effect in Catalysis, Elsevier, Amsterdam, 1993, p. 349. 19. N.-Y. Tops0e and H. Tops0e, J. Catal., 84 (1983) 386.
® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
303
Compound Formation and Hydrogen Activity at Sulfided Catalysts: A Combined Surface Science and Quantum Chemical Approach J. PauP, H. Akpati^ P. No^dlander^ W.S. Oh^ D.W. Goodman^ and B. Demirel^§ ""Bilkent University, Chemistry, Bilkent 06533, Ankara, Turkey *^Rice University, Physics, P.O.Box 1892, Houston, TX 77251-1892, USA Texas A&M University, Chemistry, College Station, TX 77843-3255, USA ^University of Utah, Chemical and Fuels Eng., Salt Lake City, UT 84112, USA HOT correspondence, present address: Center for Applied Energy Research, 3572 Iron Works Pike, Lexington, KY 40511-8433, USA ABSTRACT This paper discusses the hydrogen activity of sulfided metal catalysts supported on modified carriers. Surface characterization and theoretical modeling are combined with published studies of planar models to deduce a conceptual model for hydrogen dissociation and mobility on industrial hydrogenation catalysts. The hydrogen activity of metalsulfides is compared with that of the corresponding transition metals. TEXT Sulfided catalysts are workhorses of the refining industry. They are, among other processes, used for hydrogenation reactions at elevated temperatures. This activity as well as the activity for hydrodesulfiuization (HDS) and hydrodenitrogenation (HDN) is intimately related to the hydrogen activity of metal sulfides i.e. to the properties of adsorbed atomic hydrogen and the rate of hydrogen dissociation at the surfaces of these compounds [1]. The present work is an attempt to model these properties by advanced electron structure methods, experimental and theoretical. The work is part of a project for the production of gasoline components fi-om polyaromatic hydrocarbons (PAH) [2]. Other parts address conversion rates and turnover fi-equencies, mass balance and surface intermediates [3]. A sulfided catalyst consists of three dimensional metal sulfide particles bound to a supporting oxide. Sulfidation lowers the dispersion of an impregnated metal phase by the strong
304 driving force for sulfide formation and structural methods reveal particulates of stoichiometric compounds. This coalescence opens areas of the supporting oxide for adsorption, an important secondary result with relevance for cracking reactions at the modified sites created by electronic modifications in titania-alumina or silica-alumina catalysts. These modified sites also act as dispersion agents for the metal in its oxidic state, but this property has less significance after sulfidation than after mild reduction, sulfidation provides an extra incentive for coalescence. Modified or acidic sites are centra for dehydrogenation reactions, a property commonly used to produce unsaturated compoundsfi^omalkanes [4]. Hydrogenation of 1-methylnaphthalene was found to proceed effectively over a sulfided NiMo catalyst supported on a titanium doped alumina, but only remote quantities of cracking products were observed [2]. This indicates that the acid sites were blocked either by supported metals or carbonaceous species. This motivates our interest in isolated species on models of the carrier. The present work includes experimental characterization of the effects of sulfidation by static weight changes for stoichiometry and by photoemission spectroscopy (ESCA) for chemical characterization. Hydrogen adsorption is modeled by calculations. Transition metals have superior activity for hydrogenation reactions below the temperature for recombinative desorption of hydrogen. Above this temperature these materials instead possess dehydrogenation activity, an often activated process, due to weakening and eventual dissociation of C-H bondsfi-omoverlap with the surface electron density. Unsupported metals are less useful for high temperature hydrogenation reactions due to the low surface coverage of atomic hydrogen, a result of molecular desorption after H-H recombination, not because of an altered ability for hydrogen dissociation. Hydrogen bond-weakening and dissociation is a result of electronic overlap, a property only marginally affected by temperature. Hydrogen dissociation occurs because hydrogen orbitals overlap or hybridize with d-orbitals of appropriate symmetry when the molecule approches the surface with the H-H axis parallel to the surface. Mixing between the metal d-states and hydrogen a and o* orbitals bring the latter partly below the Fermi level. This results in bondweakening and eventual dissociation. Atomic hydrogen from the above process is trapped in the surface region. The tail of electron density gives at a certain height above the surface a most favorable electron density, well formulated in the effective medium approach [5]. Further penetration into the metal is obstructed by repulsion from the higher density. Motion away from the surface is also hindered, atomic desorption is only observed at drastically higher temperatures, a well known phenomenon for the production of atomic hydrogen in diamondfilmsynthesis. Atomic hydrogen is, as a consequence, trapped in the surface region. The adsorption potential for atomic hydrogen is predominantly a result of thefreeor sp-electron density with a smaller perturbation from d-states. The sp-density is evenly spread in the surface plane with only a small corrugation due to the atomic cores. This means that no major barriers exist for difllision along the surface, resulting in a comparatively low temperature for recombinative desorption. The maintained high dissociation rate at elevated temperatures means that atomic hydrogen can become available as long as a competitive process is more effective than
305 recombinative desorption. This is the concept of hydrogen spillover. Hydrogen spillover is promoted by small particle sizes and enhanced temperatures. The small particle size means an enhanced chance for hydrogen diffusion across the boundary between the metal particle and the supporting oxide. Large particles and poor dispersion instead promote hydrogen recombination. An enhanced temperature is necessary to boost the diffusion rate across the metal/support barrier and further via the the support material. The key feature of metal sulfides is to maintain a high coverage of atomic hydrogen at the elevated temperatures of hydrotreatment reactors. This means that recombinative desorption is quenched but not necessarily that the rate of hydrogen dissociation on the terminated metal sulfide particles is high. Different concepts for hydrogen dissociation on sulfided catalysts exist. One being the existence of special sites directly on the metal sulfide particles, these sites are often envisioned as more metallic and thus more active for hydrogen dissociation than the bulk sulfide. Another concept is a dualfixnctioncatalyst where a supported well dispersed metal provides atomic hydrogen, followed by hydrogen diffusion over to a second metal, in its sulfide form, where the hydrogenation occurs [6]. Either model is possible. Typical catalysts are bimetallic with one metal. Mo or W, forming the dominant sulfide and another metal, Co or Ni, acting as wetting agent between the metalsulfide and the support or possibly forming a separate metallic phase with high capacity for hydrogen dissociation. The hydrogenation rate usually peaks at a Mo:Co or W:Ni ratio of 3:1 [1]. The reducing environment of hydrogenation reactions can easily reduce any metal still in the oxidic state as well as create special sites at the surface of M0S2 or WS2 particles. The problem with each model is that they are often deduced from kinetics data and it is non-trivial to know which barrier a certain apparent energy term comes from. The present calculations center around cluster models of hydrogen adsorption on metal sulfide particles. The electron structure is solved in a local density approach with numerical radial parts of the atomic wavefimctions and a new dense and thus accurate integration grid. We consider surface coverages as well as penetration into the bulk with reference to the known 3D structure of the sulfide particles and the role this morphology plays for the hydrogen activity. We also refer to previous studies of hydrogen adsorption on unsupported metals. The structure of naphthalene and 1-Me-naphthalene, model compounds for heavy feeds, are modelled in a less computer intensive ab initio code [7]. The experimental part contains weight measurements to obtain the stoichiometry of sulfides at the surface and photoemission spectra to characterize these phases. The catalysts were either retrieved from intermittent steps in autoclave studies with actual feeds or prepared in parallel batches for reactor studies and surface characterization. Quantitative measurements of mass changes during catalyst pretreatments, dehydration and sulfidation, are also prerequisites for enhanced accuracy in mass balance calculations including carbonaceous residues [3].
306 REFERENCES 1. O. Weisser and S. Landa, Sulphide Catalysts, Their Properties and Applications, (Pergamon Press, Oxford, 1973); R. Prins, V. H. J. De Beer and G. Somorjai, CatalRev.Sci.Eng. 31 (1989) 1; P. Grange, CatalRev.Sci.Eng. 21 (1980) 135; J. J. F. Scholten, A. P. Pijpers, and A. M. L. Hustings A.M.L., Catal.Rev.Sci.Eng.27 (1985) 151; J. A. Rodriguez, S. Y. Li, J. Hrbek, H. H. Huang and G. Q. Xu, J.Phys.Chem. (in press); R. I. Declerek-Grimee, P. Canesson, R. M. Friedman and J. J. Fripiat, J.Phys.Chem. 82 (1978) 889 2. B. Demirel, PhD Thesis, Univ.of Utah (1996) 3. B. Demirel and W. H. Wiser, manuscript; B. Demirel and J. Paul, manuscript; §. Sayan, B. Demirel, W.H. Wiser, §. Suzer, and J. Paul, manuscript 4. G B. McVicker, G. M. Kramer and J. J. Ziemiak, J.Catal.83(1983)286; A. Corma and B. W. Wojciechowski, CatalRev.Sci.Eng. 279 (1985) 29 5. P. Nordlander, S. Holloway and J. K. N0rskov, Surface Sci. 136 (1984) 59 6. A. M. Stumbo, P. Grange and B. Delmon, in Studies in Surface Science and Catalysis (Proceedings 11th ICC) 101 (1996) 97 7. B. Demirel and J. Paul, Z.Phys.Chem., submitted
© igc)^ Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
307
Deactivation studies on NiO-MoOs/AliOs and C0O-M0O3/AI2O3 hydrodesulphurization catalysts R.Ma^mkovic-Neducin^ E.Kis", M.I>juric', J.Kiurski^ D.Z.Obadovic*' ,P.Pavlovic^ KMcic' ^University of Novi Sad, Faculty of Technology, 21000 Novi Sad, Bul.Cara Lazara 1, YU ^ University of Novi Sad, Faculty of Sciences, 21000 Novi Sad, Trg D.Obradovica 4, YU ''NIS-Rafinery Novi Sad, 21000 Novi Sad, Put Sajkaskog odreda bb., Yugoslavia The results of comparative investigation of NiO-Mo03/y-Al203 hydrodesulphurization (HDS) catalyst deactivation in industrial plant and laboratory conditions are presented. Structural (XRD, DRS, XPS), textural (LTNA) and morphological (SEM) characteristics were followed depending on temperature, time of treatment and the atmosphere of regeneration. Based on the mathematical models the of textural changes intensity depending on two independent variables (time, temperature) the role of these variables in catalyst sintering were estimated. The mechanism of deactivation process is proposed, pointing out critical conditions in processing and/or regeneration. Parallel investigation of C0O-M0O3/Y-AI2O3 catalyst was the basis for relative stabiHty estimation depending on the precursor type.
1. INTRODUCTION Broad appHcation of NiO-Mo03/Y-Al203 or CoO-Mo03/y-Al203 catalysts in hydrodesulphurization (HDS) processes generate permanent interest in investigation of different aspects of these standard industrial catalysts. Due to long-term activity of these catalysts, their deactivation received less attention, the investigations being mostiy oriented to the chemistry, structure and reaction mechanisms [1-2]. The scientific hterature in the field of HDS catalyst deactivation is mainly concemed with coldng and poisoning [3-6]. New environmental regulations impose advanced HDS/HDN processes, with improved catalysts being able to operate under more severe conditions, which together with increased interest for heavier fraction processing in more rigorous process conditions, bring about the new stimulus for catalyst aging investigations. Our previous investigations on aging of HDS catalyst and corresponding model systems have shown that sintering is one of the main processes that could cause activity decline in unpreferable regeneration conditions[7-9]. Oxidizing atmosphere in regeneration process is identified as specially critical for rapid sintering of the catalyst, the mechanism of sintering process being based on previously structurally changed active phase under the unsuitable operating conditions This study deals with comparative investigations of industrially and laboratory deactivated HDS catalyst, both based on Ni and Co as the promoters.
308 2. EXPERIMENTAL 2.1. Catalyst samples Two types of standard commercial HDS catalysts were investigated, NiO-Mo03/Y-Al203 and CoO-Mo03/y-Al203. The samples of NiO-Mo03/y-Al203 from industrial hydrotreating plant, from different reactor layers, were taken after regeneration process in steam atmosphere. Comparative investigation of the samples of both catalyst types after laboratory simulation of catalyst aging were investigated. 2.2. Laboratory simulation of catalyst aging Laboratory simulation of catalyst aging in oxidation atmosphere was reahzed in mufiQe ftimace in static air, nitrogen and steam conditions. Temperature of treatment was 500, 600, 700 and SOO^'C and treatment duration 1, 3, 6 and 9 h. 2.3. Methods Catalyst structure was characterized by: XRD (Phil^s, PW 1050 CuKa); SEM (JEOL, ISM 35); DRS (SPM-2 monochromator Veb Zeiss, Jena) with a reflection cell of the R-45/0 type; XPS (Surface Science SSX-100, small spot). Low temperature nitrogen adsorption (Micromeritics, ASAP 2000) was performed for textural properties' investigations. 2.4. Mathematical modeling Mathematical modehng was appHed m interpreting textural properties (surface area, pore volume, average pore diameter) as the ftmctions of temperature and time as two independent variables. Based on an analysis of shapes of the surfaces, Sp(T, t), Vp(T, t) and R(T, t), different polynomials models have been tested and the simplest (linear in terms of time but quadratic in terms of heating temperature) was accepted: y(T,t)=bi+b2T+b3t+b/
(1)
3. RESULTS AND DISCUSSION The activity decrease was observed after restarting hydrotreating unit with steam regenerated catalyst. The process investigation with variation of feed quahty (total sulphur level, end of distillation) confirmed pronounced activity decline in processing higher suJ^hur level, i.e. more thiophenic feeds, being still in plant limits. The problem could not be solved by increasing reactor temperature eider/or decreasmg LSHV, indicating partial deactivation of the catalyst. The characterization of regenerated catalyst samples from different reactor layers confirmed that only a part of catalyst loading was changed concerning structure and texture. The upper layer contained a considerable fraction of the sintered catalyst grains (surface area decrease of 79% in comparison to fresh sample). XRD and SEM results showed the nicrease of the crystallinity of catalyst support in that fraction, followed by formation of aluminium molybdate and M0O3 phase. The SEM investigations of the samples taken from reactor nozzle indicated formation of free M0O3 crystals, extracted from the catalyst. In the deeper catalyst layers the major part of catalyst loading did not show considerable changes concerning structure and texture. Some changes of a fraction of catalysts loading was observed, but not to the extent as
309 in upper layer, concerning both the intensity of the smtering and the ratio of the fraction changed in structure. The XPS analysis (Figure 1) of the average sample from the deeper reactor layer, with no considerable bulk properties changes, indicated fine restructuring of the active phase. The monolayer structure of molybdenum Mo(VI) phase is partially changed to multilayer structure [10, 11], while the reduced Mo(IV)/Mo(V) phase remahied stable [12]. An indication of nucleation of aluminum molybdate on the surface reveals that a part of molybdenum phase is strongly bonded to the alumina surface. These changes could be considered as an initial step in catalyst bulk restructuring, stepwise leading to loss of active component, catalyst smtering and partial hindering of active component by chemical interaction with alumina support. The changes begin on the catalyst surface and later on these processes spread through the catalyst grain causing pronounced deactivation of the catalyst.
Energy
%
235.79 232.59 234.10 230.90
30.9 44.9 9.9 14.3
242.6
Used catalyst
Fresh catalyst
224.4 243.3
225.1
Binding Energy (eV)
Figure 1. Deconvolution of Mo 3d XPS signals of NiO-Mo03/y-Al203 catalyst samples Laboratory simulation of catalyst aging in different atmospheres confirmed critical role of oxidizing atmosphere in intensive smtering and bulk restructuring. Textural and structural changes similar to that in industrially deactivated sample were observed at 800 ^C treatment in both air and steam atmosphere. A gradual shift of dominant pore size to higher values, approaching at critical temperature that of deactivated referent sample, illustrates the phenomena (Figure 2). 0.8 0.6 0.4
0.2-J
10 100 Pore diameter, (nm)
fW\ 'jM 10
— fresh — N2 air — - steam j
'300"C,9'' 100
Figure 2. Pore size distribution of MO-MOOB/Y-AIIOS catalyst samples: a) Reactor samples; b) Laboratory simulation in air atmosphere; c) Laboratory simulation in different atmospheres
310 DRS spectra indicated the change of promoter structure, with spinel formation at temperatures exceeding 700 ®C. Comparative investigation of Ni- and Co-based catalysts confirmed sHght preference of C0M0/AI2O3 concerning thermal stability. Mathematical modeling confirmed critical role of temperature concerning the intensity of smtering, indicating the change of sintering mechanism depending on temperature. The formation of fi*ee M0O3 phase, characterized by relatively low melting point (795''C), could be considered as critical for rapid sintering in oxidizmg conditions. In the same time, the formation of less reactive molybdate in the interaction of the active phase and catalyst support contribute to activity dechne. Mtensive loss of active component as an additional reason, with crystals of molybdenum oxide formed at the laboratory reactor outlet is specially pronounced in steam atmosphere.
4. CONCLUSIONS The restructuring of the original monolayer structure of active phase of the catalyst in the upper part of the reactor could be caused by formation of hot spots in reactor bed. The reducing process conditions are not critical concerning smtering, but active phase restructuring in hot spots brings about the precursors of the critical phase that initiates sintering in oxidizing conditions during regeneration. The formation of fi-ee M0O3 phase could be considered as critical for rapid sintering during regeneration. The strong interaction of the active phase and catalyst support and loss of active component contribute to activity dechne.
REFERENCES 1. KPrins, V.H.J.De Beer and G.A.Samorjai, Catal.Rev.-Sci.Eng., 31 (1989) 1 2. B.Dehnon in H.F.Lorra, P.C.H.Mitchel (eds.), Proc.3'**Int.Conf on Chem.and Uses of Molybdenum, Climax Molybdenum Co., Ann Arbor, 1979, 73 3. E.Furimsky and F.E.Massoth, CataLToday, 17 (1993) 537 4. M.Absi-Halabi, A. Stanislaus and D.L. Trimm,, Appl. Catal., 72 (1991) 193 5. RHughes, Deactivation of Catalysts, Academic Press, New York, London, 1984 6. D.S.Thakur and M.G.Thomas, ^)pLCatal., 15 (1985) 197 7. RMarinkovic-Neducin, G.Boskovic, E.Kis, G.Lomic, H.Hantsche, RMicic and P.Pavlovic, Appl. Cat, 107 (1994) 133 8. RMarinkovic-Neducin and P.Putanov, Lid.J.Eng.& MatSci., 2 (1995) 83 9. RMarinkovic-Neducin, E.Kis, J.Kiurski and RMicic, Proc.37* Int. Conf On Petroleum, Bratislava, 1995, D.43-1 10. P.Gajardo, RLDeclerck-Grimee, G.Delvaux, P.Olodo, J.M.Zabala, P.Canesson, P.Grange and B.Dehnon, J.Less-Comm.Metals, 54 (1977) 311 11. T.Edmonds and P.C.H.Mitchell, J.CataL, 64 (1980) 491 12. T.APatterson, J.C.Carver, D.E.Leyden, D.M.Hercules., J.Phys.Chem., 80 (1976) 1700
® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
CHARACTERIZATION OF AGED PETROLEUM RESIDUE.
311
CATALYST
FROM
HYDROTREATING
M.T. Martinez', J.M. Jimenez\ M.A. Callejas', FJ. G6mez^ C. Rial\ E. Carbo'. ''Instituto de Carboquimica, P.O. Box 589, Zaragoza. Spain. ^epsol Petroleo Investigacion. 28045 Madrid. ""Repsol Petroleo. Valle de Escombreras, Cartagena. Spain. A residue from Maya crude has been processed in a continuous hydroprocessing unit provided with a trickle bed reactor (TBR) at 375° C, 10 MPa H2 pressure, 1 LHSV 10.000 stdcu.ft/bbl and Topsoe TK-711 catalyst. Topsoe TK-711 catalyst life test has been performed for 7400 hours and samples of catalyst from 100, 1100, 2100, 3100, 4100, 6100 and 7400 hours on stream have been characterized by elemental analysis, S, Ni, V and Mo determination, surface area and pore volume. Aged catalyst from 100, 1100, 2100 and 3100 hours on stream were studied by scanning electron microscopy (SEM) using electron backscattering detector (BE) and microanalysis by dispersion of energies (EDS) and the distribution profiles of Ni, V, S, Mo and Fe along the diameter of the pellet have been determined. These samples have been also characterized by X-ray diffraction and the presence of sulphide phases with stoichiometry (V, Ni, Fe)~3 S4 have been found. The capacity for contaminants uptake of the Topsoe TK-711 catalyst has been of 110 g/100 g catalyst. 1. INTRODUCTION The depletion of the world's reserves of low-sulphur crude oils is leading to an increase in refinery utilisation of low-quality, heavy crudes to meet the demand for lighter, more valuable products such as motor gasoline, diesel fuel, jet fuel, and petrochemical feedstocks. These heavy crudes are characterised by sulphur contents of 1 to 6 wt. %, nitrogen of 0,1 to 0.6 wt. % and high concentrations of polyaromatic asphaltenic structures which are concentrated primarily in the 345° C + atmospheric residuum or 565° C + vacuum residuum, termed "the bottom-of-the-barrel" fraction. Upgrading the residuum from the heavy crudes must accomplish two tasks: (1) reduction of hydrocarbon molecular weight and (2) removal of detrimental non-hydrocarbon impurities including the heteroatoms N and S and the metals Ni and V that are dissolved in the oil as organometallic complexes. The technology for upgrading residuum exists today in various forms with hydroprocessing as the method of choice for maximising the yield and quality of
312 distillate range products. A major obstacle hindering the widespread utilisation of residuum hydroprocessing technology, however, has been the economic obstacle imposed the presence of Ni and V. The deposition of these metals on the hydroprocessing catalyst from hydrodemetallation (HDM) reactions, although desirable from the stand-point of product quality, accelerates the deactivation of the catalyst, particularly with residua containing more than 200 ppmw metals. The hydrodemetallation of organometallic compounds in residual oils leaves a deposit of mixed metal sulphides on and in the catalyst. The build-up of these metal deposits is irreversible and it largely determines the catalyst life. Since HDM is a diffusion-limited reaction the success of the process development and catalyst design in residuum hydrotreating is dependent on an accurate prediction of where the metal sulphides deposit and what effect the deposits have on catalyst performance. In the present paper the life test and characterization of a guard bed commercial demetallation catalyst at different stages of aging are reported. 2. EXPERIMENTAL A blend of 40% of vacuum residue (VR) and 60% of vacuum (VGO) gasoil from Maya crude has been hydroprocessed with Topsoe TK-711 catalyst. The characteristics of feedstock are indicated in Table 1. The catalyst used was Topsoe TK-711 catalyst. It is a specially developed demetallation catalyst for pre-treatment of residual oils for reduction of metals, asphaltenes and Conradson Carbon. It is utilised for guard beds, guard reactors or first stage catalyst in composite fillings. It has low HDS activity but good selectivity for demetallation, high capacity for metals uptake, high catalytic stability and high physical strength. These properties make them well suited for Table 1 Characteristics of the feedstock (40 VR/ 60 VGO) IBP - 540° c, vol. % Sulphur (%wt) Nitrogen (%wt) Nickel (ppm) Vanadium (ppm) Asphaltenes (%wt) C. Ramsbotton (%wt) Density g/ml (50° C) Cinematic viscosity, est (50° C) Dynamic viscosity, cp (50° C) B.P. > 540° C out properties Asphaltenes (% wt) C. Ramsbotton (% wt) Sulphur (%wt) Nickel (ppm) Vanadium (ppm
63.3 3.18 0.27 45.17 242.12 8.60 11.22 0.9245 111.50 103.10 15.39 30.40 4.84 142 562.40
313 front end use in high metals residue hydroprocessing and offer the necessary protection for the more active catalysts downstream in a "Composite Catalyst Filling". The Topsoe TK-711 catalyst life test was performed in TBR in the hydroprocessing unit described in (1). The catalyst was presulfured for 10 hours at 350° C and atmospheric pressure in a flowing stream of hydrogen containing 10% (v/v) of H2S. The temperature step increase until 350° was 25° C/hour. The hydroprocessing conditions were: Liquid Hourly Space Velocity (LHSV) 1 Hydrogen pressure 100 kg/cm^ Gas/Liquid ratio 10000 stdxft./bbl Catalyst Topsoe TK-711 1/20" trilobe In order to investigate intraparticule metal depositional patterns, samples of catalyst partially aged have been taken by interrupting the life-test run at 100, 1100, 2100, 3100, 4100, 6100 and 7400 hours. After 7400 hours, the test has been stopped and samples have been taken at different height from the top to the bottom of the reactor. After heptane extraction in Soxhlet for 24 hours the catalyst samples have been dried at 50° C and 61-100 mm Hg for 18 hours. Polar solvents have been avoided to minimize dissolution or redistribution of the deposits. In addition to the elemental analysis, surface area, pore volume and Ni, V, and Mo determination, the spent residue demetallation catalyst have been characterized by X-ray diffraction (XRD) and scanning electron microscopy (SEM). The difractograms were taken at 40 kvolt and 30 mamp in a Philips PW 1810 apparatus with a APD 1700 data acquisition equipment and were registered continuously in a range of 5° to 95° 26 at 0,015° 2e/s rate. The composition and elements distribution of aged catalyst were studied by scanning electron microscopy (SEM) using electron backscattering detector (BE) and microanalysis by dispersion of energies (EDS). The EDS study was performed by both "point by point" along the axis of the extrudates (distribution profiles) and by capturing composition maps. The composition maps were taken in a microscope Philips PSEM-505 provided with a EDAX-9900 at 30 kvolt. with accumulation times of 500 msg. and 500 sg for the "point by point" procedure used for the distribution profiles. The distribution profiles have been determined for the three axis, the extrudates shape is trilobe, and for two different pellets and the composition maps have been obtained for the cross section of one of them. For the distribution profiles the abscises axis corresponds to the real distance from the edge to the interior curve between two lobes, vertex. The bulk concentration determined by EDS microanalysis has been performed in a pastille prepared with several extrudated grinded and compacted and on surface. To compare the depositional patterns of the metals in cylindrical extrudate particles quantitatively the following distribution parameter were defined by Tann (2)
Mmax Jl rdr
314 where M(r) is the local concentration of metal deposit, M^ax is the concentration at the maximum and r is thefractionalradius. So 0M is the ratio of the average metal concentration in the pellet to the concentration at the maximum. Following this criterium we have calculated 9M in the following form
n
=
1
S Mi (ri
M.ax
j ; (ri
±
) Ari
-—)Ari
where Mi is the concentration measured at distance, ri from the centre of the extrudate and Ari is the width of the interval. There is a clear analogy between 9M and the classic effectiveness factor r|. Both are a measure of the relatives rates of the HDM reactions and the difiRision of the metal bearing species. The location of a metal sulphide deposit in a catalyst pellet is dependent on the relative rates of reaction and difiiision. A high distribution parameter implies that metal bearing compounds penetrate deeper into the catalyst pore system before reacting. 3. RESULTS AND DISCUSSION. In Table 2 are shown the elemental analysis, surface area, and Ni, V, Mo determination at different times of aging and the amount of contaminants, H, N, C, S, Ni and V accumulated in the catalyst calculated from the Mo percentage in fresh sample and its variation along the time.. The analysis of samples taken after 7400 hours on stream at different heights of the bed are shown in Table 3. The data of Table 2 indicate that the average contaminants absorption capacity of the TK711 catalyst is 110 g per 100 g of fresh catalyst. The amount of nickel and vanadium accumulated, 48 g per 100 g of fresh catalyst, indicates that the catalyst has a high metal retention capacity. The sulphur content seems to indicate that a big part of these metals have been deposited as sulfides. The pore volume distribution shows that after 1100 hours on stream, the catalyst pores higher than 10 nm have been totally plugged due to the constriction of the catalyst pore mouths by metal deposits and coke. After 2100 hours the catalyst has only a residual pore volume and it has reached the maximum contaminants adsorption capacity. Nevertheless from 2100 hours the content of some contaminants; S, Ni, and V increase slightly and the carbon content decreases. During the first 100 hours an important amount of carbon is accumulated in the catalyst. The accumulation goes on until 1100 hours and remains constant between 1100 and 2100 hours. After 2100 hours a decrease of coke is observed. This behaviour could also be shown in the different parts of the bed after 7400 hours of life test. The parts 1, 2 and 3 of the bed have lower carbon content than part 4 that is at the end of the life or possibly recently aged. The reversibility of coke formation has been also recognized by some other authors, (2, 3). After 7400 hours on stream only the parts 5 and 6 of the bed have active catalyst with remaining metal absorption capacity.
Table 2 Properties of fresh, presulphured and used Topsoe TK-711 catalyst at different stages of aging. Catal./Hours on stream Surface area m2/gr. Pore volumen m3/gr. Pore distrib. < 1onm 10-50 nm 50 nm Elemental Analysis N Yo w
c
"
H
"
s
"
Metals Ni % w Mo "
v
"
Contaminants/l OOg fresh catalyst N grs/100 cat. C grs/100 cat. H "
s
Ni Mo
v
" " " "
Total g/lOO.Cat. Total uptaken (g)
Fresh
Presulf
140 0.60
100 Outlet 120 0.336
1100
2100
3100
4100
6100
145 0.54
100 Inlet 116 0.33 1
94 0.12
0.16
0.023
16 0.00
11 0.00
7400 Inlet 23 0.023
0.02 0.54 0.04
0.01 0.50 0.03
0.07 0.20 0.06
0.06 0.25 0.03
0.12 0.00 0.00
0.07 0.00 0.09
0.023 0.00 0.00
---
----
0.00 0.00 0.023
--
--
0.02 0.44 0.09
0.17 0.76 3.80
0.34 12.39 1.36 3.62
0.38 12.03 1.46 3.65
0.34 13.11 1.30 10.05
0.26 11.22 0.84 19.52
0.19 6.5 0.56 8.71
0.17 7.53 0.82 22.11
0.1 1 7.59 0.80 23.90
0.11 7.37 0.84 22.56
1.43 3.95 163 ppm
1.37 5.29 75 pprn
1.13 2.99 0.78
1.05 2.86 0.73
2.07 2.18 11.82
2.91 1.86 19.25
3.16 1.83 23.19
3.38 2.05 21.76
3.72 2.02 21.88
3.62 1.97 21.06
--
--
0.02 0.44 0.09 1.43 3.95 1.63 ppm 5.93
0.19 0.85 4.25 1.37 3.52 7.5 10.18 4.25
0.45 16.36 1.79 4.78 1.49 3.05 1.03 29.85 23.92
0.52 16.61 2.02 5.04 1.45 3.95 1.oo 30.59 24.66
0.62 23.75 2.35 18.21 3.75 3.95 21.41 74.04 68.11
0.55 23.82 1.78 41.44 6.18 3.95 40.86 118.5 112.57
0.41 14.02 1.20 46.13 6.82 3.95 41.54 114.07 108.14
0.32 14.53 1.58 42.67 6.52 3.95 41.99 111.56 105.63
0.22 14.87 1.57 46.84 7.29 3.95 42.88 117.62 111.69
0.22 14.74 1.68 45.12 7.24 3.95 42.12 115.07 109.14
--
-
VI
316 It could be observed from Table 3 higher metal absorption capacity from the top to the bottom of the aged parts of the bed. This may be relationed with an improvement of the diffusion of the molecules through the catalyst down the bed that would produce a better catalyst utilization. The difractograms corresponding to the samples from 100 to 3100 hours on stream are shown in Figure 1. The sample from 100 hours of catalyst life only shows a single crystalline phase corresponding to 5-AI2O3. In the difractogram corresponding to the samples from 1100 to 3100 hours, in addition to 5-AI2O3 it is possible to identify other crystalline phases, these phases have been identified as vanadium sulfides that may contain Ni and/or Fe, in a stoichiometry (V, Ni, Fe)3S4. In Figures 2, 3 are shown the V5.45S8 (JCPDS n° 29 - 1382), V2.14 Feo.75 S4 (JCPDS n° 31-656) and V2NiS4 (JCPDS n° 36-1132) diagrams corresponding to the sample from 2100 hours. These stoichiometrics could not correspond exactly to the existing phases but they will be similars since all of them are isostructurals (space group hm) and the parameters of the unit cell are very similars (a = 5.83 - 5.87 A°; b = 3.26 - 3.30 A°; c = 11.22 - 11.33 A° due to the similitude of the ionic radius of the present cations. The difractograms from 1100 to 3100 hours of life show an increase in the sulfides proportion with a decrease of the amount of the 5-AI2O3 present. Takeuchi et al. (4) have revealed the presence of the crystalline V3S4 phase, a nonstroichiometric, polycrystalline solid with sulphur to vanadium ratios of 1.2 to 1.5. l^JUL-96
TK-711
9:47
solo
Figure 1. XRD difractograms corresponding to the samples after, 100, 1100, 2100 and 3100 hours on stream.
317 Table 3 TK-711 catalyst properties after 7400 hours on stream at different heights of the bed. Catal./Hours on stream Elemental Analysis N%w C" H"
s"
Metals Ni % w Mo" V
7400 (1) Inlet.
7400 (2)
7400 (3)
7400 (4) 7400 (5)
7400 (6) Outlet
0.11 7.37 0.84 22.56
0.11 7.17 0.83 23.17
0.11 6.82 0.80 22.38
0.12 8.20 0.96 21.63
0.13 6.55 0.78 20.78
0.17 6.50 0.79 20.33
3.62 1.97 21.06
3.59 1.85 20.91
3.58 1.79 19.72
3.41 1.89 19.26
3.35 2.06 18.08
3.24 1.99 16.62
0.22 14.74 1.68 45.12 7.24 3.95 42.12 115.07 109.14
0.23 15.31 1.77 49.47 7.66 3.95 44.64 123.03 117.10
0.24 15.04 1.76 49.38 7.90 3.95 43.52 121.79 115.86
0.25 17.13 2.00 45.20 7.12 3.95 40.25 115.90 109.97
0.25 12.56 1.50 39.84 6.42 3.95 34.67 99.19 93.26
0.34 12.90 1.57 40.35 6.43 3.95 32.99 98.53 92.60
Contaminants/1 OOg fresh catalyst Ngrs/lOOcat. C grs/100 cat. H " S " Ni " Mo " V " Total g/lOO.Cat. Total uptaken (g)
Silbemagel (5) has also suggested that at loading above 10 wt% V, the metals appear to be deposited as growing sulphide crystals with a stoichiometry close to that of V2S3. The distribution parameters 0M calculated by SEM-EDS are indicated in Table 4 and in Figures 4, 5 typical deposition profiles for Ni and V. For vanadium, an increasing concentration from the edge to the inside until a maximum followed by a decreasing concentration is observed. The BSE images (6) showed that this gradient is extended parallely to the surface. The BSE images (6) and the distribution profiles Figure 5 show the Ni in the 1100 hours sample with an increasing concentration gradient towards the centre of the pellet and some point of high concentration near the surface. Intrapellet metal profiles often provide evidence of a distribution of metal compounds reactivities. Metal species, sometimes referred to as easy metals, are readily removed (low effectiveness factor), resulting in sharp metal gradients near the pellet's edge at the reactor entrance. Metalloporphyrins in the resing fraction may be typical of such species. The more
318 ^jQ 2 I SaM)le: TK-711 2100 H (9) S.OO 4.SO 4.00 3.50 3.00 2.50 2.00 1.50 1.00 0.50
FUe; 2100H9.RD
19-JUL-96
09:19
U^w^^^
0 100.0 80.0 60.0 J 40.0 20.0 J
20.0
30.0
40.0
50.0
60.0
70.0
80.0
90.0 V5.45S8 29-1382
]
10.0 100.0 80.0 60.0 40.0 20.0 ] 10.0
1 1
20.0
u
20.0
1 .ii
i.
. 1 . 1
• 1
30.or
40.0
... 1150.0, . h 60.0
, .1
1
,
70.0
1i-H
, —
80.0
90.0 A1203 4 - 877
30.0
III I I
40.0
J
L
50.0
60.0
70.0
80.0
90.0
Figure 2. XRD diagrams of the sample from 7100 hours on stream and JCPDS diagrams
xlO
2 I Samle:
10.0
TK-711 2100 H (9)
20.0
30.0
file:
40.0
2100H9.RD
50.0
19-JUL-96
60.0
70.0
09;24
80.0
90.0
Figure 3.- XRD diagrams of the sample from 2100 hours on stream and JCPDS diagram corresponding to V2NiS4 and V2.14Feo.75S4 phases.
319
0.4
0.6
DISTANCE
Figure 4. Vanadium deposition profiles of the sample from 2100 hours on stream.
1
.: _
'——^—
AXIS 1 AXIS 2
Z
C
>( AXIS 3 :"^
> v ^
--
•
z 2:
c
I ?^j
-1/
rf
i
0
i^il
i. -j ^
^
>
•
•
•
i
•
o
H 0.2
0.4
0.6
0.8
DISTANCE
Figure 5. Nickel deposition profiles of the sample from 2100 hours on stream.
320 Table Table 4. 4. Distribution parameters of S, Ni, V, Mo and Fe for the catalyst at different stages of aging. Hours/ Element Sulphur 0 vertex 0 edge 0 average Vanadium 0 vertex 0 edge 0 average Iron 0 vertex 0 edge 0 average Nickel 0 vertex 0 edge 0 average Molibdeno 0 vertex 0 edge 0 average
100 Inlet
100 Outlet
1100
2100
3100
0.8826 0.8146 0.8486
0.8484 0.7264 0.7874
0.6305 0.7171 0.6738
0.7370 0.6557 0.6963
0.7262 0.7199 0.7230
0.7923 0.6684 0.7303
0.8205 0.7158 0.7681
0.7967 0.7884 0.7925
0.7930 0.8209 0.8069
0.8721 0.8014 0.8367
0.4181 0.2882 0.3531
0.3126 0.2824 0.2975
0.1140 0.0579 0.0859
0.0974 0.0652 0.0813
0.1086 0.0772 0.0929
0.8400 0.8199 0.8230
0.7888 0.7360 0.7624
0.8238 0.8135 0.8186
0.8081 0.7871 0.7976
0.7202 0.7120 0.7161
0.8345 0.8021 0.8183
0.8204 0.7812 0.8008
0.6641 0.5871 0.6256
0.5700 0.6788 0.6444
0.5739 0.4428 0.5083
refractory, hard compounds yield more uniform profiles (higher effectiveness factor). Metal profiles indicating the presence of easy and hard metal species have been reported for V (7) From 0M values a number of interesting features are apparent. The distribution parameters decline with the time for Ni but increase for V. Nickel in samplesfi-om100 hours on stream shows a decrease in the distribution parameterfi-omthe inlet to the outlet of the reactor and vanadium shows the reverse behaviour. Average 0M is similar for V and Ni at the reactor outlet but at the reactor inlet is much higher for Ni. 4. CONCLUSIONS TK-711 catalyst has shown to have a high contaminants absorption capacity, 110 g per 100 g of fi-esh catalyst and a high metal retention capacity 48 g of metals per 100 gfreshcatalyst. Crystalline phases of metal sulfides V5.45S8, V2.i4 Feo.75 S4 and V2NiS4 have been identified by XRD. Metal deposition profiles of vanadium show M-shaped perfiles while nickel deposition profiles show an increasing concentration gradient towards the centre of the pellet.
321 Nickel distribution parameter 6M in samples from 100 hours on stream shows a decrease from the inlet to the outlet of the reactor and vanadium shows the reverse behaviour. ACKNOWLEDGEMENT The authors v^sh to thank the European Community (Contract JOU2-CT92-0206) and the DGICYT (AMB93-1137-CE) forfinancialsupport. REFERENCES 1. M.T. Martinez, I. Fernandez, A.M. Benito, V.L. Cebolla, J.L. Miranda and H.H. Oelert, Fuel Processing Technology, 33, (1993), 159. 2. P.W.Tamm, H.F. Harnsberger, A.G. Bridge, Ind. Eng. Chem. Process Des. Dev., 20 (1981) 262. 3. H. Nitta, T. Takatsuka, S. Kodama and T. Yokoyam^, Deactivation model for Residual Hydrodesulfiirization Catalyst, 86th National AIChE Meeting, Houston Texas, April (1979). 4. C.Takeuchi, S. Asaoka, S. Nakata, and Y.Shiroto, ACS Prepr. Div. Petrol. Chem. 30 (1985) 96. 5. B.G. Silbernagel, J. Catal. 56 (1979) 315. 6. M.T. Martinez, Final Report Contract JOU2-CT92-0206, September ((1996). 7. H. Higashi, K. Shirono, G. Sato, Y. Nishimura and S. Egashira, ACS Prepr. Div. Petrol. Chem. 30 (1985) 111.
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© 7P97 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
323
Hydrotreatment of spent lube oil: Catalysts and Reactor Performance C. Yiokari^ S. Morphi^ A. Siokou^ F. Satra^ S.Bebelis^ and C.G. Vayenas^ C. Karavassilis^ and G. Deligiorgis^ ^Department of Chemical Engineering, University of Patras, GR-26500 Patras, Greece ^LPC Hellas, GR-15125 Athens, Greece The hydrotreatment of various spent lube oil fractions (light distillate) was investigated in a batch reactor at 880 psi and 330°C using a variety of laboratory prepared and commercial Ni/Mo, Co/Mo, Ni/W and Co/W catalysts supported on y-AbOs , SiOi, Zr02 or ZrOi (8% Y2O3) (YSZ) extrudates. The reactor operating time was chosen to be similar to the residence time of the commercial unit of the refinery of LPC Hellas, S.A.. The hydrodesulfrirization activity of the home-made catalysts is within 10% of the best commercial ones. The quality (viscosity index) of the product is also very good. The catalysts are compared regarding their relative hydrodesulfurization performance as well as the hydrogen consumption in parallel hydrogenolysis reactions.
1. INTRODUCTION One of the key processes in the recycling of lube oils is desulfurization, as lube oils for recycling contain high contents (0.5-1.5%) of sulfur. The removal of sulfur is dictated both by environmental concerns and by the requirements for upgrading of the lube oil, as during the desulfurization nitrogen removal, saturation of aromatics and olefins and hydrocracking is also taking place, leading to products with properties within the specifications required by the market. Many hydrodesulfurization (HDS) processes have been developed [1], while most studies on hydrodesulfurization concern model sulfur compounds such as thiophene, benzothiophene and dibenzothiophene [2-4] as well as of their alkyl-substituted derivatives [5-7]. Fewer studies on hydrodesulfurization have been carried out using feedstocks derived from refmeries [8,9] where the reaction environment is much more complex and coexisting aromatic species as well as various types of sulfur compounds compete for the same active sites on the HDS catalysts surface [8,9]. This results in different reactivities of various sulfur compounds in the feedstock compared to the corresponding ones in the case of model compounds in a pure solvent and also to hydrogen consumption in excess of the stoichiometric requirements of sulfur removal as hydrogen sulfide. The latter increases for feedstocks with high levels of complex sulfur compounds.
324 The objective of this study was to test a variety of Mo-Ni, Mo-Co, W-Ni and W-Co catalysts on different supports (y-Al203, Si02, Zr02 and YSZ) regarding their abiUty for hydrodesulfurization of vacuum distillates obtained in the process of recycling of lube oils. A commercial NiMo/y-Al203, catalyst was also tested for comparison. The above catalysts covered a wide range of active phases and supports as most of the hydrodesulfurization studies in the literature concern molybdenum or tungsten based catalysts promoted by nickel or cobalt and supported on y-Al203 or Si02 [10-12], while few reports concern W-Ni/yAI2O3, [10] or W-Ni/y-Al203 [10]. In addition to testing the above catalysts regarding their hydrodesulfurization activity, an estimation of the corresponding relative hydrogen consumption in hydrogenolysis and hydrodesulfurization reactions is also reported.
2. EXPERIIVIENTAL 2.1 Feed The lube oil feed used in the present investigation was a vacuum distillate obtained from L.P.C. Hellas with a sulfur content of 0.615% wt. Its properties are summarized in Table 1. Table 1 Properties of lube oil total sulfur (wt%) density(15°C)(g/ml) pour point C Q viscosity index
0.615 0.86 -6 102
2.2 Catalysts The catalysts that were used covered a wide range of homemade Ni/Mo, Co/Mo, Ni/W and CoAV catalysts supported on y-Al203, Zr02, Zr02 (8 mol % Y2O3) or Si02 extrudates. The catalysts were prepared by succesive incipient wetness impregnation of supports with solutions of (NH4)Mo7024-4H20, Co(N03)2H20, (NH4)io(Wi204i)-5H20 and Ni(N03)2-6H20 in distilled water. The two-phase systems were dried in a rotary evaporator at 50 °C under reduced pressure and then were succesively calcined at 110 °C overnight and at 500 °C for 5h. The homemade catalysts were compared with the commercial Ni/Mo catalyst supported on y-A^Os which is used by L.P.C Hellas and supplied by AKZO. 2.3 Presulfation and hydrotreatment Both the presulfation of the catalysts and the hydrogenation of the lube oil took place in the experimental setup shown in Figure 1. By proper adjustment of the valves either presulfation of the catalysts or hydrogenation of the lube oil is carried out.
325
> vent
H2 S scrubber
He N, H2S Ho Figure 1 : Experimental setup During presulfation the reactor was operated as a continuous flow reactor since the valves at the entrance and exit of the reactor are kept open. The presulfation procedure comprised of the following steps: Initially a stream of He was passed through the reactor at atmospheric pressure, while the reactor was heated to 360 °C at a heating rate of 5.5 °C/min. The flow of He was continued at 360 °C for Ih and then a stream of 15% H2S and 85% H2 (vol.%) was supplied to the reactor under atmospheric pressure. The presulfation time was adjusted accordingly so that the total amount of H2S passed over the catalyst contained approximately 6 times more sulfur than the stoichiometric amount needed for the sulfation of the Mo, W, Co and Ni oxides. After completing the presulfation, a flow of He was passed through the reactor for Ih at 360°C and then the reactor was cooled to room temperature at a cooling rate of 4°C/min. The hydrogenation of the lube oil was carried out in the same 300ml reactor operating as a stirred batch autoclave (Fig. 1). The catalyst to oil mass ratio in the mixture charged in the reactor was equal to 1:3. The batch reactor was pressurized with N2 and tested for leaks before each run. Each run consisted of the following steps: Pressurization of the reactor with hydrogen at room temperature up to about 600psi, setting of the stirrer speed at 120rpm, heating of the reaction mixture to 330 °C at a rate equal to 6 °C/min and fixing of the pressure in the reactor ( 850 to 950 psi) to 880 psi by H2 addition or release to or from the reactor. The reaction time was counted from the moment the temperature in the reactor had reached 330°C and the pressure was fixed to 880psi. The reactor operating time was chosen to simulate the residence time of the reactants in the industrial desulfurization reactor of the refinery of L.P.C Hellas and was equal to 3h. The consumption of the gaseous hydrogen
326 during the reaction was followed by monitoring the total pressure (Figure 2) using a pressure gauge. At the end of the run the reactor was cooled to room temperature, the gas in the reactor was removed by opening a needle valve at the exit of the reactor and the reactor was purged with He to remove the remaing and dissolved gaseous products. The liquid product was filtered to assure complete removal of suspended fines. 2.4 Analysis The sulfur content of the feed and product oils was determined using an ASOMA sulfur analyzer (200T-series) based on X-ray fluorescence analysis. The viscocity index was measured according to the ASTM D-2270 method.
3. RESULTS Since there is a variety of complex sulfur compounds in lube oils, the mechanism of the HDS reaction is quite complex . The HDS reaction that takes place can be written in a simple form as RHS + H2^H2S+R-H
(1)
For small variation in PH2 the global hydrodesulfiirization kinetics can be approximated by: dCs •=Kndt
Cs"
(2)
where Cs is the total sulfur concentration, Kn is a rate constant and n the reaction order with respect to the sulfur compounds. If n=l, integrating eq. 2 results in hi(l-xs) = Ki -t
(3)
while if n=2, integrating eq.2,results in ""^ - = K , . C o t 1-
(4)
where xs is the sulfur conversion,Co is the initial sulfur content and Ki, K2 are rate constants for first and second order HDS reaction respectively. Figure 3 shows the dependence of the sulfur concentration on reaction time when the reaction time is varied between 0 and 360 min in the case of a Mo/Ni on y-A^Os catalyst. As shown in Figures 4a and 4b the discrimination between first and second order kinetics is not easy as the corresponding fitting is equally good in both cases. As shown in Figure 4 at reaction time equal to zero the sulfur conversion, based on the sulfur concentration in the feed lube oil before heating the reactor, is not zero as during the heating period preceding the reaction time t a significant amount (almost 55%) of the sulfur is removed because the HDS
327
reaction already starts at temperatures and pressures lower than 330°C and 880psi respectively. In order to take this into account the reaction time, t, was corrected by adding the value, -t*, where t* is the time difference between the time corresponding to Xs=0 and t=0. The value of-t* is equal to 16 min regardless of the exact kinetics (Fig.4). The corrected time tc = t -1* is used throughout the present work for catalyst comparison. 900 850 800h
i-
750
0.08.
O 15%Mo03-4.5%NiO/y-Ab03 15% M0O3 - 4.5% CoO / Zr02 |- •15%W03-4.5%CoO/YSZ
700 h 650
J
40
I
I
1
I
80 120 t,inin
I
L
160
200
250
Figure 2: Dependence of pressure on reaction Figure 3: Dependence of sulfiir concentration time for three different catalysts on reaction time t and corrected reaction time tc for catalyst 15%Mo03-4.5%NiO/ y- AI2O3 (prepared) ; see text for discussion.
B
-2H
250
Figure 4a: Dependence of sulfur conversion on reaction time. The solid line corresponds to fitting according to first order kinetics (Eq. 3)
Figure 4b : Dependence of sulfur conversion on reaction time. The solid line corresponds to fitting according to second order kinetics (Eq. 4).
328 The H2 pressure in the reactor was not kept constant during the reaction time so the H2 consumption due to the hydrogenation and hydrogenolysis reactions was followed by monitoring the pressure in the reactor. Figure 5 shows the pressure corresponding to the H2 consumed by both the hydrogenation, hydrogenolysis and the HDS reactions. The latter can be calculated by measuring the sulfur conversion corresponding to each reaction time tc. The figure also shows the H2 pressure corresponding to the H2 dissolved in the lube oil. The dissolution of H2 in the lube oil was practically completed during the heating period preceding the reaction. The quantity of H2 which is dissolved in the lube oil was estimated with blank experiments, where no catalyst was used, from the difference between the measured pressure in the reactor and the pressure calculated from the ideal gas law if no H2 were dissolved in the lube oil. This difference corresponds to a H2 pressure equal to 200psi. The H2 consumed for hydrogenation-hydrogenolysis during the heating period was calculated from the difference between the pressure in the reactor in the presence of catalyst and the pressure expected from the ideal gas equation, after subtracting the H2 pressure corresponding to the amount of dissolved hydrogen. t,inm 100
150
100 150 t^ ,min
200
250
Figure 5: Dependence of the hydrogen pressure corresponding to the HDS and hydrogenolysis reactions and to hydrogen dissolution in the oil on reaction time for the catalyst 15%Mo03 - 4.5%NiO / y-Al203. A useful parameter for comparing the HDS catalysts regarding the hydrogen consumption in the hydrogenation-hydrogenolysis reactions is the ratio a defined from: a=
mole of H2 consumed due to hydrogenation and hydrogenolysis mole of H2 consumed due to HDS
(5)
329 Figure 6 shows the dependence of a on the corrected reaction time time tc for the Mo/Ni on Y-AI2O3 catalyst as calculated from the data in Figure 5. It is clear that for long reaction times the extent of the hydrogenation and hydrogenolysis reactions becomes more significant. Li a similar way the value of the parameter a for all catalysts has been calculated for a reaction time t equal to 180 min and is presented in Table 2. In the same Table we present the values of the BET surface area of the catalysts, determined by N2 adsorption at -196 °C, the viscocity index of the final product and the % conversion of the sulfiir compounds . Figure 7 shows the sulfur conversion when using the laboratory prepared Mo/Ni on y-AliOs catalyst versus LHSV"\ The LHSV (Liquid Hour Space Velocity) is calculated from the equation: LHSV = gr oil fed / (gr catalyst reaction time in hr)
(6)
The same figure compares all catalysts tested, regarding their HDS activity by comparing the sulfiir conversion at LHSV"^=1.26 h.
100 150 t^,min
Figure 6 : Dependence of the parameter a ( Eq. 5 ) on corrected reaction time for the lab prepared catalyst 15%Mo03-4.5%NiO/Y-Al203 and comparison of the a values for the various catalysts for tc equal to 196 min (LHSV •^= 1.26 h). Symbols designatmg the various catalysts are defmed in Table 2.
0.0
0.4
0.8
1.2
( gr oil feed / gr catalyst h)'^ Figure 7: Dependence of the sulfiir conversion on LHSV "^ for the lab-prepared catalyst 15%Mo03-4.5%NiO/Y-Al203 and comparison of the sulfiir conversion values for the various catalysts for LHSV •^= 1.26 h . Symbols designating the various catalysts are defined in Table 2.
4. DISCUSSION Among the various catalysts tested the one corresponding to the largest sulfiir conversion for LHSV"^=1.26 h is the commercial catalyst 15% Mo03-4.5%NiO on Y-AI2O3.
330 The homemade 15% Mo03-4.5%NiO catalyst on Y-AI2O3 results in comparable sulfur conversions. The catalysts that are supported on Y-AI2O3 exhibit the highest performance for HDS and the ones supported on YSZ the lowest. The catalysts supported on Zr02 present also good performance for HDS conversion compared to the other catalysts. Molybdenum containing catalysts give higher conversions than those containing tungsten, while nickel is a better promoter than cobalt. These comparisons, however, are based on the mass of the active component and not on their BET surface area or, better, on their active surface area. The latter can be measured via NO chemisorption and a comparison of all tested catalysts on the basis of their active surface area will be published elsewhere. Regarding hydrogen consumption the smaller a factors for LHSV"^=1.26 h correspond to the 15% W03-4.5%NiO on Y-AI2O3 catalyst. This catalyst also gives a satisfactory sulfur conversion and a very high viscosity index (Table 2). Further work is currently underway in order to optimize the catalyst composition for enhanced reactivity, HDS selectivity and final product quality. Table 2 Catalyst BET surface area. Viscosity index, parameter a (Eq. 5) and % sulfur conversion for LHSV"^=1.26 h for the various catalysts tested. Catalysts
(El) 15%Mo03-4.5%NiO/Y-Al203 (commercial) (•)15%Mo03-4.5%NiO/Y-Al203 (A)l 5%MO03-4.5%COO/YA1203
(Q)15%W03-4.5%NiO/Y-Al203 (V)l 5%W03-4.5%CoO/Y-Al203 (•)15%Mo03-4.5%NiO/Si02 (A)l 5%Mo03-4.5%CoO/Si02 ( ^ )15%W03-4.5%NiO/Si02 (4^)15%W03-4.5%CoO/Si02 (•)15%Mo03-4.5%NiO/Zr02 (1^)15%Mo03-4.5%CoO/Zr02 (A) 15%W03-4.5%NiO/Zr02 (O) 15%W03-4.5%CoO/Zr02 (+)15%Mo03-4.5%NiOA^SZ (n)15%Mo03-4.5%CoOA^SZ (0)15%W03-4.5%NiOA^SZ (e)15%W03-4.5%CoOA^SZ
BET surface (m^/g)
Viscocity Index
a
135.3 203.9 244.4 183.7 212.9 214.3 256.5 228.2 231.2 46.4 42.7 48.7 41.9 15.1 13.8 11.3 10.4
111 114 103 132 109 106 110 108 110 103 109 106 105 106 108 108 112
6.95 6.92 7.04 2.71 7.35 4.55 4.10 4.25 6.29 6.12 6.30 3.82 4.23 5.27 5.03 5.04 5.15
% Sulfur Com
97.22 96.55 87.53 77.50 52.03 67.52 48.81 44.85 19.60 84.45 81.70 51.65 52.37 48.06 42.52 52.23 27.27
331
5. REFERENCES 1. H. Topsoe, B.S. Clausen, N.Y. Topsoe, E. Pedersen, Ind. Eng. Chem. Fundam. 25 (1986) 25 2. G.H. Singhal, R.L. Espino, J.E. Sobel, G.A. Huff Jr., J. Catal. 67 (1981) 457 3. T.C. Ho, J. Sobel, J. Catal. 128 (1991) 581 4. Y. Miki, Y. Sugimoyo, S. Yamadaya, J.Jpn.Pet. Inst. 35 (1992) 332 5. M. Honalla, D.H. Broderick, A.V. Sapre, N.K. Nag, V.H.J. De Beer, B.C. Gates, H. Kwart, J. Catal. 61 (1990) 523 6. D.R. Kilanowski, H. Teeuwen, V.H.J. De Beer, B.C. Gates, G.C.A. Schuit, H. Kwart, J. Catal. 55 (1978)129 7. Y. Miki, Y. Sugimoyo, S. Yamadaya, J.Jpn.Pet. Inst. 36 (1993) 32 8. Yu-Wen Chen, Wen-Chang Hsu, Chang-Shiang Lin, Ben-Chang Kang, Shwu-Tzy Wu, Li-Jen Leu, Jung-Chung Wu, Ind. Eng. Chem. Res. 29 (1990) 1830 9. X. Ma, K. Sakanishi, and I. Mochida, hid. Eng. Chem. Res., 35 (1996) 2487 10. R. Prms, V.H.J. De Beer and G. A. Somorjai, Catal. Rev. - Sci. Eng. 31(1&2) (1989) 1 11. G. C. A. Schuit, and B. C. Gates, AIChE J. 19(3) (1973) 417 12. J. Lame, J.L. Brito and F. Severino, J. Catal. 131 (1991) 385
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® 1997 Elsevier Science B. V, All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
333
Catalytic Hydrodesulfurization of Petroleum Middle Distillate and Model Sulfur Compounds over a Series of Catalysts Activity and Scheme Emmanuel LECRENAY and Isao MOCHIDA Institute of Advanced Material Study, Kyushu University, Kasugakouen 6-1, Kasuga, Fukuoka 816, JAPAN - FAX : (81) 92 583 7798 1. S U M M A R Y Hydrodesulfurization (HDS) of Gas Oil (CO.), Light Cycle Oil (L.C.O.) and model sulfur compound 4,6-dimethyldibenzothiophene (4,6-DMDBT) was compared over a series of 6 commercial catalysts and laboratory made catalysts. The difference between the HDS activity of the model compound and the real feed stocks indicates the sensitivity of the catalyst to the intrinsic inhibitors (aromatics and H2S) present under industrial conditions. Detailed analysis of the HDS products from the model feed provided information on the reaction pathways involved over the present catalysts and how they are modified in presence of inhibitors. Complementary standard tests with other model compounds were examined to determine the Bronsted acidity and hydrogenation activity of each catalyst. Such activities explain the differences in performance of the catalysts and the relative importance of the respective reaction pathways in the HDS.
2. EXPERIMENTAL 2.1. Apparatus The catalytic HDS tests were performed in a 100 cm^ batch autoclave equipped with a two inclined-blades mechanically driven turbine (1000 r.p.m.). 2.2. Materials 4,6-DMDBT of 99.7%w purity was synthesized according to improved Gerdil&Lucken^D's method. This model compound was dissolved in n-decane, inert non-inhibiting solvent.
334 A Straight Run gas oil (1.19 %wS, ASTM D86 5-50-95 %w : 242-295-373*C, 31 %w Arom. (S.M.)) obtained from Iranian heavy crude oil in a European refinery was used as a real feed stock for HDS tests. Also, a L.C.O. cut from a Japanese Fluid Catalytic Cracking (F.C.C.) unit (0.18 %wS, ASTM D86 5-50-95 %w : 195-258-311°C, 73 %w Arom. (H.P.L.C.)) was used as a high aromatic content real feed stock. 6 commercial catalysts were supplied by different manufacturers. The available properties of these catalysts are summarized table 1. Tab. 1 : properties of commercial catalysts Catalyst Name A B C
D
E
F
Low Content NiMo AI2O3
Deep HDS CoMo
General Properties
Usual CoMo
Usual NiMo
Support
AI2O3
AI2O3
AI2O3 + zeolite
Surface area (m2/g)
268
273
220
140
220
Pore Volume (mVg)
0.53
0.52
0.44
0.60
0.45
CoO (NiO) %w
4.2
3.1
4.0
2.0
4.1
M o 0 3 %w
15.0
15.0
16.0
6.0
20.5
Deep HDS Deep HDS CoMo CoMo Amorph. Si02AI2O3
AI2O3
Also, some laboratory made Ni-Co-Mo catalysts were prepared according to the following procedure : AI2O3 —> Mo impregnation —> 130**C drying —> Ni+Co impregnation —> 130*C drying —> 500*C calcination. Prior to HDS reactions, each catalyst was presulfided for 2 h at 360*C under an atmospheric pressure of H2/H2S (molar ratio 0.95/0.05). 2.3. Analysis Sulfur species in the gas oil were analyzed by GC-FPD. HDS products of 4,6DMDBT were analyzed qualitatively by GC-MS and quantitatively by GC-FID equipped with a silicone capillary column (OVIOI; 0.25 mm*50 m). 2.4. HDS reaction conditions HDS of 4,6-DMDBT (0.01 g) in decane (10 g) was conducted for 0-60 min. at 270-360°C under 2.4-5 MPa (initial hydrogen-pressure at room temperature) over 1.0 g of presulfided catalyst. HDS with gas oil or L.C.O. (10 g) was performed for 060 min at 340^*0 under 2.4 MPa (initial hydrogen-pressure) over 1.0 g of
336 According to a series of HDS tests of 4,6-DMDBT in solvent over Ni-Co-Mo catalysts, it was confirmed that the hydrogenation activity determined by the above standard test using naphthalene has the same ranking as the hydrogenation activity for 4,6-DMDBT (first step for its HDS hydrogenation route, see reaction scheme fig. 7). Interestingly, the relative ranking of the hydrogenation activities among the Co catalysts was found in the same order to that of their acidic activities, suggesting a possible relationship between these properties. Table 2 summarizes the activities of CoMo catalysts with more amount of acid additive in the support and a spent C catalyst. The more amount of acidic additive added in the support of catalysts C and D increased their hydrogenation activity as well as the cracking activity. Tab. 2 : Cracking (Bronsted Acidity) and Hydrogenation Activities for catalysts
Catalyst C Add Additive Added C Spent C D Add Additive Added D
k (cracking) min-^' g-^ 1.97E-02 3.56E-02 9.12E-07 4.80E-05 1.54E-04
k (hydrogenation) min-^. g-^ 4.98E-03 9.10E-03 2.18E-03 4.86E-03 6.90E-03
The spent C catalyst lost considerably hydrogenation activity (1 / 2.3) and dramatically cracking activity (1 / 22 000) compared with the virgin one, inducing a different HDS routes contribution over both catalysts. Although acidity and hydrogenation activities are supposed to be linked, this relationship is not directly proportional, suggesting a non-direct participation of the acidity to the hydrogenation activity. 3.2. HDS Activity for gas oil and 4,6-DMDBT HDS tests of gas oil gave the performance ranking of A-F catalysts according to the Total Sulfur Content (fig.3) C > D > F > B > E > A , the pseudofirst order rate constant of 4,6-DMDBT (also 4-MDBT, DBT) in gas oil C > D > F > B > E = A (fig. 6b - fig. 6d). The similarity of these rankings shows that the best catalyst is most efficient for removal of refractory sulfur species. The ranking obtained with 4,6-DMDBT in decane (fig.3) is similar to that of Total Sulfur Content in gas oil ; but, definitely, NiMo catalysts exhibited a much higher activity for 4,6-DMDBT in decane than for gas oil. The inhibitors present in gas oil (aromatics, nitrogen compounds, ...) and produced by the desulfurization (H2S) retarded more severely the HDS reaction over NiMo catalysts than over CoMo catalysts. The most active C catalyst desulfurized essentially through the isomerization and cracking reactions, both acid catalyzed routes, giving a number of products. Among them, mono and poly alkylated DBTs (not desulfurized
337 products), mono and polyalkylated biphenyls and hydrogenated derivatives (desulfurized products)(2), cracked mono-ring products (benzene, toluene, xylenes, cyclohexane, methylcyclohexanes ...)^^^ were detected. k(min-1 g-1)
4,6-DMDBT in dec.
0.025
k.a (min-1 g-1) Tots in gas oil
0.1
0.08 0.06
0.015
n H
4,6-DMDBT in decane Tots in gas oil
fig.3 : HDS activity
0.04 ^^^ 9^^ ^'' (pseudo-1.65 0 02 oi'der) and for 4,6DMDBT (first 0 order) in decane A D r n f f ^^^^ constants NiMo B had a high activity through its high hydrogenation route, giving a higher ratio of (HYD/DirectDesuif.)route products at 270^ = 12 compared to 4 for CoMo D and F catalysts. Hence, the activity difference between catalysts D and F is ascribed to their different number of active sites. For spent C catalyst, the former ratio is about 1.5. At higher temperature, the DirectDesuif. route is much favored, hence this ratio decreased to 0.3 at 360°C over CoMo and NiMo catalysts. 3.3. Aromatic and H2S Inhibitors in HDS of 4,6-DMDBT in decane * Naphthalene, tetralin and isobutylbenzene (0-40%w in decane) were used as aromatic model inhibitors present in the real feed. Such aromatics inhibited the HDS in the order of Naphthalene > Tetralin > Isobutylbenzene as shown in figure 4. The HDS products showed that the hydrogenation route was more affected than the direct desulfurization route on NiMo catalyst. Nevertheless, the extent of inhibition by aromatic partners appeared similar regardless of catalysts. Another inhibitor in gas oil may deactivate NiMo catalyst more compared to CoMo catalysts. According to a series of HDS tests of 4,6-DMDBT in decane+naphthalene over some Ni-Co-Mo catalysts, it was found that the rate constant at the hydrogenolysis step (the C-S bond breaking of the tetra-hydrogenated 4,6DMDBT, see fig. 7) is proportional to the %Co on the catalyst, for a total metal loading %Co+%Ni=3%w. The acidic catalyst C is very strongly inhibited by naphthalene which eliminated its major route, that of the cracking, giving then the classical 4,6DMDBT HDS products, dimethylbiphenyl and methyl-phenyl-methyl-
338 cyclohexane. This important inhibition is explained by the high aromatic electron density which leads to a strong adsorption on acidic sites. - F+Napht.
1 X, Ratio of Activity A
0.9
m
iBuBz series
— X —
• B+Napht. • D+Napht. " F+Tetral.
A- —•
B+Tetral.
X- —•
D+Tetral. F+iBuBz
10
20
30
40
----A---
B+iBuBz
X - - - D+iBuBz %Arom. in decane Fig.4 Ratio of Activity for HDS of 4,6-DMDBT in decane + aromatics
* Figure 5 shows 4,6-DMDBT HDS inhibition by H2S which was produced by dimethyldisulfide (D.M.D.S., 0-1.5%wS) in decane. At 270°C, H2S is a strong inhibitor as aromatics. Particularly, NiMo catalyst B is more inhibited by H2S than CoMo, explaining partly the lower activity of NiMo in the gas oil. The HDS products distribution showed the hydrogenation route is more affected by H2S than the direct desulfurization route on NiMo. An opposite trend was found over CoMo. According to a series of HDS tests of 4,6-DMDBT in decane+D.M.D.S. over some Ni"Co-Mo catalysts, it was confirmed that the hydrogenation rate constant of 4,6-DMDBT (first step in its HDS hydrogenation route) is proportional to the %Co on the catalyst, for a total metal loading %Co+%Ni = 3%w. Ni has a low promoting efficiency in presence of H2S. On the other hand, the hydrogenolysis step rate constant (second step in its HDS hydrogenation route, see fig. 7) is also proportional to the %Co on the catalyst, for the same total metal loading %Co+%Ni = 3%w. At least, the direct desulfurization route is more inhibited by H2S over CoMo catalysts than over NiMo catalysts. Except for C catalyst, H2S behaved as an inhibitor like aromatics : the direct desulfurization route on CoMo catalysts and the hydrogenation route on NiMo catalysts are the most affected. On C catalyst, whereas H2S is a weak inhibitor, the aromatics are very strong inhibitors.
339 Relat. Activ, ^ -D A
m
n D
r 1
0
I
I
0.5
1
_ %w S in dec. T
1.5
Fig.5 Ratio of Activity for HDS of 4,6-DMDBT in decane + H2S 3.4. HDS of L.C.O. compared with G.O. Figs. 6a-6d show the comparison between the HDS tests of L.C.O. and gas oil over the 6 catalysts under the same conditions. k in LCO
kinGOOM Cond.: 340^; - 2.4 MPa H2 - 0-60 min.
A
B
C
D
E
F
Fig.6a Total Sulfur Content in L.C.O. and G.O. (1.65 order)
0.018 k 4,6-DMDBT 0.016 (1/min.) 0.014 0.012 0.01 0.008 0.006 0.004 0.002 0 Hg^4,6-DMDBT in L.C.O. and ' (first order)
340
k 4-MDBT
kDBT (1/min.)
