Catalysts in Petroleum Refining and Petrochemical Industries 1995 (Studies in Surface Science and Catalysis)

Studies in Surface Science and Catalysis 100 CATALYSTS IN PETROLEUM REFINING AND PETROCHEMICAL INDUSTRIES

1995

Sponsored by the Kuwait Institute for Scientific Research, the Kuwait Foundation for the Advancement ofScience, the Kuwait National Petroleum Company, the Kuwait Petroleum Corporation, Kuwait University, the GulfCooperation Council, the Public Authority for Applied Education and Training, the Petrochemical Industries Company, and the Organization ofArab Petroleum Exporting Countries.

Studies in Surface Science and Catalysis Advisory Editors: B. Delman and J.T. Yates Vol. 100

CATALYSTS IN PETROLEUM REFINING AND PETROCHEMICAL INDUSTRIES 1995 Proceedings of the 2nd International Conference on Catalysts in Petroleum Refining and Petrochemical Industries, Kuwait, April 22-26, 1995

Editors

M. Absi-Halabi, J. Beshara, H. Oabazard and A. Stanislaus Petroleum, Petrochemicals and Materials Division, Kuwait Institute for Scientific Research, Kuwait

1996 ELSEVIER Amsterdam - Lausanne - New York - Oxford - Shannon - Tokyo

ELSEVIER SCIENCE B.V. Sara Burgerhartstraat 25 ~O.

Box 211, 1000 AE Amsterdam, The Netherlands

ISBN 0-444-82381-6 © 1996 Elsevier Science B.V. All rights reserved.

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

FOREWORD

The 2nd International Conference on Catalysts in Petroleum Refining and Petrochemical Industries was held in Kuwait during the period April 22-26, 1995, under the auspises of H.H. Sheikh Saad A1-Abdullah A1-Salem A1-Sabah, Kuwait's Crown Prince and Prime Minister. The 1st conference was also held in Kuwait in 1989. The present conference was scheduled to be held in 1993; however, it was postponed due to the events that encompassed Kuwait and the Gulf region in 1990-1991. The patronage of the conference, the organizing bodies, and the selective emphasis on the role of catalysts in the petroleum and petrochemical industries reflect the keen interest of the countries in the region in actively contributing to the development of these industries. Petroleum-related industries are the main economic activities of most countries in the region. The refining capacity in the Gulf Region exceeds 5 MM barrels/day and includes some of the most sophisticated petroleum refining schemes in the world. The basic petrochemical industry has been also growing steadily in the region since the early eighties. The conference was attended by around 300 specialists in the catalysis field from both academia and industry from over 30 countries. It provided a forum for the exchange of ideas between scientists and engineers from the region with their counterparts from the industrialized countries. A total of 62 scientific papers were presented. The papers were carefully selected to include a blend of fundamental and applied research, and industrial experience. Such a blend was thought to be essential for providing the participants from both industry and academia with a chance to become familiar with the challenges facing each group and the actions taken to meet them. A number of keynote speakers, carefully selected from high ranking officials, policy makers, and multinational company representatives, were also invited to address the conference. The keynote presentations, which are published as a separate volume by the Kuwait Institute for Scientific Research, provided the participants with an overview of the directions the petroleum and petrochemical industries will take over the next decade. The program of the conference included a field visit to one of Kuwait's most modem refineries. A trip was also organized to one of Kuwait's oil fields. The partipants had a chance to observe oil lakes and the extent of the damage incurred by the blowing up of Kuwait's oil wells. The success of the conference is perhaps difficult for the organizers to assess. However, the quality of the papers in this volume provides some indication. Another indication is the keen interest and encouragement expressed by numerous participants in attending the next meeting, which will be held in Kuwait in 1998.

The Editors

vi

P

R

E

F

A

C

E

Catalysis plays an increasingly critical role in modern petroleum refining and basic petrochemical industries. The market demands for and specifications of petroleum and petrochemical products are continuously changing. They have impacted the industry significantly over the past twenty years. Numerous new refining processes have been developed and significant improvements were made on existing technologies. Catalysts have been instrumental in enabling the industry to meet the continuous challenges posed by the market. As we enter the 21st century, new challenges for catalysis science and technology are anticipated in almost every field. Particularly, better utilization of petroleum resources and demands for cleaner transportation fuels are major items on the agenda. It is against this background that the 2nd International Conference on Catalysts in Petroleum Refining and Petrochemical Industries was organized. The papers from the conference were carefully selected from around 100 submissions. The papers were refereed in terms of scientific and technical content and format in accordance with internationally accepted standards. They were a mix of reviews providing an overview of selected areas, original fundamental research results, and industrial experiences. The papers in the proceedings were grouped in the following sections for quick reference: -

Plenary Papers Hydroprocessing of Petroleum Residues and Distillates Fluid Catalytic Cracking Oxidation Catalysis Aromatization & Polymerization Catalysis Catalyst Characterization and Performance

The plenary papers were mostly reviews covering important topics related to the objectives of the conference. The remaining sections cover various topics of major impact on modern petroleum refining and petrochemical industries. A large number of papers dealt with hydroprocessing of petroleum distillates and residues which reflects the concern over meeting future sulfur-level specifications for diesel and fuel oils. The task of editing this volume was facilitated by the efforts of the International Advisory Committee and the Scientific Committee of the conference who reviewed all the papers. The editorial board gratefully acknowledge this effort; the cooperation, time and effort of all authors; and the management of the Kuwait Institute for Scientific Research for allocating the required resources to prepare the manuscript of this volume.

T h e

E d i t o r s

vii

TABLE OF CONTENTS Foreword Preface Organizing Committees Acknowledgements

v vi xi xii

PLENARY LECTURES

Control of Catalyst Performance in Selective Oxidation of Light Hydrocarbons: Catalyst Design and Operational Conditions B. Delmon, P. Ruiz, S. R. G. Carrazan, S. Korili, M A. Vicente Rodriguez and Z. Sobalik

1

Vanadium Resistant Fluid Cracking Catalysts M L. Occelli

27

Metal Clusters in Zeolites: Nearly Molecular Catalysts for Hydrocarbon Conversion B. C. Gates

49

Catalyst Deactivation D. L. Trimm

65

Preparation and Catalysis of Highly Dispersed Metal Sulfide Catalysts for Hydrodesulfurization Y O"kamoto

77

New Developments in Olefin Polymerization with Metallocene Catalysts W Kaminislry and A. Duch

91

HYDROPROCESSING OF PETROLEUM RESIDUES AND DISTILLATES

New Developments in Hydroprocessing J. W M Sonnemans

99

Optimizing Hydrotreater Catalyst Loadings for the Upgrading of Atmospheric Residues J. Bartholdy and B. H Cooper

117

Hydrotreatment of Residuals Using a Special NiMo-Alumina Catalyst A. Morales and R. B. Solari

125

Residue Hydroprocessing: Development of a New Hydrodemetallation (HOM) Catalyst o. K Bhan and S. E. George Commercial Experience in Vacuum Residue Hydrodesulfurization

135

H Koyama, E. Nagai, H Torii, andM Kumagai

147

Comparison of Operational Modes in Residue Hydroprocessing M de Wind, Y Miyauchi, and K Fujita

157

Mina Abdulla Refinery Experience with Atmospheric Residue Desulfurization (ARDS) A. AI-Nasser, S.R Chaudhuri and S. Bhatacharya.

171

Cosmo Resid Hydroconversion Catalyst - Catalyst Combination Technology Y Yamamoto, Y Mizutani, Y Shibata, Y KitouandH yamazaki

Influence of Catalyst Pore Size on Asphaltenes Conversion and Coke-Like Sediments Fomation During Catalytic Hydrocracking of Kuwait Vacuum Residues A. Stanislaus, M Absi-Halabi, and Z Khan Origin of the Low Reactivity of Aniline and Homologs in Hydrodenitrogenation M Callant, K Holder, P. Grange, and B. Deln10n

181

189 199

viii Deep HDS of Middle Distillates Using a High Loading CoMo Catalyst S. Mignard, S. Kasztelan, M Dorbon, A. Billon, and P. Sarrazin

209

Environmentally Friendly Diesel Fuels Produced from Middle Distillates Generated by Conversion Processes R Zamfirache and1 Blidisel

2 17

Factors Influencing the Performance ofNaphtha Hydrodesulfurization Catalysts J. A. Anabtawi, S. A. Ali, M A. B. Siddiqui and S. M J. Zaidi

225

Hydrocracking of Paraffinic Hydrocarbons over Hybrid Catalysts Containing H-ZSM-5 Zeolite and Supported Hydrogenation Catalyst 1 Nakamura and K Fujimoto

235

Effect of Presulfiding on the Activity and Deactivation of Hydroprocessing Catalysts in Processing Kuwait Vacuum Residue M Absi-Halabi, A. Stanislaus, A. Qamra and S. Chopra Continuous Developments of Catalyst Off-Site Regeneration and Presulfiding P. Dufresne, F. Valeri, and S. Abotteen

243 253

The Production of Large Polycyclic Aromatic Hydrocarbons During Catalytic Hydrocracking J. C. Fetzer

263

Fouling Mechanisms and Effect of Process Conditions on Deposit Formation in H-Oil Equipment M A. Bannayan, H K Lemke, and W. K Stephenson

273

Bed Expansion and Product Slate Predictions from H-Oil Process via Neural Netwrok Modelling E. K T. Kam, M M AI-Mashan, and H Dashti

283

Renewed Attention to the EUREKA Process: Thermal Cracking Process and Related Technologies for Residual Oil Upgrading T. Takatsuka, R Watari and H Hayakawa

293

FLUID CATALYTIC CRACKING

New Catalytic Technology for FCC Gasoline Sulfur Reduction without Yield Penalty U Alkemade and T. J. Dougan The Influence of Feedstocks and Catalyst Formulation on the Deactivation of FCC Catalysts R. Hughes, G. Hutchings, C. L. Koon, B. McGhee, and C. E. Snape Resid FCC Operating Regimes and Catalyst Selection P. 0' Connor and S. J. yanik Novel FCC Catalyst Systems for Resid Processing U Alkemade and S. Paloumbis

303 313 323 339

Probing Internal Structures of FCC Catalyst Particles: From Parallel Bundles to Fractals R Mann and U A. EI-Nafaty

355

Development of Microscale Acitivity Test Strategy for FCC Process Economics Enhancement O. H J. Muhammad

365

OXIDATION CATALYSIS

Partial Oxidation of C2-C4 Alkanes into Oxygenates by Active Oxygen Generated Electrochemically on Gold through Yttria-stabilized Zirconia K Takehira, K Salo, S. Hamakawa, T. Hayakawa, and T. Tsunoda

375

ix The Effects of Gas Composition and Process Conditions on the Oxidative Coupling of Methane over Li/MgO Catalyst S. M AI-Zahrani and L. Lobban Study on the Active Site Structure ofMgO Catalysts for Oxidative Coupling of Methane K Aika and T. Karasuda Various Characteristics of Supported CoPe on A1 20 3, Si02 and Si02-AI20 3 as Selective Catalysts in the Oxidative Dehydrogenation of Cyclohexene S. A. Hasan, S. A. Sadek, S. M Faramawy, and M A. Mekewi

383 397

407

Dehydrogenation of Propane over Chromia/Alumina: a Comparative Characterization Study of Fresh and Spent Catalysts A. Rahman and M Ahmed

419

Deactivation Mechanism of a Chromia-Alumina Catalyst by Coke Deposition F Mandani, E. K T. Kam, and R Hughes

427

Investigation of Synthesis Gas Production from Methane by Partial Oxidation over Selected Sream Refonning Commercial Catalysts H AI-Qahtani

437

AROMATIZATION & POLYMERIZATION CATALYSIS

Aromatization of Butane over Modified MFI-Type Zeolite Catalysts T. Yashima, S. Ekiri, K Kato, T. Komatsu, and S. Namba Development of Light Naphtha Aromatization Process Using A Conventional Fixed Bed Unit S. Fukase, N Igarashi, K Kalo, T. Nomura, and Y: Ishibashi Improvement in the Perfonnance of Naphtha Refonning Catalysts by the Addition of Pentasil Zeolite J. N Beltramini and R Fang Zeolite Catalysts in Upgrading of Low Octane Hydrocarbon Feedstocks to Unleaded Gasoline VG. Stepanov, KG. lone, andG. P. Snytnikova

447 455

465 477

Catalysts for Cyclization of C6-Alkanes N Ph. Toktabaeva, G. D. Zakumbaeva, and L. B. Gorbacheva

483

High Quality Gasoline Synthesis by Selective Oligomerization of Light Olefins and Successive Hydrogenation T. Inui and J. B. Kim Hydrogenation of Aromatic Compounds Related to Fuels over a Hydrogen Storage Alloy S. Nakagawa, T. Ono, S. Murata, M Nomura, and T. Sakai A Theoretical Study of Ethylene Oligomerization by Organometallic Nickel Catalysts L. Fan, A. Krzywicki, A. Somogyvari, and T. Ziegler IFP-SABIC Process for the Selective Ethylene Dimerization to Butene-l

499

F A. Al-Sherehy

515

489

507

CATALYST CHARACTERIZATION AND PERFORMANCE

Cobalt Containing ZSM5 Zeolites - Preparation, Characterization and Structure Simulation A. Jentys, A. Lugstein, O. El-Dusouqui, H Vinek, M Englisch and J. A. Lercher

525

Acid-base Property of Some Zeolites and their Activity for Decomposition of n-Hexane S. Tsuchiya

535

x Reduction and Sulfidation Properties of Iron Species in Fe-Treated V-Zeolites for Hydrocracking Catalysts K Inamura and R Iwamoto

Preparation of Highly Active Zeolite-Based Hydrodesulfurization Catalysts: Zeolite-Supported Rh Catalysts M Sugioka, C. Tochiyama, F Sado, and N Maesaki High-Dispersed Supported Catalysts on Basis of Monodispersed Pt-Soles in Processes Reductive Transformation of Hydrocarbons N A. Zakarina and A. G. Akkulov Infrared Spectroscopy of CO/H2 Coadsorption on NilAl20 3 Hydrotreating Catalysts: Evidence for Perturbed Metal Sites M 1 Zaki List ofparticipants Author Index

543

551

559

569 579 595

xi ORGANIZING COMMITTEE Jasem AI Besharah Khaled A1 Muhailan Mamun Absi Halabi Abbas Ali Khan Anwar Abdullah Taher A1 Sahaf Mohammad Ali Abbas Abdul-Karim Abbas Bader AI Safran Faisal Mandani Hassan Qabazard Mubarak AI Adwani AI Tayeb Wenada

Chairman Rapporteur Coordinator Member Member Member Member Member Member Member Member Member Member

KISR KFAS KISR KFAS GCC KU KPC KNPC PIC PAAET KISR KISR OAPEC

INTERNATIONAL ADVISORY COMMITTEE Mamun Absi Halabi David L. Trimm Bernard Delmon Burce C. Gates Walter Kaminsky Yasuaki Okamoto Mario L. Occelli Henrik Topsoe

Chairman Member Member Member Member Member Member Member

Kuwait Australia Belgium USA Germany Japan USA Denmark

SCIENTIFIC COMMITTEE Taher A1 Sahaf Anthony Stanislaus Abdullah S. A1 Nasser Jaleel Shishtary Erdogan Alper Mustafa A. A. Gholoum Faisal Mandani Ezra Kam

Chairman Rapporteur Member Member Member Member Member Member

KU KISR Mina Abdulla~NPC Mina A1 Ahmadi/KNPC KU Shuaiba/KNPC PAAET KISR

. ~

Xll

ACKNOWLEDGEMENTS The Organizing Committee was deeply honored by the patronage of//. H. The Crown Prince and Prime Minister Sheikh Saad A1-Abdullah A1-Salem AI-Sabah, which reflects his keen interest in science and technology. The Committee is also grateful for the financial support of the Kuwait Institute for Scientific Research, the Kuwait Foundation for the Advancement of Science, the Kuwait National Petroleum Company, the Kuwait Petroleum Corporation, Kuwait University, the Gulf Cooperation Council, Public Authority for Applied Education and Training, the Petrochemical Industries Company and the Organization of Arab Petroleum Exporting Countries. The Committee would like also to express gratitude for the efforts of the Japan Petroleum Institute in coordinating and supporting the participation of prominent Japanese scientists in this event. The Committee would like also to extend its deep appreciation for the effort and time put forth by the distiguished keynote speakers, namely H.E. Mr. Hisham Al-Nazer, H.E. Mr. Erwin Valera, H.E. Mr. Lulwanu Lukman, Mr. Abdullatif AI-Hamad, Mr. Charles DiBona, Mr. John Yimoyines, Mr. J. Kent Murray, Mr. Mahmoud Yusef, Mr. Moayad Al-Qurtas, Mr. Khalaf A1-Oteibeh, Mr. Khaled Buhamra, and Mr. Nader Sultan. The Organizing Committee are also appreciative of the efforts of the members of the International Advisory Committee and the Scientific Committee for their thorough work in selecting and refereeing the submitted papers. The Committee also acknowledges the help and guidance provided by Elsevier Science Publishing Company and the advisory editors of this series in preparing this proceedings. We would like to thank our colleagues at the Kuwait Institute for Scientific Research, the Kuwait Ministry of Oil, and the chairmen and cochairmen of the sessions, who provided unlimited assistance at times when it was badly needed. Finally, we feel deeply indebted to the participants who enriched the meeting with their serious discussions till the end. DR. J A S E M B E S H A R A

CHAIRMAN, ORGANIZING COMMITTEE

Catalysts in Petroleum Refining and Petrochemical Industries 1995

M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All rights reserved.