B
C
i
B
Fig.6c 4-MDBT in L.C.O. and G.O. Fig.6d DBT in L.C.O. and G.O. (first order) (first order) In L.C.O., all the catalysts had about the same activity for the total sulfur content, except catalyst D showed a definitely higher activity. For the identified sulfur species, it appeared that the ratio kin L.co. / kin G.O. is decreasing from DBT («0.5) > 4-MDBT («0.25) > 4,6-DMDBT («0.68). It means that the high aromatic content of L.C.O. affected more the hydrogenation route (main route for 4,6DMDBT) than the direct desulfurization route (main route for DBT). To distinguish the catalysts, NiMo catalysts showed a slightly higher activity than CoMo catalysts for HDS of DBT ; on the contrary, for 4,6-DMDBT, CoMo are superior to NiMo. It confirms that the direct desulfurization route is more inhibited by aromatics on CoMo than on NiMo ; on the contrary, the hydrogenation route is more inhibited by aromatics on NiMo than on CoMo. Acidic catalyst C had a comparable activity with the other ones, showing that isomerization & cracking routes are probably become minor routes in L.C.O. desulfurization because of large aromatic content in L.C.O.
4, D I S C U S S I O N The present study compared the catalytic activities of 6 catalysts for HDS of a gas oil and aim to clarify the high activity of CoMo on acidic supports. One can examine the HDS reactivity of 4,6-DMDBT in decane with inhibitors such as aromatic compounds and H2S to evaluate the performances and schemes for the available catalysts. The present comparisons of the catalysts clarified several points of discussion. The catalysts achieved the deep desulfurization by the desulfurization of 4,6-DMDBT tiirough the hydrogenation and acidic routes, which are both inhibited severely by aromatics partners, similarly to basic compounds. Both routes can be evaluated by standard tests using hydrogenation of naphthalene and dealkylation of isopropylbenzene
341
Hydrodesulfurization Reaction Scheme for 4^6-DiMethylDiBeiizoThiophene
rt^^S— r^^H
UJr^^
TT tion
ca
& Direct Desulfurization
Cracking & Isomerization
VA CH3
CH3
4,6.DMDBT
CH3 CH3 H-4,6DMDBT
Direct Desulfurization
CH3
CH3
C Product
Hydrogenolysis
CH3
CH3
B Product
CH3
CH3
A Product
J
Cracking •> Benzene, Toluene, Xylenes, Cyclohexane, Methylcyclohexanes Fig. 7: HDS reaction scheme for 4,6-dimethyldibenzothiophene
342 respectively. Very acidic zeolite-containing catalyst allows the isomerization and cracking of sterically hindering methyl groups, such as for 4,6-DMDBT in decane. However such catalysts suffer the decrease of these properties with the time on stream by coking and are inhibited by aromatic partners. So, the importance of acidic routes is lower for HDS of real feed stocks than for that of 4,6-DMDBT in decane. Silica-alumina supported and zeolite containing acidic catalysts exhibited a very higher HDS activity for 4,6-DMDBT in decane compared to the other catalysts, but their activity is leveled off for HDS of high aromatic content L.C.O. real feed stock. Inhibition by H2S is important for the studies on model molecule at 270°C, but this effect may be lower at higher temperature, under practical conditions, owing to the H2S adsorption constant decreasing with the temperature. Interestingly, acidity of the catalyst appears to enhance the hydrogenation activity of CoMo which accelerates the desulfurization. Lower coking acidity obtained with silica-alumina support provides a high hydrogenation which is less inhibited by aromatics, explaining the high activity for HDS of G.O. and the highest activity for HDS of L.C.O. Strong adsorption of both aromatics and H2S inhibitors is involved on alumina-supported catalysts. The same adsorption phenomena could explain the similarity of qualitative inhibition effects ; indeed, aromatics and H2S affected severely the direct desulfurization route on CoMo and the hydrogenation route on NiMo. Deep desulfurization (ex. 4,6-DMDBT desulfurization) going through mainly hydrogenation route, CoMo suffer much less inhibitions by aromatics partners and H2S than NiMo, being superior in the practical desulfurization where the inhibitors are always present. The inhibition of H2S appears less on the acidic catalysts. Hence, the catalyst which is active and selective for the hydrogenation of 4,6-DMDBT in presence of inhibitors can be a target of better performance. The support for CoMo catalysts can be thus explored in terms of controlled acidity and better dispersion of CoMo.
Acknowledgments:
We gratefully acknowledge ELF-ANTAR FRANCE for supporting this work, Japan Energy Co. and Haldor Topsoe A/S for supplying experimental and reference commercial catalysts. 1. R. Gerdil and E. Lucken, J. Am. Chem. Soc. 87 (1965) 213 2. T. Isoda, S. Nagao, X. Ma, K. Sakanishi, I. Mochida Japanese Petroleum Institute bi-annual Conference, October 1994 3. D. Yitzhaki, M.V. Landau, D. Berger, M. Herskowitz Applied Catalysis A: General 122 (1995) 99
® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
343
HYDROTREATING OF COMPOUNDS AND MIXTURES OF COMPOUNDS HAVING MERCAPTO AND HYDROXYL GROUPS T.-R. Viljava and A. O. I. Krause Helsinki University of Technology, Department of Chemical Technology Kemistintie 1, FIN-02150 Espoo, Finland ABSTRACT Simultaneous hydrodesulfurization (HDS) and hydrodeoxygenation (HDO) of mercapto and hydroxyl group containing benzenes was studied using a commercial presulfided C0M0/7AI2O3 catalyst under hydrotreating conditions (150-280 °C, 7 MPa). Mercaptobenzene, phenol and 4-mercaptophenol were used as model compounds, and CS2 was used as precursor for HjS. The HDS rate of a mercapto group in the presence of a hydroxyl substituent in the para position was higher than that for the molecule containing only a mercapto group. When the hydroxyl group was present as phenol, the HDS rate of the mercapto group was about 30% lower than that for mercaptobenzene without an oxygen-containing additive. The decrease in the HDS rate was independent of the initial molar ratio of sulfur and oxygen within the ratios studied (5:1-1:1). The HDO rate of a hydroxyl group was suppressed by the mercapto group present either in the same or in a separate molecule. HDO reactions did not start until HDS conversion was almost complete. CS2 also decreased the HDO rate of phenol. When compared to the reactions of phenol alone, the rate of the hydrogenolysis route to benzene was decreased in the presence of a sulfur additive more than the hydrogenolysishydrogenation route to cyclohexane. 1. INTRODUCTION Hydrodesulfurization (HDS) is of great importance in the oil-refining industry. In the long term, it is likely that biomass will be used as an alternative raw material for liquid fuels and chemicals as such or in mixtures with traditional feedstocks. A new kind of hydroprocessing is needed to treat the feeds containing considerable amounts of both oxygen and sulfur. Reactions taking place during hydroprocessing have mostly been studied using model compounds [1-3]. Benzothiophenes and dibenzothiophenes represent the least reactive organic sulfur compounds in fossil fuels, and thus their reactions have been most intensively investigated. Phenol and naphthol derivatives and heterocyclic oxygen compounds have generally been used as oxygen-containing model compounds [1]. Interactions between different heteroatoms during hydrotreating have typically been studied with mixtures of model compounds. HDS reactions have turned out to be slightly inhibited by the oxygen-containing compounds [1,3].Only a few studies deal with the hydrotreating of
344 compounds containing both sulfur and oxygen in the same molecule [4-10]. In the cases where the model compound contained both oxygen and sulfur in different substituents of a benzene ring, the HDS reactions were promoted considerably in the presence of the oxygencontaining substituent compared to the corresponding oxygen-free compound. This phenomenon has been explained by the increased electron density of the sulfur atom in the presence of an oxygen-containing substituent [8-10]. Organosulfur compounds seem to have only a weak effect on hydrodeoxygenation (HDO) [1]. However, the role of sulfur, especially the role of H2S [3], in HDO is not quite clear so far. Sulfur is, to some extent, needed in HDO to maintain the sulfidation of the catalyst. On the other hand, competitive adsorption of the sulfur compound and HjS formed from the sulftir compound may have an inhibiting effect in HDO [1,11]. In addition, sulfur-containing compounds may alter the selectivity of HDO [11,12]. So far, the hydrotreating of binary mixtures of sulfur and oxygen compounds has only been studied with a thiophenic compound, or H2S as a sulfur-containing reactant. No reports have been presented concerning interactions in hydrotreating of mixtures of sulfur and oxygen compounds with similar structures, e.g. mercaptobenzene and phenol. In order to get deeper on understanding, we have carried out a series of experiments with compounds and binary mixtures of compounds containing a hydroxyl group and a mercapto group either in the same or in a separate molecule. The effect of H2S, formed from CS2 precursor, on the HDO of a phenolic hydroxyl has also been investigated. The kinetic parameters of the HDS and HDO reactions of the substituents are compared in this paper. 2. EXPERIMENTAL 2.1. Catalyst The catalyst was a commercial hydrodesulfurization catalyst, Ketjenfine 742-1.3Q, which contained 4.4 wt-% of CoO and 15 wt-% of M0O3 on 7-AI2O3. The catalyst was crushed and sieved to a fraction of 0.75-1.0 mm, and presulfided off site with CS2/n-hexane at 280 °C under hydrogen. 2.2. Hydrotreating procedure Hydrotreating of mercaptobenzene (Merck, >98%), 4-mercaptophenol (Aldrich, >90%), phenol (Carlo Erba, >99.5%) and carbon disulfide (Merck, 99.99%) was studied using a 50 ml stainless steel autoclave at 150-280 °C and at a total pressure of 7.2-8.4 MPa. The substrate or the mixture of substrates in m-xylene (Merck, >99%) was added to the preheated reactor via a high pressure feed vessel. Decane (Fluka, >99.5%) was used as a tracer during the reactions. The amount of solvent, tracer and catalyst in the experiments was constant, being 15 ml, 200 /xl and 0.5 g, respectively. CS2 was used as a precursor for H2S. 4-10 runs were carried out with each model compound or a mixture of model compounds. Four to six samples of 100-200 mg were withdrawn from the reactor liquid phase and analyzed by gas chromatography (HP 5890 A, flame ionization detector, capillary column DB-1). The accuracy of the analysis method was within + 5% for sulfur-containing aromatics, and ±2% for phenol and hydrocarbons. A more detailed description of the hydrotreating procedure is presented in our previous paper [10].
345 2.3. Phase equilibrium in the reactor To compensate for the lack of quantitative gas phase analysis, the gas phase composition in equilibrium with the analyzed liquid phase was estimated as described in detail earlier [13]. 2.4. Conversions and kinetic parameters The total conversions of HDS and HDO were defined as conversion of sulfur or oxygen in the reactant to HjS and H2O. The reactions were assumed to be first-order with respect to the sulfur or oxygencontaining reactant and the concentration of hydrogen in the reaction mixture was assumed to be constant. The following rate equation was used for the HDS and HDO reactions: rate = r^m^^Cj,
(1)
where k' is the pseudo reaction rate constant, m^at the amount of catalyst and C^ the concentration of the reactant. The parameters for the reactions of the known intermediates were fitted separately. All the reaction data available for one reactant or a mixture of reactants were processed simultaneously. The MODEST model estimation program [14] was used for calculations. 3. RESULTS Examples of the composition of the hydrotreated product at 225 °C are given in Table 1. On the basis of the product compositions, the simplified reaction networks for mercaptobenzene, phenol and 4-mercaptophenol can be presented as shown in Figure 1. Direct hydrogenolysis of the aromatic carbon-sulfiir bond was the main reaction of the mercapto group. No ring hydrogenated products were detected for mercaptobenzene, and no compounds containing sulfur but not oxygen were found in the reaction product of 4mercaptophenol. The HDO of phenol proceeded via both hydrogenolysis of the carbonoxygen bond to form benzene and the hydrogenation-hydrogenolysis route to form cyclohexane and cyclohexene [10]. HDS of mercaptobenzene was much easier than HDO of the oxygen compound of similar structure, phenol, under the same experimental conditions. The HDS conversions were almost complete in the reaction times in which the HDO conversions were still below 10%. The HDS conversion the disubstituted model compound, 4-mercaptophenol, was clearly higher than that of the corresponding oxygen-free compound, mercaptobenzene under the same reaction conditions (see Figure 2a). On the other hand, the conversion of HDO of 4mercaptophenol was much lower than that of phenol (see Figure 2b). In studies with binary mixtures of phenol and a sulfur-containing compound, mercaptobenzene or CS2, the same reaction products for HDS and HDO were detected as with mercaptobenzene and phenol alone. However, the HDO started at higher temperatures and the HDO conversion was much lower than that for phenol alone under the same reaction conditions. The conversions as a function of reaction time at 250 °C for pure phenol, 4mercaptophenol and phenol in the presence of CS2 arepresented in Figure 2b. It was not possible to determine the HDO conversion in the experiments with mixtures of phenol and
346 |SH
(a)
0^ —
+ H2,-H2S
0 ^ ^
<
-1/2*H 2S
DiphenylsuMde
Mercaptobenzene
(b)
X 0
+H2,-H20
j^
0
+n*H2,-H20^^ Phenol
^^
Benzene Cyclohexane + Cyclohexenes
r (c)
Benzene
OH
H2,-H2S
^ ^
+H2, -H2O
^^--^
SH +n*H2,-H20 4-Mercaptophenol
Phenol
0 0-0 Benzene Cvclohexane + Cvclohexenes
Figure 1. Simplified reaction networks for the hydrotreating of (a) mercaptobenzene, (b) phenol and (c) 4-mercaptophenol. mercaptobenzene, because the reaction product, benzene, was formed from both the sulfur and the oxygen-containing reactant. Significant side reactions of phenol to coke and high molecular weight products, as well as inaccuracies in the analysis of phenol made it impossible to use the phenol concentration data in calculations. The selectivity of phenol reactions changed significantly in the presence of sulfur (see Figure 3). The hydrogenolysis route to benzene was clearly retarded in the presence of CS2. When the reactant contained sulfur in the same molecule (4-mercaptophenol), the selectivity of the hydrogenolysis route first decreased, but increased again later as the HDO conversion increased, and an even higher selectivity was achieved than in the absence of sulfur. This differs from the behaviour of phenol without sulfur additives [10]; selectivity to benzene in the hydrotreating of phenol decreased slowly with increasing conversion at any temperature studied. In studies with binary mixtures of phenol and mercaptobenzene, the comparison of
347 Table 1 Hydrotreating products at 225 °C Mercaptobenzene, 7.3 MPa Time, h: mol/1 - Mercaptobenzene - Benzene - Diphenylsulfide Phenol, 7.8 MPa Time, h: mol/l - Phenol - Benzene - Alicyclics^ 4-Mercaptophenol, 7.9 MPa Time, h: mol/1 - 4-Mercaptophenol - Phenol - Hydrocarbons^ Mercaptobenzene + Phenol,, 7.7 MPa Time, h: mol/1 - Mercaptobenzene " - Phenol - Benzene - Diphenylsulfide - Alicyclics^ CS2 + Phenol, 1:2 molimol , 8.2 MPa Tune, h: - Phenol mol/1 ti
II
II
II
II II
II II
II
0 0.562 0 0
0.25 0.339 0.133 0.023
0.5 0.239 0.213 0.031
1.0 0.074 0.334 0.031
0 0.603 0 0
0.5 0.569 0.004 0.003
1.0 0.564 0.007 0.010
2.0 0.527 0.013 0.021
0 0.530 0 0
0.12 0.173 0.338 0
0.5 0.005 0.497 0
1.0 0.004 0.486 0.001
0 0.600 0.665 0 0 0
0.41 0.352 0.665 0.140 0.034 0
0.75 0.236 0.661 0.238 0.048 0
1.42 0.097 0.655 0.315 0.049 0
0 0.607
1.0
1.5
2.0
_C
_C
_C
^ Alicyclics = cyclohexane + cyclohexenes. ^ Hydrocarbons = benzene -h cyclohexane + cyclohexenes. ^ No reaction detected at 225 °C. selectivity is not reliable due to the benzene reaction product which is conmion for both reactants. 4. DISCUSSION The HDS rate of the mercapto group in the benzene ring was significantly enhanced by a hydroxyl substituent in the para position. When the hydroxyl group was present as phenol in the reaction mixture, HDS was slightly retarded. On the other hand, the HDO of a phenolic hydroxyl was clearly suppressed in the presence of sulfur either in the same molecule or in a separate compound. The presence of a sulfur additive altered the selectivity of HDO as well. Under all reaction conditions studied, HDS was easier than HDO of the corresponding structure.
348 jp
100 7 5 -\ ' lfQrca.ptobenzene - 4-HQi:captophQnol - © — Her cap t ob QRZ e n e : P h e n o l 2 : 1 inol:inol -^— Hex cap t ob QHZ e n e : P h e n o l 5:1 inol:inol
(a)
Time^ h
(b) U Phenol - B — 4-Hez:captophenol - 0 — P h e n o l : C S 2 2 : 1 inol:inol - • — P h e n o l : C S 2 1 0 : 1 inol:inol
Time, h Figure 2. Conversion of (a) HDS at 225 °C for mercaptobenzene, 4-mercaptophenol and mixtures of mercaptobenzene and phenol and (b) HDO at 250 °C for phenol, 4-mercaptophenol and mixtures of phenol and CSj as a function of time. 4.1. Effect of oxygen on HDS As mentioned earlier, the main reaction in the hydrotreating of mercaptobenzene alone or in the presence of phenol is the hydrogenolysis of the sulfur-carbon bond to benzene and H2S. For 4-mercaptophenol, phenol is formed correspondingly as the main product of HDS. The parameters for these HDS reactions (marked with (*) in Figure 1) at 225 °C are presented in Table 2. The HDS rate of a mercapto group was about one order of magnitude higher for the compound having a hydroxyl substituent in the para position than for the oxygen-free compound. This kind of enhancement of the HDS rate was also observed earlier,
349
=B= Phenol, 250 C
- 4-lfezcaptophenol, 250 C - Phenol :CS2 10:1 molimol, 250 C 4-Mei:captophenol, 225 C
HDO c o n v e r s i o n , %
Figure 3. Selectivity towards benzene as a function of HDO conversion for phenol, 4-mercaptophenol and a mixture of phenol and CSj (initial molar ratio 10:1). when the reactivity of 4-(methylmercapto)phenol was compared to that of the corresponding oxygen-free compound [10] It has often been proposed that there are at least two kinds of active sites on a hydrotreating catalyst, one of which is responsible for hydrogenolysis and the other for hydrogenation reactions [1,11,12,15-18]. The active sites for hydrogenolysis are thought to be more electrophilic in nature than those for hydrogenation. The electron-releasing properties of the hydroxyl substituent increase the x-electron density of the sulfur atom in 4mercaptophenol compared to the sulfur in mercaptobenzene. Thus 4-mercaptophenol adsorbs preferentially via the sulfur atom on the hydrogenolysis sites, and the direct hydrogenolysis of the sulfur-carbon bond to form phenol is favoured over the reactions of the hydroxyl group. The apparent activation energy for HDS of 4-mercaptophenol is clearly higher than that of mercaptobenzene (see Table 2). This may indicate a change in the mechanism of the hydrogenolysis reaction compared to the mechanism for the corresponding oxygen-free compound. It is possible that the bonding between the sulfur of 4-mercaptophenol and the active site for hydrogenolysis differs from that of mercaptobenzene due to the aforementioned electronic properties of the sulfur in 4-mercaptophenol. This may also explain the difference in the reaction networks of mercaptobenzene and 4-mercaptophenol; no disulfide structures were found in the hydrotreating products of 4-mercaptophenol. When the hydroxyl was present as a phenol additive, the HDS rate of the mercapto group in mercaptobenzene decreased by about 30% when compared to the reactions of pure mercaptobenzene. The decrease was independent of the initial molar ratio of mercaptobenzene and phenol for the ratios studied, 1:1, 2:1 and 5:1. The activation energy for the HDS of the mercapto group in binary mixtures was even 50% higher than that for mercaptobenzene alone. Competitive adsorption of the reactants on the active sites of the catalyst is a probable reason for the retardation of HDS and the high apparent activation
350 Table 2 The effect of a hydroxy 1 group on HDS rate of a mercapto group at 225 °C
k', (gcat*h)' EAct, kJ/mol Explained, %""
Mercaptobenzene
4-Mercaptophenol
30.8 84 99.1
314 100 99.7
Mercaptobenzene + Phenol Initial molar ratio 5:1 2:1 1:1 20.5 20.9 20.7 -^ -^ 126 99.4 99.2 98.8
^ No temperature dependence studied. ^ Data explained by the model. energy of HDS. Since no alicyclic reaction products were detected in the reactions of mercaptobenzene, the hydrogenated ring structures in the reactions of mixtures of mercaptobenzene and phenol originate solely from phenol, and they can thus be an indication of HDO. However, hydrogenated oxygen-free products of phenol were only detected in very small amounts in the reaction mixture, even at higher reaction temperatures. Thus HDO via the hydrogenation-hydrogenolysis route has not proceeded markedly. It seems that as long as the reaction mixture contains unreacted mercapto groups, the adsorption via sulfur is prevailing over the adsorbed species of phenol on the active sites for both hydrogenolysis and hydrogenation. Because of the low HDO rate of phenol compared to the HDS rate of mercaptobenzene /13/, the HDO can not be observed, and only the decrease in the HDS rate indicates the competition in adsorption. 4.2. Effect of sulfur on HDO The HDO reactions of the phenolic hydroxy 1 were markedly retarded in the presence of sulfur, either in the same or in a separate molecule (see Figure 2b). In the presence of CS2 in the reaction mixture, the HDO rate at 250 °C decreased when the initial concentration of CS2 was increased. In the presence of mercaptobenzene, alicyclic compounds, indicators for HDO, were only detected in the reaction products at HDS conversions higher than 90%. Also, with 4-mercaptophenol, HDO only started at almost complete HDS conversion. This period, when HDO does not proceed, may be linked to the type of sulfur compound present: HDO does not proceed to any detectable degree in the presence of mercapto groups but starts when H2S is the dominant form of sulfur in the reactor. In the presence of mercapto groups, both active sites are preferentially covered by a surface complex of the mercapto sulfur. For CS2, the reaction to H2S is relatively fast, and proceeds via the methylmercaptan intermediate [19]. In the early stages of the reaction with binary mixtures of phenol and CS2 at the low temperatures used, the amount of methylmercaptan intermediate may limit HDO reactions. About 50% of the H2S formed from CS2 goes to the gas phase of the reactor under the reaction conditions used. Thus the concentration of sulfur in the liquid phase of the reactor decreases with conversion of CS2. Further, it seems that the adsorption constants for mercapto group containing compounds are higher and the inhibition effect in HDO reactions therefore stronger than for H2S. So far we have not enough experimental data to test Langmuir-Hinshelwood type of rate equations for the reactions. The pseudo reaction rate
351 constants for the HDO of phenol, 4-mercaptophenol and mixtures of phenol and CS2 are presented in Table 3. As can be seen in Table 3, the activity of the hydrogenolysis sites on the catalyst is almost totally lost in the presence of CS2, and at higher CS2 concentrations also the hydrogenation sites become inactive. It has been suggested [11,12,15] that changes in the H2S-H2 ratio in the gas phase during hydrotreating causes interconversion of the active sites on the catalyst for hydrogenolysis and hydrogenation. At low concentrations of H2S, the hydrogenation route can even be promoted to some extent, due to the conversion of hydrogenolysis sites to hydrogenation sites. At higher concentrations, conversion of HDO will clearly be decreased, since sites inactive for both adsorption and reaction are formed. If so, interconversion of the active sites may explain the decrease in the selectivity of the direct hydrogenolysis route of HDO in the presence of CS2, but because the total HDO activity of the catalyst is strongly decreased in the presence of sulfur, the sites may just have become inhibited to a different degree. However, the selectivity of hydrogenolysis of 4-mercaptophenol increases with HDO conversion and reaches an even higher level than with phenol in the absence of any kind of sulfur additive. A slow reactivation or resulfidation of the oxide state of the catalyst formed in HDO reactions may be one possible explanation. In a batch reactor, the stability of the catalyst during the reactions cannot be tested, and this has potentially affected the reactions in our study. We have now started an investigation concerning the stability of the catalyst during HDO in a continuous flow reactor. Table 3 The effect of different types of sulfur compound on the HDO rate of a hydroxyl group at 250 °C
k'HG,(gcat*h)-^^ k'HYD, (gcat*hr''
Explained, %""
Phenol
4-Mercaptophenol
1.78 1.15 99.9
-0.27 0.13 92.9
Phenol + CS2 Initial molar ratio 10:1 2:1 --0.06 --0.07 0.77 0.24 99.4 88.4
"* HG = Hydrogenolysis route. ^ HYD = Hydrogenation route. *" Data explained by the model.
CONCLUSIONS The rate of the HDS of a mercapto group in a benzene ring is clearly affected by the presence of a hydroxyl substituent, either in the same or in a separate molecule. The HDS rate is about an order of magnitude higher in the presence of a hydroxyl group in the same molecule in tht para position to a mercapto group. This increase in the HDS rate is probably caused by the increased electron density of the sulfur atom, and the apparent activation energy of HDS of the compound is also increased. This indicates a change in the reaction
352 mechanism of HDS, potentially caused by a different type of bonding between sulfur and the catalyst in the presence of the hydroxyl substituent. When the hydroxyl group is in a separate molecule, phenol, the rate of HDS is decreased, probably due to competitive adsorption of the reactants on the active sites of the catalyst. The HDO rate of a hydroxyl group is suppressed in the presence of a mercapto group, either in the same or in a separate reactant molecule, and ahnost complete HDS conversions are needed to start HDO. The presence of CS2 also decreases the HDO rate of phenol. The reaction selectivity of HDO is strongly affected in the presence of sulfur: in the presence of H2S the hydrogenationhydrogenolysis route of HDO is favoured over direct hydrogenolysis. REFERENCES 1. 2. 3. 4. 5. 6. 7. 8. 9. 10. 11. 12. 13. 14. 15. 16.
17. 18. 19.
Girgis, M.J., Gates, B.C., Ind. Eng. Chem. Res., 30 (1991) 2021. Topsoe, H., Clausen, B.S., Massoth, F.E., Hydrotreating Catalysis, Science and Technology, Springer-Verlag Berlin Heidelberg, Germany 1996. Schulz, H., Schon, M., Rahman, N.M., Stud. Surf. Sci. Catal., 27 (1986) 201. Aubert, C , Durand, R., Geneste, P., Moreau, C , Bull. Soc. Chim. Belg., 93 (1984) 653. Aubert, C , Durand, R., Geneste, P., Moreau, C , J. Catal., 97 (1986) 169. Nagai, M., Sakikawa, N., Bull. Chem. Soc. Jpn., 51 (1978) 1422. Toropainen, P., Bredenberg, J. B-son, Appl. Catal., 52 (1989) 57. Konuma, K., Hasegawa, H., Itabashi, H., Appl. Catal., 38 (1988) 109. Konujna, K., Takase, S., Kameda, N., J. Mol. Catal., 79 (1993) 229. Viljava, T.-R., Krause, A.O.I., Appl. Catal. (Accepted), 1996. Laurent, E., Delmon, B., Ind. Eng. Chem. Res., 32 (1993) 2516. Delmon, B., Bull. Soc. Chun. Belg., 104 (1995) 173. Viljava, T.-R., Krause, A.O.I., Appl. Catal. A: General, 135 (1996) 317. MODEST Version 1.0, Haario, H., Profmath 1994. Delmon, B., Froment, G.F., Catal. Rev.-Sci. Eng., 38 (1996) 69. Moreau, C , Durand, R., Geneste, P., Olive, J.L., Bachelier, J., Comet, D., Duchet, J.C., Lavalley, J.C, Bonnelle, J.P., Grimbolt, J., Kasztelan, S., Payen, E., Breysse, M., Cattenor, M., Decamp, T., Frety, R., Gachet, C , Lacroix, M., Leclerq, C , De Mourgues, L., Portefaix, J.L., Vrinat, M., Engelhard, P., Gueguen, C , Toulhoat, H., Prepr.-Am. Chem. Soc, Div. Pet. Chem., 32 (1987) 298. Moreau, C , Joffre, J., Saenz, C , Geneste, P., J. Catal., 122 (1990) 448. Gevert, B.S., Otterstedt, J.-E., Massoth, F.E., Appl. Catal. 31 (1987) 119. Weisser, O., Landa, S., Sulfide Catalysts, Their Properties and Applications, Pergamon Press, New York, 1973.
® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
353
Influence of High Mo Loading on the HYD/HDS Selectivity of Alumina Supported M0S2 Catalysts Pedro Da Silva, Nathalie Marchal and Slavik Kasztelan Kinetics and Catalysis Division, Institut Frangais du Petrole, B.P. 311, 92506 Rueil-Malmaison, France High loading Mo/alumina catalysts have been tested for the simultaneous hydrodesulphurisation of dibenzothiophene and hydrogenation of 1-methyl naphthalene in presence of isoquinoline at SMPa, 623 K. In these conditions, it is found that the HYD/HDS activity ratio is increasing when the Mo loading increases. All used catalysts have been found well sulphided as indicated by X-ray photoelectron spectroscopy. Transmission electron microscopy analysis indicated that the detected M0S2 slabs have a mean length varying from 25 to 40 A and a number of slab per stack increasing from 1.1 to 1.4 when the Mo loading increases. Assuming that the detected slabs are representative of the active phase morphology, only the HYD activity is found to correlate well with the number of edge or edge+corner surface sites in the range of Mo loading investigated. 1. INTRODUCTION The influence of the Mo loading of sulphided Mo/alumina hydrotreating catalysts on their catalytic properties has been often investigated in the literature at atmospheric pressure using simple reactants such as thiophene, pyridine, propene etc. [1-7] and seldom at high pressure on real feed or complex mixtures [8]. The HDS and/or HYD specific activity have been usually found to increase when the Mo loading increases and to follow a S shaped curve whereas the HYD/HDS activity ratio has been found in general constant at high Mo loading [1-8]. In their working state these catalysts are known to contain poorly crystallised M0S2 particles with one or several slabs in interaction with the alumina surface [9-13]. Two morphological parameters are important regarding the activity and selectivity of these particles namely the length and the stacking of the slabs. It has been proposed that the Mo loading of Mo/alumina catalysts has an effect on the mean length of the M0S2 slabs measured by TEM [14,15]. On the other hand, the influence of Mo loading on the stacking does not seem to have been reported. The activity of M0S2 slabs is assumed to be linked to peripheral coordinatively unsaturated Mo sites and both corner or purely edge Mo sites
354
have been proposed as active sites [8,9,11,14]. Therefore the size of the two dimensional slab will determine the number of edge sites per slab whereas the number of corner sites will be constant. Recently, it has also been proposed that the stacking of the slabs influences the number of sites accessible for different reactants and therefore affects the HYD/HDS selectivity [16,17]. The influence of the Mo loading on the activity and selectivity for the different hydrotreating reactions is therefore related to the amount, morphology and electronic properties of the alumina supported M0S2 particles. From an academic and applied point of view, it is important to establish clearly the activity-structure relationship in the case of supported catalysts. In this work, a preliminary study is reported whereby a set of catalysts with high Mo loading was prepared, sulphided and tested to determine simultaneously the HDS and HYD activities in presence of a nitrogen compound and under hydrogen pressure. The used catalysts have been characterised by TEM with the hope that they will reveal differences in morphology of the supported M0S2 particles that could be related to differences in activity and selectivity. 2. EXPERIMENTAL A set of M0O3/AI2O3 catalysts with Mo loading increasing from 9.8 to 22.5 wt% Mo has been prepared by impregnation of ammonium heptamolybdate ((NH4)6M07O24*4H2O from Merck) on a y-Al203 support (Rhone-Poulenc, 230 m ^ g ' l , 0.75 cm^g"^). These catalysts have been dried at 393 K overnight and calcined in air at 773 K for 2 hours. Mo loading from 14.8 to 22.5 wt % have been obtained by two successive impregnation with an intermediate drying step at 393 K overnight, followed by calcination in air at 573 K for 2 hours. These oxidic Mo/alumina precursors have been characterised by several techniques and the results are reported elsewhere [18]. The HDS and HYD activities have been determined using a mixture of model compounds namely dibenzothiophene (DBT) for HDS and 1-methyl naphthalene (MelN) for HYD. The feed contained 10 wt% of DBT, 20 wt% of MelN, 8 wt% of isoquinoline, 5 % of dodecane and 57 wt% of cyclohexane. Activities have been measured in a fixed bed reactor at 623 K and 3 MPa. In these conditions, it has been found that isoquinoline, originally employed to determine the hydrodenitrogenation activity, is totally converted. Hence it has not been possible to determine the influence of the Mo loading on the C-N bond cleavage [19]. Isoquinoline is therefore an in situ generator of ammonia. The HYD and HDS reactions have been found to be first order and activities are expressed in the following as pseudo first order rate constant in mole/g/h. Prior to test, the samples have been sulphided at the same pressure using a feed containing 12 wt% of dimethyldisulphide, 20 wt% of toluene and 68 wt% of cyclohexane during 2 hours at 623K. After catalytic testing, the samples have been kept under n-decane to avoid contact of the surface with air. They have been
355 analysed by X-ray photoelectron spectroscopy (XPS), X-ray diffraction (XRD) and transmission electron microscopy (TEM). Conventional transmission electron microscopy (TEM) images have been recorded in order to evaluate changes in the morphology of the observed M0S2 particles. This has been done using a JEOL 2010 electron microscope operated at 200 kV. TEM samples have been ultrasonically dispersed in ethanol and the suspension was deposited on carbon coated copper grids. Each sample analysed by TEM showed the presence of layered M0S2 particles on the support and their mean length and mean number of layers in a stack has been determined according to rules described elsewhere [18]. In order to ensure a representative sampling of the catalyst morphology, at least ten images have been recorded per sample and about 300 slabs have been counted and measured. 3. RESULTS AND DISCUSSION The variation of the specific activity for the HDS of DBT and HYD of MelN as well as the variation of the HYD/HDS specific activity ratio versus the Mo loading are reported in Figures 1 and 2 respectively. In Figure 1, the specific HYD activity increases continuously when the Mo loading increases whereas the specific HDS activity, within experimental uncertainties, tends to remain constant in this range of Mo loading. When the specific HDS and HYD activities are expressed in mole/mole M o / h and plotted versus Mo loading, a continuous decrease of the activity per Mo versus Mo loading is obtained as reported in the literature [1-7]. In Figure 1, the HYD activity increases whereas the HDS activity remains constant. Therefore the HYD/HDS activity ratio is depending on the Mo loading and increases when the Mo loading increases up to 15 wt % Mo as shown in Figure 2. This tendency is notably different from literature reports where usually the HYD/HDS activity ratio is constant versus Mo loading [1-3,5]. However, these works have been performed using reactants such as thiophene, propene, butene etc. at atmospheric pressure. Our results suggest that the effect of Mo loading on the HYD/HDS activity ratio is different when a mixture of S, N and aromatic compounds is employed. The used catalysts recovered after testing have been found almost completely sulphided by XPS as indicated in Table 1. XRD analysis showed that in samples with high Mo loading namely 18.1 and 22.5 wt% Mo, M0O3 and Al2(Mo04)3 phases which were present in the oxidic precursor totally disappeared after sulphiding and testing. The XRD patterns showed traces of M0O2 and of M0S2 present as large particles in accordance with literature [7]. The used samples have been analysed by TEM in order to determine the change in the distribution of length and stacking of the observed M0S2 slabs as a function of Mo loading. The pictures obtained show the typical black lines corresponding to the projected images of M0S2 layers oriented in line with or slightly tilted from the electron beam [6-8, 15, 20, 22-25]. Therefore, only a small
356 fraction of the total number of slabs is detected [8, 20, 22]. It can be assumed, however, that M0S2 stacks as well as alumina crystallites are randomly oriented toward the electron beam and therefore that the slab length distribution of observed slabs is representative of the whole population of slabs. The M0S2 slabs not seen by the TEM can be assumed to have the same mean length distribution as what is detected. It has been proposed that very small clusters, not detectable by TEM, could be responsible for the catalytic activity [8]. Although it is difficult to imagine that clusters can survive in these conditions as suggested by thermodynamic considerations [26], this hypothesis cannot be ruled out. However, bulk M0S2 is known to be active suggesting that, in supported catalysts, M0S2 slabs should very likely be responsible for the catalytic activity [11]. 213 • AHDS
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Figure 1. DBT HDS and IMeN HYD activities vs Mo loading.
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Mo Loading (wt%)
Figure 2. HYD/HDS activity ratio vs Mo loading.
In this work we have attempted, as a first step, to interpret activity data using the observed TEM features. The smaller slabs detected in this work have been 1012 A long and it is likely that their number is underestimated. On the other hand it must be considered that the black lines observed on TEM pictures correspond to the thicker and most structurally organised part of the slab and therefore the length of the black lines underestimates likely the real larger dimension of the slab. For a given used catalyst, a large variation of slab size, from 10 to 80 A was found as shown by the histograms reported in Figure 3 for two different Mo loading. The mean length of the observed M0S2 slab calculated from such histograms are reported in Table 1 and is found to increase almost Unearly with the Mo loading in the range studied. The increase of the mean length of the slabs versus Mo loading confirms previous observations made after H 2 / H 2 S sulphiding at atmospheric pressure [15] and extends them to experimental conditions where high pressure testing
357 using S, N and aromatic compounds has been employed. However, the mean length remains rather small (25 to 40 A) and, surprisingly, the mean number of layers is close to 1 and does not change significantly in the studied Mo loading range except for an increase to 1.4 for the highest Mo loading. This number of layer in a stack is lower than previously reported data for sulphided Mo/alumina catalysts [6,7,15]. a)
20
30
40
50
60
Length of MoS2 slabs (A)
70
80
o
10 20 30 40
50 60 70 80
Length of MoS2 Slabs (A)
Figure 3. M0S2 slab length distribution for a) 9.8 wt%, b) 22 wt% M0/AI2O3 catalysts. The sulphided samples analysed by TEM and XPS in this work are those recovered after testing. The XPS analysis indicated that almost all of the Mo is sulphided. Hence there is no hidden assumption about the amount of M0S2 sulphide phase. The main assumption in the search for a relationship between morphological features and activity lies evidently in the TEM measurements which we assume to be representative of the total M0S2 population of particles in the sample. Table 1. Sulf. Mo Mo L n M N Nb of Slab Loading XPS TEM TEM 10'^ mol/g cat wt% Mo (% mol) (±2A) (± 0.2) 9.8 95±5 25 1.2 5 58 1.8 11.8 95±5 25 1.1 58 2.1 5 14.8 95±5 17 1.1 67 5 2.3 18.1 95±5 32 1.2 91 6 2.1 22.5 95±5 38 1.4 7 1.9 125 L: mean length of black lines on TEM pictures N : mean number of layers in a M0S2 stack n: mean number of Mo ions on one side of hexagonal M0S2 slab M : mean number of Mo ions in hexagonal M0S2 slab
358 Knowing the mean slab length, and assuming a symmetrical hexagonal shape for the M0S2 slabs, it is possible, using formulas previously reported [14], to estimate the number of Mo ions on one side of the hexagon (n) with L=6.4(n-1) in A and the number of Mo (M) in a M0S2 slab with M=3n2-3n+l. Assuming that all of the Mo ions are in the form of M0S2 particles, the number of slabs per gram of catalyst can be computed. Results from these calculations are indicated in Table 1. From the size of the hexagonal slab, the number of different types of surface Mo ions can be computed i.e. edge (Me=6n-12), corners (Mc=6) or pairs of adjacent Mo ions (Mp=(6n-6)) which is equal to the sum edge + corner Mo ions : (Mp=Me+c) [27]. IQ •
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3 4 5 6 7 8 Sites (10-4 mol/g) Figure 4. HYD activity vs number of edge (Me) or pair (Mp) of Mo sites.
2
3
4
5
6
7
8
Sites (10-4 mol/g)
Figure 5. HDS activity vs number of edge (Me) or pair (Mp) of Mo sites.
The specific HYD and HDS activities are plotted in Figures 4 and 5 respectively versus the calculated number of edge (Me) and pairs of Mo ions (Mp). The number of corner sites is equal to 6 times the number of hexagonal slabs and it can be seen in table 1 that the number of slabs per gram remains rather constant within the experimental uncertainties. Therefore, the HYD and HDS activities do not correlate with the number of corner sites alone. In Figure 4 it can be observed that a correlation is obtained for the variation of the specific HYD activity versus the amount of Me or Mp sites except for the highest Mo loading samples. For these samples the number of active slabs is Hkely overestimated because a fraction of the Mo in sulphided form is Hkely inaccessible due of the presence of large aggregates of M0S2 and some M0S2 layers wrapping M0O2 particles as observed by TEM. In Figure 4, it is not possible to differentiate between the two types of sites because of the scatter of the data. Nevertheless, it can be concluded that, within the range of Mo loading investigated, the HYD activity is well correlated to the number of surface sites of the M0S2 slabs.
359 Figure 5 indicates however that the HDS activity measured in our conditions is not correlated to the number of edge or edge+corner sites. This is also indicated by the change of the HYD/HDS ratio vs Mo loading previously underlined. This result is different from previous results reported by Bachelier et al. [5], which indicated that the HYD/HDS selectivity calculated from HYD of propene and HDS of thiophene activities, measured separately, did not vary as a function of Mo loading. While these results may suggest that there are different active sites for hydrogenation and hydrogenolysis reactions whose number varies differently with the length of M0S2 slabs, it is too early to reach such a conclusion. As a matter of fact, in our work the HYD and HDS activities have been measured simultaneously using a mixture of model compounds and complex inhibiting or kinetic effects may arise and interfere with the morphological effect. A comparison of the effect of Mo loading on the catalytic activities for the mixture of model reactants with the catalytic activities for the single reactants in the same operating conditions should provide an answer concerning this hypothesis. One more difficulty to establish a clear-cut activity-structure relationship is the potential effect of the stacking of M0S2 layers because the various slabs in a stack may not have the same activity and selectivity for a given catalytic reaction [16,17]. In addition, the intrinsic activity of each type of site of one M0S2 slab may also vary with the size of the slab. 4. CONCLUSION The results reported in this work indicate that high Mo loading on alumina does influence the HYD/HDS selectivity when HDS of DBT and HYD of MelN are measured simultaneously in presence of a nitrogen compound under hydrogen pressure. It is found that increasing Mo loading favours HYD activity. In parallel, increasing Mo loading leads to an increase of the length and perhaps the stacking of the TEM observed M0S2 particles. Assuming that the M0S2 phase observed by TEM is responsible for the activity, the HYD activity is found to correlate well with the number of edge or edge+corner sites, whereas the HDS activity is not correlated to the number of surface sites. Further work is needed to prepare samples with lower Mo loading, with well defined morphologies of M0S2 particles and to establish more precisely the relationship between the morphology of the M0S2 particles and the activity for the various hydrotreating reactions, specially in complex mixtures of reactants and under hydrogen pressure. Acknowledgement The financial support of the JNICT of Portugal, program PRAXIS XXI, to P. Da Silva is gratefully acknowledged.
360 REFERENCES 1. 2. 3. 4. 5. 6. 7. 8. 9. 10. 11. 12. 13. 14. 15. 16. 17. 18. 19. 20. 22. 23. 24. 25. 26. 27.
V.H.J. De Beer, C. Bevelander, T.H.M. van SintFiet, P.G.A.J. Werter, C.H. Amberg, J. CataL, 43 (1976) 68. }. Medema, C. van Stam, V.H.J. De Beer, D.C. Koningsberger, J. CataL, 53 (1978) 386. Y. Okamoto, H. Tomioka, T. Imanaka, T. Teranishi, J. CataL, 66 (1980) 93. R. Thomas, E.M. Van Oers, V.H.J. De Beer, J. Medema, J.A. Moulijn, J. CataL, 76 (1982) 241. J. Bachelier, M.J. Tilliette, J.C. Duchet, D. Cornet, J. CataL, 76 (1982) 300. J.L. Portefaix, M. Cattenot, J.A. Dalmon, C. Mauchausse, Stud. Surf. Sci. CataL, 50 (1989) 67. K.C. Pratt, J.V. Sanders, V. Christov, J. CataL, 124 (1990) 416. S. Eijsbouts, J.L. Heinerman, H.W. Eljerman, AppL CataL, 105 (1993) 69. F.E. Massoth, G. Muralidhar, Proceedings of the Fourth International Conference on the Chemistry and Uses of Molybdenum; Barry, H.F.; Mitchell, P.C.H.; Eds.; Climax Molybdenum Co. Ann Arbor, MI, 1982, p. 343 R. Prins, V.H.J. De Beer, G.A. Somorjai, CataL Rev. Sci. Eng., 31 (1989) 1. R.R Chianelli, M. Daage, M.J. Ledoux, Adv. CataL, 40 (1991) 177. B. Delmon, Stud. Surf. Sci. CataL, 53 (1990) 1. H. Tops0e, B.S. Clausen, F.E. Massoth, Catalysis Science and Technology; Anderson, J. R.; Boudart M.; Eds.; Springer Verlag Berlin, Vol. 11 (1996). S. Kasztelan, H. Toulhoat, J. Grimblot, J.P. Bonnelle, AppL CataL, 13 (1984) 127. E. Payen, S. Kasztelan, S. Houssenbay, R. Szymanski, J. Grimblot, J. Phys. Chem., 93 (1989) 6501. R.R. Chianelli, M. Daage, A.F. Ruppert, Stud. Surf. Sci. CataL, 75 (1993) 571. R.R. Chianelli, M. Daage, J. CataL, 149 (1994) 414. P. Da Silva, unpublished results. L. Vivier, V. Dominguez, G. Perot, S. Kasztelan, J. Mol. CataL, 67 (1991) 267. J.V. Sanders, Catalysis Science and Technology; Anderson, J. R.; Boudart M.; Eds.; Springer Verlag, Berlin, 7 (1995) 51. T.F. Hayden, J.A. Dumesic, J. CataL, 103 (1987) 366. J. Ramirez, S. Fuentes, G. Diaz, M. Vrinat, M. Breysse, M. Lacroix, AppL CataL, 52 (1989) 211. S. Srinivasan, A.K. Datye, C.H.F. Peden, J. CataL, 137 (1992) 513. P.T. Vasudevan, S.W. Weller, J. CataL, 99 (1988) 235. H. Toulhoat, S. Kasztelan, in M.J. Phillips and M. Ternan, Eds., Proceeding of the 9th International Congress on Catalysis, Calgary, The Chemical Institute of Canada, Ottawa, 1 (1988) 152. S. Kasztelan, A. Wambeke, L. Jalowiecki, J. Grimblot, J.P. Bonnelle, J. CataL, 124 (1990) 12.
® 1997 Elsevier Science B. V, All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P, Grange, editors
361
Low temperature hydrocracking of paraffinic hydrocarbons over hybrid catalysts I. Nakamura, K. Sunada and K. Fujimoto Department of Applied Chemistry, School of Engineering, The University of Tokyo 7-3-1, Hongo, Bunkyo-ku, Tokyo 113 Japan A hybrid catalyst, which was prepared by physically mixing of a H-ZSM-5 and a Pd/SiOi, showed an excellent activity for the hydrocracking of paraffins at low reaction temperature (503 K). In the n-heptane cracking, the hybrid catalyst gave only isomerized heptane and propane and equimolar amount of i-butane whereas the products on H-ZSM-5 alone distributed from C3 to C9 and C4 products contained all kind of paraffins and olefins. The wide product distribution for H-ZSM-5 system should be attributed to the reaction path comprising oligomerization and cracking of the oligomer. The simple products for the H2-hybrid system should be formed through no other reaction path than the primary cracking reaction on H-ZSM-5, which proceeds through a intermediate containing the protonated dialkylcyclopropane structure, while the hydrocracking reaction over Pd-mordenite hybrid catalyst proceeds via 6-scission of the carbenium ion. 1. INTRODUCTION Hydrocracking of petroleum heavy hydrocarbons have been practiced extensively conmiercially in petroleum refining to produce high quality gasoline, jet fuel, gas oil and lubricants. Many hydrocracking catalysts of commercial importance are dual functional catalysts containing both hydrogenation components such as sulfided Ni-Mo or Ni-W and acidic components such as zeolites. The most predominant reaction mechanism for the hydrocracking of paraffin over a dual functional catalyst is as follows: (1) the dehydrogenation of paraffin to olefin on the supported metal; (2) proton addition to the olefin to form carbenium ion on the acidic component; (3) 6-scission of the carbenium ion to form smaller carbenium ion and olefin on the acid component; (4) hydrogenation of the cracked olefin to paraffin on the supported metal [1]. On the other hand, a new mechanism for the hydroisomerization of n-pentane, which is also catalyzed by dual functional catalysts, has been proposed. The role of supported metal is not dehydrogenation-hydrogenation but the porthole of hydrogen spillover. A hydrogen molecule is dissociated on the metal particle and spills over onto zeolite surface as a proton and a hydride, where the proton promotes the acid catalyzed reaction such as skeletal isomerization. And hydride ion stabilizes cationic intermediate to prevent secondary reactions such as oligomerization and cracking to improve isomerization selectivity [2]. The spillover proton can also promotes hydrocracking.
362 And recently, Sie proposed a new mechanism for acid-catalyzed hydrocracking, which assumes a protonated dialkylcyclopropane carbonium ion as intermediate instead of a classical carbenium ion [3]. In the hydrocracking of paraffins with a noble metal containing catalyst such as Pd/HZSM-5, the olefinic-cracked products should be hydrogenated easily to be converted into the less reactive smaller paraffins. And if the hydride ion supplied by hydrogen spillover also reacts with the cationic-cracked products, secondary cracking reaction should be also suppressed and primary cracked products can be observed. In the present work, mechanism of paraffin hydrocracking was studied using hybrid catalysts containing zeolites and a supported palladium from the standpoint of hydrogen spillover. 2. EXPERIMENTAL 2.1 Catalyst preparation Zeolites used were a H-ZSM-5 (Toso, HSZ-840NHA SiOi/AlzOs ratio=44) and a dealuminated mordenite (DAM, Si02/Al203 ratio=44 ). The dealuminated mordenite was prepared by treating a H-mordenite (Toso, TSZ-640 Si02/Al203 ratio=19) with 4.5 N of hydrochloric acid at 343 K for 1 h so that its Si02/Al203 ratio fell into the H-ZSM-5 range. Pd was introduced to zeolite by the method of ion-exchange with aqueous solution of Pd(NH3)4a2. The ion-exchange was carried out at 373 K for 6 h with 0.1 wt% Pd(NH3)4Q2 aqueous solution under stirring, the palladium supported zeolite was washed by water until no chloride ion was detected. Silica-supported palladium was prepared by impregnating a commercially available Si02 (Aerosil 380, BET specific surface area 380 m2/g) with PdCh from its aqueous hydrochloric solution which was followed by the calcination in air at 723 K for 3 h and the reduction in flowing hydrogen at 723 K for 1 h. Hybrid catalyst was prepared by co-grinding the mixture of one weight part of zeolite with one weight part of Pd/SiOa (2.5 wt%) and pressure molding the mixture to granules to 20/40 mesh. Catalysts were activated in air at 723 K for 2 h and reduced inflowinghydrogen at 673 K for 1 h, before use. 2.2 Reaction apparatus and procedure The hydrocracking of paraffins was conducted with a continuous flow type fixed bed reaction apparatus under pressurized conditions (1.1 MPa). The reactor was a stainless steel tube with an inner diameter of 6 mm. The feed material which had been deeply desulfurized was fed by a liquid pump. The mole ratio of H2/C7H16 or N2/C7H16 in the feed was 9:1. Products were withdraw in the gaseous state and were analyzed by a high-resolution capillary gas chromatograph. 3. RESULTS AND DISCUSSION 3.1. n-Cy cracking on Pd containing DAM (dealuminated mordenite) catalysts Hydrocracking of n-heptane, which should give only C3 and C4 hydrocarbons as primary cracked products in acid catalyzed cracking, was studied. Fig. 1-a shows the changes of catalytic activities of a variety of catalysts containing Pd/Si02
363
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563 K, H2/n-C7=9, 1.1 MPa, D Pd/Si02:DAM=l:l, W/F=2.4 g h mol-i, mol-i
2) W/F=1.2 g h
Figure 1. Hydrocracking of n-heptane over Pd-DAM hybrid catalyst; a) conversion as a function of time on stream, b) C-number distribution (TOS=2.5 h), c) distribution of C4 hydrocarbons formed in the hydrocracking of n-Cv (TOS=2.5 h).
and/or DAM for n-heptane hydrocracking. Pd/DAM, which is a typical dual functional catalyst, showed excellent activity for the hydrocracking. The catalytic activity of DAM was not affected by the atmosphere and decreased quickly. Pd/Si02 showed little activity for both dehydrogenation and cracking of n-C-jKi^. On the other hand, the catalytic activity of a hybrid catalyst comprising Pd/Si02 and DAM was the highest and its activity was kept constant under hydrogen atmosphere while it was much lower and decreased quickly under nitrogen atmosphere. This phenomenon clearly shows that the presence of hydrogen is essential in order to generate hydrocracking activity. It is well known that the supported platinum shows a high catalytic activity for the
364 dehydrogenation of paraffin whereas the supported palladium does not The results shown in Figure 1 suggest that the dehydrogenation activity of supported metal is not essential for the appearance of the paraffin hydrocracking activity. The essential point is that the hydrogen-activating component intimately contacts with acidic catalyst. The present authors have pointed out that the skeletal isomerization of lower paraffin is effectively promoted by the hybrid catalyst composed of Pd/Si02 or Pt/Si02 and H-ZSM-5 and that the hydrogen spillover is the key step of isomerization reaction [2]. In the present case also, the hydrogen migration from Pd/Si02 or supported palladium to DAM should be essential for the high and stable catalytic activity. The characteristic feature of the product distribution is that the reaction products of H2-hybrid catalyst system are only isomerized heptane and propane and equimolar amount of isobutane (small amount of n-C4Hio was formed), whereas the products on DAM alone distributed from C3 to C9 and C4 products contained all kind of paraffins and olefins as shown in Fig. 1. The wide product distribution for DAM system should be attributed to the reaction path comprising oligomerization of cracked fragments and cracking of the oligomers. In the hydroisomerization of n-pentane over hybrid catalyst containing H-ZSM-5 and supported noble metal catalyst, it was proposed that a hydrogen molecule spills over onto zeolite surface as a proton and a hydride [2]. In hydrocracking, the proton also promotes cracking reaction as shown in Fig. 3. The hydride ion reacts with the cationic-cracked products and olefiniccracked products should be hydrogenated by palladium to be converted into the less reactive smaller paraffins. And the secondary cracking reaction should be suppressed. The simple products for the H2-hybrid system should be formed through no other reaction path than the primary cracking reaction on DAM. In hydrocracking of normal paraffin with metal supported acid catalyst, the iso/normal ratios in the paraffinic products generally exceed the thermodynamic equilibrium. It proves that at least some of the branched paraffins are primary products of the cracking and not a results of the post isomerization. This is particularly true in the case of C4, since n-butane cannot be isomerized under typical hydrocracking conditions with a zeolite catalyst. Especially the fact that isobutane selectivity in the Q products in the Pd-hybrid catalyst system reached 90 % as shown in Fig. 1-c suggests that the hydrocracking should not proceed through the cracking of n-paraffin to form smaller n-paraffin and another n-paraffin. One probable path is the skeletural isomerization of n-paraffin to branched paraffin and its cracking. Fig. 2 shows the results of hydrocracking of C7 isomers. 2,4-Dimethylpentane was more reactive than 2-methylhexane and 3-methylhexane and n-heptane was less reactive than the other branched isomers. The reactivities of the C7 isomers can be explain by stability of the corresponding carbenium ion. At low conversion level, isomerization selectivities of 2-methylhexane, 3-methylhexane and n-heptane were vary low and are estimated by extrapolation to be 0 at W/F=0, while 2,4dimethylpentane was isomerized and cracked, parallelly. These facts indicate that n-heptane, 2-methylhexane and 3-methylhexane are first isomerized to 2,4-dimethylpentane and then cracked to give propane and iso-butane. It has been suggested that formation of multibranched isomers from the feed and cracking are consecutive reactions [4]. Cracking of a normal paraffin must thus proceed through the stage of formation of monobranched isomers, dibranched isomers such as 2,4-dimethylpentane and finally cracked as shown in Fig. 3, because the high energy barrier for 6-scission of monobranched carbenium ion to form primary carbenium ion. The isobutylene, which is one of a pair of the primary cracked
365
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523 K, H2/n-C7=9,1.1 MPa, Pd/Si02:DAM=l:l, Figure 2. Hydrocracking of C7 isomers over Pd-DAM hybrid catalyst. products of the 2,4-dimethylpentane hydrocracking, will be hydrogenated to isobutane over palladium catalyst in the presence of hydrogen. As mentioned before, the hydride (Hso) as a counter anion of the proton (H+so), which is formed in hydrogen spillover process, stabilizes C7 carbenium ions and the propyl-carbenium ion to give C7 isomers and propane, respectively. Thus oligomerization of the cracked fragments and consecutive cracking reacdon are prevented in the H2-Pd-hybrid system. The hydrocracking of C7 paraffins over Pdmordenite hybrid catalyst can be explained by 6-scission of the carbenium ion and hydrogen spillover.
366 p-scission
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No p-scission
Hgure 3. Reaction model of n-heptane hydrocracking over Pd/Si02-DAM hybrid catalyst 3.2.
n-C7 cracking on Pd containing H-ZSM-5 catalysts It was suggested that hydrocracking of C7 isomers over DAM containing catalyst proceeds by 6-scission mechanism for breakages of a C-C bond in a classical type of carbenium ion and 2,4-dimethylpentil carbenium ion is the only intermediate to be cracked. HZSM-5 is also suitable for paraffin hydrocracking such as dewaxing[5-7]. H-ZSM-5 contains two intersecting channel systems the size of which are determined by ten-membered oxygen rings. The pore diameters are slightly below 0.6 nm [8-9]. Weitkamp et al. claimed that reaction involving tertiary carbenium ion intermediates do not take place in H-ZSM-5 [10]. The 2,4-dimethylpentil carbenium ion is a typical tertiary carbenium ion. Hydrocracking of C7 paraffins with H-ZSM-5 containing catalyst was studied. The results of n-heptane cracking with H-ZSM-5 hybrid catalyst were very similar to those of the DAM containing catalysts system as shown in Fig. 4. As was the case of the cracking with DAM hybrid catalyst, the catalytic activity was the highest for the hybrid catalyst under hydrogen pressure. The characteristic feature of the product distribution is that the reaction products of H2-hybrid catalyst system are only isomerized heptane and propane and equimolar amount of isobutane (a litfle n-C4Hio was formed), whereas the products over H-ZSM-5 alone distributed from C3 to C9 and C4 products contained all kind of paraffins and olefins. Only the difference is the composition of isomerized C7. C7 isomers were limited to be 2-methylhexane and 3-methylhexane over PdH-ZSM-5 catalyst, while both 2,4-dimethylpentane and methylhexanes were formed over PdDAM hybrid catalyst as shown in Fig. 4-d. However, the reactivities of branched C7 were different from those of the DAM containing catalysts system as shown in Fig 5.
367
O Pd-hybridH2^> •
Pd-hybridN2^^
D H-ZSM-5H2^> •
H-ZSM-5N2^>
A PdJSiO^ll^^^ 0
0.5 1 1.5 2 2.5 Time on Stream (h)
3
Pd-hybridH2
. . .. 1
,.. '
Pd-hybrid N2
1
H-ZSM-5 H2
A
H-ZSM-5 N2 — 1 — 1 — 1
1 2 3 4 5 6 7 8 9101112 Carbon Number
I
— 1
m 1
1
1
1
1
0 10 20 30 40 50 60 70 80 90 100
Distribution (mol%)
Pd/Si02-H-ZSM-5 Pd/Si02-DAM 0
10 20 30 40 50 60 70 80 90 100 Distribution of C7 Product (%)
503 K, H2/n-C7=9, 1.1 MPa, D Pd/Si02:H-ZSM-5=l:l, W/F=2.4 g h mol-i,
2) W/F=1.2 g
hmol-i Figure 4. Hydrocracking of n-heptane over Pd-H-ZSM-5 hybrid catalyst; a) conversion as a function of time on stream, b) C-number distribution (TOS=2.5 h), c) distribution of C4 hydrocarbons formed in the hydrocracking of n-Cv (TOS=2.5 h), d) distribution of C7 hydrocarbons (DAM hybrid catalyst; the same experiment as Fig. 1).
368 The reactivity of the C7 isomers was in the order n-heptane > 2-methylhexane > 3methylhexane > 2,4-dimethylpentane. 2,4-Dimethylpentane was the least reactive in the isomers. The reactivities of the isomers may be explained by diffusion limited caused by medium-pore zeolites, whereas H-ZSM-5 sorb both n-paraffms and monomethyl-substituted paraffins at fast rate even at room temperature [8]. The following results can not be explained reasonably by the generally accepted classical theory of acid-catalyzed cracking, which assumes a 6-scission mechanism for breakages of a C-C bond in a classical type of carbenium ion. The selectivity for cracking reaction was in the
n-heptane
Urn
o
>
c o U
90 80 70 60 50 40 30 / 20 - m 10 1 01 L 0.5
2-methylhexane
^ —
1 1
1 2
\ 1.5
1
1 2.5
0.5
W/F(ghmorM
'•'
'
' \ '
^^(g^-""' Conversion(%)
2
2.5
W/F(g h mol"^)
3-methylhexane
°
1.5
1
2,4-dimethylpentane
'•' >
• Sel. of Cracking(%)
0
0.5
1
1.5
2
2.5
W/F(ghmor') A Sel. of Isomerization(%)
503 K, H2/n-C7=9, 1.1 MPa, Pd/Si02:H-ZSM-5=l:l, Figure 5. Hydrocracking of C7 isomers over Pd-H-ZSM-5 hybrid catalyst.
369 order n-heptane > 2-methylhexane > 3-methylhexane > 2,4-dimethylpentane. The cracking selectivity of the n-heptane conversion was more than 40 % even at very low conversion level. The result means that n-heptane cracking dose not needed preisomerization. It can be explained reasonably by the reaction path shown in Fig. 6. The essential point of this reaction mechanism is that the cracking reaction proceeds through a intermediate containing the protonated dialkylcyclopropane structure which has been proposed by Sie [3]. According to his theory, the smallest paraffm chain which can be cracked is n-Cv. 2,4-dimethylpentane, 2methylhexane and 3-methylhexane must be isomerized to n-heptane before cracking. And cracked fragments must be isobutyl-carbenium ion and propylene. The propylene can be hydrogenated to propane over palladium catalyst in the presence of hydrogen. The hydride (Hso) as a counter anion of proton (H+so), which is formed in hydrogen spillover process, stabilizes the isobutyl-carbenium ion to give isobutane. Thus oligomerization of the cracked fragments and consecutive cracking reaction is prevented in the H2-Pd-hybrid system and hybrid catalyst gives only isomerized heptane and propane and equimolar amount of i-butane.
H
H
1 H
H
1
H
H
1
H
1
H
H
1
l/C\i/9^i/C^i 1 1
H- - c
1 H
1
H
H
H
+H"
H
H I H
-H
I
H
I
^C—H
I
H
I
H
H
H
I
H
H
I H
^?>.l
H I
H
Intramolecular Hydride Shift H
H
H
H I H I
H
H
H I
H
Hydride Shift H~X H
C*
Hydride Shift H / _ Scission ^ i / ° v ....-C. I -C. I ,=**^\"^ H-C .A H " I H I H
.+H
H^ ^H
H
Pd
Figure 6. Reaction model of n-heptane hydrocracking on Pd/SiOa-H-ZSM-S hybrid catalyst.
370 4. CONCLUSION A hybrid catalyst, which was prepared by physically mixing of a H-ZSM-5 and a Pd/Si02, showed an excellent activity for the hydrocracking of n-paraffins. Hydrogen gas is dissociated on the palladium on Si02 and spills over onto the H-ZSM-5. The spillover hydrogen presumably exist on the zeolite surface as proton and hydride. The proton promotes cracking reaction. The hydride generated simultaneously stabilizes intermediate carbenium ion to prevent over-cracking and promote isomerization of alkane. In the n-heptane cracking, the hybrid catalyst gave only isomerized heptane and propane and equimolar amount of i-butane whereas the products on H-ZSM-5 alone distributed from C3 to C9 and C4 products contained all kind of paraffins and olefins. The simple products for the H2-hybrid system should be formed through no other reaction path than the primary cracking reaction on H-ZSM-5, which proceeds through a intermediate containing the protonated dialkylcyclopropane structure, while the hydrocracking reaction over Pd-mordenite hybrid catalyst proceeds via 6-scission of the carbenium ion.
REFERENCES 1. 2. 3. 4.
B.S. Greensfelder, H.H. Voge and G.M. Good, Ind. Eng. Chem., 41 (1949) 2573. K. Fujimoto, K. Maeda and K. Aimoto, Applied Catal., 91 (1992) 81. S.T. Sie, Ind. Eng. Chem. Res. 31 (1992) 1881 . M. Steijns, G. Froment, P. Jacobs, J. Uytterhoeven and J. Weitkamp, Ind. Eng. Chem. Prod. Res. Dev., 20 (1981) 654. 5. P.B. Weisz, Pure Appl. Chem., 52 (1980) 2091. 6. J.D. Hargrove, G.J. Elkes and A.H. Richardson, Oil and Gas J., 77 (No.3, Jan. 15, 1979) 103. 7. P.B. Weisz, Proc. 10th Worid Petr. Congr., Vol. 4, 1980. 8. N.Y. Chen and W.E. Gawood, J. Catal., 52 (1978) 453. 9. P.A. Jacobs, H.K. Beyer and J. Valyon, Zeolites, 1 (1981) 161. 10. J. Weitkamp, P.A. Jacobs and J. A. Martens, Appl. Catal., 8 (1983) 123.
® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
371
Tail-selective Hydrocracldiig of Heavy Gas Oil in Diesel Production M.V.Landau, L.O.Kogah and M.Herskowitz The Blechner Center for Industrial Catalysis and Process Development, Ben-Gurion University of the Negev, Beer-Sheva 84105, Israel Hydrocracking of straight run heavy atmospheric gas oil (HAGO) was carried out in a batch high pressure autoclave and fixed-bed high pressure minipilot at 50-55 atm. and temperatures 350-390^0 with Ni-Mo-Al catalysts containing zeolite HY, H-ZSM-5 and mixtures of those zeolites. Addition of zeolite H-ZSM-5 to the Ni-Mo-Al-HY catalyst increased its activity significantly beyond corresponding increase of HY zeolite loading. It reduced the 90 and 95% boiling points (BP) of diesel fraction by 17-30OC and pour point (PP) by about 12-20^0 depending on zeolite H-ZSM-5 loading and testing conditions. The overall conversion of HAGO with the bi-zeolite catalyst was not additive yielding lower gasoline yields compared to the sum of the yields obtained with mono-zeolite catalysts. Based on GC-MS analysis of the HAGO, tail fractions those effects were attributed to increased conversion of the least reactive paraffins as a result of H-ZSM-5 zeolite addition. 1. INTRODUCriON The growing demand for diesel fuels requires expension of the diesel pool. One of the solutions here involves HAGO with 95% BP >390OC. Adjustment the 95% BP ( < 370^0 , PP (< 5 -^ -20OC) and density ( 0.800 - 0.845 kg/L) by normal hydrotreating in this case is still a technical challenge. One of the schemes includes hydrotreating of HAGO to reduce the sulfur content mixing it with up to 50 vol.% light atmospheric gas oil (LAGO) with FBP < 320^0 (hydrotreated or not, depending on diesel sulfur content requirements) to reduce the 95% BP, PP and density. The disadvantage of this scheme is the limited contribution of HAGrO in fingd diesel fuel that limits the potential resources in diesel production. Tail-selective hydrocracking of HAGO could increase the resources for diesel saving LAGrO as high quality component of diesel pool that needs only desulfurization. Mild hydrocracking of middle distillates with zeolite catalysts adjusts their distillation patterns, pour point and density [1]. The problem here is to produce a catalyst that converts mostly the tail part of HAGO boiling out at 350^0+ into products boiling in middle distillate range so that the gasoline yield will be much less than the amount of LAGO needed for adjusting the diesel quality. It is known that with most multi-function catalysts containing acidic zeolite Y, the hydrocracking reactivity of different hydrocarbon groups existing in gas oils varies in range of an order of magnitude where the paraffins and especially normal paraffins are the least reactive compounds [2]. The tail fractions of hydrocracked gas oil which composition is almost not affected by hydrocracking products are enriched with paraffins . Hydrocracking catalysts based on H-ZSM-5 zeolite selectively convert normal paraffins [3,4]. Combination of those two zeolites in one catalyst could improve its hydrocracking efficiency increasing selectivity to the tail fraction.The scope of this study was to test the effect of zeolite H-ZSM-5 addition on Ni-Mo-Al-HY catalysts performance and properties of diesel fraction.
372
2. EXPEEHMENTAL SECTION. Feedstocks., Two kinds of atmospheric gas oils from Ashdod ORL (Israel) boiling in range 157-396 and 175-343^C were used in present study. Their properties and composition are summarized in Table 1. HAGO-1 with 90% BP 3720C, 95% BP 390OC and PP +110C needs substantial upgrading of tail fraction to meet the specification requirements: 90%BP < 357^0, 95% BP <370OC , while HAGO-2 fits those requirements. HAGO-2 was the first feedstock used in this study because of the low nitrogen content that increases the hydrocracking efficiency of the catalysts making the effects of zeolite composition more pronounced . Table 1 Characteristics of atmospheric gas oils Feedstock # Density (150C),k^ Potir point, ^C Sulfur content, wt% Nitrogen content, ppm Distillation, OC: IBP 10% 30% 50% 90% 95% FBP
1 0.841 +11 0.95 480
0.832 -2 0.73 86
157 251 298 320 372 390 396
175 242 278 290 319 337 343
Chemical composition according GC-I Paraffins Naphthenes Aromatics
44.1 39.1 16.8
37.7 44.0 18.3
Catalysts. The catalysts used in present study were two commercial Ni-MoAl (hydrotreating) and Ni-Mo-Al-HY (hydrocracking) and six proprietary samples prepared by introduction of HY gmd HZSM-5 zeolites to Ni-Mo-Al composition or of Pd/HY to alim^Lina. Their properties and composition are listed in Table 2. HY zeolite LZ-Y62 manufactured by Linde AG and H-ZSM-5 zeolite CBV-3020 manufactured by PQ Corp. were used for catalysts preparation. They were mixed with aluminum hydroxide, formed by extrusion and the pellets were then impregnated with ammonical solution of nickel nitrate and ammonium paramolybdate. Psdladium was loaded to zeolite HY by impregnation with Pd(NH3)4Cl2 water solution, after that Pd/HY zeolite was formed by extrusion with gdumintmi hydroxide binder. Catalysts testing, The catalysts testing was performed in a high pressure, 3/4 gallon, autoclave model E-12-0554, manufactured by Autoclave Engineers Inc and fixed-bed high pressure minipilot at 50-55 atm and 390OC.
373 Table 2 Characteristics of catalysts in oxide form Catalysts
Commercial Commercial Ni-Mo-Al Ni-Mo-Al-HY
Chemical composition,wt% NiO 4.0 M0O3 19.2 Pd Zeolites: HY H-ZSM-5 Surface area, m2/g 190 Pore volume, cm3/g 0.42
HCN-1 HCN-3 HCN-4 HCN-5 HCN-6 HCN-7
3.3 16.7
7.3 13.8
5.2 13.9
7.1 15.6
25
37
20 18
18
320
350
330
0.64
0.69
0.65
3.0
7.6 7.3 16.2 15.3
70
20
280
420
320 310
0.54
0.81
0.72 0.42
25 7
The batch experiments were run for 6 hours at stirring speed of 900 rpm with a catalyst- to-oil vol. ratio 0.05. The total reaction pressure was controlled at the designed level by adding fresh hydrogen into autoclave to compensate for its consumption. All Ni-Mo-based catalysts were presulfided before loading to the reactor while the Pd-catalyst was reduced in hydrogen flow at SOO^C. Additional testing at P = 55 atm, T = 3550C, LHSV = 0.5 h-l and VH2 = 1500 NL/L were performed in a fixed-bed high pressure minipilot equipped with high and low pressure separators, gas flow controllers at the inlet and outlet and online GC gas analyzer. 50 cm^ of presulfided catsdysts were loaded between two layers of silicon carbide pellets. Feed/products analysis. The liquid products were analyzed for distillation distribution (ASTM D-86), density at 150C (ASTM D-1298), pour point (ASTM D97) and sulfur content (X-ray flourescence, Oxford Lab X 1000) before and after distilling off the light gasoline fraction IGO^C-. The efficiency smd selectivity of HAGrO mild hydroracking was calculated according to yield of diesel fraction leO^C-h, gasoline content in the liquid product, 90% or 95% BP and PP of diesel fraction. The chemical composition of liquid products was carried out by GCMS group anedysis allowing to define the chemical compovmds according to characteristic ions [5].The GC-MS analyzer "HP MSD 5917A" with capillar coltmin containing SE-30 phase was used by programming the thermostate from 70 to 250^C at a heating rate 20^C/min, injector temperature 250^C and temperature of trgmsfer line 200^0. The scanning range was 40-^450 mass, registration time 1.5 sec. In order to obtain the mass-spectrum of the whole sample the intensities were summarized for all time of analysis [6]. Calculations of the group composition were carried out according to [7]. The relative contents of high molecular hydrocEirbons responsible for boiling points of the tail fractions were estimated from the intensities of molecular peaks of paraffins, naphthenes and gdkylbenzenes relative to the full ion current:
374
Cin=(Im-Im-l-k)/FIC, where Im - intensity of molecular ion, Im-i intensity of ion with mass (m-1), k isotopic correction coefficient, FIC - full ion current. The gas products C1-C5 were analyzed by GC method ( GOMAC 580) with TCD detector, packing column VZ-7. 3. RESULTS AND DISCUSSION. Four catalysts were tested in hydrocracking/hydrotreating of HAGO-2 to study the effects of Ni-Mo-Al matrix and acidic zeolite components inserted in this matrix: separate zeolites HY and ZSM-5 and the random zeolite mixture. The distillation patterns of the total liquid products obtained with Ni-Mo-Al catalyst and three zeolite catalysts containing zeolite HY, H-ZSM-5 and HY-HZSM-5 mixture are compared in Fig.l with corresponding data for the feedstock. Ni-Mo-Al matrix caused a significant change in BP distribution producing about 8% of gasoline and shifting the BP to the lower temperatures by 54(20%) -f-
130
180
230
Temperature, oC
Fig.l. Combination of HY and H-ZSM-5 zeolites increases the "tailselectivity" in hydrocracking of HAGrO-2. 11(90%)OC so that 90% BP was reduced from 319 to 308OC. That is a result of combination of hydrocarbons hydrocracking reactions with desulfurization and saturation of aromatics, producing lower boiling compounds. Introduction of zeolite HY yielded an additional shift of BFs compared with Ni-Mo-Al matrix as a result of hydrocracking reactions - almost evenly over the full boiling range. That accounts for about equal conversion of hydrocarbons boiling out at the overall feedstocks range. The gasoline content in the product reached 15 vol.% and 90% BP was 304OC. Zeolite H-ZSM-5 showed much deeper conversion of the feedstock - the gasoline content in the product became 25%, but the 90% BP was reduced only to 3140c. The efficiency of this zeolite for the tail fraction of the feedstock
375 was very low comparing with zeolite HY. Combination of two zeolites (catalyst HCN-3) gave 18% vol. gasoline in the liquid product with maximum depresion of 90% BP to 2970C. That is evident for higher tail-selectivity of the bizeolite catalyst compared with both mono-zeolite catalysts. The three Ni-Mo catalysts containing zeolites HY, H-ZSM-5 and their r a n d o m mixture were tested in hydrocracking of heavier HAGO-1 in a fixed-bed reactor. The results are listed in Table 3 and the GC charts of HAGO-1 £ind diesel fractions obtained in fixed-bed experiments with HY- and (HY+H-ZSM-5) containing catalysts are shown in Fig. 2. Fig. 3 demonstrates the changes in the group composition of the HA GO tail fraction (hydrocarbons containing >25 and >26 carbon atoms in a molecule that were identified in fraction boiling out at >360OC). The same trends in selectivity to the tail fraction were observed as in batch e3q)eriments with HAGO-2. Table 3 Fixed-bed pilot plant testing results with HAGO-1
Feedstock)
HCN-6
Products yields, wt.%: __ - diesel fraction 160^0+ 88.0 — - gasoline IBP-160^0 8.9 3.1 - gas C1-C5 Properties of diesel fraction: 0.841 0.830 - density ( 1 5 ^ 0 , g/cm^ 372 359 -90%BP,oc 390 373 - 95% BP, oc - pour point, ^C +11 ^ - sulfur content, wt.% 0.95 0.04 Content of heavy hydrocarbc IS in diesel fraction, au.: -C25+ 100 77
100 80 Content of paraffins in heav hydrocarbons, a.tL: - in C25+ 100 129 100 141 -inC26+
HCN-4
HCN-3
76.4 14.3 9.3
73.2 16.0 10.8
0.832 366 382 -9 0.09
0.820 352 363 -11 0.03
88 86
62 64
68 72
70 66
The two effects recorded in those experiments: i) introduction the H-ZSM-5 zeolite to Ni-Mo-Al-HY catalyst increased the activity and hydrocracking selectivity to the tail fraction of feedstock and ii) the conversion of the feedstock with bi-zeolite catalyst (gasoline yield) was less than the sum of the conversions obtained with mono-zeolite catalysts. The former effect could be explained considering that different groups of hydrocgirbons in petrolevim gas oil are converted in hydrocracking process with different rates. According to [2] the first-order rate constants for hydrocracking of sulfur compounds (SO), polycycioaromatics (PGA), mono-cycloaromatics (MCA), poly-cycloparaffins (PCP), mono-cycloparaffms (MCP), iso- (IP) and normal paraffins (NP) measured at standard conditions by changing the space time decrease in the following sequence: SC (1.00) >MCA (0.72) >PCP (0.65) >PCA(0.61)> MCP(0.36) > IP (0.26) > N P (0.12),
376 IGglag^;
aHAGO-1 1.56+07-^
|\bundance 1.2e+07
b. diesel fraction with Ni-MoAl-HY catalyst
jrime ->
LuiJM 5.00
c. diesel fraction withNi-MoAl-HY-HZSM-5 catalyst
Fig.2 . GG charts of HAGO-1 (a) and diesel fractions obtained in fixed-bed experiments with HCN-6 (b) and HCN-3 (c) catalysts where the relative values of the rate constants are shown in the brackets. It means that hydrocracking products obtained with zeolite HY-based catalyst and especially their tedl fraction whose composition is almost unaffected by the fragments of hydrocracked molecules should be enriched with paraffins as it was measured by GC-MS (Table 3). Zeolite HZSM-5 in hydrocracking catalysts displays strong shape selectivity to normal paraffins and long-chain containing
377
3
Feedstock
HY
ZSM
0.01 y
Feedstock
MY
ZSM
HY + ZSM
Fig.3. Distribution of paraffins (P), mono-naphthenes (MN), binaphthenes(BN) and alkylbenzenes (AB) in tail parts of diesel fractions obtained in fixed-bed experiments with HAGO-1 molecules [3,4]. Therefore the other groups of hydrocarbons in tail fraction remain less converted (Fig.3) yielding higher overal concentration of heavy hydrocarbons (Table 3) and minimum depression of 90% BP. The combination of two zeolites - HY and H-ZSM-5 in one catalyst, strongly increases the hydrocracking tail-selectivity yielding less heavy hydrocarbons in diesel fraction (Table 3) and maximum depression of the 90% BP. Zeolite H-ZSM-5 in addition to non-selective hydrocracking of HY zeolite selectively removes the remaining paraffins in the tail fraction (Fig.3). The good resolved peacks on the GC chats shown in Fig.2 at retention times > 13 min. which intensity was substantialy reduced after addition of H-ZSM-5 zeolite to Ni-Mo-Al-HY catalyst belong to normal paraffins according to MS analysis. Addition of H-ZSM-5 zeolite to Ni-Mo-Al-HY catalyst increased the gasoline yield much less than it could be expected taking into accoimt the comparatively high gasoline yield with Ni-Mo-Al-H-ZSM-5 catalyst. That is a result of a well established fact that the conversion in hydrocarbons cracking reactions with those two zeolites used in series or in rajcrdom mixtures is not additive [8,9]. In this case additivity is the crackability defined as conversion/(100-conversion) [9]. Table 4 illustrates the effect of zeolite loading on the catalyst performance in hydrocracking/hydrotreating of heavier HAGO-1 measured in batch runs. Increasing the HY-zeolite content in Ni-Mo-Al matrix from 20 to 37% and especially to 70% in combination with palladium gradually increases the gasoline yield, decreases the 90% BP and PP. But even after two-stage hydrocracking with Pd-catalyst, at the second stage, in spite of high conversion of the feedstock (35% gasoline in the liquid product), the 90% BP did not reach 357^0 corresponding to specification requirement. Increasing the content of second zeolite H-ZSM-5 also increases the conversion of the feedsock yielding lower values of diesel 90% BP ( < 3570C) and lower PP as a result of more eJBfective removal of normal paraffins. Those data conform to the results obtained in [10] for hydrocracking of vacuum gas oil with Ni-Mo-Al-zeolite catalysts. In that case the addition of zeolite H-ZSM-5 to Ni-Mo-Al-HREY increased the conversion of the feedstock by about twice compared with the
378 Table 4 Hydrocracking of HAGO-1 in batch reactor: effect of zeolite loading Catalysts
Results: Gasoline content in liquid product,vol%
-- (Feedstock) HCN-6 (Ni-Mo20%HY) Commercial (Ni-Mo-25%HY) HCN-1 (Ni-Mo37%HY) HCN-5 (Pd70%HY*)) HCN-7(Ni-Mo25%HY-7%HZSM-5) HCN-3 (Ni-Mo20%HY-18%HZSM-5)
Diesel fraction 160OC+: Density 90% BP, Pour Sulfur (150C), oc point, content g/cm3 oc wt%
—
0.841
372
+11
0.95
10
0.828
363
+10
0.08
13
0.824
362
+10
0.05
17
0.829
360
-tS
0.05
35
0.810
358
-+3
0.03
13
0.826
357
-1
0.04
19
0.818
355
-9
0.05
*) Feedstock: the product obtained with catalyst HCN-1 at the same conditions, corresponding rise of conversion yielded from increasing the HREY zeolite loading. The results obtained in this study demonstrate an interesting option in production of diesel fuels from heavy atmospheric gas oil using the bi-zeolite catalyst where its composition and operating conditions should be optimized. REFERENCES 1. L.R.Aalund and A.Cantrell, Oil & Gas J., 89(13),63,1981. 2. R.Krishna, Erdol und Kohle-Erdgas-Petrochem., 42(5), 194,1989. 3. N.Y.Chen and W.E.Garwood, Catal.Rev.-Sci.Eng., 28(2&3), 185,1986. 4. B.K.Nefedov, M.V.Landau and L.D.Konovalchikov, Chemistry and Technology of Fuels and Oils, 10,4,1988 (Russian). 5. I.Dridic, H.A.Petersen, P.A.Wadsworth and H.V.Hart, Anal. Chem., 64,2227,1992. 6. A.A.Polyakova, L.O.Kogan and G.L.Vasilenko, Scientific Papers of the Prague Institite of Chemical Technology, Prague, 1982, p.105. 7. RS.Brodsky, Neftechimia, 40,57,1985 (Russian). 8. Z.Zainuddin, F.N.Guerzoni and J.Abbot, J.Catal., 10,150,1993. 9. R.W.Mott, Oil &Gas J., 26,73,1987. 10. B.K.Nefedov, "Adsobents, Supports and Catalysts for Oil Refining", Proc. of the All Union Research Institute for Oil Refining, ISSN 0202-398, Moskow, 1990,v.62. p.108 (Russian).
® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
379
Influence of the hydrocarbon chain length on the kinetics of the hydroisomerization and hydrocracking of n-paraffins. B.Debrabandere and G.F. Froment Laboratorium voor Petrochemische Techniek, Universiteit Gent, Krijgslaan 281, B-9000 Gent, Belgium The product distributions obtained by hydrocracking of several n-paraffins were determined on a Pt/US-Y catalyst in a gas phase reactor with complete internal mixing. The experimental results indicate that the components of the mixtures react independently: they only influence each other by competitive physisorption. Reactivity and strength of adsorption in the zeolite pores increase with chain length of the paraffins. At higher temperatures and lower pressures, the product distribution is no longer a unique function of conversion. The hydrocracking is no longer 'ideal'. The isomerization conversion decreases with increasing temperature and decreasing pressure. The longer the chain length, the more important this phenomenon. It is attributed to a relative decrease in the rate of hydrogenation, compared to the rate of the carbenium ion reaction. The experiments were modeled by means of two lumped kinetic models. A first model assumes that the carbenium ion reactions are rate determining. This gives a good fit for all experiments where the hydrogenation reaction is in pseudo-equilibrium. A second model also explicitly accounts for the rates of dehydrogenation. This gives a good fit for all experiments, also when the isomerization conversion is no longer a unique function of conversion. 1. INTRODUCTION. Hydrocracking is widely practised in modern petroleum refining for the conversion of heavy destillates and residua into more valuable fractions, such as gasoline, diesel and jet fuel. The development of kinetic models for hydrocracking is important for catalyst development, for reactor design and simulation. It is generally accepted [1-3] that hydrocracking proceeds through physisorption, then dehydrogenation on the metal sites of the bifunctional catalyst, followed by protonation on the acid sites. The carbenium ions thus formed can then undergo isomerization and /5-scission, followed by deprotonation and hydrogenation. Recentiy, a single event kinetic model was developed [4]. The model takes into account the rates of the elementary reactions of the carbenium ions - which are assumed to be rate determining. Parameters were determined from experiments with nCg and isomers [5]. However, more information is needed on the influence of the chain length on the kinetics. Therefore, experiments were performed with several normal paraffins and mixtures of them. In a first step, the experiments were modeled by means of a lumped model, so as to simplify the parameter estimation.
380 2. EXPERIMENTAL SECTION. 2.1. Description of the equipment. The experiments were performed in a Berty reactor, a gas phase reactor with complete internal mixing. On-line gaschromatographic analysis was used to quantify the products. The equipment is already described elsewhere [3]. The only modification is, that in the present work methane was used as an internal standard. 2.2. Experimental conditions. The experiments were performed with a Pt/US-Y catalyst, containing 0.68%Pt. Several pure n-paraffins (nCg, nC^o, nCi2, ^C^^) and also mixtures of n-paraffins (nCg/nC^Q, nCjo/nCjjL/^^12/^^13 (Parapur) and nCi2/nCi4) were fed. Data were collected over a wide range of experimental conditions: temperatures between 200-280°C, pressures between 5 and 70 bar, space times of 20 to 200 g catalyst.h/mole feed and hydrogen to hydrocarbon ratios (y) of 30 to 400. 2.3. Hydrocracking of pure n-paraffins. The following general conclusions can be drawn from the experiments: - Conversion increases with increasing temperature, space time and decreasing molar ratio hydrogen to hydrocarbon (y). - At lower temperatures and higher pressures an increase of the pressure causes a decrease in conversion, while for high temperatures and low pressures, the conversion increases. The longer the chain, the more important this phenomenon. - For a certain conversion, the conversion to isomers decreases when the temperature increases above a certain temperature and/or the pressure is below a certain value. The longer the chain length, the more these effects are pronounced. The phenomenon is illustrated for nCi2 in Fig. 1. The squares are experiments for which the carbenium ion reactions are rate determining. The curves defined by other symbols relate to conditions whereby this is not so.
40
X(%)
Figure 1: Isomerization conversion versus total conversion for nC|2-
381 These curves are very similar to those of Giannetto et al. [6] and Degnan and Kennedy [7], which describe the influence of Pt% on the isomerization selectivity of nCy. In that case, lower selectivities are obtained when the Pt-loading is lower. Such a behavior would be observed when the rate of the carbenium ion reactions on the acid sites is no longer rate determining. In the terminology of Weitkamp [8] the hydrocracking would no longer be 'ideal'. The reason that the effects are more pronounced for heavier feed, is that longer chain paraffins are more active on the acid sites, because they undergo more reactions. - For the experiments for which 'ideal hydrocracking' occurs, monobranched isomers are the only primary products, dibranched isomers are secondary products (Fig 2). For the other experiments, dibranched isomers and cracked products can be formed directly out of the feed. - For all experiments, the distribution of the monobranched isomers is a function of conversion only. This indicates that the dehydrogenation rate is the same for all monobranched isomers. The same distributions were obtained by Steijns et al.[3]. - The distribution of cracked products is symmetrical: no secondary cracking is observed (Fig.3) lUU -
CM 8 0
o
N.
•D
(D
^ \ ^
(D O60
0 0 0
^^fc.
^^.
f-
^iL^
^^»
o
^^nC12 ^^__^MBC12
?40
^
o
^^+ij>4*--**F# rzz----^
^^,#^^^^o.«^''*^^
E20
DBC12
E
0 0
(D 0
O-
20
40
X (%)
60
80
Figure 2: Distribution of the C^2 fraction in C^2 hydrocracking.
100
Xcr (%)
Figure 3: Distribution of the cracked products in €^2 hydrocracking.
2.4. Hydrocracking of mixtures. - The reactivity increases with chain length. This is illustrated in Figure 4, which shows the conversion of nCg versus the conversion of nC^g ^^^ experiments, performed with the pure components at the same experimental conditions and for experiments, performed with an equimolar mixture of both. The conversion of nC^g is always higher than the conversion of nCg, which means that nC^g is more reactive: more reactions are possible for heavier paraffins. The curve of the mixture falls below the curve of the pure components, indicating that the heavier hydrocarbons are more strongly adsorbed in the zeolite pores. - For mixtures, in which one component cannot be formed by cracking of the other, the product distributions of the individual components are identical to those obtained for the pure components. This means that the components influence each other only by competitive physisorption. Bimolecular reactions, such as hydride transfer or alkyl shifts do not proceed.
382 50 + pure feed (same conditions) 40
• mixture
1
XC10(%)
60
80
T"
100
Figure 4: Conversion of nCg versus the conversion of nC^Q for experiments with the pure
3.KINETIC MODELING. 3.1. Reaction scheme and general rate equations. The reaction scheme retained e.g. for nCj2 ^^ given in Fig.5. In this scheme monobranched isomers are lumped, since they were shown to be at equilibrium. Physisorption in the zeolite pores is described by the Langmuir isotherm, which was shown by Steijns and Froment [10] to give the best fit of the experimental data. (1)
with Ky the physisorption constant for component i. The rate of dehydrogenation on the metal sites is written, according to Dumez and Froment[8], but in terms of partial pressures. k ^DH,i
with:
\PP:
P0;PH,
KDH; D'M
(2)
383 MBC
multl a12
MB(
multl C. '12
'12
^1
12 11 13 nC+ ^ ^ = ^ MB C+2^=^^ multl C+2
14
C+^ ^
15
16
^t
10
Figure 5: Partially lumped reaction scheme for hydrocracking of nCj2.
^i = DM
^DH^^f^Mf^U
i.E<
'' ^^H^LPi
^^Y^lPi
(3)
The protonation of the olefins is so fast, that it can be considered to reach pseudo-equilibrium [5]:
0;^H* % c:
(4)
^C, =
'^OpH Kcj is the chemisorption constant of the component i. A balance on the acid sites yields: i
^
(5)
i
The rate equations of the various carbenium ion reactions numbered as in Fig. 5 are given
384 in (6), (7), (8). rn = ^ i i ( V ; 3 - ^ A / < / ^ i i >
(6) (7)
'•12 = kn(<=MBc:,-'muUic:J^n) « = 13,14,15,16
'i =''i'nuuuc:.
'12
(8)
Combining equation (6) with equations (4), (5) and (1) leads to:
^11
l+E^L,/'i
(9)
^TPOKC,
with:
K-li = Ki^K^
IK^
Kr
i
/in\
IKj ^12
"^12
i
The other reaction rates can be expressed in the same way. Finally, equations are written for the net rate of formation of paraffins and olefins: Paraffins: RnC,,
= -h(PnC,,-P„c:PHjK{)ID'^
(11)
Olefins: Kcl,
- h^nC,,-P„clPHjK,)ID^
-^'nKc-PMBClJ^'n)IDA
(12)
Since the concentrations of the olefins are very low, pseudo-steady state can be assumed for these components: Ro. = 0
(13)
From equation (13), the partial pressures of the paraffins can be expressed as a function of the partial pressures of the olefins, e.g. for nCy^.
385
This leads to the following set of non-linear equations. (15)
P = AO The partial pressures of the olefins can then be calculated from:
(16)
0=A-^P
D^ still contains the concentrations of the olefins, so that the system has to be solved iteratively. The partial pressures of the olefins are then substituted into equation (12), yielding the reaction rates of the paraffins. For the special case, that the carbenium ion reactions are rate determining, the dehydrogenation reactions are in pseudo-equilibrium. Substitution of equation (14) into equation (11) leads to: PMBCI^ ^«c.
= -*I'I(^A)"'
''^12
(17)
r^f
K'n
Replacing the partial pressure of the olefins by means of the expression for the dehydrogenation equilibrium, derived from equation (2) leads to: ^«C,, = -
- ^ '' , ^ff3(i+E^i/'i)^E^c,/'i i
(18)
i
This is the same rate equation as developed by Steijns and Froment [10]. 3.2. Parameter estimation when the carbenium ion reactions are rate determining. The experiments performed with the different feeds, in conditions where carbenium ion reactions seemed to be rate determining, were modeled with the model from 3.2. Parameter estimations were performed by both the Rosenbrock [11] and the Marquardt [12] routines. The following observations were made: - Chemisorption constants were not significant - Physisorption of the cracked products was never significant. - Only one physisorption constant could be estimated for each carbon number, i.e. no distinction can be made between the isomers. - One activation energy was estimated for isomerization and one for cracking and this for each carbon number.
386 - Only one heat of adsorption was estimated for all the carbon numbers. The parameters values are given in Table 1. These are all significant. Table 1: Parameter estimates. Cg
^10
Cii
^12
^13
Cl4
0.35 0.29 0.35 0.17 0.37
2.4 2.5 1.9 0.36 0.98 1.2
3.0 3.5 2.4 0.41 1.3 1.5
5.1 7.0 4.7 0.46 1.5 1.8 1.8
10.7 14.2 11.8 1.2 2.9 5.1 5.2
5.4 107
18.6 94 95
21.7 123 86 61
24.8 100 85
62.7 118 106
16.3 23.0 17.6 0.94 3.7 4.9 5.3 5.5 96.4 101 75
ki (240°C)
nq ^ MBq MBCj -H^muiq muiq 'H. Cr mulC- ^ C3 muiq H ' ^ C4 muiq H ' ^ C5 muiq H ' ^ €5 muiq -^ Cy KL (bar-l) Ea (iso) (kJ/mol) Ea (cr) AH (all Ci)
The parameter values increase with carbon number (Figure 6).
70 60 /—N
-j-
50 JCQ ^^ 40 -L
.fi
0
s
30
„^
?() ^ 10 0
-jJ •
100 An-MB 0 MB-multi • multi-cr • KL
80
•/
ou
40
m^^^.^^
-^^-^"^
20
1 ffr—|""-f 1 ^ i 10
1
11
12
\ 13
[ 14
C-number Figure 6: Parameters as a function of carbon number.
Parity plots are given in Fig. 7
0
u
OS
J
387 3E-02
8E-03
OE+00 OE+00
OE+00 OE+00
1E-02 2E-02 FMBC12 (exp.)
3E-02
2E-03
4E-03 6E-03 FC7 (exp.)
8E-03
Figure 7: Parity plots for the 'ideal hydrocracking'.
3.3. Parameter estimations for the general case. Parameters were estimated for the experiments with nCi2 ^^^ nC^4. For nCg, nC^g ^^^ Parapur, the number of experiments with 'non ideal behaviour' was insufficient. Kj^ and K^.^ were not significant. The resulting parameter estimates are given in Table 2. For the isomerization and cracking reactions, the parameter given is k|Kj)j|^. These were estimated instead of Iq, as otherwise the parameters would differ too much in magnitude. Table 2: Parameter estimates (general case) ki (240 °C) nq "H. n q = MBq ^ MBq= muiq 'f* muiq^
nq=
^MBq=
MBq= muiq= muiq= muiq= muiq= muiq=
'H^muiq^ "H. C3 H ' ^ C4 'H. C5 "H. C6 'H. C7
KL (bar-l) E^ (DH) (kJ/mol) Ea (iso) Ea (cr) AH
nCi2
nCi4
0.61 0.21 0.075 5.2 2.3 0.066 0.24 0.28 0.13
0.9 0.86 0.65 17.6 7.3 0.31 1.3 1.7 1.8 0.96 36.5 74 169 165 18
12.3 24.4 158 152 18
388 Parity plots are given in Figure 8. The fit is generally good, although the net rate of isomer formation is sometimes overestimated. 2E-02
1E-01
OE+00 OE+00
5E-02 FnC12(exp.)
1E-01
OE+00 O.OE+00
1.0E-02 FC7 (exp.)
2.0E-02
Figure 8: Parity plots for the general case. Again, the parameters increase with carbon number. Those for the dehydrogenation are lower than those for the isomerization. The activation energies for the latter are much higher so that at higher temperatures the dehydrogenation will be intrinsically slower than the carbenium ion reactions, leading to 'non-ideal' behavior. The parameter values for the isomerization and cracking reactions differ from those determined by means of the specific 'ideal' model, possibly because of a slight correlation between the rate parameters and the physisorption constants or because the experiments with nCg and nC^Q were not used in the parameter estimation based upon the general model, to avoid the predominance of 'ideal' behavior. 4. CONCLUSION. The behavior of n-paraffins in hydrocracking significantly depends upon the chain length. A model was developed that explicitly accounts for both the rates of dehydrogenation and of the carbenium ion reactions. This model fits the results of experiments with normal paraffin and mixtures thereof, also those obtained at lower pressure and/or higher temperatures, where the isomerization conversion is lower and the carbenium ion reactions are no longer rate determining. The model can be used to describe the influence of temperature and pressure on the isomerization selectivity of hydrocarbon conversion on a given catalyst, but also to investigate the influence of the relative strength of hydrogenation and acid sites of the catalyst on this selectivity. For a weaker hydrogenation component than Pt the dehydrogenation parameters will be lower than those reported here, for equal dispersion at least. Non ideal behavior is even more likely in that case.
389 REFERENCES 1. Martens J . A . , Jacobs P . A . , Theoretical aspects of Heterogeneous Catalysis, N e w York, 1990 2. Coonradt H . L . , Garwood W . E . , Ind. Eng. Chem. Proc. Des. D e v . , 3 (1) (1964)38 3 . Steijns M . , Froment G . F . , Jacobs P . , Uyterhoeven J., Weitkamp J . , Ind. Eng. Chem. Prod. Res. Dev. 20 (1981)654 4. Vynckier E. and Froment G . F . , Kinetics and Thermodynamic Lumping of Multicomponent Mixtures, Elsevier, Amsterdam, 1991 5. Svoboda G . D . , Vynckier E . , Debrabandere B . , Froment G . F . , Ind.Eng.Chem.Res.,34 (1995)3793 6. Giannetto G . E . , Perot G.R., Guisnet M . R . , Ind.Eng.Chem.Prod.Res.Dev.25 (1975)481 7. Degnan T . F . and Kennedy C.R., A . I . C h . E . 39 (1996) 8. Pichler H . , Schulz H . , Reitemeyer H . O . , Weitkamp J., Erdol, Kohle, Erdg. Petroch., Brennst. ch., 25 (1972) 9. Dumez F . S . , Froment G . F . , Ind. Eng. Chem. Proc. Res. Dev. 15 (1976) 2 9 1 10. Steijns M . and Froment G . F . , Ind. Eng. Chem. Prod. Res. Dev, 20 (1981)660 11. Rosenbrock H . H . , Storey C , Computational Techniques for Chemical Engineers, Pergamon Press Oxford (1966) 12. Marquardt D . W . , Soc. Ind. Appl. Math. J.,11 (1963)431
SYMBOLS. CpCj E^ kj K^,^ ^DH Kj Ky K I Pi R rj Rj T Xj AH y
concentration of paraffin Pj in the zeolite pores total concentration in the zeolite pores activation energy (kJ/mol) parameter chemisorption constant equilibrium constant for dehydrogenation equilibrium constant fysisorption constant (bar"^) adsorption constant for adsorption on the metal phase partial pressure (bar) gas constant = 8.314 J/mol/K reaction rate of reaction i(mol/h/g catalyst) rate of formation of component i (mol/h/g catalyst) temperature (K) conversion (%) adsorption constant (kJ/mol) molar ratio hydrogen to hydrocarbon
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® 1997 Elsevier Science B. V, All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
391
AROMATICS HYDROGENATION OVER SUPPORTED PLATINUM CATALYSTS: THE INFLUENCE OF SULFUR ON THE KINETICS OF TOLUENE HYDROGENATION OVER PtA^-ZEOLITE CATALYSTS. Hakon Bergem, Edd A. Blekkan, and Anders Holmen Department of Industrial Chemistry, The Norwegian University of Science and Technology (NTNU), N-7034 Trondheim, Norway.
In this paper we report some experimental results with Pt/Y-zeolite catalysts, using the hydrogenation of toluene as a model for aromatics hydrogenation in hydrotreating. The hydrogenation of toluene was studied under conditions close to realistic industrial conditions. The findings include a complex temperature behaviour with a maximum in the reaction rate close to 325 °C. The pressure dependence of the rate with respect to toluene and H2 was found to be sensitive to the level of sulfur in the liquid feed. 1. INTRODUCTION Increasing environmental concerns are leading to new regulations for fuel composition. For diesel fuels the main changes are lower sulfur levels and limitations of the contents of aromatics in diesel fuels. An example of this trend is the Swedish taxation class I (City Diesel), limiting the aromatics contents to 5 % by volume. Possible sources to such a diesel oil, either as a blending stock or as a pure fuel, can be synthetic diesel e.g. from Fischer-Tropsch wax hydrocracked into the diesel range. This gives a diesel fuel with excellent properties [1]. Biofuels with good properties are also produced, e.g. rapeseed oil methyl ester, coproducing glycerol in the process [2,3]. But the likely source to the bulk of the diesel fuel even at more stringent environmental requirements will for the foreseeable future be petroleum based, and hydrotreating and hydrocracking will be the key processes to their production. Hydrotreating over traditional sulfide catalysts has as its main target the removal of heteroatoms from the fuel. Sulfur and nitrogen can be removed with high efficiency, and usually the catalysts have been designed to perform the operation with the lowest possible hydrogen consumption. Co-Mo is the preferred system for hydrodesulfurization (hds) due to the high activity and selectivity for the hds reaction. Ni-Mo or Ni-W are reportedly more active for hydrodenitrogenation (hdn), and also give a higher degree of the usually unwanted hydrogen consumption due to aromatics saturation [4,5]. However, these catalysts have low activities for aromatics saturation, and even at pressures exceeding 100 bars and using long residence times
392 there is a large fraction of unconverted aromatics in the product. High temperatures are not beneficial in the highly exothermic and equilibrium limited hydrogenation reactions; with increasing reaction temperature a minimum in the aromatic content of the product oil is observed, usually around 350 - 380 °C [4,5]. Traditional aromatics hydrogenation catalysts are metallic nickel or noble metals, supported on high surface area supports. These catalysts are very active, even at low temperatures. The problem with these systems is their sensitivity to sulfur poisoning [6], since sulfur is an inherent feature of these feedstocks. Altemative hydrogenation catalysts that can be suggested include more active sulfides: Group Vm transition metal sulfides like the sulfides of Ir, Ru, Rh and Os are reported to be more active than Mo for hds [7]. A similar trend has not been reported for aromatics hydrogenation, but Ledoux et al. [8] did find a similar trend for hdn. This is relevant since hdn of aromatic amines mechanistically passes through a hydrogenation step before the C-N bonds are broken [9,10]. In a recent paper it was shown that sulfided Ft/y-Al20^ was about 20 times as active for toluene hydrogenation per g metal compared to a conventional Ni-Mo catalyst. Both systems were inhibited by H2S and NH3, but the reaction orders with respect to the poisons were higher for Pt than for NiMo [11]. It is also possible to speculate about other, less conventional, catalyst types. High surface area carbides and nitrides are now available [12,13]. Transition metal nitrides have recently been shown to be excellent hdn catalysts [14,15], again with the mechanistic implication that hydrogenation of the aromatic ring is a necessary part of the hdn mechanism. Carbides are known for their sulfur tolerance, but their hydrogenation activity is less known. They are reported to be active dehydrogenation catalysts [12]. However, industrially it is the zeolite-supported sulfur tolerant noble metal catalysts that have made an inpact so far. This development is based on the recognition of the special properties of Y-zeolite-supported Pt and Pd [16-19]. Hydrogenation processes utilizing these catalysts are now available, and the catalysts have been shown to tolerate high levels of sulfur and still maintaining their aromatics saturation activity. Examples of these processes have been given by Lucien et al. [5], Marchal et al. [20], and Cooper et al. [21]. The papers do not give details of catalyst preparation or properties, but indicate that the catalysts are "sulfur tolerant noble metal catalysts", in some cases also indicating that the support is zeolite based [5]. These papers report very high sulfur tolerance data, giving aromatics saturation at lower pressures than with mixed sulfide catalysts even in the presence of up to 1000 ppm sulfur in the feedstock. But so far very little is known about the properties of these commercial catalysts. The sulfur tolerance of supported metal catalysts has been studied for several other applications. With Y-zeolites the enhanced sulfur tolerance was recognized early [16-19]. The effect has been attributed to an electronic influence on the metal particles in the zeolite. The poisoning by sulfur is due to the strong metal-sulfur bond, and it is speculated that electronic (positive) charge transfer to the metal reduces the metallic character of the particles and hence weakens the metal sulfur bond [22]. Similar theories have been used to account for differences in sulfur tolerance when supporting noble metals on conventional supports. Both the support properties and the particle size have been reported to influence the sulfur tolerance strongly. For the hydrogenation of ethylbenzene over supported Pd it has been shown that the sulfur tolerance
393 correlates well with the electron deficiency of the metal, measured as the shift in binding energy in the XPS spectrum [23,24]. Hoyos et al. [25] studied the dehydrogenation of cyclohexane over supported Pt and Pd. The poisoning effect of thiophene and H2S (up to 3 ppm H2S) was found to depend on the acidity of the support and was found to be lower when the metal was supported on Si02-Al203 compared to AI2O3 as the support. We have earlier shown that Y-zeolite supported metals (Pt, Pd, Ir and Ru) have improved suliur tolerance in cyclohexane dehydrogenation [26]. In this paper we report some results from our work using the hydrogenation of toluene as the model for the aromatic saturation of diesel fuel [27,28]. The kinetics and mechanisms of hydrotreating reactions over traditional sulfide based catalysts have been widely studied (for a recent review see e.g. [29]). However, the properties of this new class of catalysts when used in the presence of sulfur are not well described. Toluene was chosen as the model aromatic compound, since the last aromatic ring is kinetically the most difficult to hydrogenate. In the present experiments benzothiophene (BT) was used as the sulphur source. In earlier communications we have reported that under these conditions BT gives the same poisoning effect as does dimethyldisulfide (DMDS), which is easily decomposed to H2S. This indicates that BT is desulfurized to give H2S in the experiments reported here.
2. EXPERIMENTAL 2.1. Catalysts Supported Pt catalysts were prepared on different Y-zeolite supports from PQ zeolites. The supports had bulk Si/Al ratios rangingfrom2.6 to 40 and were used as received in the H- or NH4 -forms. The supports were inpregnated with 0.3 wt% Pt by ion-exchange, using Pt(NH3)4Q2. On the most acidic sample (CBV-500) a competitive ion-exchange technique was used, by adding NH4NO3 (NH4^:Pt = 200). This was done to ensure a good distribution of Pt in the sample. On the other zeolites no competing ion was used. The ion-exchange was performed at room temperature with stirring for 24 hours followed by settling for 48 hours and filtering, washing and drying at 80 - 100 °C overnight. The degree of exchange was checked by analysis of the exchanged solution. After drying the samples were calcined in flowing dry air (200 cc/g cat., min), using a slow temperature increase to avoid autoreduction by the evolved ammonia [16]. The heating rate was 1 °C/min up to 180 °C, holding at this temperature for 3 hours, followed by heating to 400 °C at 1 °C/min and holding at this temperature for 6 hours. The catalysts were reduced in situ using flowing H2 at 1 bar. The reduction procedure consisted of a drying step (flowing N2, heating at 0.8 °C/min to 240 °C, hold 4 hours) followed by the reduction. The gas was switched to H2 (30 cc/min) at 240 °C, the heating rate was 10 °C/min, up to 500 °C, and hold for 2 hours. A 0.3%Pt/Y-Al2O3 sample was used as a reference. This sample was given the same pretreatment as the zeolite based samples. The acidity of the supports and one catalyst was measured by ammonia TPD using a traditional microbalance [30]. The amount of NH3 desorbing at temperatures higher than 250 °C was taken as a measure of the acidity of the samples. The Pt metal dispersion was characterized by H2 adsorption using the pulse technique, as described elsewhere [31].
394 2.2. Activity measurements The catalyst activity for toluene hydrogenation in the presence of sulfur components of varying concentration was studied in a fixed bed tubular reactor. Figure 1 shows the experimental set-up for the activity measurements. The unit consists of a reactor (9 mm i.d.), two liquid feed lines and several gas feed lines. The liquid feed was dosed with 2 separate HPLC pumps, while the gases were metered using Flow-Tek mass flow controllers. The catalyst (approx. 0.15 g) was diluted with SiC up to 1.3 g total weight, to give a bed height of approximately 20 mm in the reactor. The liquid feed was injected into the gas feed at the reactor inlet. Separate reservoirs, pumps, and feedlines for clean feed and sulfur-spiked feed was used to avoid history effects due to strongly adsorbed sulfur components desorbing into the clean feed. The exit gas was analyzed by on-line GC using a flame ionization detector and a CP-5 (Chronpack) column. The hydrocarbon products collected in the condensation pot could be sanpled and analyzed off-line for detailed composition.
Vent
i
-CD-
-•H E D C M
±'
]M
s&e
••-^
Vent
m^m
K2 H
Q •
0
: ball valve : pressure reduction valve
^ Condenser
. Condenser : Cold lines : Heated lines
: manometer
Figure 1. Scheme of the experimental set-up for the activity measurements. Al, A2: H2 and N2 cylinders, ultrahigh purity, B: Gas purifier (oxisorb), C: Flow controller (Bronkhorst Hi-Tech), D: 3-way valve, E: HPLC pumps (range 1 |il/min to 10 ml/min), F: liquid reservoir, G: tubular reactor placed in vertical furnace, H: On-line GC, I: Back pressure regulator.
395 The toluene used was of p.A. quality (minimum purity 99.5 %), containing < 0.5 ppm S (by analysis). Benzothiophene (BT, Aldrich, >95%) was used as the sulfur compound. The sulfur concentration in the liquid feed was controlled by analysis (ANTEK 7000 Elemental Analyzer). The collected liquid products were also analyzed and found to be virtually sulfur free ( < 0.5 ppm S), confirming that the BT molecule is desulfurized completely under these conditions. Before the activity measurements the catalyst was reduced in situ as described above. After the conviction of the reduction the sample was cooled to the desired reaction temperature in H2, then flushed with N2 and pressurized to the reaction pressure. The gas composition was adjusted to the desired flow before the liquid pumping was started.
3. RESULTS AND DISCUSSION 3.1 Catalyst characterization Table 1 shows the important characteristics of the catalysts. The Y-zeolite supports were characterized by NH3-TPD. As expected the number of strong acidic sites, defined here as the amount of NH3 desorbed above 250 °C, decreases with increasing Si/Al ratio. The TPD profiles also showed a broad peak centered around 180 °C, attributed to physically adsorbed NH3 after comparing with the desorption profile of a non-acidic Y-zeolite. The desorption peaks are wide, indicating a large distribution of acid strength. The amount desorbed is low, and only constitutes about 30-40% of the theoretical number of acid sites based on calculations of framework Al per Table 1 Characterization data for the catalysts Support Catalyst
Metal function
Si/Al (bulk)
acidity mmoles/g^
H :Pt^
Part, size^ (nm)
3PtC500
2.6
1.28
0.28
3.9
3PtC712
6.0
0.56
0.06
18
3PtC720
15
0.27
0.29
3.7
3PtC740
20
0.23
0.19
5.7
3PtC780
40
0.19
0.35
3.1
0.51 2.1 3Pt/Al203 1. NH3-TPD, amount of NH3 desorbed from the pure supports in the range 250 - 550 °C. 2.Measured by H2-O2 titration using the pulse technique at 25 °C after reduction (10 °C/min up to 500 °C, hold 120 min.) followed by degassing in flowing Ar at 500 °C, 120 minutes. 3. Assuming spherical particles.
396 unit cell [30]. This is in line with other reports [32-35]. The catalysts were prepared with the goal of making well-dispersed, active hydrogenation catalysts. The conditions chosen were selected to give a high Pt dispersion, both during the metal introduction step (ion exchange) and during the calcination step [16,36]. The metal function of these catalysts was characterized by H2 chemisorption. The results of the chemisorption and the calculated Pt particle sizes are also given in Table 1. The chemisorption results show a large variation in the measured chemisorption of hydrogen, with 3 samples falling in the range 3-4 nm particle diameter, while the other samples have large particles. The larger values would indicate that the metal might be located in mesopores formed during the steaming of the supports, since the particle sizes are too large to be accomodated in supercage cavities. As mentioned the preparation was directed towards obtaining well-dispersed Pt, both through the ion-exchange and the subsequent treatments. However, the competing ion was only added during the ion-exchange of the CBV-SOO support, while for the other supports no conpeting ion was used since they were considered to have lower acidities. This could have led to a less even distribution of Pt, particularly in the case of the rather strongly acidic CBV 712 support. But it is also possible that the dispersion is underestimated somewhat by the chemisorption, since the chemisorption of H2 has been noted to be an unreliable tool for dispersion measurements in these catalysts [36]. This effect is however reported to be important for systems with very high dispersion, whereas larger Pt particles normally are easier to characterize. The reference catalyst demonstreted the dispersion level expected for such catalysts.
3.2. Temperature effect A series of experiments were conducted where the reaction temperature was varied over a wide range (160 - 400 °C). Figure 2 shows a plot of the toluene conversion as a function of the reaction temperature. The catalyst used was Pt/CBV-712. Upon increasing the reaction temperature there is detectable conversion of toluene from about 160 °C. A further increase in temperature leads to increasing toluene conversion up to about 300 - 325 °C, where there is a maximum in the conversion. The conversion reaches about 50% at these conditions. Upon a further increase in the temperature to 400 °C the conversion drops to about 30%, which is very close to the equilibrium at these conditions. Similar curves were obtained for the other catalysts investigated, although the alumina-supported sample gave rise to a maximum in rate at a somewhat lower temperature (275 °C). A simiUar trend can be seen when using a toluene feed containing 20 ppm S (100 ppm BT). The activity is now much lower, and a reaction temperature of 240 °C is necessary to obtain a detectable conversion. Again a maximum in the conversion is observed around 300 - 325 °C. The maximum conversion is about 35 %. At 400 °C the conversion is again close to the equilibrium. This phenomenon has been observed earlier over supported Ni-catalysts. Van Meerten and Coenen [37] observed this behaviour in the hydrogenation of benzene over Ni/Si02, and excluded effects such as poisoning, diffusion limitations, and approach to equilibrium as the cause of this effect. In their studies of toluene and ethylbenzene hydrogenation. Salmi et al. [38,39] attributed such behaviour to a reduction in the hydrogen coverage with increasing temperature. The
397 100 Equilibrium
80 H
O
60 i
>
40
o
Clean feed
o
20 H
100
200
300
400
500
Temperature °C Figure 2. Temperature response in toluene hydrogenation over PtC720. Conditions: P,„, = 31 bar, P,„, =1,6 bar, PH2=PN2=14,7 bar, WHSV = 40hr'. maximum in rate is then a consequence of the combined effects of the temperature dependence of the rate constant k, and the heat of adsorption, leading to a reduction in the coverage of adsorbed reactants with increasing temperature. The effect observed here is very similar. The position in the rate maximum is a function of the reaction pressure, higher pressures lead to a shift in the reaction rate maximum to higher temperatures, as was also observed earlier [37]. 3.3. Pressure effects To further explore the kinetic behaviour of these catalysts a simplefied kinetic study was conducted, using power rate law kinetics (eq. 1): ,
m
n
r = k-pH2PT
(1)
Figure 3 shows plots of the data obtained over the Pt/CBV712 catalyst, keeping either the partial pressure of toluene or that of H2 constant and varying the pressure of the other component, always maintaining differential conditions in the reactor (conversion below 10%). The observed kinetic parameters are summarized in Table 2 below, also including data for some other catalysts. The
398 -1U -
-9
a
0-11-
c
"
o -10-1 L.
n=0
-12-13-
-0,5
"1
1
0,0
0,5
•
m = 1.5
-11 H -12
1,0
2
I
I
2,5
3,0
3,5
'"PH.
-11
-10 -11 H
-12 H
n = -0.7
-13 H
-12 H -13 0,5
1
1
0,0
0,5
m = 1.5
-14 2
1 -14
e
o -14 ^
2,5
3,0
f
m = 0.7
-15-16-
-0,5
3.5
n = -0.7
•i
•
1
InPh,
' " Ptol -13 1
1
-]
0,0
-17
1
0,5
1
—1
2,5
1
3,0
3.5
""PH, Figure 3. Kinetic plots of toluene hydrogenation over PtC712. Conditions: Pj^t = 31 bar, T = 240 °C. a, b : Clean toluene, c,d: 20 ppm S, e,f: 200 ppm S. a,c,e: order in toluene with constant partial pressure of hydrogen = 14,7 bar, b,d,f: order in hydrogen with constant partial pressure of toluene =1.7 bar.
399 Table 2 Observed kinetic parameters in toluene hydrogenation over supported Pt catalysts Clean feed
Clean feed
20 ppm S
200 ppm S
n
m
n
m
n
m
-0,7
1,5
-0,7
0,7
-0,7
1,5
Catalyst
E (kJ/mol)
TMAX(°C)
Pt/Al203
45
275
PtCSOO
35
325
0
1,5
PtC712
63
325
0
1,5
PtC780
75
activation energies were estimatedfromthe temperature response experiments with clean toluene feed, taking the data in the region 160 - 240 °C. The activation energies are in the region from 35 - 75 kJ/mol. With clean toluene feed the observed reaction orders are 0 in toluene and 1.5 in H2. These values can be compared to hydrogenation kinetics reported in the literature. Lin and Vannice [40] studied toluene hydrogenation at total pressures close to atmospheric and reported activation energies of 12 ± 2 kcal/mol (50 ± 8 kJ/mol) for toluene hydrogenation over Pt supported on a range of supports. They found a zero order dependency in toluene, and a pressure dependency in H2 that varied with the temperature. At low temperatures (60 °C) they reported an order m < 0.7, whereas at 100 °C m was found to be > 1. A general reaction model involving the addition of hydrogen atoms to an adsorbed hydrogen deficient species as the rate determining step was developed [41]. The results reported here, although obtained at much higher pressures, are comparable, showing very similar trends, and would support such a model. The range of acidity used here does not seem to influence the kinetic behaviour of these catalysts. With 20 ppm sulfur present in the liquid feed (added as 100 ppm BT) a slightly negative order in toluene is observed (n = -0.7),whereas the order in H2 is unchanged. Increasing the sulfur concentration to 200 ppm S leads to noftirtherchange in the observed order in toluene (n = -0.7), but reduces the observed order in H2 to m = 0.7. Although a comprehensive study has not been completed, the summary in Table 2 indicates that these trends are independent of the support within the range used here. The reference catalyst looses too much of its activity with 20 ppm S present to allow kinetic investigations [27]. The changes in reaction order with the sulfur addition to the feed would indicate a change in the mechanism or the rate determining step involved. The addition of sulfur to the feed has the effect of partly or completely sulfiding the surface, not only covering the active Pt sites but leading to different adsorption sites with different adsorption properties. From a thermodynamic point of view a complete sulfidation of the Pt to bulk PtS or PtS2 is not expected at these conditions. A surface sulfide is formed, first on the high coordination sites followed by the development of a 2dimensional surface sulfide [6]. Chiou et al. [42] used IR spectroscopy of chemisorbed CO to
400 characterize sulfur-poisoned Pt/Al203. A shift in the CO stretching frequency from 2056 cm-^ to 2065 cm-^ was interpreted as a weakening in the Pt-CO bond strength, and a similar weakening in the Pt-H bond strength was inferred. Similar conclusions were drawn from model studies on Pt (110) [43] and Pt(lOO) [44] single crystals. The observed negative orders in toluene are surprising, since zero order behaviour, corresponding to a complete surface coverage of the aromatic molecule, is often observed both on clean metals and on sulfides in aromatic hydrogenation [45]. Further studies are therefore needed to improve the understanding of these systems.
4. CONCLUSIONS The hydrogenation of toluene over Pt supported on Y-zeolites show a kinetic behaviour similar to that reported for supported Ni and Pt catalysts, with a maximum in reaction rate with increasing temperature. The reaction orders with respect to toluene and hydrogen change when trace amounts sulftir is added to the feed. The order with respect to toluene is more sensitive to the level of sulfur in the feed than the order in hydrogen. Acknowledgements. The work was done with support from the Norwegian Research Council and Statoil under the MILRAF programme. We thank Ame Gr0nvold for performing some of the characterization work.
REFERENCES 1. 2. 3. 4.
V.M.H. Van Wechem and M.M.G. Senden, Stud. Surf. Sci. Catal. 81 (1994) 43. F. Staat, Lipid Technology (1993) 88. F. Staat and E. Vallet, Chemistry & Industry (1994) 863. J.H. Gary and G.E. Handwerk, Petroleum Refining: Technology and Economics, Marcel Dekker Inc., New York, 1994. 5. J.P. Lucien, J.P. Van den Berg, G. Germaine, H.M.J.H. Van Hooijdonk, M. Gjers, and G.L.B. Thielemans, in M.C. Oballa and S.S. Shih (Ed.), Catalytic hydroprocessing of petroleum and distillates: proceedings of th AIChE Spring National Meeting, Houston, Texas, March 28 - April 1, 1993, Marcel Dekker, Inc., New York, 1993, p. 291. 6. C.H. Bartholomew, C.K. Agrawal, and J.R. Katzer, Adv. Catal. 31 (1982) 135. 7. T.A. Pecoraro and R.R. ChianeUi, J. Catal. 67 (1981) 430. 8. M.J. Ledoux and B. Djellouli, J. Catal. 115 (1989) 580. 9. J.F. Cocchetto and C.N. Satterfield, Ind. Eng. Chem. Process Des. Dev. 20 (1981) 49. 10. L. Vivier, V. Dominguez, G. Perot, and S. Kasztelan, J. Mol. Catal. 67 (1991) 267. U . S . Mignard, N. Marchal, and S. Kasztelan, Bull. Soc. Chim. Belg. 104 (1995) 259. 12. M.J. Ledoux and C. Pham Huu, Catal. Today 15 (1992) 263. 13. S.T. Oyama, Catal. Today 15 (1992) 179. 14. D.J. Sajkowski and S.T. Oyama, Preprints, ACS Div. Petr. Chem. 35 (1990) 233. 15. C.C. Yu, S. Ramanathan, F. Sherif, and S.T. Oyama, J. Phys. Chem. 98 (1994) 13038.
401 16. 17. 18. 19. 20.
21.
22. 23. 24. 25. 26. 27. 28. 29. 30. 31. 32. 33. 34. 35. 36. 37. 38. 39. 40. 41. 42. 43. 44. 45.
P. Gallezot, Catal. Rev.- Sci. Eng. 20 (1979) 121. P. Gallezot, in L.V.C. Rees (Ed.), Proc, 5th Int. Conf. on Zeolites, 1980, p. 364. T.M. Tri, J. Massardier, P. Gallezot, and B. Imelik, Stud. Surf. Sci. Catal. 5 (1980) 279. P. Gallezot, in D. Olson and A. Bisio (Ed.), Proc, 6th Int. Conf. on Zeolites, 1983, p. 352. N. Marchal, S. Kasztelan, and S. Mignard, in M.C. Oballa and S.S. Shih (Ed.), Catalytic hydroprocessing of petroleum and distillates: proceedings of the AIChE Spring National Meeting, Houston, Texas, March 28 - April 1, 1993, Marcel Dekker, Inc., New York, 1993, p. 315. B.H. Cooper, P. S0gaard-Andersen, and P. Nielsen-Hannerup, in M.C. Oballa and S.S. Shih (Ed.), Catalytic hydroprocessing of petroleum and distillates: proceedings of the AIChE Spring National Meeting, Houston, Texas, March 28 - April 1, 1993, Marcel Dekker, Inc., New York, 1993, p. 279. B. Bhatia, A. Beltramini, and D.D. Do, Catal. Rev.- Sci. Eng. 31 (1990) 431. A. Arcoya, X.L. Seoane, N.S. Figoli, and P.C. L'Argentiere, Appl. Catal. 62 (1990) 35. X.L. Seoane, P.C. L'Argentiere, N.S. Figoli, and A. Arcoya, Catal. Lett. 16 (1992) 137. L.J. Hoyos, M. Primet, and H. Prahaud, J. Chem. Soc. Faraday Trans. 88 (1992) 113. H. Bergem, A. Gitlesen, A. Holmen, and E.A. Blekkan, in G. Centi, S. Perathoner, E. Christiani, and P. Forzatti (Ed.), Environmental Catalysis. Proceedings of the 1st World Congress, Pisa (Italy), May 1-5 1995, Societa Chimica ItaUana, Rome, 1995, p. 411. H. Bergem, A. Holmen, and E.A. Blekkan, EuropaCat-II, 3-8 September 1995, Maastricht, The Netherlands, Book of Abstracts p.779. H. Bergem, A. Holmen, E.A. Blekkan, and A. Gr0nvold, 11th International Congress on Catalysis, June 30 - July 5, 1996, Baltimore, USA, Poster no. 058. B.S. Clausen, H. Tops0e, and F.E. Massoth, Hydrotreating catalysis, Springer Verlag, Berlin, 1996. A. Gr0nvold, T. Dypvik, K. Moljord, and A. Holmen, submitted, Appl. Catal. A. E.A. Blekkan, A. Holmen, and S. Vada, Acta Chem. Scand. 47 (1993) 275. A. Aurox, M. Muscas, D.J. Coster, and J.J. Fripiat, Catal. Lett. 28 (1994) 179. B. Chauvin, M. Boulet, P. Massiani, F. Fajula, F. Figueras, and T. DesCourieres, J. Catal. 126 (1990) 532. C. Hidalgo, H. Itoh, T. Hattori, M. Niwa, and Y. Murakami, J. Catal. 85 (1984) 362. M. Iwamoto, M. Tajima, and S. Kagawa, J. Chem. Soc, Chem. Commun. 28 (1986) 598. W.M.H. Sachtler and Z. Zhang, Adv. Catal. 39 (1993) 129. R.Z.C. van Meerten and J.W.E. Coenen, J. Catal. 37 (1975) 37. L.P. Lindfors, T. Salmi, and S. Smeds, Chem. Eng. Sci. 48 (1993) 3813. S. Smeds, D. Murzin, and T. Salmi, Appl. Catal. A 125 (1995) 271. S.D. Lin and M.A. Vannice, J. Catal. 143 (1993) 554. S.D. Lin and M.A. Vannice, J. Catal. 143 (1993) 563. J.F. Chiou, Y.L. Huang, T.B. Lin, and J.R. Chang, Ind. Eng. Chem. Res 34 (1995) 4277. H.P. Bonzel and R. Ku, J. Chem. Phys. 58 (1973) 4618. T.E. Fischer and S.R. Kelemen, J. Catal. 53 (1978) 24. A. Stanislaus and B.H. Cooper, Catal. Rev.- Sci. Eng. 36 (1994) 75.
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® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
403
KINEnC STUDY OF THE HYDRODENTmOGENATION OF PYRIDINE AND PIPERIDINE ON A NiMo CATALYST Ragnhild Pille and Gilbert Froment Laboratorium voor Petrochemische Techniek, Universiteit Gent Krijgslaan 281, B-9000 Gent, Belgium The gas phase hydrodenitrogenation of pyridine and piperidine was studied on a sulfided NiMoP/AljOs catalyst in a continuous stirred tank reactor of the Berty type operated at temperatures between 573 K and 633 K and pressures in the range of 15-40 bar. The molar H2S/(H2+H2S) ratio was varied between 10""^ and 0.4. Rate equations of the Hougen-Watson type were developed for the hydrogenation of pyridine into piperidine and the hydrogenolysis of piperidine into pentane. Two different types of active sites were considered: r sites for hydrogenation and a sites for hydrogenolysis. The surface reaction between the reactants and competitively adsorbed hydrogen was found to be the rate determining step. To describe the HDN of pyridine and piperidine accurately with a single set of model parameters, the Freundlich isotherm had to be introduced for the piperidine adsorption on a sites. l.INTRODUCTION Hydrodenitrogenation (HDN) of oil fractions is of paramount importance in petroleum refining. Research in this area has focused on the HDN of model components, mainly of heteroaromatic nitrogen components which are among the most difficult to denitrogenate. In this study, pyridine was chosen because of its simple reaction network and the clear functionality in each reaction step. Rate equations for the pyridine hydrogenation and piperidine hydrogenolysis were reported by Mcllvried (1971), Hanlon (1987) and Moravek (1991). In the three cases, rate coefficients and adsorption constants were estimated at one temperature only (589 K, 583 K and 573 K respectively) and no activation energies and heats of adsorption were reported. While Mcllvried (1971) and Moravek (1991) considered different types of active sites for the hydrogenation of pyridine and the hydrogenolysis of piperidine, Hanlon (1987) used a one site model for all the reactions in the pyridine network. The effect of hydrogen sulfide and hydrogen on conversions and selectivities in the HDN of pyridine has received much attention. Hydrogen sulfide was found to enhance the hydrogenolysis of piperidine (Goudriaan et al., 1973; Satterfield et al., 1975, Hanlon, 1987). According to Hanlon (1987), the hydrogenation of pyridine is not affected by hydrogen sulfide. On the other hand, Satterfield et al. (1975) observed a decrease in hydrogenation rate at increasing hydrogen sulfide partial pressure. Hydrogen was found to increase the pyridine hydrogenation rate and decrease the piperidine hydrogenolysis rate (Hanlon, 1987). 2.EXPERIMENTAL 2.1.Experimental set-up The experiments were carried out in a Berty type continuous flow stirred tank reactor
404 described previously (Pille et al., 1994). The nitrogen components, pyridine or piperidine, were diluted with hexane. The feed was pumped into a packed bed, where it was vapourized and mixed with a Hj, HjS and CH4 gas stream, before entering the reactor. The reactor temperature and pressure were controlled by a PI controller and a back pressure regulator respectively. A gas chromatograph with FID and a fused silica column with CP-Sil 5CB film was used for the on-line analysis of the effluent gas. 2.2.Procedure The catalyst used was a commercial catalyst (Procatalyse EC/22/216/91) containing 12.6 wt% Mo, 2.9 wt% Ni and 2.9 wt% P on an alumina support. To avoid internal diffusional limitations it was crushed to the size of 0.63 - 0.99 mm. 0.715 gram of catalyst was diluted with nonporous alumina. The catalyst was dried for at least 4 hr by means of argon at 423 K. Subsequently, it was presulfided in situ by a 15 mole% H2S in H2 mixture at 1 bar and 673 K for 4 hr. The total gas flow during the sulfiding procedure was 20 Nl/hr. After sulfidation, the catalyst was stabilized with the feed containing pyridine, H2 and H2S at 573 K and 30 bar until a constant activity was reached. The experiments were carried out at temperatures between 573 and 633 K and pressures in the range of 15-40 bar. The hydrogen to pyridine (or piperidine) ratio ranged from 80 to 600, the molar ratio of hexane to pyridine (or piperidine) was 40. The space time, W/F°p or W/F °pp, was varied between 100 and 500 kgc^t.hr/mole. H2S was fed as a 5 mole% H2S in Hj mixture. Through the addition of pure H2, the molar ratio RH2S=H2S/(H2S-I- Hj) was comprised between 0.0001 and 0.04. The hydrogen partial pressure was kept constant at varying HjS partial pressure through the addition of methane. The total number of experiments amounted to 178. 3.RESULTS Preliminary experiments whereby the reactor was packed with AI2O3 only, demonstrated the absence of conversion of both pyridine and piperidine in the absence of the catalyst. The absence of heat and mass transfer limitations was also experimentally verified. The stability of the catalyst was periodically checked by performing replicate experiments. No deactivation was observed during the whole period of experimentation. The reaction products of the hydrodenitrogenation of pyridine were piperidine, pentane and pentenes (1-pentene, trans-2-pentene and cis-2-pentene). In Figure 1 a typical conversion versus space time plot is shown. With increasing HjS partial pressure and at constant H2 partial pressure, the reaction rate for the hydrogenation of pyridine to piperidine is almost invariant, while the reaction rate for the hydrogenolysis of piperidine increases. This is due to the promotional effect of hydrogen sulfide on the hydrogenolysis of C-N bonds. As a consequence, the selectivity for piperidine decreases. With increasing H2 partial pressure, all conversions increase, both at constant H2S partial pressure and constant ratio RH2S=H2S/(H2S + H2). The reaction products of the hydrodenitrogenation of piperidine were n-pentylamine, pentane, pentenes and pyridine. Pyridine was not observed at 573 K. At higher temperatures the conversion of piperidine into pyridine amounted to 17 %. The maximum conversion into n-pentylamine amounted to 18 % and was observed at 573 K. At higher temperatures the amount of n-pentylamine was very small. The effect of the H2S partial pressure on the conversions is plotted in Figure 2. The piperidine conversion and the conversion into pentane, pentenes and n-pentylamine increase with increasing H2S partial pressure, while the conversion into pyridine very slightly decreases.
405 0.8
T=548K p^ = io bar
•
«tfei=..:0^^
0.6
i
a^ f 500
C
A^^''^
L^y^
m^^
y^X
"yv^
o
8
/yXp^Xpg
g "°
i
i
O O
0.2
0-4
0.2
T=613K I .Pj = 3Qbarj W/FpJ3=20^g.h/niiol >200|
L/^L^ ""pfL i 0
100
200
300
Xp
400
500
0.1
600
W/F^p (g.h/mol)
PH2S
Figure 1 - Conversions as a function of space time for the HDN of pyridine. The curves correspond to simulations based on the kinetic modeling.
0.15
0.2
0.3
(b^)
Figure 2 - Conversions as a function of the H2S partial pressure for the HDN of piperidine. The curves correspond to simulations based on the kinetic modeling.
At 573 K the effect of the H2 partial pressure at both constant H2S partial pressure and RH2S was negligible. At 598 K and a constant H2S partial pressure, the conversions to piperidine and pentenes decrease, while the conversions to n-pentylamine and pentane increase. The resulting piperidine conversion was invariant. A limited amount of experiments was conducted with both pyridine and piperidine in the feed. Pyridine was found to decrease the hydrogenolysis rate of piperidine at constant piperidine concentration. No conclusions concerning the effect of piperidine on the hydrogenation of pyridine at constant pyridine concentration could be drawn. 4.KINETIC MODELING The kinetic modeling was based on the following reaction scheme. Two different types of active sites were considered: r sites for hydrogenation and o sites for hydrogenolysis. DHP
THP
PP.T^.L
\>P.a
PN 'PP.a
'P.T
NH^
^PN,a
PA
406 A number of plausible rate equations of the Hougen-Watson type were considered in an attempt to correlate the experimental data. The surface reaction on both types of active sites was assumed to be the rate determining step. The reaction mechanisms differ by the adsorption mode of hydrogen only: competitive or not, molecularly or atomically. Reaction of the nitrogen components with Hj from the gas phase (a so called Eley-Rideal mechanism) was also considered. Consequently, 13 rival models were derived. Since the experiments were conducted in a reactor with complete mixing, yielding directly the rates, the differential method of kinetic analysis was used (Froment and Bischoff, 1990). A nonlinear least-squares regression was applied for the parameter estimation. The following reparameterization of the rate and adsorption coefficients was carried out:
A:=A:^expr--A,l = k*
X EA
1.
exp -—(—-—)
^' R T TJ
exp -^
-{—-—)
T^ is the average temperature of the experiments and equals 602 K. The significance of the parameter estimates was evaluated by means of their calculated n-values, defined as the ratio of the parameter estimate and the standard deviation. The significance of the regression was tested by means of the F value, defined as the ratio of the mean regression sum of squares to the mean residual sum of squares. Discrimination between the rival models was based on the F value, on the mean residual sum of squares R and on physicochemical criteria. The model discrimination was based on the set of data for the HDN of pyridine only. The selected model was then used and extended to describe the experiments with piperidine feed. For the ultimate parameter estimation both sets of experiments were used simultaneously. Since in all experiments the conversion to pentenes was very small, the pentenes were lumped with pentane. For the promotional effect of HjS on C-N hydrogenolysis, the approach of Massoth et al. (1990) was used as a first approximation. According to Massoths model, C-N hydrogenolysis can proceed by two paths, one unpromoted and one promoted by dissociative adsorption of H2S. Consequently, the overall rate of hydrogenolysis is multiplied by (1+(IQi2s,«Ps)^^^)> assuming that the same sites are involved in both paths. 4.1.HDN of pyridine The discrimination and estimation based on the experiments with pyridine led to the model in which molecular hydrogen adsorbs competitively with the nitrogen compounds on both types of active sites. The corresponding reaction mechanism and rate equations are given below, -for the hydrogenation of pyridine into piperidine + T
//.
+ T
P.r
+ H^.x
^ DHP,T
+ r
RDS
407
DHP,T
+ H^,r
THP.r
+ H,.T
THP.T ^ PP.T
+ T +r
^PP
PP.r
*P.T ^i».T ^H.X ( PP PH Vr
=
( 1 + ^ ^ , pp
+
-i )
(1)
KPa
K, "if.x PH * ^PF,x PFP + ^NH^x PNB, + ^HJS,. PH^ * ^DHP.x PP/^PH
+ ^IHP.T
PpplPs)^
-for the hydrogenolysis of piperidine into pentane PP
+a
^ PP.o
//.