C O N T R O L OF CATALYST P E R F O R M A N C E IN SELECTIVE OXIDATION OF L I G H T H Y D R O C A R B O N S : C A T A L Y S T D E S I G N AND O P E R A T I O N A L CONDITIONS B. Delmon, P. Ruiz, S.R.G. Carraz~in, S. Korili, M.A. Vicente Rodriguez, Z. Sobalik Catalyse et Chimie des Mat6riaux Divis6s, Universit6 Catholique de Louvain, Place Croix du Sud 2/17 - 1348 Louvain-la-Neuve, Belgium This paper is an attempt to summarize the situation with respect to the selective catalytic oxidation of light alkanes using heterogeneous catalysts. Methane oxidation reactions and the oxidation of butane to maleic anhydride will only be alluded to occasionally, because they have been reviewed in detail in a large number of papers. We shall first show that it is still far from clear which are the families of catalysts to be used for the various reactions: mainly oxidative dehydrogenation or oxidation to oxygen-containing molecules of ethane, propane or isobutane. Much research is still necessary for understanding the mechanisms leading to high selectivity. In this context, we shall suggest that many concepts inherited from the development in selective oxidation and ammoxidation of olefins are probably of little use. Conversely, much emphasis has to be laid on new data which opens promising perspectives, namely (i) the occurrence of cooperation effects between two (or several) separate phases and especially the role of spillover oxygen and the so-called "remote control" and (ii) the occurrence of homogeneous non-catalysed reactions which occur at temperatures only slightly higher than the catalytic ones and correspond to similar selectivities. This suggests that research on selective catalytic oxidation, to be effective, should be comprehensive: it should continue to involve a search for new active phases and efforts to improve the already known catalysts. But research should also include investigations on the role of spillover oxygen, the nature of this oxygen (more or less electrophilic), the donors that can generate it, and the way this spillover oxygen reacts with the catalytic surface. Research should also contemplate the problem of how homogeneous and heterogeneous reactions proceed simultaneously or consecutively. In parallel with these research lines, chemical engineering must develop new concepts and new reactors. Recent spectacular results in methane coupling or oxidative dehydrogenations show that considerable progress can be made if the problem of light alkane selective oxidation benefits from a multifacetted approach. 1. I N T R O D U C T I O N Making valuable products from light hydrocarbons is presently one of the major challenges for the petroleum and petrochemical industries. Among the various processes able to transform light hydrocarbons to useful products, catalysis has a major role to play. Conceptually, the cheapest and easiest route is through catalytic oxidation. The reason is that oxygen (pure or in air) is cheap and possesses the high reactivity necessary to activate saturated hydrocarbons. For that type of activation, heterogeneous and homogeneous catalysis are competing. Nevertheless, the preference in principle goes to heterogeneous catalysis, especially if very large quantities have to be transformed, as in the case of methane.

On the whole, a continuous progress towards a more selective oxidation of light saturated hydrocarbons is observed, and recent announcements demonstrate that dramatic progress can be made even in the very difficult case of methane activation, using either heterogeneous or homogeneous catalysts. The activation of light saturated hydrocarbons becomes increasingly more difficult as the molecules become smaller, with methane reactions being the most difficult to control. On the other hand, the occurrence of non-catalysed gas phase oxidation makes selectivity control very complicted. This is a problem common to almost all oxidations, unless one of the products is extremely stable 9examples are unsaturated nitriles (e.g. acrylonitrile in the ammoxidation of propane) or maleic anhydride (in the oxidation of butane). There is a parallel trend in the changes of reactivity with molecular weight in catalytic and non catalytic (gas phase) oxidation. The challenge to catalysis to achieve selective reactions at lower temperature is thus equally important for all light hydrocarbons. The activation of very light hydrocarbons (propane, ethane and methane) in the presence of oxygen has been achieved only at temperatures substantially or much higher than those used in the reactions of other hydrocarbons. There is however little doubt that some mechanistic similitudes exist and that the vast body of knowledge accumulated on the reaction of other hydrocarbons (including unsaturated ones) with oxygen will be useful for improving the efficiency of these difficult reactions. Nevertheless, the outstanding commercial success of the oxidations and ammoxidations of light olefins and that of the oxidation of butane to maleic anhydride has directed the fundamental research of the largest number of investigators to topics which are probably not the most relevant to the new challenges set by the selective oxidation of light alkanes. A much broader approach has certainly to be taken, compared to that used in former investigations. It is the aim of this contribution to highlight a few promising directions for research in the area of selective reactions of light alkanes with oxygen (oxidation and oxidative dehydrogenation). We shall emphasize three aspects: (i) new concepts have been recently developed in a field which seemed to be well established, namely the catalytic oxidation of olefins and butane, but where new powerful methods of action have been discovered. We shall show that these new concepts are applicable to the catalytic oxidation of the light saturated hydrocarbons, namely containing from one to five carbon atoms. We shall present, in some cases for the first time, results which strongly suggest that a cooperation between distinct phases in oxidation catalysts could play an important role in the oxidation of light hydrocarbons, even perhaps in the coupling of methane. (ii) we shall suggest, on the basis of new results from our and other laboratories, that the intervention of non catalysed gas phase reactions must be accounted for and should be investigated carefully. (iii) we shall also show that catalyst discovery and development in the field of heterogeneous oxidation of light hydrocarbons should be accompanied by innovative developments on the chemical engineering side. Before examining specifically these points, we shall "set the stage", namely attempt to give an overview of the results published in literature on the selective reactions of light alkanes with oxygen. The largest part of the contribution will consist in a critical overview of the parameters traditionally believed to be crucial for activity and selectivity. We shall show that one parameter, which probably has the largest importance, has been almost completely forgotten: this is the ability for separate phases, inactive or poorly active, to enhance the activity of potentially active and selective phases, via an oxygen spillover process. Results will be presented which strongly suggest that the same sort of cooperation between phases can operate in the reactions of light alkanes. At the end, we shall suggest that the existence of gas phase oxidation reactions, the occurrence of the phase cooperation mentioned above and the other particularities of light alkane oxidation are about to trigger new developments in chemical engineering which will probably be as innovative and crucial for viable processes as the development of fluidized bed reactors for oxidation or ammoxidation, and riser reactors (in the

case of butane oxidation) has been during the remarkable development of catalytic oxidation in the last 25 years. 2 . C A T A L Y S T S A C T I V E IN T H E ALKANES W I T H OXYGEN

SELECTIVE

REACTION

OF

LIGHT

The variety of catalysts which have been claimed to activate light alkanes is very large. The only conspicuous exception concerns the reaction of butane to maleic anhydride; this is, however, a special case considering the high stability of the product, namely maleic anhydride. But this large diversity of formulations exists even in the ammoxidation of propane to acrylonitrile, although the product is also particularly stable in this case. It cannot be therefore concluded that given oxidation reactions take place only on a single family of catalysts. In what follows, we present a series of tables concerning various reactions of light alkanes with oxygen. We wish, however, to underline the fact that the data contained in the tables are by no means comprehensive. We have selected them in view of our objectives, namely (i) to underline the variety of formulations proposed for a single reaction, (ii) to extract from these data a few conclusions and (iii) to speculate on the possible importance of some parameters. We have avoided to overburden the tables with information on reaction conditions. These are indeed very different, and correlating them with catalyst composition has little usefulness for the moment (except perhaps for propane ammoxidation, where investigation is more advanced). We do not present data concerning either methane or butane. In the case of methane oxidation and oxidative coupling, innumerable articles (more than 1000) have been published, together with many review papers. Concerning butane, the numerous articles and review papers dealing with oxidation of maleic anhydride obscure the few scattered articles dealing with oxidative dehydrogenation; dehydrogenation of butane has mainly been done in reactions without oxygen. In the tables, we omit the chemical symbol of oxygen and list only the elements combined with oxygen in the catalysts, or oxygen when it is present in a phase indicated as such by the authors (e.g., supports: MgO, SIO2), except if there is good ground to believe that well defined metal oxide entities are crucial for catalytic activity (e.g., VO...). In addition to the systems listed in Table I for the oxidative dehydrogenation of ethane, other systems have been tested because they have proven to be active in other alkane oxidations; this is particularly the case of many catalysts used in the oxidative coupling of methane, VPO and magnesium phosphate catalysts (butane oxidation and propane dehydrogenation, respectively) and MoVO catalysts. Various zeolites have also been tested. This table, the largest to be presented here, perfectly illustrates the fact that no formulation seems convincingly better than the others. In the oxidative dehydrogenation of propane (Table II), the various magnesium vanadates have been the object of many studies, but other systems seem to have comparable performances (systems based on cerium, niobium, or vanadium, molybdates and noble metals on monoliths used with very short contact time). Because the direct dehydrogenation of isobutane to isobutene is now in operation industrially, it is not surprising that relatively few publications deal with the corresponding oxidative dehydrogenation to isobutene (about 20 in the past 6 years). On the whole, the catalysts used are similar to those mentioned in the previous tables: phosphates, chromates, molybdates. Active carbon has also been mentioned, but it is hard to imagine that the catalyst could work a long time in the presence of oxygen. Table III gives two examples of the results mentioned in literature. Mention has been made of the selective oxidation (yield = 65%) of isobutene on UV activated TiO2 [50].

Table I. Ethane oxidative dehydrogenation to ethylene

Catalyst Ca-Ni ceramic foam monoliths + Pt, Rh, Pd Cd-La-A1 MgO based catalysts Ce2(CO3)3 Mo-Si, Si-W or P-W/A1203 Cr-Zr-P Li-Na-Mg Li/MgO Sr-Ce-Yb Na-Mn zeolites La203-B aF2 heteropolyacid Pt-cordierite (electrocatalytic) Mo-V-Nb-Sb Mo-V-Nb-Sb-M Na-K-Zr Li-Ti+Mo, Sn or Sb Li-Ti-Mn V-P-U Zn2TiO4+Bi Co-P+promoter Mo-Te Mo-Bi-Ti-Mn-Si Li/M~D+promoter

Conversion % 25 80

Yield %

35 45 20-30 38 75-79

Selectivity % 93.6 70 84 73.7 90 90 50-60 86 70

49 76.8

86.5 86-90 84.7 76-98 72 96.9 72-82 86 86.4 86.3 74.6

59.1 70 10.6 22-57 34 85 54.9 46.9 22.7 71.2 68.1 75

67.5 80.5 100 100 76

Ref. 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27

It seems that very few investigations concern the oxidation or oxidative dehydrogenation of C5 alkanes. Oxidative dehydrogenation of isopentane to isoprene has been mentioned. Two articles deal with MnO2, CoO/CaO3, NaOH/A1203, but in the presence of HI [51,52]; this obviously suggests the intervention of gas-phase reactions. The yields (Y) in isobutene were relatively high (e.g., Y = 50-60% with a selectivity of 65 to 95%). Pentane can also produce maleic anhydride and phthalic anhydride [53-57]. Considering in a general way the activation of light alkane by oxygen, the ammoxidation of propane has certainly not to be forgotten. This process is already under industrial development. If we try to get an overview of the recent work on the selective reactions of light alkanes with oxygen, two remarks may be made: 9 Several lines have been followed, all inspired by former successful lines of research. It is striking that the proposed catalysts are generally similar to those previously used in the selective reaction of alkanes with oxygen: oxidative coupling of methane or oxidation of butane to maleic anhydride. Many of them are also similar to catalysts used for the reactions of olefins with oxygen (molybdates) or for dehydrogenation without oxygen (chromium containing catalysts). Because of the success of vanadyl phosphate in butane oxidation, investigators tend to focus on vanadium containing catalysts also in the case of other alkanes. Nevertheless, the data available do not seem to exclude any other formulation. 9 On the other hand, the reaction of ethane, propane, isobutane, and pentanes with oxygen described until now are poorly selective at high, and even at moderate conversions. One cannot

exclude the empirical discovery of completely new catalysts with outstanding performances. However, a more systematic approach may also help find satisfactory catalysts. An in-depth understanding of the principles involved in catalytic selective oxidation is necessary to improve activity, selectivity and resistance to ageing of catalysts. This is true as well for the catalysts to be perhaps discovered as for those already cited. Table II. Propane oxidative dehydrogenation Reactant Product Catalyst Propane Propane

Propene Propene

Propane

Propene

Propane

Propene Ethene

Propane Propane

Propene Propene

Nb based catalysts VMg, VMg+Ag, Electrochemical pumping of oxygen (EOP) VMg and chloride of Cu +, Li +, Ag+, Cd2+ noble metals (Pt,Pd) on ceramic foam monoliths at short contact time, 5

Conversion % 7 10

Yield Selectivity Ref. % % 85 28 84, 86.9 29

23.1

30

100

65 (total olefins)

31

19

60

32 33

23 23 23 25 20 43 41.3 40.3 50

46 59 49 60 62 34.5 81.1 66.2 50

ms

Propane Propane

Propene Propene

Co0.95MoO4 V/Mg= 2/1 2/2 2/3 VMgffiO2 NiMoO4

Propane Propane Propane Butane Hexane Propane Propane Propane Propane

Propene Propene Alkenes

CeO2]2CeF3/Cs20 FeV-supported Nd203 Vanadate catalysts

Propene Propene Propene Propene* Acrylonitrile**

V-Fe-Nd-A1 VMg CeO~CeF3 (NH3)3PO4+ in(NO3)3+ Vanadyl phosphate NiMoOx (a=0.6-1.3; x=number determined by Ni or Mo valency) A1203 supported Pt/Cs/Sm MgV206 (50% V2Os+MgO calcined at 610 ~ CoMoO4/SiO2 NaHO/Na3VO4/A1

Propane

Propene

Propane

Propene

Propane

Propene

Propane Propane

Propene Propene

40.3 10 53.4 12 29

4.1 20.9

12.5 14.8 33.5

66.2

26.7 65

3 6.7 35* 36.7** 18.1

16.6

34 35 36 37 38 39 40 41 42 43

91

44

71

45

77.9 79.8

46 47

The next sections will therefore indicate some of these fundamental aspects and suggest the perspectives that some new f'mdings are opening. Table HI. Isobutane oxidative deh~,drosenation to isobutene. Catalyst Y203 + CeF3 Ni2P207 Zn2P207, Cr4(P207)3, M~2P207

Selectivity (S)

Ref.

high conversion S = 82 % S = 60-70 %

48 49 49

3. PARAMETERS TRADITIONALLY CONSIDERED IN S E L E C T I V E OXIDATION A very large amount of work has been devoted in the past to the oxidation of olefins Callylic" oxidation to unsaturated aldehydes) and butane (to maleic anhydride). This has led to the development of ideas and concepts which are quite naturally used in the new investigations concerning light alkanes. It is necessary to examine these ideas and concepts and to evaluate in a critical way their potential for discovering or improving catalysts in the new field that oxidation of light alkanes constitutes. This will be done here shortly on the basis of classical books or articles [53,58-62].

3.1.

Doping

The idea is to add foreign ions as a solid solution in already active oxide structures. This is logical. The oxidation of hydrocarbons involves oxygen from the catalyst lattice and replenishment of the latter by molecular oxygen after the hydrocarbon molecule has been dehydrogenated or oxidised. This is an oxido-reduction mechanism. Doping by elements of other valencies can in principle change the oxido-reduction level of the surface. More precisely, the really important parameters in the processes are the rates of (i) removal of oxygen by the reaction with the hydrocarbon and (ii) reoxidation by 02. In principle, doping can alter these rates, but very few measurements have been made along this line. Doping can also change surface acidity, a parameter essential for the activation of alkanes. Doping is certainly a good approach for modifying a catalyst. It should however be underlined that it has seldom been verified that the doping elements were really incorporated in the host oxide and did not spontaneously segregate out. There are indeed conspicuous instances of such segregations. For example, it had been claimed that antimony in solid solution in tin oxide SnO2 explained the high activity of Sb-Sn-O catalysts in oxidation. Actually, Sb has a strong tendency to segregate out of SnO2 during the catalytic reaction [63-65]. But in other reactions, there seems to be indeed an effect of doping elements to alter the extent of oxidationreduction in the near surface layers (e.g., cobalt in V-P-O catalysts) [66]. It is therefore advisable to use the doping elements in quantities compatible with complete solubility in the host oxide, and to check that they do not segregate during the catalytic reaction. Cobalt, mentioned above as a useful dopant, could exert a catastrophic effect if segregated as cobalt oxide, because of the high activity of the latter in complete oxidation.

3.2.

Supports

It seems that supports have been considered with much circumspection in the early days of allylic oxidation. Progressively silica began to be used, but it is considered as being generally inert, and permitting only a better dispersion or a higher mechanical strength. However, real supports are progressively appearing in the field of catalytic oxidation, as suggested by the tables presented above. A conspicuous and well known example is TiO2 as a support for V205. The advantage of using TiO2 (e.g., in o-xylene oxidation to phthalic anhydride) is probably not to give isolated surface vanadium atoms, but rather to stabilise islets of a sub-oxide of vanadium, V6013 over a broader range of oxido-reduction conditions [67-69]. This

stabilization has to be attributed to the strong interaction existing between vanadium oxide and the support. But another new factor should probably be taken into consideration. Surface mobile oxygen (spillover oxygen) has an important role in selective oxidation, as will be shown below. Silica is at the bottom of the scale with respect to oxygen surface mobility [70]. ZrO2 is much better, so could presumably be TiO2. We believe that supports could play a more important role in the oxidation of light alkanes than it did in allylic oxidation. But this role will be complex, and include better dispersion of the active phase, stabilisation of the selective phase, control of oxido-reduction, and/or facilitation of oxygen spillover.

3.3.