+ o
^H^.o
PP,a
+ H^.o
^ PN.a
+a
PN.o
+ H^,o
^ PA,o
+ NHyO
^PA
+ o
PA,o
V o ^PP,0 =
( 1 + K„^
p„
+ K^„
Pa
RDS
(2)
^PP,a ^H,a PPP PE (^ ^ >j^H^,o PH^ ^
+ Kj,^ Pj, + K^^^
Pffjj^ + K,^„
+ p,^
(p,^
+ K^t^^^ (Pp^ + Pp^
PNHJPJ"
The parameters that could be estimated significantly are shown in Table 1. The parity plot (Figure 3) shows a very good agreement between the experimental and calculated conversions. .
Ixp
l-^PP
c
'^PN
g
• D P / •* «J/..
U i
*
§
c o o
•D 0 ®
j / o 0.4
\j^
£
Q.
0.2
0.4
0.6
0.8
experimental conversions
1
Figure 3 - Parity plot for the kinetic modeling of the HDN of pyridine (Eqs. (l)-(2)).
0.2
0.2
0.4
0.6
0.8
experimental conversions
1
Figure 4 - Parity plot for the kinetic modeling of the HDN of piperidine (Eqs. (3)-(4)-(5)).
408 Table 1 - Parameter estimation using (i) the experiments with pyridine feed, (ii) the experiments with piperidine feed and (iii) all experiments (l%L^,p=kp^yI^^, lq,p_»p=kppyKppy, kppL^pN=kppoI^o^
k* E
kp_pp
Pyridine feed
Piperidine feed
Eqs. (1)(2)
Eqs. (3)(4)-(5)
Pyridine and piperidine feed Eqs. (l)-(2)(3)
2.74 10-^ 2.66 10^ n.s. n.s. n.s. n.s. n.s.
7.19 6.49 2.37 1.20 1.75 8.71 n.s. 1.34 n.s.
6.10 10-^ 5.78 10^
kp_»pp
kpp_*p
kpp_»p
Kp,T KH,T '^PP.r *^H2S,T K^fH3,^
(-AH°)
%P-PN
l^PN-PA
k* E 573 K K* (-AH°)
1.50 n.s. 2.72 1.92 1.24 2.68
^H2S,a
F-value RSSQ
3055 11.5
^K,a ^NH3,
Kp,r
10^ 10-1
'^PP.r
10^
5.30 10-1 7.38 10' 8.47 2.46 9.01 n.s. n.s. n.s. 1.14
^PP,a
10^
10^ 10' 10-2
I^H,T
^H2S,T *^NH3,T
10-1
2.66 10'
^H2S,g
2.37 101 n.s. 3.97 10^ 4.24
n.s. 8.93 101 n.s. 8.53 10-1
F-value RSSQ
2296 7.4
1097 43.3
^^PN-»PA/'Ql,a
^??,J^,a
^?,J^K,a KpN,o/KH,a
k* E
10-' 10' 10-2 10' 101 10-'
3.54 9.77 10' 4.73 10-^ 4.06 10^
%P-PN/*^H,
^NH3,o/*^H,a
10-1
k* E k* E
1.56 101 1.03 10'
4.2.HDN of piperidine The selected model for the pyridine HDN was adapted to describe the HDN of piperidine. For the dehydrogenation of piperidine to pyridine, the same mechanism as for the hydrogenation of pyridine was used, only the rate determining step was changed from the P •^ DHP reaction to the PP «^ THP reaction. Further, an additional equation was introduced to describe the conversion into n-pentylamine at 573 K. For this conversion there was no promoting effect of HjS. Since 2KiPi was always much larger then 1 on the o sites, 1 was omitted in the denominator of the equations and the adsorption coefficients were estimated relative to KH ,,. ^FF.s ^PF.x ^PF,x (( Ppp PPP -Pp -Pp ^^ Pa Pa ) ) ^FF,t ^PF.x
( 1 •*• ^P,x Pp * ^H,x PE + ^PP.x Ppp * ^NHyX Pm^ * ^H^.x PE^ + ^DEP.x PPFI^H + ^IHP.t PpplPs )^
(3)
409
( - p ^ PPF ^PH^IT-PP*
^PS.a
-
^
PNH, + -p^
* Pp^
* -f^
Ppi^^
^PN,a ^PN.J^M.a PpN PH ^NH^o < - p r - Prr "^ PE * - p — /'i' + -p—
^PJLa, ^ Pm^ "^ "p—(Pw + Pp^
^niM * -p—
vi PPNT
y-^^
The parameters that could be estimated significantly are shown in Table 1. The parity plot (Figure 4) shows that there are no significant deviations between the model predictions and the experimental values of the conversions. 4.3.HDN of pyridine and piperidine Finally, the parameters in equations (l)-(3)-(4) were reestimated using both the experiments with pyridine and piperidine feed. The parameters that could be estimated significantly are shown in Table 1. From the F value and the mean residual sum of squares, it follows that the use of both sets of experiments in the estimation of the parameters results in a much poorer fit. While the experiments with pyridine feed are reproduced quite well, the model predictions for Xpp and Xp^+XpE+XpN for the set of experiments with piperidine feed are poor. A similar discrepancy in experiments with piperidine and pyridine feed was encountered by Mcllvried (1971) and Hanlon (1987). According to Hanlon, the value of lq,p^ calculated from a pure piperidine feed is about 34 % lower than that calculated from a pure pyridine feed. Mcllvried reported a 25 % difference between both rate coefficients. Both authors fixed the adsorption coefficients on the values obtained with a pure piperidine feed. The difference in the rate coefficients was attributed to (i) a difference in catalyst activity, (ii) a strong dependency of kpp^ on the fixed value of Kpp^ and (iii) a variation of Kpp^ with the reaction conditions. Since in the present study, the same catalyst batch was used for both the experiments with pyridine and piperidine feed, explanation (i) can be rejected. Explanation (ii) is not valid neither since the adsorption coefficients on the o sites were not fixed on the values obtained from the experiments with piperidine feed. A possible variation of the adsorption coefficients for pyridine and piperidine with the reaction conditions is examined in the next paragraph. Also changes in the reaction mechanism or rate determining step were considered as possible causes for the observed difference between experiments with pyridine and piperidine feed. The existence of a variation of the adsorption coefficients with the reaction conditions was examined by dividing the experiments into three categories, corresponding to different piperidine or pyridine partial pressures. The parameters were reestimated using different adsorption coefficients for pyridine and piperidine on both the o and T sites for the three categories. From Table 2, it follows that Kpp,,decreases with increasing piperidine partial pressure. No correlation was found between Kp ^ or Kp^ and the pyridine partial pressure. The variation of Kpp^ with piperidine partial pressure could not be checked since Kpp^ could not be estimated significantly. The dependence of Kpp^ on the piperidine partial pressure can be attributed to two factors: first, the existence of active sites with different binding energies and, second, an interaction between adsorbed piperidine molecules (Boudart and Djega-Mariadassou, 1984).
410 Table 2 - Variation of Kpp^ with the piperidine partial pressure. PPP (bar)
Kpp,, (bar^)
0.000 - 0.006 0.006-0.016 0.016-0.101
5.08 10^ 4.17 10' 2.47 10'
Since in the present case, the piperidine partial pressure was very low and always smaller than the pyridine partial pressure, an interaction between adsorbed piperidine molecules is very unlikely, especially in the absence of an interaction between pyridine molecules. The existence of active sites with a distribution of strengths of adsorption is more likely and was already mentioned by Massoth and Miciukiewicz (1986). It dominates at low coverages and is therefore only detectable through the piperidine adsorption coefficient. In the presence of a distribution of adsorption strenghts, one of the assumptions of the classical Langmuir theory, namely the independence of the heat of adsorption on the surface coverage, is no longer valid. Therefore two other adsorption isotherms (Froment and Bischoff, 1990) were tried to describe the piperidine adsorption on o sites. The Freundlich isotherm is based on a logarithmic dependence of the heat of adsorption on the surface coverage.
C
c.. It leads to the following rate equation for the piperidine hydrogenolysis on the o sites. ;;
^pp,o
-j^
-
(6)
-
PH * -jp- Pp + -^— Pm^ * ~K~~^'^ ^'^
~T~
^^
The Temkin isotherm assumes a lineair dependence of (-AH) on the surface coverage.
i-hH,,) = i-AH,,)^
(1 - a • ^ ) tr,t
C^P.. ^ -
(
/
RT
Arr X
)^<^PP,aPpp)
a,t
It leads to the following rate equation. (7) PM^Y^PP"^ H,a
-J^
H,9
Pm, * -J^^PA * Pp^ ^ ^ H,o
PPN
H,o
411 Both the Freundlich (F-valiie=2886, RSSQ=16) and the Temkin isotherm (F-value=2412, RSSQ = 17) for the adsorption of piperidine on the o sites give much better results than the Langmuir isotherm and the Freundlich isotherm is superior to the Temkin isotherm. The resulting parity plot, given in Figure 5, shows a good fit. The agreement between the model predictions and the experimentally measured conversions after the introduction of the Freundlich isotherm is also shown in the Figures 1-2. The parameter values are given in Table 3. In an attempt to improve the fit while 0.2 0.4 0.6 0.8 1 accepting the Langmuir isotherm, the reaction experimental conversions mechanism was modified by assuming that part of the piperidine molecules do not desorb Figure 5 - Parity plot for the kinetic modeling before reaction into n-pentylamine. The of the HDN of pyridine and piperidine after piperidine hydrogenolysis then occurs on both the introduction of the Freundlich isotherm r and o sites. The resulting F value of 1240 for the PP adsorption on o sites (Eqs. (l)-(3)and mean residual sum of squares of 42 indicate (6)). an improvement of the fit. The fit remains inferior to the fit obtained by the introduction of the Freundlich isotherm. A switch of the rate determining step from the surface reaction to the adsorption of piperidine did not improve the fit (F-value = 163, RSSQ=303). Table 3 - Parameter estimates, lower and upper limits of the approximate 95% confidence intervals and n-values for the selected model (Eqs. (l)-(3)-(6)) after the introduction of the Freundlich isotherm for the piperidine adsorption on o sites (k in kmole/(kg^t.hr), K in bar'^).
kp,T
^PP,T
k* E k* E
Kp,r ^H.r
^ms,T %P,
^VY,a
y ^?,J^U,a Ks.a
k* E
Parameter estimate
Lower limit
Upper limit
n-value
7.85 10-^ 6.98 10* 7.62 10-^ 1.86 10^ 2.01 10^ 9.37 10-' 8.15 10-^ 2.21 10-^ 9.29 10* 3.46 10-^ 3.39 10-^ 6.93 10^ 5.90 10-^
6.76 10-' 6.74 10* 3.96 10-' 1.29 10^ 1.79 10^ 6.62 10-' 1.03 10-^ 1.95 10-^ 8.86 10* 1.77 10-^ 3.18 10-1 4.88 10^ 3.18 10-1
8.94 10' 7.22 10* 1.13 10-^ 2.12 10^ 2.18 10^ 1.21 10-^ 1.53 10-1 2.47 10-^ 9.72 10* 5.15 10-1 3.59 10* 9,02 101 8.62 10-1
14.4 57.4 4.2 6.6 18.4 6.8 2.3 17.0 42.9 4.1 33.1 6.6 4.4
412 5.CONCLUSION A rigorous kinetic study led to the selection of a kinetic model based on the existence of two kinds of active sites and a molecular and competitive Hj adsorption. To describe the HDN of pyridine and piperidine accurately with a single set of model parameters, the Freundlich isotherm had to be introduced for the piperidine adsorption on a sites to account for the existence of active sites with different binding energies. ACKNOWLEDGEMENTS R. Pille is grateful to the Belgian Nationaal Fonds voor Wetenschappelijk Onderzoek for a Research Assistantship and a NFWO-Petrofina Grant. NOMENCLATURE concentration of vacant sites s concentration of total number of sites s Cs,t E activation energy FC molar feed flow rate of component i (-AH) heat of adsorption rate coefficient of component i on s sites Ks adsorption coefficient of component i on s sites Ki.s partial pressure of component i Pi r rate molar ratio H2S/(H2S+H2) RH2S RSSQ mean residual sum of squares W catalyst weight W/Fi° space time X conversion Subscripts DHP dihydropyridine P pyridine PA pentaan PE pentenen PN n-pentylamine PP piperidine THP tetrahydropyridine Greek symbols a,Y parameters in Freundlich and Temkin isotherm X active site for hydrogenation Oyd* active sites for hydrogenolysis Cs
kmole/kg^t kmole/kg«,t kJ/kmole kmole/hr kJ/kmole kmole/(kg^t.hr) bar^ bar kmole/(kgcat.hr)
kgcat
kgcafhr/kmole
REFERENCES M. Boudart, G. Djega-Mariadassou, Kinetics of Heterogeneous Catalytic Reactions, Princeton University Press, 1984. G.F. Froment, K.B. Bischoff, Chemical Reactor Analysis and Design, John Wiley, New York, 1990. - F. Goudriaan, H. Gierman, J.C. Vlugter, J. Inst. Petrol., 59 (1973) 40. - R.T. Hanlon, Energy & Fuels, 1, 5 (1987) 424.
413 F.E. Massoth, J. Miciukiewicz, J. Catal., 101 (1986) 505. F.E. Massoth, K. Balushami, J. Shabtai, J. Catal., 122 (1990) 256. H.G. Mcllvried, Ind. Eng. Chem. Process Des. Develop., 10, 1 (1971) 125. V. Moravek, J.-C. Duchet, D. Cornet, Appl. Catal., 66 (1990) 257. R.C. Pille, C.-Y. Yu, G.F. Froment, J. Mol. Catal., 94 (1994) 369. C.N. Satterfield, M. Modell, J.F. Mayer, AlChe J., 21, 6 (1975) 1100.
® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
415
Kinetic Modelling of H D N Reaction over (Ni)Mo(P)/Al203 Catalysts M. Jian and R. Prins Laboratory for Technical Chemistry, Swiss Federal Institute of Technology, 8092 Ziirich, Switzerland
ABSTRACT A kinetic modelling method with a two-site Langmuir-Hinshelwood rate expression for the hydrodenitrogenation of orr/io-propylaniline at 623 K and 3.0 MPa over sulfided (Ni)Mo(P)/Al203 catalysts is discussed. Kinetic parameters were determined by varying the space time at constant initial reactant pressure as well as by varying the initial reactant partial pressure at constant space time. An elimination mechanism was found mainly responsible for the C-N bond cleavage reaction. Nickel strongly enhanced the hydrogenation activity of the catalyst, phosphorus increased only the adsorption constants over Mo(P)/Al203 catalysts and exhibited no influence on the HDN rate constant, but increased both HDN rate and adsorption constants over NiMo(P)/Al203 catalysts. INTRODUCTION Kinetic modelling of a hydrodenitrogenation (HDN) network is difficult, and kinetic data are often treated in an oversimplified manner. Reaction steps have been assumed to take place on one and the same catalytic site, and/or a pseudo-zero or first order reaction has been assumed to calculate the model parameters (1). Although it is well known that hydrogenation and C-N bond cleavage reactions in the HDN network take place over different catalytic sites (2-5), even in the kinetic modelling of the HDN networks of pyridine and quinoline, one-site Langmuir-Hinshelwood rate expressions for all reaction steps of the network were used (6, 7). Although a two-site model was used for the quinoline HDN reaction by Shih et al. (8), they assumed that all nitrogen-containing species, including ammonia, had the same adsorption parameters. Our recent work has shown that kinetic parameters in the network of quinoline can be obtained separately (9), and once a substantial number of such parameters is known, an effective modelling of the entire quinoline HDN network is possible. Orr/w?-propylaniline (OPA) is a major intermediate in the HDN network of quinoline, and one of the representative nitrogen-containing hydrocarbons in petroleum distillates. A recent study showed that the transformation of anilines to hydrocarbons is of primary importance in a practical HDN process, as the accumulation of anilines as intermediates is substantial over a commercial NiMo/AUOs catalyst (10).
416 In the present work, OPA was employed as the target reactant to model the HDN network. The reaction was carried out over Mo(P)/Al203 and NiMo(P)/Al203 catalysts at 623 K and 3.0 MPa. Different ways of determining the kinetic parameters of the HDN network were explored, and the modelling results were used to interpret the nature of the HDN catalytic sites, reaction mechanisms and the function of the catalyst components. EXPERIMENTAL Catalysts used in this work had a composition of 0 or 3 wt% nickel, 8 wt% molybdenum and 0 or 2 wt% phosphorus. Details of the catalyst preparation can be found elsewhere (11). The HDN reactions were carried out in a continuous-flow microreactor. A sample of 0.1 g catalyst diluted with 9.5 g SiC was suMded in situ with a mixture of 10% (mol) H2S and H^ at 643 K and 1.5 MPa for 4 h. Dimethyldisulphide (DMDS) was added to the liquid reactant to generate H2S in the reaction stream (PHIS = 6.5 kPa). Reaction products were analysed by on-line gas chromatography with a 30 m DB-5 fused silica capillary column and a flame ionisation detector. Samples were taken after 100 h on stream at 643 K when the activity of the catalyst was relatively stable, with n-nonane and n-dodecane as internal standards. The space time (r) was changed by varying the liquid and gaseous reactant flow rates while keeping their ratios constant. RESULTS AND DISCUSSION The HDN reaction network of OPA is shown in Fig. 1. In the presence of H2S, the direct C(sp^)-N bond cleavage of OPA to PB occurs only to a minor PQHE extent (12, 13), and the HDN reaction takes place mainly through hydrogenation of the aromatic ring to propylcyclohexylamine (PCHA), followed either by PCH an elimination of NH3 to propylcyclohexene (PCHE) with subsequent Fig. 1 HDN network of ortho-propylaniline reaction to propylbenzene (PB) and propylcyclohexane (PCH), or by direct hydrogenolysis to PCH. In the present study, the PCHA concentration in the product was very low, and therefore the intermediate PCHA was neglected in the kinetic analysis. Furthermore, it was demonstrated that hydrogenation of PB hardly occurs in the presence of OPA, and therefore this reaction step was omitted from the network analysis as well (Fig. 1).
417 /. Modelling of the HDN network At least three different kinds of reactions may be involved in the HDN process of Fig. 1: direct C(sp^)-N bond cleavage of OPA to PB, hydrogenation of the phenyl ring of OPA to PCHA followed by C(sp^)-N bond cleavage to PCHE or PCH, and (de)hydrogenation of PCHE to PCH (PB). It has been suggested that these reactions may occur on different kinds of catalytic sites(13-16), in that case different adsorption constants of OPA and NH3 have to be assumed. However, since the direct C(sp^)-N bond cleavage in the HDN of OPA occurred to a small extent only, it was assumed that OPA has the same adsorption constant on the catalytic sites used in the first reaction steps of the HDN network. Assuming that the HDN of OPA and hydrogenation of PCHE take place on two different catalytic sites, and that the adsorption of the nitrogen-free hydrocarbons in the reaction stream can be neglected, we have (see Fig. 1) __
^'OPA
(^1 "*" ^ 2 "^ ^ 6 ^'^OPA'OPA 1 + KQP^PQP^
at
""'PB
'^\'^OPA'^OPA
'^A^ PCHE ^PCHE
1 + KQP^PQP^^ + Kf^ff2'NH3
ax
d'pCHE
1 "*" ^ 0/>A 'OPA "*" ^ NH3 'NH3
1^2'^OPA'OPA
K'^^'^'^S)'^
1 + l\ OPA'OPA "^ '^NHl'NH3
UT
^'PCH
'^6'^OPA'OPA 1 + KQPA PQP^ + ^NHYNHl
dX
^ 5 ^ PCHE 'PCHE
'OPA
(2) (3) (4)
1 "^ ^ OPA 'OPA "*" ^ NHl 'NH3
(5)
= P° — P
'NHi
PCHE PpCHE
1 + / ^ QPA 'OPA "*" ^ AW3 'NHZ
. +
—
P
(1)
+ Kf^if^Pj^u^
'OPA
Direct simulation of the experimental data with the above rate expressions proved impossible. With a variety of combinations of the model parameters which may contain 10 fold variation of the individual parameters, the model fitting was equally good and the confidence ranges of the parameter values were very wide. Therefore, other possibilities to calculate the kinetic parameters had to be explored. 2. Determination of the kinetic parameters If the HDN reaction is performed at low OPA conversion, equation 1 can be written as i«n
^
X - (^1 '^^I'^^C^^oPA
- ln(l - XoPA) -
—r.
-To
_
T
.^x
(6)
^"^ '^ OPA'OPA
By varying the initial partial pressure of OPA (P'^QPAX a series of XQPA values is obtained at a fixed space time. Thus, the sum of the rate constants fc+fo+fe and the adsorption
418 constant KOPA can be determined through regression of equation 6 (Fig. 2). The individual values of ki and fe+fe can be obtained in combination with the initial slope of the 1-XOPA versus XPB (conversion to PB) plot which gives the ratio of (k2+k6)/kj (13, 14).
Fig. 3 Regression
of
the
OPA
network
parameters by variation of the initial reactant partial over NiMo/Al203 at 623 K and 3.0 MPa through equation 6.
initial partial pressure. kPa
The inhibition adsorption effects of OPA and NH3 on the olefin hydrogenation reaction step in the HDN of OPA (PCHE-^PCH) were studied by using cyclohexene (CHE) as the model reactant, and pentylamine as the source of NH3. With the aid of equation 7, the adsorption constant of OPA and NH3 can be determined.
(7)
-ln(l-jcc;,,) = ^^^^^^^^^ V
CHE J
1 + /^. P
Once the adsorption constants of OPA on the HDN and hydrogenation sites, as well as the adsorption constant of NH3 on the hydrogenation site were known, the remaining parameters in the complete HDN network could be uniquely fitted. The 95% confidence ranges of the fitted parameters were usually within ±20% variance of the parameter values, demonstrating the effectiveness of the fitting to the experimental data. The kinetic parameters obtained by direct determination, as well as by fitting are given in Table 1. Table 1. Kinetic parameters of the HDN network of OPA at 623 K and 3.0 MPa over different catalysts (k in [kPas'^] and K in [kPa"^]) parameters
M0/AI203
M0P/AI2O3
NiMo/AbOa
NiMoP/AbOa
k,
0.2
0.5
k2
0.9
0.3 0.8 1.2
0.8 6.4
6.2 0.2
k^KpCHE
0.9 7.6 0.2
5.8 1.1
0.5 0.4
K'oPA
0.3 0.4 0.2
0.3
38 0.5 0.8 0.5 0.4
K'NH3
0.4
0.6
0.5
ksKpCHE
ks KOPA KNH3
0.9 49 0.6 1.2 1.0 0.9 1.0
419 In the case of M0/AI2O3 and M0P/AI2O3, a direct reaction route from OPA to PCH (ko) had to be added to the rate equations to obtain a good regression of the experimental data (Fig. 3). In the case of the NiMo/Al203 and NiMoP/Al203 catalysts, addition of such a reaction route did not substantially influence the quality of the regression.
Fig. 5 Regression of the kinetic parameters in the HDN of OPA over M0P/AI2O3 catalyst at 623 K and 3.0 MPa
3. Application of the modelling results Table 1 shows that the presence of nickel in the catalysts increases the rate constant of fo+fcs (which is limited by a hydrogenation reaction) as well as that of kj, the C(sp^)-N bond cleavage. However, the rate constant of fe+fe is increased by a factor of 6-7, while the rate constant of kj is only increased by a factor of 2-3. A similar conclusion can also be drawn from the hydrogenation of PCHE. The results suggest that nickel strongly enhanced the hydrogenation activity of the catalyst. In the absence of nickel, the rate constants of fe+fo are the same over M0/AI2O3 and M0P/AI2O3, suggesting that the number and/or the intrinsic activity of the hydrogenation site are not influenced by the presence of phosphorus. However, the adsorption constant of OPA was increased substantially by the presence of phosphorus. Therefore, the promotional effect of phosphorus over M0/AI2O3 catalysts in the HDN of OPA is only due to the favourable adsorption of the reactant over the phosphorus-containing catalysts. A similar observation has also been reported in the HDN of piperidine and pyridine where phosphorus neither influences the conversion of piperidine, nor the C-N bond cleavage of piperidine and pyridine (11). It only exhibits a minor influence on the hydrogenation conversion of pyridine which may also be due to a favourable adsorption of the reactant over the phosphorus-containing catalyst. The results in Table 1 confirm that the majority of the C-N bond cleavage in the HDN of OPA takes place through C(sp^)-N bond cleavage of PCHA, There can be two ways of C(sp^)-N bond cleavage: elimination and hydrogenolysis. By studying a series of amines, Portefaix et al. demonstrated that hydrogen atoms in p-C position are involved in the
420 formation of hydrocarbons and that an E2 Hoffmann type elimination mechanism is mainly responsible for the C-N bond cleavage reaction (17). However, the role of catalyst has not been clearly explained in the C-N bond cleavage mechanism yet (18-20). In the HDN network of OPA, the C-N bond cleavage of PCHA directly to PCH can be regarded as hydrogenolysis, while the C-N bond cleavage of PCHA to the formation of PCHE, which is further hydrogenated to PCH, can be regarded as elimination. The present results show that whether C-N bond cleavage takes place via an elimination or a hydrogenolysis mechanism also depends on the type of catalyst used. Over the M0/AI2O3 and M0P/AI2O3 catalysts 20-30% of the C-N bond cleavage reaction takes place through hydrogenolysis, while in the case of NiMo/AlaOs and NiMoP/AlaOa catalysts this portion is less than 10%.
REFERENCES 1. M. J. Girgis and B. C. Gates, Ind, Eng. Chem. Res. 30, 2021 (1991). 2. C. N. Satterfield and S. Gultekin, Ind. Eng. Chem. Proc. Des. Dev. 20, 62 (1981). 3. R. Prins, in Knozinger, Ertl and Weitkamp (eds), Encyclopedia of Catalysis, to be published. 4. Y. Okamoto, A. Maezawa and A. Imanaka, J. Catal. 120, 29 (1989). 5. E. Laurent and B. Delmon, Ind. Eng. Chem. Res. 32, 2516 (1993). 6. C. N. Satterfield and S. H. Yang, Ind. Eng. Chem. Proc. Des. Dev. 23, 11 (1984). 7. R. T. Hanlon, Energy Fuels 1, 424 (1987). 8. S. S. Shih, K. N. Mathur, J. R. Katzer, H. Kwart and A. B. Stiles, Prep. - Am. Chem. Soc, Div. Pet. Chem. 22, 919 (1977). 9. M. Jian and R. Prins, Stud. Surf. Sci. Catal. 101, 87 (1996). 10. S. Kasztelan, T. Courieres and M. Breysse, Catal. Today 10,443 (1991). 11. M. Jian, J. L. Rico Cerda and R. Prins, Bull. Soc. Chim. Belg. 103, 225 (1995). 12. C. Moreau, J. Joffre, C. Saenz and P. Geneste, J. Catal. Ill, 448 (1990). 13. M. Jian and R. Prins, Catal. Today 30,127 (1996). 14. B. S. Gevert, J.-E. Otterstedt and F. E. Massoth, Appl. Catal. 31, 119 (1987). 15. V. Stuchly and L. Beranek, Appl. Catal. 35, 35 (1987). 16. B. Delmon, Bull. Soc. Chim. Belg. 104, 173 (1995). 17. J. L. Portefaix, M. Cattenot, M. Guerriche, J. Thzivolle-Cazat and M. Breysse, Catal. Today \{S,Al'i{\99\). 18. R. M. Laine, Catal. Rev. Sci. Eng. 25,459 (1983). 19. H. Schulz, M. Schon and N. M. Rahman, Stud. Surf. Sci. Catal. 11,201 (1986). 20. T. C. Ho, Catal Rev. Sci. Eng. 30, 117 (1988).
® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
A KINETIC MODEL FOR HYDRODESULFURISATION
M. Sau, C.S.L. Narasimhan & R.P. Verma Indian Oil Corporation Limited, Research & Development Centre Sector-13, Faridabad - 121 007, India
Abstract
Due to stringent environmental considerations and related insistence on low sulfur fuels, hydrodesulfurisation has emerged as an important component of any refining scheme globally. The process is used ranging from Naphtha/Kerosine hydrotreating to heavy oil hydretreating. Processes such as Deep gas oil desulfurisation aiming at reduction of sulfur levels to less than 500 ppm have emerged as major players in the scenario. Hydrodesulfurisation (HDS) involves parallel desulfurisation of different organo- sulfur compounds present in the complex petroleum mixtures. In order to design, monitor, optimise & control the HDS reactor, it is necessary to have a detailed, yet simple model which follows the reaction chemistry accurately. In the present paper, a kinetic model is presented for HDS using continuum theory of lumping. The sulfur distribution in the reaction mixture ii treated as continuum and parallel reaction networks are devised for kinetic modelling using continuum theory of lumping approach. The model based on the above approach follows the HDS chemistry reasonably well and hence the model parameters are almost feed invariant. Methods are also devised to incorporate heat and pressure effects into the model. The model has been validated based on commercial kero-HDS data. It is found that the model predictions agree with the experimental/commercial data.
Introduction Hydrodesulfurisation (HDS) of petroleum fractions has gained considerable importance in the international refining industries in order to meet the stringent environmental regulations. Hydrodesulfurisation of products e.g. Kerosine/ATF, Diesel and Cracked Gas Oils are carried out to upgrade the products quality and render them environmentally friendly. Apart from the above, HDS is also used to pretreat the feed stocks for secondary processing e.g. FCC, Hydrocracking, Delayed coking etc.
*
Author to whom correspondence to be addressed.
421
422 The hydrodesulfurisation process involves removal of sulfur from the complex petroleum mixture through hydrogenation. The organo sulfur compounds loose sulfur on hydrogenation. Sulfur is released in the form of hydrogen sulfide. The reaction is carried out under moderate pressure in a trickle bed reactor. The feed petroleum mixture consist of a distribution of different classes of organo-sulfur compounds depending on the boiling range (TBP). The complex sulfur compounds are generally in the form of mercaptans, sulfides, disulfides, thiophenes, benzothiophenes, dibenzothiophenes etc. As different feedstocks have different distribution of the above compound classes, they greatly influence the reactor operation and performance. The rate of hydrodesulfurisation is significantly inhibited by formation of hydrogen sulfide as the reaction progress. Also the reactor performance is influenced by trickle bed hydrodynamics. Mathematical modeling of any reaction system is very important because it is a powerful tool in their design, optimisation & operation/control. Models are also useful in evaluation of catalysts and feed stocks. As proper design and optimisation of process parameters, selection of suitable catalysts etc., are absolutely necessary for profitable HDS operation, it is necessary to have a reliable and fairly accurate model for the purpose. There are different approaches reported in the literature on modeling of HDS. Majority of them consider a single lump model with first order or power law kinetics. These are limited with the fact that the kinetic parameters are feed variant and they can not account for hydrogen sulfide inhibition. Yui(1989) developed a single-lump power law model which can approximately include hydrogen sulfide inhibition. Korsten & Hoffmann (1996) have presented a single lump model with Langmuir - Hinshelwood (L-H) kinetics to account for HoS inhibition. Although the performance of single-lump models are reportedly good for specific feed stocks, they are not able to account for effects of feed variant. In order to reasonably augment the effects of feed variation, a novel HDS model is presented in this paper which applies continuum theory of lumping. The model considers a feed characterisation based on TBP and broad compound class distribution (e.g. mercaptans, sulfides, disulfides, thiophenes, benzothiophenes, dibenzothiophenes etc.). The model also considers HnS inhibition through L-H kinetics to individual components of the reacting mixture. As temperature & pressure effects are quite critical, the model describes a novel method to determine the heat & pressure effects. A separate hydrodynamic model for trickle bed determines the flow regime & pressure drop. The model is tested with a commercial Kero-HDS unites data.
423 Continuum Theory of Lumping ks t h e petroleum reaction mixture for hydrodesulf urisation is so complex t h a t it is no longer worth d i s t i n g u i s h i n g individual species. The continuum t h e o r y of lumping (Chou & Ho, 1988) c o n s i d e r s t h e r e a c t i v e mixture to form a c o n t i n u o u s mixture with r e s p e c t to i t s species t y p e , boiling point, molecular weight & so forth. This is in c o n t r a s t to d i s c r e t e lumping t e c h n i q u e wherein t h e reaction mixture is divided into finite lumps based on p r o d u c t slate or boiling point r a n g e e t c . The major a d v a n t a g e of continuum t h e o r y of lumping(CTL) is i t s closeness to c h e m i s t r y of p r o c e s s . The a p p r o a c h is elegant a n d r e s u l t s in r e d u c t i o n of number of model p a r a m e t e r s . I n c o n t r a s t t o d i s c r e t e lumping a p p r o a c h , CTL formulation t r e a t s individual nonlinear kinetics e.g. L-H kinetic with ease. Recently, continuum t h e o r y of lumping a p p r o a c h h a s been applied to h y d r o c r a c k i n g p r o c e s s (Browarzik & Kehlen, 1995; Laxminarasimhan e t ah, 1995 & 1996). Feed characterisation The petroleum feed mixture for HDS mainly c o n t a i n s different d i s t r i b u t i o n of t h e following o r g a n o - s u l f u r compound c l a s s e s d e p e n d i n g upon boiling r a n g e and t y p e of mixture. (a) (b) (c) (d) (e) (f)
Mercaptans Sulfides Disulfides Thiophenes Benzothiophenes Dibenzothiophenes & i t s d e r i v a t i v e s
A normal Kerosene fraction would contain mainly m e r c a p t a n s , sulfides and t h i o p h e n e s , while Diesel & h i g h e r r a n g e fraction would contain additionally d i s t r i b u t i o n of b e n z o t h i o p h e n e s & d i b e n z o t h i o p h e n e s (Murray & Gray, 1994; Swaty Mohanthy, 1989; Froment et al., 1994). Feed is t h e r e f o r e r e p r e s e n t e d a s a mixture of individual TBP(True Boiling Point) d i s t r i b u t i o n of all t h e a b o v e compound c l a s s e s . I t may b e noted h e r e t h a t a combination of all t h e individual TBP d i s t r i b u t i o n of t h e a b o v e compound c l a s s e s will r e s u l t in total sulfur d i s t r i b u t i o n of t h e feed with r e s p e c t to TBP.
424 The individual TBP distribution can be represented in functional form as follows. (i)
Mercaptans
(ii)
Sulfides
: C /e) : c^ce)
(iii)
Disulfides
••
(iv)
Thiophenes
: c^ce)
(V)
Benzothiophenes
:
(vi)
Where
Dibenzothiophenes
C^(6)
:
Cj^Ce) C^^O)
6 is normalised TBP defined as follows TBP
e =
~
TBP(l)
CD
TBP(h) - TBP(l)
where TBP(l) is the lowest TBPCInitial Boiling Point) & TBP(h) in highest TBP(Final Boiling Point) in the fraction. It may be noted that in all the above distribution C^CO) dG indicates weight fraction of the respective compound class 'x' (where xE[m,s,ds,t,b,dbl) present within 6 & 0+d0 .
Reaction Scheme The following kinds of reactions take place with various classes in the petroleum mixture during hydrodesulfurisation. Mercaptans RSH +
H2
->
RH + H2S
Sulfides 2 RH "»• HjS
R^S + 2 ^2 Disulfides
2 RH + 2 H^
RS2 + 3 H2 ThioDhenes C^ H^ S + 4H2
-^
Ci Hin + HraS
compound
425 The above reaction is similar to all class of thiophenes, benzothiophenes, dibenzothiophenes etc. Although some amount of hydrogenation and hydrocracking reaction take place during hydrodesulfurisation (HDS), they would be insignificant as the severities are quite low. Hence these reactions are not considered in the scope of present scheme.
Model formulation The reactor modeling involves consideration of both kinetic effects and hydrodynamic effects. It has been found by several workers that commercial reactor operates close to plug flow conditions with respect to wetting efficiency (Korsten & Hoffmann, 1996). Therefore, a pseudo-homogenous plug flow model would closely represent the commercial HDS. The present paper, hence, develops a scheme for pseudo-homogenous plug flow model. However, for pilot plants the partial wetting & mass transfer effects are predominant. The model can be easily modified to incorporate the intracatalyst mass transfer & partial wetting effects, in case of scale-ups of pilot plant results to commercial scale. Several workers have shown large inhibiting effects of H~S on the conversion (Gates et al., 1979; Vrinat, 1983; Parijs & Froment, 1986; Parijs et al., 1986; Stephan et al., 1985, Papayannakos & Marangozis, 1984, Korsten k Hoffmann, 1996). To account for the above, the rate of individual components of compound classes is shown to follow Langmuir-Hinshelwood (L-H) kinetics. Reaction Rate Constants The model formulation is done applying CTL for non-linear kinetics (Chou & Ho, 1998). Based on the reported literature data, it is clear that the HDS reactivities of individual component in a compound class, monotonically decrease with increase in the molecular weight (or TBP). The individual reactivities can be mathematically represented as a function of 0 as follows
\
=Ve>
(2)
where 'x'denotes the compound class, the choice of k (0) would be based on experimental results of different model compounds. It is found that most of the reported data follow the following functional form.
k. = k^-k„^ In [ e-He-'-l)»]!^ ]
O,
426 Species T y p e Distribution Function I t would be matheniatically c o n s i s t e n t to formulate t h e model equation in kj plane r a t h e r t h a n '6' plane (Chou & Ho, 1988). The following c o - o r d i n a t e transformation is effected Cj (e,t) d e where
=
Cj (k^,t) D(kj) dkj
D(kJ =
di
dS
d0
(4) (5)
dkj
i is defined a s species index D(kJ is termed a s species t y p e d i s t r i b u t i o n function where D(kJ dk d e s c r i b e s t h e number of species between k a n d (k + dk_) If t h e r e a r e N^ species of compound class 'x' (where N ->«) assuming t h a t t h e y a r e equally s p a n n e d in '0' space,
and
D(k_) may be written a s ^^
DCkj) = N^
dk
(6) X
it may be elegant a n d simplified r e p r e s e n t a t i o n to divide '6' plane into 'N' i n t e r v a l s (N ->«) a n d consider N = N. Therefore,
de dk.
Material Balance Equation Following material balance equation along t h e bed h e i g h t *z* would be valid for each compound class
dCJk^) '^
- K C#^) [1 * ^/^ CHM
(8)
427
'E^ . _ J where,
K
[ Erry^-i^^^k;, *,) =
Adsorption
rate
constant
(9)
of HgS
HgS Y|
is stoichiometric constant for ILS.
Yj
=1
for Mercaptans
Yj
= 2
for Disulfides
j E (m,s,ds,t,b,db) M(kj) = molecular weight relationship with kj. This can be obtained from standard correlations of molecular wt. vs, TBP(8) & transferred to M(kj)
Hydrogen consumption is estimated a s follows.
(10)
where
Ha
Yj = stoichiometric co-efficient of Hj
Effect of Pressure As H2 partial p r e s s u r e would be quite high, t h e total p r e s s u r e can be approximated to H, pressure. The effect of p r e s s u r e on the reactivity may be written as follows
428 k = k (£.)»•' ^ Pr
(11)
where p^ is base case pressure and p is pressure of interest.
Effect ot tenpm^ure A simple Arrheneous relationship is considered to explain the effect of temperature :^ll.±,
(12)
where E(kj)is a functional relationship between Arrheneous constant and k. For simplicity ECk^ can be considered as a constant £ , as each compound class can be expected to have similar temperature effect. T^, is base case temperature and T is any other temperature.
Heat c/l Reaction Hydrodesulfurisation is an exothermic reaction and is related to number of moles of BJ^ formed. Based on the above the heat balance is formulated as follows.
" ^ . ^ dz
- A f f - ^ dz
(13)
where AH = Heat of reaction, 1ii = Total mass flow rate of reaction mixture, cp = Av. heat capacity. The above equation should be solved in conjunction with temperature & pressure effects.
All the above equations are required to be solved simultaneously to obtain the temperature & concentration profile along the bed.
429
Sifflpiified Model for Kero-HDS While the above generalised complex model would be quite useful for diesel & heavy petroleum fractions, the model may be simplified in case of Kero-HDS considering the presence of limited compound classes e.g. mercaptans, sulfides & thiophenes. Added to the above, it is shown by different researchers that all the above compound classes have similar HDS activity (Frank Buyan, 1995). Also in the TBP distribution the thiophenes occur towards the end of the curve (i.e. at the FBP side), while most of the mercaptans & sulfides are distributed before thiophenes (Thomson et al., 1995). From the above it can be observed that for kerosine-HDS, representation of the total *TBP* sulfur distribution as a single pseudo-compound class distribution would depict the phenomenon reasonably well. With the above understanding the generalised equations written above can be simplified for Kero-HDS as follows.
^^^
= yY
*^<*'^>
m
(15)
where Y ^^ the average stoichiometric co-efficient to account disulfides which has a value between 1 and 2, is written as follows
C/o) * 2 CJp) * Cj(p)
for
(16)
MW is the average molecular weight and 0^(0), C^ (0), 0^(0) are initial distribution of sulfides, disulfides & thiophenes, respectively.
Heat balance equation can be formulated as follows
(17) dz
di
430 The functional fca-m for reactivities to describe effect of temperature & pressure would remain as described earlier. They may be rewritten for convenience as follows.
(18)
*= Kif)'
(19)
^i^jr]
The above set of coupled integro-differential equations are solved numerically using forward 'z* marching. The resultant Cj(k,z) can be integrated to obtain the total sulfur in the feed at any z as follows
J C/k,z)D(k)iBc
where
D(k) =
(20)
di de
.
dS dk
dO dk
where 'N* (N ->•) is number of *s* components in the feed.
The above model has been applied to monitor the performance of an operating commercial Kero-HDS reactor. The results of this exercise are discussed in the following section.
431 Results & Discussions The model consists of five parameters which has been estimated during tunning with base case consisting of three data s e t s of commercial reactor accounting for LHSV, temperature and pressure effects. The parameters obtained are given below : a = 5,2768; k^, = 32.88; k ._ = 9.414; K = 42000; Y = 2.0; E = -25316
The parameter estimation is carried out using Levenberg-Marquardt algorithm and the accuracy of tunning is depicted in figures 1. The above tuned model has been used to predict the plant performance different feed rates and t h e predictions are found to be reasonably agreement with the plant data a s can be seen in figures 2 &. 3. The effects pressure and temperature on kero-HDS is depicted in figures 4 and respectively.
at in of 5
Figures 6 and 7 provide reactor profiles for total sulfur and H^ respectively at different operating conditions. The trend given by the model is in agreement with the observed trend in the plant and t h e H^S generation is well predicted by the model.
Conclusions In summary a novel approach has been formulated to model Hydrodesulfurisation reactor using pseudo-homogenous approach and continuum theory of lumping for kinetics. It is found that the approach follows the process closely and the model can be used for monitoring commercial plant performance, process optimisation and design. Also t h e model can be shaped for performing effective catalysts evaluation.
432 EFFECT OF RESIDENCE TIME ON KERO-HDS (RESIDENCE TIME • 0.S436 HR)
0.06
WT % OF SULFUR
EFFECT O F RESIDENCE TIME O N KERO-HDS (REStOENCE TIME • 0 3 8 0 9 HR)
0.0S
WT % OF SULFUR
0.04 h
0.04 h
0.03 p
0.08
0.02
0.02 h
0.01
0.01 h
100
128
160
176
200
225
260
276
TBP IN DEO C MODIL PRIDCTION
0.06
MOML PREDCTtON
« COMMIUCIAL Wlh
FIQUm NO. - 1
FIOUKE NO. - S
EFFECT OF RESIDCNCE TIME ON KERO-HDS (RESIDENCE TIME - 0.8111 HR)
EFFECT OF PRESSURE ON KERO-HDS (FRE8S-27 ATM, BASE CASE PRESS-28 ATM)
WT % OF SULFUR
0.06
» 0.04
* OOMMtRaAL DMBI^
WT % OF SULFUR
•
/* 0.03
0.03
/* 0.02
/* 0.01
0.01
0 » —r^^"""* 100 126 160
J
176
_i
200
» 226
1
1
260
276
TBP IN DE(]I C MODIL Pf»«tCnON
»^ COMMCRCtAL OMIA
PIQURE NO. - S
100
126
160
176
200
226
260
TBP IN OEQ C MODEL FRBOCnON
* COMMERCIAL DMOk
FIQURE NO. - 4
276
433 EFFECT OF TEMPERATURE ON KERO-HDS (TEMP-315 OEQ CiBASECASE TEMP-326 OEQ G) WT % OF SULFUR
0.M
H2S CONCENTRATION PROFILE IN THE REACTOR (AT OlFFEnENT OPERATING CONDITIONS)
12
1.0E-04* H28 CONC. MOLE/CJTS 1
0.03
[. ''>^^^^ i "^^r^
0.02 h
jm 1 im \ Im
1-> KMltftno* T i m * •0.8480 hr a-» f H l l < t n o » Tlw»i0.88Ot Hf a-» R M M W I O * Tim* • 0.8111 hr
m 1j ^
4>* P r » M - >7.0 Aim < * • • • C M * f f « M • 88.0 A|M>
f 1
S-'Ytemp. - 818 OEQ C f B m * C M * l^mp • 886 DEO C>
10
0.00
0.04 h
0.01
100
125
160
17«
200
226
260
276
MODIL PRBDCnON
WT % OF SULFUR 1 -^ fl**id*no* Ttm* -0.8488 hr 2-> RMtd*no* TlBt*^0.3808 hr 8-> l l * * l d * n o * Thn* » 0.8111 hr 4-» P r * M • 87.0 AlBI ( E M * C M * h * M • 88.0 Atm) «->ikm^ - 818 0 8 « c
0.6
0.4
1
C B M * C « M TlMip - 888 DBQ C) mmm, 8
-
9
^^T^
^
5
0.2 1
4 L_
1
2
^ X.
3
nOURE NO. - e
S CONCENTRATION PROFILE IN THE REACTOR (AT DIFFERENT OPERATING CONDITIONS)
0.8
2
• COMMBKCIAL Wth
FIOUKE NO. - 6
1
1
t .. ,
3
1
4
REACTOR LENGTH. M
FIQURE NO. - 7
4
REACTOR LENOTTH. M
TBP IN DEG C
X
1
6
434 References 1.
Murray R. Gray,'Upgrading petroleum residues and heavy oils*, Marcel Decker, inc, 1994.
2.
H. Korsten and U. Hoffmann, 'Three Phase Reactor Model for hydrotreating in pilot trickle-bed reactors', ACHE Journal, 42, 5 (1996) 1350.
3.
M. Chou and T.C. Ho, AICHEJ, 34, (1988) 1519.
4.
D. Browarzik and H. Kehlen, 'Hydrocracking process of n-Alkanes by continuous kinetics', Chem. Eng. Sci., 49, 6 (1994) 923.
5.
C.S. Laxminarasimhan, R.P.Verma, and P.A.Ramachandran, 'Continuous lumping model for simulation of hydrocracking',Submitted to AICHEJ, 1996.
6.
C.S. Laxminarasimhan, M. Sau, and R.P.Verma, 'A model for hydrocracking kinetics', Presented in ACS Symposium, Chicago (1995).
7.
S. Mohanti, PhD Thesis,'Modelling and Simulation of Hydrocracking and Hydrotreating Processes', Dept. of Chem. Engg., IIT Kanpur, India, June, 1989.
8. G. F. Froment, G.A. Depauw, and V. Vanrysselberghe, 'Kinetic modelling and recator simulation in hydrodesulfurisation of oil fractions', Ind. Eng. Chem. Res., 33, 12 (1994) 2975. 9.
B.C. Gates, J.R. Katzer, and G.C.A. Schuit, 'Chemistry of catalytic processes', McGraw-Hill, NewYork (1979).
10. M.L. Vrinat, 'The kinetics of Hydrodesulfurisation processes - A review', Appl. Catal., 6 (1983) 137. 11. I.A.V. Parijs, and G.F. Froment, 'Kinetics of hydrodesulfurisation on a CoMo/ -AljOj Catalyst :1. Kinetics of the hydrogenolysis of thiophene', Ind. Eng. Chem. Prod. Res. Dev., 25, (1986) 431. 12. I.A.V. Parijs, L.H.Hosten and G.F. Froment, 'Kinetics of hydrodesulfirisation on a CoMo/y-AUO- Catalyst :2. Kinetics of the hydrogenolysis of Benzothiophene', Ind. Eng. Chem. Prod. Res. Dev., 25, (1986) 437. 13. R. Stephan, G. Emig and H. Hofmann, 'On the kinetics of hydrodesulfurisation of gas oil', Chem. Eng. Proc, 19 (1985) 303. 14. N. Papayannakos and J. Marangozis, 'Kinetics of catalytic hydrodesulfurisation of a petroleum residue in a batch recycle trickle bed reactor', Chem. Eng. Sci., 39 (1984) 1051. 15. Frank Buyan, 'Mobil catalytic hydrodesulfurisation process for distillate', Mobil Technology Seminar, November 1995, New Delhi. 16. J.S. Thomson, J.B. Green, T.B. McVilliams, and G.P. sturm,Jr, 'Analysis of sulfur compound in light distillates', Presented in ACS Symposium, Chicago (1995).
435 17, S.M. Yui, 'Hydrotreating of bitumen derived coker gas oil: Kinetics of hydrodesulfurisation, hydrodenitrogenation, and mild hydrocraoking and correlation to predict product yields and properties', AOSTRA J, Res. 5 (1989)211-224.
This Page Intentionally Left Blank
® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Frotnent, B. Delmon and P. Grange, editors
437
HYDROTREATING OF GAS-OILS: A COMPARISON OF TRICKLE-BED AND UPFLOW FIXED BED LAB SCALE REACTORS* Rune Myrstad^, Jorunn Steinsland RosvoU^, Knut Grande** and Edd A. Blekkan^ ^SINTEF Applied Chemistry, N-7034 Trondheim, Norway. Statoil Research Centre, Postuttak, N-7005 Trondheim, Norway. ^Department of Industrial Chemistry, The Norwegian University of Science and Technology (NTNU), N-7034 Trondheim, Norway. E-mail: [email protected] 1. INTRODUCTION Hydrotreating reactions of distillates and gas-oils are industrially performed in trickle bed reactors, on the scale of several metres in diameter and height. For the development of new catalysts and processes in the laboratory, a realistic scale down of this process is necessary [1]. The smaller the scale of the laboratory unit, the lower the investment and operating costs, smaller amounts of materials to handle and dispose of, and increased safety of the operation. Lab-scale reactors are usually much shorter than the industrial reactors, and thus have very different height to diameter ratios, leading to very much lower linear gas and liquid velocities when operated at similar residence times [2]. For catalyst testing with the purpose of discriminating between commercial catalysts not only should the ranking of catalysts be correct, the lab-scale or pilot testing operation should also allow prediction of the full scale operation with respect to feedstock/operating conditions, catalyst and product quality. To fulfil this ambition it follows that the testing must be done with real feedstocks. The choice of reactor type is then limited to either using trickle bed reactors, or up-flow fixed bed reactors. The trickle bed reactor emulates the physical conditions in the industrial reactor, whereas the upflowfixedbed reactor assures complete wetting of the catalyst. For a trickle bed reactor the scale down must obey several important criteria [1-3], of which plugflowoperation and catalyst wetting are the most important. Plug flow is generally assumed if the Mears criterion (eq. 1) is fulfilled [4]. However, Gierman [3] points out that this is a conservative estimate, and recommends a slightly less stringent criterion, using a factor 8/Pe in the same equation as the criterion for plug flow. L 20 dp Pe
1 1 l-X
where L is reactor length, dp is the catalyst particle diameter, Pe is the Peclet number
*We thank Statoil for financial support and permission to publish this work.
^^^
438 representing axial dispersion, Pe = (u-dp/DJ, u = liquid mass velocity, D^^ is the effective axial dispersion coefficient, X is conversion, and n is the reaction order. The minimum reactor height will consequently increase with the order of reaction, and is sensitive to the level of conversion at very high conversions. When operating small reactors with commercial catalysts with rather large particles it is therefore necessary to dilute the catalyst bed with small, inert particles to avoid deviations from plug flow. This has the effect of reducing the average dp and hence the catalyst particles in the form of cylinders or other shapes can be used without crushing. This was verified experimentally by van Klingen and van Dongen [5], who also found that dilution improved the catalyst wetting. Gierman [3] points out that the important phenomenon in this respect is irrigation of the catalyst particles, that is the continuos renewal of the liquid inside the catalyst pores. The irrigation is improved by a higher liquid viscosity, hence this is a critical parameter for the study of light fractions. A proposed criterion [1,3] is given in eq. (2). The dimensionless number W compares the frictional force of the flow with the gravity
W =-
^
> 5-10-^
(2)
where r\^ is the dynamic viscosity of the liquid, PL is the density of the liquid and g is the gravity constant. force. For a 1 m reactor used for LGO or HGO hydrotreating an average dp of 1-2 mm can be estimated from this correlation [1,3]. Harold and Ng [6] developed a model for a phenomenon they term dewetting due to vaporization of the liquid. This is a problem in strongly exothermic processes such as the hydrogenation of olefines. The laboratory reactors are usually operated in an isothermal mode, and by using diluent particles with a high heat conductivity, such as SiC, the heat transfer through the bed can be improved and a flat radial temperature profile obtained [1]. As seen, the scale down of trickle bed reactors is difficult and care has to be taken to avoid deviations from plug flow and poor catalyst utilization. Instead, the use of concurrent upflow reactors could be considered [7,8]. This involves a different flow pattern (bubble flow) and a high liquid holdup. De Wind et al. [8] recommends this mode of catalyst testing, and shows that it gives results comparable to those of commercial hydrotreating units. The purpose of this paper is to give some experimental results comparing catalyst testing in bench scale/small pilot scale units employing trickle bed as well as upflow reactors, and comparing with activity data from commercial use of the same catalyst. 2. EXPERIMENTAL 2.1. Trickle bed unit The trickle bed unit is shown schematically in Fig. 1. The liquid is pumped from one of 3 reservoirs with different feedstocks, using a reciprocating metering pump. The pump calibration is checked regularly by pumping from a burette. The unit has two gas feed lines, one for H2 and one for Ar used as internal standard for on line gas analyses. Both lines are equipped with FlowTech flow controllers. The Ar line is also used for O2 or air if the catalyst is regenerated in the
439
JJoUE
GAS CHROMATOGRAPH
Jh-Q—> WGM
N^
/ BALANCE \
J 11-
N2
-iS3-
H2
LOW PRES. SEPARATOR \1/
[Qg MFC
MASS FLOW CONTROLLER
Q
WET GAS METER
WGM H
ROTAMETER
A PCV
PRESSURE CONTROL VALVE
[J
PRESSURE TRANSMITTER 3-WAY VALVE
OIL RESERVOIR, / BALANCE \
^
MANOMETER
PUMP
Figure 1. Schematic drawing of the trickle bed unit. reactor. A separate line for pretreatment gas (H2S/H2) is metered using a gas rotameter. Downstream of the reactor the products are quenched and condensed in a cooler, and liquids and gases are separated in a high-pressure separator. The gases are sent to vent after depressurizing and scrubbing. A side stream allows the analysis of the gas composition by on-line GC. The liquids are collected in a product tank (atmospheric pressure) where they are stripped with N2 to remove dissolved H2S and NH3. The reactor consists of a stainless steel tube with an i.d. of 37.5 mm. The first section (length 0.51 m) is a preheater,filledwith ceramic rings. The reactor part (length 0.78 m) has a total volume of approx. 800 cm^ The catalyst volume used is 400 cm^, diluted with SiC (particle size range 500 - 710 ^m) in a volume ratio of 1:1. The top of the reactor is filled with SiC (50 cm^) to aid in liquid dispersion over the cross section of the catalyst bed. The reactor is placed in a vertical electrical fumace formed as two half shells that can be clamped closely around the reactor. The fumace is divided into 4 segments, each controlled by a separate thermocouple and temperature controller. 2.2. Up-flow unit The up-flow unit (Fig. 2) is constructed in a fashion very similar to the trickle-bed unit. The reactor is a stainless steel tube (1 m x 19.3 mm i.d.) with a total volume of 290 cm^ The length of the catalyst bed is adjusted using the diluent particles (SiC, 500 - 710 |Lim), and the catalyst is diluted in a volume ratio 1:1 with SiC. In the present experiments a catalyst volume of 75 cm^ was used, giving a bed height of about 535 mm. The reactor is placed vertically in a heating block made of bronze and heated in an electric fumace with two separate heating zones. The reactor is equipped with an axial thermowell containing 4 thermocouples, 3 of which are in the catalyst bed.
440
H55OIL RESERVOIR
Q.
- ^
PUMP
_f
1. PREHEAT
3 11 11 11
J ] QUE
A{
GAS CHROMATOGRAPH
H53 HIGH PRES. SEPARATOR
1.
14-^
a
MFC
MASS FLOW COrfTROLLER
B
ROTAMETER
PC
PRESSURE CONTROLLER
^
3-WAY VALVE
?
MANOMETER
A
PUMP
Figure 2. Schematic drawing of the upflow fixed bed unit. 2.3. Materials The catalyst was a commercial C0-M0/AI2O3, used as received from the suppliers.The feed used in the laboratory experiments was a light gas oil (LGO) drawn from the commercial unit. The feed contained 0.16 wt% S, 49 wt ppm N and 27 wt% aromatics (measured using HPLC). The boiling range was 240 - 341 °C, and the density was 855 kg/ml 2.4. Procedures The laboratory reactors were operated using identical procedures. After closing the reactor and pressure-testing the unit, the system was flushed with N2 and sulfided in 3% H2S in H2 (P = 1 bar, GHSV = 300 hr'\finaltemperature 350 °C held for 4 hours). The temperature was then reduced to 200 °C, the gas changed to the H2 feed gas and the pressure increased to that of the first planned experimental point. The liquid feed was introduced, and the conditions were then adjusted to those of the first experimental point (gas flow, liquid flow and finally temperature). The system was allowed to stabilize for 48 hours before collecting the products for analysis of the first point. Products were subsequently collected once every 24 hours, followed by adjustment of the conditions to a new set point. Sulfiir and nitrogen contents were analyzed using an ANTEK 7000 analyzer and corrected for product density. 3. RESULTS AND DISCUSSION The results from hds and hdn of LGO in the two laboratory reactors are reported as plots of residual S and N in the liquid products. Fig. 3-a (hds) and Fig. 3-b (hdn) shows a comparison of
441 the trickle bed unit and the upflow unit. Any differences in the performance of the reactors, particularly with respect to deviation from plug flow, should be observed at the high conversions obtained here (%hds in the range 90-97%). The virtually coinciding curves demonstrate that the units have the same catalyst efficiency. Due to the fundamental difference in operation, this can be interpreted as proof that both reactors operate as plug flow reactors. Furthermore, the results also show that there are no problems with catalyst wetting, with complete irrigation of the catalyst in the trickle bed. This is not surprising, taking into account that all the reported guidelines for design have been followed closely. Furthermore, both reactors also have a controlled temperature profile, since the reaction rates are very sensitive to the temperature. Temperature measurements along the reactor axis confirm this, but we have no measure of the radial temperature gradient in the reactors. However, the calculations reported by Sie [1] for a case very similar to the trickle bed used here, shows that the maximum expected temperature difference between the reactor centre and reactor wall is about 1-2 K, which is within the uncertainty of the temperature measurement. The smaller upflow unit has should have an even smaller radial temperature gradient. Fig. 4 shows a comparison of the trickle bed unit and the commercial unit using the same catalyst. The data points from the commercial unit were recorded shortly after start-up with a new catalyst batch. In spite of different pretreatment methods and comparing isothermal operation in the laboratory to the weighted average bed temperature (WABT) of the adiabatic industrial scale reactor, the data points fall close to the experimental curve from the laboratory. Some of the scatter in the points in the commercial unit is due to small variations in the crude oil base over time
200 1
40
g 150
£ 30
a a.
a
Q.
Z
75
0) 100
3
15
3 T3 CC 50 H
20
0)
O : Upflow fixed bed reactor D : Trickle bed reactor
00
OC 1 0
00
O : Upflow fixed bed reactor D : Trickle bed reactor
1
\
1
Base-10 Base Base+10
Base-10 Base Base+10
Reactor temperature
Reactor temperature
Figure 3. Comparison of laboratory reactors in hds (a) and hdn (b) of LGO. Conditions: P = 28.5 bar, LHSV = 1.5 hf\ H2:oil = 69 NmVml
442
100 1
O
"55 o >
O
96
o :£
D Trickle bed lab. reactor ^ Commercial unit
94
(0
92 Base-10
Base
Base+10
Reactor Temperature °C Figure 4. Hydrodesulfurization of LGO at high pressure: Comparison of desulfurization in the laboratory trickle bed unit and a commercial unit. in the refinery. The results reported her illustrate that the chosen experimental setups, both the trickle bed and the upflow fixed bed, are well suited for catalyst testing, and that the results obtained can be used not only in catalyst ranking, but are also valid in terms of production planning and product quality. REFERENCES 1. S.T. Sie, in P. O'Connor, T. Takatsuka, and G.L. Woolery (Ed.), Deactivation and testing of hydrocarbon-processing catalysts (ACS Symp. Ser. no. 634), ACS, Washington, DC, 1996, p.6. 2. S.T. Sie, Rev. Inst. Francais Petr., 46 (1991) 501. 3. H. Gierman, Appl. Catal., 43 (1988) 277. 4. D. Mears, Chem. Eng. Sci., 26 (1971) 1371. 5. J. Van Klinken and R.H. Van Dongen, Chem. Eng. Sci., 35 (1980) 59. 6. M.P. Harold and K.M. Ng, Ind. Eng. Chem. Res., 32 (1993) 2975. 7. C.N. Satterfield, AIChE Journal, 21 (1975) 209. 8. M. de Wind, F.L. Plantenga, and J.J.L. Heinerman, Proceedings, Symposium on catalyst performance testing. Unilever Research Laboratories, The Netherlands, March 28-29 1988. p. 29.
® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
443
Trickle-bed reactor modeling for middle-distillates hydrotreatment C.G. Dassori^, N. Femdndez^, R. Arteca^, A. Diaz'^ and S. Bxiitrago* ^Intevep S.A., Apdo.76343, Caracas 1070A, Venezuela ^Universidad Sim6n BoKvar, Apdo.89000, Caracas 1080A, Venezuela
A comprehensive approach is presented for trickle-bed reactor modeling, which covers proper characterization of reactive species, the discrimination of major reactions and the development of a reactor model within process simulators' architecture. The reduction of aromatics content in middle-distillates is presented as a case study. Inlet and outlet streams are represented in terms of pseudocomponents determined from mass spectrometry, boiUng point distribution, molecular weight and elemental analysis. A global optimization algorithm determines the mixture composition and distribution of pseudocomponents, discriminated as paraflfins, naphthenics, monoaromatics and poUaromatics of different carbon nimibers. Thermod3niamic and transport properties of these species are also predicted. This analysis is applied to a set of experimental data obtained from a laboratory reactor. Inlet and outlet streams described in terms of pseudocomponents are used to determine, the most important reactions during the process from raw data and using an optimization program. Kinetic parameters for selected reactions are determined using a trickle-bed reactor model that is viewed as a collection of repetitive cells. The computational program is developed within a commercial process simulator. 1. E«*RODUCTION The reduction in aromatics content of diesel fuels is strongly driven by new regulations. These processes are commonly carried out in fixed-bed reactors where hydrogen is contacted with a liquid hydrocarbon feed. Depending upon operating conditions, partial or complete vaporization of inlet Uquid stream can be achieved within the reaction vessel. This paper presents a reactor model that incorporates liquid vaporization through the use of a process simulator for the vapor-Hquid equilibriimi. Complex hydrocarbon feeds have to be characterized in terms of pseudospecies that retain chemical composition information in order to develop a model for hydrotreating reactors. Mass spectrometry, elemental analysis and molecular
444
weight are employed as raw data sources from which pseudospecies are bxiilt in this work. A bank of pure hydrocarbon compounds is used to match the analj^ical information. The resulting mixture fits raw analjrtical data. This large set of molecules is lumped into compoimd classes, each one divided into carbon number groups. Using this approach, a manageable group of limiped species is obtained, from which a reaction mechanism is determined. The reactor model is used to predict conversion of reactants and product distribution as a fimction of inlet conditions, temperature and pressure for non-ideal mixtures. The process simulator provides thermodynamic information through appropriate vapor-Uquid equiUbriimi models. Pressure drop, phase holdups and catalyst wetting are incorporated into the reactor model through correlations available in the literature[2,4]. 2. EXPERIMENTAL An isothermal fixed-bed reactor was used for obtaining experimental data. Its length was 65 cm and it had a 25mm internal diameter with three axial points where internal temperature was measured. Liquid and gas feeds were introduced from the top. Pure hydrogen was used as inlet gas stream. An infrarred jacket was used to heat up the bed. Effluents were separated into gas and Uquid phases at reactor outlet and analyzed as reported in Table 1. The following parameters were varied: space velocity, pressure, hydrogen/ hydrocarbon inlet ratio and temperature. Catalyst C-448(Criterion) was used for the experiments and was diluted 1:1.5 with SiC.
Table 1 Analytical Characterization Methods Analysis
Norm
Sulfur API Gravity Digital Densimetry Simulated Distillation Mass Spectrometry Molecular Weight (osmometry) C (atomic spectroscopy) H (atomic spectroscopy) Gas analysis
ASTM/D-1552 ASTM/D-1298 ASTM/D-2887 ASTM/D-2425 VPO/INTEVEP Combustion Combustion GC
445
3. REACTOR MODEL The fixed-bed reactor was modeled as a group of repetitive cells as shown in Figure 1. It is implied that gas-liquid equilibrixmi prevails along the bed. Kods and Ho[l] have made similar assxmiption when a purely liquid phase reaction was considered. Pressure drop across cells and liquid holdup, ei, are computed using correlations by Larachi et al.[2]. Once ei is computed at a given step, liquid CSTR volume is calcxolated by: CSTRiiq.volume = eiAV
(1)
AV = V T / n
(2)
where V^ is total reactor volume and n is the nimiber of cell units. A similar procedure is followed to compute CSTR gas volxmie. Both flash and CSTR imits are computed by using PRO/II® (release 4.01)[3] process simulator of Simulation Science Inc. Catalyst wetting fraction, f^ was computed using the correlation proposed by Mills and Dudukovic[4]: f - a L (3s-l)/2
(3)
In the present work a = 0.85 and s = 0.52; LQ, is the liqtiid mass superficial velocity.
i_J. FLASH
T7. CSTR [^ [S] CSTR FLASH CSTR
CSTR
Figure 1. Reactor model scheme
446 4. PSEUDOCOMPONENT EVALUATION A bank of pure molecniles was used to fit analytical data using a global optimization program that gives as a result their respective concentrations in the sample. There were used more than 500 different molecules belonging to three compoxmd classes: paraffins, naphthenes and aromatics. From the selected optimxim set, different pseudospecies groups were constructed: paraffins(Pi), naphthenes(Ni), monoaromatics (MAi) and poliaromatics(PAi) of different carbon numbers; index i indicates carbon number group. That is, i=l, carbon atoms from 1 to 6; i=2, carbon atoms from 7 to 12, etc. Each pseudocomponent has a complete thermodynamic characterization in terms of critical and transport properties that are computed using proper mixing rules for the composition set determined in the optimization program. 5. REACTIONS By analyzing lumps distribution at reactor outlet and comparing them to feed composition a set of reactions was selected for representing the reaction network. This selection was done from a group of sixty possible reactions among the pseudocomponents obtained previously. Based on a global optimization technique[5] the integral mass balance for pseudopecies in the laboratory reactor is used to discriminate the largest reaction rates imder the conditions considered in this work. The following reactions were finally selected:
N03 PA02 PA03 PA04 PA05
•
^ ^
•^ -M
^ '—^^ ^ • •
N02 + POl
MAn9.
MA03
^
IVLrlUrt
^
MAHJ.
MA05
^
... ^^ ^
N02 N03 N04
Rate laws for each reaction were assumed to be first order in both hydrogen and hydrocarbon species for each phase. The following expressions were used: for the liquid phase rijL = ^ijKjWf^CH2^CiL
(4)
and for gas phase: rijV = ^ijKjW(l-f^)CH2^CiV
(5)
where ry is the reaction rate for i-th component in j-th reaction, Kj is the kinetic consteint in j - t h reaction, W is catalyst weight in cell, xy is stoichiometric coefficient for reactant i in j - t h reaction, f^ is the wetting factor, C H 2 is hydrogen concentration, Ci is hydrocarbon concentration. Superscripts L and V stand for liquid and gas phase.
447
6. RESULTS AND DISCUSSION Experimental data obtained in the laboratory were used to determine kinetic constants for the reactions described above. Figure 2 shows that 25 is an adequate niimber of cells for performing the computations because increasing this value to 150 has minor effects on mixture composition in terms of global limips. Pseudocomponents concentrations are sensitive to number of cells when this value is smaller than 25. This is the mmierical evidence that the chemical reaction and gas-liquid mass transfer are adequately considered as simultaneous events and that an adequate discretization is employed (number of cells). Figure 3 shows the effect of liquid hourly space velocity (LHSV) on poliaromatics conversion of group 5 (PA5, more than 24 carbon atoms). Figure 4 shows the natural trend of increasing conversion with increasing hydrogen inlet pressure according to reaction rates employed (eqns.4-5). Figure 5 shows the difference in predicted conversion of group 5 poliaromatics depending on the equation of state selected to use in the process simulator(PR=Peng-Robinson; SRKM=Modified Redlich-Kwong-Soave).
H150 cells • 25 cells PRESSURE (atm};68,04S LIQUID aOWRATE (KgA^): 2.4587E-02 GAS FLOWRATE(Kg/h);7,9977E-02
Figure 2. Product distribution in terms of compound families
0,3
0,4
0.5
0,6
0.7
Dimensionless Axial Position
Figure 3. Effect of LHSV. T=650K;P=68atm;FH2/Fc=500Nm^/m3
448
0.3
0.4
0.5
0,6
0.7
Dimensionless Axial Position Figure 4. Effect of Pressure. T=655K;LHSV=lh-^FH2/Fc=544NmW;SRKM
- T.657.15 K.P.68.(M5irtm.SRKM. | -T-657.15 K. P-68.045«tm. PR. - EXPERIMENTAL VALUE
0.3
0,4
0.5
0.6
Dimensionless Axial Position
J— 0.7
Figure 5. Effect of Equations of state. LHSV=1.64h-^;FH2/Fc=500Nm3/m3 General results obtained with the present model indicate that proper thermodjntiamic properties and compound class limiping can be effectively used when modeling hydrotreating reactors within process simulator's architecture. REFERENCES 1. G.R. Kocis and T.C. Ho, Chem.Eng.Res.Des., 64 (1986) 288. 2. F. Larachi, A. Laurent, N. Midoux and G. Wild, Chem.Eng.Sci., 46 (1991) 1233. 3. PRO/II is a registered mark of Simulation Sciences Inc. 4. P.L. Mills and M.P. Dudukovic, AIChE J., 27 (1981) 893. 5. S. Buitrago and C.G. Dassori, Intevep Internal Report (1996).
® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P, Grange, editors
449
Petroleum residua hydrotreating on Co and/or Ni containing catalysts V. M. Kogan* and N. M. Parfenova^ ^N.D.Zelinsky Institute of Organic Chemistry, Russian Academy of Sciences, 47 Leninsky prospect, Moscow 117913, Russian Federation; ^stitute of Chemistry, Turkmenian Academy of Sciences, 92 Pogranichnicov st, Ashgabat 744012, Turkmenistan. A comparative radioisotope study into alumina- and silica supported Co, Ni and Co+Ni catalysts has been carried out. Catalysts were characterized by physical and chemical methods and tested during hydrotreating (HTR) of oil residua of WestSibirian petroleum. Some dependencies among catalyst composition, carrier support, the number, type and productivity of active sites, and catalyst fimctioning under oil residua HTR have been found. The study has permitted us to put forward some criteria to evaluate the results of radioisotope testing with the aim of designing a catalyst composition optimum for a definite refining process. Some cheap catalysts-adsorbents for preliminary treating of heavy crudes, before they are applied to the main oil refining processes have been designed. 1. INTRODUCTION Involving heavy residual oils into refining processing makes it of major importance to have effective catalysts for oil HTR. When developing such catalysts many parameters are taken into accoimt - the type of residua, the range of products obtained, their quality as well as HTR degree. All attempts to create universal catalysts that would effectively catalyze processes of desulfurization, demetallization, denitrogenation and hydrocracking have failed so far because of catalyst deactivation which comes very soon. To use multy-zone reactors or multy-reactor systems in which different catalysts are to perform their ownfimctionshas proved to be more efficient. As a rule, at the first stage catalysts for heavy crudes demetalization are applied. These are more efficient as they have rather low metal concentration ( 7 - 8 %), do not contain expensive Mo or W and efficiently carry out demetallization and desulfiuization. Thus pretreated without cracking, crudes can be fiuther effectively refined. Some major concepts of HTR reactions and architecture of the supported catalysts have been already formulated [1 - 22]. Nevertheless, studies into supported catalyst composition and principles of their fimctioning are still very urgent to which fact testifies a great number of articles and reviews [10, 11, 23 - 29] which are mainly focused on the issue of catalyst composition and not so often on reaction mechanisms.
450 Radioisotope technique offers us extra opportunities to obtain unique information on active sites number and their functioning and heterocompound transformation mechanisms. Earlier, using radioisotopes ^^S and ^H we studied thiophene hydrodesulfurization (HDS) on various Co(Ni)/Mo sulfide catalysts supported on alumina and activated carbon [30-36], We found that catalyst surface sulfiir participates in H2S formation. In the course of reaction some part of catalyst sulfiir (mobile sulfiir) is replaced by thiophene sulfiir. The amoimt of sulfide catalyst mobile sulfiir, in its turn, depends on catalyst composition and sulfidizing technique. An original forcing out scheme of thiophene HDS mechanism was suggested. Due to it, when thiophene is adsorbed on anion vacancy catalyst SH groups interact with molecular hydrogen, which results in H2S formation. This step is limiting for the reaction as a whole. An original radioisotope technique for sulfide catalyst testing was developed. Our studies were made both on model catalytic systems and commercial catalysts. There has been no experience in comparing the results obtained with the data on real crudes conversion. In this paper the comparison of the results of catalyst radioisotope testing during model reaction of thiophene HDS with the data of residual oil HTR on the same catalysts is being undertaken. It aims at adequate evaluating of both methods potentials to design the best catalyst composition for an actual HTR process. 2. EXPERIMENTAL 2.1. Catalyst preparation, characterization and pretreatment More than 70 Ni- and Co-containing catalysts were synthesized and studied. Catalyst samples were prepared by wet impregnation of y-Al203 (specific surface - 212 m^/g, pore volume - 0.8 cm^/g, pore average D - 98 A). Impregnation procedure was carried out (i) using water solutions of Ni or Co nitrates for preparing Ni or Co catalysts, or (ii) using join solution of Co and Ni nitrates - for Ni+Co catalyst samples preparing. After impregnation the catalysts were air dried for 24 h at room temperature and 2 h at 110°C. Metal surface amount in the samples varied in the range of 1 - 25 %. All the carriers and catalysts were studied by physico-chemical methods. Pore structure was determined by Hg porometry technique using Cultronix automatic porometer. Metal reduction degree was determined by volumetric technique. Metal concentration in the catalysts was measured by atomic absorption technique. Two forms of the catalysts were used - reduced and sulfidized. Sample reduction was carried out in the flow of purified H2 for 10 h at 400°C. Sulfidation of preliminary reduced catalysts was made by elemental sulfur in hydrogen atmosphere (3 MPa, 380°C, 1 h). 2.2. Residual oil hydrotreating technique and product characterization Residual oils of West-Sibirian petroleum contained: asphaltenes - 4.3 % wt, sulfiir - 2.3 % wt, V - 40 g/t, Ni - 60 g/t. Residual oil HTR was carried out in a 250 cm^ stainless steel autoclave (5 MPa H2 pressure, 380°C, 5 h). Sulfiir amount in crudes was determined analytically. Derivatographic measurements of coke deposits on the catalyst were made by Q-1500 D installation (Hungary).
451 2.3. The technique of HDS experunents with radiochromatographic analysis of the products. 2.3.1. Catalyst sulfidation was carried out in a pulse microcatalytic installation. The microreactor in the form of U-tube made of pyrex was loaded with 100 mg of the catalyst and linked with the gas-liquid (GL) chromatograph. The rest of the reactor volume was filled with quartz. All catalysts were preliminary subjected to He flow at 200°C for 2 h, then by H2 flow, 400°C, 4 h. After that they were sulfidized. As sulfidizing agent either 5 % H2S/H2, in which H2S was labeled by ^^S isotope, or thiophene-^^S was used. H2S/H2 sulfidation lasted for 1 h at 450°C. Further on the catalyst was again treated by He at the same temperature for 30 min. to get rid of adsorbed H2S. Sulfidation by thiophene-^^S was made by 3 jiil thiophene pulsing into 001 the reactor at 360°C with fiirther \c4H4S chromatographic analysis of the 80 products (Fig. 1). After H2S curve o became constant thiophene-^^S pulsing 60 - \ 3 C4-HC was stopped. After sample sulfidation O 40 was over, reactor temperature - J^ stabiUzed at 360°C. Then the catalyst 20 either was taken out of the reactor for ^jsr^s sulfide sulfiu- amount measurement or 1 ^ 2 4 6 8 10 was tested in the reaction of thiophene Pulse number HDS. Figure 1. Thiophene HDS product distribution vs. number of miophene- S pulses. 2.3.2. The technique of HDS experiments with radiochromatographic analysis of the products and mathematical treatment of experiment results. After sulfidation procedure catalyst samples labeled by S were tested in the thiophene HDS reaction. Unlabeled thiophene was injected into the reactor by pulses, 1 jiil each. To measure reaction product radioactivity aflowproportional counter installed in the outlet of the chromatograph was used. Based on the results obtained some curves of H2S molar radioactivity (MR) dependencies on H2S amount (cm) are built equivalent to the dependencies on reacted thiophene amounts. H2S MR values are normalized relatively to the value of initial sulfiir MR on the catalyst (%). Mathematical treatment of the experimental curves allows us to approximate the dependencies we obtained by exponential equations of the following type: n a = ll Aj exp(- Aj x) (1) i=l
452 The exponential fonn of equation (1) is determined by the fact that the processes of isotope exchange and replacement are described by equation of the first order. Every item of (1) describes H2S MR changes with intensity determined by the value of exponential powder Xj, this intensity being characteristic of active sites of the given type. The fact that it is impossible to describe the experimental curve by monoexponential equation testifies that in this case some replacement processes that differ in their intensity but are simultaneous take place. Thus, there are some types of active sites on the catalyst. They are responsible for H2S formation and differ in SH group mobility. In (1) preexponential multiplier Aj determines thiophene share converted on the site of the given type. If there is only one type of the sites (/=7) on the catalyst Aj = 100%. Proceeding from equation (1) it is possible to estimate catalyst sulfiu" amount (in mg) that can participate in H2S formation (mobile part of sulfiu") on the sites of the given type: S^ =
1 Aj expi-J^ x)dx 22.4x100 0
(2)
Integration of (2) gives:
sr—^4
(3)
22.4x100 ^ If the curve is described by more than one exponent, i.e. several types of active sites take part in the reaction, the total amoimt of mobile sulfiir is estimated by the sum: Smob=Si+S2+ (4) If conversion is known, it is possible to estimate thiophene amount converted on each type of the sites (the ratio of the nimiber of thiophene molecules to that of SH groups of the sites of given type), i.e. site productivity: Pi = 100
:
100
X
Ns
(5)
Then Pi = 100'<
Ai
vp_ /Sj_ >^ 84 / 32
(6)
100 If we; put (3) into(6) and make some simple transformations we will have:
i^=2.84xl0-V,A
(7)
453
3. RESULTS AND DISCUSSION 3.1. Radioisotope testing Preliminary experiments with various amounts of active metal on the catalyst have shown that catalytic activity curves of the samples under study pass through a maximum, depending on an active component amount. The maximum area for all the contacts under study is in the range of 7.0 - 7.6 % (Fig. 2). So, those catalyst samples that contain the given amount of active metal have been studied in detail. The results are shown in Tables 1 and 2. Measurements of radioactivity of thiophene hydrogenolysis products on all the catalysts containing sulfide sulfiir labeled by ^^S show that radioactivity is present only in H2S formed. Thiophene that leaves the reactor does not contain radioactivity, which points to the absence of isotope exchange between thiophene sulfur and catalyst sulfide sulfiir. In the intervals of thiophene ^ 4 6 8~ pulses into the reactor or under special Co loading, % wt. Iexperiments any noticeable amount of Figure 2. Thiophene HDS conversion (1) H2S formed is not foimd, i.e. any marked and mobile sulfiir amount (2) v^. Co loading sulfidized catalyst reduction does not on Co/Si02 catalysts. occur. H2S is formed only as a product of thiophene hydrogenolysis. The amount of sulfide sulfiu" on all the samples with 2.5 - 11 % Ni, subjected to sulfidation by H2S/H2 reach values close to stoichiometrically possible relative to NiS. In Co catalysts some excess of sulfide sulfiir over stoichiometric amount in CogSg is observed. For the samples with low Co content this excess approaches 100 %. For instance, a sample with 3.4 % Co/Si02 contains 3.7 % sulfide sulfiir while stoichiometric amount in CogSg, with the given Co loading is equal to 1.64 mg. It permits us to suppose the presence of C0S2 phase on the catalyst. For higher loadings such as 5.7 % and 11 % Co on Si02 sulfide sulfiir content is 1.5 time higher than stoichiometric for CogSg - 5.7 % and 8.63 % but stoichiometrically possible amounts 3.57 % and 5.3 % correspondingly. The data observed can give us ground to assume that under given loadings in Co/Si02 catalysts sulfidized by H2S two Co-sulfide compounds might present in active phase - 50 % CogSg and 50 % C0S2. Table 2 gives equations for the curves of dependencies of H2S molar radioactivity (MR) on the amount of H2S formed under thiophene HDS on the catalysts sulfidized by radioactive sulfiir. One can see that these equations are monoexponential, i.e. Co or Ni catalysts or Co+Ni catalysts have one type of H2S formation site while Co(Ni) promoted Mo catalysts have two types of sites differing on the mobility of sulfiir, i.e. productivity [35].
454 Table 1 Results of radioisotope testing of catalysts in the reaction of thiophene HDS (100 mg sample loading, 360°C, H2 flow, pulses of thiophene 1 ^1) No smp
Catalyst composition
Pretreatment agent
Stotal
Smob
Sm/St
%wt
%wt
% 100
(7.0%Co)/Al2O3
thiophene
1.16
1.16
lb
(7.0%Co)/Al2O3
H2S/H2
5.50
0.89
2a
(7.4%Co)/Si02
thiophene
2.42
2.42 0.63
la
16.2 100
p** % mol
X 10^
5.40
1.89
5.34
2.43
8.00
1.34
4.00
2.56
8.43
3.33
15.70
4.63
4.25
2.10
7.25
2.34
7.62
2.30
17.00
3.19
5.35
1.73
2b
(7.4%Co)/Si02
ri2S/ri2
5.70
3a
(7.6%Ni)/Al203
thiophene
1.03
1.03
3b
(7.6%Ni)/Al203
H2S/H2
4.00
1.38
4a
(7.3%Ni)/Si02
thiophene
0.82
0.82
4b
(7.3%Ni)/Si02
H2S/H2
3.76
1.23
5a
(3.4%Co+4. l%Ni)/Al203
thiophene
1.34
1.34
5b
(3.4%Co+4. l%Ni)/Al203
H2S/H2
4.82
2.16
6a
(3.5%Co+3.5%Ni)/Si02
thiophene
1.25
1.25
6b
(3.5%Co+3.5%Ni)/Si02
H2S/H2
4.10
0.95
23.2
4.15
2.03
Ni-Mo/Al203
H2S/ri2
6.08
2.49
41.0
59.00
20.1
y***
11.1 100 34.5 100 32.7 100 45.0 100
5.4 ^) y - thiophene conversion; **) P - active site productivity; ***) commercial catalyst Table 2 Change of H2S molar radioactivity (a) in the course of thiophene HDS on sulfide-^^S catalysts (100 m g sample 1 loading, 360°C, H 2 flow, pulses of thiophlene 1 ^l). No smp
Catalyst
H2SMR --100 exp(-1.23x)
No smp
Catalyst
H2SMR
4a
Ni/Si02
a = 100 exp(-1.74x)
a == 100 exp(-1.60x)
4b
Ni/Si02
a = 100exp(-1.16x)
Co/Si02
a =-- 100exp(-0.59x)
5a
(Co+Ni)/Al203
a = 100 exp(-1.07x)
2b
Co/Si02
a =-- 100 exp(-2.26x)
5b
(Co+Ni)/Al203
a = 100 exp(-0.66x)
3a
Ni/Al203
a =-- 100 exp(-1.39x)
6a
(Co+Ni)/Si02
a = 100exp(-1.14x)
3b
Ni/Al203
= 100 exp(-1.04x)
6b
(Co+Ni)/Si02
a = 100 exp(-1.50x)
7
NiMo/Al203
la
C0/AI2O3
lb
C0/AI2O3
2a
a = 52.08 exp(-1.20x) + 47.92 exp(-0.32x)
455 Sulfur distribution, mg
C0/AI203
C0/SIO2
NI/AI203
Nl/8i02
CoNl/AI203
CoNI/Si02
Figure 3. Sulfiir distribution v^. catalyst composition and pretreatment procedure. Left bars - samples sulfidized by thiophene; right bars - by H2S/H2. Light parts of bars - mobile sulfiir; dark ones - immobile sulfiir (see experimental conditions in Tables 1 or 2). Figure 3 shows that treating the reduced catalyst by H2S/H2 always results in much deeper sulfidation as compared to thiophene. Share of mobile sulfiir in samples sulfidized by H2S/H2 ranges within 11 - 50 % and in Ni or Co+Ni catalysts it is higher than in Co ones. It is true both for alumina- and silica supported catalysts and may be due either to high dispersion of Ni-sulfide particles as compared to Co-sulfide ones or to probable C0S2 phase formation. It is also essential that on alumina supported catalysts, sulfidized by H2S/H2, mobile sulfiir share is larger than on analogous catalysts supported on silica. It might also be accounted for by higher dispersion of sulfide particles supported on alumina as compared to that of particles supported on silica. Unlike in catalysts sulfidized by H2S/H2, in catalysts sulfidized by thiophene sulfidation does not reach maximum possible values and all sulfide sulfur formed under sulfidation by thiophene is mobile. Major portion of mobile sulfur is on Co/Si02 sulfidized by thiophene (sample 2a in Table 1) - 2.42 % and the least (0.63 %) - on the same catalyst sulfidized by H2S/H2 (2b). In the first case mobile sulfiir equals 100 % of all catalyst sulfide sulfur and in the second - only 11 %. Productivity of Co/Si02 catalyst active sites is also much dependent on a sulfidation procedure - under H2S/H2 sulfidation it is twice as high as under thiophene. It can be also noted here that this difference is common to other catalysts. In some cases it is noticeably large while in others not (for instance, for Ni/Si02 it is not higher than 10 %). We may suppose that under full sulfidation the formation of a substantial amount of immobile sulfur encourages active site formation that have more mobile SH groups as compared to a phase where all sulfiir is mobile.If compare Co/Si02 and C0/AI2O3 catalysts, sulfidized in the same way, one can see that productivities (P) of active sites of these are close. It can be caused by an insignificant effect of carrier nature on reactivity of active sites and by an essential effect of a carrier on the number of the sites.
456 An effect of active phase metal nature on catalyst active site productivity is clearly seen in the sequence Co - (Co+Ni) - Ni. In all the catalysts supported by one and the same carrier and sulfidized in the same way the productivity grows in the sequence Co < (Co+Ni) < Ni. Thus, SH groups of Ni-sulfide particles demonstrate higher reactivity in H2S formation under thiophene HDS conditions than analogous SH groups connected with Co. A possible explanation of these dependencies might be searched within Bond Energy Model, recently developed by Tops0e et al [37]. The results of radioisotope testing permit us to single out three catalysts with the best characteristics - Co/Si02 (sample 2a), (Co+Ni)/Al203 (5b) and Ni/Al203 (3b). Sample 2a is characterized by the highest amount of mobile sulfur, though its active site productivity is the lowest and it demonstrates no high activity in thiophene HDS. Sample 5b demonstrates the highest thiophene conversion, its mobile sulfur amoimt is actually the same as in sample 2a, and P is considerably higher than in most of catalysts under study. Finally, sample 3b has the highest P, highest thiophene conversion and relatively small amount of mobile sulftu*. To evaluate the effect of the given parameters on the catalyst fimctioning under real conditions we have compared them with the data obtained in the course of hydroconversion of residual oil. 3.2. Hydrotreating of residual oil In the coiu*se of the experiments it has been found that, the same as in the above described experiments, most effective have proved catalysts containing 7.5 % metal. Catalysts with 3.4 % metal loading were ineffective. Catalytic activity of samples with metal loading 10 % and higher did not exceed that of the catalysts with 7.5 % metal content. The data about hydrotreating of residual oil is given in Table 3. First it should be stressed that all the catalysts show a rather high degree of selectivity towards residual oil containing metals (Ni, V) that are poisons for cracking catalysts. The catalysts demonstrate different degree of activity towards the desulfurization reaction, depending on the nature of an active metal and support (Fig. 4). The most active is Co/Si02 in reduced form (sample 2a, Table 3) - residua desulfiuization degree is 55.7 %. Somewhat less active are reduced catalysts C0/AI2O3 (la) and (Co+Ni)/Si02 (6a) - 26 % and 27 % correspondingly. Ni reduced catalysts on alumina (3 a) and silica (4a) show much less desulfurization activity - 18.0 % and 16.5 % correspondingly. Reduced Co/Si02 shows the highest activity in relation to deasphaltization and the lowest towards coke formation. In this catalyst on the silica surface metallic Co of fine dispersion with crystallite size 25 A has been found. No formation of surface compoxmds with support is observed. The catalyst has a high specific surface 307 m^/g pore volume 0.78 cm^/g, pore average D 46 A. Co reduction degree is 97 %. If compare pore structure of the support and catalyst one can see that Co loading does not much change the value of specific surface and pore size. It is possible that in the course of the reaction with sulfur containing crudes Co-sulfide particles formation takes place and they are active in hydrogenation (HYD) reactions. Catalyzing residua asphaltenes HYD they practically completely adsorb metals (Ni and V), which leads to a high degree of residua demetallization. Co sulfides as well as pure Co on silica do not
457 catalyze HYD process of hydrocarbons, which is proved by a low coke deposition on these catalysts. Tables. Results of hydrotreating residual oil of West-Sibirian petroleum (initial residual oil contained: asphaltenes - 4.3 %, S - 2.3 %, V - 40g/ton, Ni - 20 g/ton). No smp
Catalyst composition
Pretreatment*
Desulfiir. degree, %wt.
Demetall degree, %wt.
Residual asphaltenes %wt.
Coke deposits %wt.
la
(7.0%Co)/Al2O3
R
26.0
80.0
2.1
2.8
lb
(7.0%Co)/Al2O3
S
19.4
80.0
2.0
3.5
2a
(7.4%Co)/Si02
R
55.7
95.6
0.5
2.0
2b
(7.4%Co)/Si02
S
13.6
89.0
1.3
3.5
3a
(7.6%Ni)/Al203
R
18.0
82.0
1.2
5.0
3b
(7.6%Ni)/Al203
S
30.8
80.0
2.0
6.8
4a
(7.3%Ni)/Si02
R
16.5
90.5
1.0
3.0
4b
(7.3%Ni)/Si02
S
28.0
78.0
1.8
3.4
5a
(3.4%Co+4. l%Ni)/Al203
R
21.5
76.0
2.7
3.2
5b
(3.4%Co+4. l%Ni)/Al203
S
46.0
61.0
3.8
4.6
6a
(3.5%Co+3.5%Ni)/Si02
R
27.0
79.0
2.4
2.4
6b
(3.5%Co+3.5%Ni)/Si02
S
20.5
76.0
2.6
2.8
7
(Ni+Mo)/Al203**
S
32.9
75.5
2.8
10.5
*) R - reduction; S - sulfidation after reduction;
**) commercial catalyst
C0/AI2O3 in its reduced form (la) has a sharp difference from Co/Si02: desulfiuization degree is twice as low and asphaltene disintegration degree is about 50 %. Demetallization degree is lower too. In all probability, incorporation of some part of Co particles in the AI2O3 carrier material decreases the ability of C0/AI2O3 to interact with the sulfiir of residua heterocompoimds. Ni/Si02 reduced catalyst (4a) differs with Co one on its catalytic action. Its desulfiuization degree is more than three times lower but it is active in asphaltene HYD, demetallization and coke formation. These data testifies to the fact that metallic Ni on silica weakly interacts with the sulfiir of sulfiu"organic compounds from residual oil and more actively in HYD and deHYD reactions. Reduced Ni/Al203 (3 a) is more active than Ni/Si02 in reactions of C-C bond break and coke formation and less active in asphaltene HYD.
458 Desulfurlzation degree, % wt
C0/AI203
Co/Si02
Ni/AI203
NI/SI02
CoNl/AI203
CoNi/SI02
Figure 4. Desulfiirization degree of residual oil v^. catalyst composition and pretreatment procedure. Left bars - reduced catalysts; right ones sulfidized catalysts. Coke deposits, % wt
C0/AI203
Co/Si02
Ni/AI203
Nl/SiO
C0NI/AI203
CoNi/Si02
Figure 5. Coke deposition on the catalyst vs. composition and pretreatment procedure. Left bars - reduced catalysts; right ones sulfidized catalysts. Sulfidized forms of Co and Ni catalysts are essentially different from reduced forms of this catalysts on their catalytic activity, which supposes other nature of their interaction with oil residua components. All the sulfide catalysts have proved to be less active in deasphaltization but more active in coke formation (Fig. 5). Sulfidized Co/Si02 (2b) shows the least degree of desulfiirization -13.6 % and highest degree of demetallization among the sulfide catalysts. Obviously, Co sulfide particles in 2a are not so active in desulfiirization than reduced Co particles of sample 2b. Sulfidized Ni/AlaOs catalyst (3b) shows rather low activity in oil residua desulfiirization, the same
459 as in demetallization and tends to coke formation. A high activity in oil residua desulfurization has been shown by sulfidized (Co+Ni)/Al203 catalyst (5b) - 46 % and at the same time its degree of demetallization is the lowest - 61 %. Its coke deposition is 4.6%. 3.3. Comparative study of radioisotope testing and residual oil hydrotreating data While comparing the data of radioisotope testing with the results of residual oil HTR one can see a correlation between residual oil desulfurization degree and the amount of mobile sulfttr on the catalyst [Fig. 6]. Previously we observed a linear Oil residua desulfurization, % wt
~G3 ] 13 2~ Amount of mobile sulfur, mg
Coke deposits, % wt.
1 2 3 ^ Active site productivity, X 100
Figure 6. Oil residua desulfiirization degree v^. amoimt of mobile sulftir on the catalysts under study. Figure 7. Coke deposition on the catalyst samples vs. active site productivity. dependency between thiophene conversion and the amount of mobile sulfur in CoMocatalysts [33, 34]. However, as it was found later [35], this dependency was true not in all cases but only for the catalysts active sites of which are characterized by close values of productivity. In Table 1 one can see no correlation between thiophene conversion and the amount of mobile sulfur due to the fact that active site Ps of these catalysts differ from each other as much as two - threefold. In case of residual oil, however, such correlation is seen. One can suppose that a large number of contained in crudes and subjected to desulfurization sulfurorganic compounds which possess different degrees of stability in destruction processes as if levels the values of P of different catalysts. Hence, it seems we could conclude that such value as productivity is not essential when we deal with processes based on real crudes. But actually it not so at all. It was rather surprising to establish another linear dependency - between productivity of H2S formation active sites and coke formation on the catalyst - a secondary process in relation to HDS and H2S formation (Fig. 7). To some extent this dependency can be explained by carrier acidity on coke formation and hydrodesulfurization. The interdependency between carrier acidity and coke formation is commonly known and the dependency between catalytic activity and carrier acidity is marked by Welters et al [38, 39]. Tables 1 and 3 show that the samples of analogous composition and
460 sulfidized in the same way but supported on different carriers differ on the amount of coke deposition and site productivity - for Al203-supported catalysts these values are higher than for their Si02-supported analogs. But this fact can only partially serve as an explanation to the marked dependency. For instance, by acidity one cannot explain higher tendency of Ni-containing catalysts towards coke formation. It seems probable that the higher site productivity the more hydrocarbon fragments are formed in the course of destruction of sulftirorganic molecules and the higher coke formation is. This supposition is confirmed when we compare the results of residual oil HTR on the above described catalysts non-containing Mo with the results of the same procedure on the commercial Ni-Mo/Al203 sulfidized catalyst (see Table 3, sample 7). This catalyst has a high desulfiirization ability, being inferior only to catalysts 2a and 5b, but it displays rather low demetallization activity and the strongest tendency towards coke formation. The results of radioisotope testing of Ni-Mo/Al203 were submitted and discussed in [35]. In this paper we present radioisotope testing data only for one, mentioned above, commercial catalyst (the last fines of Tables 1 and 2, sample 7) for comparison with data obtained for non-containing Mo catalysts. Sample 7 is characterized by the presence of two types of active sites, "rapid" and "slow", Ps of which are 20.1x10"^ and 5x10"^ correspondingly and the number of the 'rapid" sites is 25 % of the total number of the both types of the sites. The amount of mobile sulfiir is 41 % of the total sulfide sulfiu* in catalyst 7. On this catalyst imder the same experimental conditions as for the other catalysts thiophene HDS conversion is 59 % i.e. much higher than on the catalysts without Mo. We believe that on the Mo catalyst promoted by Ni or Co the "rapid" sites are related to Mo and the "slow" ones - to Ni or Co. It was shown [33 - 35] that the unpromoted M0/AI2O3 catalyst has one type of the sites of low P (« 5x10"^), low catalytic activity in the thiophene HDS (conversion « 12 %) and the share of mobile sulfiir is about 20 % of the total sulfiir amount on the catalyst. Introducing promoter (Ni or Co) increases P of active sites related to Mo, though some part of these are blocked by Ni(Co)-sulfide particles. The total of the active sites on the promoted catalyst grows as compared to the unpromoted one because of an increase in "slow" sites. Evidently, the role of the "rapid" sites in catalystfimctioningis dual: on the one hand, the sites increase catalytic activity, on the other - they encourage coke deposition of the catalyst. It is particularly important when processing heavy crudes. It is this dual role of "rapid" sites that we observe dxuing residual hydrotreating on sample 7: the total mobile sulfiir is high (actually is the same as in sample 5b), which leads to a high desulfiirization activity, and a high P of some part of active sites results in strong coke formation. Thus, there seems to be some ground to state that when selecting most efficient catalysts for residual oil preliminary hydrotreating one should consider as preferable the systems with great amount of mobile sulfiir. i.e. relatively high number of active sites and low productivity of these. The study discussed in this paper has permitted us to put forward two catalysts 7.4 % Co/Si02 in its reduced form (sample 2a) and (3.4 % Co + 4.1 % Ni)/Al203 in its sulfide form (sample 5b) as catalysts-adsorbents for the first stage of hydrotreating of
461 heavy crudes. An important advantage of these catalysts is their relative cheapness due to the absence of expensive Mo in them. LIST OF SYMBOLS Aj - preexponential multiplier; Ms - sulfiir atomic mass = 32; Mfh ' thiophene molecular mass = 84; Nxh = vp /Mxh - the number of thiophene mmol in the pulse; ISi^ = Sf /Ms - the number of mmol of mobile sulfiir of the given type; Pf - productivity of active site of given type durind 1 thiophene pulse; Sf - mobile sulfur amount on the sites of the given type z, [mg]; X - H2S amount formed in the course of thiophene HDS reaction, [cm^]; a - H2S molar radioactivity normalized to the value of initial one, [%]; Yl - thiophene conversion [% mol]; D - thiophene pulse volume = 1 ltd; p - thiophene density = 1.064; Xj - exponential power REFERENCES 1. J.M.J.G. Lipsh and G.C.A. Schuit, J. Catal., 15 (1969) 163, 174, 179. 2. S.C. Schuman andH. Shalit, Catal. Rev., 4 (1970) 245. 3. R.J.H. Voorhoeve and J.C.M. Stuiver, J.Catal., 23 (1971) 243. 4. G. Hagenbach, Ph. Courty andB. Delmon, J.Catal., 23 (1971) 295. 5. G. Hagenbach, Ph. Courty andB. Delmon, J.Catal., 31 (1973) 264. 6. G Hagenbach, P. Menguy andB. Delmon, Bull. Soc. Chim. Belg., 83 (1974) 1. 7. C.H. Amberg, Less-Common Met, 36 (1974) 339. 8. P. Grange and B. Delmon, Less-Common Met. 36 (1974) 353. 9. T. Ohtsuka, Catal. Rev., 16 (1977) 291. 10. F.E. Massoth, Adv. Catal., 27 (1978) 265. 11. P. Ratnasamy and S. Sivasanker, Catal. Rev. - Sci. Eng., 22 (1980) 401. 12. H. Tops0e, B.S. Clausen, R. Candia, C. Wivel and S. Morup, Bull. Soc. Chim. Belg., 90 (1981) 1189. 13. H. Tops0e, B.S. Clausen, R. Candia, C. Wivel and S. Morup, J. Catal, 68 (1981) 433, 453. 14. N.-Y. Tops0e andH. Tops0e, Bull. Soc. Chim. Belg., 90 (1981) 1311. 15. H. Tops0e, R. Candia, N.-Y. Tops0e and B.S. Clausen, Bull. Soc. Chim. Belg., 93 (1984) 783. 16. S. Kasztelan, H. Toulhoat, J. Grimblot and J.P. Bonnelle, Appl. Catal., 13 (1984) 127. 17. B. Delmon, Surface and Interface Anal., 9 (1986) 195.
462 18. S. Kasztelan, L. Jalowiecki, A. Wambeke, J. Grimblot and IP. Boimelle, Bull. Soc. Chim. Belg., 96 (1987) 1003. 19. N.-Y. Tops0e, H. Topsoe andF.E. Massoth, J.Catal., 119 (1989)119, 252. 20. L. Viver, P. D'Araujo, S. Kasztelan and G. Perot, J. Mol. Catal., 67 (1991) 267. 21. L. Viver, P. D'Araujo, S. Kasztelan and G. Perot, Bull. Soc. Chim. Belg., 100 (1991) 807. 22. U.S. Oskan, L Zhang, S. Ni andE. Moctezuma, J.Catal., 148 (1994) 181. 23. M. Zdrazil, Appl. Catal., 4 (1982) 107. 24. M.L. Vrinat, Appl. Catal., 6 (1983) 137. 25. R. Prins, V.H.J. de Beer and GA. Somorjai, Catal. Rev. - Sci. Eng., 31 (1989) 1. 26. B. Delmon, Catal. Lett., 22 (1993) 1. 27. B. Delmon, Bull. Soc. Chim. Belg., 104 (1995) 173 28. AN. Startsev, Catal. Rev. - Sci. Eng., 37 (1995) 353. 29. H. Tops0e, B.S. Clausen and F.E. Massoth, Hydrotreating Catalysis - Science and Technology, Edited by J.R. Anderson and M. Boudart, Vol. 11, ISBN 3-54060380-8 Springer-Verlag Berlin Heidelberg, 1996. 30. G V Isagulyants, AA. Greish and V.M. Kogan, Kinetics and Catalysis, Engl, tr., 28 (1987) 220. 31. GV. Isagulyants, A A Greish, V.M. Kogan, GM.V'unova. GV.Antoshin, Kinetics and Catalysis, Engl, tr., 28 (1987) 550. 32. G V Isagulyants, A A Greish and V.M. Kogan, Kinetics and Catalysis, Engl, tr., 28 (1987) 555.. 33. GV. Isagulyants, AA. Greish and V.M. Kogan, in: "Proc. 9-th Int.Congress on Catalysis", (Calgary, 1988), eds. M.J. Philips andM. Teman, 1988, Vol. 1, P. 35. 34. V.M. Kogan , A A Greish and G V Isagulyants, Catal. Letters, 6 (1990) 157. 35. V.M. Kogan, Nguen Thi Dung and VI. Yakerson, Bull. Soc. Chim. Belg. 104 (1995) 303. 36. V.M. Kogan , A.A. Greish and G V. Isagulyants, in "Proc. 2-th European Congress on Catalysis", Maastricht, 3-8 September 1995, P.23. 37. H. Tops0e, B.S. Clausen, N.-Y. Topsoe, J.K. Norskov, C.V Ovesen and C.J.H. Jacobsen, Bull. Soc. Chim. Belg., 104 (1995) 283. 38. W.J.J. Welters, G Vorbeck, H.W. Zandbergen, J.W. de Haan, V.H.J. de Beer and R A van Santen, J.Catal., 150 (1994) 155. 39. W.J.J. Welters, VH.J. de Beer and R A van Santen, Appl. Catal., 119 (1994) 253.
® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
463
Saturation of Aromatics in Diesel Fuels: The Catalytic Toxicities of Sulfur and Nitrogen Compounds P. Kokayeff and G.J. Antos UOP 25 £. Algonquin Road Des Plaines, Illinois 60017-5017 U.S.A.
Emission considerations have precipitated legislation limiting the concentration of aromatic components in diesel fiiels or setting a minimum value for the fuel's cetane number. The saturation of aromatics meets both requirements because aromatics have lower cetane numbers than the corresponding saturates. Saturation processes for diesel fuels have been based on the application of catalysts that contain one or more noble metals. The choice of metals is made to obtain the best activity in the presence of sulflir as well as to achieve target stability. In addition to sulfur, hydrotreated diesel fuels also contain organo-nitrogen compounds, which also attenuate the activity of saturation catalysts. Although the poisoning effects of sulfur have been well documented, those of nitrogen have received rather scant attention. Because diesel feedstocks for saturation processes will inevitably contain both sulfiir and nitrogen, an understanding of the catalytic toxicities of each poison, both separately and in combination, is essential for both process design and catalyst development. This paper presents the results of an investigation of the toxicities of sulfur and nitrogen compounds on a noble metal saturation catalyst. The relative toxicities of sulfur and nitrogen compounds have been determined and successfully modelled by the introducing an "adsorption" term into the rate equation. The presence of both types of poisons, that is, both sulfur and nitrogen, was found to lead to a much greater attenuation of saturation activity than predicted from the catalytic toxicities of the individual poisons. In addition to a suppression of saturation activity, the presence of both types of poisons has an impact on stability. 1. INTRODUCTION Although the poisoning effects of sulfiir on the activity of noble metal saturation catalysts have been well documented, the toxicity of nitrogen compounds has received considerably less attention. This lack of a fundamental knowledge of the effect of nitrogen compounds leads to misconceptions as to their catalyst toxicity. Because most hydrotreated diesel feedstocks suitable for processing over saturation catalysts for the reduction of aromatics are likely to contain both sulfur and nitrogen species, an investigation was undertaken to determine the relative toxicities of organic sulfur compounds (as exemplified by dibenzothiophene), organic nitrogen compounds (acridine and quinoline), and combinations of both types of these poisons acting in concert.
464 2, EXPERIMENTAL The experimental program was conducted in a bench-scale, downflow trickle-bed reactor. Liquid hydrocarbon was delivered to the reactor by a Milton Roy minipump, and hydrogen was metered by a Brooks mass flow controller. The combined hydrogen-hydrocarbon stream was routed to a tubular reactor located in a three-zone clam-shell furnace. Three thermocouples were located within a central thermowell and positioned at the top, middle, and bottom of the catalyst bed. The catalyst, a noble metal catalyst on a proprietary support, was placed into service following a start-up procedure. Properties of the feed used during this experimental program are: Table L Properties of Diesel Feedstock •Gravity: 36.9 API •Distillation: - ffiP: 216 T - 5 vol-%: 227 T - 50 vol-%: 264 T -95 vol-%: 386 T - Endpoint: 392 T •Sulfur, wppm: < 5 •Nitrogen, wppm: < 1 •Aromatics, wt-%: 20.3 This feedstock was used "as is" and also spiked to various concentrations of sulfur and nitrogen using dibenzothiophene, acridine, and quinoline. Process conditions were within the range typically used in processes for the saturation of aromatics in diesel fuels. Aromatics in the reactor effluent were determined by supercritical fluid chromatography. 3. RESULTS AND DISCUSSION The initial experiment was conducted with the deeply hydrotreated feedstock containing less than 5 wppm sulfur and less than 1 wppm nitrogen. The resuUs of the experiment were modelled successfully by a simple first-order rate law. Since aromatics saturation is a reversible process, and neglecting the reverse reaction tends to introduce some error into the model, but the error is not significant at the relatively low temperatures used for this study. Subsequent experiments, which were carried out with new charges of catalyst, used the same feed doped with 60 wppm sulfur as dibenzothiophene and 30 wppm nitrogen as acridine. The results of these experiments were modelled by introducing an adsorption term for the poison, sulfur or nitrogen, into the denominator of the rate equation. The results of the experiments with a contaminantfree feed as well as with the feed containing the poisons are presented in Figures 1 and 2.
465
Temperature, "C
Temperature, ^C
Figure 2. Calculated and Observed Conversion as afiinctionof TemperaturePresence of N or S
Figure 1. Calculated and Observed Aromatics Conversion as a Function of TemperatureClean Feed Case
In both cases the models fit the data quite well over the entire conversion range. The temperature dependence of the rate constant was described by an Arrhenius type of exponential dependency. The adsorption constants for the poisons also followed an exponential dependence; however, with a positive sign for the temperature coefficient reflecting the expected form of an adsorption equilibrium constant, that is, a positive value for both the enthalpy and entropy of adsorption. The adsorption constants for both poisons exhibit the expected decrease with increasing temperature, a phenomenon expected of adsorption processes (Figure 3). An inspection of Figures 2 and 3 reveals that at lower temperatures of operation, nitrogen appears to be the more-potent poison (lower conversion and higher value of the adsorption constant). The reverse is true at higher temperatures, where sulfur appears to be the more potent of the two poisons. This apparent dependence of the relative catalytic toxicities of the two poisons may in part be due to the different conversion rates of dibenzothiophene and acridine to HjS and NH3, respectively (Figure 4).
Clean Feed - no S or N
L ——— Calculated
^^itf^
^___-J
\-r = kCA/(1+KPx)
Calculated
r = kCA
l^^^'^^y^
• Dibenzothiophene '> Afvidine
r Temperature, °C
Figure 3. Adsorption Constant (Ka) Variations as a Function of Temperature
Temperature, "C
Figure 4. Product S and N Concentrations Fit with a First-Order Rate Law
466 Because HjS and NH3 are less toxic to the catalyst than dibenzothiophene and acridine, the conversion of the more-toxic forms of the poisons to the less-toxic ones as a function of temperature determines the relative toxicities and introduces a complication into the analysis. The conversion of dibenzothiophene and acridine to HjS and NH3 was modelled by a first-order rate law, (Figure 4). The rate constants determined for the conversion of each were then used to estimate the concentration of dibenzothiophene or acridine as a fiinction of bed depth for three temperatures. The results are presented in Figures 5 and 6.
§
i
55 50 45 40 35 30 25 20 15 10 5
'
W
1
N.
1
1
Increasing Temperature
1——r
40
-
I
50
I
1
I
1
60
% Bed Depth
Figure 5. Organic Sulfur (Dibenzothiophene) Concentration Remaining at Varying Bed-Depth
1
30
40
50
60
70
% Bed Depth
Figure 6. Organic Nitrogen (Acridine) Concentration Remaining at Varying Bed-Depth
An inspection of Figures 5 and 6 shows that even at the lowest temperature, dibenzothiophene has undergone substantial conversion to HjS while acridine is essentially unconverted. The bottom half of the catalyst bed is exposed to sulfur mostly in the form of H2S, with some dibenzothiophene remaining, and to nitrogen in the form of acridine, with little nitrogen in the form of NH3. Thus, at lower temperatures, dibenzothiophene manifests its toxicity primarily as HjS while acridine exerts its fiill toxicity as an organo-nitrogen poison. As temperature is increased to higher levels the conversion of acridine increases until the conversion to NH3 is nearly the same as the conversion of dibenzothiophene to HjS. So, at lower temperatures, acridine appears to be the more-toxic agent because dibenzothiophene undergoes substantial conversion to the less-toxic HjS. At higher temperatures, where the conversions of dibenzothiophene and acridine become more nearly equal, sulfur in the form of H2S is more toxic than nitrogen in the form of NH3. Because diesel feedstocks contain both sulfur and nitrogen, the next experiment was conducted with a feedstock containing 40 wppm sulfur as dibenzothiophene and 10 wppm nitrogen as acridine. The results are compared to the previous experiments done with feeds containing either dibenzothiophene (60 wppm) or acridine (30 wppm), as shown in Figure 7, and to the prediction generated by the model developed from the previously described experiments, which used each catalyst poison separately to isolate the toxicities of each poison (Figure 8).
467 • 60f^m $ (DBT) » 40ppm S/ lOf^m N • 30ppm N (Acr)
Pr«dlet«d
r = kCA/(1-hl^Cs-hKf^
I • 40ppm S/ lOppm N I I I I
Temperature, *C
Figure 7. Observed and Calculated Aromatics Conversion at varying Temperatures
Temperature, "C
Figure 8. Observed and Calculated Aromatics Conversion at varying Temperatures
An inspection of Figures 7 and 8 reveals that the presence of both sulfur and nitrogen in the feed results in a much greater attenuation of saturation activity than predicted from the toxicities of each poison acting by itself A possible explanation for this effect is the known inhibition of desulfurization by nitrogen compounds. If the desulfurization of dibenzothiophene to form the lesstoxic HjS is inhibited by the presence of acridine, then the catalyst is exposed to the more-toxic dibenzothiophene, and the saturation activity is attenuated to a greater degree than expected from a consideration of each poison acting independently. At higher temperatures, both dibenzothiophene and acridine are converted to a substantial degree, and the catalyst is exposed to the less-toxic HjS and NH3 The attenuation of saturation activity approaches that predicted from the toxic effects of each poison acting independently. The presence of both poisons, dibenzothiophene and acridine, affects not only the activity but also the stability of the catalyst. Figure 9 shows the stability of the catalyst, that is, the maintenance of saturation activity over time, in the presence of 300 wppm sulfur as dibenzothiophene, 150 wppm nitrogen as quinoline, and the combination of both poisons, 200 wpppm sulfur as dibenzothiophene and w50 ppm nitrogen as quinoline.
468
• 300 ppm S (DBT) « f 50 ppm N (Quin) • 200 ppm S (DBT)/SO ppm N (Quin)
iT
•
Time On-stream, Days
Figure 9. Aromatics Conversion Stability in the Presence of Sulfur or Nitrogen Although the presence of either poison, sulfiir (dibenzothiophene) or nitrogen (quinoline), even at high concentrations, does not cause a marked decline in stability, the presence of both simultaneously results in a decay in saturation activity with time-on-stream. 4. CONCLUSIONS The attenuation of the saturation activity of noble metal catalysts by sulfur or nitrogen compounds may be modelled successfully by the introduction of an adsorption term into the denominator of the rate equation. This simple procedure affords a goodfitto the data in spite of the fact that the organic sulfur or nitrogen species undergoes varying extents of conversion to H2S or NH3 depending on the operating temperature. When both organic sulfur and nitrogen compounds are simultaneously present in the feed, the simple model presented in this paper is inadequate for representing the attenuation of saturation activity. The presence of both sulfur and nitrogen compounds results in a much greater attenuation of saturation activity than predicted by the action of either catalyst poison by itself In addition to a marked effect on catalyst activity, the presence of both organic nitrogen and sulfur compounds in the feed causes a decrease in catalyst stability.
® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
469
Production of high octane gasoUne components by hydroprocessing of coalderived aromatic hydrocarbons B. Demirel* and W. H. Wiser Department of Chemical and Fuels Engineering, University of Utah, Salt Lake City, UT 84112 The main objective of this work"^ was to convert aromatic compounds, representative of coal-derived liquids, into high octane compounds low in aromatics. For this study, 1-methylnaphthalene was chosen as a model compound. Experiments were performed in single and twostage systems. The model compound was treated with various catalysts at different reaction conditions in the single-stage operation. The highest conversion to isoparaffins and substituted cyclohexanes and cyclopentanes, 40.1% was achieved with a NiW/Si02-Al203 catalyst at 325 °C and lOOOpsig in 10 h with a feed to catalyst ratio of 10 to 1. In the two-stage operation, 1methylnaphthalene was hydrogenated to methyldecalins at 325 °C with almost 100% conversion in the first stage, using a NiMo/Ti02-Al203 catalyst. In the second stage, the methyldecalins underwent hydrocracking reactions at different reaction conditions using various catalysts. Pd/REX exhibited the best result for the production of high octane gasoline components without aromatics at 300-325°C. 1. INTRODUCTION Coal is the most abundant fossil fuel resource available in the world. It was the dominant energy source until replaced by petroleum and natural gas a few decades ago. However, coal is a main source of electric power in the US and still supplies 25% of the world's energy. It is widely agreed that the availability of petroleum in international trade is likely to diminish in the future. Nuclear power, natural gas, petroleum, solar energy and new technologies will not be sufficient to meet the growing energy needs of the world, and coal is expected to return to its position of dominance. Coal is not only used for a direct source of heat or producing steam for generation of electricity and a raw material in the chemical industries but also for the production of a range of gaseous and liquid fuels. During World War 11, over 100,000 barrels per day of high octane fuels were produced from coal in Germany [1]. Concerns continue to mount relative to the use of large
* present address: University of Kentucky, Center for Applied Energy Research, 3572 Iron Works Pike, Lexington, KY 40511 USA "^ This work was supported by the US Department of Energy.
470 quantities (up to 40%) of aromatic compounds in gasoline. These concerns are based on demonstrated toxicity to humans during handling of these fuels prior to combustion, as well as environmental impacts resulting from combustion of these highly aromatic fuels. Present evidence suggests that toxic components of fuels are the aromatic hydrocarbons [2]. Today's gasoline is subject to more demands and expectations than ever before. Regulatory agencies around the world are increasingly exercising their powers to regulate motor gasoline characteristics and mandate the use of deposit-control additives. The main goals are to increase octane number without addition of aromatics and to reduce emissions. The quality of gasolines can vary quite widely depending on crude source, sulfur content and octane number. It is desired to improve octane number, which is a measure of resistance to preignition, by using organic additives. Octane number or antiknock quality of a gasoline primarily depends on the type of hydrocarbons present. Aromatics have the best antiknock quality and isoparaffins, naphthenes, olefins and n-paraffins follow in succession. Refinery processes for the upgrading of gasolines aim at the conversion of other hydrocarbons into aromatics and isoparaffins. The objective of this research is to produce transportation fuel low in aromatics but exhibiting a satisfactory octane rating, from coal-derived liquids. Since coal-derived liquids are high in aromatics, the work has first focussed on conversion of aromatics to other compounds, which are observed to exhibit high octane numbers. The studies were initiated using 1methylnaphthalene as a model compound representative of aromatic compounds found in coalderived liquids. Different hydrogenation, hydroprocessing and hydrotreatment catalysts, NiMo/Ti02-Al203, NiW/Si02-Al203, NiW/Ti02-Al203, NiW/Al203, Pt/REX and Pd/REX, were examined with this feed material. Different temperatures, ranging from 200°C to 450°C and different pressures, ranging from 600 psig to 1500 psig, were investigated. Experiments were conducted in two forms in a stirred batch reactor: single-stage operation and two-stage operation, the model compound was treated with different catalysts at different temperatures. In two-stage operation, 1-methylnaphthalene was first hydrogenated to methyldecalins and subsequently treated with different hydrocracking catalysts at different reaction conditions to obtain compounds having satisfactory octane numbers. 2. EXPERIMENTAL 2.1. Catalyst preparation NiMo/Ti02-Al203 catalyst prepared by incipient impregnation method contained 25 mmol of Ni and 77 mmol of Mo per 100 g of Ti02-Al203 (Engelhard). The support impregnated with an aqueous solution of ammonium molybdate ((NH4)6Mo7024.4H2), Alfa Product) followed by a solution of nickelous nitrate (Ni(Ni03)2.6H20, Fisher Scientific Co.). NiW/Ti02-Al203 and NiW/A^ Q were impregnated in a single step with an aqueous solution containing ammonium meta tungstate ((NH4)6H2Wi204o.3H20, Cerac Inc.) and nickelous nitrate. Resulting material contained 88.9 mmol of Ni and 165.6 mmol of W per 100 g of support. Each support was ground, sieved to -100 mesh and calcined at 540°C for 16 h. The impregnated sample was oven-dried at 120°C overnight and finally calcined at 540°C for 16 h. NiW/Si02-Al203(Harshaw Catalysts), Pt/REX and Pd/REX (proprietary) are commercial catalysts used for the experiments.
471 2.2. Presulfidation All catalysts except Pt/REX and Pd/REX were sulfided in a tubular reactor prior to use. Nitrogen with a flow rate of 60 ml/min purged the reactor to remove air. The catalyst was heated to 400°C under nitrogen flow and kept at this temperature and inert conditions for 1 h. The gas flow then changed to 10% (by wt) hydrogen sulfide in hydrogen. The reactor was maintained at 400 °C at the same flow rate for 2 h. Subsequently it was purged with nitrogen at the same conditions for 1 h to remove the residual amount of hydrogen sulfide and then cooed to ambient temperature. 2.3. Reaction studies Reaction experiments were conducted in a 300 cc batch reactor equipped with a magnetic drive stirrer, heating assembly, gas inlet, and gas collector. Weighed amount of 1methylnaphthalene (Aldrich Chemicals, 98%) and the sulfided catalyst in a feed to catalyst ratio of 10 to 1 were quickly duped into the reactor. The sealed reactor was repeatedly purged with nitrogen and subsequently with hydrogen to replace air, and pressurized with hydrogen and heated to a reaction temperature over about 20 minutes a slow stirring rate (80 rpm). Once the desired temperature was reached, the stirring speed was increased to 800 rpm and the reactor was remained at reaction temperature for 1-10 h. Sulfided fresh catalyst was added to the reaction mixture after 5 h for experiments longer than 5 h. After completion of the reaction, the reactor was cooled to ambient temperature and the products were analyzed. The autoclave was operated either as a sealed-off system, back filled hydrogen at room temperature, or under a constant external hydrogen pressure. Two principle conversion schemes were examined. 2.3.1. Single stage operation 1-Methylnaphthalene was fed into the autoclave with different hydroprocessing, hydrotreatment or hydrogenation catalysts. Reactions were studied under the following conditions: temperature, 325 °C; initial hydrogen pressure, 1000 psig; reaction time, 10 h; feed to catalyst ratio, 10 to 1; stirring rate, 800 rpm. 2.3.2. Two-stage operation This procedure comprised hydrogenation and hydrocracking stages. In the first stage, 1methylnaphthalene was hydrogenated to methyldecalins using a NiMo/Ti02-Al203 catalyst at 325 °C and 1000 psig hydrogen initial pressure. In the second stage, methyldecalins underwent cracking in hydrogen atmosphere at temperatures of 300 to 450 °C and hydrogen initial pressure of 700 psig. using Pt/REX and Pd/REX catalysts. 2.4. Product analysis Liquid products were identified by GC and GC/MS (Hewlett Packard 5890). DB-5 column (30 m x 0.25 mm i.d. x 1.0 /um, J&W Scientific) with a temperature program from 40 260°C. Gas products were analyzed by a Shimadzu GC-14A gas chromatograph, equipped with a flame ionization detector and Chromosorb 102 80/100 column (6'xl/8"x0.0085", Supelco). The column temperature ranged from 40 to 200°C. Scotty standard gases were used for calibration. Methyldecalins from the hydrogenation of 1 -methylnaphthalene were also identified by ^^C-NMR spectroscopy (VXR-400).
472 3. RESULTS AND DISCUSSION Reaction conditions were determined after screening studies with various hydroprocessing, hydrotreatment and hydrogenation catalysts. All conditions remained the same for all experiments changing only the catalyst. The total conversion was defined as Feed - Feed in product %Total Conversion = 100 x — Feed Liquids were weighed just after completion of the reaction, and gas amounts were calculated from gas chromatograms. Liquid product distribution was calculated based on converted material. The products subdivided into four groups. (a) Cycloalkanes (methyldecalins, cyclohexanes and cyclopentanes which are mostly alkyl substituted and 'other cycloalkanes' which include decalins, octahydroindenes, bicycloheptanes) (b) Alkanes {normal and branched) (c) Alkenes and cycloalkenes (d) Aromatics {Methyltetralins, and 'other aromatics' which include bigger than one ring aromatics). For the single stage operation, the reaction temperature was chosen as 325 °C, remaining the pressure and reaction time the same, since screening test reactions at high temperatures resulted in cracking to aromatics. Results are shown in Figure 1. Almost 100% conversion was achieved with each catalyst. The only difference observed was in the depth of ring hydrogenation with the catalysts used. NiMo/Ti02-Al203 catalyst exhibited the hydrogenation activity followed by NiW/Al203 and NiW/Ti02-Al203 catalysts. The aromatics observed with these catalysts were methyltetralins ranging from 0.8 to 28.6%. Pt/REX and Pd/REX exhibited very high cracking activity. NiW/Si02-Al203 catalyst presented quite different results from the other catalysts and showed both hydrogenation and cracking activities. The apparent difference lies in the support material, where Si02-Al203 gives the highest yield of monocyclics. In terms of molecular size or boiling range and octane number, substituted cyclopentanes and cyclohexanes are very desirable components for gasoline. In general, cycloalkanes exhibit higher octane numbers than most of the isoparaffins. The highest octane numbers are observed when the substituent is a methyl group, with the octane number dropping rapidly as the length of the substituent group increases beyond a single carbon atom. The octane number is slightly higher when a cyclic compound contains only one methyl group rather than two methyl groups, and when both methyl groups are attached to the same carbon atom than when the methyl groups are bound to separate carbon atoms [3]. The most effective catalyst for the production of the desirable compounds in a single stage is NiW/Ti02-Al203 catalyst (Harshaw Catalysts). With the conversions to liquids being essentially 100%, this catalyst yielded 38.5% of the 1-methylnaphthalene by weight of cyclopentanes and cyclohexanes in 10 h at 325°C besides forming methyldecalins (15.3%) and 'other' cycloalkanes (17.2%), which are mostly octahydroindenes, decalins and bicyclo[2.2.1]heptanes. Only 1,6% of branched alkanes were observed. Methyltetralins, one ring aromatics and 'other' aromatics are 1.2%, 12.2% and 1.6%, respectively.
473 1-MethyInaphthalene, 3 2 5 ^ ; , 1000 psig, 10 h D Methyldecalins H Cyclohexanes + Cyclopentanes 11 Other Cycloalkanes n Normal Alkanes D Branched Alkanes HI Alkenes + Cycloalkenes M Methyltetralins • Other Aromatics
NiMo/Ti02-Al203 NiW/Ti02-Al203 NiW/Al203 NiW/Si02-Al203 Pt/REX Pd/REX _L 20
J40
-1. 60
80
100
wt% Of liquid
Figure 1. Yield and distribution of liquid products from hydrogenation/hydrocracking of 1methylnaphthalene using different hydroprocessing catalysts. Difference to 100% is the feed. For the two-stage operation, the data clearly indicate that methyldecalins are the intermediate compounds in the formation of either isoparaffins or cycloalkanes from methylnaphthalenes. Therefore methyldecalins are preferred feed for the second stage processing. In the first stage of the two stage operation, 1-methylnaphthalene was hydrogenated to methyldecalins with the catalyst, NiMo/Ti02-Al203, shown the highest hydrogenation activity at single stage experiments given in Figure 1. Figure 2 summarizes the results under different reaction conditions, keeping only the temperature constant at 325 °C. A total conversion of 100% was achieved and the yields of methyldecalins were over 92% for each experiment. The yield of methyldecalins only changed from 97.2% to 98.6% as the pressure was increased from 1000 psig to 1500 psig. When only the feed to the ratio was changed from 10/1 to 5/1, 99.5% conversion to methyldecalins was accomplished. In each case, the amount of cyclohexanes and cyclopentanes decreased with an increase in pressure and a decrease in the feed to catalyst ratio due to cracking to small hydrocarbons. Very small proportions of methyltetralins were observed. In the second stage of the two stage operation, methyldecalins from hydrogenation of 1methylnaphthalene were used as feeds with Pt/REX and Pd/REX catalysts after several catalyst screening experiments. Figure 3 shows the yield and distribution of liquid from hydrocracking of methyldecalins as a function of temperature, using Pt/REX. Figure 4 also shows the results with Pd/REX catalyst under the same conditions. As apparent from Figures 3 and 4, these catalysts show almost identical performance and the following discussion is valid for both metals. For both catalysts, total conversion increased as the temperature increased from 300°C to 450°C. The amounts of gas products also increased as a result of cracking reactions at higher temperatures. The yields of cyclohexanes and cylopentanes reached a maximum value (55.2% with Pt/REX and 68.8% with Pd/REX) at 325 °C and started to decrease above this temperature. The decrease of 'other' cycloalkanes and cycloalkenes indicates that they underwent cracking at
474 1-Methylnaphthalene, NiMo/Ti02-Al203, 325°C D Methyldecalins U Cyclohexanes + Cyclopentanes M Other Cycloalkanes M Aikanes M Alkenes + Cycloalkenes SI Methyltetralins • Other Aromatics
1000 psig.1 Oh, 10/1
1500psig, 10h, 10/1
lOOOpsig, 10h.5/1
]
1500psig, 8h.5/1
1500psig, 8h,4/1 0
J
I
20
40 60 wt % of liquid
L--J
\
^
I
80
^
1 L
100
Figure 2. Effects of different variable combinations on the yield and distribution of liquid products from the hydrogenation of 1-methylnaphthalene using NiMo/Ti02-Al203. Difference to 100% is the feed.
Methyldecalins, Pt/REX, 700 psig, 1 h 300 325 350
400
450 40 60 wt % of liquid
Figure 3. Yield and distribution of liquid products from hydrocracking of methyldecalins as a function of temperature using Pt/REX catalyst. Difference to 100% is the feed.
475 Methyldecalins, Pd/REX, 700 psig, 1 h
o
o
2 3
s. E
|2
40
60 wt % of liquid
80
Figure 4. Yield and distribution of liquid products from hydrocracking of methyldecalins as a function of temperature using Pd/REX catalyst. Difference to 100% is the feed. higher temperatures faster than cyclohexanes and cyclopentanes, and their yields were less than 1% at 400-450°C. The amount total alkanes, particularly branched, increased with temperature for both catalysts, but the increase in aromatics is a prominent result of higher reaction temperature. This can be understood from the increase in cyclohexane dehydrogenation and intermediate methyltetralin hydrocracking rates with an increase in temperature. The above results could form the basis for the effective production of desirable components for the reformulated gasoline from a feed rich in two-ring and possibly multiring aromatics. It was observed that the gaseous fractions from the hydrocracking of 1 -methylnaphthalene and methyldecalins are predominantly composed of butanes and pentanes, depending on the type of catalyst and the reaction temperature. The C4-C5 hydrocarbons may be alkylated or polymerized to form isoparaffins in the gasoline boiling or used directly as fuel gas [4]. In one stage operation, the highest yields of methane were obtained with NiMo/Ti02AI2O3, NiW/Ti02-A^0^ and NiW/AJ Q , because these catalysts have higher hydrogenation activity compared to NiW/Si02-Al203, Pt/REX and Pd/REX cracking catalysts as discussed in connection with Figure 1. The highest amount of propane was achieved with Pt/REX and Pd/REX, and isobutane was the highest with NiW/Si02-Al203. In the first step of the two-stage operation, the high yield of methane was reached because of the strong hydrogenation activity of the NiMo/Ti02-Al203 catalyst. In the second step, Pt/REX and Pd/REX showed the highest amounts of isobutanes at 300°C and 325°C. These amounts started to decrease when the temperature was increased. At the optimum condition for the production of cyclohexanes, cyclopentanes and isoparaffins, 325 °C and 700 psig, the amount of gases and the yields of isobutane were about 4% and 64%, respectively with these catalysts.
476 3.1. Potential for high octane gasoline without aromatics It was shown that 38.5% substituted cyclohexanes and cyclopentanes were produced by single stage operation at 325 °C over NiW/Si02-Al203 along with 1.6% of isoparaffins and 12.2% alkenes and cycloalkenes. For the two-stage operation, the highest yield of substituted cyclohexanes and cyclopentanes was also achieved at 325 °C with Pd/REX as catalyst and methyldecalins as feed. These results for both operations are summarized in Table 1. Alkenes and cycloalkenes in the product from the single stage operation can be hydrogenated and converted to isoparaffins, cyclohexanes and cyclopentanes. With recycling of methyldecalins from the first stage to the second stage, these intermediates can be converted in the presence of Pd/REX as catalyst to high octane substituted cyclohexanes and cyclopentanes. In this case, 6.4% cycloalkanes, 0.9% isoparaffins and 0.5% hydrogenated alkenes and cycloalkenes will be formed. For the potential production of high octane gasoline, the yield will reach 60.1% of methylnaphthalene converted to high octane compounds. The results are summarized in Table 2.
Table 1. Summary of the results from single and two-stage conversion of 1-methylnaphthalene over NiW/Si02-Al203 and Pd/REX catalysts.
Liquid Compounds
Total Conversion, wt % Cycloalkanes Methyldecalins Cycloalkanes and Cyclopentanes Others Alkanes Normal Branched Alkenes and Cycloalkenes Aromatics Methyltetralins Others
Product,
wt%
One stage Operation
Tw(HStage Operation
1 -Methylnaphthelene NiW/Si02-Al203 325°C, lOOOpsig, lOh
Methyldecalins Pd/REX 325°C, 700 psig, 1 h
99.9
75.9
15.3 38.5 17.2
68.4 10.7
0.2 1.6 12.2
1.3 9.6 7.6
1.2 13.8
0.4 2.0
477 Table 2. Potential for the production of high octane gasoline components from 1-methylnaphthalene.
From First Stage;
wt%
(T = 325°C, NiW/Si02-Al203) Substituted Cycloalkanes and Cyclopentanes
38.5
IsoparafFms
1.6
Hydrogenated alkenes and cycloalkenes
12.2
From Second Stage; (T = 325°C, Pd/REX) Substituted Cycloalkanes and Cyclopentanes
6.4
IsoparafFms
0.9
Hydrogenated alkenes and cycloalkenes
0.5
Total
60.1
4. CONCLUSION The highest conversion to isoparaffins and substituted cyclohexanes and cyclopentanes, 40.1%, was achieved at 325°C with NiW/Si02-Al203 catalyst for the single stage operation. Almost 100% conversion to methyldecalins was accomplished at 325 °C by hydrogenation of 1-methylnaphthalene, using NiMo/Ti02-Al203 in the first stage of the two-stage operation. Pt/REX and Pd/REX catalysts are very promising catalysts for hydrocracking of methyldecalins without forming aromatics, and Pd/REX appeared the most effective catalyst for the production of reformulated gasoline in a two-stage operation. The total yield of isoparaffins and substituted cyclohexanes and cyclopentanes is 78% at the same temperature with a negligible amount of aromatics. Gas products from single stage operation range from mostly methane to propane, depending on reaction temperatures and type of catalyst. NiW/Si02-Al203 catalyst produced the highest yield of isobutane for the single stage operation whereas Pd/REX catalyst presented the highest yield of isobutane in the second stage of the two-stage operation. In the first step of the latter, gas products are very rich in methane as a result of the strong hydrogenation activity of the catalyst. Considering the yields of isobutanes and isopentanes for possible alkylation to gasolinesize isoparaffin molecules, and the yields of substituted cycloalkanes, the experimental results showed promise to reduce the aromatic content of coal-derived liquids to very low values, while yielding a gasoline product of octane number of 85 or higher. Such an octane number rating is
478 based on the substituted cycloalkanes and isoparaffins, both produced directly from aromatics, and with alkylation of the isobutanes and isopentanes to follow. A potential conversion of 60.1% appears possible for the production of high octane gasoline from 1-methylnaphthalene in a twostage operation. REFERENCES 1. Kleinpeter, J.H. in 'Coal Liquefaction Products Volume T, (Ed H.D. Schultz), John Wiley, New York, 1983, p.5 2. Barbir, F. and Veziroglu, T. N. in 'Clean Utilization of Coal', (Ed Y. Yuriim), Kluwer Academic, Dordrecht, 1992, p. 131 3. 'Knocking Characteristics of Pure Hydrocarbons', American Petroleum Institute Research Project 45, ASTM, Philadelphia, 1958 4. 'NESTE - From Oil to Plastics', (Ed K. Hastbacka), Neste OY, Porvoo, 1993
® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
479
Process Developments in Gasoil Hydrotreating R C Lawrence, D H McKinley and M A Wood Davy Process Technology Limited; 30, Eastbourne Terrace, London W2 6LE, United Kingdom Changing demand patterns and legislation increase the pressure upon hydrotreating capacities at many refineries. To meet these pressures, improvements have been and will be necessary not only in catalysts, but also in the hydrotreating process. On the basis of its hydrogenation experience, Davy Process Technology has developed and tested a number of concepts aimed at improving the effectiveness of the basic process - enabling economic deep desulphurisation and opening up the potential for an integrated HDS/HDA flowsheet using sulphur tolerant HDA catalysts. 1. DRIVING FORCE There are two essential elements to the demand on hydrotreating capacity for transportation fiiels and particularly diesel. Increased environmental awareness and changing demand patterns are both working in the same direction which is unfortunate for an industry suffering from generally tight margins as is the case in refining. Some of the economic pressures can be relieved through technology improvements. It is the aim of all responsible process developers to provide economic solutions to these issues. 1.1. Environmental Legislation The worldwide trend in diesel fijel specifications is to lower sulphur, nitrogen and aromatics, particularly poly-aromatics. In the wake of US legislation, the European Governments are generally committed to meeting clean air goals as indicated by the various studies and reports prepared in relation to the Auto Oil Programme and other initiatives. The investment costs associated with the potential resulting legislation are widely reported and represent a major problem for refiners at a time of tight margins. Such statistics are, of course highly dependant upon the measuresfinallyadopted, but it is essential that costs are maintained within a reasonable limit whilst recognising the environmental needs. Already in meeting the recent measures which came into effect in October 1996 (ie Diesel: 500 ppm S), it is estimated that the investment cost has been ECU 3-5 bn - the potential investment to meet the Auto Oil Programme measures are estimated at up to ECU 26 bn (diesel only). 1.2. Demand/Supply Patterns Compounding the problem is the impact of changes in the demand/supply patterns. Generally speaking, the demand barrel is becoming lighter, due in part to the general increase in transportadonfijelsconsumption, particularly jet and diesel. The increasing use of natural gas for power generation is driving fuel oil into the middle distillate pool via upgrading. Increasing need to blend Light Cycle Oil into the diesel pool also increases the pressure on hydrotreating capacity
480
as a means of reducing aromatics. With these challenges in mind, DPT initiated a programme of work to enhance the effectiveness of those catalysts already available for desulphurisation and to investigate the potential for a truly integrated HDS and de-aromatisation unit for diesel applications. 2. TECHNOLOGY BACKGROUND Hydrotreating of gasoils is a complex system with many reactions proceeding at the same time, consuming hydrogen, evolving heat and producing inhibitors. Fig 1 shows diagrammatically that in a typical desulphurisation of a gasoil down to moderate levels (say 2000 ppm), the overall reaction rate is effectively first order because the large quantity of relatively easy sulphur compounds mask the reaction kinetics of the smaller quantities of refractory compounds. To achieve deep desulphurisation, below about 500 ppm, these refractory compounds must be broken down, and it is their kinetics which now set the overall rate. Thus to meet even more stringent diesel sulphur specifications would require an extremely large catalyst 10000 volume even with the more active Easy catalysts now available. ppm S
Refractory
1000
100
Residence Time
Fig 1 - Sulphur Reactivity
The diagram highlights the fact that the bulk of the desulphurisation (approximately 95%) is carried out in the first half of the catalyst bed. The remaining half of the catalyst is severely inhibited by the significant quantities of H2S which is released by the reactions. In the second half of the catalyst bed, only a small amount of hydrogen is consumed and therefore the heat release is low.