Epitaxy

Most active catalysts in oxidation contain several phases which act synergetically. This led to the widespread assumption that an epitaxy at the contact between two different phases was of crucial importance. This is undoubtedly a hypothesis to consider. The above example of V205fI'iO2 catalysts indeed suggests that a strong interaction between two phases could make one of them more stable, more active or more selective. But epitaxy should not be taken as a universal explanation, because there are very few proofs (if any) of such epitactic contact between the phases detected in allylic oxidation, even in the case of Sb204-FeSbO4 mixtures whose activity has often been attributed to epitaxy. The explanation of the activity of V-P-O catalysts has long been believed to involve such an epitaxy between two types of vanadium phosphates. But no such proof could be found [71,72]. The explanation of the activity of V-PO is now that a special local structure on the surface of vanadium pyrophosphate, namely twin flat pyramids in adjacent positions oriented in opposite directions is the active sites (fig. 1) [7375], and the epitaxy hypothesis is leaving the scene.

P

.o

I I o~176 I , ~..... ,.;o-',,,-It'.,,

3.4. Formation of monolayers

When the cooperating phases in catalytic oxidation have been found to be clearly \ / ;',, II ,,-'_/.-" i separated in no epitaxial 0 - , . I. , , ~ , O' o O I position, another traditional o .......' II.---'7 i q explanation was put forward, t o p j t namely that an element of one phase migrated to the other P t I .,tOP phase for making a \ : o ..-;, )j',,"'--.. I contamination layer of P , I o ~/-- i1-',, ,)o molecular thickness, or monolayer. The idea has been based on the observation that o . , ....... o MoO3 spreads spontaneously I 1 on 1,-A120 3 and, to a certain P P extent, on bismuth molybdates during calcination in air. But a Figure 1. Structure of vanadium pyrophosphate (VO)2P207. review of literature shows that MoO3 has much lower ability to spread on many other oxides [76]. A contaminating layer is intrinsically fragile, and stable only when its adhesion energy on the other phase is higher than the cohesion energy inside the bulk contaminant. The stability is extremely sensitive to the oxido-reduction conditions. A monolayer appearing upon calcination may not be stable in the conditions where catalysis takes place. A conspicuous example is the

,,:,X,...

v,,,- ..->:t'..--"

"'o

'[/P

v,,/

case of MoO3 mixed with Sb204. Even if dispersion of one element on the oxide of the other is realised, the contamination may disappear during catalysis [63,76]. The common teaching of section 3.3 and the present one, is that it is not excluded that epitaxy or mutual contamination could explain the high activity of oxidation catalysts, but that this has not been proven and that there are good reasons, experimental as well as theoretical, for thinking that such effects are not common.

3 . 5 . Role of the traditional parameters A comprehensive view of the parameters playing a role in the selective oxidation reactions investigated until now is presented elsewhere [77]. When considering all the traditionally discussed parameters, it is clear that very few lines appear for controlling in a comprehensive way catalysts activity, selectivity and resistance to ageing. This is true even with the control of acidity. Removal of undesirable acidic sites leading to poor selectivity is possible to a certain extent [77]. But creating the acido-basic properties necessary for activating alkanes has not been possible until now. The idea which emerges from recent results is rather different, as the example of butane oxidation to maleic anhydride suggests. In full agreement with the new concepts developed in catalysis, the reaction takes place at special sites on the surface (e.g. the twin flat pyramid in VPO shown in fig. 1). This permits a special activated conformation of the reactant in the adsorbed state and makes possible the complicated concerted mechanism necessary for selective transformation. The emphasis is on surface structures, well determined at the atomic scale, which possess the adequate catalytic activity. This is obvious and should have been obvious for many years. What has been overlooked in the past is that surface structures do not necessarity reflect bulk structure: this result has been emphasized by the progress of surface science. Bulk structures, long range order or collective electronic behaviour influence only partially the structure and properties of the limited number of atoms in a special configuration which constitutes the active center. Another teaching of surface science has also been forgotten, namely that surfaces change according to the molecules they are contacted with on the one hand, and all other experimental conditions on the other hand. Position of doping elements, epitaxy, or monolayer depend crucially on all experimental conditions. The conclusion is thus that attention should be given to the local arrangement of limited numbers of atoms which permit the selective reaction and to mechanisms which maintain these structures intact in spite of the oxido-reduction process which continusouly tends to put this structure upside down. The next section will show some typical results of our work in reactions involving oxygen. These results strongly support the correctness of the above views. Our work has permitted to point to the crucial role of hydroxyls, an aspect almost completely ignored before, and to suggest the structure of molybdenum containing phases during catalysis. We have discovered a mechanism by which the steady-state surface can be controlled. The consequences of this discovery will be very briefly outlined. In the subsequent section, we shall suggest how a more comprehensive view of selective oxidations can foster progress in alkane activatien. This will be illustrated by some of our recent results. 4. COOPERATION HYDROCARBONS STRUCTURES

BETWEEN PHASES IN T H E WITH OXYGEN: CONTROL

REACTION OF OF SURFACE

It is well known that the catalysts used for oxidation reactions such as those of propylene to acrolein, isobutene to methacrolein, or for ammoxidations (propylene to acrylonitrile, methylsubstituted benzenic rings to the corresponding aromatic nitriles) contain many components. This complexity in elemental composition is reflected by a complexity in phase composition.

The so-called "multicomponent catalysts" used in selective oxidation are oxides, and they represent the vast majority of catalysts used in this field. All multicomponent industrial catalysts contain several phases. We discovered about ten years ago that simple mechanical mixtures of two oxides had much better performances than those of the two constituents [63-69,7172,76,78]. This is illustrated by fig. 2 in the case of the oxidation of isobutene to methacrolein over mixtures of micron-size MoO3 and t~-Sb204 particles. All experiments were made with the same total quantity of catalysts. The arrows show the increase of yield compared to the simple addition of the individual contributions of the catalyst components. 20

This phenomenon is due to the action exerted by surfaceIsobutene ~ Methacrolein mobile oxygen on the surface of one of the phases, which we call the acceptor (i.e., acceptor of surface-mobile, or spillover e s oxygen: this is MoO3 in the s example of fig. 2). Spillover (3 s oxygen Oso reacts with the surface of the acceptor and, thanks to this reaction, keeps the catalytic sites active and selective. 10 The other phase, often not active or poorly active catalytically, produces the Oso species. This is the donor of spiUover oxygen: aS b 2 0 4 is a typical donor. A comprehensive characterization l of the mixtures before and after catalytic test permitted to exclude any other explanation, such as mutual contamination, formation of new solid phases, bifunctional 0.0 0.5 1.0 catalysis, bulk diffusion, etc., in the majority of cases investigated Moo3 [63]. The occurrence of a surface (mass) migration of oxygen from a s eo ..oo3 donor (t~-Sb204) to an acceptor (MOO3) has beeen shown directly Figure 2. Synergy between o~-Sb204 and MoO3 particles using labelled oxygen [79-82]. in the selective oxidation of isobutene to methacrolein. The Another example, that of figure concerns yields (namely conversion x selectivity) in experiments where conversion was always below 25%. mixtures of o~-Sb204 with SnO2, The catalysts were prepared by mixing the powders of a- very conspicuously shows that Sb204 and MOO3, prepared separately, as a stirred the action of spillover (donated suspension in n-pentane, and evaporating n-pentane. The by a- S b 2 0 4) modifies the same overall weight of mixture was used for all selectivity of the active sites compositions and the experimental conditions were identical situated on SnO2 (the acceptor in [63,78]. the present case) (fig. 3).

I

10

It had been believed for long that the best oxidation catalysts were oxides associating two or several elements in a ~. b,lethacrolein Isobutene given mixed oxide structure, like bismuth or iron molybdates. Fig. 4.a and 4.b [84,85] show that these compound oxides benefit from the flb contact with a donor of Oso (0t-Sb204 3O is a typical donor, as it has no activity of its own). The figures we present here are simplified, just showing that an important synergy (increase compared to the straight line joining the C~ (b two extreme points) occurs when the powders of the two compounds are C3 2O mixed with each other (simply by suspending them in n-pentane, agitating and evaporating n-pentane; please note that the same weight has been used in all experiments of the series). ~n \ We showed that the same 10 synergetic effect occurs in a broad variety of reactions: 9 oxidative dehydrogenation of butene (C4=) to butadiene (BDE) (fig. 4.c and 4.d) [85,86] 9 oxidation of alcohols: methanol to formaldehyde (fig. 4.e) [87], ethanol to 10 acetaldehyde (as shown in fig. 4.f) 0.5 [88]; an almost identical figure is sno 2 obtained when a-Sb204 is mixed with (mass) MoO3 instead of Fe2(MoO4)3 [87]) and , s.o 2 ethanol to acetic acid using a mixture of Figure 3. Synergy (selectivity) between 0~-Sb204 and three phases: MoO3 + SnO2 + a-Sb204 SnO2 particles in the selective oxidation of isobutene [88]oxygen-aided (fig. 4.g). transformation of to methacrolein. The preparation of the sample formamides to nitriles: an example mixtures and the experimental conditions are among more than 15 cases is shown in described in the legend of fig. 2. More details are fig. 4.f [89]; in that case, the selectivity found in the original articles [63-65,83]. remains always high, the most dramatic effect concerning activity. A very interesting observation is that the action of spillover oxygen protects the active phase from deactivation [63,90,91]. On the basis mainly of results obtained in the oxidation of isobutene to methacrolein, the oxidative dehydrogenation of butene to butadiene and the oxygen-aided dehydration of formamide to nitriles, it was possible to show that oxides present in catalysts are located on a scale reflecting donor-acceptor properties (fig. 5). Some oxides are essentially acceptors (e.g., MOO3, some tellurates)" they can potentially carry active and selective sites, provided they receive spillover oxygen. Others are essentially donors; a-Sb204, in this respect, is typical: it produces spillover oxygen but carries no sites active for oxidation. Other oxides have mixed properties. The acceptors are relatively covalent, the donors are more ionic [63,77]. 40

_

9

11 Our work, and especially the comparison of results obtained with different types of reactions (see above) but using exactly the same catalyst mixtures, coupled with methods aimed at identifying active sites, also led to the demonstration that one of the consequences of the action of Oso was the creation or regeneration of acid hydroxyl groups (on MOO3) [63,92,93]. It was also shown directly that the deactivation and loss of selectivity of catalysts was associated with the fact that their surface got slightly reduced during the catalytic reaction. This does not occur when donors are present in the catalysts constituted of mixtures of donor and acceptor phases. The beneficial action of spillover oxygen is thus to keep the surface of the catalysts (acceptors) in a higher oxidation state [63,90,91,94,95]. All the phenomena observed can be explained by considering the full mechanism of the reaction, namely the simultaneous changes undergone by the reacting molecule and the acceptor part of the catalyst [91,94,95]. To make the argument as simple as possible, let us consider a very schematic structure of the surface of MoO3 (fig. 6). Octaedra composed of a central Mo ion and 6 oxygen ions surrounding it are the building blocks of the structure. They are normally linked together by comers, where an oxygen ion is shared by two neighboring octaedra: fig. 6 shows the real picture (a) together with the simplified representation we shall use in the following (b). The surface oxygens which react with the organic molecule might in principle be free "tips" (on top of our representation) or connecting O ions linking two surface octaedra. But theoretical and steric considerations [96] rule out the possibility that linking oxygens could come close enough to the hydrocarbon to react with it: only "tips" remain as likely candidates (fig. 7). The reaction of oxygen from the catalyst with the hydrocarbon thus brings about the formation of a reduced site which, in the MoO3 structure, corresponds to octaedra linked by one edge (namely by 2 oxygen ions, instead of one). We mentioned that acceptors not irrigated by Oso coming from donor tend to reduce. At the atomic scale, this means that oxygen is taken out of the surface by the hydrocarbon HC faster than molecular oxygen 02 from the gas phase can restore the corner-sharing structure (fig. 8). It ensues that the surface contains many more edge-sharing octaedra than corner-sharing ones. The role of Oso is to prevent this inbalance (fig. 9). The full argument is actually more elaborate and involves non-linear responses of the equilibrium as suggested in this figure [94,95]. The inbalance in the case where Oso is absent corresponds to a diminution of the number of active selective sites (the corner-sharing octaedra), and the appearance of non-selective sites (group of edge-sharing octaedra). The location of the acidic OH centers mentioned above is not yet clearly identified: they are likely to be present on the tip of a certain proportion of the corner-sharing octaedra at the surface of the catalyst. The transformation to edge-sharing pairs leads to their disappearance and the loss of activity. The accumulation of an excessive number of edge-sharing octaedra leads to bulk reduction and long-standing deactivation. This picture (or more precisely the complete elaborate picture resting on the ideas presented here in a schematic way) points to the necessity to have a well-defined architecture on the surface, which constitutes a demand for the elaborate concerted mechanism in selective oxidation. The conclusion is that spillover species permit that the correct coordination of atoms and groups of atoms at the surface of oxide catalysts be kept, thus permittting high activity and selectivity, and avoiding deactivation. The phenomenon by which a donor distinct from the real catalytic phase controls the catalytic properties of the latter is what we call a remote control.

12

(:3

(:3 I.,')

(%1 X#,~#~oloS

II

(%! ,O!~!laalaS

...c. (b 0 r.j

I

o,i

(%) X,z/A!laalaS

(%) ,OM!laalaS

0

(:b

N (~3,,,1.

(:b',,~

0

.-.~ 0

o

o

(%) plo!,~

o Q)

I

\

(%) Xl!A/,ZoalaS

/

(%) X,qA!~ooloS

\

-~e

c5

cb

9 r

c:b

d~

I

13 Figure 4. Examples of synergy between phases in various oxidation reactions. The mixtures were made by suspending the starting powders in n-pentane and evaporation under stirring; rm is the weight ratio in the mixture of the oxide mentioned at the fight of the figure. 4a.: oxidation of isobutene to methacrolein on SnO2-Bi2MoO6 mixtures (460 ~ [84]. 4b.: oxidation of isobutene to methacrolein on a-Sb204-FeSbO4 mixtures (400 ~ [85]. 4c. oxidative dehydrogenation of 1-butene to butadiene on tx-Sb204-ZnFe204 mixtures (400 ~ BiPO4 has an effect almost identical to that of a-Sb204 [86]. 4d.: oxidative dehydrogenation of 1-butene to butadiene on BiPO4-Fe2(MoO4)3 mixtures (400 ~ [851. 4e.: oxidative dehydrogenation of methanol to formaldehyde on a-Sb204-MoO3 mixtures

(350 ~

[871.

4f.: oxidative dehydrogenation of ethanol to formaldehyde on a-Sb204-Fe2(MoO4)3 mixtures (350 ~ [881. 4g.: oxidation of ethanol to acetic acid on mixtures of a-Sb204, MoO3 and SnO2 (240 ~ MoO3 and SnO2 were mixed (mass ratio MoO3/(MoO3+SnO2)---0.4) before the addition of ot-Sb204 [88]. 4h.: oxygen-aided dehydration of N-ethyl-formamide to propionitrile on a-Sb204-MoO3 mixtures (370 ~ The selectivity of the reaction is higher than 98%. The figure presents the variation of propionitrile yield [891.

/9

~q,,+ Figure 5. Donor-Acceptor scales for oxides used in selective oxidation (adapted from ref 63 or 77).

a

b

Figure 6. MoO3 octaedra and their normal linking by comers (or tips) (a). Picture b is the usual schematic representation of octaedra in the description of the structure.

14

12

corner sharing

"t:l

b

"'tip'" vacancy

i,,,,

d edge sharing k not likely in oxidation catalysis

"bridge'" vacancy

Figure 7. Representation of vacancies created by the reaction of an oxygen of the lattice with a hydrocarbon. As "tip" oxygens (corner oxygens above the surface) are the only ones accessible, at the exclusion of the bridging oxygens, the vacancies formed should be "tip" vacancies. The surface structure tends to spontaneously rearrange to create an edge sharing pair.

hydrocarbon

02 Figure 8. Inbalance in the rates of the antagonistic reduction of the surface by the hydrocarbon reactant and the reoxidation by molecular 02 in selective oxidation.

15

a . Spontaneous

b . S u r f a c e kept

ox,do-reduct,on

more

state

by spillover oxygen

of s u r f a c e

oxidised

Figure 9. Schematic representation of the surface at steady state a. when spillover oxygen is not present b. when spillover oxygen flows over the surface. 5 . ROLE OF HOMOGENEOUS REACTIONS

Contrary to the case of olefins, homogeneous catalytic oxidations of light alkanes occur at temperatures similar to those of the catalytic reaction. This certainly led to misinterpretation of supposedly catalytic data in certain cases. Two examp!es will illustrate the role of homogeneous reaction: the oxidative dehydration of propane and the reactions of pentane with oxygen. Burch and Crabb investigated in detail the role of homogeneous and heterogeneous reactions in the oxidative dehydrogenation of propane [97]. The reaction needs a temperature about 130 ~ lower for the catalysed reaction, but the difference depends somewhat on the oxygen/hydrocarbon ratio. The quite unexpected result of Burch and Crabb is that there are similar conversion vs. selectivity relationships for both the homogeneous and most of the heterogeneous reactions [97]. The authors add that even the best catalysts are only as good as no catalyst at all (but at higher temperature in this last case). This could seem pessimistic, but does not exclude that other catalysts could give a decisive advantage to catalysed reactions. A very interesting finding can perhaps modify the vision we have presently of the reaction. In the case of the homogeneous reaction, we found that a partial pressure of water in the feed promotes propane conversion. Fig. 10 shows the dramatic difference [98]. This makes the performance of the homogeneous reaction at a given temperature very close to those of the catalysed reaction at this temperature. An interestiag observation is that the production of byproduct ethylene is very little affected by conversion and almost not at all by the presence of water [98]. Fig. 11 gives propene selectivity as a function of propane conversion [98]. This seems to exceed the performances indicated by Burch and Crabb. It is not yet known whether similar effects could take place in catalysed reactions.

16 ~

100

Water added .9 L.

"Dry"

w

50

q.

a..

480

530

580 Inlet temperature ~

Figure 10. Influence of water on the homogeneous oxidative dehydrogenation of propane: propane conversion. Quartz reactor: internal diameter 9.3 mm; length of the void zone: 7 cm; Feed: propane, 4% vol; oxygen 9.3% vol; when water added: 15% vol; the balance was helium; flow: 50 cm3.min -1 [98]. 100

Water added

50

~aa aa 0

50

100 Conversion (%)

Figure 11. Influence of water and temperature on the homogeneous oxidative dehydrogenation of propane: selectivity to propene. Conditions as in fig. 10 [98].

17 A new work based on old patented data and which adds much to the interest of homogeneous oxidation shows that propylene oxide can be formed in certain conditions [99]. With respect to heterogeneous or hetero-homogeneous reactions, a very special system, constituted of lithium hydroxide/lithium iodide melts gives considerably higher propene yields at higher propane conversion than other homogeneous reactions or reaction catalysed by solid catalysts [ 100]. It is therefore very difficult to take without restriction the pessimistic view of Burch and Crabb. But conversely, the last remark in their abstract is certainly very relevant: "A combination of homogeneous and heterogeneous contributions to the oxidative dehydrogenation reaction may provide a means of obtaining higher yields in propene" [97]. Another interesting case is that of n-pentane oxidation. The reaction has been studied in the presence of vanadium phosphate catalysts around 330 ~ [100-103]. Maleic anhydride and phthalic anhydride are produced. It should be mentioned, however, that the homogeneous reaction begins to be significant above 300 ~ (fig. 12). The extent of conversion increases with the oxygen partial pressure [104]. By using reactors with empty spaces of different volumes (lengths), it is possible to evaluate the relative influence of the heterogeneous and homogeneous reaction (table IV) [ 104]. The non-selective homogeneous reaction increases the n-pentane conversion, but the surprising finding is that the maleic/phthalic anhydride selectivity varies substantially. This suggests two conclusions. The first is that the homogeneous reaction can play an important role in the oxidation of n-pentane in the range of temperature where catalysts like VPO are active (around 350-400 ~ The second is that the occurrence of the homogeneous reaction in parallel with the heterogeneously catalysed one might modify selectivity. ~

5O -20

~ 25

4

200

300

400

500 T ~

Figure 12. Non catalyzed reaction of n-pentane in an empty reactor (quartz; internal diameter 0.93 cm; length of the void zone: 7 cm; the rest of the reactor space is filled with SiC particles); gas feed: n-pentane 1% vol; 02:10 or 20% vol; balance: helium; total flow 30 cm3.min -1 [101].

18 Table IV. Influence of the homogeneous reaction on the oxidation of n-pentane. The reactor was a U-tube (inner diameter 9.3 mm) in which a section of the length indicated in the table was left void. After this section, the reactant flow passed through a frit and the catalyst (0.2 g, bed height 3 mm). The remainder of the tube was f'tlled with carborandum. The catalyst was vanadium phosphate with P/V=l.26, surface area 44 m2.g -1. The gas composition (volume) was: pentane 0.7%; oxygen 20%; helium 5%; balance nitrogen. Total flow 30 cm3.min -1. (Hourly Space Velocity 6000 h-l). T = 375 ~ CTOT is the conversion obtained with the above arrangement (void section + catalyst). The homogeneous conversion CHOM was determined with the same empty section but without the catalyst (replaced by carborandum). SMA and SPA are the selectivities to maleic and phthalic anhydride, respectively [104]. Void section cm CHOM % CTOT % SMA % SPA % 0 0 27 62 34 0.2 12 33 37 20 0.5 27 60 22 6.5 1.5 35 68 20.4 5.0 These results question the validity of many previous results on catalytic oxidation of light alkanes. One should reassess the data concerning the relative reactivity of the various alkanes [105] and selectivity. The general conclusion of this section is that the problem of the selective oxidation of alkanes must unavoidably involve consideration of homogeneous reactions in parallel with the catalysed processes. This is obviously necessary for understanding the phenomena and progressing in the selection of better catalysts. If new processes are the goal of investigations, the interaction between homogeneous and heterogeneous processes must be taken into account. The kinetics will be different. The relative importance of the two kinds of phenomena, homogeneous and heterogeneous, depends necessarily on the shape and size of the catalyst, the form of the reactor, and the overall design of the reactor. Progress in the oxidation of alkane thus needs a comprehensive approach, where catalysis chemists and chemical engineers should work in fight cooperation. 6. CONTROL OF CATALYST ACTIVITY IN ALKANE OXIDATION There are very good reasons to believe that the new phenomena discovered in the selective oxidation of olefins, in oxidative dehydrogenations and the other reactions mentioned in section 4 also occur in the reactions of alkanes with oxygen. This clearly breaks open the way to a better control of these reactions. We have indeed shown that the concept of a control of catalytic activity thanks to the addition of a spillover oxygen donor applies to reactions of alkanes. A conspicuous case is the oxidation of butane to maleic anhydride. We have discovered that a typical oxygen donor, namely a-Sb204, acts synergetically with the VPO phases which are responsible for the reaction [72]. BiPO4, although less good for enhancing selectivity, substantially increases activity. Thermoreduction and thermoreoxidation measurements show that, as in the cases of section 4, the surface oxido-reduction is affected by the presence of a donor [72]. We speculate that spillover oxygen coming from a-Sb204 or BiPO4 protects the special structures necessary for the concerted reaction of butane to maleic anhydride on vanadium pyrophosphate (fig. 1). In a cooperative work of our laboratory with Mamedov and Baidikova, it was also demonstrated, for the first time, that 2-phase catalysts are more efficient than single phase ones in the oxidative coupling of methane [106]. The oxide catalyst contained bismuth and manganese, which can form a well defined phase, Bi2Mn4010. This phase decomposed partially to give a-Bi203 (and a-Mn203) during the catalytic test. Using a catalyst containing

19 mainly Bi2Mn4010 , the C2 yield slowly increased to a plateau in the course of the fast hour of reaction and a-Bi203 was simultaneously formed. A mixture of a-Bi203 and Bi2Mn4010 reached the steady-state activity in a short time, and this activity was higher than in the previous experiment. Higher yields were observed when intimately mixed a-Bi203 and a Bi-depleted phase, Bi2.xMn4010-y were present. This result leads us to speculate: on a possible control of another factor not yet mentioned in this article. Several oxygen species can be present on the surface of oxides: O2", 022", O', 02". Their respective surface concentrations depend on the nature of the oxide, gas partial pressure and temperature. These various species have different reactivities [61-63,77]. It is believed that 02- (nucleophilic) is necessary in aUylic oxidation, and that the other species (electrophilic) are detrimental, by bringing about complete oxidation. On the other hand, some of these electrophilic species are very likely necessary for removing the first hydrogen of the saturated hydrocarbons (oxidation of butane to maleic anhydride and selective reactions of methane with oxygen). We tentatively explain the results concerning methane oxidative coupling by supposing that a-Bi203 and Bi2-xMn4010-y are complementary in providing the fight surface oxygen species. Manganese oxides have a high activity for complete oxidation. This implies that they produce strongly electrophilic species. The presence of bismuth, together with manganese, in Bi2. xMn4010-y should diminish the aggressiveness of the electrophilic species: Bi203 is a good oxygen donor, which produces mild' (i.e., nuc!eophilic) oxygen. The combination could provide the adequate balance of the various oxygen species necessary for the oxidative coupling reaction [107]. Recent results of our laboratory also show that the kind of concepts we are developing applies to other reactions of alkanes. We selected the oxidative dehydrogenation of propane to propene. Based on previous investigation with pure magnesium phosphate phases [33], we mixed a-Sb204 with the pyrovanadate (Mg2V2OT, written here MgV2/2 in short) and the orthovanadate (Mg3V208 or MgV3/2). According to cases, the yield or the selectivity are improved [108]. If we refer to the remote control concepts and the various effects that spiUover produces, we can interpret the results in the following way: 9 spillover oxygen produced by a-Sb204 essentially creates additional sites of approximately the same selectivity (probably the same geometry) on magnesium pyrovanadate MgV2/2. 9 this spillover oxygen modifies favorably the selectivity of surface sites on magnesium orthovanadate MgV3/2 (probably by slightly modifying the surface structure). If we reason in this way, we may conclude, by reference to the donor-acceptor scale shown in fig. 5, that MgV 3/2 behaves as a typical acceptor, because its selectivity is increased by spillover oxygen. Along this line, MgV 2/2 should have a lesser degree of acceptor character and more of a donor character. If this was correct, mixing MgV 2/2 with MgV 3/2 would lead to a syngergetic effect. This is what we observe: the selectivity gets enhanced [ 109]. A similar reasoning had led us to the prediction that two VPO catalysts with different P/V ratios could act synergeticaUy in butane oxidation to maleic anhydride, and this was also verified [71]. Concluding, it seems that the concepts concerning cooperation between phases and the role of spillover oxygen can be extended to the field of selective reaction of light alkanes with oxygen. But the control is more subtle, because more reactive oxygen species are necessary. The challenge, for producing useful molecules from saturated hydrocarbons and oxygen, is to avoid complete oxidation to CO2 and H20. It seems that electrophilic species are necessary for the first step, probably the removal of the first hydrogen from the saturated molecule. But there should not be too large a quantity of these species on the surface, and their reactivity should not be excessive (O2-, O22-, O have certainly different electrophilicity and different reactivities). These electrophilic species are probably detrimental for the subsequent steps of the reaction. Then, nucleophilic species are necessary. They may be necessary just for diminishing the concentration of the harmful electrophilic species through mutual competition for sites on the surface. They are very likely necessary, as in the cases mentioned in section 4, for maintaining

20 the adequate oxidation state of the surface and, consequently, avoid the destruction of the arrangement of surface atoms demanded by the concerted mechanism necessary for selective reaction. They may also be necessary as reactant for certain steps. Although the demands concerning the active oxygen species seem conflicting, the experimental conditions can be selected to achieve a compromise. The oxidation of butane to maleic anhydride, widely industrialized now, shows that this compromise can be achieved and lead to economically attractive processes. Fortunately, experimental conditions do not constitute the only control parameter. SpiUover of oxygen can play a crucial role. This is what is observed in the examples mentioned above. Spillover takes place from an adequate donor to the active phase (or acceptor) namely VPO or MgVO in the case of butane or propane reactions, respectively, or possibly Bi2.xMn4010-y for methane coupling. In this context, the present situation suggests that research should be directed in priority along two lines. The first one would be to detect solids which, under given conditions, can develop the active and selective surface structures (the equivalent, for other reactions, of the inverted flat square pyramids necessary for butane oxidation to maleic anhydride). The second one would be to understand what kind of solids may generate the adequate spillover species in good proportion at adequate temperatures. The scales presented at the end of section 4 seem to concern essentially donors of nucleophilic spillover oxygen 02-. It can be expected that more ionic solids would produce more electrophilic species [63]. The higher the temperature necessary for the reaction, the more ionic will be the donors necessary for achieving the good balance of oxygen species. 7. PROSPECTS: COMPREHENSIVE APPROACH TO F U N C T I O N A L I Z A T I O N OF L I G H T ALKANES BY R E A C T I O N OXYGEN

THE WITH

Letting light alkanes react selectively to give valuable products is one of the main goals of petroleum chemistry nowadays. If the selective oxidation of methane is considered, this even appears by far as the most important issue in the very next years. This is clear when remembering that methane represents about one-third of the hydrocarbon resources of the world during this decennia. It is therefore not surprising that all chemists and particularly catalysis chemists have devoted much effort to functionalize methane and the light alkanes. Progress since the industrialization of the butane to maleic anhydride until 1994 has been extremely modest. It is therefore worthwhile to assess critically the approach taken by the various investigators. In the present article, we suggested some critical considerations. But one aspect was almost left aside until now, namely the role of chemical engineering. We shall now attempt to suggest how the various pieces of science are probably assembling together and are progressively unveiling a new, more comprehensive and more realistic approach. The functionalization of light alkanes and particularly their reactions with oxygen necessarily involve, roughly speaking, both the chemical and the chemical engineering aspects (in addition, of course, to economic considerations and the now associated environmental aspects). The chemical approach itself is composed of two distinct but narrowly interconnected lines: the purely catalytic and the homogeneous aspects. The latter is obviously of considerable importance as commented above and proven by the case of methane oxidative coupling. But it is striking that, even in methane coupling, an overwhelming fraction of research has been directed to the discovery of new catalysts (perhaps over 90%) with only a very small fraction trying to take homogeneous phenomena into account. The progress has been deceivingly modest. This has allowed respected scientists, even industrial scientists, to discourage further research on the topic. They were right in mentioning that the results obtained were very far from being economically attractive. But, instead of discouraging research, they should have spurred research, while specifying "on different lines". Among these lines, new developments in chemical engineering were obviously to be considered. The growing importance of chemical

21 engineering is clear in all the field of catalysis, as shown by the overview of the new catalytic processes developed in the world during the 80's [ 110]. It should have been perceived as still more proeminent in selective reactions of alkanes with oxygen, just if one had considered the possibility that homogeneous reactions could occur in conditions identical, or very close to, those of catalysis. It is therefore easy to predict that the research and pre-development work aimed at alkane functionalization using oxygen should incorporate in comparable amount various ingredients. Recent developments announce these changes. These ingredients are: 9 continuation of the approach traditionally taken in catalysis, namely search for new phases able to permit the initial attack of alkanes by some form of oxygen; 9 the new approach described in section 5, considering the role of surface-mobile oxygen in catalysis and the special reactivity of such species when produced by separate phases (donors); 9 the understanding of homogeneous reactions: initiation in the gas phase or on the catalyst surface, propagation in the gas phase, inhibition of propagation thanks to radical trapping on adequate surfaces, etc.; 9 the design and building of new types of reactors and equipment, in order to compensate for poor conversion (if high selectivities are desired), to cope with homogeneous reactions and, probably, to permit extremely fast reactions. In this last section of our contribution, little has to be added concerning the first and second points, which have already been discussed in detail. Following Mamedov [ 111], we wish to underline the role of the reactive atmosphere in selective oxidation. A new result obtained with gold deposited in a proper way on TiO2 supports this assertion [ 112]. The authors show that C3 and C4 hydrocarbons can be selectively oxidised at very low temperatures (50-80~ using simultaneously molecular oxygen and hydrogen: examples are propane to acetone and isobutane to tert.butanol with selectivities of, respectively, 14.6 and 46% (at, understandably, low conversions). In this context, we remark that the role of water (steam) in selective reactions with oxygen has not been given proper attention in general. The role of CO2 should perhaps be also studied. A systematic search for oxides able to donate the appropriate spillover oxygen species at high temperatures is highly desirable. Very recent results certainly reinforce the conclusion that the occurrence of homogeneous oxidation reactions of light alkanes must be considered with attention. The example of ethane discussed above shows that the homogeneous reaction can be as selective, or almost as selective, as the catalysed one [97]. In the reported experiments, the homogeneous reaction was controlled by none of the techniques well known in the field of combustion and radical gas phase reactions (artificial genesis of radicals, trapping of radicals, presence of foreign inert molecules, etc.). Oppositely, the catalysts used for comparison were the result of a selection and some optimisation. This could suggest that simple homogeneous reactions might be the basis of economically viable processes in the future. Other very recent results reinforce the validity of this prospect, like the recent observation that the gas phase reaction of methane with oxygen can give methanol in selectivities exceeding 30% at methane conversion of 5 % [113]. If we now consider the role of chemical engineering, the impressive results of Huff and Schmidt cited above demonstrate that employing a type of reactor not used previously in oxidation and very short residence time can lead to promising prospects [31]. But chemical engineering is not only the science of reactors. It has to consider the whole plant. Two very recent results dramatically demonstrate that integrating recycle and separation features with a catalytic reactor lead to very impressive yield. Tonkovich et al. reached a 50% yield in C2 hydrocarbons in the oxidative coupling of methane using a moving bed reactor, thus permitting a sort of chromatographic separation [114]. The problem indeed is the high reactivity of ethylene compared to CH4. But the reaction of methane to ethylene can be extremely selective at very low methane conversions. Considering these particularities, the group of the University of Patras led by C.G. Vayenas achieved an ethylene yield of 85% (calculated on the carbon contained in CH4) [115]. The key to success is highly selective adsorption of ethylene, ethane and CO2 on a 5A molecular sieve from which they are periodically released. Conversion

22 is kept very low, and the non-reacted methane recycled. It is remarkable that the catalyst is not one of those giving the best performances in conventional reactors. In addition to silver on which oxygen is "pumped" electrolytically, the authors used a Sn203-CaO-Ag catalyst using conventionally molecular oxygen from the gas phase (both catalysts gave approximately the same results). Not only does the reactor concept offer new perspectives, but the composition of the catalyst used suggests other developments. This spectacular result obtained in the most extensively investigated reaction, namely the oxidative coupling of methane, is a clear indication that we are still at the beginning of really innovative research in the functionalization of light alkanes. This is also a clear indication that the discoveries triggering these innovations will involve all the relevant scientific fields, catalysis as well as homogeneous reactions, discovery of new catalysts as well as new che~,ical engineering concepts. Finally, we would like to underline that this article is not intended to be a comprehensive review paper. Many extremely valuable contributions have not been cited. We would particularly like to give credit to the recent papers to be found in references 116 to 123. REFERENCES

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B. Delmon, Surface Reviews and Letters, 2 (1995), in press. K. Hermann, private communication. R. Burch andE.M. Crabb, Appl. Catal., 100 (1993) 111. Z. Sobalik, P. Ruiz and B. Delmon, unpublished results. T. Hayashi, L.-B. Han, S. Tsubota and M. Haruta, submitted. I.M. Dahl, K. Grande, K.-J. Jens, E. Rytter and ~. Slagtern, Appl. Catal., 77 (1991) 163. Z. Sobalik, P. Ruiz and B. Delmon, to be published. J.T. Gleaves and G. Centi, Catal. Today, 16 (1993) 69. F. Trifir6, Catal. Today, 16 (1993) 91. Z. Sobalik, P. Ruiz and B. Delmon, to be published. A. Aguero, R.P.A. Sneeden and J.C. Volta, in "Heterogeneous Catalysis and Fine Chemicals" (M. Guisnet, J. Barrault, C. Bouchoule, D. Duprez, C. Montassier, G. PErot, eds.); Elsevier, Amsterdam, (1988) 353. I. Baidikova, M. Matralis, J. Naud, Ch. Papadopoulou, E.A. Mamedov and B. Delmon, Appl. Catal. A, 89 (1992) 169. B. Delmon, Symposium on Production and Processing of Natural Gas (A. Fakeema, A. Omar, eds), Riyadh, 29 Febr. - 2 March 1992, Preprints, 32-1. S.R.G. Carraz~in, C. Peres, J.-P. Bernard, P. Ruiz and B. Delmon, submitted. Xingtao Gao, P. Ruiz, Qin Xin, Xiexian Guo and B. Delmon, J. Catal., 148 (1994) 56. A. Cahuvel, B. Delmon and W.G. Htilderich, Appl. Catal., 115 (1994) 173. E.A. Mamedov, Appl. Catal., 116 (1994) 49. T. Hayashi and M. Haruta, private communication, to be submitted. L.B. Han, S. Tsubota, T. Kobayashi and M. Haruta, J. Chem. Soc., Chem. Comm., accepted. A.L. Tonkovich, R.W. Carr and R. Aris, Science, 262 (1993) 221. Y. Jiang, I.V. Yentekakis and C.G. Vayenas, Science, 264 (1994) 1563. G. Centi, F. Trifiro, J.R. Ebner and V. Franchetti, Chem. Rev., 28 (1989) 400. B.K. Hodnett, Catal. Rev. Sci. Eng., 27 (1987) 373. E. Bordes and P. Courtine, J. Catal., 57 (1977) 236. G. Centi (guest ed.), Catal. Today, 16 (nr. 1) (1993), pp. 1-153. G. Bergeret, M. David, J.P. Broyer and J.C. Volta, Catal. Today, 1 (1987) 37. F. Ben Abdelouahab, R. Olivier, N. Gilhaume, F. Lefebvre and J.C. Volta, J. Catal., 134 (1992) 151. M.T. Sananes, A. Tuel and J.C. Volta, J. Catal., 145 (1994) 251. G. Centi and F. Trifiro, Appl. Catal., 12 (1984) 1.