In a conventional hydrotreater (Fig 2), recycle and make-up hydrogen and gasoil pass co-currently in down-flow over the catalyst. Hydrogen is consumed in the process, giving off heat and producing inhibitors, particularly H2S. The result is that at the reactor exit, where the most difficult sulphur compounds are trying to react, the conditions are at their worst for effective catalyst operation : H2S is at its maximum Hydrogen partial pressure is at its lowest Inerts have been concentrated The only positive parameter for the reactions at this point is the higher temperature - although this promotes cracking, as does the lower hydrogen partial pressure. Conventional hydrotreating flowsheets therefore have distinct drawbacks in the context of minimising capital cost. Our
481 developments are therefore aimed at removing or at least reducing the limitations and improving the operating conditions for the catalyst to break down the more refractory compounds.
•^oh
@Feed
Recycle Hydrogen^ Lights
c^ Yl
I'roduct Make-up Hydrogen
Fig 2 - Conventional Hydrotreater
3. SuperTreet Concepts The principles of the developments carried out by DPT are illustrated in Fig.3. The key element is a hot, high-pressure separation between two reaction zones. For convenience, this is shown as two separate reactors although it is possible to arrange the system using one vessel. The
Recycle Hydrogen
@Feed
Lights
A
XA L J ^ ^
Product
Make-up Hydrogen
Fig 3 - SuperTreet Concept
482 intermediate separation system facilitates removal of H2S upstream of the second stage. Conditions for reaction of the most difficult compounds are significantly improved since the HjS inhibition is reduced and hydrogen partial pressure is maximised. Note also that the recycle gas only passes through the first reactor, so that there is no pressure drop penalty due to the additional catalyst. 4. DEMONSTRATION FACILITY A purpose-built pilot plant, operating at near-commercial superficial liquid velocities has been operated by DPT for over 2000 hours to demonstrate the application of these developments. The unit is designed for maximum flexibility to enable DPT to simulate a wide range of hydrotreater flowsheets to enable comparison of performance and to provide kinetic data for computer modelling. The key features of the unit are listed below: Four reactors provide the opportunity for inter-bed sampling and analysis which enables "tracking" of the progress of reactions. The reactors can be arranged in sequence to reflect a single stage conventional hydrotreater, or in groups to allow multiple stage operation. Capability to operate in adiabatic or isothermal modes, enables data collection for development of kinetic model as well as commercial demonstrations. Full hydrotreating flowsheet includes gas and liquid recycle, acid gas removal and product stabilisation. Full distributed control and data collection system. 5. RESULTS SUMMARIES 5.1. Desulphurisation A variety of feeds supplied by refineries in the UK, Italy and Japan have been processed. A summary of feedstock characteristics is provided in Table 1 together with operating conditions and product sulphur. The tests were performed with a standard CoMox HDS catalyst under operating conditions selected to reflect "typical" refinery configurations chosen by feedstock suppliers and compared the perfomiance of a conventional HDSflowsheetwith the SnperTreet flowsheet. The effects of the normal variables such as temperature, hydrogen partial pressure and LHSV were tested and plots of some of the data obtained for LHSV and temperature effects are provided in Figs 4 and 5.
483
Table 1: Test Data and Results Summary (HDS) CASE Feedstock Data Type
IBP FBP Sidpliiir Aroiixitics
SRGO 2(M°C 355 T 1.65 %\t. 22.P/VOI.
SRGOUCO8(y20 180 T 350 °C 1.36 %\t. 27.7«/vol.
SRCXyiXX) 50/50
201T 356 T 0.87 %M 37.7 "/M)1.
Operating Conditions: Ptessiine Tenperatiire LHSV
GOR
585psig 340 T 1.4 h-i 200:1
585psig 330 T 1.4 Ir^ 200:1
585 psig 330 T 1.4 Ir^ 300:1
Product Siilpluir Content: Conventional SiificrTreet
450ppni 160ppni
610ppin 200ppni
280 ppm 170ppni
Inprovement:
+64%
+67%
+40%
Fig.4 provides an insight into the progress of HDS reactions and confirms the value of interbed sampling. The lower temperature effect shown in Fig. 5 can be used to good advantage in extending catalyst life. Fig 4 Typical sulphur profile through the reactor (LHSV Effect)
Fig 5: Temperature Effect
100000
"^ ~. _^ 50() ppiii
~~ — .^
- "^-'•*~-'=:::::-C ^~'"-'"*<~->.,__^ S — •Conventional
Conventional
"^^ ^ ^
""• J
^^''''*'**~«««,i._
SuperTreet^""^
-^—Ailvancetl SMpcrTreet
Average reactor temperature (°C)
484 A "spin-ofF' benefit of removing the hydrogen sulphide at an intermediate stage is a slightly enhanced reduction in poly-aromatics. This can be seen in Table 2 which summarises data for a 50/50 blend of straight run and light cycle oils as described above, using CoMox. This issue has been the subject of more detailed testing, described below. Table 2 Comparison of Aromatics (wt%) Component
Feed
Product Conventional
SuperTreet
Tri-Aromatics
9.4
3.1
2.3
Di-Aromatics
18.3
7.7
7.0
Mono-Aromatics
10.3
26.9
28.4
5.2. De-Aromatisation The SuperTreet process, being two-stage, is suited to a combined HDS/HDA treatment, with HDS catalyst in the first stage and HDA in the second. Interstage removal of H2S, as indicated above, has the effect of improving HDA activity. DPT has recently completed a series of tests aimed at demonstrating the potential to provide refiners with several options in revamping to meet lower aromatics and higher cetane specifications. The tests used CoMox catalyst in stage 1 and a sulphur tolerant HDA catalyst in stage 2. As a comparison, the performance of NiMo in stage 2 was also tested. Table 3 summarises the results. Table 3: Dearomatisation Test Results Summary Feedstock (to 2nd Stage):
Operating Conditions
Type
SRGO/LCO 70/30
Pressure
900 psig
IBP
180T
Temperature
320T
FBP
350T
LHSV
1.0
Sulphur
500ppmw
GOR
750
Mono-Aromatics
18.25wt%
%Dearomatisation:
Di-Aromatics
11.26wt%
Pt/Pd
53%
Tri-Aromatics
2.77wt%
NiMo
15%
Total Aromatics
32.78wt%
Product S ppmw:
<50
® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B, Delmon and P, Grange, editors
485
BENZENE HYDROGENATION OVER TRANSITION METAL CARBIDES Carlos Marquez-Alvarez/ John B. Claridge, Andrew P. E. York, Jeremy Sloan and Malcolm L H. Green The Catalysis Centre, Inorganic Chemistry Laboratory, University of Oxford, South Parks Rd., Oxford OX1 3QR, U.K. ' On leave from: Institute de Cat^lisis y Petroleoqufmica, CSIC, Campus Cantoblanco, 28049 Madrid, Spain
ABSTRACT The hydrogenation of benzene has been carried out at atmospheric pressure on high surface area transition metal carbides. 100% selectivity to cyclohexane was achieved, with benzene conversions ranging from 0 to 100%. Activities were comparable to that of a 5% Ru/AljOg catalyst, but deactivation due to carbon deposition was observed on the carbide catalysts. a-MOjC showed the highest initial activity and deactivation rate, while a-WjC, and mixed MoTa and Mo-Nb carbides had improved stability. INTRODUCTION In recent years, increased public awareness to environmental issues has forced the chemical industry to improve existing technology and search for more efficient processes, in order to comply with the strict environmental regulations being implemented. In the petroleum industry, controls on exhaust pipe emissions have resulted in the search for new ways of producing cleaner fuels for use in the internal combustion engine. This includes the need to remove aromatic compounds, which are considered to represent a significant health risk and are present at elevated levels in diesel fractions. It is known that current hydrotreating (hydrodesulphurisation, HDS, and hydrodenitrogenation, HDN) units are not able to achieve the required degrees of aromatics hydrogenation. However, little work has been carried out on the development of adequate aromatic hydrogenation catalysts, and the habitual systems are those used in HDS/HDN, i.e. noble metals and promoted Mo or W sulphides [1]. This work explores the potential application of group V and VI metal carbides as efficient and economically viable alternatives to noble metal catalysts for the hydrogenation of aromatics in industrial feedstocks. Transition metal carbides have been shown to be active for a number of reactions, including HDN/HDS [2,3], hydrocarbon isomerization [4-8], methane reforming [9], and CO hydrogenation and Fischer-Tropsch synthesis [10,11]. Further, electronic similarities have been observed between WC and Pt [12,13], while it has been reported that TaC is highly active for ethylene hydrogenation [14] and aluminasupported MOjC exhibits catalytic activity similar to that of Pt or Ru for benzene hydrogenation [15]. This, coupled
to the known sulphur resistance of MOjC [16,17], makes these systems promising candidates for the aromatic hydrogenation functions of hydrotreating catalysts. Several groups have developed methods of synthesizing high surface area carbides (> 150 m^ g'^). These include gas phase reactions of metal compounds [18,19], pyrolysis of organometallics [20], and the reaction of gaseous reagents (e.g. CH^, CO, etc.) with solid state metal compounds [21-23]. The catalysts used in this study were prepared by temperature programmed reduction of the relevant oxide precursor, under a flow of ethane/hydrogen mixture, and the resulting high surface area carbides were tested for the hydrogenation of benzene at atmospheric pressure. Benzene was chosen since: i) it represents an important proportion of the total aromatic content of gas oils and the majority of those present in hydrotreated oils; and ii) the hydrogenation of monoaromatics is thermodynamically less favourable than that of multi-ring aromatic compounds. The activity of the various transition metal carbides for the benzene hydrogenation reaction are discussed by comparison with a known active catalyst for this reaction, namely 5% Ru/Al203.
EXPERIMENTAL Catalyst Synthesis Carbides were prepared by temperature programmed reduction (TPR) of the corresponding oxides, with a flow of 10% v/v ethane in hydrogen (150 ml min'^), and at a heating rate of 1 K min'^ from room temperature to 900 K. The reaction was held for 3 h at 900 K, and then the gas flow changed to pure hydrogen and maintained at 900 K for a further 30 mins. so that any surface carbon could be removed. Finally, the samples were cooled to the reaction temperature under
486 flowing argon. Tungsten carbide was treated at lower temperature (623 K) under H^, to avoid formation of tungsten metal [24]. The oxide precursors were high purity commercial M0O3, and WO3 (Johnson-Matthey), and MOgTagO^, MOaNbjO,^ and \NgHb^O„. The mixed oxides were prepared by intimately mixing various combinations of M0O3, WO3, NbgOg or TaPs powders (JohnsonMatthey), and firing in a sealed silica ampoule at 1073 K for 7 days (Mo3Nb20,,) or 1173 K for 15 days (MoJajO^), as described by Ekstrom [25]. In the case of WgNbgO^y, the oxide powders (Johnson-Matthey) were crushed together, pelletized, and placed in a platinum crucible at 973 K for 10 days, followed by 1473 K for a further 4 days, as described by Viccary and Tilley [26]. The composition of the mixed oxides was verified by XRD. For comparison, a 5% Ru supported on alumina catalyst was prepared by impregnating yAl203 (Akzo; SBET = 190 m^ g') with an aqueous solution of RuCig. The solid was dried and calcined at 773 K for 22 h. Before use, the catalyst was treated under a flow of pure Hj at 473 K for 1 h. Catalyst Testing Benzene hydrogenation was carried out at atmospheric pressure and 373 K in the same reactor used for the preparation of carbides, avoiding any contact of the catalyst with air. The reaction mixture (10% v/v benzene in hydrogen) was prepared by flowing 28 pmoi H2 s' through a saturator containing benzene at 293 K. A GHSV of approximately 3.5 x 10" h' was used. The effluents from the reactor were analyzed using a PerkinElmer 8600 gas chromatograph, fitted with a FID detector and a 25 m /0.32 mm diameter Chrompack PLOT CPAI2O3/KCI capillary column. After reaction, the carbides were passivated by allowing air to diffuse into the reactor at room temperature overnight. Characterization of the post-reaction catalysts was then possible. Catalyst Characterisation X-ray diffraction (XRD) patterns of the oxide and carbide powders, in the 3° to 70° 29 range, were acquired with a Philips PW 1729 diffractometer using Cu Ka radiation at 40 kV and 30 mA. Samples were mounted on an aluminium plate. The diffraction peaks were fitted to Lorentzian curves and the average particle size calculated from the broadening of the main diffraction peak using the Scherrer equation. The aluminium peaks were used as reference for the experimental broadening. High resolution transmission electron microscopy (HRTEM) was carried out using a JEOL 4000FX TEMSCAN electron microscope, with an accelerating voltage of 400 kV. Samples were ground to a fine powder and dispersed in AR grade chloroform. The dispersed sample was then placed in an ultra-sonic bath for 10 mins., before a drop of the suspension was placed on a copper grid, and the grid positioned in the microscope specimen holder. N„ BET isotherms were determined using an allglass high vacuum line, and the catalyst surface areas calculated.
RESULTS AND DISCUSSION Catalyst Synthesis and Characterization The XRD patterns of the starting metal and mixed metal oxides all exhibited sharp diffraction lines, in accordance with the low surface area (< 5m^ g"^) of these materials. However, after reaction under the ethane/hydrogen mixture, only very broad signals were present in the XRD. The d spacings agreed with the literature values for the formation of phase pure hexagonal close packed MOjC and WjC, and face centred cubic MoTaC,, MOgNbgC, and WgNbgC,; none of the starting oxide could be seen. The transformation was accompanied by an increase in surface area (see Table 1), as determined by BET. The average particle size, r, of the various carbides was also calculated from the XRD patterns and are shown in Table 1. For the binary systems, these values are in good agreement with the values calculated from the BET surface area (values shown in Figure 4), however, for the ternary systems the XRD values are smaller than those from BET, indicating another source of broadening in the temary systems; probably some form of disorder.
TABLE 1. Surface area and average particle size of the carbides used for benzene hydrogenation. Carbide (phase^)
Precursor
a-MOjC (hex)
M0O3
a-W^C (hex)
Sppr/m'g'
r^Rn* /nm
170
4.9
WO3
70
4.0
MoTaC, (cub)
Mojap,,
80
1.9
M03Nb2C, (cub)
MOgNbp,,
150
2.7
W . N b A (cub)
W,Nb«0,,
100
2.7
^ As identified from the XRD pattern. * From the [101] (hexagonal phases) or [220] (cubic phases) peak width in the XRD.
Figure 1 presents electron micrographs of WgNbgO^^ and the corresponding carbide synthesized under ethane/hydrogen. The ordered, large crystallites of the oxide have been broken up by the ethane/hydrogen treatment, leading to the formation of carbide particles made up of agglomerates of smaller particles. This can also be seen by the variable thickness of the sample, seen in Figure lb as a mottling of the particles. The change in structure noticeable in the micrographs can be explained by the simultaneous reduction of the oxide and substitution of gas-phase carbon into the lattice, as also proposed by Delporte et al. [27]. This blocks the collapse of the oxide to its lowest oxidation state form or total
487 reduction to the metal, and results in the formation of the carbide. Electron diffraction (not given) shows that the carbide crystallites are aligned [28].
Benzene Hvdrooenation The metal and mixed metal carbides were tested for the benzene hydrogenation reaction, and the results obtained are presented in Figure 2, along with those for the 5% RU/AI2O3 catalyst. The only product detected In all experiments was cyclohexane, except In the case of a-MOgC, which gave minute amounts of cyclohexene (<0.1% selectivity) at high benzene conversions (>90%).
40 H
^
30 •
e a
n
<
\
20 •
\ \
FIG. 1. Representative high resolution electron micrographs of: a) W9Nb8047; and b) WgNbgC, prepared by TPR in an C^HJH^ mixture (Scale bars are the same).
ooooo^^ 0 0 0 0 0 0 0
For molybdenum oxide, the relationship between the oxide and carbide formed from the reaction with ethane can be easily obtained since M0O3 exhibits a strong tendency to form plates with (010) perpendicular to the plate. Using electron diffraction (not given), it is possible to show that molybdenum carbide was formed with (2 1 10) perpendicular to the plate, i.e. parallel to the (010) of the original M0O3. For the other oxides, no such preferred orientation is observed, however, the measurement of many electron diffraction patterns on crystallites of WgNbgO^^ showed that, though not as pronounced as for M0O3, most oxide particles were oriented close to perpendicular to (001), and that most of the carbide particles were oriented perpendicular to (110).
1
2
3 4 Time (h)
5
6
7
FIG. 2. Activity (pmol of benzene reacted per second per gram of catalyst) for benzene hydrogenation at 373 K. (•) 5% Ru/AIA. (V) a-Mo^C, (D) a-W^C, (A) MoTaC,, (0) MOgNb^C,, and (•) WgNb^.
The most active of all the carbides was MOgC, which had an initial activity of the same order of magnitude as the ruthenium catalyst (at t = 0, =29 ^imol g ' s' compared with =45 |imol g ' s"'). However, the molybdenum catalyst deactivated rapidly. Since benzene is a well known carbon precursor [29,30], the large decrease in the MOgC catalyst activity can be attributed to the formation of carbonaceous residues. Indeed, HRTEM showed that large amounts of amorphous deposits, probably carbon, were present on the post-catalytic Mo^C
488
^-^l^i;
FIG. 3. Representative high resolution electron micrographs of post-catalysis samples: a) Mofi carbon; "M" = metal carbide).
and b) W j C . ("C"
sample (Figure 3a). In comparison, the rate of catalyst deactivation over the ruthenium catalyst was much slower
and, using
electron
microscopy,
no carbon
deposits could be seen on the ruthenium particles (not given), although there may be some carbon deposition on the acidic alumina support. W2C had a lower activity on a per gram basis, compared with MOjC, but when calculated per mole of catalyst the activity was only slightly lower.
In addition,
catalyst
than
deactivation
molybdenum
was much
carbide.
observable carbon
HRTEM
deposition
slower showed
over
that no
''BET = 3-7; rxRD = ~10
had occurred on the
tungsten carbide sample, although this does not rule out the presence of a thin layer of carbon slowly building up on the surface. The micrograph also showed that there '•BET = 5 . 4 ; rxRD = 5.8
was no significant change in the general morphology of the WgC catalyst after reaction (Figure 3b), as was the case for all the carbides. The initial activities of the ternary carbides were all lower than for the molybdenum sample. whereas
••BET = 6.2; rxRD = 4.9
However,
Mo^C deactivated to zero activity, all the
AI
ternaries maintained some activity at the end of the experiment, although
and the deactivation
the W^NbgC, sample
quickly.
Again,
using
rate
deactivated
HRTEM,
morphological changes were observed. on
the use of the ternary
carbides
no
tt>tnt^t^
relatively significant
Further studies for benzene
rBET=8-2: ''xRD = 3-''
JWWNM*,IW/
hydrogenation are currently being carried out. XRD
of the post-catalytic
^ j '•BET = 4.5; rxRD = 6.6
was slower,
carbide
samples,
vliw
" ^ — I — ^ 20
30
shown in Figure 4, indicates that no phase change has
40
I 50
60
70
29
occurred. For the ternary carbides, the average particle size, r, calculated using the Scherrer equation, is now in
FIG. 4. XRD patterns of the carbides after benzene
fairly good agreement with those calculated from the BET
hydrogenation: (a) a-Mo^C, (b) a-W^C, (c) MoTaC,. (d)
values, indicating that the materials are more ordered.
MOgMbjC,, and (e) WgNbgC,.
In the case of Mo^C, there is some
corresponding to the aluminium sample holder, r values
disagreement
between the two values, but this could be explained by
given are in nm.
Al labels show signals
489 non-spherical particle shape or the inherent errors in fitting Lorentzian peaks to a group of three XRD signals. In the diffractograms for MoTaC, and WgNbgC,, the small, sharp peaks at around 20-25 29 are probably due to the formation of metal oxide during removal of the samples from the reactor. This work has shown that group V and VI transition metal carbides, particularly those of molybdenum and tungsten, are almost as active as the highly active supported noble metals, such as Ru, Pt, Pd and Ni, which are generally accepted as the best catalysts currently available [1]. However, it has been shown previously that metal carbides, such as those tested here, have much greater resistance to sulphur and nitrogen poisons commonly present in industrial feeds than the noble metals. In addition, the carbides have been shown to be highly active for HDN and HDS reactions, while modified carbides are active and highly selective for hydrocarbon isomerization, meaning that the combination of two or more processes may be possible, and therefore that the use of carbides may be advantageous for hydrotreating processes [1]. The group V and VI transition metals offer further advantages, since they are cheaper, and available in greater abundance than the noble metals. The major problem with the metal carbides is their poor stability due to deactivation by carbon deposition. However, it may be possible to overcome this by: i) operation at higher pressure; ii) operation with higher hydrogen partial pressure; or iii) modification of the catalyst. This may be achieved by the use of a support or by the addition of dopants, such as those commonly employed in industry for noble metal catalysts, e.g. rhenium, iridium or tin, resulting in a possible increase in the stability of the catalysts to a number of poisons. Further work on improving the catalyst stability is currently being carried out [31].
CONCLUSIONS Various binary and ternary group V and VI transition metal carbides, synthesized by ethane TPR, have been tested for the benzene hydrogenation reaction. Molybdenum carbide exhibited an initial activity of the same order of magnitude as ruthenium supported on alumina, a well known active catalyst for this reaction, although catalyst deactivation occurred over the molybdenum catalyst. The other carbides had lower activities than molybdenum carbide, but deactivated more slowly. These carbide catalysts may provide advantages over the current aromatics hydrogenation catalysts based on supported noble metals, due to their ability to catalyse several of the important reactions
involved in hydrocarbon upgrading, and thus providing an opportunity to combine processes. ACKNOWLEDGEMENTS CM.A. wishes to thank the Spanish "MInlsterio de Educacion y Ciencia" for a Postdoctoral Fellowship. REFERENCES 1. A. Stanislaus and B. H. Cooper, Catal. Rev.-Sci. Engf., 36,75-123(1994). 2. J. S. Lee and M. Boudart, Appl. Catal., 19, 207-210 (1985). 3. J. C. Schlatter, S. T. Oyama, J. E. Metcalfe, ill and J. M. Lambert, Ind. Eng. Chem. Res., 27,1648-1653 (1988). 4. E. Iglesia, F. H. Ribeiro, M. Boudart and J. E. Baumgartner, Catal. Today, 15, 307-337 (1992). 5. E. Iglesia, F. H. Ribeiro, M. Boudart and J. E. Baumgartner, Catal. Today. 15, 455-458 (1992). 6. C. Pham-Huu, M. J. Ledoux and J. Guille, J. Catal., 143,249-261 (1993). 7. E. A. Blekkan, C. Pham-Huu, M. J. Ledoux and J. Guille, Ind Eng. Chem. Res., 33.1657-1664 (1994). 8. M. J. Ledoux. C. Pham-Huu, A. P. E. York, E. A. Blekkan, P. Delporte and P. Del Gallo, in "The Chemistry of the Transition Metal Carbides and Nitrides" (ed. S. T. Oyama), p.373-397, Blackie Academic & Professional, 1996. 9. A. P. E. York, J. B. Claridge, S. C. Tsang and M. L H. Green, J. Chem. Soc, Chem. Common., submitted (1996). 10. K. Y. Park, W. K. Seo and J. S. Lee, Catal. Lett., 11, 349-356(1991). 11. H. C. Woo, K. Y. Park, Y. G. Kim, l.-S. Nam, J. S. Chung and J. S. Lee, Appl. Catal., 75, 267-280 (1991). 12. L H. Bennett, J. R. Cuthill, A. J. McAlister, N. E. Erickson and R. E. Watson, Science, 184, 563-565 (1974). 13. R. J. Colton, J. J. Huang and R. W. Rabalais, Chem. Phys. Lett., 34, 337-339 (1975). 14. I. Kojima, E. Miyazaki, Y. Inoue and I. Yasumori, J. Cafa/., 73, 128-135(1982). 15. J. S. Lee, M. H. Yeom, K. Y. Park, l.-S. Nam., J. S. Chung, Y. G. Kim and S. H. Moon, J. Catal., 128, 126-136(1991). 16. S. Ramanathan and S. T. Oyama. J. Phys. Chem., 99,16365-16372(1995). 17. A. P. E. York, C. Pham-Huu. P. Del Gallo, E. A. Blekkan and M. J. Ledoux. Ind. Eng. Chem. Res., 35,672-682(1996). 18. M. J. Ledoux and C. Pham-Huu. Catal. Today, 15, 263-284(1992). 19. R. J. O'Brien, L Xu, X. X. Bi, P. C. Ekiund and B. H. Davis, in "The Chemistry of the Transition Metal Carbides and Nitrides" (ed. S. T. Oyama), p.362372, Blackie Academic & Professional, 1996. 20. J.-M. Giraudon, L. Leclerq, G. Leclerq, A. Lofberg and A. Frennet, J. Mater. ScL, 28. 2449-2454 (1993). 21. L. Voipe and M. Boudart. J. Solid State Chem., 59, 332-347(1985). 22. L. VoIpe and M. Boudart, J. Solid State Chem., 59, 348-356(1985).
490 23. L. Leclerq, M. Provost, H. Pastor, J. Grimblot, A. M. Hardy, L. Gengembre and G. Leclerq, J. Catal., 117, 371-383(1989). 24. G. Leclercq, M. Kamal, J. F. Lamonier, L Feigenbaum, P. Malfoy and L. Leclercq, Appl. Catal. A-Gen., 121, 169-190(1995). 25. T. Ekstrom, Acta Chem. Scand, 25, 2591-2595 (1971). 26. M. W. Viccary and R. J. D. Tilley, J. Solid State C/?em., 104, 131-148(1993). 27. P. Delporte, F. Meunier, C. Pham-Huu, P. Vennegues, M. J. Ledoux and J. Guille, Catal. Today, 23, 251-267(1995). 28. J. B. Claridge, D. Phil. Thesis, University of Oxford (1996). 29. J. Barbier, L. Elassal, N. S. Gnep, M. Guisnet, W. Molina, Y. R. Zhang, J. P. Bournonville and J. P. Franck, Bull. Soc. Chim. Fr., 1. 245-249 (1984). 30. J. Barbier, L. Elassal, N. S. Gnep, M. Guisnet, W. Molina, Y. R. Zhang, J. P. Bournonville and J. P. Franck, Bull. Soc. Chim. Fr., 1, 250-254 (1984). 31. J. Biswas, G. M. Bickle, P. G. Gray, D. D. Do and J. Barbier, Catal. Rev. -Sci. Eng., 30,161-247 (1988).
® 1997 Elsevier Science B. V. All rights reserved, Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
491
Hydrodesulfiirization of Dibenzothiophene in a micro trickle bed reactor
D. Letoumeur, M. Vrinat, R. Bacaud, Institut de Recherches sur la Catalyse, C.N.R.S., 2, Avenue Albert Einstein, 69626 Villeurbanne Cedex. France 1. ABSTRACT Hydrodesulfiirization (HDS) of dibenzothiophene (DBT) solutions has been studied in a micro trickle bed reactor using commercial supported Mo-based catalysts. A systematic study of the influence of the following parameters has been performed: flow of reactants, DBT concentration, catalyst granulometry, flow direction in the reactor, weight of catalyst, H2S partial pressure. It has been observed that, provided conversion is not determined by transport phenomena, the rate of conversion is first order with respect to DBT concentration. Hydrogen sulfide inhibits conversion. Flow direction is critical but does not modify the influence of hydrogen-to-liquid flow ratio upon conversion.
2. INTRODUCTION Catalytic hydrodesulfiirization (HDS) of refined products has been extensively studied and this process is the matter of a renewed interest arising from the more stringent products specifications refiners must deal with. A strict limitation of SOx emissions implies a drastic reduction of sulfur in fuels which can only be obtained through an improvement of present production processes. An increased severity (higher pressure or temperature) can effectively yield higher sulfur conversion levels. However, the range of variation of these parameters is limited, for cost considerations concerning hydrogen pressure, or for catalyst deactivation and hydrocracking of the feed when processing is performed at high temperature. The conventional supported Mo-based sulfide catalysts have been highly optimized, in terms of composition, additives, preparation, texture of supports, and any further improvement of the existing formulations requires a large amount of catalysts preparation, testing and screening. Since a given catalyst will have to process feeds of widely varying composition and properties, the determination of catalytic activity is a complex task. A catalyst which is highly active for the conversion of a given feed may result less active in the hydrotreatment of a different one. This situation arises for example, when the influence of
492 pore size distribution upon activity is considered : the impact of this parameter is strongly dependent from the reactivity of the feed [1]. For the purpose of a rapid comparative evaluation of catalytic activity, HDS of model compounds is generally preferred since it reflects the absolute activity of the considered solid in the desired reaction of carbon-sulfur bond cleavage. Large differences in reactivity of sulfur-containing substrates are well documented. Thiophene is more reactive than benzothiophene, which in turn is more easily converted than dibenzothiophene. The position of alkyl substituents affects the reactivity of DBT. These differences are probably not catalyst dependent. Therefore, catalysts ranking can indifferently be performed with anyone of these reactants, provided, experimental conditions are representative of real feeds processing. This point implies that, not only temperature and pressure range must be realistic (280-3 80°C and 2-6 MPa), but also that the distribution of reactants between the liquid and gas phases must reflect the conditions prevailing during the processing of a refinery feedstocks such like a light cycle oil or a gasoil [2]. Sulfur-containing compounds are in majority represented by benzo and dibenzothiophenic structures which exists predominantly as a dissolved phase in the considered range of pressure and temperature. The comparison of activity measurements performed in the gas phase and in the presence of various solvents indicates that a "solvent effect" must be invoked in order to explain large difference in conversion. However, the respective impact of aromatic and paraffinic solvents is not easily justified by mere considerations concerning a competitive adsorption [3,4]. The present paper reports the study of HDS of DBT, dissolved in dodecane. Activity measurements were performed in a micro reactor operated under biphasic flow conditions in the presence of commercial supported Mo-based catalysts. A systematic investigation of the impact of operating parameters upon conversion has been conducted in order to determine the conditions required for a significant measurement of activity and comparison of catalysts. The purpose is the development of a tool for a comprehensive evaluation of catalysts performance in HDS. 3. EXPERIMENTAL Activity measurements were performed with two alumina-supported catalysts prepared by impregnation of alumina pellets (250 m^.g"^ before impregnation) : a Ni-Mo (w/w% : Mo03=14, NiO=3) and a Co-Mo (Mo03=14, CoO=3). The calcined solids were grounded and sieved. Three fractions were considered : pellets, 80-125 |im and <80 |Lim. The reactant was a solution of DBT in dodecane containing 0.4 weight % of S. The gas phase consisted in a mixture of hydrogen sulfide in hydrogen, the composition of which could be varied between 0 to 5% H2S. The micro-reactor consisted in 2 cm^, up-flow tubular reactor which was filled with catalyst between two layers of inactivated alumina. Prior to HDS reaction, the catalysts were sulfided under 5% hydrogen sulfide in hydrogen for 11 h, at 400 °C, 0.1 Mpa total pressure. After that, the reactor was pressurized and a mixed flow of DBT solution and hydrogen was established. Liquid hourly space velocity could be varied in the range 2-40 h"^ The gas-toliquid flow ratio was between 50 and 300 cm^/cm^. Total pressure was limited to 5 Mpa. A gas-liquid separator allows a periodical sampling of liquid effluent. Conversion was defined
493 as percent of DBT converted into non sulfur products. Total sulfur in the effluent was determined by X-ray fluorescence. 4. RESULTS AND DISCUSSION If the number of parameters which may be involved in DBT conversion is considered, an experimental planning would be required in order to produce a model involving cross contributions of each parameter. Although this approach is justified when the description of a totally unknown system is considered, it is not pertinent in the present case since some relationships must be previously examined, such as the influence of reactant concentration. We therefore chose to follow a step-by-step process, keeping in mind that some crosscontributions could be ignored in this approach. 4.1. Influence of granulometry The influence of catalyst granulometry upon conversion is illustrated in Figure 1. The results are obtained in the presence of a constant mass of solid and constant flows of reactants.
^
o^
>
o U
< 80 ^m
80-125 |im
pellets
Figure 1. Influence of catalyst granulometry upon conversion Temperature = 300°C, weight of catalyst (NiMo/Al203) = 0.2 g, LHSV = 22.5 h" , gas-toliquid flow ratio = 200 cm'^/cm^, pressure = 3 MPa A clear difference is observed when the reactor is filled with powdered solid instead of pellets. A reduction of the mean particle size of the catalyst under 80 [im does not cause any additional increase in conversion. The 80-125 \xm fraction was utilized for all subsequent experiments. 4.2. Influence of residence time Using a constant weight of catalyst, constant flow and composition of gas phase, residence time was varied. The corresponding values are plotted in Figure 2. A linear relationship is observed up to 90 % conversion. Considering that DBT concentration undergoes large variations in the considered range of conversion, this result implies that the rate of DBT conversion is independent of DBT concentration, in other words, the reaction is
494 zero order vs DBT. Considering a plug flow reactor and a partial order, n, different from unity, the rate constant, k is expressed as :
k =
X
X
1 (I-T)' n-l
-I
in which, F is the flow of reactant, Q the initial concentration, m the weight of catalyst and T the conversion. Zero value for n yields a linear relationship of x vs 1/F, which is effectively observed. 100 -n 80 -
Experiment Theoritical curve with first order reaction vs DBT
60 40 C
o U
20 0 -<^
00
1
1
1
0.1
0.2
0.3
0.
1/liquid flow (h/g)
Figure 2. Influence of liquid flow upon conversion Temperature = 290°C, weight of catalyst (NiMo/Al203) = 0.4 g, gas flow rate = 1500 cm^/h, 1% H2S, pressure = 3 Mpa. HDS studies of DBT performed in the gas phase, generally report that this reaction is first order vs DBT [5]. Changing from gaseous to dissolved phase is not likely to modify completely the reaction scheme and to alter the apparent order of reaction rate [2]. A positive determination of the dependence of reaction rate vs DBT concentration has been performed, comparing the conversion of solutions containing variable amounts of DBT. The values of conversion collected in table I clearly indicate a first order relation. Table 1. Influence of DBT concentration upon conversion Temperature = 290°C, weight of catalyst (NiMo/Al203) = 0.4 g, gas flow rate = 1500 cmVh, liquid flow rate = 10.8 g/h, 1% H2S, pressure = 3 MPa. Weight % of Sulfur Conversion (%)
0.4 30
A first order reaction yields the following expression for the rate constant:
0.12 33
495 k = - (-ln(l-T) C„m A plot of the calculated value of x vs \/F is presented in Figure 2. Discrepancies between calculated and experimental values appear for higher residence times. A change in DBT concentration remaining in the solution might explain this behavior if a partial vaporization of the solvent is considered. Higher hydrogen flow rate favors dodecane vaporization, and as a consequence, induces an increment of DBT concentration. This effect, which would increase reaction rate if a first order is assumed, could justify the apparent contradiction between the observed first order reaction evidenced through a variation of DBT concentration and the linearity of the relation of conversion vs. residence time which reflects a zero order reaction rate. 4.3. Influence of flow conditions The effect of a variation in the total flow of gas phase is evidenced in Figure 3, where conversion is plotted against residence time for two different gas flow rate and a constant gas phase composition (1% H2S). In both cases, conversion varies linearly with residence time, but a decrease of gas flow rate produces a reduction of conversion.
Cocurrent upflow Cocurrent downflow Cocurrent upflow with mixer Cocurrent upflow Cocurrent downflow Cocurrent upflow with mixer
>
a o
0.0
0.1
0.2
0.3
H2flowrate = 1500 cm7h
H2 flow rate = 800 cm7h
0.4
1/liquid flow (h/g) Figure 3. Conversion as a fiiction of residence time Temperature = 290°C, weight of catalyst (NiMo/Al2G3) = 0.4 g, 1% H2S, pressure = 3 Mpa. This dependence of conversion with respect to hydrogen flow rate indicates that, in limited flow conditions, the reaction could be limited by hydrogen availability in the dissolved phase. The dissolution of hydrogen in the liquid phase may be influenced by mixing conditions in the reactor. Therefore, parameters affecting gas-liquid contact, such as the ratio of hydrogen-toliquid flow and reactor arrangement, might affect conversion. Two type of modifications were explored : an inversion of flow direction and the insertion of a mixing tank before the reactor.
496 The observed conversion is plotted as a function of residence time for both arrangements and for tw^o values of hydrogen flow rate in Figure 3. The presence of an additional gas-liquid contactor does not modify conversion. In contrast, switching flow direction from upflow to downflow improves conversion by about 30 % but does not affect the influence of hydrogen flow rate. Thus hydrogen dissolution is apparently not modified by a change in flow direction but the contact between the reactants and catalyst particles is affected. The variation of conversion as a function of catalyst weight is illustrated in Figure 4 for two values of hydrogen flow rate. In these experiments, performed at constant gas-to-liquid flow ratio, the composition of the liquid phase flowing in the reactor is constant. The resulting curves are conveniently described by a first order equation. The ratio of conversion values, respectively observed at 0.8 and 1.5 l.h"' hydrogen flow rate, is constant over the range of catalyst weight. This observation indicates that the influence of hydrogen flow rate upon conversion reflects the variation of DBT concentration in the liquid phase rather than transport limitation.
100 80 H c o
>
o U
60
•
Flow ofH2= 1500 cm-'/h
•
Flow ofH2 = 800 cm^/h
40 H 20 H 0.0
0.1
0.2
0.3
0.4
0.5
Weight of catalyst (g) Figure 4. Conversion as a fiunction of catalyst weight Temperature = 290°C, liquid flow rate = 5.4 g/h, 1% H2S, pressure = 3 MPa. 4.4. Influence of hydrogen sulfide During reaction, hydrogen sulfide is produced and its partial pressure is related with conversion. Since this product is reported as an inhibitor of HDS, a comparison of various catalytic formula may be obscured if HDS reaction is performed at different conversion levels. In order to get some insight into the relative impact of hydrogen sulfide partial pressure upon conversion, we measured its influence in the presence of two distinct catalysts. The corresponding results are presented in Figure 5. In the presence of low additional H2S partial pressure, NiMo catalyst is more active than CoMo. Increasing H2S concentration causes a considerable decrease in the observed difference.
497
NiMo/Al203
^
o^
•
o
C0M0/AI2O3
o U
% H2S Figure 5. Influence of H2S upon conversion for two catalysts (NiMo and C0M0/AI2O3) Temperature = 280°C, liquid flow rate = 5.4 g/h, H2flowrate = 1.2 1/h, pressure = 3 MPa. 5. CONCLUSION The aim of the present study was to explore the experimental basis for a pertinent comparison of catalysts performance in HDS of a model compound. DBT was chosen as a representative model and the conversion in the liquid phase has been measured. The influence of various experimental parameters upon observed conversion has been investigated. In contrast with reported kinetics data, which are generally concerned with gas phase conversion, the present results indicate that flow conditions in the reactor are of paramount importance. Gas-liquid contact determine hydrogen dissolution, which in turn affects conversion. Although we did not investigated the simultaneous effects and cross-contributions of porosity, total pressure and gas-to-liquid flow ratio, the observed tendencies indicate that these parameters must produce a variable impact upon conversion when distinct catalysts are compared. The influence of H2S partial pressure upon conversion provides an illustration of the changing impact of one parameter when two distinct catalysts are compared: a CoMo catalyst is less sensitive than NiMo to increasing partial pressure of H2S. Consequently, a single measurement of activity in fixed experimental conditions does not reflects the performance of a catalyst and may lead to an erroneous comparative ranking when a screening of various catalytic formulations is performed. Aknoldedgement The present work was carried whithin the framework of the program «HDS of Gasoils» supported by ELF, IFF, TOTAL and CNRS-Ecotech.
REFERENCES
1. B.M. Moyse, B.H. Cooper, A. Albjerg, Am-84-59, National Petroleum Refiners Association, (1984) Annual Meeting, Texas. 2. M.J. Ledoux, C.P. Huu, Y. Segura and F. Luck, J. Catal., 121 (1990) 70. 3. A. Ishihara, T. Itoh,T. Hino, M. Nomura, P. Qi and T. Kabe, J. Catal., 140 (1993) 184. 4. S.Y. Lee, J.D. Seeder, CM. Tsai and F.D. Massoth, Ind. Eng. Chem. Res., 30 (1991) 607. 5. M. Vrinat, Applied Catal, 6 (1983) 137.
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® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
Use of Dispersed Catalysts for Fossil Fuel Upgrading A. S. Hirschon and R. B. Wilson SRI International, 333 Ravenswood Ave, Menlo Park, CA 94025, USA 1.
ABSTRACT
Because of dwindling supplies of premium oil reserves, we are striving to find economical methods to utilize altemative fossil fuels, such as coal and heavy oils, which are far more abundant. Conventional hydrotreating catalysts are not adequate for these types of feedstocks because of the heavy asphaltenes and metal contaminates. One approach to convert these feedstocks is to use dispersed catalysts. The advantage of dispersed catalysts is that they should not be subject to deactivation as readily as conventional supported catalysts, while at the sametime,should provide more catalytic activity with a lower metal loading. We have tested the effect of several preparations of dispersed catalysts to convert feedstocks such as coal, recycle solvents, and heavy oils. These dispersed catalysts include those containing Fe, Mo, and Ni metals or metal sulfides. We found that the dispersed catalysts are more effective than those of conventional catalysts, with the nK)st effective ones being those that are easily activated under the reaction conditions. 2.
INTRODUCTION
Worldwide crude oil production is already close to capacity and is expected to decline due to diminishing reserves. Estimates of the decline range from 2015 to 2020.1*2 Yet estimates of demand for premium oil is expected to increase by over 30% over this time pericxi and the demand could exceed the supply before the year 2005. Due to these factors and frequent political uncertainties, the cost of oil is expected to be significantiy higher in the near future and the need to economically utilize altemative fossil fuels to usable feedstocks is essential. Vast amounts of altemative feedstocks exist world wide, but because of the difficulties in upgrading these feedstocks, they are not fully utilized. Coal, resids, and heavy oil reserves could account for decades of future energy usage. The problem with these materials is that they contain heavy amounts of asphaltenes and metals which rapidly deactivate conventional supported catalysts. The asphaltenes contain coke precursors which tend to accumulate in pores of conventional hydrotreating or liquefaction catalysts. Furthermore, metals, especially those associated with porphrins such as Ni and V, deposit as the sulfides in the pores of the catalyst support.^"^ Our approach to upgrade and utilize these altemative feedstocks is to use dispersed catalysts. The advantage of dispersed catalysts is that since they are not supported they are not subjected to deactivation by pore plugging, and since they are highly dispersed, there is an intimate contact between the oil and/or cod and solvent with the hydrogen atmosphere.
499
500 In the ideal case, they can be activated and transfer hydrogen directiy to the asphaltene molecules even in the absence of a good H-shuttiing solvent Thus a more efficient hydrogen utilization can be used, potentially lowering pressure and perhaps temperature requirements. Furthermore, with proper design, cat^ysts with high surface areas with less metal loadings will be required. One of the problems in developing dispersed catalysts is to be able to produce the catalyst in the most active state. The reason fortfiisrequirement is so that retrogressive reactions or coking will not occur while waiting for the catalyst to activate. In our work we try to design the dispersed catalyst so that it is already in the most active stoichiometry and oxidation state, and is therefore more easily activated. The following work describes our efforts to determine the most effective catalysts for the various altemative feedstocks. 3,
EXPERIMENTAL
The hydrotreating catalyst Shell 424 NiMo alumina supported catalysts was received as a gift from Criterion Catalyst Co. and was ground and sieved to 60-200 mesh. The catalyst was presulfided under flowing 10% H2S/H2 at 400°C prior to use. Ammonium tetrathiomolybdate (NH4)2(MoS4) was obtained from Alfa Chemicals and is referred to as M0S4. The organometallic molybdenum catalyst was (C5H5)2Mo2(^SH)2(^-S)2, referreid to as Mo(OM), and was prepared by modifying the method of Cowens et al.^ Coal conversions and oil upgrading experiments were conducted in 300 mL Autoclave Engineers stirred reactor. The coals tested were either Illinois #6, North Dakota Lignite, or Black Thunder (obtained from the Wilsonville coal liquefaction facility). The Black Thunder recycle solvent as well as the MolyVanL oil (used as a dispersed catalyst) were also obtained from Wilsonville. Maya ATB was used to test the effectiveness of the various catalysts for heavy oil upgrading. The feedstock or solvent containing the desired amount of catalyst was charged with 1200 psia H2 containing 3% H2S. The coal conversions in synthetic solvents were conducted using approximately 5 g of coal in 30 g of solvent. Fortiieconversions in recycle solvents, 20 g of coal (Black Thunder) and 50 g of solvent were used. In the case of the heavy oil, 50 g of the oil was used. The Mo loading for the coal and oil experiments was approximately 500 ppm. In the case of the NiMo supported catalyst conversion of the Maya, approximately 1.0 gram of catalyst was used. After the autoclave was charged it was heated to 400°C or 425°C for up to one hour. Microdistillations were conducted on Perkin-Elmer TGA apparatus. Mass spectra were obtained on a Field Ionization Mass Spectrometer (FIMS) built by SRI. The distribution of preasphaltenes, asphaltenes, and oils, were determined by solubilities in THF, toluene, and hexane, respectively. Coal and oil conversions were based on THF solubilities. 4.
RESULTS AND DISCUSSION
The purpose of our work was to examine dispersed catalysts for several types of feedstocks, each of which has the characteristic in that they are hard to upgrade using conventional catalysts. In this paper we compare conventional dispersed catalysts such as M0S4, MolyVanL, and a thiolato molybdenum catalyst precursor. These feedstocks have different properties in that they have varying degrees of H/C ratio, and varying amounts of aliphatic and aromatic components in the hydrocarbon material. Thus these feedstocks offer a wide spectrum of properties to which we can test our catalysts. The H/C ratio, ranging from 0.79 for the coals, 0.96 for the recycle solvent to 1.49 for the ATB oil is less than what would be desired for a premium petroleum feedstock. Furthermore, these materials
501 contain high degrees of asphaltenes, which easily forms coke. During the upgrading process these heavy asphaltenes tend to phase separate, making the coking process even more of a problem. Furthermore, these feedstocks have high degree of meml contents which tend to rapidly deactivate conventional supported catalysts. In order to examine the various dispersed catalysts we conducted a series of experiments designed to simplify the assessment and comparison of the catalysts. This step was necessary due to the complexity of fossil fuel feedstocks. For our fct approach to compare the dispersed catalysts, we examined coals in synthetic solvents such as tetralin, and n-hexadecane (CI6). This comparison allows us to determine the relative effects of a good H-donation solvent with a totally inert solvent in conjunction with the dispersed catalysts. The more inert solvent allows us to magnify the comparison of the dispersed catalysts. Furthermore, a catalyst which works well in conjunction with such a solvent would most likely be useful in cases where we wish to upgrade heavy oils. We tested two ranks of coals. A subbituminus coal with a relative high H/C ratio, and moderately easy to convert, and in an addition, a lignite coal which is difficult to convert, and contains many precursors to retrograde reactions such as free radical chemistry during the loss of carboxylates. Of the various feedstocks, coal, recycle solvent, and heavy oils, Ae coal is the most challenging due to the low H/C content and the large amounts of metal and ash which can readily deactivate conventional supported catalysts. The retrogressive reactions can occur even under moderate temperatures so dispersed catalysts tiiat can be rapidly activated during the heat-up period of the feedstock is essential. Table 1 compares the effects of the two molybdenum dispersed catalysts and a supported CoMo catalyst for the conversion of Illinois #6 coal in both tetralin and in the more inert solvent, C16. Noted in this figure is that in the tetralin solvent, even though more metal catalyst was used for the supported catalyst conversion, the conversion of the coal to liquid products was far worse witii the alumina support. In contrast, both the dispersed catalysts gave much better yields to soluble products, with the Mo(OM) catalyst giving the better results. Also note that for the thiolato complex the conversion is almost as good in the inert C16 solvent as is in the tetralin, suggesting that this dispersed catalyst should be useful in the conversions using recycle solvents, resids, or heavy oils. The fact that the supported catalyst behaved so poorly suggests that the ^umina support also promoted retrogressive reactions, perhaps centered on the acidic functionalities of the support. Table 2 shows the results where we substituted authentic heavy fossil solvents for the synthetic solvent. In the case of the coal conversion experiment, the ratio of solvent to coal was approximately 2.5 to 1 by weight. In each case the amount of Mo catalyst added was approximately 500 ppm. The coal conversion (Black Thunder Coal) was calculated by determining the THF soluble fraction, and allowing form the contributions of the heavy recycle solvent We found the hexane soluble fractions (oils) to be approximately the same in all cases (approximately 20-25%). The products (from the Maya ATB conversion) was taken up in hexane and filtered through a medium porosity filter. Again, the hexane fractions or oils were approximately the same in all cases (12%). The hexane insoluble portion was then taken up and filtered in THF. We compare the Maya conversions in terms of THF solubility. As seen in both cases, the coal conversion and Maya conversion, the thiolato catalysts gave superior results. In the case of the coal conversion, the value increased from 45% using the MolyVanL catalyst to 99% for the Mo(OM) catalyst. Likewise in the case of the heavy oil upgrading, the final product using the Mo(OM) was essentially totally THF soluble, indicating very littie coking, if any, had occurred. In the
502 case of the MolyVanL and NiMo catalyzed runs material, the THF fraction could only be filtered difficulty, most likely do the high molecular weight species formed after reaction. TGA distillation curves of the oils were quite similar, however the Mo(OM) treated product showed a higher proportion of lower boiling point materials. FIMS spectral data of the two fractions likewise showed an improvement for the Mo(OM) product, with a higher fraction of low molecular weight products in the range of 300 mass units. The number average molecular weights of the various products were fairly similar with values of 679,645, and 725 for the Mo(OM), MolyVanL and NiMo catalyzed conversions, respectively.
Table 1 Effect of catalyst for conversion of coal to toluene-soluble products in synthetic solvents^^ Catalyst
Coal
Solvent
None
Illinois # 6
Tetralin
M0S4
Illinois #6
Tetralin
M0/AI2O2
Illinois #6
Tetralin
Mo(OM)
Illinois #6
Tetralin
None
Illinois #6
Hexadecane
M0S4
Illinois #6
Hexadecane
Mo(OM)
Illinois #6
Hexadecane
None
Lignite
Hexadecane
Mo(OM)
Lignite
Hexadecane
T(°Q
400 400 400 400 400 400 400 425 425
%"TS—
48 51 41 61 25 41 61 24 49
^Reaction conducted in a 300-mL autoclave with 5 g coal, 3 mmol catalyst, 30 g solvent and 500 psi H2 for 20 min at temperature. 'fields calculated on daf basis for Illinois #6 coal and on cartx)n basis for the lignite.
Table 2 Effect of catalyst for conversion of coal or oil in petroleum products^ Catalyst MolyVanL Mo(OM)
Feedstock Black Thunder Coal Black Thunder Coal
% Conversion 45 99
^Reaction conducted in a 300-mL reactor with 500 ppm metal In catalyst, 50 g solvent for 1 hr at425X.
503 5.
CONCLUSIONS
Dispersed catalysts appear to be an effective way to convert these difficult to process feedstocks, coal, resids, and heavy oils. These catalysts are less susceptible to catalyst deactivation, they can increase distillate yields while at the sametimedecreasing coke formation. The dispersed catalysts precursors that are akeady in the correct stoichiometry and oxidation state of the active catalyst are more easily activated during the reaction, and give superior conversions. In future work we hope to explore these catalysts to determine the most active metals and combination of metals that will further increase their performance during the hydrocarbon conversion process. The authors gratefully acknowledge the partial support of this work by the Department of Energy under Contract No. DE-AC22-91PC91039.
REFERENCES 1.
Report from DOE Direct Liquefaction Contractors' Review Meeting, Pittsburgh, Pennsylvania, September 1995.
2.
Personnel communication 1996, S. Leiby, SRI Intemational.
3.
J. Reynolds, and W. Biggs, Ace. of Chem. Res., 21 (1988) 319-326.
4.
C. Galarraga, and M. M. Ramirez de Agudelo, J. of Catalysis, 134 (1992) 98-106.
5.
I. A. Wiehe, Am. Chem. Soc. Div. Petroleum Prepr., 38(2) (1993) 428-433.
6.
B. A. Cowens, R. C. Haltiwanger, and M. R. DuBois, Qrganometallics 6 (1987) 995-1004.
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® 1997 Elsevier Science B. V. All rights reserved, Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
505
HYDRODESULPHURIZATION OF GAS OIL USING C0-M0/AI2O3 CATALYST A.S. Nasution and E. Jasjfi "LEMIGAS R/D Centre for Oil and gas Technology, P.O. Box 1089 JKT, Jakarta 12230, Indonesia Three types of gas oil has been hydrodesulphurized using Co-Mo/A1203 catalyst to improve the performance of diesel fuel product. Sulphur removal increases with high temperature and pressure but it decreases with space velocity. Cracked gas oil is more difficult to hydrotreat than straight run gas oil. 1.
INTRODUCTION
Gas oil generally consist of predominantly straight-run hydrocarbons obtained from fractional distillation of crude oil; however compounds such as thermally and catalytically cracked gas oils, and hydrocracked gas oil are sometimes included. Sulphur containing molecules in the heavier petroleum distillates can be either benzothiophene or dibenzothiophene, naphthalene, anthracene, and heavier compounds are also found. In the case of cracked stocks, either thermal or catalytic cracked gas oil, they content a higher persentages of aromatic and olefin hydrocarbons; they have a higher specific gravity and a lower cetane number than the other two types of gas oils. Hydrodesulphurizing gas oil for diesel engine reduces pollution; and some degree of aromatic and olefin saturation leads to an improvement in the cetane number and stability. The present paper discusses the hydrodesulphurization of straight-run, thermal and catalytic cracked gas oils using C0-M0/AI2O3 catalyst using a catatest unit operated in continous system. 2.
EXPERIMENTAL
Straight-run, thermal and catalytic cracked gas oils have been used as feedstock in this experiment. Electrolytic hydrogen used was purified before uset from oxygene compound by passing this hydrogen into deoxo catalyst and then followed by drying in the molecular sieve. The Co-Mo/AI2O3 catalyst used is of the commonly employed in the hydrotreating process. The hydrotreating of gas oil was carried out in the catatest unit which can be operated in continous system. The volume and inside diameter of reactor are 220 cc and 19 mm, respectively. 80 cc of the catalyst was charged to the reactor. Operating conditions are temperature from 270 to 340° C, pressure: from 20 to 50 kg/cm^ and Hj/HC = 100 ItAt. Sulphur content, specific gravity and cetane number of feedstock and product were analyzed by using ASTM D-1551, ASTM D-1298 and ASTM D-613.
506 3.
RESULT AND DISCUSSION
Sulphur removal increases with the sulphur contents of gas oils (Figure 1). The residual sulphur content from any catalytic hydrodesulphurized products is observed almost directly proportional to the space velocity or throughout per catalyst volume, so that doubling the space velocity doubles the residual sulphur content (Figure 2).
100
i
100 ^
i
75
50
1.0 1.5 2.0 SULPHUR CONTENT, %wt
Figure 1: Relation between sulphur content and sulphur removal, note: straight-run (*), thermal cracked (o) and cat. cracked (•) gas oils.
75 r
2.5 LHSV
5.0
vol/vol/h
Figure 2: Relation between space velocity and sulphur removal note: straight-run (*), thermal cracked (o) and cat. cracked (•) gas oils.
tower-boiling compounds are desulphurited more easily than higher-boiling ones. The various molecules have different reactives, or with mercaptane sulphur much easier to eleminate than resonant sulphur molecules like thiophene or dibenzothiophene founded in cracked gas oil. The structural differentes between the various sulphur-containing molecules make it impractical to have a single rate expression applicable to all reactions m hydrodesulphurization, as was possible for the hydrogenation of aromatics and olefins. Each sulphur-containing molecule has its own hydrogenolysis kinetics that is usualy complex because several successive equalibrium stages are involved; and these are often controlled by internal diffusion linwttations, at least under industrial operating conditions. Cracked gas oil is more difficult than straight run gas oil to hydrotreat because of its higher thiophenic sulphur and the multi-ring organic sulphur content. Increasing of temperature and pressure increase sulphur removal and aromatic saturation (Figure 3 dan 4). Temperature and pressure effects are to some extent mtearelated. The use of high temperature and low pressure promotes dehydrogenations reactions, will increase laydown of carbon and thereby shortens the hours on stream before regeneration. Combination of high pressure and low tempersature induce sulphur removal and saturation of hydrocarbon. The spesific gravity increases 1-4 API, and the cetane number undergoes an increasses of 1-6 number.
507 100
100
I
&
I
g
75 h
75 I
50
50 260°
310**
TEMPERATURE , **C
Figure 3: Relation between temperature and sulphur removal, note: straight-run (*), thermal cracked (o) and cat. cracked (•) gas oils.
10
35
60
PRESSURE , kg/cm2
Figure 4: Relation between pressure and sulphur removal. note: straight-run (*), thermal cracked (o) and cat. cracked (•) gas oils.
The removal of a fairly heavy molecule sulphur and its replacement by hydrogen results in a small but drops in specific gravity, proportional to the amount of sulphur removal and aromatic saturation. Cracked gas oil has a higher improvement of both cetane number and specific gravity than straight run gas oil because most of its higher heavy aromatics are saturated, although its total aromatic content may not decrease significant. Hydrogenation of dinuclear aromatics to tetraline converts these refractory dinuclear aromatics into compounds which crack readly to monocyclic aromatics. The most refractory substituted heterocyclic compounds are usually associated with aromatics which has higher density. Sulphided compounds of a given structure are more refractory as their molecular weight gets higher. The compounds produced from the breakdown of sulphur and other contaminant are of high aromatics content when processing catalyst cycle oil because the sulphur exists largely as mono and dibenzothiophenes. This mono aromatic diesel oil usually has a high cetane number. With the increasing both the cracked stock portion in the diesel oil pool and the stringent diesel oil specifications, i.e. sulphur content < 0.05% wt; total aromatic = from 10 to 20% vol with polycyclic aromatic ^ 1.4% wt and cetane number ^ 48, the gas oil hydrotreating process will play a more important role. The traditional hydrotreating process is still dependent on a single-stage reactor system that combines severe operating conditions with hydrogenation using a single catalyst type such as C0-M0/AI2O3. The other process is a dualstage reactor in which the hydrodesulphurizing step with Co-Mo/AI2O3 catalyst, is followed by hydrocarbon saturation step using Ni-Mo-ZAljOj or Ni-W/AlzOs catalyst. The pressure of this dual-stage system is much lower at around 50 kg/cw? as comparred to above 80 kg/cm^ for the single-stage system.
508 4.
CONCLUSION
Sulhpur removal increasse with the temperature and pressure but decreases with space velocity. Specific gravity increase 1-4 API, and the cetane number undergoes an increase of 1-6 number. The various types of sulphur containing molecules in gas oil feeds have very different reactivity, with mercaptan sulphur easier to eleminate than resonant sulphur compounds like benzothiophenic, usually associated with aromatic which have higher feed densities. If the purpose is the saturation of aromatics and olefins the pressure must be higher than that for hydrodesulphurization. Combination of high pressure and low temperature induce sulphur removal and saturation of hydrocarbons. With the increasing both the cracked stock portion in the diesel fuel and the stringent diesel oil specifications, the hydrotreating process plays an important role. A dual-stage reactor needs a much lower operating pressure compared to a single-stage reactor, for the hydrotreating of gas oil into high quality diesel oil. REFERENCES 1.
I.E. Germain, Catalytic Conversion of Hydrocarbons, Academic Press, New York, (1969).
2.
B. Delmon and G.F. Froment, Catalytic Deactivation, Elsvier Scientific Publ. Co, New York (1980), 271.
3.
J.F. Le Page, Applied Heterogenous Catalysis, Edition Technip (1987), 357.
4.
A.S. Nasution, XI Simposio Iberoameracano de Catalysis, Guanajuato, Mexico, (1988).
5.
S.L. Lee et. al., Fuel Reformulation, Vol. 3, No. 3, (1993), 26.
6.
B.H. Cooper, et al.. Hydrocarbon Processing (June 1993), 83.
® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
509
The application of cobalt containing acidic zeolites as catalysts for hydrodesulfurization reactions Tamas I. Koranyi^*', Ngan H. Pham'', Andreas Jentys^ and Haimelore Vinek*' Institute of Isotopes of the Hungarian Academy of Sciences, P.O. Box 77, H-1525 Budapest, Hungary ^'Institut fiir Physikalische Chemie, Technische Universitat Wien, Getreidemarkt 9/156, A-1060 Wien, Austria Cobalt containing HY and HZSM5 zeolites were prepared by impregnation and solid state ion exchange methods. The type and concentration of acid sites of non sulfided and presulfided faujasites, HZSM5 and one CoHZSM5 zeolite prepared by impregnation were studied by adsorption and temperature programmed desorption (TPD) of pyridine followed by infrared spectroscopy. The concentration of Bronsted acid sites of unsulfided cobalt containing zeolites was low compared to the starting materials, indicating reaction between the OH groups of the support and the cobalt species. Three types of Lewis acid sites were observed on the CoHY samples. They were assigned to weakly acidic Lewis sites (1440 cm'^), extralattice Al (1448 cm'^) and Co (1448 cm"^) species. After sulfidation the Bronsted acidity of cobalt containing zeolites highly increased and the number of Lewis acidic cobalt species strongly decreased, which means that cobalt is sulfided at high extent. The faujasite samples are more active in the thiophene hydrodesulfurization (HDS) reaction compared to CoHZSM5 because of their higher Co dispersion. The overall catalytic activity can be divided into C-S bond hydrogenolysis activities without C-C bond breaking, cracking and to the reaction to coke formation. Coking and cracking are comiected with the acidity, cracking is a consecutive reaction of the hydrogenolysis products. The C-S bond hydrogenolysis activity can be attributed to sulfided cobalt species. 1. INTRODUCTION Enviromnental constraints will further decrease the levels of sulfur, nitrogen and metals tolerated in petroleum products, while in parallel the concentration of these impurities will increase in the crude oil available in the ftiture. Co and Ni containing molecular sieves are possible materials for replacing the routinely used alumina supported Co-Mo, Ni-Mo or Ni-W sulfide catalysts in hydrotreating reactions, where heteroatoms like S, N and metals are removed from petrochemical hydrocarbon feeds (1). By using CoS, NiS or CoMoS containing molecular sieves we expect to constrain the size of the sulfide particles and to obtain an interaction of the clusters with the zeolite host lattice which will lead to a structural stabilization of the clusters. Moreover, the shape of the zeolite pores will restrain the enviromnent available for the reactants, thus enabling shape selective reactions over these materials. Finally, the acid base properties of the zeolite can be fine tuned to specific needs of the desired reaction.
510 In this contribution we compare the acidity and the thiophene hydrodesulfurization (HDS) activity of cobaU containing HY and HZSM5 zeolites prepared by impregnation and solid state ion exchange. The effects of preparation method, metal loading and acidity are studied. We try to separate the C-S bond hydrogenolysis without C-C bond breaking of thiophene within the overall HDS activities and to find a connection between selectivities and structural properties. 2. EXPERIMENTAL Cobalt containing faujasites (Y) and ZSM5 (Z) zeolites were prepared by impregnation and by solid state ion exchange. The preparation method and metal loading of the catalysts are shown in Table 1. The CoHY zeolites were prepared by impregnation of NH4Y, prepared from NaY [FCA/K Degussa] with Co(N03)2 solutions. After impregnation the samples were dried at 383 K overnight. The calcination was carried out in air in two steps. First by heating up the sample by 2 K / min to 653 K, keeping the temperature for 12 h, and then heating up to 723 K and staying there for 2 h. HZSM5 (Si/Al =42) was solid state ion exchanged with C0CI2 heating the mixture in He at 773 K for 6 - 10 h. Co4Z was prepared by impregnation of HZSM5 with Co(N03)2 solutions. The calcination procedure was the same as in CoHY zeolites. Table 1 Preparation method and metal loading of the zeolite samples. Sample Starting zeolite Preparation method Co2Y, Co4Y, C06Y, C08Y HY impregnation ColZ, Co2Z HZSM5 solid state ion exchange Co4Z HZSM5 impregnation
Co^^ (wt%) 2,4,6,8 1,2 4
The structure of the zeolites was verified by XRD measurements. The acidic properties of the zeolites were studied by adsorption and temperature programmed desorption (TPD) of pyridine in their non sulfided and presulfided state by infrared (IR) spectroscopy. The samples (-'5mg) were prepared as self supporting wafers and placed inside an in situ flow cell. The samples were heated up to 673 K (10 K / min) and treated for one hour either in N2 (non sulfiding, NON) or in 8 % H2S/H2 (presulfidation, SUL) gas stream (30 cmVmin). After cooling down to 423 K the gas stream was changed to nitrogen. 4 /xl pyridine were injected into the nitrogen gas stream within 15 min. The type (Bronsted or Lewis) and concentration of their acidity was characterized by IR spectroscopy at 423 K followed by TPD (up to 773 K) of adsorbed pyridine. The IR spectra were recorded in a BOMEM FTIR spectrometer equipped with a DTGS detector (32 scans, 4 cm"^ resolution). The concentration of Bronsted and Lewis acid sites was calculated with the integral molar extinction coefficients of pyridine IR absorption bands determined by Emeis (2). For thiophene HDS test reactions the catalysts were heated up to 673 K in He flow (30 cm^/min) at a rate of 6 K/min and pretreated at this temperature in situ with a mixture of 5 % H2S in H2 (presulfidation) for 1 h. The reaction was carried out at 673 K and atmospheric pressure in a flow reactor system using 30 cmVmin flow of 3 % thiophene in H2. Thiophene and the products were detected using a GC equipped with a 3 m 1/8" n-octane / Porasil C column. The catalytic activities are described by - first order reaction rate constants, (knos =- (F/W) * In(l-X) where F =gas flow [mVs], W = catalyst weight [kg], X = conversion of thiophene)
511 - conversion, (X = 1 - unreacted thiophene / thiophene in feed) - yield, (products / (products +unreacted thiophene)) - C-S bond hydrogenolysis activity without C-C bond breaking, (C4 products / (products 4-unreacted thiophene)). We assigned the difference between conversion and yield to the quantity of thiophene remaining on the catalyst as coke. The difference between yield and C-S bond hydrogenolysis activity was used to calculate the amount of cracked products. Product selectivities were calculated from the composition of the product mixture. The C4 selectivity (C4 products / products) describes the ratio of uncracked products in the product mixture. The isomerization (isobutane / (isobutane +n-butane)) and hydrogenation ((isobutane -hn-butane) / C4 products) selectivities characterize the extent of isomerization and hydrogenation reactions over the catalysts. 3. RESULTS AND DISCUSSION The pyridine IR absorption bands (1700-1400 cm'^) related to unit (1 mg) sample weight of non sulfided HY, Co2Y, Co4Y and C06Y zeolites saturated with pyridine at 423 K are shown in Fig. 1. The band at 1541 cm"^ is assigned to pyridinium ions adsorbed on Bronsted acid sites (3). The bands at 1440 and 1448 cm'^ are assigned to pyridine adsorbed on Lewis acid sites (3). The doublet indicates that at least two kind of Lewis acid sites are present in the faujasites. The band at 1608 cm'^ in the spectra of CoHY is assigned to pyridine coordinatively bonded to cobalt cations (4). We observed that the intensity of the bands at 1608 and 1448 cm"^ increased in parallel with the cobalt loading, therefore we assigned the latter band - at least partially (see next paragraph) - to pyridine adsorbed on cobalt species. This assignment is in agreement with an other finding (5): the IR spectrum of a NaY zeolite [free from extralattice alumina] contains a singlet Lewis acidic band at 1441 cm"^ but new Lewis acidic band appears at 1453 cm'^ following partial ion-exchange of sodium with nickel. The pyridine IR absorption bands of non sulfided faujasites during TPD of pyridine at 573 K are shown in Fig. 2. Only a single Lewis acid site was observed at 1448 cm'^ at this temperature. The absence of the band at 1440 cm'^ indicates the weak adsorption of pyridine on this site. We assign the band at 1440 cm"^ to weakly Lewis acidic sites, and that at 1448
1600
WAVENUMBERS
Ccm
Figure 1. Pyridine IR absorption bands of non sulfided faujasites after saturation with pyridine at 423 K.