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Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All fights reserved.

27

V A N A D I U M R E S I S T A N T F L U I D C R A C K I N G CATALYSTS* Mario L. Occeili

Zeofites and Clays Research Program, Georgia Tech Research Institute, Georgia Institute of Technology, Atlanta, Georgia 30332, USA.

ABSTRACT 29Si and 27AI MAS NMR spectroscopy has shown that on calcination cracking components such as FlY, Si-enriched FlY, REHY, and CREY undergo dealumination and that steam-aging increases the presence of extra-framework A1 in these zeolites. Dealumination is more severe in HY-type materials. The zeolite V resistance seems to decrease when RE ions are present and to increase with increasing extra-framework A1 (generated during steam-aging). At the high temperatures used for FCC regeneration, oxycations of vanadium (VO2§ or VO 2+ could attack the A1-O-Si bond in HY and cause lattice collapse. In REHY and CREY crystals, it is believed that Ce4§ ions, present as an oxycerium complex, undergo a redox reaction with oxyvanadyl cations (VO>), and form a stable orthovanadate. Removal of other charge-compensating cations (such as Na + ions) in the form of vanadates further destabilizes the crystal lattice, thus promoting zeolite destruction. Atomic Force Microscopy (AFM) can provide details of the surface topography of an FCC with unprecedented resolution, and can be used to rationalize the deleterious effects that metals such as Ni and V have on the properties of FCC. The deleterious effects of V deposits on zeolite-containing FCC can be greatly reduced by addition of certain materials (metal scavengers) capable of minimizing direct V-FCC interactions by selectively (and irreversibly) sorbing migrating V compounds such as H4V2OT. Dual-function cracking catalyst (DFCC) mixtures have been prepared that can retain most of their useful cracking activity (at MAT condition), even in the presence of 1.0% to 1.5% V. Thus, DFCC systems appear to have the necessary metal tolerance to crack residual oil as well as heavily V-contaminated crudes and may offer cost (as well as coke selectivity) advantages over conventional FCC. 1. I N T R O D U C T I O N During the cracking process, metal-containing heterocyclic compounds decompose leaving metal contaminants on the fluidized cracking catalyst (FCC) surface. Of the 28 elements identified in several domestic crudes, Ni and V are generally the most abundant [ 1]. The deleterious effects of these two elements on FCC activity and cracked product selectivities have long been recognized and are the subject of many patents and publications [2-21]. The study of metal effects and the preparation of metal-resistant FCC have been reviewed in two articles published in the last three years[23,24]. * based in part on a paper presented at an AKZO Catalyst Symposium in Scheveningen, The Netherlands.

28 Metals such as Ni (and to a lesser extent Fe) have little effect on catalyst activity, but they can catalyze the secondary cracking of gasoline with generation of high quantities of hydrogen and coke. Vanadium effects on catalyst properties are more severe because this metal can irreversibly destroy the catalyst cracking centers, thus eliminating the FCC's useful cracking activity. Vanadium in crudes is present mainly in the W 4 state as part of organometallic compounds such as porphyrins and naphthenates. During the cracking reaction in a FCCU, these compounds deposit V (probably in the form of VO 2+ cations) on the catalyst surface. Then, after steam-stripping and catalyst regeneration, formation of V +5 surface species occurs. In the regenerator, the oxidative decomposition of carbonaceous deposits on the FCC surface forms V205. and vanadia-like compounds. This oxide has a low melting point (658~ and is therefore capable, during regeneration, of diffusing within the FCC microstructure where it can cause pore blockage in addition to irreversibly destroying the zeolite crystallinity. Lowervalance vanadium oxides melt at temperatures (> 900~ much higher than those encountered in an FCC regenerator [24]. Thus, V oxidation to V +~ should be minimized to improve the FCC vanadium resistance [25]. The nature of the species formed when V-contaminated FCC are exposed to steam remains somewhat controversial. When immersed in water (at room temperature), vanadium (supported on solids) undergoes complex hydrolysis-condensation-polymerization reactions that form H2VzO7 "2, HV207 "3 and H2V10028"4 ions [22,26]. V concentration, surface composition, and liquid pH control the nature of the polyanions formed and their degree of protonation. Different reactions and reaction products are expected to occur when the same Vcontaminated materials are exposed to steam. However, it is believed that the same parameters (such as surface compositions, V-levels, and residence times) that influence the nature of the polyanions formed when V-contaminated solids are exposed to water will also affect the nature of the volatile V-compound formed when the same catalyst is exposed to steam. Yannopoulos [27] has proposed that vanadia reacts with steam to form vanadic acid: V/Os(s) + H20(v) = 2H3VOa(v). Vanadic acid was believed capable of leaching AI from the zeolite framework in the form of A1VO4, thus causing reduction in crystallinity and therefore cracking activity [15]. However, experimental evidence for A1VO4 formation could not be obtained by x-ray diffraction (XRD), laser Raman spectroscopy (LRS) [28-30], secondary ion mass spectroscopy (SIMS) [31 ], or by 51V-NMR [32,33]. This vanadate is not thermally stable at the temperatures existing in a typical cracking unit [30]. Thus, its role during zeolite deactivation must not be important. LRS characterization of DFCC systems tested at microactivity test (MAT) conditions, has indicated that in steam-aged catalysts containing more than 1% V, [V207] units are formed [29,30]. In the presence of a layered magnesium silicate (such as sepiolite), formation of 13Mg2V207 has been verified by LRS [28] as well as by 51V-NMR [34]. These results support the experimental work of Glemser and Muller [35]. in which the reaction: V2Os~s,l)+ 2H20~v) = HnV2OT~v) was reported. Thus, it is believed that HaV207 is one of the volatile V-compounds that can be generated in the steam-stripping zone of an FCC unit. It is the purpose of this paper to review vanadium-zeolite interactions and define all the major effects to consider when studying and preparing V-tolerant FCC.

29 2. EXPERIMENTAL

2.1 Catalyst Preparation The sample of calcined rare earth (RE) exchanged Zeolite Y (CREY) and the highactivity cracking catalyst (GRZ-1) used in the present study were obtained from the Davison Company. Davison's GRZ-1 is an FCC that contains an estimated 35% CREY which, after steam-aging, has a BET surface area of 161 m2/g. The CREY sample had a SIO2/A1203 ratio of 5.0, contained 7.6% Ce203, 4.0% La20, 2.8% Nd203, or 0.9% Pr203, and had a BET surface area of 749 mE/g. A residual 3.5% Na20 was found in these crystals. The HY sample (Linde LZY-82) had a bulk 5iO2/A1203 ratio of 5.4 and BET surface area of 761 mE/g. After calcination at 540~ in air, the two silicon-enriched HY used (Linde LZ210 type crystals) had BET surface area of 625 m2/g and 629 mE/g and bulk SIO2/A1203 ratio of 10.1 and 13.2, respectively. Solutions of vanadyl naphthenate in benzene were used to metal load the various materials according to an established procedure [19]; the naphthenate was obtained from Pfaltz and Bauer, Inc. and contained 1.9 wt% V. The vanadium loaded materials were first air-dried for 10 hours, slowly brought to 540~ (in flowing air) and then kept at this temperature for 10 hours. Steam-aging was accomplished by passing steam at 760~ (at latm) over the calcined catalysts for five hours. 2.2. Catalyst Characterization Vanadium in several aqueous extracts was determined by flame atomic emission spectrometry [22]. Powder diffraction measurements were obtained with a Siemens D-500 diffractometer at a scan rate of 0.01~ using 3 sec/step monochromatic Cu-ka radiation; CaF2 was used as an internal standard for angle calibrating. Raman spectra were recorded on a Spex Ramalog 1403 spectrometer (Spec Industries, Metuchen, NJ) equipped with a cooled RCA GaAs photomultiplier tube (CA 31034-02). The 4880A line of a model 165At laser (Spectra Physics, Mountain View, CA) was used to generate Raman scattered light [22]. Silicon-29 NMR spectra were recorded at 53.7 MHz on an IBM AF-270 FINMR spectrometer equipped with Doty Scientific MAS solids probe. Samples were spun in cylindrical 7mm alumina (sapphire) rotors equipped with vespel caps at 3.0 to 3.5 kHz. Experimental Silicon-29 NMR signals were deconvoluted into Gaussian components using the LINESIM program (courtesy of Dr. Peter Barton, Gritiity University, Natham, Australia) that was written for an ASPECT-3000 computer. The best-fitting simulated spectra were obtained using an iterative simplex routine. The Si/A1 ratios were calculated from the derived line 4 intensities using the relation:

Si/Al"l,,,,/ ~ 0.25n/s,t.an n--1

where ITOTis the total intensity of the spectrum and Isi~,AJ)isthe intensity contributions from Si atoms with nA1 neighbors in their second coordination sphere. The same spectrometer was also used to record Aluminum-27 MAS NMR spectra at 70.4 MHz. Typical scan conditions involved 18~ lasec) pulse with a recycle delay of 2 secs to obtain near quantitative results [22]. In calculating AI(VI)/AI(IV) ratios, it was assumed that spinning side bands (SSB) were of equal intensifies. Justification for this assumption rests on the observation of equal intensity SSB in the V-loaded (calcined in air) HY crystals.

30 To obtain images with the atomic force microscope (AFM), the FCC microspheres were sprinkled over a steel disk covered with a film of epoxy resin. After the glue dried, the AFM tip was placed onto the microspheres. The AFM used for these experiments [36] was a contact mode microscope based on the optical lever cantilever detection design of Amer and Mayer [37] and Alexander, et al. [38]. The AFM works like a record player. An xyz piezoelectric translator raster scans a sample below a stylus attached to a cantilever. The motion of the cantilever, as the stylus moves over the topography of the surface, is measured by reflecting a laser beam off the end of the cantilever and measuring the deflection of the reflected laser light with a two-segment photodiode. A digital electronic feedback loop keeps the deflection of the cantilever, and hence the force of the stylus on the surface, constant. This is accomplished by moving the sample up and down in the z direction of the xyz translator as the sample is scanned in the x and y directions. The images presented in this paper contain 256 x 256 data points and nearly all images were acquired within a few seconds. The Si3N4 cantilevers (with integral tips) used for imaging were 120~m in length and possessed a spring constant of approximately 0.6 N/m. The force applied for these images ranged from 10 to 100 nN. Approximately 900 images were acquired by examining a variety of microsphere surfaces. 2.3. Catalyst Testing Catalyst evaluation was performed with a microactivity test (MAT) using conditions described elsewhere [22]. Conversions are on a vol% fresh feed (FF) basis and have been defined as [Vt- Vp/Vt] X 100, where Vt is the volume of feed and Vp is the volume of product with b.p > 204~ [22]. 3. VANADIUM INTERACTIONS WITH HY-TYPE ZEOLITES In recent years, HY-type crystals have replaced in importance calcined rare-earth exchanged Y zeolites (CREY, with 10% to 20% RE203) in the preparation of FCC. Although oil prices during the 1987-1992 period have oscillated between $10 and $25 per barrel, nickel concentrations on equilibrium FCC from North American refineries have decreased from near 900 ppm to 700 ppm. Similarly, vanadium concentrations during the same period have decreased from 1300 ppm to about 1100 ppm, (Figure 1). However, a worldwide survey of metals concentration on equilibrium FCC that begun in 1992, has indicated that this trend is now reversed and that the mean Ni and V concentrations for the second quarter of 1994 have reached the 1029 ppm and 1608 ppm levels, respectively. Therefore, an understanding of VHY zeolite interactions is essential to the design of novel metal resistant FCC for the 1990s. When calcined at high temperatures (540 ~ to 760~ in flowing air, HY type crystals (Linde's LZY-82) are stable even in the presence of 4% V. With 5% V, the faujasite structure collapses only when the calcination temperature is raised from 540~ to 760~ forming mullite and some silica [22]. Recently, Marchal and coworkers [39] have reported that V205, can interact with NaY crystals even at low (410 ~ to 480~ temperatures and that when the V/(Si + A1) atomic ratio reaches 0.2, a collapse of the faujasite structure occurs with formation of a sodium-vanadate-like phase. Thus, even in the absence of steam, V-loaded Y zeolites collapse when calcined in air with an ease dependent largely on calcination temperatures, Na and V levels.

31 1800 1700 1600 1500 . ~ 1400

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92

94

96

98

2000

YEAR

Figure 1. Metal level trends on equilibrium fluidized cracking catalysts. Before 1992 data is based on Davison analysis of samples from cracking units in the USA and Canada. After 1992 the survey is world-wide (full symbols). Electron paramagnetic resonance (EPR) studies have indicated that vanadium (when introduced in the form of vanadyl naphthenate) is stabilized on the zeolite primarily as octahedrally coordinated VO 2§ cations even atter calcination [40]. In contrast, after calcination an amorphous aluminosilicate gel stabilizes vanadium mainly in the form of V205 [40]. Thus, it is believed that during calcination at 760~ in air, VO 2§ cations can attack Si-O-AI bonds causing de-alumination and lattice collapse. When present as V205, vanadium during calcination reacts with charge-compensating Na § cations to form stable vanadate-like phases that destabilize the faujasite structure [39]. In the presence of steam, the ease with which Na and V destabilize the faujasite lattice increases drastically. The deleterious effect of Na ions on the hydrothermal stability of zeolites have been well-documented in the literature [41,42]. In the absence of V impurities, hydrothermal stability depends on steam-aging temperature and, most importantly, Na levels (Figure 2). Thus, when studying V effects on these crystals, the presence of residual Na ions must be carefully considered. It has been found that when the Na20 level is reduced to 0.14% Na20, HY can retain most of its crystallinity when steam-aged (100% steam, 1 atm, 5 hr) in the 760 ~ to 815~ temperature range (Figure 2). EPR [37] as well as XPS results [29,30,43] have shown that atier steam-aging, V is present mainly as a V+Lspecie. It is believed that residual VO § together with VO2+~ cations and acids (such as I-hV207) resulting from hydrolysis reactions between steam and surface V-

32 impurities, are responsible for the ease with which HY crystals lose their crystallinity in the presence of about 2% V [22]. Sodium collapses the faujasite structure, leaving an x-ray amorphous residue (Figure 3). In contrast, the destruction ofHY crystallinity by V results in mullite and silica (tridymite) formation (Figure 4). A third-phase, vanadia, can be easily identified in the laser Raman spectra (LRS) of V-loaded HY crystals (Figure 5). The calculated orthorhombic unit cell parameters for several mullites, crystallized with and without V, have indicated that (in a qualitative sense) V causes an expansion of the unit cell volume resulting mainly from an increase in the a dimension [22]. Thus, incorporation of V into the crystal lattice of this mineral can occur during mullite formation. Crystallinity, together with surface area retention data, suggests that silicon-rich MFI crystals are generally more V-tolerant than HY crystals at hydrothermal conditions (see Table 1). The percent Na20 level in the two pentasils is less than 0.01%; in the two HY crystals it is less than 0.5%. The destruction of the pentasil structure by V generates crystobalite, indicated by the arrows in Figure 6. As observed for HY, vanadia formation can be seen only by LRS (Figure 7). Results in Table 1 suggest that by increasing framework AI, V tolerance decreases, indicating that V preferentially attack Si-O-AI bonds in these zeolites. Lattice degradation from thermal or hydrothermal treatments also can be followed by NMR. Silicon-29 NMR spectra of calcined and steam-aged HY-type crystals are characterized by a single resonance between -108 ppm and -110 ppm attributed mainly to the presence of Si[OAI] units generated by dealumination [22]. In Figure 8 there is an additional weak and broad shoulder near -115 ppm. For HY crystals (with % Na20 between 0.9 and 1.0%), the relative signal intensity of the upfield resonance near -115 ppm increases with V-levels (Table 2) suggesting formation of extra frame-work silicon resulting from lattice degradation. Table 1. Surface Area Retention for Several Steam-Aged Zeolites in the Presence of 0-5 wt% Vanadium. % Surface Area Retention 1.0 2.0 3.0 4.0

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

Vanadium, wt%: Zeolite SiO2/A1203 Silicalite 422 ZSM-5 98 HY 6.5 HY 4.7

0.0

100 100 100 100

81 76 67 45

69 52 37 10

38 32 6 5

Table 2. Percent Signal Assignable to Silica Formation Resulting from Lattice Degradation. Vanadium, wt%"

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Figure 4. X-ray diffractograms of HY crystals with 5% V calcined at: A) 540~ in air, B) 760~ air and C) 760~ steam. Silica (tridymite) formation is indicated by arrows (22).

Figure 5. Laser Raman spectra of steam aged crystals containing 5% V: (A) HY (Linde's LZY-82), (B) Silicon-Enriched HY (Linde's LZ-210) with Si/A1 = 6.1 and Na20 = 0 . 1 8 wt%.22

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-140

(ppm)

Figure 8. Silicon-29NMR spectra at 74 MHz for a set of HY crystals (Linde's LZ-210, with Si/A1 = 5.5) steam-aged in the presence of: A) 0.0%V, B) 0.5% V, C) 1.0% V, D) 1.5%V and E) 2.0% V.

CHEMICAL

SHIFT (ppm)

Figure 9. Aluminum-27 NMR spectra at 70.4 MHz from HY (Linde' s LZY-82), A) after calcination at 540~ in air and B) after steam-aging at 760~ 100% steam at 1 atm.

35 Table 3. Formation of Extra Framework AI in V-loaded HY type Zeolites alter Calcination at 540~ Oh Dry air . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .

Vanadium, wt%" wt% Na20

0.0

0.5

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

1.0

1.5

2.0

0.17 0.22

u 0.28 0.18 0.38

0.53 * 0.22 0.42

Si/A!

0.14 4.7 2.46 3.8 0.90 5.3 1.01 11.0 * Could not be computed with

0.20 0.18 0.14 0.14 accuracy ,

0.17 0.23 0.10

,,,

Table 4. Effects of V on HY (Si/A1 = 11.0 and Na20 = 0.19 wt%) Crystal Properties After Calcination in Air at 540~ Vanadium, wt%:

0.0

0.5

1.0

1.5

2.0

% Cryst. Ret. BET SA (m2/g) ao(A) + 0.003

100 641 24.377

100 631 24.368

94 629 24.368

92 619 24.366

90 618 24.363

Aluminum-27 NMR spectra contain an intense asymmetric peak in the 55 to 57 ppm region associated with framework A1 atoms and one near 0 ppm associated with extra lattice AI(VI) (Figure 9). Spectra (at 70.4 MHz) of HY, as well as of Si-enriched HY crystals, indicate that calcination (at 540~ in air) induce dealumination in all the samples examined [22]. For the set ofHY zeolites with 0.1AI-OH + VO +2 = (>A1-O)2-VO + 2H +) complexes capable of sequestering V, thus enhancing the crystals ~ V resistance. In fact, surface area (and crystallinity) data indicate that the enhanced stability is maintained also when the crystals are steam-aged in the presence of vanadium [41] (Table 6).

36 Table 5. The effects of steam-aging on the crystal properties of V-flee HY crystals. (C = Calcined, S = Steamed) BET Surface Na20 wt%

Si/AI

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

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

C

3.8 2.5 5.3 0.9 11.0 1.0 9 Signal too broad to be

S

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

C

24.581 24.26 607 24.474 24.301 625 24.272 24.272 629 integrated with accuracy

..... i .....

S

C

S

77 494 574

0.18 0.21 0.14

* * *

Table 6. The effect of steam-aging on the surface area of V-loaded, silicon-enriched, HY crystals. Surface Area (m2/g) ....................................................................................................................................................

Vanadium, wt%"

0.0

0.5

1.0

1.5

2.0

Si/A! 5.3 11.0

494 575

336 556

260 486

136 418

42 286

In agreement with Silicon-29 NMR results, V addition induces further dealumination of the HY lattice (Figure 9). A third peak near 30 ppm appears in the Al-27 NMR spectra of these HY crystals (Figure 9). A line near 30 ppm has been attributed to the presence of highly distorted extra framework AI(IV) [45] or to pentacoordinated AI nuclei [46]. Therefore, V attack on the faujasite structure can cause both silica and extra framework AI formation (Figure 10). Silicon-enriched HY crystals (with Na20 between 0.9-1.0% and Si/Al = 11.0) are more resistant to V attack at hydrothermal conditions than HY (Linde LZY-82 with Si/A1 = 4.7) containing 0.14% Na/O [22]. Thus, ideally, a zeolite to be used in the preparation of metal-resistant FCC should contain low residual Na (<0.5% Na/O) and high framework Si/AI ratio. The presence of extra framework A1 is also believed to benefit the zeolite V tolerance. 4. VANADIUM INTERACTIONS WITH CREY-TYPE ZEOLITES Vanadium-loaded CREY (and HY) crystals have high thermal stability (at 760~ air) and 5% V loading is required to collapse the crystal structure (Table 7). When this thermal treatment is repeated in the presence of steam (760~ 100% steam at 1 atm), the collapse of the faujasite structure readily occurs with only 0.4% V [29]. In steam-aged CREY samples containing more than 1% V, CeVO4 formation can be observed by either LRS or XRD (see the arrows in Figure 10); Vanadate levels increase linearly with V concentrations [29]. Although the CREY crystals used contained 4% La203, La compounds could not be observed [29]. However, LaY crystals (with 8.1% La203),when steam-aged in the presence of V, collapse with formation of a phase identified by LRS to be LaVO4 [29].

37

c~i

o

~1

.,.

,~

:~

~

T-

~...~"

-'~" "

.

r

"

"+-'"

''

"~"

D

(:L.

C

~"

",,

,

-

,

-

_

,

-

,

+

,

.

.

,

-

TWO -- THETA (DEGREES)

I

IJJ

+,.,...

B

z_

. r

J"-200

150

100

50

0

A -50

-100

-lS0

-200

o

O.

,~.

8.

1'2.

16.

2b.

2~4.

2'8.

32.

3'8.

40.

C H E M I C A L SHIFT (ppm)

Figure 10. Aluminum-27 NMR spectra at 70.4 MHz for HY (Linde's LZ-210, with Si/A1 = 5.5) steam-aged in the presence of: A) 0.0% V, B)0.5% V, C) 1.0% V, D) 1.5% V and E) 2.0% V.

Figure 11. X-ray diffractograms of CREY crystals (containing- 15% Re203 and 3.5% Na20) after steam-aging in the presence of: A) 0.0% V, B)0.4% V, C)0.8% V and D) 1.2% V. k %,

<{

I

W rr"

<~

,

~

k~

I" "Y(s. . . . . . . . ~+. N,,O:O,+! IO"'(s ......~ ,,,~o.... ' /

~~

In

CREY( . . . . . 0 . . . . . )

80

0

3S)

ILl

(J <~

~:

60

z 0 z W p. W rr

(5 0.

4.

8.

12.

16.

20.

2'4.

28.

32.

36.

40.

44.

20

0.5

1.0

1.S

2.0

V A N A D I U M L O A D I N G (WT%) TWO -- THETA (DEGREES)

Figure 12. X-ray diffractogram of CREY crystals (containing- 15% Re203 and 0.11% Na20) after steam-aging in the presence of: A) 0.0% V, B) 0.4% V, C) 0.8% V and D) 1.2% V.

Figure 13. Vanadium tolerance of steam-aged CREY containing: A) 3.5% Na20, B) 0.11% Na20 and HY with C) 0.14% Na20 and Si/A1 = 4.7 and D) HY with 0.11% Na20 and Si/A1 = 11.0.

l

]

44.

38 In RE-containing Y crystals (and in CREY containing FCC) it is believed that Ce, present as an oxycerium complex, is preferentially located near the supercages where it can more readily react with the stable (V=O) § cations and form CeVO4 by the following redox reaction [29,30]: [Ce/O \ 0 ) " Ce]+4 + 2 VO+2 + 4H20 --->2 Ce VO4 + 8H+ The CREY resistance to V deactivation can be improved by decreasing the crystals' Na levels. By ion exchanging commercially available CREY crystals with NI-hNO3 solutions, it is possible to reduce the Na level to 0.1% Na20 from 3.50% Na/O without altering the crystals' surface area or Re203 content [44]. Then, in the presence of 0.4% V, a CREY sample containing 0.11% Na20 can retain 61% of its original crystallinity after steam--aging; even with 0.8% to 1.2% V, these crystals show a residual (10-20%) crystallinity (Figure 12). As observed previously [22] HY is generally more V resistant at hydrothermal conditions than CREY crystals containing comparable Na levels (Figure 12). Lanthanides (and chargecompensating Na-ions) removal and vanadate formation, together with the enhanced framework dealumination (owing to the acidity generated during steaming), are probably the main causes of the rapid and total collapse of the CREY crystals when steam-aged in the presence of surface V impurities. Table 7. The effect of calcination (at 760~ and Linde's HY (LZY-82).

air) on the surface area retention (%) ofDavison CREY % Retention of Zeolite Surface Area

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

Vanadium, wt%:

0.0

1.0

2.0

3.0

4.0

5.0

HY CREY

100 100

95 91

94 75

94 69

76 64

14 12

5. ATOMIC FORCE MICROSCOPY The atomic force microscope (AFM) can provide details of the surface topography of an FCC with unprecedented resolution [47,48]. Atomic-scale imaging has allowed the identification of surface openings (or pores) with variable length (L) and width (W) irregular in size and shape [47,48]. Slits with LAV >>> 1 and width in the 6 to 9 nm range appear most frequently; valleys, kinks, and cracks are believed to represent the major sources of the FCC microporosity [47,48] (Figures 14A- 14D). Large pits 3.0~tm in size are frequently observed. The architecture and structural features of these pores are defined by the mode of plates aggregation (Figure 14A- 14B). The walls of large pits offer terraces separated by steps about 0.2-0.3~tm high that could act as docking sites for hydrocarbon adsorption and cracking [47,48] (Figure 14B). Large pores with diameter in the 1 to 51.tm range could facilitate surface retention of gas oil to be removed during steamstripping or burned during FCC regeneration. It is believed that in coke-selective microspheres,

39

Figure 14. A) Large scale AFM image showing a micropore on the FCC surface. B) Details of a micropore wall. C) Opening or pores formed by missing plates. D) Narrow slits, or cracks, 6.0-9.0nm wide are otten observed on the FCC surface. E) Large scale AFM image of steamed GRZ-1 containing 2%V. F) Vanadia filled slits representing pore blockage [47,48].

40

Figure 14. (Continued)

41

Figure 14. (Continued)

42 the catalyst components form a house-of- cards-structure that minimizes macropore formation on the surface. When contaminated with 2% to 4% V, during steaming at 760~ part of the catalyst surface appears to lose some of its surface roughness (Figure 14E). The smaller pores (slits) appear blocked (Figure 14F). Pore blockage and crystallinity losses are probably the two main causes of the drastic reduction in surface area and cracking activity suffered by GRZ-1 (and by other FCC) in the presence of 2% to 4%V. It is believed that the gain in cracking activity exhibited by spent equilibrium FCC atter reactivation with the DEMET process [49] is probably due to the removal ofvanadia impurities from the catalyst's microporous structure. Since the microscope cantilever does not damage or alter the surface, the AFM can provide the actual images of working catalysts. The AFM's major limitation appears to be its inability to provide chemical composition data of the surface. 6. VANADIUM-RESISTANT FCC Should vanadium impurities on equilibrium catalysts remain at the present level (i.e. near 0.16 wt%), losses in catalytic activity in a FCC could probably be avoided by prudent unit design and increased fresh catalyst makeup rates. For V levels in the 0.1 to 0.5 wt% range, cracking activity in an FCC can be maintained by incorporating, into the FCC matrix, metal scavengers (metal traps) capable of forming inert V compounds. Calcium titanates with the perovskite structure are examples of effective V passivators [50]. The advantage of using strontium titanate (SrTiO3) has been demonstrated in pilot plant and commercial trials [51 ]. E. Kugler [52] has also proposed the use of Sr compounds (SrCO3) as V passivators. Sacrificial zeolites [25], sepiolite [53], and the use of anionic clays [54] are other examples of additives that could be incorporated into the FCC matrix to improve the catalyst V tolerance. For V levels above 0.5 wt%, a different approach has been recommended to avoid FCC deactivation by high V levels. It has been reported in the patent literature that vanadium (and nickel) resistance in an FCC can be significantly enhanced by the addition of certain diluents (metal scavengers) capable of selectively sorbing V[55-57] impurities. In fact, sepiolite addition to a high-activity commercial FCC can generate dual-functional cracking catalyst (DFCC) mixtures that, even when metal contaminated with as much as 1.5wt% vanadium, have been found capable of retaining useful cracking activity (70% conversion) when cracking a light gas oil (with an API gravity of29.6)at microactivity test (MAT) conditions; improved coke and hydrogen selectivity were also observed. Transport experiments at MAT conditions, together with pilot plant results [58], have indicated that the DFCC enhanced vanadium resistance can be attributed to the gas phase transport of this metal from the host catalyst to the diluent (sepiolite) where it is sorbed and passivated. Metal scavengers reported to date have few cracking properties. Thus, their addition to a commercial FCC initially cause a decrease in cracking activity. However, as metals are selectively deposited on the scavenger surface, deactivation rates are significantly reduced and a cross-over point is reached at which the FCC mixture is more active (in the presence of high metals) than the parent (undiluted) FCC (Figure 15). Cross-over points depend on gas oil composition, feed metal levels, and zeolite concentration (and type) in the host FCC and on the properties of the diluents; this data must be obtained experimentally. Other important concepts for residuum catalyst development have recently been reviewed by O'Connor and coworkers[59].

43

Ms

DFCC CROSSO V E R / ~

POINT A

i

~

"

i i | a

,

~176 J

METALS (Ni AND V)

Figure 15. The enhanced preservation of cracking activity in a DFCC mixture (A) is attributed to the metal scavenger's (MS) ability to irreversibly sorb migrating V (B)thus minimizing direct V-FCC interactions. 7. S U M M A R Y AND CONCLUSIONS

Results from this and other studies [22] have indicated that FCC deactivation by V contaminants can occur by two different mechanisms depending on the way the FCC cracking component (the zeolite) has been stabilized. Residual Na-ions, RE cations, framework Si, A1 composition and extra framework A1 are parameters believed to influence FCC resistance to Vinduced deactivation. During thermal treatment (in air) of V contaminated faujasite crystals, the following have been observed: 1) Calcination (540~ in air, induced dealumination and unit cell contraction in all the HY crystals examined without affecting their crystallinity or surface area. 2) The oxidative decomposition (at 540~ of HY saturated with solutions ofVO 2+naphthenate (in benzene or toluene) promote dealumination. Extra framework A1 and contraction of the HY unit cell dimension increase with V levels (% V < 5.0); however, these changes in the crystals' lattice have little effect on the crystallinity (as measured by XRD) and surface area measurements. 3) For a set of HY (with 0.1 < % Na20 < 2.5 and 3.8 < Si/AI < 11.0), V-resistance during calcination in air does not seem to depend on Na20 levels or Si/A1 ratios. 4) After increasing the calcination temperature in air to 760~ from 540~ HY with 5% V collapse forming mullite and silica. A third phase, vanadia, can be identified by LRS. 5) NaY crystals readily collapse in the presence of vanadia when heated in air at relatively low calcination temperatures [36]. 6) Therefore, V-loaded HY-type crystals can collapse when heated in air with an ease that depends on calcination temperature, vanadium and sodium levels. 7) When calcined in air at 540~ levels up to 5%.

CREY crystals have similar V-tolerance to HY for V

44 When the thermal pretreatment is performed in the presence of steam (at 1 atm)the stability of the faujasite structure is greatly affected by the presence of residual Na ions [41,42]. Thus, the V tolerance of different zeolites should be compared at similar Na-levels. In the presence of steam, the deleterious effects of V impurities on zeolite properties are greatly enhanced. Specifically: 1) In the presence of steam (at 760~ HY crystals containing only about 2% V collapse, forming mullite, tridymite, and a vanadia-like phase. Incorporation of V into the mullite structure is believed to occur [21 ]. 2) HY resistance to V attack increases with increasing framework Si/AI ratios, and with decreasing Na levels. 3) Zeolites with the pentasil structure are more V-tolerant than HY (or CREY) and collapse when steam-aged in the presence of V, forming cristobalite and vanadia. Thus, it is believed that vanadium (VO § or VO2§ preferentially attack Si-O-AI bonds in zeolites. 4) During steaming, hydrolysis products of V § compounds (such as V205) generates acids (such as H4V207) that further promote Si-O-A1 bond breakage, dealumination reactions, and therefore lattice collapse. 5) The reaction between residual Na § ions in HY and V impurities (such as V205) can lead to the formation of stable sodium vanadates [36]. Removal of these charge-compensating cations destabilizes the zeolite lattice. 6) In the presence of V, CREY crystals collapse during steaming forming a REVO4 phase. There is a nearly linear relationship between REVO4 formation and V levels up to 5% [29]. 7) The V tolerance of CREY crystals can be improved by removing Na ions in these commercially available materials with NH4NO3 solutions. However, for a given Na-level, HY type crystals seem to be more V tolerant than CREY. This difference has been attributed to the ease with which V impurities can react with the various RE ions (especially Ce +4) present in CREY samples. Examination of the topography of a fluid-cracking catalyst using atomic force microscopy has revealed the presence of a unique surface architecture characterized by valleys, ridges, crevices, dislodged plates, and narrow slits 6 to 9 nm wide. When V was added at the 2% to 4% level, the catalyst surface roughness decreased. Furthermore, AFM images indicate the formation of vanadia islands and coating of the surface with vanadia causing blockage of the narrow slits and cracks responsible for most of the catalyst's microporosity. In conclusion, it is proposed that heavily V-contaminated crudes should be cracked with DFCC mixtures that, in addition to having an effective metal scavenger (such as sepiolite or attapulgite), contain a host FCC in which the cracking centers are provided by Si enriched HY type zeolites with high framework Si/AI, extra framework AI, low (<0.5% Na20) residual Na ions, and low levels of RE cations. The AFM can provide details of the surface topography of a FCC with unprecedented resolution. Because all measurements were performed in air, the AFM allowed the characterization of solid surfaces without the possibility of altering their molecular features. Thus, in addition to FCC, the AFM could be particularly useful in studying other commercially important materials such as hydrotreating catalysts.