WAVENUMBERS
1500
(cm
~^)
Figure 2. Pyridine IR absorption bands of non sulfided faujasites during TPD of pyridine at 573 K.
512 cm'^ to pyridine adsorbed on Lewis acidic extralattice AP^and on Lewis acidic Co^^species. The IR spectra of pyridine adsorbed on sulfided faujasites at 423 K are shown in Fig. 3. The position of Bronsted and Lewis acidic bands did not change during sulfidation. The relative concentration of Bronsted acid sites compared to Lewis acidic sites is much higher after the sulfidation procedure (see also Table 2). We attribute the changes of Br / Le ratios to the regeneration of OH groups of the support and to the decrease of Lewis acidic cobalt species by the formation of cobalt-sulfide during sulfidation. The pyridine IR absorption bands of sulfided faujasites during TPD of pyridine at 573 K are shown in Fig. 4. Similar to the non sulfided faujasites (Fig. 2) the weakly Lewis acidic bands at 1440 cm'^ already disappeared at this temperature.
Co6HY
Co4HY
1600
WAVENUMBERS
1500
Tern - 1 )
Figure 3. Pyridine IR absorption bands of sulfided faujasites after saturation with pyridine at 423 K.
1600
WAVENUMBERS
1500
Ccm - 1 ^
Figure 4. Pyridine IR absorption bands of sulfided faujasites during TPD of pyridine at 573 K.
The pyridine IR absorption bands of non sulfided and "sulfided" (pretreated in HjS / Hj) HZSM5 zeolites are shown in Fig. 5. The bands at 1612 and 1452 cm'^ are assigned to Lewis acidic sites, the bands at 1638 and 1545 cm"^ to Bronsted acidic sites. The intensities of Bronsted and Lewis acidic bands show small changes during sulfidation and the concentration of Lewis acidic sites decreases during TPD at 573 K. The pyridine IR absorption bands of non sulfided and sulfided Co4Z zeolites are shown in Fig. 6. The position of the bands is the same as in the HZSM5 sample, new bands do not appear. The concentration of the Lewis acidic band at 1452 cm'^ compared to the Bronsted acidic band at 1545 cm"^ is much higher than in the case of HZSM5 showing the presence of Lewis acidic Co^"" species. The intensity of Lewis acidic bands decreases, that of Bronsted acidic bands increases during sulfidation, therefore we assume that cobalt is sulfided at a high extent. The band at 1452 cm'^ does not disappear completely during sulfidation similar to the band at 1448 cm"^ in the faujasites, therefore we assume that the band at 1452 cm"^ also contains at least two components: Lewis acidic Co^^and some extralattice AP"^ species. The concentration of acid sites calculated from the intensity of the bands at 1540 cm"^ and 1440-1450 cm'^ by the method described by Emeis (2) is shown in Table 2. The Bronsted acid sites of HY and HZSM5 (HZ) decrease due to impregnation by cobalt indicating reaction between the OH groups of the support and the cobalt species. The concentration of Lewis acid sites increases parallely with the cobalt loading in the non sulfided faujasites. This increase is
513
SUL 573 K
LiJ
SUL 423 K
m g
SUL 573 K
z< if)
1600
WAVENUMBERS
1500
WAVENUMBERS
Ccm - 1 )
Figure 5. Pyridine IR absorption bands of non sulfided (NON) and sulfided (SUL) HZSM5 zeolites after saturation with pyridine at 423 K and during TPD at 573 K.
Tern - 1 ^
Figure 6. Pyridine IR absorption bands of non sulfided (NON) and sulfided (SUL) Co4Z zeolites after saturation with pyridine at 423 K and during TPD at 573 K.
Table 2 Bronsted (Br) and Lewis (Le) acidity (mol/kg) as well as Br/Le concentration ratios of non sulfided (NON) and presulfided (SUL) zeolites after saturation with pyridine at 423 K and during TPD of pyridine at 573 K. Sample Br
saturated at 423 K Le Br/Le
Br
TPD at 573 K Le
Br/Le
HYNON Co2Y NON Co4Y NON Co6Y NON HZ NON Co4Z NON
0.68 0.39 0.35 0.39 0.14 0.07
0.66 0.81 0.96 1.55 0.05 0.18
1.04 0.48 0.36 0.25 3.02 0.37
0.68 0.48 0.41 0.43 0.12 0.05
0.20 0.27 0.59 1.19 0.02 0.09
3.35 1.80 0.70 0.36 7.75 0.57
HYSUL HYS+N Co2Y SUL Co4Y SUL Co6Y SUL HZ SUL Co4Z SUL
0.62 0.63 0.56 0.57 0.82 0.16 0.09
0.40 0.52 0.49 0.47 0.44 0.04 0.05
1.55 1.21 1.14 1.21 1.88 4.33 1.86
0.58 0.62 0.57 0.51 0.77 0.13 0.07
0.11 0.14 0.14 0.24 0.14 0.01 0.02
5.22 4.28 4.33 2.11 5.44 10.1 2.75
much higher in the Co4Y zeolite (0.30 mol/kg) than in the Co4Z sample (0.13 mol/kg) indicating a much lower dispersion of cobaltin the latter. Presumably the strong Bronsted acidic SiOHAl groups are covered partially by Lewis acidic cobalt species. During sulfidation the initial Bronsted and Lewis acidity of HY decreased (Table 2) indicating some adsorption of H2S on the strong Bronsted acidic SiOHAl groups and also on
514 the Lewis sites. After one hour sulfidation at 673 K a HY sample was purged in nitrogen for one hour at 673 K, then its acidity was studied (HY S+N).The concentration of Bronsted and Lewis acidic sites in the HY S+Nsample was between those in the HY NON and HY SUL zeolites (Table 2) indicating partial desorption of H2S during purging at 673 K. Ratnasamy and Knozinger (6) observed that pyridine competes with thiophene for the adsorption sites in HDS catalysts. We assume a sunilar competion between H2S and pyridine for the acid sites of HY causing the apparent decrease of Bronsted and Lewis acidity after sulfidation. In sulfided CoY and Co4Z zeolites the Bronsted acidity increases and the Lewis acidity decreases leading to an increase of the Br/Le ratios of at least twofold after sulfidation (Table 2). The increase of the concentration of Bronsted acid sites is attributed to the regeneration of SiOHAl groups and to the creation of new Bronsted acid sites ((CoOH)^ and/or (CoSH)'^(7)) during sulfidation at high cobalt loading (in C06Y). The Lewis acidity of the non sulfided and sulfided zeolites highly decreases, but the amount of pyridinium ions adsorbed on Bronsted acid sites hardly changes during TPD at 573 K compared to the concentrations at 423 K (Table 2). It can be explained by the stronger adsorption of pyridine on the Bronsted than on the Lewis acid sites and that H2S blocks rather Lewis than Bronsted acid sites. Three types of Lewis acid sites were observed on the CoY samples. The band at 1440-1450 cm"^ was separated to three contributions assigned to weakly acidic Lewis sites (1440 cm'^), extralattice Al (1448 cm'^), and Co (1448 cm'^) species (Table 3). For calculating the number of Co sites a constant concentration of extralattice Al sites was assumed. The concentration of Lewis acidic Co species was correlated with the cobalt loading in the non sulfided catalysts. After sulfidation the number of Lewis acidic cobalt species decreased to a very similar value (0.07 YO.Ol), therefore we conclude that cobalt is sulfided at high extent in the faujasites. The 0.07 mol/kg concentration of Lewis acidic cobalt species after sulfidation can be attributed to non sulfided, presumably oxidic cobalt species, because similar acidity did not remain after sulfidation of Ni2NaY and Ni4NaY zeolites prepared by ion exchange (5). The decrease of weak and extralattice Al Lewis acidity during sulfidation is similar, about 1/3 of the original value. The concentration of weak (0.28 mol/kg) and extralattice Al (0.25 mol/kg) Lewis acidity of the HY S+Nwas between those of HY NON and HY SUL samples, therefore we assume a strong adsorption of H2S on these sites. The first order thiophene HDS reaction rate constants of the zeolites and of some alumina supported reference catalysts (Eurocat C0/AI2O3, C0M0/AI2O3 and NiMo/Al203) are shown in Table 4. The catalytic activities of faujasite (Y) based samples were higher or comparable with those of Mo containing alumina supported catalysts, while the CoZ zeolites had lower activities. Table 3 Separation of the pyridine IR band (1430-1450 cm'^) of non sulfided (NON) and presulfided (SUL) zeolites after saturation with pyridine at 423 K into Lewis acidity (mol/kg) of weak (weak), extralattice Al (exAl) and cobalt ion (Co) sites. SUL NON Sample weak exAl weak exAl Co Co 0.27 0.21 0.39 0 HY 0.18 0 0.36 0.27 0.24 Co2Y 0.18 0.18 0.06 0.37 0.27 Co4Y 0.31 0.23 0.18 0.08 0.37 0.27 0.66 C06Y 0.18 0.18 0.07
515 Table 4 First order reaction rate constants (k^s) ^^^ reaction product selectivities (C4, isomerization, and hydrogenation) of presulfided reference and zeolite catalysts. C4 selectivity Isomerization Hydrogenation %DS selectivity (%) Catalyst selectivity (%) (10-'*mV(kg*s)) (%) 93 0 6 C0/AI2O3 15 51 99 1 22 C0M0/AI2O3 2 99 61 NiMo/Al203 20 47 24 90 Co2Y 31 62 85 Co4Y 93 42 82 136 65 C06Y 46 66 80 120 45 C08Y 24 55 7 ColZ 10 13 30 40 7 Co2Z 27 12 59 Co4Z 3 The activities of the CoY zeolites increased with the metal loading, while they were practically constant for the CoZ catalysts. However, the first order reaction rate constants reflect the overall catalytic activies only. Determining the contributions of coking, cracking and C-S bond hydrogenolysis without C-C bond breaking to the total activity gives more information on the reaction mechanism of thiophene HDS. The ratio of coking, cracking and C-S bond hydrogenolysis activity within the total activities is shown in Fig. 7. The C0M0/AI2O3 and NiMo/Al203 reference Eurocat HDS catalysts exhibit C-S bond hydrogenolysis without C-C bond breaking only. For the C0/AI2O3, Co2if and CoZ catalysts coking is higher than the hydrogenolysis activity. The amount of coking (formation of presumably sulfur containing carbonaceous deposits) changes between 15 and 30 % of the conversion level for the CoY samples. We do not observe any correlation between coking and metal loading. A rough estimate may exist between the relatively high Bronsted and Lewis acidity of sulfided CoY zeolites compared to those of sulfided Co4Z (Table 2) and the relatively high coking on the former compared to the latter. We found some H2S adsorption on the faujasites but minor if any on the ZSM5 zeolites. Parallely assuming higher amount of adsorbed thiophene on the CoY than on the CoZ zeolites may lead to higher amount of coking on the former samples. The C-S bond hydrogenolysis activity without C-C bond breaking and the cracking increased with the metal loading in the Y zeolites (Fig. 7), therefore these activities should be connected with the presence of sulfided cobalt species. Presumably cracking is a consecutive reaction of the C-S bond hydrogenolysis products and proceeds on acid sites. Consequently coking and C-S bond hydrogenolysis take place on different sites; the former on acidic, the latter on sulfided metal species. The Bronsted acidity of sulfided alumina supported NiW and NiMo catalysts is negligible (8) which explains the lack of coking and cracking reactions on C0M0/AI2O3 and NiMo/Al203 catalysts (Fig. 7). Subtracting coking from the overall activities, the C-S bond hydrogenolysis activities of Co4Y, C06Y and C08Y zeolites are still comparable with those of the C0M0/AI2O3 and NiMo/Al203 catalysts (Fig. 7). The selectivities to the formation of C4 products are shown in Table 4. The C4 selectivities of the alumina supported catalyst are higher than 90 %, those of the zeolites are lower than 70
516
•
Hydrogenolysis Cracking
I
o
o
5
o
o o
5
z o
>
o
o
lO
-<
O
1
o
o
-<
o
o
o
o
-<
NO C
00
-A
S
Ol
o
o
1^
%
o
CO
s Ol
Coking
CO
s
Ol
Figure 7. C-S bond hydrogenolysis without C-C bond breaking (Hydrogenolysis), cracking and coking activities of reference and zeolite catalysts. % leading to a high amount of cracked and coked products over the zeolites compared to the reference catalysts. Very high isomerization selectivities (Table 4) are observed on the zeolite based catalysts, while on the reference catalysts they are close to zero. The high isomerization selectivities of the zeolites can be connected with the Bronsted acidity of the zeolite lattice, but not with the active, metal sulfide species. The hydrogenation reaction might be important, because it is suggested (9), that in thiophene HDS hydrogenation of the aromatic ring takes place before the C-S bond hydrogenolysis step. Comparing the hydrogenation selectivities with the first order reaction rate constants (Table 4) they change parallel, but conclusions to the reaction mechanism cannot be taken. 4. CONCLUSIONS The CoHY zeolites contain Bronsted and three kind of Lewis acidic sites. Increasing the cobalt loading of CoHY samples the Bronsted acidity decreases, while the Lewis acidity attributed to cobalt species increases. The strong Bronsted acidic SiOHAl groups of the zeolites are partially covered by Lewis acidic cobalt species. After sulfidation the amount of pyridine adsorbed on HY slightly decreases due to some competitive adsorption of HjS on the Bronsted and Lewis acidic sites. The increase of Bronsted acidity in cobalt containing zeolites is attributed to the regeneration of SiOHAl groups and to the creation of new Bronsted acid sites (presumably (CoOH)"^ and/or (CoSH)^)at high cobalt loading. The other effect of sulfidation is the decrease of Lewis acidic cobalt species to a similar low value which means that cobalt is sulfided at high extent in the zeolites. The sum activities (knos) of cobalt containing faujasites are comparable with those of alumina supported Eurocat catalysts. The CoZSM5 zeolites show very low activities, which indicate a low dispersion of the metallic species both in the non sulfided and sulfided samples. The high activities of Y zeolites and their linear increase with the metal loading indicate, that
517 the metal species are sulfided and highly dispersed. The overall catalytic activities can be divided into coking, cracking and C-S bond hydrogenolysis without C-C bond breaking activities. Coking cannot be directly connected with the metal content but presumably with the acidity of the zeolites due to the strong adsorption of thiophene on the acid sites. Cracking is a consecutive reaction of the hydrogenolysis products which proceeds also on acid sites. The C-S bond hydrogenolysis activities are connected with the presence of sulfided cobalt species. We attribute the differences between the HDS activities of zeolite and alumina supported catalysts (Table 4) to their different acidities and to the weak interaction between cobalt and the zeolite compared with the strong cobalt - alumina interaction (if Mo is absent) which can hamper sulfidation of Co species (10). We expect, that the use of zeolites as supports for transition metals as catalysts for hydrotreating reactions is a promising perspective for future applications. The high cracking within the total HDS activity of Y zeolites as compared to the Eurocat reference catalysts (Fig. 6) can be advantageous or even desirable in the HDS of heavy compounds (e.g. dibenzothiophene derivatives). The high isomerization selectivity of the zeolites can also be favourable in thiophene HDS, because isobutene and isobutane are valuable starting substances in the synthesis of methyl-tertbutyl-ether. ACKNOWLEDGEMENTS The work was partially supported by the "Fonds zur Forderung der Wissenschaftlichen Forschung" under project FWF P9167 and by the "Hungarian-Austrian Intergovernmental Science and Technological Cooperation Programme" under project No. A-20. T.I.K. thanks the Lise Meitner postdoctoral fellowship for the Austrian Science Foundation (M00091-CHE). REFERENCES 1. C.W.Ward, Stud. Surf. Sci. Catal., 16 (1983) 587. 2. C.A. Emeis, J. Catal., 141 (1993) 347. 3. P.A. Jacobs and H.K. Beyer, J. Phys. Chem., 83 (1979) 1174. 4. M.I. Vazquez, A. Corma and V. Fomes, Zeolites, 6 (1986) 271. 5. T.I. Koranyi, submitted for publication. 6. P. Ratnasamy and H. Knozinger, J. Catal., 54 (1978) 155. 7. T.I. Koranyi, V.V. Rozanov and E.A. Rozanova, Book of Abstracts of 10th Intern. Conf. on Fourier Transf. Spectr., Poster No. B5.10, Budapest, Hungary, 1995. 8. T.I. Koranyi, M. Dobrovolszky, K. Matusek, Z. Paal and P. Tetenyi, Fuel, in press. 9. M.E. Bussell and G.A. Somorjai, J. Catal., 106 (1987) 93. 10. B. Scheffer, P.J. Mangnus and J.A. Moulijn, J. Catal., 121 (1990) 18.
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® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
519
Hydrodearomatization of naphtenic base machine oils G. Kons, H.-J. Miiller, M. Vicari, E. Schwab and M. Walter BASF AG, D-67056 Ludwigshafen Low sulphur- and polycyclic aromatics-contents are required for high-performance lube oils. To reach these specifications, naphtenic based feedstocks can be hydrotreated using a NiMocatalyst which shows high desulphurization and dearomatization rates, with low cracking activity. 1. INTRODUCTION Part of the lube oils are produced from naphtenic based hydrocarbons, contaihmg high contents of sulphur and polycyclic aromatics (PCA). For their performance as well as for environmental and safety reasons it is necessary to reduce the PCA and sulphur contents [1-2]. The conventional process for this purpose is solvent extraction, followed by clay- or hydrofinishing (Fig. 1). Drawbacks of this technology are low yields, disposal of by-products and the need to recycle the solvent. Fig. 1: conventional process for the up-grading of naphtenic based hydrocarbons to lube oils Extracts
Vacuum Destination
Extraction
Hydrofinishing
520 BASF has developed an alternative for the traditional technology, namely a catalytic hydrotreatment process (Fig. 2) which offers a most elegant solution with many advantages over the conventional technique: higher yields, improved quality and no by-products to be disposed of. Precondition is the development of robust, sulphur-resistant and performant catalysts in term of dearomatization, but with low cracking activity. There is a large fund of experience with lube oil processes at BASF which goes back to the time when BASF was operating an own lube oil refinery through its subsidiary Wintershall. These processes have continuously been updated and are today licensed through engineering contractors. In the case of hydrodearomatization of naphtenic base machine oils, a BASF developed NiMo catalyst run in a trickle bed reactor shows an excellent performance, not only by deep dearomatization with low cracking activity but also with high desulphurization rates. Fig. 2: BASF process for the up-grading of naphtenic based hydrocarbons to lube oils ^SSSEBS^
Vacuum Destination
Hydrotreating
2. EXPERIMENTAL Several machine-oil-feedstock-grades - 8 W, 56 W, 320 W, 680 W - have been processed in a 1,5 1 trickle bed reactor under following reaction conditions: Pressure 80 - 95 bar Temperature (reactor inlet) 250 - 360 °C 0,2.0,4kg/(l.h) WHSV 0,75-1,0 Nm^/kg H2/oil - ratio Catalyst NiO-Mo03 on activated AI2O3 1.5mmextrudates sulfided at 300 °C with Dunethyldisulfide
521 Pure hydrogen (99,9 vol.-%) was led through the reactor in oncethrough operation. The liquid product was made free of H2S and NH3 after the high pressure separator with a washing column using NaOH-solution. The light ends of the products were topped off in order to reach the required flash point specifications. Samples were characterized by following standard methods: IP 346 ASTM D-2622 ASTMD-1500 ASTM D-445
PCA Sulphur Colour Viscosity
3. RESULTS The obtained results are listed in following table:
PCA (wt-%) Sulphur (wt-ppm) Colour Viscosity (mmVs) 40 °C 100 °C Yield (wt-%)
Feed 5.7
8 W 56 W 320 W 680 W Product Feed Product Feed Product Feed Product 2.4 10.2 2.4 9.5 2.4 9.3 2.8
890
6
1850
60
2200
80
2500
145
L 1.0
L 0.5
L 2.5
L 0.5
L 4.0
L 0.5
L 7
L 2
8.63 2.20
8.99 2.26 93.1
55.05 5.70
47.65 5.36 99.2
326.1 14.2
222.6 12.1 98.6
742.7 24.1
439.8 19.1 98.2
The decisive criterion was to lower the PCA-content in the feedstocks to less than 3 wt-% (PCA-content according IP 346). This specification is being discussed in different countries as limit for base oils. By these test-runs, BASF proves that naphtenic base machine oils with a PCA-content (according IP 346) of max. 3 wt-% can be produced at high yields by hydrotreatment alone, with also high desulphurization rates (94 to 99 %). These pilot plant tests provide a good basis to realise the transfer of the results to the technical scale. One plant using this technology has been constructed m Japan and came on stream with a capacity of 70 000 t/y in mid 1995. REFERENCES 1. 2.
D. Klamann, in Ullmann's Encyclopedia of industrial chemistry, Vol. A15, p. 423 VCH, Weinheim, 1990 J.F. Le Page et al.. Applied heterogeneous catalysis, p. 435, Ed. Technip, Paris, 1987
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523
Highly dispersed metal sulfide catalysts for the hydroconversion of vacuum destination residues K. Bilker*, H. Bemdt*, B. Liicke* and W. Kotowski^ ^Institut fiir Angewandte Chemie Berlin-Adlershofe.V., Rudower Chaussee 5, 12484 Berlin, Germany ^Institut fiir Schwere Organische Synthese, Ul. Energetykow 9, 47-225 Kedzierzyn-Kozle, Poland 1. INTRODUCTION The present situation on the global oil market is characterized by a steadily increasing demand for middle destiUates such as gasoline, diesel fuel, kerosine and a significant decrease of the quaUty of crude oils. Depending on the deposit, the crude oils may contain remarkable amounts of coke precursors and high concentrations of sulfur, nitrogen, oxygen, in form of heterocychc organic compounds as weU as metals such as vanadium and nickel. These impurities not only comphcate further upgrading of heavy feeds by poisoning catalysts during refinery processing but lead also to the severe problem of environmental pollution due to emmission of SO2 and NOx. For these reasons, it wiU be necessary in future to set up more efficient catalytic processes for the hydrotreatment of oils with low quahties, such as heavy feedstocks and/or vacuum destination residues (gudron). During the last years a number of papers have been pubUshed dealing with the use of microcrystalline hydrotreating sulfide catalysts dispersed in the oil residue [1-5]. In this contribution it will be reported about the promising use of oil- and water soluble precursors for the formation of highly dispersed metal sulfide catalysts for the hydrocracking of gudron under batch and flowing conditions. 2. EXPERIMENTAL 2.1. Catalyst precursors and formation of the active phase The catalyst precursors used in the experiments described below were oil soluble naphtenates of nickel, cobalt, tungsten or molybdenum. In principal water soluble ammonium salts of Mo and W and nitrates of Ni and Co can be used as weU.
524
Analysis of the solid residue (catalyst and coke) by XRD revealed that under hydrocracking conditions the metal naphtenate precursors are transformed into NiS, CoS, M0S2 and WS2. As it was observed by electron microscopy the average diameter of the crystallites was 4-5.5 M^m and their lenght 4-6 |xm. 2.2. Batch experiments Different amounts of metal naphtenates were solved in oil and added to 85 g gudron resulting in solutions containing 500 to 5000 ppm metal. In the case of NiMo, CoMo, and NiW sulfides, the metal naphtenates were mixed in such a way that an atomic ratio of the metals of 1:1 was adjusted. The mixture was heated in an autoclave (V = 0.25 1) under constant pressure (max. 20 MPa) and flowing hydrogen (250 l^h^ up to temperatures of max. 455**C. The reaction was quenched after different times by rapid cooling the autoclave with ice water. Gaseous reaction products were analyzed gas chromatographically, Uquid products were extracted with benzene and separated by destination. 2.3. Continous flow experiments [5] Hydrocracking was performed in a two stage reactor cascade under a pressure of 14 - 15 MPa, a space velocity of gudron of 0.8 h-i, and an supply of 1500 m^ H2 /m^ gudron. Catalyst concentrations were variied between of 600 ppm (NiMo) to 2000 ppm (NiW). The maximum temperature of the first and second reactor was 455**C but to achieve an optimum balance of the velocities of crackand hydogenation processes the temperature in the first reactor had to be kept 40''C below the value of the second one. 2.4. Properties of the vacuum destination residue (gudron) Gudron starts boiling at 494''C and is mainly composed of Carbon (85.60 wt.%) and Hydrogen (11.55 wt.-%). The amount of cokable materials (Conradson Carbon Residue, CCR) is 14.1 wt.-%. However, as can be seen in Table 1, gudron contains a number of heteroatoms. Table 1 Content of heteroatoms of gudron Sulfur
2.38 wt.-%
Oxygen
0.87 wt.-%
Nitrogen
0.46 wt.-%
Vanadium
114 ppm
Nickel
59 ppm
Iron
42 ppm
525
3. RESULTS AND DISCUSSION Batch experiments showed that among others the investigated NiW and CoMo sulfide catalysta are mostly efficient in terms of high hydrocracking conversion, low formation of coke and high degree of hydrodesulfurization (HDS). In figure 1. the influence of the residence time on the conversion of gudron and the formation of coke in presence of the NiW catalyst is displayed as one representative example.
[ 10
^ c o
c
>
a o
B u
o
U
1.0 1.5 Residence time / h Fig 1. Influence of the residence time on the conversion of gudron and the formation of coke at 14 MPa and 430*'C in batch autoclave experiments with flowing hydrogen. 1- conversion without and 2 - in the presence of 2000 ppm NiW-catalyst. 3- formation of coke without and 4- in the presence of catalyst. As can be deduced from this figure, presence of the NiW sulfide catalyst leads to a decrease of conversion of gudron. At the same time a significant decrease of the formation of coke can be observed. It was also found that without and in the presence of catalyst a linear correlation exists between the amount of hydrogen consumption and the conversion of gudron. However, in the presence of catalyst the amount of hydrogen consumption is about twice as high. Obviously, radicals produced by
526 thermal cleavage of higher hydrocarbons are saturated by activated hydrogen and as a result condensation reactions leading to precursors of coke are efficiently suppressed. In addition asphaltenes which are mainly responsible for coke formation are degraded by the action of the catalyst. Further, experiments were performed under conditions of continous flow of gudron in a cascade of two reactors. Figure 2a illustrates the high degree of desulfurization of the vacuum destination residue in presence of 600 ppm NiMo sulfide and figure 2b the influence of temperature on the gudron conversion. From these results, it can be deduced that even in the case of low amounts of catalyst and moderate temperatures efficient catalytic upgrading of vacuum destination residues can be achieved. Further it was found out that the catalyst concentration had almost no influence on the gudron conversion, indicating that cleavage of C-C bonds is mainly a thermal reaction.
/ 1
.2 '-C
Jc
s "3 TS
\r'
70' 65-
^ /
60
>> H
p
55 -
zr^
B O
U
p-
15
(a) 45 Conversion / wt.-%
75
^vCf / r " ^O-'O / A2 ^ - O -
70
P/
*»^ .2 50*5 >
A.
/
GQ Q>
O
^^
/
(i
1 .-^^-t
90
/
/4/
soii' ii
lo 'i
(b)
1.0 0.75 0.5 Volume velocity of gudron / h"^
Figure 2. Results of gudron hydrocracking at 15 MPa in a reactor cascade; temperature of first reactor: 415*'C. (a) Dependence of the degree of desulfurization on the conversion of gudron, temperature of second reactor: 445*'C. 1: 600 ppm NiMo; 2: without catalyst (b) Influence of temperature of second reactor and volume velocity of gudron on its conversion. 1: 450*^0; 2: 440**C; 3: 430°C.
527 4. CONCLUSIONS The oil soluble metal naphtenate precursors are transformed under the conditions of hydrocracking of gudron in highly disperse metal sxdfide catalysts consisting of microcrystallites with particle size below 6 |j,m. During hydrocracking hydrocarbon radicals are initially formed which are hydrogenated in further steps. The concentration of the catalyst therefore has little effect on the conversion of gudron but enhances the selectivity for generation of hydrocarbons with a higher H/C ratio and suppresses formation of coke. The presented upgrading process results in the conversion of the vacuum oU residue with a boiling point of 494**C into products which can be destOled afterwards, such as vacuum oU, crude gasoline and diesel. Highly dispersed NiW and CoMo sulfide catalysts have been proven as most efficient for both the conversion of gudron, high degree of desulfurization and low amounts of coke production. In future research activities wiU be focussed on the two stage hydroconversion process where the temperature of the first reaction vessel is kept about 40°C below the temperature of the second reactor to achieve an optimum balance of the velocity of crack- and hydogenation processes.
REFERENCES 1. A. Del Bianco, N. Panariti, S. Di Carlo, J. Elmouchnino, B. Fixari und P. Le Perchec, Applied Catalysis A: General, 94 (1993) 1 2. A. Morales, A. Salazar, C. Ovalles, and E. Filgueiras, 11th. Intern. Congr. Catal., 30. June - 5. July, Baltimore, USA 3. R. J. Bearden and C. L. Aldrige, Energy Process: 1 (1981) 44 4. A. Del Bianco, N. Panariti, S. Di Carlo, P.L. Beltrame, P. Camiti, Energy & Fuels 8 (1994) 593 5. W. Kotowski, H. Bemdt, K. Bilker and W. Fechner, Chem.-Ing.-Tech, submitted for publication
This Page Intentionally Left Blank
® 1997 Elsevier Science B. V. All rights reserved, Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
529
Hydrogenation of tetralin over a sulfided ruthenium on Y zeolite catalyst : comparison with a sulfided NiMo on alumina catalyst. J.L. Lembertona, M. Cattenot^ V. Kougionas^ M. Mhaouer^, J.L. Portefaix^ M. Breysse^ and G. Perota 3 Laboratoire de Catalyse en Chimie Organique, URA CNRS 350, Universite de Poitiers, 40 avenue du Recteur Pineau, 86022 Poitiers Cedex, France ^ Institut de Recherches sur la Catalyse, CNRS, 2 avenue Albert Einstein, 69622 Villeurbanne Cedex, France Abstract Hydrogenation of tetralin at 300°C, under hydrogen pressure, and in the presence of hydrogen sulfide was used to modelize the hydrogenation of aromatics in gasoils. The activity for this reaction of a sulfided ruthenium on dealuminated Y zeolite catalyst (RuKY^) was found to be much higher than that of a sulfided NiMo on alumina. Moreover, the activity of the RuKY^ catalyst was dependent very much on the sulfidation method : owing to a very significant carbon deposition initiated by the zeolite acid sites, the catalyst sulfided by dimethyldisulfide in solution in n-heptane was less active than when sulfided by a H2S/H2 mixture. This effect was not observed in the case of the NiMo/alumina catalyst, which is not acidic. On the other hand, whatever the catalyst, the presence of n-heptane as a solvent during the hydrogenation test decreased significantly the conversion of tetralin. This was due to the competitive adsorptions between the two molecules. 1. INTRODUCTION Pollution caused by diesel motors in cars and trucks (NOx, soots) is a problem of great concern, particularly in Europe. Tests have demonstrated that the principal element responsible for soot is a low Cetane Number of gasoil, connected with the aromatic content [1]. Therefore, a decrease in the aromatic content in gasoil, preferably through hydrogenation, would improve the quality of the gasoils (higher Cetane Number), and decrease pollution. The feeds to be treated, generally mixtures of straight run gasoil and light cycle oils, contain c.a. 1.5 wt.% sulfiir, which imposes the use of sulfide catalysts, for example the sulfided NiMo on alumina catalysts widely used in hydrotreating. However these catalysts, which were initially designed to achieve simukaneously hydrodenitrogenation and hydrodesulfidation of the gasoils, would not allow to obtain a satisfactory level for the hydrogenation of aromatics, unless the refiners increase significantly the size of the reactors or the rigour of the operating
530
conditions. Consequently, new catalytic formulations are required to achieve the hydrogenation of aromatics in the presence of hydrogen sulfide. From the numerous studies on the catalytic properties of transition metal sulfides for hydrotreating reactions, the most active sulfides are apparently those of group VIII metals, and particularly ruthenium sulfide [2-8]. On the other hand, it could be expected that the use of a zeolitic support instead of alumina could modify very much the activity of the sulfide phase : better dispersion, protection against poisoning by bulky molecules, destruction though acid cracking of certain molecules capable of inhibiting hydrogenation. In the present work, the hydrogenation of aromatics in gasoil was modelized by the hydrogenation of tetralin, at 300°C, under hydrogen pressure, in the presence of hydrogen sulfide. The catalysts used were a classical NiMo on alumina, and ruthenium dispersed into a potassium-exchanged dealuminated Y zeolite. Both catalysts were tested under two completely different experimental conditions : sulfidation by H2S or by a dimethyldisulfide/n-heptane mixture, transformation of tetralin alone or in the presence of a solvent (n-heptane). This procedure allowed to compare the hydrogenating activities of the two catalysts as a function of the sulfidation method and of the reaction conditions. 2. EXPERIMENTAL 2.1. Catalysts The starting material was a commercial dealuminated HY^f zeolite (Conteka CBV 712, Si02/Al203 molar ratio = 6), calcined at 530°C under dry oxygen. The KY^j zeolite was prepared through two successive exchanges with a 1 M KNO3 aqueous solution at 60°C, followed by a calcination at 530°C under dry oxygen. Ruthenium was introduced into the zeolite by ion exchange with an aqueous [Ru(NH3)6]Cl3 solution (Johnson-Mattey), for 48 hours at room temperature. Lastly the catalyst was washed three times with water, then dried overnight at 120°C under air. The RuKYj so obtained contained L8 wt.% ruthenium and 0.8 wt.% potassium. Ruthenium sulfide was shown to be well-dispersed within the zeolite lattice [9]. NiMo/Al203 was a commercial catalyst, containing 9.3 wt.% molybdenum and 2.4 wt.% nickel. 2.2 Catalysts sulfidation Before reaction, the catalysts were sulfided for 12 hours at 400°C using either H2S (15 vol.% in H2) or dimethyldisulfide (5 wt.% in n-heptane). In the first case, the sulfidation was carried out at atmospheric pressure, under a 5 ml.mn'^ H2S/H2 flow (H2S pressure =15 kPa). In the second case, the total pressure was 3.5 MPa, the H2S pressure obtained from dimethyldisulfide decomposition being 240 kPa; n-heptane pressure was 1.48 MPa, and hydrogen pressure was 1.54 MPa. The n-heptane/dimethylsulfide mixture was injected by a high pressure pump (Gilson). 2.3. Hydrogenation of tetralin The hydrogenation of tetralin was carried out at 300°C in two different flow reactors, both operating in a dynamic mode. In the first one, pure tetralin was introduced into the reactor
531 by means of a gas phase saturator, together with hydrogen and H2S. The total pressure was 4.5 MPa, the tetralin, H2S and hydrogen pressures being 2.6 kPa, 84.4 kPa and 4.41 MPa, respectively. In the second reactor, tetralin was diluted in a solvent (n-heptane) to which dimethyldisulfide was added in order to generate H2S, the mixture being injected by a high pressure pump. The reaction was carried out under a total pressure of 7 MPa. The different partial pressures were tetralin = 10 kPa, H2S = 240 kPa, n-heptane =1.57 MPa and hydrogen = 5.19 MPa. 2.4. Carbon analyses After the sulfidation, or after the tetralin hydrogenation experiment (6 hours), the spent catalyst was sent to the "Service Central d'Analyses CNRS" to have the carbon content measured. 3. RESULTS AND DISCUSSION 3.1. Reaction products The first difference between the two catalysts is found in the reaction products (figure 1). (1)
00 (3)
- _
(2)
00 '
oo'to Figure 1. Reaction scheme for the transformation of tetralin. Indeed, with the NiMo/alumina catalyst, the only products of the transformation of tetralin were decalins (1, hydrogenation products) and a small amount of naphthalene (4, dehydrogenation product). These products were also formed on the RuKYj catalyst, but together with significant amounts of CjoHig molecules (2), which resulted from the isomerization into methyl cyclopentanes of one or two of the cycles of decalins, or from the hydrogenation of small amounts of methyl indans (3) also observed. Small amounts of methyl and dimethyl tetralins (5) were also formed with the RuKY^j catalyst, when tetralin was converted in the presence of n-heptane and dimethyldisulfide. These molecules had the methyl groups on the aromatic ring and resulted most likely from a reaction between tetralin and the methylfragmentsproduced by dimethyldisulfide decomposition [10].
532
3.2 Overall activity The specific rates for the conversion of tetralin on the two catalysts are shown in Table 1. All the values were obtained after the catalysts had reached a steady state of activity. With the RuKYj catalyst, when the feed was pure tetralin, the steady state was observed after a slight deactivation, whatever the sulfidation method. When tetralin was used in solution in nheptane, the RuKY^ catalyst had a very stable activity when sulfided by H2S/H2, whereas an activation period was observed during the first 3 hours of reaction after sulfidation by dimethyldisulfide in n-heptane. The NiMo/alumina catalyst was always stable, whatever the sulfidation method or the feed. The resuhs in Table 1 show that the RuKY^j catalyst is much more active than the NiMo/alumina catalyst, under the same sulfidation conditions and with the same feed : for example, RuKY^i is 50 times more active than NiMo/alumina after sulfidation by H2S/H2, with a pure tetralin feed. This difference is even greater if the activity per atom is considered instead of the activity per gram, since there is about 5.4 times less Ru than Mo on the catalysts.
Table 1 Influence of the sulfidation method and of the reaction conditions on the rate of tetralin conversion on RuKY^j and on NiMo/alumina. Sulfidation Feed
H2S/H2 THN
DMDS/n-Cy
THN/n-Cy
THN/n-Cy
THN
RuKYd Activity %C before reaction %C after reaction
4.2
5.1
2.1
11.5
NiMo/alumina Activity
1.1
0.21
1.3
0.27
52
1.3
24
0.35 5.4
<0.1
Reaction rates in 10"^ mol.g"l.s"l. DMDS = dimethyldisulfide, THN = tetralin
3.3 Influence of the sulfldation method Table 1 shows that the RuKY^j catalyst is more active for the conversion of a given feed when it is sulfided by H2S/H2 than when it is sulfided by dimethyldisulfide in n-heptane. The value of the activity ratio is between 2.2 and 3.7, depending on the nature of the feed. On the other hand, this ratio is always close to 1 in the case of the NiMo/alumina catalysts. These results are certainly due to a deposition of carbon ("coke") on the RuKYj catalyst during the sulfidation in the presence of hydrocarbons (dimethyldisulfide in n-heptane). Indeed, we observed that this method of sulfidation left a 5.4 wt.% carbon deposit on the catalyst. Of
533 course, this was not the case with the sulfidation by H2S/H2 (less than 0.1 wt.% C). Since "coking" is generally initiated by the acid sites of the catalysts, it is quite normal to observe this phenomenon with a zeolitic catalyst, and not with the NiMo/alumina which has not a strong acidity. The "coke" molecules formed during the sulfidation most likely poison the catalyst by blocking part of the zeolite porosity, thus decreasing the number of accessible ruthenium sites. 3.4. Influence of the nature of the feed Whatever the sulfidation method, and whatever the catalyst, the conversion of tetralin is much higher when it is processed alone than when it is used in solution in n-heptane. When the catalyst was sulfided by H2S/H2, the amount of carbon deposited on the catalyst after reaction was roughly the same, whatever the feed (4.2% vs. 5.1%). However, the comparison between the amount of carbon deposited on the catalyst and the activity of the catalyst for tetralin conversion shows that the carbon/activity ratio obtained is about 4 with the tetralin/n-heptane mixture, whereas it is only 0.08 with pure tetralin. A very significant carbon deposition is observed with the tetralin/n-heptane feed when the catalyst is sulfided by dimethyldisulfide in n-heptane, since a 6 wt.% carbon increase is observed during the reaction (this increase is similar to the one observed when the catalyst was sulfided by H2S/H2). On the contrary, the amount of carbon remaining on the catalyst after sulfidation by the dimethyldisulfide/n-heptane mixture (5.4 wt.%) decreases during the reaction of pure tetralin (2.1 wt.%). Such an elimination of carbon during the reaction has already been invoked to explain the initial activation of the catalyst observed with the tetralin/n-heptane mixture [11]. In the present case, "coke" would be totally eliminated by pure tetralin, whereas it would be only partially eliminated by the tetralin/n-heptane mixture, perhaps due to the significant amount of n-heptane when compared to that of tetralin. It must also be pointed out that the carbon/activity ratio measured after pure tetralin conversion is 0.08, i.e. the same as the one measured after the H2S/H2 sulfidation. Consequently, it seems that the same amount of "coke" was formed during reaction of pure tetralin, whatever the sulfidation method. Lastly, it must be noted that the carbon/activity ratio is much higher with the tetralin/nheptane feed than with pure tetralin. It has been shown that the carbonaceous molecules deposited on the catalyst during reaction in the presence of n-heptane were saturated cyclic compounds formed only through tetralin condensation [11]. Consequently, this "coking", if due to tetralin conversion, would have been more significant with pure tetralin, which is not the case. This indicates that the lower activity of the catalyst for the conversion of tetralin in the presence of n-heptane is not only due to "coking", but also to a competition between tetralin and n-heptane for the adsorption on the active sites : indeed, it must be remembered that the nheptane partial pressure is 157 times greater than that of tetralin (see experimental part). In the case of the NiMo/alumina catalyst, where "coking" does not play an important role owing to the weak acidity of the catalyst, the activity for pure tetralin conversion is 5 times greater than the activity for the conversion of tetralin in the presence of n-heptane, whatever the sulfidation method. The difference in activity observed with the NiMo/alumina catalyst can simply be explained by a competition between tetralin and n-heptane for the adsorption on the active sites.
534 4. CONCLUSION The sulfided ruthenium in zeolite catalyst (RuKYj) is much more active for the hydrogenation of tetralin than a classical sulfided NiMo/alumina catalyst. However, its activity depends very much on the sulfidation method, which is not the case with the NiMo/alumina catalyst. The sulfidation by dimethyldisulfide in solution in n-heptane leaves a very significant amount of carbon on the RuKY^ catalyst, initiated by the zeolite acid sites. This causes a partial blocking of the porosity, resulting in a decrease in the catalytic activity. Lastly, the presence of a solvent (n-heptane) during the tetralin hydrogenation experiments decreases the activity of both catalysts. This decrease can be explained by a simple competition between tetralin and n-heptane for the adsorption on the active sites. 5. ACKNOWLEDGEMENTS This work was carried out within thefi^ameworkof the research program " Aromatics hydrogenation", with the financial support of ELF, IFP and TOTAL, and of the CNRS ECOTECH.
6. REFERENCES 1 2 3 4 5 6 7 8 9 10 11
H.J. Lovink, Symposium and Tutorial on Impact of Reformulated Fuels, 208th National Meeting, American Chemical Society, Washington, 1994, p. 513. T.A. Pecoraro and R.R. Chianelli, J. Catal., 67 (1981) 430. M. Ledoux, O. Michaux and G. Agostini, J. Catal., 102 (1986) 275. M. Lacroix, N. Boutarfa, G. Guillard, M. Vrinat and M. Breysse, J. Catal., 120 (1989) 473. S. Eijbouts, V. de Beer and R. Prins, J. Catal., 109 (1988) 217. Z. Vit and M. Zdrazil, J. Catal., 119 (1989) 1. T.G. Harvey and T.W. Matheson, J. Chem. Soc, Chem. Comm. (1985) 188. T.G. Harvey and T.W. Matheson, J. Catal., 101 (1986) 253. B. Moraweck, G. Bergeret, M. Cattenot, V. Kougionas, C. Geantet, J.L. Portefaix and M. Breysse, J. Catal., under press. J. Leglise, A. Janin, J.C. Lavalley and D. Comet, J. Catal., 114 (1988) 388. M. Mhaouer, J.L. Lemberton and G. Perot, Catalysis Today, 29 (1996) 241.
® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
535
CATALYSTS OF PHOSPHOTUNGSTIC OR PHOSPHOMOLYBDIC ACIDS ON DIFFERENT SUPPORTS FROM DIMETHYLFORMAMIDE SOLUTIONS L. R. Pizzio, P. G. Vazquez, M. G. Gonzalez, M. N. Blanco, C. V. Caceres and H. J. Thomas. Centro de Investigacion y Desarrollo en Procesos Cataliticos (CINDECA), UNLP, CONICET, Calle 47 N° 257, 1900 - La Plata, Argentina. Catalysts PW(Mo)Ni/Al203(Si02) were prepared from phosphotungstic or phosphomolybdic acids in dimethylformamide solutions. The characterization of the species present in the solutions and in the catalysts is shown. The performance of the catalysts for hydrodenitrogenation (HDN) and hydrodesulfiirization (HDS) reactions are related to the precursor-support interaction and to the shape of concentration profiles in the pellets. 1.- INTRODUCTION The importance of the petroleum hydrotreating processes is in relationship to the objective of producing clean-funning fuels, this objetive has carried to a reappearance in recent years of the investigation on the chemistry and the engineering of hydrotreating reactions. Many of the papers are related to the characterization of the catalysts by physical methods and to the study of reactions of compounds as thiophene, at low pressure, on supported catalysts as CoMo, NiMo or NiW, using as support generally alumina, frequently doped with additives, such as F [1] or P [2]. Fundamentally, the sulfurized form of the hydrotreating supported catalysts is studied. However in these last times, the performance of these catalysts has attempted to be improved by applying new preparation methods, for example, replacing the conventional precursors, employing different metals from the usually used promotors, adding a third element to the catalysts or adopting new supports that replace y-alumina. Many researchers have studied the use of transition metals as Ni [3] and Zn [4] among others [5-7], as promotors or additives to the support to modify the performance of the conventional catalysts Co-Mo/alumina or Mo/alumina. Lee et al. [8] reported that the addition of small amounts of W into a commercial CoMo/alumina catalyst brought about advantageous effects on the catalyst performance for asphaltene HDS and on the conversion of heavier to lighter oil fractions during the high pressure hydrotreatment of an atmospheric residual oil feed. Lee et al. [9] studied a series of catalysts with different W content and other series was prepared by changing the impregnation order of Co and W. They accomplished the characterization of the catalysts series previously studied. If W is incorporated in a content of 0.025, expressed as W/(W+Mo) atomic ratio, the dispersion of M0S2 in sulfided catalysts was considered maximum, resulting in the maximum catalytic activity promotion. In order to obtain new generations of HDS and HDN catalysts, our group has begun studies some years ago. In a first instance the objetive of these researchs was to incorporate
536 simultaneously P and Mo on the alumina, by impregnation of P2Mo5023^", which is the species present in the solution of phosphoric acid (PA) and ammonium heptamolybdate (AHM) with P/Mo=0.4 molar ratio [10]. Thereinafter, the adsorption of Mo and P integrating the same structure was carried out using heteropolyacids (HPA), being the most prominent those that present primary structure of the Keggin type [11], with a general formula [XMi204o]^^""^ where n is the valency of X. The structure is formed by a central tetrahedron of XO4 surrounded by twelve octahedra of MOe. The central atom or heteroatom X can be either P, As, Si, Ge or B while most peripheral atoms, called poly or addenda atoms, are Mo or W. The HPA catalytic activity is directly related to the nature and subsequent stability of adsorbed species in the impregnation stage. So, it was studied the nature of the species present in solution and on the solid for the phosphotungstic acid from ethanol-water solutions and different supports, as silica, titania, carbon and different aluminas [12]. For the phosphomolybdic acid it was studied, among other variables, the behavior of species so much from solutions in organic solvents as in water and for alumina as for different supports [13]. In this work it is presented a new method of hydrotreating catalysts preparation, for HDS as well as for HDN, where phosphotungstic and phosphomolybdic acids are dissolved in dimethylformamide, using alumina and silica as support. It is presented a detailed analysis of the nature of the present species in the impregnating solutions as well as the characterization of the present species in the obtained catalysts. The activities of different catalysts prepared are measured, using thiophene for HDS and pyridine for HDN as test molecules. The influence on these catalytic activities of the active phase distribution in the pellets is studied. 2.- EXPERIMENTAL Materials. The supports used were commercial spheres of y-Al203 and Si02. The anal5^ical reagents used were phosphomolybdic acid ^PMo/^ MERCK pa H3PMoi204o.5^IA phosphotungstic acid (PW/^ HDHpa H3PWi204o.?i^ and dimethylformamide (IMF) ANEI»A Preparation of Catalysts. PMo(W)Ni/y-Al203(Si02) catalysts were prepared by means of pore filling two step impregnation: first, PMoA^WA) acids were impregnated on the alumina (silica) by employing DMF solutions. In the second step and after drying the catalysts, these were impregnated with DMF solution of nickel nitrate. Solids thus obtained were dried at room temperature during 24 h and then calcined at 350°C during 1 h. After acid dissolution of the solids. Mo, W and Ni contents were analyzed by atomic absorption spectrometry; an IL457 model double beam spectrophotometer (Instrumentation Laboratory Inc.) was used. Mo and W contents in the different catalysts were in the range of 6 10 to 11 10 mol metal/g catalyst and the Ni contents were in the range of 4 10 to 5 10 mol Ni/g catalyst. The distribution of Mo, W and Ni along the sphere radius was obtained by electron microprobe analysis according to a previously described technique [14]. NMR spectroscopy. The solutions of both Mo and W were analyzed by ^^P NMR. A Bruker MSL-300 equipment linked to a "SOLIDCYC.DC" pulse program, operating at afrequencyof 121.496 MHz for ^^P at 25°C was utilized. PA 85% was used as external reference. DR spectroscopy. The DRS study of the solid samples was made with an UV-visible Varian Super Scan 3 spectrometer, fitted with a diffijse-reflectance chamber with inner surface of BaS04. Spectra of wet and calcined samples were registered in the range of 200-600 nm, yalumina and silica being used as references.
537 FTIR spectroscopy. A Bruker IFS 66 equipment, pellets in BrK and a measuring range of 400-1500 cm"^ were used to obtain the spectra of silica, bulk acids and calcined samples. Catalytic activities. The activity measures in HDS and HDN reactions were accomplished in a fixed bed reactor, at 3 MPa, 280°C, a H2/HC molar ratio=5 and LHSV=53 K\ The catalysts in sphere form and ground samples were sulphurized in situ at 280°C. The liquid feed was a mixture of ciclohexane-pyridine(2000 ppm N2)-thiophene(6000 ppm S). The activity was determined by following the conversion degree of thiophene and pyridine, using GC analysis. 3.- RESULTS AND DISCUSSION Figure la shows the ^*P NMR spectra of the initial PWA solution in DMF (I) and those of the solution after the impregnation on Y-AI2O3 (II) and after the unpregnation on Si02 (III). The spectra corresponding to PMoA solutions are shown in Figure lb. Only the chemical shifts of ^^P in phosphotungstic and phosphomolybdic anions (-15.6 and -4.1 ppm, respectively) are observed. These results indicate that the species present in all solutions are the Keggin phases unaltered, for both the H3PWi204oand H3PM012O40 solutions in DMF. So, the stabilizing effect of non-aqueous solvents was observed. -4.1
\
-11
-13
-15
-17
ppm
-19
m_
JU
2.0 0.0 -2.0 -4.0 -(.0 -1.0
ppm
Figure 1. NMR spectra of PWA (a) and PMoA (b) solutions in DMF. I) initial solution, II) solution after contact with alumina and III) after contact with silica. FTIR spectrum of bulk H3PW12O40 (Figure 2a) shows bands at 1081, 982, 888, 793, 595 cm"\ which coincide with those referred to in the literature for this acid. Also by FTIR on the PWNi/Si02 catalyst, it was found that the [PWi204o]^" anion is present because, as it can be seen in Figure 2a, the spectrum corresponding to this catalyst shows the bands at 982 and 888 cm"\ The bands at 1081 and 793 cm"^ are masked by the silica bands, meanwhile the band at 595 cm"^ has a low intensity. The shifting of the band at 982 to 959 cm"^ and the appearance of a shoulder at 730 cm"^ in the spectrum of this catalyst should indicate a partial transformation into the [PWii039]^" species. With regard to FTER spectra of bulk H3PM012O40, which is shown in Figure 2b, bands at 1066, 961, 872 and 782 cm"^ are observed. In this figure it is also observed the spectrum of the catalyst prepared from this heteropolyacid on silica, which presents a band at 943 cm"^ and a
538
1300 900 500 Wavenumbcr (cm"*)
Wavcnumbcr (cm )
Figure 2. FTIR spectra of bulk heteropolyacid (I), silica (II) and calcined catalyst (III) obtained from PWA (a) and PMoA (b). shoulder at 895 cm'^ while the other bands are masked by the support bands. The shifting of the band maxima in the catalyst PMoNi/Si02 with respect to that in the bulk acid is attributed to partial transformation of [PMoi204o]^'species into [PMonOsp]^" anion. The above-mentioned FTIR results were also confirmed by DRS. Figures 3a and 3b show the DRS spectra of the wet and calcined catalysts supported on silica and prepared from PWA and PMoA, respectively. The wet samples show the characteristic band of the species with undegraded Keggin structure. This band is less extended in the calcined samples. This indicates that the Keggin phase is partially degraded. DRS spectra of PWNi/y-Al203 and PMoNi/y-Al203 catalysts (Figures 3a and 3b) indicate that the Keggin unit undergoes a greater depolymerization than in the catalysts on silica, both at calcination temperature of the samples (350<^C).
300
X(nm)
400
500
200
300
X(nin)
400
Figure 3. a): DRS spectra of wet (I) and calcined (III) samples of PWNi/y-Al203 and wet (II) and calcined (IV) samples of PWNi/Si02. b): idem for PMoNi/y-Al203 and PMoNi/Si02.
539
4
•
• •
s
A
o
A 2
A
A
II S
T * •—f~-«—i f •—#-f—•-A -1,0
-0,5
0,0 r/R
• • i ^ •
A 4
0,5
1,0 -1,0
S
-+-
-0,5
ft
i i"
M
0,0
-+-
0,5
B *
p
1,0
r/R
Figure 4. W, Mo and Ni concentration profiles in PWNi/y-Al203(Si02) (a) and PMoNi/y Al203(Si02) (b) catalysts: •:W(Mo)/ SiOa, A:W(Mo)/yAl203, •:Ni/Si02and •:Ni/yAl203 The distribution of W(Mo) and Ni along the y-Al203(Si02) sphere radius was obtained by electron microprobe analysis. The profiles are shown in Figures 4a and 4b, for catalysts prepared fi-om PWA and PMoA, respectively. The PWNi/Si02(y-Al203) catalysts show W profiles where the concentration diminishes slowlyfi'omthe surface to the center of the sphere. On other hand. Mo concentration in PMoNi/y-Al203 catalyst showed an egg shell profile, this have a high concentration on the surface and low values in the rest of the pellet. In PMoNi/Si02 catalyst the Mo concentration profile is almost flat. All these profiles are fimdamentally the result of different adsorption strength of the precursor on the respective support. PMoA-alumina interaction is greater than the corresponding to PMoA-silica and PWA-alumina. For all catalysts, Ni concentration profiles are nearly constant along the pellet. The HDN and HDS specific activities, expressed as conversion/metal concentration, of PW(Mo)Ni/y-Al203(Si02) catalysts prepared are shown in Figures 5a and 5b, respectively. The HDN activity of the ground catalysts prepared fi'om PWA as well asfi'omPMoA is greater for the catalysts supported on silica than those on alumina. This could be due to the smaller interaction of PWA and PMoA with silica as compared to those with alumina. This would involve that during the sulphurization the active phase would probably be more easily formed in the catalysts supported on silica. Similar results were obtained for the HDS activity of the mentioned catalysts. Comparison of the catalysts HDN activity determined in pellets and in ground samples allows to observe that the catalysts PMoNi/y-Al^Og is more active in pellets than in ground samples. This is due to that this catalyst presents an egg shell profile and that the diffusion of reactant towards the pellet interior controls the rate of the HDN reaction of pyridine. As a consequence of the last process and of the Mo or W profiles of the catalysts PWNi/y-Al203 and PW(Mo)Ni/Si02, which are ahnost flat, the HDN activity of the pellets is smaller than that of the ground samples. The thiophene HDS activity of the catalysts supported on alumina determined in pellets or in ground samples are similar. This implies that in these catalysts the HDS reaction is not controlled by the diffusion of the reactant. However, for the catalysts supported on silica the activity in pellets is smaller than in the ground samples. This is a consequence of that in these catalysts, thiophene diffiisional control exists, probably due to the small pore diameter of silica compared with the corresponding to alumina, and that W and Mo profiles are almost flat.
540
Si gro
Si sph
Al gro
Al sph
Figure 5. HDN (a) and HDS (b) specific activities of PWNi/y-Al203(Si02) and PMoNi/yAl203(Si02) catalysts, sph: spheres, gro: ground samples. Taking into account all results above-mentioned, a direct relation between the species present in the catalysts and the HDN and HDS activities was not found. It appears that the Keggin phases, both fi-om PWA and PMoA, even degraded in different extents on silica or alumina, have a similar catalytical behaviour. Instead, these activities fundamentally depend on the PMoA^Wi^upport interaction and on the shape of the concentration profiles. Besides, it can be concluded that the impregnation of heteropolyacids from dimethylformamide solutions provides an adequate way for obtaining catalysts with different profiles and performances. ACKNOWLEDGEMENTS The authors thank L.Osiglio, G. Valle, D.Peiia and F.Iborra for their experimental contribution. REFERENCES 1.- Ch.Papadopoulou, A.Lycourghiotis, P.Grange and B. Delmon, Appl. Catal. 38 (1988) 273. 2.- CPoulet, R.Hubaut, S.Kasztelan and J.Grimblot, Bull.Soc.Chim.Belg. 100(11) (1991) 857. 3.- KRodriguez, C.Caceres, M.Blanco and H.Thomas, Adsorp. Sci. Tech. 4 (3) (1987) 162. 4.- H.Thomas, C.Caceres, M.Blanco, J.L.G.Fierro and A.L6pez Agudo, J.Chem.Soc.Faraday Trans. 90 (14) (1994) 2125. 5.- ASaini, B.Johnson and F.Massoth, Appl. Catal. 40 (1988) 57. 6.- P.Mitchell and C.Scott, Bull. Soc. Chim. Belg. 93 (1984) 619. 7.- M.Teman, J. Catal. 104 (1987) 256. 8.- D.Lee, LLee and S.Woo, Appl. Catal. A: Gen. 109 (1994) 195. 9.- K.Lee, H.Lee, HLee, S.Park, S.Bae, S.Kim and S.Woo, J. Catal. 159 (1996) 219. lO.-P.Vazquez, C.Caceres and M.Blanco, XVI Simp. Iberoamer. Catal., Chile (1994) 1347. ll.-J.Keggin, Proc. R. Soc. London, Ser A, 144 (1934) 75. 12.- L.Pizzio, C.Caceres and M.Blanco, XVII Simp. Iberoamer. de Catal., Argentina (1996). 13.- M.Castillo, P.Vazquez, M.Blanco and C.Caceres,!. Chem Soc Faraday Trans., in press. 14.- P.Vazquez and J.Riveros, X-Ray Spectr. 21 (1992) 197.
® 1997 Elsevier Science B. V. All rights reserved, Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
541
Infrared study on the acid sites of nitrided molybdena-alumina catalysts Masatoshi Nagai, Osamu Uchino, Takuya Kusagaya, and Shinzo Omi Department of Applied Materials, Graduate School of Bio-applications and Systems Engineering, Tokyo University of Agriculture and Technology, Koganei, Tokyo 184, Japan The acid sites of the nitrided 12.5% M0/AI2O3 catalysts have been studied using a diffuse reflectance FTIR spectroscopy. The activities of the nitrided catalysts for the hydrodesulfurization (HDS) of dibenzothiophene at 573 K and at 10.1 MPa were also determined. The nitride catalysts were prepared by the temperature-programmed reaction of 12.5% M0O3/AI2O3 with NH3 from 573 K to 773 K, 973 K, and 1173 K at 1 Klmin and then kept at each temperature for 3 h. On the basis of the FTIR spectroscopy, it was found that Lewis acidity was predominant on the surface of the nitrided catalysts. The Lewis and Br0nsted acidities were correlated with the HDS activity of the catalysts for the HDS of dibenzothiophene. 1. INTRODUCTION An increasing interest has developed in exploring the catalytic properties of Mo nitrides. Mo nitride is an active catalyst for various reactions to compete with the catalytic properties of noble metals. Unsupported Mo nitride has been reported to be active for CO hydrogenation, NH3 synthesis, and quinoline HDN (1-3). Recently, Markel and Zee (4) studied the HDS of thiophene on an unsupported Mo nitride and reported that the Mo nitride promoted more hydrogenation than the breakage of the C-S bond of thiophene. In contrast to the efforts devoted to studies of catalytic activity, much less attention has been given to examining the surface properties of Mo nitrides supported on alumina, compared to the unsupported Mo nitrides for industrial applications. Particular attention has been paid to identifying the acid sites on the surface of the nitrided M0/AI2O3 catalysts. In this study, DRIFT spectra of NH3 are taken for the 12.5% M0/AI2O3 nitrided at 773 K, 973 K, and 1173 K to determine the nature of acid sites on the surface of the nitrided M0/AI2O3 catalysts. The relationship between the surface acidities and the activities of the nitrided catalysts for the HDS of dibenzothiophene is discussed. The mechanism of the dibenzothiophene HDS is proposed on the basis of the combination of catalytic activity with the IR data of the catalysts.
542 2.
EXPERIMENTAL
2.1.
Materials and catalyst preparation Hydrogen, argon, and helium (99.999%) were dried by passing them through a Deoxo unit (SUPELCO Co. Oxysorb) and a Linde 13X molecular sieve trap prior to use for characterization. Hydrogen (99.9%) was dried by passing it through a Linde 13X molecular sieve trap for activity measurement. Other gases were used without further purification. For activity measurement, dibenzothiophene (ultra-extra pure) and xylene (mixed, extra pure, Kishida Chem. Co.) were used without further purification. A 12.5% M0O3/AI2O3 was prepared using a mixture of ammonium paramolybdate and 7 -alumina and calcined in air at 823 K for 3 h (Nikki Chemicals Co.). The 12.5% M0O3/AI2O3 pellets were crushed and sieved to 10-20 mesh granules for activity measurement and to ground for FTIR measurement. The catalyst was held in place by means of a fritted quartz disk of 10 mm o.d., attached with both sides of 50 mm long quartz tubing of 6-mm o.d. Nitriding of the 12.5% M0O3/AI2O3 was carried out by temperature-programmed reaction with NH3. The 12.5% M0O3/AI2O3 precursor was oxidized at 723 K for 4 h, cooled to 573 K, and then treated in flowing pure NH3 at 4 liters/h, from 573 K to 773, 973, and 1173 K at a rate of 1 K/min and held for 3 h at 773 K (LTN), 973 K (MTN), and 1173 K (HTN). The catalysts were purged in flowing He for 1 h at the nitriding temperature, cooled to room temperature inflowingHe, and then passivated in 1% 02/He gas. To compare the catalytic properties of the nitrided catalysts with the sulfided catalysts, the fresh catalyst was presulfided in flowing 10% H2S/H2 at 573 and 773 K for 3 h. The specific surface area of the nitrided catalysts was measured by nitrogen adsorption using a standard BET volumetric apparatus after the catalysts were evacuated at 473 K and 1 Pa for 2 h. The BET surface area of the nitrided catalysts are shown in Table 1. 2.2. FTIR study Diffuse reflectance FTIR (DRIFT) spectra of the ground nitrided M0/AI2O3 catalysts were recorded on a DRIFT Instrument (Nicolet, Model 740), together with a liquidnitrogen-cooled MCT detector. Before the IR measurement, the catalyst was reduced at 773 K for 2 h to remove water. The DRIFT cell in a flow system was equipped with a Harrick diffuse reflectance accessory fitted with a controlled environmental chamber. DRIFT spectra of the nitrided 12.5% M0/AI2O3 were presented in Kubelka-Munk units and recorded at room temperature under atmospheric pressure. DRIFT spectra were recorded (128 scans, 4 cm"^ resolution) before and after NH3 was admitted to the chamber, which was heated to the desired temperature, 2.3. Measurement of HDS The HDS of dibenzothiophene on the nitrided 12.5% M0/AI2O3 was studied using a high-pressure microreactor at 553 K and at 10.1 MPa. For the activity measurement, 2.0 g of the nitrided catalyst (granule) was packed in the stainless steel reactor, as described in detail elsewhere (5). The liquid feed, consisting of 0.25wt% dibenzothiophene dissolved in xylene, was introduced into the reactor at 20 ml/h with a He flow of 6 1/h. The HDS rate of the M0/AI2O3 catalysts was calculated, based on the rate for disappearance of dibenzothiophene.
543
3.
RESULTS AND DISCUSSION
3.1. DRIFT spectra Figure 1 shows the spectra of the surface species of the nitrided M0O3/AI2O3 catalysts after dosing with 5% NH3 in He at 553 K. The catalyst reduced at 673 K and alumina are also shown for comparison (6-8). The band at ca. 1260 cm'^ for the LTN, reduced, oxidized catalysts was assigned to the bending modes of coordinated NH3, These bands are ascribed to the presence of surface Lewis acid sites (surface-bonded species of NH3). However, a very weak band at 1480 cm"^ due to the deformation vibrational mode of NH^"^ was observed, showing the presence of surface Bronsted acidity on the nitrided and reduced M0/AI2O3 catalysts. Very small proportion of the Bronsted acid sites was present on the catalysts. The spectra for the MTN and HTN catalysts were lower in intensity than those for the oxidized and reduced M0/AI2O3 catalysts. Small peaks were observed for the nitrided alumina. The Lewis acid sites were predominate on the Mo nitride catalysts. Figure 2 shows the changes in the IR spectra corresponding to the Lewis acidity (ca. 1260 cm'^) and the Bronsted acidity (1480 cm"^) as a function of nitriding temperature. The Lewis and Bronsted acidities decreased with increasing nitriding temperature. Therefore, nitridation of the 12.5% M0/AI2O3 above 973 K lowered the surface Br0nsted and Lewis acidities of the catalysts. Nitriding of the 12.5% M0/AI2O3 catalyst with NH3 was accompanied by the creation of a large number of Lewis acid sites. Less Lewis and Brensted acid sites were created on the HTN catalyst than on the LTN catalyst. L
3.2. HDS of dibenzothiophene The major product was biphenyl alone with small amounts of tetrahydrodibenzothiophene, cyclohexylbenzene and bicyclohexyl. In previous papers (9,10), the network of dibenzothiopheneHDS on sulfided and nitrided M0/AI2O3 catalysts involves hydrogenation as well as Sremoval reactions, as shown in Figure 3. In this scheme, biphenyl was formed by the direct sulfur removal from dibenzothiophene, and cyclohexylbenzene was produced during the hydrogenation of dibenzothiophene through hexahydrodibenzothiophene. The scheme of dibenzothiophene HDS also takes place on the Mo nitride catalysts. Figure 4 shows the
nitrided at 1173K
a
2030
1830 1630 1430 1230 Wavenumber/cm'^
1030
Figure 1. Diffuse reflectance FT-IR spectra of NH3 adsorption on 12.5wt% M0/AI2O3 at 553 K.
544
0.1
3.5 3
Bronsted acidity
0.08
2.5 0.06
2 1.5
0.04
Lewis acidity
1 o^
0.02
0.5 700
Figure 2. Changes in the IR spectra of NH3 on the 12.5wt% M0/AI2O3 catalyst corresponding to Bronsted acidity (1480 1200 cm*^) and Lewis acidity (ca. 1260 cm'^) as a function of nitriding temperature.