45 8. ACKNOWLEDGEMENTS

The support received from the Unocal Analytical Department staff is gratefully acknowledged. Special thanks are due to Dr. P. S. Iyer for providing NMR data and for participating in many useful discussions. REFERENCES

1. 2. 3. 4. 5. 6. 7. 8. 9. 10. 11. 12. 13 14. 15 16. 17 18 19. 20. 21. 22. 23. 24. 25. 26. 27.

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Catalysts in Petroleum Refining and Petrochemical Industries 1995

M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All fights reserved.

49

Metal Clusters in Zeolites: Nearly Molecular Catalysts for Hydrocarbon Conversion B. C. Gates Department of Chemical Engineering and Materials Science University of California, Davis, California 95616, U.S.A. ABSTRACT Metals supported in the pores of zeolites have long been used as catalysts, but only recently have samples been prepared and characterized that incorporate metal particles smaller than about 1 nm in diameter, referred to here as clusters. Nearly uniform zeolite-supported metal clusters are the focus of this review, which addresses their synthesis, physical characterization, and catalytic properties. Supported clusters represented as Ir4, Ir 6, Rh4, and Rh6 were prepared in the supercages of faujasite zeolites by ship-in-a-bottle synthesis of the metal carbonyl cluster precursors [Ir4(CO)12], [Ir6(CO)16], [Rh4(CO)12], and [Rh6(CO)16], respectively, which were decarbonylated. Alternatively, Ir4 and Ir 6 clusters in zeolite NaX were prepared by synthesis and then decarbonylation of [HIr4(CO)ll]- and [Ir6(CO)15] 2-. Metal clusters that are nearly as small and uniform as these have also been formed from metal salt precursors in various zeolites, including zeolites NaY and KLTL. The most important of these are clusters of 5-6 Pt atoms each, on average, in basic zeolite LTL. These were made from platinum ammine salts by calcination and reduction in hydrogen. Similarly, clusters of iridium nearly this same size were prepared from an iridium ammine salt. The supported platinum clusters are active and selective industrial catalysts for dehydrocyclization of naphtha to give aromatics, whereas the supported iridium clusters are unselective. The unique selectivity and activity of the zeolite-supported platinum clusters is associated with the basicity of the support and the smallness of the platinum clusters; the role of the zeolite is not fully understood. Opportunities for discovery of supported metal clusters with new catalytic properties appear to be excellent. 1. INTRODUCTION: ZEOLITE-SUPPORTED METALS Zeolites have been used for years as supports for metal catalysts [1-5]. Such catalysts are typically made by impregnation of the zeolite with an aqueous solution of a metal salt, followed by calcination and reduction in hydrogen. Because the metal particles in such catalysts are typically extremely small and nonuniform in size and shape, often being present both inside and outside the zeolite pore structure, their structures are not well understood. This structural complexity provides a fundamental motivation for preparing and investigating structurally simple zeolite-supported metals, those that are so small and uniform as to be nearly molecular in character and located almost entirely within the zeolite pores; investigations of well-defined

50 clusters in zeolites could determine how catalytic properties depend on cluster size and interactions with the support. There is also an important practical motivation for investigating such materials, related to the recent discoveries that extremely small clusters of platinum in the pores of basic zeolite LTL are highly active and selective catalysts for dehydrocyclization of straight-chain paraffins. These catalysts are now applied industrially for naphtha reforming to make aromatics, and there is only a partial understanding of why they work so well. The average size of the metal particles in a supported metal catalyst is usually estimated from the dispersion (fraction of metal atoms exposed) measured by titration of the metal surface sites by hydrogen or CO chemisorption. Complementary measurements giving structural information are made with transmission electron microscopy and extended X-ray absorption fine structure (EXAFS) spectroscopy, among other methods. Very small supported metal particles have often been modeled as clusters or crystallites having the symmetries of bulk metals [6]. For example, a six-atom platinum particle may be represented as an octahedron with an edge length of about 0.55 nm and a dispersion of 1. A 19-atom platinum particle with such a symmetry has an edge length of about 1 nm and a dispersion of 0.95. Supported metals with diameters < 1 nm are the focus of this review; they are referred to here as metal clusters to distinguish them from the larger particles. Clusters as small as those with diameters < 1 nm are comparable to true molecular metal clusters, which are stabilized by ligands such as CO. Examples of such clusters are [Ir4(CO)12] and [Ir6(CO)16]. Ligand-stabilized metal clusters on supports are often good precursors of the supported metal clusters of interest here. Virtually all the preparations of nearly uniform supported metal clusters have involved decarbonylation of metal carbonyl cluster precursors, most with robust metal flames, e.g., [Ir4(CO)12] [7]. For example, [Ir4(CO)12] in the supercages of zeolite NaY, upon treatment in 1-12 at 300~ [8], gives decarbonylated clusters that are modeled as Ir4, as described below. However, there are still only a few examples of supported clusters formed from metal carbonyl clusters without destruction of the metal frame. Most attempts to prepare supported metal clusters by decarbonylation of supported metal carbonyls have led to poorly defined mixtures, including mononuclear (single-metal-atom) complexes, and/or clusters, and/or metallic particles. Nonuniform materials such as these, which have been reviewed extensively [ 1-3], are largely ignored here. 2. ZEOLITE SUPPORTS FOR METAL CLUSTERS Crystalline aluminosilicates (zeolites) have pores with diameters of the order of 1 nm. The smallness and regularity of these pores account for shape-selectivity and many of the important applications of zeolites in acid catalysis. Zeolite frameworks consist of linked TO 4 tetrahedra (T = Si, AI). The zeolites that have been most often investigated as supports for metal clusters are faujasites (zeolites X and Y), which have three-dimensional pore structures incorporating nearly spherical cages with diameters of about 1.2 nm connected by apertures that are 12membered oxygen tings, with diameters of about 0.75 nm. Zeolite LTL, which is used as the support for industrial aromatization catalysts, has a two-dimensional pore structure consisting

51 of parallel, nonintersecting channels incorporating ellipsoidal cages with dimensions of about 0.48 • 1.24 • 1.07 nm; the cages are connected by 12-membered oxygen rings. Zeolites are unique as supports for metal clusters because the steric restrictions offered by their cages limit the sizes of clusters that can form in them, and the restrictions offered by the apertures separating the cages limit what can enter and leave the cages. Thus clusters can be formed from small precursors in the cages and trapped there. The cages are small enough to exert solvent-like effects on clusters within them, and thus the cages may cause the cluster structures and properties to be different from those of clusters in solution or on the surfaces of amorphous supports, which usually have pores much larger than those of zeolites [5]. Confinement in cages may hinder cluster interactions and thereby increase cluster stability. 3. PREPARATION OF ZEOLITE-SUPPORTED METAL CLUSTERS

3.1. Ship-in-a-Bottle synthesis of metal carbonyl clusters Metal carbonyl clusters on supports are important to the subject reviewed here because they are the best known precursors of structurally simple supported metal clusters, which are formed by decarbonylation of the precursors. The routes for preparation of molecularly or ionically dispersed metal carbonyl clusters on zeolite and metal oxide supports include syntheses from mononuclear precursors on the support surface [5,9]. Ship-in-a-bottle syntheses of this type take place when the clusters formed in zeolite cages are trapped there because they are too large to fit through the apertures. Syntheses in the nearly neutral NaY zeolite are similar to those occurring on the nearly neutral ~t-Al203 and in nearly neutral solutions. Examples are the syntheses of [Ir4(CO)12] [8] and of [Ir6(CO)16] [10] from [Ir(CO)2(acac)] in the presence of CO. Syntheses in the more basic NaX zeolite are similar to those occurring in basic solutions and on the basic surface of MgO, e.g., those of [HIr4(CO)l 1]- and [Ir6(CO)15]2- [11]. 3.2. Decarbonylation of supported metal carbonyi clusters The preparation of supported metal clusters by decarbonylation of supported metal carbonyl clusters is exemplified by the removal of the CO ligands from the metal frame of [Ir4(CO)12] dispersed in NaY zeolite by treatment in hydrogen at 300~ [12]. Decarbonylation of [Ir6(CO)16] in NaY zeolite cages occurs similarly [ 10,12]. The decarbonylations of iridium carbonyl clusters in NaY zeolite are reversible [ 12]. Infrared spectra show that [Ir6(CO)16] in the supercages was decarbonylated by treatment in hydrogen at 300~ [12]. When CO was adsorbed at -196~ on the decarbonylated clusters formed from [Ir6(CO)16] (or from [Ir4(CO)12]) and the temperature raised with the sample under CO, mononuclear iridium carbonyls formed at about -30~ These were converted at about 50~ into [Ir4(CO)12] and at about 125~ into [Ir6(CO)16] [12]. In contrast, the decarbonylations of the iridium cluster carbonyl anions [HIr4(CO)ll]- and [Ir6(CO)15]2- supported on MgO were found to be irreversible [13 ].

52 The decarbonylation of the zeolite-supported iridium carbonyl clusters takes place with little or no disruption of the metal frame, as discussed below [ 14]. Details of the decarbonylation chemistry are still unknown, but the simple decarbonylation may be relatively unsurprising in view of the stability of the Ir4 frame indicated by electrospray mass spectrometry showing that CO ligands are peeled off one by one from a salt of [HIr4(CO)ll]-, giving clusters represented as HIr4(CO)ll-x (x = 0, 1, 2, ....11), although the charge of these gas-phase clusters and whether they incorporate hydrogen is unknown [ 15]. Chemistry similar to that described above for iridium clusters has also been observed for rhodium clusters. Several authors [16-18] have prepared [Rh6(CO)16] in NaY zeolite; [Rh4(CO)12] has also been formed [18], and each of these has been decarbonylated with minimal changes in the metal frame, as shown by EXAFS spectroscopy [18]. Thus there appears to be some generality to the method of forming small clusters in zeolite cages by synthesis of stable metal carbonyl precursors followed by decarbonylation. However, the method is limited. Attempts to use it to prepare zeolite-supported platinum clusters that are structurally simple and uniform have apparently not been successful. The literature of platinum carbonyl clusters in zeolites is not considered here because it is still contradictory. Because, with some generality, zeolites seem to favor the formation of small clusters and stabilize them, zeolites may be the optimum supports for small, nearly uniform metal clusters. Perhaps the confining environments of the zeolite pores stabilize the clusters and minimize aggregation of the metal into larger clusters and particles. However, these ideas are speculative, and more evidence is required. 3.3. Reduction of exchange cations Zeolite-supported metal clusters have most commonly been prepared by ion exchange followed by reduction in hydrogen [1-3]. Usually the metals are introduced as cationic complexes, (e.g., [Pt(NH3)4]2+),which replace cations such as Na + in the zeolite and are then decomposed by heating in oxygen or air (calcination) followed by reduction in hydrogen. Alternatively, the metal complex may be introduced by impregnation of the zeolite with an aqueous solution, e.g., by the incipient wetness method, whereby just enough solution is added to fill the zeolite pores. Sample preparation is not easily reproduced; it depends, for example, on the nature and crystallographic locations of the cations, the presence of oxidizing agents such as hydroxyl groups, and the presence of residual water. Reduction of metals in zeolites often leads to metal particles that are too large to fit in the cages, and the result is breakup of the zeolite framework. Often much of the reduced metal then resides outside of the intracrystalline space. Not uncommonly, literature reports of zeolite-supported metals have not provided evidence to determine whether the metal is present inside the pores, outside the pores, or both.

Cations of noble metals in zeolites are easily reduced by hydrogen, but activation and reduction treatments to give the highest metal dispersions have not o~en been determined. Treatments of ion-exchanged zeolites to give maximum metal dispersion should minimize migration and sintering of the noble metal. For example, prior to reduction, it is usually thought to be necessary to eliminate ammonia produced by thermal decomposition of

53 ammonium cations or ammine complexes of the noble metal, because reduction of noble metal cations in the presence of evolving ammonia often leads to formation of agglomerated metal [19]. Direct reduction of noble metal cations by hydrogen at high temperatures may lead to formation of mobile species, causing metal agglomeration and low metal dispersions [20]. Water may also be detrimental to metal dispersion [20]. Activation in flowing oxygen prior to reduction in hydrogen gives highly dispersed platinum clusters in zeolites. The temperature of oxidation is critical in determining the metal dispersion, as shown in the early work of Dalla Betta and Boudart [20]. These authors investigated a sample prepared from Pt(NH3)4CI2 in CaY zeolite, finding that calcination in oxygen at 350~ followed by treatment in hydrogen at 400~ gave highly dispersed platinum in the zeolite. Reagan et al. [ 19] found that ammonia from the platinum complex reduced the platinum, concluding that reduced platinum is always the product of the thermal decomposition (at 300~ or higher temperatures) of platinum ammines in Y zeolite, even in air. These authors recommended 300~ as the optimum calcination temperature in air, and their result was supported by subsequent results [21 ]. Similarly, highly dispersed platinum (with about 6 atoms per cluster, on average) in the pores of H-mordenite was prepared from Pt(NH3)4(OH)2, with calcination at 350~ and reduction in hydrogen at 350~ [22]. Somewhat different values seem to pertain to platinum clusters in zeolite KLTL. Extremely small clusters were formed in this zeolite by aqueous impregnation with tetraammineplatinum (II) nitrate followed by calcination at 260~ and reduction in hydrogen at 500~ [23]. The patent literature [24] refers to this same low calcination temperature, which is inferred to favor the formation of extremely small clusters in zeolite LTL. The apparent differences between the zeolites with respect to the optimum calcination temperature are not resolved. More experiments are needed. Alternatively, extremely small clusters of platinum have been formed in zeolite LTL without a calcination step. Vaarkamp et al. [25] prepared platinum clusters of about 5 or 6 atoms each, on average, from Pt(NH3)4(NO3) 2 in zeolite BaKLTL simply by reduction at 500~ TriantafiUou et al. [26] similarly prepared iridium clusters from [Ir(NH3)5C1]CI 2 in zeolite KLTL by reduction in hydrogen at 300 or 500~ also without calcination; the average cluster contained about 5-6 iridium atoms. 4. STRUCTURAL CHARACTERIZATION OF ZEOLITE-SUPPORTED M E T A L

CLUSTERS Metal clusters in zeolites are difficult to characterize structurally because the clusters are small, often nonuniform, and present in only low loadings. Only recently have persuasive structural data become available, as a result of investigations by methods including EXAFS spectroscopy, high-resolution transmission electron microscopy, and complementary methods, including 129Xe NMR spectroscopy, hydrogen and CO chemisorption, and wide angle X-ray scattering. EXAFS spectroscopy [27] has emerged as the most reliable single technique, but it is best when complemented by other methods, particularly high-resolution transmission electron

54 microscopy. EXAFS provides the most precise structure data characterizing supported clusters when the clusters are very small and nearly uniform [28]. The technique is limited because it provides only average structural information and relatively imprecise values of coordination numbers, with errors of typically +20%. EXAFS results have been reported for several zeolite-supported clusters derived from metal carbonyl clusters (Table 1). For example, supported clusters formed by decarbonylation of NaX zeolite-supported iridium carbonyls believed to be [HIr4(CO)ll]-, which has a tetrahedral metal frame, have an average Ir-Ir first-shell coordination number of 3, the value for a tetrahedron. Transmission electron microscopy is valuable because it allows imaging of even the smallest supported clusters of a heavy metal such as platinum. Dark field microscopy has been used to advantage for platinum dusters consisting of < 20 atoms each in zeolite KLTL [30]. Even single platinum atoms can be detected on zeolite supports thinner than about 20 nm, but the precision with which clusters can be pinpointed in the structure is limited by beam damageinduced distortion of the zeolite framework [30]. Hydrogen chemisorption is not yet a routine characterization method for supported metal clusters because stoichiometries of chemisorption on the clusters are not well known and are different from those of chemisorption on metal crystallites [ 13]. Chemisorption of CO is also of limited value because CO typically reacts with supported clusters, leading to changes in cluster structure. Research is needed to clarify these matters. Although it has long been recognized that very highly dispersed metals in zeolites can be prepared, the literature of these materials indicates the difficulty of the characterizations and the recent advances resulting from the use of modem physical methods, especially EXAFS spectroscopy. These points are illustrated by the research with zeolite-supported platinum reported over a period of years by Boudart and coworkers. A succession of experiments characterizing platinum in zeolite Y with essentially all the available techniques, summarized by Boudart et al. [31 ] and by Samant and Boudart [32], has led to estimates of the average number of platinum atoms per cluster ranging from 6 to about 40. It is difficult to provide precise statements of cluster sizes when the clusters are present in mixtures and when the clusters contain more than a few atoms. A summary of zeolite-supported metal clusters prepared from metal salts is given in Table 2. Included here are only samples that have been characterized by EXAFS spectroscopy and incorporate extremely small clusters. The additional literature of zeolite-supported metals is reviewed elsewhere [ 1-5]. The results of Table 2 show that with well chosen conditions it is possible to prepare extremely small clusters in zeolites with conventional methods. It is not yet quite clear whether the zeolites are different from amorphous supports in giving such highly dispersed metal clusters. Perhaps the cages and apertures limit the sizes of the clusters by hindering the interactions of species that migrate and sinter readily on amorphous supports.