800 900 1000 1100 Nitriding temperature / K
1.2.3.i|.i»a.9a-Hexahyclrodibenzothiophene
mo—
/i^.2.3.4-Tetrahydro dibenzothiophene Dibenzothiophefie
\
Cyclohexylbenzene
lCg0)-O-O PerhydroBlcyclohexyl dibenzothiophene
**' (C)]—(T)] Biphenyl
Figure 3. The network of the hydrodesulfurization of dibenzothiophene. TABLE 1 Catalyst Activity for HDS of dibenzothiophene at 553 K Catalyst
Surface HDS rate^ C-S hydrogenolysis Area^( m^g'^) ii mol/gh /Hydrogenation''
LTN MTN HTN LTS HTS
226 195 138 ~ —
97.6 66.6 40.9 81.5 42.1
4.7 4.3 7.2 1.5 1.7
^ evacuated at 473 K and at a Pa. ^HDS activity is based on the rate for disappearance of dibenzothiophene. ^ calculated by the molar ratio of biphenyl to cyclohexylbenzene after the HDS of dibenzothiophene starts in 10 h. 1
J
••'
"•
545 HDS activity of the catalysts nitrided at various nitriding temperatures in addition to those of the sulfided catalysts at 553 K. The activity of the MTN catalyst was the highest, while the LIS catalyst second. The HTS catalyst showed extremely low activity for the HDS. The HDS activity of the M0/AI2O3 catalysts decreased with increasing nitriding temperature. Furthermore, the selectivity of the catalyst for desulfurization to hydrogenation was evaluated by the molar ratio of biphenyl to 700 800 900 1000 1100 1200 cyclohexylbenzene as shown in Nitriding temperature / K Table 1. The selectivity of the MTN catalyst for the direct Figure 4. Change in HDS rate for dibenzothiophene desulfurization is about 3 times HDS on the catalysts nitrided at three nitriding greater than that of the sulfided temperatures. catalyst at 553 K. As a result, the nitrided catalyst was significantly active toward direct sulfur removal from dibenzothiophene with less consumption of hydrogen. This result was in disagreement with the data obtained by Markel and Zee, in the HDS of thiophene at atmospheric pressure (4). 3.3. Active sites for HDS of dibenzothiophene Figure 5 shows the rate for dibenzothiophene HDS is plotted against the IR spectra peak area of Bronsted and Lewis acidity. The 40 60 80 100 120 HDS rate linearly decreased with HDS rate /fimol h'^ g'^ increasing number of the Lewis and Bronsted acid sites. This Figure 5. The rate for dibenzothiophene HDS result showed that the HDS plotted against the IR peak area of ( • ) NH3 activity of the nitrided catalyst Lewis acidity and NH3 Bronsted acidity ( • ) . was correlated to the Lewis and Br0nsted acid sites. From the results, the mechanism of HDS of dibenzothiophene on the Mo nitrides is
546 proposed. The HDS of dibenzothiophene needs both Lewis acid sites (Mo sites) and Bronsted acid sites. The electron-rich sulfur atom of dibenzothiophene is adsorbed on more reduced Mo unsaturated coordinated Mo sites (Lewis acid sites). This step is likely to be a rate-controlled step during the reaction. The C-S bond is weakened by the adsorption of dibenzothiophene on the Mo site, and then broken by a transfer of hydrogen from an adjacent hydroxyl group (Bronsted sites). The C-S bond fission precedes ring hydrogenation. The sulfur on the Mo atom is hydrogenated to form H2S and/or is accumulated on the surface to form M0S2. 4. CONCLUSIONS (1) The activity of the 773 K-nitrided catalyst was the highest, the 623 K-nitrided catalyst second. The HDS activity of the M0/AI2O3 catalysts decreased with increasing nitriding temperature. The 773 K-sulfided catalyst is 0.7 times less active than the nitrided catalysts, even though the 623 K-sulfided catalyst was active during the HDS of dibenzothiophene. (2) The Mo nitride catalyst was extremely active for the selective C-S bond breakage of dibenzothiophene to produce biphenyl. (3) Nitriding of the 12.5% M0O3/AI2O3 catalyst with NH3 at high temperature lowered a number of Lewis and Br0nsted acid sites. (4) The IR data showed that the HDS activity of the nitrided catalysts was correlated to the Lewis acid sites (reduced unsaturated coordinated molybdenum) and Bronsted acid sites on the surface of the catalysts. REFERENCES 1. G. S. Ranhotra, A. T Bell, and J. A. Reimer, J. Catal., 108 (1987) 40. 2. L. Volpe and M. Boudart, J. Phys. Chem., 90 (1986) 4874. 3. J. C. Schlatter, S. T Oyama, J. E. Metcalfe, III, and J. M. Lambert, Jr., Ind. Eng. Chem. Res., 27 (1988) 1648. 4. E. J. Markel and J. W. Van Zee, J. Catal, 126, (1990) 643. 5. M. Nagai, T. Masunaga, and N. Hanaoka, J. Catal., 101 (1986) 284. 6. J. Yalyon, R. L, Schneider, and W. K. Hall, J. Catal., 85 (1984) 277. 7. M. C. Kung and H. H. Kung, Catal. Rev-Sci. Eng., 27 (1985) 425. 8. R R Groff, J. Catal, 86 (1984) 215. 9. M. Nagai, T Sato, and A. Aiba, J. Catal, 97 (1986) 52. 10. M. Nagai, T Miyao, and T Tsuboi, Catal. Lett., 15, 18 (1993) 9.
® 1997 Elsevier Science B, V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
547
Synthesis and characterization of zirconia-alumina mixed oxides F. Dumeignil«, P. Blanchard«, E. Payen«, J. Grimblot«* and 0. Poulet^ «Laboratoire de Catalyse Heterogene et Homogene, Batiment C3, USTL / CNRS URA 0402, 59655 Villeneuve dAscq Cedex, France ^Centre de Recherches TOTAL-FRANCE, Gonfreville TOrcher, 76700 Harfleur, France The incorporation of zirconium in alumina has been studied by several methods : sol-gel synthesis, impregnation and grafting. Sol-gel synthesis was performed by hydrolysis of a mixture of zirconium and aluminium alkoxides with subsequent drying and calcination. Impregnation of alumina was performed with various solutions in which the zirconium salt was dissolved. Different salts were used such as ZrOCk, ZrO(N03)2, and a commercial zirconium polymer solution. Grafting by a controlled hydrolysis of a zirconium alkoxide by the alumina surface hydroxyl groups was also performed. The effect of the incorporation of Zr on the texture and the structure of the soUds was studied by BET surface area measurements, XPS and XRD. Depending on the preparation method, quite different types of dispersion were evidenced.
1. INTRODUCTION Improvement of the Co(Ni)-Mo-Al203 HDS catalysts has been driven by the need to produce clean fuels, based on the pressing requirement for environmental protection. Although the alumina is the most widely used support, in the last few years much attention has been paid to the beneficial role of other supports. Indeed, other soUds such as Ti02 [1] and Zr02 [2] have been proposed as carriers of the active CoMoS phase. Introduction of additives to alumina such as phosphorus [3] or fluorine [4] has also been proposed to enhance the activity of the classical C0M0/AI2O3 catalysts. Some studies have revealed that zirconia could be an interesting alternative to alumina. For example, for Mo-Zr02 based catalysts, the intrinsic activity was higher [5,6]. Nevertheless, this oxide support has very low textural properties [7] compared to alumina namely in terms of specific surface area and mechanical stabiUty. In order to improve these properties, the preparation of zirconia-alumina supports has already been investigated by several authors [8-11]. * To whom correspondance should be addressed
548 In this work, we studied various methods of preparation of a mixed Zr-Al oxide i.e. sol-gel preparations, dry impregnation with aqueous Zr solution and grafting with Zr alkoxides of alumina surface. The solids have been characterized by several techniques such as X-ray diffraction (XRD), X-ray microprobe, laser Raman spectroscopy (LRS) and X-ray photoelectron spectroscopy (XPS) in order to precise the nature and location of the Zr entities.
2. EXPERIMENTAL 2.1. Preparations 2.1.1. Sol-gel preparations
The alkoxide used were 70% zirconium-n-propoxide in n-propanol (ZrP : Fluka Chemika), 78% zirconium-n-butoxide in n-butanol (ZrB : Strem Chemicals) and aluminium-tri-sec-butoxide (ASB : Merck). Solvent 85°C
Vsolvent =4 Valkoxides
Addition ASB Addition Zf alkoxide
^M^f}^'^^WS.^ Hydrolysis
1
Stirring 3h 85°C
|
Freeze drying
|
[Water] = 10 [ASB] [Water] • = [Zr alkoxide]
5
Calcination 500Xor450X air or oxygen Carried out in ^^loye-box^
Figure 1. Preparation of sol-gel non-chelated samples (Procedure I) As described in Figure 1, mixed zirconia-alumina powders were prepared from Zr alkoxides (ZrP or ZrB) and ASB in the parent alcohol of the zirconium alkoxide under inert atmosphere. It is well known that the Zr alkoxides are more sensitive to hydrolysis than Al alkoxides [12] and upon dissolution of ASB in
549 n-propanol, an alcoholysis reaction occurs. The modified Al alkoxide product should be more sensitive to hydrolysis [13]. The required amounts of ASB and zirconium alkoxide were successively added to the solvent and maintained at 85**C under reflux for 1 hour. The solution was then hydrolysed with distilled water under reflux for 3 hours at 85°C. A gel more or less clear is thus obtained. As the hydrolysis properties of alkoxides of Al and Zr are different, prehydrolysis for 1 hour at 85 ®C of the ASB was also performed before the addition of the Zr npropoxide. The use of a chelating agent such as P-diketones has often been proposed [14] to modify the textural properties of the materials. Therefore, preparations with acetylacetone (acac) have been investigated according to the process schematically represented in Figure 2. Each alkoxide was stirred with the chelating agent for one hour before being used in the preparation (sample 10P-P(H)). acac 2-butanol 85°C
Vsolvent = Valkoxides
4
|
Addition ASB Stirring 1 h 8 5 X
| Zr n-propoxide + acac, RT
Partial hydrolysis [H201/[ASB]=1
[Water] = 10 (total) [ASB] [Water] [Zr n - propoxide]
Stirring 1 h 8 5 X
|
Stirnngir>gyC.KT,
Hydrolysis (remaining H2O) I Stirring: 1 h 8 5 X , 1 h R T I Freeze drying Calcination 500°Cor450X air or oxygen
Carried out in
Figure 2. Preparation of chelated sol-gel samples (Procedure II) Preparations with n-propanol as the solvent were attempted. They did not give a gel but a solution from which aluminium acetylacetonate (identified by IR spectroscopy), precipitated after about one day, so further preparations using
550
acac were carried out in 2-butanol and clear gels were obtained within half an hour. It has been shown that freeze drying was the more efficient drying method for this kind of preparation because it allows the lowest carbon content on the calcined samples [15]. So, whatever the method of preparation, the solvent was then separated from the gel by freeze-drying. The nomenclature of these so obtained sol-gel carriers is detailed in Table 1.
2.1.2. Impregnations
The solids as extrudates or powders were prepared by incipient wetness impregnation of a commercial y-alumina (pore volume : 1 cm^.g*^ specific surface area : 350 m^.g-i) with solutions containing the appropriate amount of zirconium salt. The Zr loading **X" will be expressed hereafter as wt% of Zr02. The commercial alumina was bruised in a mortar and sieved to 50 ^m. Dry impregnation of the powder was performed with "Bacote 20'*, a commercially available zirconia precursor which is a solution containing 20% wt% Zr02 (Magnesium Elektron). Two types of impregnation were performed, one with a solution at its natural pH and another one in basic medium with ethylene diamine (En) as this solution is stable at a basic pH. Impregnated alumina was kept 2 hours without exposure to air (maturation) before drying at 100 ®C overnight. The nomenclature of the so obtained soUds recalls the preparation procedure i.e. IiX or I2X for samples prepared respectively without or with En in the impregnating solution. Owing to the slow diffusion of the large species present in these zirconia based solutions, impregnation of extrudates was performed in order to check the mass transfer in the pores (macrodispersion). Three different precursors of Zr were used to prepare the solutions, i.e. zirconyle chloride or nitrate and bacote 20. The two former have a Umited solubility so different agents were introduced to increase it. Zirconyle nitrate was dissolved in an aqueous hydrogenocarbonate solution in which En was added. These solids are called I3X. The ZrOCb based impregnating solution was prepared by using a tartaric acid-H20 solution in which ZrOCl2 was added. The dissolution of ZrOCk was achieved by adding En. These soUds are referred hereafter I4X. The procedure used for the impregnation of extrudates with bacote is the same as described above. These new solids will be referred as I5X and leX, the latter one being prepared with ethylene diamine.
2.1.3. Grafting
Grafting, based on the controlled hydrolysis of a zirconium alkoxide by surface hydroxyl groups of AI2O3 [16] was also performed. The alumina powder previously dried during 24 hours in an oven at 100 ^'C was added to n-propanol and maintained at 85 °C under agitation. Then, the desired amount of Zr-n-propoxide, 4 cm^ of fuming nitric acid and acac with [Zr n-prop]/[acac]=l were added. The mixture was maintained under inert atmosphere during 24 hours. Then, the solvent was eliminated in a rotary evaporator and the
551 powder was dried at 100 ^'C. A grafting at room temperature was also performed. The same procedure as above was used except that the n-propanol was changed by n-hexane and no complexing agent was used. Two types of supports, corresponding to both preparation modes were thus obtained which will be hereafter identified respectively by GR1I6 et GR2I6 (both are containing 16 wt% Zr02 to be consistent with ref [16]). Whatever the preparation method, after drying, all the solids were calcined under air at 500 **C during 4 hours, this temperature being reached at a rate of 72 *»C.h-i
2.2. Characterizations
Bulk atomic compositions were determined using X-ray fluorescence by the "Service Central d'Analyses du CNRS, Vernaison, France". X-ray diffraction (XRD; Siemens D5000 diffractometer equipped with a monochromator and a Cu X-ray tube) and laser Raman spectroscopy (LRS; XY from DUor) were used to investigate the structure of the samples. Specific surface areas (SSA) were measured using the BET method with a Quantasorb Junior (Ankersmit). The surfaces of the samples were examined by X-ray photoelectron spectroscopy. Binding energies (BE) were measured by reference to the AI2P peak at 74.8 eV except for the pure zirconia samples for which the reference was the Cis (BE = 285 eV). The surface atomic composition was determined using the following formula: ^E ^'^'' ^Kx',n'l'j' x
x,nrj
x,nlj
(1)
Kx, nlj
Where Ix,nij et Ix.nTj' are the X rays photoelectron peak intensities (peaks areas) for elements x and x' considering the nlj and nTj' orbitals, EK is the kinetic energy, and a is the corresponding cross section.
3. RESULTS AND DISCUSSION 3.1. Texture 3.1.1. Sol-gel preparations
The SSA are reported in Table 1. The SSA of the pure zirconia samples were found to be 54 m^.g-i and 17 m^.g-i for the samples respectively prepared from the propoxide and butoxide precursors through the procedure I whilst lower SSA are obtained for solids prepared through procedure II (8.4 m^.gi with n-propanol as solvent and 2 m^.g-i with 2-butanol as solvent). Moreover, these latter samples were heavily contaminated by incompletely burned organic substances. In contrast, the SSA of AlP and A1P(A) solids were respectively 400 m^.g-i and
552 375 m^.g-i. The sample prepared with a mixture of alkoxides do show a decrease in SSA when there is a high level of Zr and whatever the Zr content, the use of ZrP in its parent alcohol allows to have soUds with higher SSA. Whilst the chelated preparations gave samples with larger SSA up to 30 %, non-chelated preparations offer a higher SSA and high water pore volume (4.5 cm^.g-i for the 30P-P solid) for samples of at least 60 wt% Zr02. Table 1. Nomenclature and specific surface areas of sol-gel samples (calcined at 500 ®C under O2) Name %Zr02 Zralk. Solvent SSA (m2.g-i) Non chelated preparations lOP-P
10
ZrP
propanol
409
lOB-B
10
ZrB
hutanol
380
20P-P
20
ZrP
propanol
415
30P-P
30
ZrP
propanol
366
30B-B
30
ZrB
hutanol
352
60P-P
60
ZrP
propanol
432
95P-P
95
ZrP
propanol
125
ZrP-P
100
ZrP
propanol
54
ZrB-B
100
ZrB
hutanol
17
AlP
0
-
propanol
404
10P-P(H)
10
ZrP
propanol
404
Chelated preparations with acac 30P-2B(A)
30
ZrP
2-hutanol
496
60P-2B(A)
60
ZrP
2-hutanol
95
95P-2B(A)
95
ZrP
2-hutanol
7.4
ZrP-2B(A)
100
ZrP
2-butanol
2
A1P(A)
0
-
propanol
375
ZrP-P(A)
100
ZrP
propanol
8.4
553
3.1.2. Impregnated and grafted samples
The SSA of these samples are reported in Table 2 and Table 3.
Table 2. SSA of impregnated and grafted samples on alumina powders SoUd Al2O3/50^ Ii2 I22 GR1I6 SSA(m2.g-i)
373
336
336
GR2I6
373
294
The SSA of Ii, I2 and GR1I6 samples are similar to the value of the bare alumina. This shows that the textural properties of alumina are not modified by the impregnation. A small decrease is observed for GR2I6. This could be ascribed to the high Zr loading of this sample which is likely obstructing the pores. Table 3 also shows, for impregnated extrudates at the considered loadings that the impregnation does not affect the specific surface area.
Table 3. SSA of some impregnated samples on extrudates Solid
AI2O3
I3I.7
I4I.7
I43.4
I5I.7
Is5
lel.T
l65
SSA (m2.g-0
350
340
335
342
350
346
343
340
3.2. S t r u c t u r e 3.2.1. X-ray diffraction a n d laser R a m a n s p e c t r o s c o p y The XRD patterns of samples prepared by the sol-gel method are reported in Figure 3b-d. The rather amorphous character of these solids is partly responsible for the peak broadening. The XRD features of pure zirconia correspond either to the tetragonal or the cubic form with a small quantity of monoclinic form (by comparison with JCPDS data). It is not possible to distinguish between these two forms by XRD. On the contrary, LRS is a suitable technique for differentiating between the various aUotropic forms of zirconia [17, 18]. The Raman spectra of these phases show obvious differences : tetragonal Zr02 shows lines at 260 c m ^ 321 cm-i, 462 cm^ and 636 cm-i, the monoclinic form shows six weU defined lines at 638 cm-i, 561 cm-i, 538 cm-i, 504 cm-i, 476 cm-i, 381 cm-i, 335 cm-i, 191 cm-i, whereas the cubic form exhibits two broad lines at 600 cm^ and 300 cm-i. As the use of a Raman microprobe allows to analyze particle by particle, it is clearly shown in Figure 4 that monoclinic and quadratic phases are present, in good agreement with the XRD data. However, the existence of cubic Zr02 could not be rejected.
554
1400 ^
1200 t
S* 1000 + •IN
^
800
d
J
o
600
u
S
400 f 200 H
1
1
1
\
\
1
1
15
25
35
45
55
65
75
29 0 Figure 3. XRD patterns of some sol-gel samples (a) Y-AI2O3, (b) ZrP-P, (c) 60P-P, (d) 30P-P, (e) I28, (f) I18, (g) GR2I6, (h) GR1I6, (*) cubic zirconia lines The volumic proportion (Vm) of monoclinic Zr02 can be estimated from the XRD results with the equations used by Mercera and al [19]. Since the peaks referred in the equation were poorly resolved, a computer simulation was used, after corrections for background and Ka radiation, to access to the intensity line, giving Vm = 0.23 (ie. ^ 25% monoclinic Zr02) in general agreement with Mercera and al, who report Vm ~ 0.3 for a sample calcined sample at 500 °C. The XRD patterns of the unchelated Zr-Al samples (Figure 3) show that these samples are amorphous and no crystalline features of zirconia are observed even at 60 wt% Zr02. Examination by Raman spectroscopy also failed to find any definite traces of Zr02 in the unchelated, calcined Zr-Al preparations. The XRD patterns of these samples shows the features of y-Al203 at low Zr loadings whereas at high Zr loadings broad peaks at 20 « 31° and 15° are observed. These two peaks are also observed for pure zirconia and should correspond to an amorphous zirconia phase.
555
1000
1500
50.0
lO'OO
500
v(cm~*) Figure 4. Raman spectra of two different particles of ZrP-P For the leS solid only the XRD features of alumina are observed (XRD pattern not shown). Indeed, the Zr loadings of the la to le samples are below the detection limit of this technique. Whatever the samples, impregnated or grafted alumina powders, the XRD patterns (reported in Figure 3 e-h) show a broad peak at 20 ~ 31** characteristic of an amorphous zirconia and peaks corresponding to gamma-alumina (20 ~ 36°, 45°, 66°). For GR2I6 the XRD features of tetragonal or cubic zirconia (marked as "*" on Figure 3) are observed on the broad underlying peaks of amorphous zirconia and alumina. Such formation is in agreement with the lowering of the SSA. The absence of a chelating agent induces a higher sensitivity of ZrP to hydrolysis and thus the formation of bulk zirconia which obstructs the pores.
3.2.2. X-ray p h o t o e l e c t r o n s p e c t r o s c o p y a n d X-ray m i c r o p r o b e For all the analyzed solids the binding energy (BE) of the Zr3d5/2 level never characterizes bulk zirconia (BE = 182.2 eV). The BE is observed at 182.8 eV and characterizes Zr^v j ^ interaction with alumina.
3.2.2.1 Sol-gel samples XPS is a surface sensitive technique which allows the determination of the surface atomic ratio nzr/uAi. For homogeneous samples, if the XPS surface atomic ratio is plotted versus the bulk atomic ratio a straight line with a 0.5 slope should be observed. This curve is plotted on Figure 5.
556 A • • —
0
0,2
0,4
0,6
unchelated propoxide unchelated butoxide chelated propoxide homogeneous
0,8
nZr/nAl (bulk) Figure 5. Surface atomic ratio nzr/uAi (XPS) as a function of the bulk atomic ratio It shows that the zirconium is homogeneously dispersed in the chelated propoxide samples. It also indicates that the unchelated 10% Zr02 sample is approximately homogenous, whereas for the corresponding 20% and 30% samples there is a tendency for Al enrichment in the surface. The unchelated samples prepared from Zr n-butoxide generally show a lower nzr/uAi ratio than similar samples prepared using the n-propoxide. The unchelated sample prepared from the n-propoxide with partial pre-hydrolysis of the ASB (not reported here) produces an homogenous material and thus this method does not offer any advantage over preparations in which both alkoxides are hydrolysed simultaneously. So, no surface enrichment with zirconia is evidenced for these solids prepared by a sol-gel method. Whatever the Zr content, the difference in BE between the Al2p and Ois electron levels for non-chelated samples prepared from Zr n-prop is the same (-- 457.0 eV). In contrast, up to 60 wt% Zr02, the BE difference between Zraa and Ois levels is constant (A « 348.9 eV). This suggests that these samples essentially consist of an AI2O3 system containing Zr species. This is supported by the XRD and LRS data which failed to show any significant presence of Zr02. The non-chelated samples prepared from butoxide precursors and the chelated samples did not show the same trends. With the data available, no definite explanation was found to interpret the results for these two types of sample.
3.2.2.2. Impregnated samples Figure 6 shows the repartition of the zirconium entities in the extrudates, observed by X-ray microprobe. Whatever the zirconium loading, two types of concentration shapes are observed. The first one (Figure 6a) shows that zirconium entities well diffuse in the extrudates when the impregnation is performed with ZrOCb and ZrO(N[03)2 solutions. In the reverse, when using bacote with or without En, the profile shows a mass transfer limitation during
557
the impregnation. Thus the nature of the zirconium precursor i.e. the steric hindrance of the zirconium polymer present in the impregnating solution drastically affects the diffusion in the alumina extrudates. These solids have then been analyzed after bruising. The dispersion of the Zr has been studied by XPS using the Moulijn and Kerkhof model [20] and the characteristic curve is reported in Figure 7 in which is also reported the theoretical line. Indeed, the thickness U of the alumina layer (given by t« = 2/p8.So -where p« is the volume mass of the alumina carrier and So its SSA-) is 14 A and is inferior to the mean free path of the photoelectrons in the range of studied kinetic energies.
Figure 6. Zirconium concentration shape of cut alumina extrudate impregnated with solutions containing Zr : (a) from a ZrOCb or a ZrO(N03)2 solution, (b) from a bacote solution or a bacote + En solution The Zirconyle based solids present a Moulijn and Kerkhof t5rpe repartition. It confirms that Zr atoms are well dispersed in the macrostructure but also in the micropores. When the extrudates are impregnated with bacote without En, the ratio IZT/IAI is higher than the predicted values by the Kerkhof-MouHjn model. This shows that there is a coating of all the particles which form the extrudates. This phenomenon is far less important for samples prepared with En. This is confirmed by the XPS results obtained on the IiX. The XPS surface intensity ratio shows that a good dispersion is obtained for the I2 solids whilst the Ii ratio are above the theoretical line indicating that this carrier has a small surface zirconia enrichment. Thus, the use of En allows a better dispersion of the Zr entities whereas a coating of the alumina grains by an amorphous zirconia should exist in the solid Ii. We can formulate the hypothesis that En reduces the
558
size of the zirconium based polymeric species present in the solution and so increases their diffusion in the alumina pores. In conclusion, two types of Zr repartition have been evidenced for samples prepared with bacote: i) A macroscopic diffusion which induces the repartition of Zr in the alumina extrudates. In both cases, the migration of Zr entities is dependent of the steric hindrance of the Zr precursor used. ii) A coating of the particles.
•
1.2
•
1-0.8
•
^
^0.6-
•
0
0
^ y ^
u
0.4 0.2
k
0
—\
0.02
\
0.04
H 0.06
0.08
(^ZI/^\l)bulk
Figure 7. Results relative to the Kherkof-Moulijn model [20] : plot of IZr3d/IAl2p XPS intensity ratio as a function of the Zr/Al bulk composition. Ii (•);I2 (A);I3 (-);l4 (A);l5 (o); l6(n); GR2 (•); GRi is not represented on this graph as the result is much higher as these presented here.
3.2.2.3. Grafted samples
GR2I6 solid, prepared with En in the impregnating solution shows a MouUjn and Kerkhof type repartition but the GR1I6 has a characteristic ratio (2.9, not reported in Figure 7) which is higher than the theoretical one. In that case, a dip-coating of the alumina grains should likely occur which hides the aluminium atoms. The difference between the two procedures of grafting could be ascribed to the lower sensitivity to hydrolysis of the chelated ZrP.
559
4. CONCLUSION Various synthesis methods have been used to prepare mixed Zr-Al oxides which can substitute pure zirconia as support of catalysts. Sol-gel synthesis gives mixed Zr-Al oxides in which amorphous zirconia is dispersed in the alumina. However, due to their interesting textural properties (high SSA and high water pore volume), the non chelated soUds could be used as carriers for the preparation of HDS oxidic catalyst precursors. The impregnation, i.e. the classical method of doping an alumina carrier, gives samples in which the repartition of the zirconium depends drastically on the precursor. The use of the bacote allows the coating of alumina grains by an amorphous zirconia phase. A "dip-coating'' could also be achieved by grafting with the Zr-n-propoxide. These supports constituted by coated alumina should have structural and textural properties which are well required for being used as carriers for hydrodesulphurization catalysts.
REFERENCES 1. J. Ramirez, M. Vrinat, S. Fuentes, G. Diaz, M. Breysse, M. Lacroix, Appl. Catal., 52 (3), 211 (1989) 2. US Patent n^ 76-690254, Exxon Res. and Eng. Comp. (1976) 3. P. Atanasova, T. Halachev, J. Vchytil, M. Kraus, Appl. Catal., 38, 235 (1988) 4. G. MuraHdhar, F. E. Massoth, J. Shabtai, J. Catal., 85, 44 (1984) 5. D. Hamon, M. Vrinat, M. Breysse, B. Durand, M. Jebruni, M. Roubin, P. Magnoux and T. des Courieres, Catal. Today, 10, 613 (1991) 6. B. S. Parekh, S. W. Weller, J. Catal., 55, 58 (1978) 7. J. C. Duchet, M. J. TilUette, D. Cornet, Catal. Today, 10, 507 (1991) 8. P. Ruin, G. MeHn, R. Guinebretiere, A, Lecomte and A. Danger, J. of Sol-Gel Sci. and Techn., 2, 539-544 (1994) 9. L. Chiuping, C. Yu-Wen and Y. Tzu-Ming, J. of Sol-Gel Sci. and Techn., 4, 205215 (1995) 10. M. L. Rojas-Cervantes, R. M. Martin-Aranda, A. J. Lopez-Peinado and J. de D. Lopez-Gonzalez, J. of Material Science, 29, 3743-3748 (1994) 11. J. G. Weissman, E, C. DeCanio and J. C. Edwards, Catal, Lett,, 24, 113-122 (1994) 12. J. Livage, M. Henry and C. Sanchez, Prog. SoUd St. Chem., 18, 259-341 (1988) 13. E. Ponthieu, E. Payen, J. Grimblot in Sol-Gel Processing and Applications, A. Attia Ed., Plenum Press (1994) 14. M. Guglielmi, G. Carturan, J. Non Cryst. SoHds, 100, 16-30 (1988)
560 15. E. Ponthieu, E. Payen, G. M. Pajonk, J. Grimblot, J. Sol-Gel Sci. & Techn. (under press) 16. C. Marquez-Alvarez, J. L. G. Fierro, A. Guerrero-Ruiz, and I. RodriguezRamos, J. of Colloid and Interface Science, 159, 454-459 (1993) 17. P. D. L. Mercera, J. G. Van Ommen, E. B. M. Doesburg, A. J. Burggraaf and J. R. H. Ross, Appl. Catal., 57, 127 (1990) 18. E. Payen, L. Gengembre, F. Mauge, J.-C. Duchet, J.-C. Lavalley, Catal. Today, 10,521 (1991) 19. P.D.L. Mercera, J. G. Van Ommen, E. B. M. Doesburg, A. J. Burggraaf and J. R. H. Ross, Appl. Catal., 57, 127 (1990) 20. F. P. J. M. Kerkhof, J. A. Moulijn, J. Phys. Chem., 83, 1612 (1979)
® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors
561
Selective synthesis of methylcyclopentane from cyclohexane using Pt-zeolite hybrid Catalyst I. Nakamura, A. Zhang and K. Fujimoto Department of Applied Chemistry, School of Engineering, The University of Tokyo 7-3-1, Kongo, Bunkyo-ku, Tokyo 113 Japan Isomerization of cyclohexane to methylcyclopentane was studied on a series of ZSM-5 catalyst in the presence of hydrogen . A Pt-ZSM-5 catalyst, which is typical dual functional catalyst for skeletal isomerization, showed high activity and high methylcyclopentane selectivity. A hybrid catalyst prepared by physically mixing H-ZSM-5 and Pt/Si02 showed high activity and selectivity equal to those of Pt/ZSM-5 while a H-ZSM-5 and a Pt/Si02 showed a poor catalytic performance for this reaction. With the hybrid catalyzed reaction, the mixing state of H-ZSM-5 and Pt/Si02 had a strong influence on the activity and methylcyclopentane selectivity. In the absence of hydrogen, the conversion of cyclohexane was drastically reduced and cracking reaction became dominant. These results suggest that spilt-over hydrogen plays an important role in cycloparaffm activation and the stabilization of intermediates to give high isomer yield.
1. INTRODUCTION Hydrogenation of benzene to cyclohexane followed by isomerization of cyclohexane to methylcyclopentane appears to be very attractive to reduce the benzene content in gasoline. Since research octane number (RON) get 8 points increment when cyclohexane (RON 83) is converted to methylcyclopentane (RON 91), and the loss in octane number caused by the deduced amount of aromatics can be compensated pardy by isomerized product. It is generally known that the hydrogenation of benzene occurs very easily on noble metal such as Pt or Pd, therefore in this study, we just focus our attention to the selective isomerization of cyclohexane. Isomerization reaction is usually carried out using zeolite supported noble metals catalysts such as Pt/H-ZSM-5 or Pt/mordenite in the presence of hydrogen. Steinberg et al. studied conversion of cyclohexane over a Pt/H-ZSM-5 or a mechanical mixture of a H-ZSM-5 and a Pt/Al203 using a reactor with recirculation in the presence of hydrogen at 0.1 MPa [1-2]. Mechanical mixing of small amount of Pt/Al2C^ (1 wt%) to H-ZSM-5 promoted dehydrogenation of cyclohexane to give benzene. Over Pt/H-ZSM-5, MCP, benzene, nhexane and i-hexane were formed as intermediates and the final products of hydrocracking are observed increasingly with reaction time. They claimed that the large supply with spillover hydrogen opens the route to benzene, parallel to methylcyclopentane and products of the ring opening reaction.
562 On the other hand, in the hydroisomerization of n-pentane over hybrid catalyst containing H-ZSM-5 and supported noble metal catalyst, it was proposed that a hydrogen molecule spills over onto zeolite surface as a proton and a hydride [3-4]. The proton promotes isomerization reaction even at low reaction temperature. The hydride ion generated simultaneously stabilizes intermediate carbenium ion to prevent oligomerization and cracking of the oligomers and promote isomerization of alkane. In this paper, cyclohexane isomerization over a hybrid catalyst containing H-ZSM-5 and supported Pt catalyst was studied using a continuous flow type fixed bed reactor under atmospheric and pressurized conditions. 2. EXPERIMENTAL Pt(0.5 wt%)/H-ZSM-5 catalyst was prepared using a H-ZSM-5 (HSZ-840NHA, supplied by Toso Ltd.) with silica/alumina ratio of 44. Platinum was loaded to the zeolite from the aqueous solution of tetraammineplatinum (II) chloride by ion-exchange. Pt(2.5 wt%)/Si02 was prepared by wet-impregnating a commercially available Si02 (Aerosil 380, BET specific surface area 380 m^/g) with aqueous solution of H2PtQ6. Pt-hybrid catalyst was prepared by co-grinding a mixture of four weight parts of H-ZSM-5 and one weight part of Pt/Si02 and pressure molding granules. Catalysts (20/40 mesh) were calcined in air at 550 "C for 2 h and reduced in flowing hydrogen at 400 "C for 1 h. The isomerization of cyclohexane (CH) and the conversion of benzene (BZ) to methylcyclopentane (MCP) was conducted with a continuous flow type fixed bed reactor under atmospheric and pressurized conditions. 0.5 g of catalysts were packed in a stainless steel tube with inner diameter of 4 mm. CH or BZ was fed to the reactor by a liquid microfeeder. The mole ratio of H2/CH or BZ in the feed was 9:1. Products were withdraw in the gaseous state and analyzed by a high-resolution capillary gas chromatograph. 3. RESULTS AND DISCUSSION 3.1. Isomerization of cyclohexane (CH) Table 1 shows CH conversion and product distributions for four catalysts (H-ZSM-5, Pt/Si02, Pt/H-ZSM-5 and Pt-hybrid) at 290 "C under atmospheric pressure. Pt/H-ZSM-5, which is a typical dual functional hydroisomerization catalyst, showed high catalytic activity and significantiy high selectivity of MCP. Pt/Si02 also showed high catalytic activity, however, dehydrogenation of CH was proceeded exclusively to give BZ. H-ZSM-5 showed poor activity in CH conversion and MCP selectivity, cracking reactions occurred more seriously than isomerization. However, Pt-hybrid catalyst, which is a physical mixture of Pt/Si02 and H-ZSM-5, showed higher activity and selectivity for MCP formation than Pt/HZSM-5. Although the low BZ yield on the hybrid catalyst should be attributed to the lower content of Pt/Si02, the drastic increase in MCP yield by hybridization is hard to be explained. In Table 1 are shown how the method of mixing of Pt/Si02 with H-ZSM-5 affects their isomerization activity. The simple mixture of granules (20-40 mesh) of each catalyst showed the catalytic performance similar to that of Pt/Si02. This means that the intimate contact of Pt/Si02 and H-ZSM-5 is essential for the generation of the isomerization activity and the H-
563 ZSM-5 was almost inactive under these conditions when it is separated from Pt/Si02. The effects of atmosphere are also shown in Table 1. It should be also noted that the catalytic activity of the intimately mixed hybrid catalyst was extremely low and the yield of MCP was lower than that of H-ZSM-5 under nitrogen atmosphere. This clearly shows that hydrogen gas is essential for the generation of the isomerization activity. As an important reason that Pt-hybrid catalyst also including Pt/H-ZSM-5 promotes tiie isomerization reaction in the presence of hydrogen, the effect of hydrogen spillover has been claimed [1-4]. In Fig. 1 are shown the main roles of the spillover hydrogen, perhaps existing on zeolite surface as proton and hydride, are that (1) the spillover proton promotes acid catalyzed reactions on Br0nsted acidic site, (2) isomerized carbenium ion intermediate is stabilized rapidly by the hydride addition to suppress acidic cracking reactions effectively. Table 1 Isomerization of cyclohexane on ZSM-5 catalyst Catalyst Atmosphere Conv. (%) Yield (%)
HZSM-5
Pt/Si02 Pt/HZSM-5
Pt-hybrid^^
Pt-hybrid**^ H2
3.7 1.4
54.4 0.5
51.5 16.7
53.1 29.6
5.8
0.2 0.4 0.5 0.0 0.1
0.3 1.4 0.8 4.6
0.5 2.3 0.9 3.1
4.7 0.7
55.8 3.0
3.2
0.3 0.4 0.6 0.4 5.4
Selectivity (%) C1+C2 C3 C4+C, QH,4
21.3 13.2
MCP BZ
36.9 12.2
others
5.8
4.8
98.8
0.0
32.5 60.1
0.3
55.8 37.3
0.1
11.9
7.4 3.2
20.7 50.9
2.7
92.7
0.2
290 ^C, 0.1 MPa, W/F = 5.0 g h mo\-\ UjCH = 9 : 1 (mole ratio), a) powdery mixture, b) granular mixture,
+ B
+ Hi
BH-'
Figure 1. Model of hydrogen spillover in cyclohexane isomerization over Pt-hybrid catalyst.
564 BZ was formed via dehydrogenation of CH as a main by-product over Pt/H-ZSM-5 and Pt-hybrid under atmospheric pressure. However, dehydrogenation of CH to form BZ should be suppressed by pressurized hydrogen. Fig. 2 shows the effect of reaction pressure. As it is clear from the figure, BZ yield decreased drastically with increasing reaction pressure to reach almost zero at 0.6 MPa. The yield of MCP increased with an increase in reaction pressure. Cracked products of CI to C5 hydrocarbons increased slightly with increasing reaction pressure. They were mostly composed propane and the formation of methane was negligible small for Pt/H-ZSM-5. These results suggest that the cracking reaction occurs mainly on acidic site and the hydrocracking on metal site was negligible even under high hydrogen pressure. C6 paraffms were mostly composed of n-hexane, 2-methylpentane and 3methylpentane.
0.1 0.2 0.4 0.6 0.8 1.1 Reaction pressure (MPa) n
CH H others •
BZ Q C6 ^
C1-C5 M MCP
Catalyst; Pt/H-ZSM-5, W/F=0.5 g h mof^ 290 ^C Figure 2. Composition of products as function of reaction pressure in CH conversion. The products composition of CH conversion were plotted against the reaction temperature in Fig. 3. The reactions were carried out under 0.6 MPa, because over Pt/H-ZSM-5, BZ formation was suppressed completely and isomerization of CH proceeded selectively in hydrogen flow as mentioned before. At all reaction temperature, H-ZSM-5 showed poor activity in CH conversion and MCP selectivity. Pt/Si02 also showed poor activity in CH conversion, however it showed high catalytic activity for the dehydrogenation under 0.1 MPa. The dehydrogenation of CH to give BZ on platinum was suppressed by pressurized hydrogen. WTith Pt-hybrid catalyst or Pt/H-ZSM-5 and up to 260 *C, the yield of MCP increased with an increase in reaction temperature. At higher reaction temperature, the conversion of CH increased with an increase in reaction temperature, however, due to the occurrence of cracking reaction, the MCP yield no longer coincide with the CH conversion. The products compositions of CH conversion were plotted against the contact time (W/F) in Fig. 4. The reactions were carried out under 0.6 MPa at 260 "C with Pt hybrid catalyst or Pt/H-ZSM-5. Both conversion level and MCP yield increased with an increase in contact time with either Pt hybrid catalyst and Pt/-ZSM-5. Up to 7.5 (g h mol-i) of W/F, the cracking reaction was suppressed effectively, however, small amount of cracked products were formed atlO.O(ghmol-i)ofW/F.
565 H-ZSM-5
Pt/SiO^
200 230 260 290 320
200 230 260 290 320 Temperature (°C)
100 0 80 h E
i
60 |-
1 40 I|20
5 0 Temperature (°C) Pt/H-ZSM-5
100 o 80 £ y. 60
CH
_100 •S 80 B ^ 60
J^
1 40
140
1 20 O
^
1 20
^^gjg^^^^m
o ^
0 200 230 260 290 320 Temperature (°C)
D
Pt-hybrid catalyst
CH
H others
•
BZ
0 200 230 260 290 320 Temperature (°C)
E3 C6
S
m MCP
C1-C5
W/F=0.5 g h mof , reaction pressure 0.6 MPa. Figure 3. Composition of products as function of reaction temperature in CH conversion.
Pt/H-ZSM-5
2.5
5
_100
7.5
2.5 5 7.5 W/F(ghmoH)
W/F(ghmol-^) n
CH
H others
Pt-hybrid catalyst
BZ
H C6
S
C1-C5
m MCP
Figure 4. Composition of products as function of contact time in CH conversion at 260 ^C under 0.6 MPa.
566 3.2. Hydroconversion of BZ into MCP The results of BZ conversion to MCP using Pt/H-ZSM-5 or Pt-hybrid catalyst are shown in Fig.5. At all reaction temperature and over Pt/H-ZSM-5 or Pt-hybrid catalyst, little BZ was detected in outlet gas and compositions of products for the both catalysts were very similar to those of the CH conversion. BZ should be hydrogenated on Pt particle rapidly, because the reactions were carried out under 1.1 MPa in hydrogen flow. And the reaction becomes essentially the isomerizationof CH. Pt/H-ZSM-5
230 260 290 Temperature CQ n CH
^^
Pt-hybrid catalyst
200 230 260 290 320 Temperature (°C)
B others H C6 BZ -1 W/F=0.5 g h mol" , reaction pressure 1.0 MPa.
0
C1-C5
m MCP
Figure 5. Composition of products as function of reaction temperature in benzene conversion over H-ZSM-5 containing catalysts. 4. CONCLUSION Cyclohexane was converted to methylcyclopentane selectively over Pt-hybrid catalyst, which is a physical mixture of Pt/Si02 and H-ZSM-5 in the hydrogen flow, while H-ZSM-5 and Pt/Si02 showed a poor catalytic performance for this reaction. The mixing state of HZSM-5 and Pt/Si02 had a strong influence on the activity and methylcyclopentane selectivity. These results suggest that spilt-over hydrogen plays an important role in cycloparaffm activation and the stabilization of intermediates to give high isomer yield. Further more, benzene was converted methylcyclopentane selectively over Pt-hybrid catalyst or Pt/H-ZSM-5 under pressurized hydrogen. REFERENCES 1. K.H. Steinberg, U. Mroczek and F. Roessner, Proceeding of the Second International Conference on Spillover, Leipzig, pp. 150-166 (1989). 2. F. Roessner, U. Mroczek and A. Hagen, in "New Aspects of Spillover Effect in Catalysis" (T. Inui, K. Fujimoto, T. Uchijima and M. Masai, eds.), Elsevier, Amsterdam, 1993, pp. 151-157. 3. K. Fujimoto, K. Maeda and K. Aimoto, Appl. Catal., 91, (1992) 81. 4. A. Zhang, I. Nakamura and K. Fujimoto, Ind. Eng. Chem. Res. 34 (1995) 1074.
® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G,F. Froment, B. Delmon and P. Grange, editors
567
Hydrocracking activity of NiMo - USY zeolite hydrotreating catalysts B. Egla'', J.F. Cambra'', B. Guemez^ P.L Arias'', B. Paweiec^ and J.LG.Fierro^. 1 Departamento de Ingenieria Quimica y del Medio Ambiente. Escuela de Ingenieros. Alda Urquijo s/n. 48013 Bilbao. Spain. 2 Instituto de Catalisis y Petroleoquimica (CSIC). Campus U.A.M. Cantoblanco. 28049 Madrid. Spain. 1.
INTRODUCTION
Hydrotreating reactions are usually carried out with bimetallic catalysts containing Mo or W and Ni or Co depending on the function to be promoted [1,2]. As hydrogenation seems to be an important step In heteroatom removal from polyaromatic compounds the use of Ni as promoter is a growing tendency. As the carrier plays an important role on the activity of the final sulfided catalyst [2], the behaviour of different supports is an important field of research. For example, alumina is the most extended carrier for this kind of reactions, but it has a moderate to low hydrocracking activity when severe hydrotreatment of the feedstock is needed. Research on the application of zeolites in heteroatom removal processes is fairly recent [3], being their exceptional properties as catalityc activity and resistance to poissoning by sulfur and nitrogen containing organic compounds an incentive for its use as hydrotreating catalyst support. In previous works [4-6] the HDS and HDN activities of Ni, Mo and MoNi zeolite supported catalysts were studied, and a remarkable initial deactivation by cocking was detected. Thus in the present work the hydrocracking activity of these catalysts in the sulfided state has been studied. 2. EXPERIMENTAL 2.1. Catalysts The catalyst compositions used in this work are shown in table 1, as measured by atomic absortion. The Ni loaded ultraestable Y (USY) zeolite catalysts were prepared by ion exchange with Ni(N03)2.4H20. These catalysts will be refered to as Nix, where x is the theoretic content of Ni atoms per unit cell, (x = 2,5,9,14,23). The three molibdena loaded ultraestable Y zeolite catalysts were prepared using various procedures and different precursors. One catalyst was prepared by solid ion exchange by mixing the USY zeolite with M0CI5 [7-11]. These catalyst will be referred to hereafter as MoCI. Catalyst MoCO was prepared by wet impregnation of USY zeolite with neutral Mo(CO)6 complex [12-17]. Catalyst 4MoA was prepared by conventional aqueous impregnation of USY zeolite with (NH4)6Mo7024.H20 (HMA) followed by removal of water in a rotary evaporator
568 [18-20]. The Impregnates were dried in vacuum for 8 h and in air at 383 K. Calcination was achieved at 773 K, and for the 4MoA catalyst at low and constant pressure of water extended over a period of time of 800 h [19]. Table 1 Chemical composition of the calcined catalysts and HC activity. Total conversion (wt %) Catalysts
M0O3 (wt %)
NiO (wt %)
Ni2 Ni14 4MoA MoCI MoCO 4MoNi2 4MoNi5 4MoNi9 4MoNi14 4MoNi23
-
0.8 5.0 0.8 1.5 2.3 5.3 10.2
3.9 6.3 3.1 5.2 5.3 5.6 6.0 6.3
548 K 14.3 2.5 7.6 8.2 11.6 19.7 11.5 18.9 13.0
598 K 15.9 35.1 12.8 23.9 30.5 34.2 28.6 25.4 46.4 21.5
648 52.8 21.7 45.7 48.4 66.5 63.1 55.6 65.4 43.4
The five binary MoNi catalysts that will be refered to as 4MoNix, where x is the theoretic content of Ni atoms per unit cell, were prepared by incorporating Ni first by ion exchange following impregnation from aqueous solutions of Ni(N03)2.4H20, being the procedure the one decribed for Ni/USY catalysts. Mo was subsequently incorporated in a second step by solid-solid ion exchange with M0CI5 precursor using the same method described previously for MoCI catalyst. These catalysts were characterised by atomic absortion, nitric oxide and pyridine adsorption and X-ray photoelectron spectroscopy, and results are included in previus works [4-6]. 2.2. Activity tests Hydrocracking of n-decane was performed in a stainless steel fixed bed catalytic reactor. This reactor was filled with glass beads (1 mm diameter) to the desired level for the catalytic bed. The bed, which was between two thin layers of SiC, was made of 0.3 g of catalyst diluted with SiC (1:3) of the same size (0.42-0.75 mm). The activation procedure consited of heating to the sulfidation temperature (673 K) in a nitrogen flow at atmospheric pressure. When this temperature was reached, a flow os H2:H2S (10:1 molar) mixture was passed through the catalyst bed for 4 h. The reactor was then purged in a nitrogen flow at 673 K for 0.5 h, and then cooled to room temperature. Finally, the nitrogen pressure increased up to that of the experiment and the catalytic bed was heated up to the temperature of experiment. Hydrogen and n-decane, with a 1% of dimethyl disulfure to mantain the sulfided state of the metals in the catalysts, were then passed through the reactor, with reaction products analyzed by a on-line HP 5890 A gas
569 chromatograph provided with FID detector. Reactions conditions were: P = 3 MPa, T = 548, 598 and 648 K and W/F = 19.5 Qcat-h/mol. 3. RESULTS The n-decane HC model reaction have been used to establish the perfomance of the Ni, Mo and MoNi zeolite supported catalysts. Results for Mo-zeolites tested in HC as a function of the time on-stream for a temperature of 598 K are shown in figure 1. Figure 1 shows a strong deactivation during the first hours on stream for all catalysts. The MoCO catalysts displays the best HC activity. The second one in activity is the MoCI catalyst and finally the 4MoA, where Mo is well dispersed in zeolite internal channels, but as Mo content is relatively low, and probably diffusional hinderance is present, the concentration of the active sites available for reaction is small and these could explain its low activity. 100
100
4M0-A 80
c
^ 4MoNi5 801
Mo-CI
c o •(0
O 4MoNi2
Mo-CO
60
60
Bo
40
C
A
•
• •
O
4MoNi9
^ 4MoNi14 o 4MoNi23
A
A A
Q
I
>
Q>
B 40
o
D
o
20
-1—1
2
• •
« • •
• - _
1
1
1
4
1
6
O X
A D
D
4
« fi i I 20 H
r-
8
10
Time on-stream (h)
Figure 1. Mo/USY catalysts. HC activity vs time on stream (T=598 K)
2
4
6
8
10
Time on-stream (h)
Figure 2. MoNi/USY catalysts. HC activity vs time on stream (T=598 K)
For Ni/USY catalysts, as can be seen in table 1, the HC activity increases with increasing Ni content, in agreement with the results obtained in HDS and HDN reactions for these catalysts [5]. The n-decane HC results for MoNi/USY series as a function of the time onstream for a temperature of 598 K are shown in figure 2. These data show a continuous decrease in conversion during the first hours on stream. The most active catalyst is the 4MoNi14 and then the catalysts with the lower Ni content, 4MoNi2 and 4MoNi5. The catalyst that exhibits the lowest HC activity is the
570 4MoNi23, with the highest Ni content, this behaviour could be explained in terms of a blockage of the zeolite porous structure by Ni atoms. From these data, it appears that the Ni content is an important but not a determinant factor in HC activity. 4. DISCUSSION The comparison of the n-decane HC total conversion data for the most active MoNi/USY catalyst, 4MoNi14, with the monometalic precursors, MoCI and Nil4, indicates that, for all the temperatures studied HC activity is higher for the binary system than for the monometallic ones, but it is lower than the corresponding to the addition of their activities, suggesting that there is not a promotional effect between the Ni and the Mo, probably due to the blockage of Ni atoms by the Mo incorporated in a second step As can be seen in table 1, HC activity for Ni catalyst is higher than that for Mo catalyst when comparing catalyts with the same metal content, MoCI and Nil4. as distinguished from the alumina supported in which the Ni sulfided catalysts present a lower activity than the Mo sulfided ones. This finding is in agreement with the work of Welters [21], who stated that these differences in activity are influenced for the metal phases dispersion over the support due to the different metal-support interactions. -J!
HC
I
I
I
S HDS
L
Q HDN |
zzzzzzzzzzzz o o o o o o o o o o o o o t o
Figure 3. Mo/USY catalysts. Comparison of HDS, HDN and HC activities
Figure 4. MoNi/USY catalysts. Comparison of HDS, HDN and HC activities
In order to compare activities in heteroatom removal (HDS and HDN) and HC reactions, the normalized steady state conversions corresponding to the different model compounds for the Mo/USY and MoNi/USY catalysts series are plotted in figures 3 and 4, respectively. Run conditions are different for each reaction, so data
571 are normalized to the total conversion of the MoCO catalyst in Mo series catalysts and to the 4MoNi14 catalyst for the binary systems. For Mo/USY catalysts, figure 3 shows that the activity trend in HDS and HC reactions is the same, that is: MoCO < MoCI < 4MoA, while in HDN reaction, the best catalyst was MoCI, being the catalyst which shows the highest acidity. For MoNi catalysts, the most active catalyst in HDS and HC reactions was 4MoNi14 catalyst with a high Ni content. HDS activity increased with increasing Ni content although for the catalyst with the highest Ni content, 4MoNi23, the conversion is lower, probably due to a blockage of the porous structure with metal atoms. The explanation for the difference in the order of activities in HDS, HDN and HC reactions is based on the fact that the active sites for these reactions are not the same, being the NiMoS phases active for the first two reactions, while the balance between hydrogenating (metal) and acid (support) functions is critical for HC reactions. 5.
CONCLUSIONS
From the results of this work, it may be concluded that: The bimetallic system supported on HY zeolite is more active in HC reaction than the monometallic Ni catalysts and these more than the Mo catalysts. The lack of sinergy in bimetallic catalysts may be due to the blockage of Ni atoms by Mo that is added in the second step. There is an initial deactivation of all the catalysts tested due to cocking reactions. For the binary systems there are not a correlation between metal content and HC activity, although apparently a maximum is found for the 4MoNi14 catalyst. Other factors as residual acidity after deactivation or pore blockage seem to play a significant role. ACKNOWLEDGMENTS This work was supported by the EC- Research Programme JOULE-2 (Contract 0049) and the University of the Basque Country.
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AtTTHOHS INPEK Akpati H. Antos G.J. Arias P.L. Arteca R. Bacaud R. BaudonA. Bebelis S. Bergem H. Berndt H. BlanchardP. Blanco M.N. BlekkanE.A. Bougeard D. Breysse M. Buitrago S. Buker K. Caceres C.V. CallejasM.A. Cambra J.F. CarbdE. Cattenot M. Chadwick D. Claridge J.B. Cornet D. Da Silva P. Dassori C.G. de Beer V.H.J. Debrabandere B. de Jong A.M. Deligiorgis G. Delmon B. Demirel B. Depauw G.A. Diaz A. DjuricM. DonkerR.. Dubois J.L Dumeignil F. Egia B. El Qotbi M. Faye Ph. Fernandez N. Flerro J.L.
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577
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Preparation of Catalysts I.Scientific Bases for the Preparation of Heterogeneous Catalysts. Proceedings of the First International Symposium, Brussels, October 14-17,1975 edited by B. Delmon, P.A. Jacobs and G. Poncelet TheControloftheReactivityof Solids. A Critical Survey of the Factors that Influence the Reactivity of Solids, with Special Emphasis on the Control of the Chemical Processes in Relation to Practical Applications by V.V. Boldyrev, M. Bulens and B. Delmon Preparationof Catalysts II. Scientific Bases for the Preparation of Heterogeneous Catalysts. Proceedings of the Second International Symposium, Louvain-la-Neuve, September 4-7,1978 edited by B. Delmon, P. Grange, P. Jacobs and G. Poncelet Growth and Properties of Metal Clusters. Applications to Catalysis and the Photographic Process. Proceedings of the 32nd International Meeting of the Societe de Chimie Physique, Villeurbanne, September 24-28,1979 edited by J. Bourdon Catalysis by Zeolites. Proceedings of an International Symposium, Ecully (Lyon), September9-11,1980 edited by B. Imelik, C. Naccache, Y. Ben Taarit, J.C. Vedrine, G. Coudurier and H. Praliaud Catalyst Deactivation. Proceedings of an International Symposium, Antwerp, October 13-15,1980 edited by B. Delmon and G.F. Froment New Horizons in Catalysis. Proceedings of the 7th International Congress on Catalysis, Tokyo, June 30-July4,1980. Parts A and B edited by T. Seiyama and K. Tanabe Catalysis by Supported Complexes by Yu.l. Yermakov, B.N. Kuznetsov and V.A. Zakharov Physicsof Solid Surfaces. Proceedings of a Symposium, Bechyhe, September 29-October 3,1980 edited by M. Laznicka Adsorption at the Gas-Solid and Liquid-Solid Interface. Proceedings of an International Symposium, Aix-en-Provence, September 21-23,1981 edited by J. Rouquerol and K.S.W. Sing Metal-Support and Metal-Additive Effects in Catalysis. Proceedings of an International Symposium, Ecully (Lyon), September 14-16,1982 edited by B. Imelik, C. Naccache, G. Coudurier, H. Praliaud, P. Meriaudeau, P. Gallezot, G.A. Martin and J.C. Vedrine Metal Microstructures in Zeolites. Preparation - Properties-Applications. Proceedings of a Workshop, Bremen, September 22-24,1982 edited by P.A. Jacobs, N.I. Jaeger, P. Jim and G. Schulz-Ekloff Adsorption on Metal Surfaces. An Integrated Approach editedbyJ. Benard Vibrations at Surfaces. Proceedings of the Third International Conference, Asilomar, CA, September 1-4,1982 edited by C.R. Brundle and H. Morawitz Heterogeneous Catalytic Reactions Involving Molecular Oxygen byG.I.Golodets
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Methane Conversion. Proceedings of a Symposium on the Production of Fuels and Chemicals from Natural Gas, Auckland, April 27-30,1987 edited by D.M. Bibby, CD. Chang, R.F. Howe and S. Yurchak Innovation in Zeolite Materials Science. Proceedings of an International Symposium, Nieuwpoort, September 13-17,1987 edited by P.J. Grobet, W.J. Mortier, E.F. Vansant and G. Schulz-Ekloff Catalysis 1987. Proceedings of the 10th North American Meeting of the Catalysis Society, San Diego, CA, May 17-22,1987 edited by J.W.Ward Characterization of Porous Solids. Proceedings of the lUPAC Symposium (COPS I), Bad Soden a. Ts., April 26-29,1987 edited by K.K. Unger, J. Rouquerol, K.S.W. Sing and H. Krai Physics of Solid Surfaces 1987. Proceedings of the Fourth Symposium on Surface Physics, Bechyne Castle, September 7-11,1987 edited byJ. Koukal Heterogeneous Catalysis and Fine Chemicals. Proceedings of an International Symposium, Poitiers, March 15-17,1988 edited by M. Guisnet, J. Barrault, C. Bouchoule, D. Duprez, C. Montassier and G. Perot Laboratory Studies of Heterogeneous Catalytic Processes by E.G. Christoffel, revised and edited by Z. Paal Catalytic Processes under Unsteady-State Conditions byYu.Sh.Matros Successful Design of Catalysts. Future Requirements and Development. Proceedings of the Worldwide Catalysis Seminars, July, 1988, on the Occasion of the 30th Anniversary of the Catalysis Society of Japan edited by T.lnui Transition Metal Oxides. Surface Chemistry and Catalysis byH.H.Kung Zeolites as Catalysts, Sorbents and Detergent Builders. Applications and Innovations. Proceedings of an International Symposium, Wurzburg, September 4-8,1988 edited by H.G. Karge and J. Weitkamp Photochemistry on Solid Surfaces edited by M. Anpo and T. Matsuura Structure and Reactivity of Surf aces. Proceedingsof a European Conference, Trieste, September 13-16,1988 edited by C. Morterra, A. Zecchina and G. Costa Zeolites: Facts, Figures, Future. Proceedings of the 8th International Zeolite Conference, Amsterdam, July 10-14,1989. Parts A and B edited by PA. Jacobs and R.A. van Santen Hydrotreating Catalysts. Preparation, Characterization and Performance. Proceedings of the Annual International AlChE Meeting, Washington, DC, November 27-December 2,1988 edited by M.L. Occelli and R.G. Anthony New Solid Acids and Bases. Their Catalytic Properties by K. Tanabe, M. Misono, Y. Ono and H. Hattori Recent Advances in Zeolite Science. Proceedingsof the 1989 Meeting of the British Zeolite Association, Cambridge, April 17-19,1989 edited by J. Klinowsky and P.J. Barrie Catalyst in Petroleum Refining 1989. Proceedings of the First International Conference on Catalysts in Petroleum Refining, Kuwait, March 5-8,1989 edited by D.L Trimm, S. Akashah, M. Absi-Halabi and A. Bishara Future Opportunities in Catalytic and Separation Technology edited by M. Misono, Y. Moro-oka and S. Kimura
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New Developments in Selective Oxidation. Proceedings of an International Synnposium, Rimini, Italy, September 18-22,1989 edited by G. Centi and F. Trifiro Volume 56 Olefin Polymerization Catalysts. Proceedings of the International Symposium on Recent Developments in Olefin Polymerization Catalysts, Tokyo, October 23-25,1989 edited by T. Keii and K. Soga Volume 57A Spectroscopic Analysis of Heterogeneous Catalysts. Part A: Methods of Surface Analysis edited by J.L.G. Fierro Volume 57B Spectroscopic Analysis of Heterogeneous Catalysts. Part B: Chemisorption of Probe Molecules edited by J.L.G. Fierro Volume 58 Introduction to Zeolite Science and Practice edited by H. van Bekkum, E.M. Flanigen and J.C. Jansen Volume 59 Heterogeneous Catalysis and Fine Chemicals II. Proceedings of the 2nd International Symposium, Poitiers, October 2-6,1990 edited by M. Guisnet, J. Barrault, C. Bouchoule, D. Duprez, G. Perot, R. Maurel and C. Montassier Volume 60 Chemistry of Microporous Crystals. Proceedings of the International Symposium on Chemistry of Microporous Crystals, Tokyo, June 26-29,1990 edited by T. Inui, S. Namba and T. Tatsumi Volume 61 Natural GasConversion. Proceedings of the Symposium on Natural Gas Conversion, Oslo, August 12-17,1990 edited by A. Holmen, K.-J. Jens and S. Kolboe Volume 62 Characterizationof Porous Solids II. Proceedings of the lUPAC Symposium (COPS II), Alicante, May 6-9,1990 edited by F Rodriguez-Reinoso, J. Rouquerol, K.S.W. Sing and K.K. linger Volume 63 Preparation of Catalysts V. Scientific Bases for the Preparation of Heterogeneous Catalysts. Proceedings of the Fifth International Symposium, Louvain-la-Neuve, September 3-6,1990 edited by G. Poncelet, P.A. Jacobs, P. Grange and B. Delmon Volume 64 New Trends in CO Activation edited by L. Guczi Volume 65 Catalysisand Adsorption by Zeolites. Proceedings of ZEOCAT 90, Leipzig, August 20-23,1990 edited by G. Ohlmann, H. Pfeifer and R. Fricke Volume 66 Dioxygen Activation and Homogeneous Catalytic Oxidation. Proceedings of the Fourth International Symposium on Dioxygen Activation and Homogeneous Catalytic Oxidation, Balatonfijred, September 10-14,1990 edited by LI. Simandi Volume 67 Structure-Activity and Selectivity Relationships in Heterogeneous Catalysis. Proceedings of the ACS Symposium on Structure-Activity Relationships in Heterogeneous Catalysis, Boston, MA, April 22-27,1990 edited by R.K. Grasselli and A.W. Sleight Volume 68 Catalyst Deactivation 1991. Proceedings of the Fifth International Symposium, Evanston,IL, June 24-26,1991 edited by C.H. Bartholomew and J.B. Butt Volume 69 Zeolite Chemistry and Catalysis. Proceedings of an International Symposium, Prague, Czechoslovakia, September 8-13,1991 edited by P.A. Jacobs, N.I. Jaeger, L Kubelkova and B. Wichterlova Volume 70 Poisoning and Promotion in Catalysis based on Surface Science Concepts and Experiments by M. Kiskinova
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Catalysis and Automotive Pollution Control II. Proceedings of the 2nd International Synnposium (CAPoC 2), Brussels, Belgium, September 10-13,1990 editedbyA.Crucq New Developments in Selective Oxidation by Heterogeneous Catalysis. Proceedings of the 3rd European Workshop Meeting on New Developments in Selective Oxidation by Heterogeneous Catalysis, Louvain-la-Neuve, Belgium, April 8-10,1991 edited by P. Ruiz and B. Delmon Progress in Catalysis. Proceedings of the 12th Canadian Symposium on Catalysis, Banff, Alberta, Canada, May 25-28,1992 edited by K.J. Smith and E.C. Sanford Angle-Resolved Photoemission.Theory and Current Applications edited by S.D. Kevan New Frontiers in Catalysis, Parts A-C. Proceedings of the 10th International Congress on Catalysis, Budapest, Hungary, 19-24 July, 1992 edited by L. Guczi, F. Solymosi and P. Tetenyi Fluid Catalytic Cracking: Science and Technology edited by J.S. Magee and M.M. Mitchell, Jr. New Aspects of Spillover Effect in Catalysis. For Development of Highly Active Catalysts. Proceedings of the Third International Conference on Spillover, Kyoto, Japan,August 17-20,1993 edited by T. Inui, K. Fujimoto, T. Uchijima and M. Masai Heterogeneous Catalysis and Fine Chemicals III. Proceedings of the 3rd International Symposium, Poitiers, April 5 - 8,1993 edited by M. Guisnet, J. Barbier, J. Barrault, C. Bouchoule, D. Duprez, G. Perot and C. Montassier Catalysis: An Integrated Approach to Homogeneous, Heterogeneous and Industrial Catalysis edited by J.A. Moulijn, P.W.N.M. van Leeuwen and R.A. van Santen Fundamentals of Adsorption. Proceedings of the Fourth International Conference on Fundamentals of Adsorption, Kyoto, Japan, May 17-22,1992 edited by M.Suzuki Natural Gas Conversion II. Proceedings of the Third Natural Gas Conversion Symposium, Sydney, July 4-9,1993 edited by H.E. Curry-Hyde and R.F. Howe New Developments in Selective Oxidation II. Proceedings of the Second World Congress and Fourth European Workshop Meeting, Benalmadena, Spain, September 20-24,1993 edited by V. Cortes Corberan and S. Vic Bellon Zeolites and Microporous Crystals. Proceedings of the International Symposium on Zeolites and Microporous Crystals, Nagoya, Japan, August 22-25,1993 edited by T. Hattori and T. Yashima Zeolites and Related Microporous Materials: State of the Art 1994. Proceedings of the 10th International Zeolite Conference, Garmisch-Partenkirchen, Germany, July 17-22,1994 edited by J. Weitkamp, H.G. Karge, H. Pfeifer and W. Holderich Advanced Zeolite Science and Applications edited by J.C. Jansen, M. Stocker, H.G. Karge and J.Weitkamp Oscillating Heterogeneous Catalytic Systems by M.M. Slin'ko and N.I. Jaeger Characterization of Porous Solids III. Proceedings of the lUPAC Symposium (COPS III), Marseille, France, May 9-12,1993 edited by J.Rouquerol, F. Rodriguez-Reinoso, K.S.W. Sing and K.K. Linger
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Catalyst Deactivation 1994. Proceedings of the 6th International Symposium, Ostend, Belgium, October 3-5,1994 edited by B. Delmon and G.F. Froment Catalyst Design for Tailor-made Polyolefins. Proceedings of the International Symposium on Catalyst Design for Tailor-made Polyolefins, Kanazawa, Japan, March 10-12,1994 edited by K. Soga and M. Terano Acid-Base Catalysis II. Proceedings ofthe International Symposium on Acid-Base Catalysis II, Sapporo, Japan, December 2-4,1993 edited by H. Hattori, M. Misono and Y. Ono Preparation of Catalysts VI. Scientific Bases for the Preparation of Heterogeneous Catalysts. Proceedings ofthe Sixth International Symposium, Louvain-La-Neuve, September 5-8,1994 edited by G. Poncelet, J. Martens, B. Delmon, P.A. Jacobs and P. Grange Science and Technology in Catalysis 1994. Proceedings ofthe Second Tokyo Conference on Advanced Catalytic Science and Technology, Tokyo, August 21-26,1994 edited by Y. Izumi, H. Arai and M. Iwamoto Characterization and Chemical Modification ofthe Silica Surface by E.F Vansant, P. Van Der Voort and K.C. Vrancken Catalysis by Microporous Materials. Proceedings of ZEOCAT'95,Szombathely, Hungary, July9-13,1995 edited by H.K. Beyer, H.G.Karge, I. Kiricsi and J.B. Nagy Catalysis by Metals and Alloys by V. Ponec and G.C. Bond Catalysis and Automotive Pollution Control III. Proceedings of the Third International Symposium (CAPoC3), Brussels, Belgium, April 20-22,1994 edited by A. Frennet and J.-M. Bastin Zeolites: A Refined Tool for Designing Catalytic Sites. Proceedings of the International Symposium, Quebec, Canada, October 15-20,1995 edited by L. Bonneviot and 8. Kaliaguine Zeolite Science 1994: Recent Progress and Discussions. Supplementary Materials to the 10th International Zeolite Conference, Garmisch-Partenkirchen, Germany, July 17-22,1994 edited by H.G. Karge and J. Weitkamp Adsorption on New and Modified Inorganic Sorbents edited by A. D^browski and V.A. Tertykh Catalysts in Petroleum Refining and Petrochemical Industries 1995. Proceedings ofthe 2nd International Conference on Catalysts in Petroleum Refining and Petrochemical Industries, Kuwait, April 22-26,1995 edited by M. Absi-Halabi, J. Beshara, H. Qabazard and A. Stanislaus 11th International Congress on Catalysis - 40th Anniversary. Proceedings ofthe 11th ICC, Baltimore, MD, USA, June 30-July 5,1996 edited by J. W. Hightower, W.N. Delgass, E. Iglesia and A.T. Bell Recent Advances and New Horizons in Zeolite Science and Technology edited by H. Chon, S.I. Woo and S. -E. Park Semiconductor Nanoclusters - Physical, Chemical, and Catalytic Aspects edited by P.V. Kamat and D. Meisel Equilibria and Dynamics of Gas Adsorption on Heterogeneous Solid Surfaces edited by W. Rudzihski, W.A. Steele and G. Zgrablich Progress in Zeolite and Microporous Materials Proceedings ofthe 11th International Zeolite Conference, Seoul, Korea, August 12-17,1996 edited by H. Chon, S.-K. Ihm and Y.S. Uh Hydrotreatment and Hydrocracking of Oil Fractions Proceedings ofthe Istlnternational Symposium/6th European Workshop, Oostende, Belgium, February 17-19,1997 edited by G.F. Froment, B. Delmon and P. Grange