55 Table 1 Zeolite-supported iridium clusters formed by decarbonylation of supported iridium carbonyl clusters: characterization by EXAFS spectroscopy a

Zeolite support

Precursor

Cluster modeled

N

R, nm

ref

as

NaY

[Ir4(CO)12] formed from [Ir(CO)2(acac)]

Ir4 tetrahedra

3.4

0.270

8

NaY

[Ir6(CO)16] formed from [Ir(CO)2(acac)]

Ir6 octahedra

3.6

0.271

10

CaA

Pd(NH3)4(NO3) 2

Pd6 ?

3.5

not stated

29

NaX

[HIr4(CO)ll]- formed

Ir4 tetrahedra

3.0

0.271

11

from [Ir(CO)2(acac)] aNotes: N is coordination nmnber, R the average absorber-backscatterer distance. Typical experimental errors in N and R are approximately+20% and +2%, respectively.

It must be assumed that the samples listed in Table 2 have distributions of cluster sizes, although the available methods do not provide good evidence of the distributions. It seems likely that zeolite-supported metal clusters made from metal carbonyl clusters (Table 1) incorporate more nearly uniform clusters than samples made by conventional methods from metal salts, but this suggestion is not yet tested. In summary, zeolite-supported metal clusters have now been prepared that are so small and apparently nearly uniform in size that they are regarded as nearly molecular. Preparations with metal carbonyl cluster precursors are the best known for making nearly uniform and thus nearly molecular supported clusters, but it is clear that conventional preparation methods based

56

Table 2 Zeolite-supported metal clusters formed from salt precursors a

Zeolite support

Precursor

Characterization methods

Metal-metal coordination number

Approximate average cluster nuclearity

Comment

ref

BaKLTL

Pt(NH3)4(NO3) 2

EXAFS, TEM, H2 chemisorption

3.7

5-6

Sample not calcined after addition of Pt; reduced at 500~

25

KLTL

Pt(NH3)4(NO3) 2

EXAFS, H2 chemisorption, TPD of hydrogen

4.0 + 0.1

6

Sample not calcined after addition of Pt; reduced at 300~

33, 34

KLTL

Pt(NH3)4(NO3) 2

EXAFS, H2 chemisorption, TPD of hydrogen

4.8 + 0.2

10

Sample not calcined after addition of Pt; reduced at 500~

33, 34

KLTL

Pt(NH3)4(NO3)2

EXAFS, H2 chemisorption, TPD of hydrogen

4.9 + 0.1

12

Sample not calcined after addition of Pt; reduced at 600~

33, 34

KLTL

[Ir(NH3)5CI]CI 2

EXAFS

3.2

4

Sample not calcined after addition of Ir; reduced at 300~

26

KLTL

[Ir(NH3)5C1]CI2

EXAFS

4.2

6

Sample not calcined after addition of Ir; reduced at 500~

26

57 HLTL

Pt(NH3)4(NO3) 2

EXAFS, H2 chemisorption, TPD of hydrogen

4.1 + 0.1

6

Sample not calcined after addition of Pt; reduced at 300~

33, 34

HLTL

Pt(NH3)4(NO3) 2

EXAFS, H2 chemisorption, TPD of hydrogen

4.4+0.1

9

Sample not calcined after addition of Pt; reduced at 500~

33, 34

HMAZ

Pt(NH3)4(NO3) 2

EXAFS, H2 chemisorption

2.9 • 0.2

4

Sample not calcined after addition of Pt; reduced at 500~

33

HMOR

Pt(NH3)4(OH)2

EXAFS, TPD of hydrogen

3.7

6

Sample calcined at 350~ after addition of Pt; reduced at 350~

22

KHMOR

Pt(NH3)4(OH)2

EXAFS, TPD of hydrogen

3.9

6

Sample calcined at 350~ after addition of Pt; reduced at 350~

22

on the use of salt precursors can be manipulated to give clusters as small as those made from the metal carbonyls. The results for supported iridium are striking in that the cluster sizes in samples made by these two types of procedure overlap. These samples offer a long-envisioned opportunity to vary metal cluster size and to determine how the catalytic properties depend on it. Results of some of the first experiments with zeolite-supported metal cluster catalysts are summarized below. 5. CATALYSIS BY ZEOLITE-SUPPORTED METAL CLUSTERS It is likely that metal clusters have been present for years in conventional supported metal catalysts, such as those used for naphtha reforming, but because of the difficulty of distinguishing the small clusters from larger metal crystallites, it has not been possible to identify and define the roles of clusters. Evidence of catalysis by supported metal clusters has arisen only recently in work with catalysts containing the metal almost entirely in the form of clusters. The following section is a summary of catalytic results for zeolite-supported metal clusters that have been characterized by EXAFS spectroscopy.

58 5.1. Catalysis by zeolite-supported Ir4 and Ir6 clusters Catalytic properties of supported clusters identified as primarily Ir4 or Ir6 were reported by Xu et al. [15], who investigated a structure-insensitive reaction, toluene hydrogenation. The support was NaY zeolite or, for comparison, MgO. EXAFS spectroscopy (Table 3) showed that the first-shell Ir-Ir coordination numbers characterizing both the flesh and used MgOsupported catalysts made by decarbonylation of supported [Ir4(CO)12] or [HIr4(CO)ll]- are indistinguishable from 3, the value for a tetrahedron, as in [Ir4(CO)12] and [HIr4(CO)ll]-. The decarbonylated clusters retained or nearly retained this metal frame. EXAFS data show that the decarbonylated Ir6 clusters had metal frames indistinguishable from the octahedra of the precursor hexairidium carbonyls, indicated by the coordination number of approximately 4. Catalytic activities of the zeolite-supported clusters (Table 4) are reported as turnover frequencies; these are rates per total iridium atom for such small clusters. Rates were also reported for conventional (structurally nonuniform) supported catalysts consisting of aggregates of metallic iridium on supports; these rates, per unit of metal surface area, are markedly greater than those observed for the supported clusters [15]. Changing the support from zeolite NaY to MgO had little effect on the activities of the decarbonylated clusters. Although the Ir4 and Ir6 clusters catalyze the same reactions as metallic iridium particles, their catalytic character is different, even for structure-insensitive hydrogenation reactions. It is inferred [ 15] that the clusters are metal-like but not metallic; consistent with the structural inferences stated above, we refer to them as quasi molecular. Thus these data show the limit of the concept of structure insensitivity; it pertains to catalysis by surfaces of structures that might be described as metallic, i.e., present in three-dimensional particles about 1 nm in diameter or larger. This conclusion suggests that supported metal clusters may be found to have catalytic properties superior to those of conventional supported metals for some reactions. The suggestion finds some support in the results observed for platinum clusters in zeolite LTL, as summarized below. 5.2. Catalysis by clusters of Pt and of Ir in zeolite LTL Supported metal clusters as small as those described in the preceding paragraphs are now important in catalytic technology. These catalysts are used commercially for naphtha reforming for production of aromatics [35,36]. The catalysts consist of platinum clusters in zeolite LTL made basic by the presence ofK + or K + and Ba 2+ exchange ions [37-39]. Several industrially prepared catalysts of this type had almost all the clusters in the zeolite pores, as shown by EXAFS spectroscopy, transmission electron microscopy, and hydrogen chemisorption (Table 2). EXAFS spectra indicate first-shell Pt-Pt coordination numbers of 45, indicating clusters of, on average, about 5-12 atoms. Dark field electron micrographs [30] have led to similar conclusions for industrially prepared catalysts.

59 Table 3. EXAFS results characterizing fresh and used supported metal cluster catalysts

sample number

Catalyst precursor

'

Zeolite support

Treatment/ catalysis a

N, b fresh catalyst

R,c~nm, fresh catalyst

N, b used catalyst

R,Cnm, used catalyst

ref

3.9

0.270

15

1

[Ir6(CO) 16]

NaY

decarbonylation and catalysis of toluene hydrogenation

2

[Ir6(CO)16]

NaY

decarbonylation, treatment in I-I2at 300~ and catalysis of toluene hydrogenation

3.6

0.271

4.1

0.269

10, 15

3

Pt(NH3)4(OH)2HMOR calcinationat 350~ treatment in H2 at 350~ and catalysis of nhexane isomerization

3.7

0.272

not stated, but nearly 3.7

not stated

22

aConditions of catalysis experiments stated in Table 4. bN is the first-shell metal-metal coordination number; CRis the metal-metal distance, both determined by EXAFS spectroscopy.

The performance of platinum supported in LTL zeolite catalyst is well illustrated by data of Lane et al. [23] for conversion of n-hexane in the presence of excess H2 at 330-440~ and atmospheric pressure (Table 4). Primary products were observed to form both from one-six and one-five ring closure, giving benzene and methylcyclohexane, respectively. The catalyst is remarkable for its high benzene selectivity, which increases with increasing conversion of nhexane because some of the primary products are further converted into benzene. For example, ultimate benzene selectivities (defined as the hexanes converted to benzene divided by the hexanes converted to benzene and light, C1-C5, hydrocarbons) as high as 93% were observed at 420~ Several explanations have been advanced for the unique performance of Pt/LTL zeolite catalysts. There is a consensus that dehydrocyclization is catalyzed by the platinum clusters alone, with the support providing no catalytic sites [23,38,40]. The support must be nonacidic to prevent acid-catalyzed isomerization and hydrocracking as side reactions [41-43]. The aromatic selectivity increases with the basicity of the LTL zeolite support. The interaction of the platinum clusters with the basic support has been suggested to result in an increase in the

60

Table 4 Catalytic activities of supported metal clusters

S~i~ie..............cataiyst ................. Zeolite . . . . . . . number a

modeled

Catalytic reaction b

support

H2 treatment

103 x TOF, c

temperature of

s-1

ref

catalyst, ~

as

1

Ir6

NaY

toluene hydrogenation

no 1-12treatment

0.25

46

2

Ir6

NaY

toluene hydrogenation

300

0.52

46

3

approximately H M O R Pt6, on average

n-hexane isomerization

350

not stated; rate influenced by diffusion

58

4

approximately KLTL Pt6, on average

n-hexane aromatization

500

2

23

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

- . . . . . . . . . . . . . . .

,7-

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

". . . . . . . . . . . . . . . . . . . . . . . . . .

"

.

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" . . . . . . . . . . . . . . . . . . .

-

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~

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- . . . . . . . .

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aSample numbers match those of Table 3. bReaction conditions: Vapor-phase reactants in a once-through fixed-bed reactor operated at atmospheric pressure. The molar ratio of hydrogen to toluene was 14.2. The molar ratio of hydrogen to hexane was 7.8. CTurnover frequency for toluene hydrogenation at 60~ and for n-hexane conversion at 420~

electron density on platinum that favors the catalysis [44-47]. The steric environment of the platinum clusters may also be important; the one-dimensional pore structure of LTL zeolite has been suggested to orient the straight-chain paraffin parallel to the pore axis, thereby increasing the probability of terminal adsorption [48]. Because well-prepared catalysts incorporate extremely small clusters, with almost no platinum outside the pores, the high selectivity is associated with the low hydrogenolysis activity of platinum clusters smaller than about 1 nm [43]. The small zeolite cavities favor small, stable entrapped clusters. Iridium clusters in zeolite KLTL, like the platinum clusters, consisting of 4 to 6 atoms on average, have also been prepared by hydrogen reduction of [Ir(NH3)sCI]CIE in the pores at temperatures >300~ [26]. Even though the iridium clusters were as small as the selective platinum clusters in the same basic zeolite support, they were found to be unselective catalysts, being similar to other iridium catalysts for conversion of n-hexane and hydrogen principally into hydrogenolysis products. It is inferred that the combination of cluster size, electronic

61 modification of the iridium by the support, and the steric environment near the metal clusters was not sufficient to suppress the high intrinsic hydrogenolysis selectivity of iridium. Thus a comparison of these results with the results for platinum clusters supported in KLTL zeolite suggests [26] that two criteria are critical for a selective aromatization catalyst: a nonacidic support and a metal with low hydrogenolysis selectivity, like platinum. The uniquely excellent performance of platinum supported in KLTL zeolites as dehydrocyclization catalysts suggests that there is more to the story than is stated above, but there is too little information available now to tie all the loose ends together. The structure of the LTL zeolite support is inferred to be important in providing some, albeit rather small, improvements in activity and selectivity. In these catalysts, the hydrogenolysis activity is suppressed by alkali [47,49], which has been shown by infrared [45,50,51] and X-ray photoelectron spectroscopies [52] to increase the electron density on the metal clusters. The hydrogenolysis selectivity is also cluster size dependent, decreasing with decreasing cluster size [43]. Furthermore, the specific pore geometry of the LTL zeolite increases the selectivity for 1,6 ring closure relative to 1,5 ring closure. This has an indirect effect on the hydrogenolysis selectivity, which becomes important at high conversion. For example, catalysts with a high 1,5 ring closure selectivity produce high yields of methylcyclopentane, 2-methylpentane, and 3methylpentane. Before these can back-react to give benzene, they must first be converted into n-hexane through the methylcyclopentane intermediate. Each of the 1,5 ring closure products is also subject to hydrogenolysis. The higher the selectivity to 1,5 ring closure, therefore, the higher the hydrogenolysis selectivity in the limit as all the hexanes are converted into benzene, i.e., at high conversion. In summary, maximization of aromatization selectivity results from the choice of a catalytic metal with a low intrinsic activity for hydrogenolysis (platinum) and optimization of the support alkalinity and steric environment, which favors small clusters (which are highly selective) and perhaps otherwise favors the desired ring closure. Just how the zeolite affects the performance is not fully understood. 6. CONCLUSIONS Zeolite-supported metal clusters are a new class of catalyst made possible by syntheses involving organometaUic chemistry and by precisely controlled treatment of metal complexes in zeolite cages. Elucidation of the preparation chemistry would not have been possible without the guidance of EXAFS spectroscopy. Clusters such as Ir4, Ir6, and Pt n (where n is about 6) are small enough to be considered quasi molecular rather than metallic. Their catalytic properties are distinct from those of metallic particles, even for structure-insensitive reactions. The zeolite pores seem to confer some properties on the clusters that are not yet well understood. The results suggest that it may be fruitful to search for reactions for which supported metal clusters have catalytic properties superior to those of conventional supported metals. The important opportunity in catalysis may be to find reactions for which the activity or selectivity of supported metal clusters is superior to those of conventional supported metals. The high selectivity of Pt/LTL zeolite catalysts for paraffin dehydrocyclization, which is now exploited

62 commercially, is the most persuasive indication of the value of supported metal cluster catalysts. The high selectivity of this catalyst for dehydrocyclization is related to its low selectivity for hydrogenolysis, which may be related to the smallness of the platinum clusters, the basic environment of the zeolite pores, and the steric environment of these pores. ACKNOWLEDGMENT

This work was supported by the U. S. National Science Foundation (CTS-9300754). REFERENCES

1. P. Gallezot, Catal. Rev.--Sci. Eng., 20, (1979) 121. 2. P. Jacobs in B. C. Gates, L. Guczi, and H. KnOzinger (eds.), Metal Clusters in Catalysis, Elsevier, Amsterdam, 1986, pp. 357-414. 3. P. Gallezot in M. Moskovits (ed.), Metal Clusters, Wiley-Interscience, New York, 1986, pp. 219-247. 4. W. M. H. Sachtler and Z. Zhang, Adv. Catal., 20, (1993) 129. 5. S. Kawi and B. C. Gates in G. Schmid (ed.), Clusters and Colloids, VCH, Weinheim, 1994, p. 299. 6. O. M. Poltorak and V. S. Boronin, Russ. J. Phys. Chem., 40 (1966) 1436. 7. B. C. Gates, Chem. Rev. (1995) in press. 8. S. Kawi, J.-R. Chang and B. C. Gates, J. Phys. Chem., 97 (1993) 10599. 9. B. C. Gates and H. H. Lamb, J. Mol. Catal., 52 (1989) 1. 10. S. Kawi, J.-R. Chang, and B. C. Gates, J. Am. Chem. Soc., 115 (1993)4830. 11. S. Kawi, J.-R. Chang, and B. C. Gates, to be published. 12. T. Beutel, S. Kawi, S. K. Purnell, H. KnOzinger, and B. C. Gates, J. Phys. Chem., 97 (1993) 7284. 13. F.-S. Xiao, O. Alexeev, and B. C. Gates, J. Phys. Chem., 99 (1995), 1548. 14. S. Kawi, J.-R. Chang, and B. C. Gates, J. Phys. Chem., 97 (1993) 5375. 15. Z. Xu, F.-S. Xiao, S. K. Purnell, O. Alexeev, S. Kawi, S. E. Deutsch, and B. C. Gates, Nature (London), 372 (1994) 346. 16. E. Mantovani, N. Palladino, and A. Zanobi, J. Mol. Catal., 3 (1977-78) 385. 17. (a) E. J. Rode, M. E. Davis, and B. E. Hanson, J. Catal., 96 (1985) 574; (b) L.-F. Rao, N. Fukuoka, H. Kosugi, H. Kuroda, and M. Ichikawa, J. Phys. Chem., 94 (1990) 5317. 18 S. Kawi and B. C. Gates, to be published. 19 W. J. Reagan, A. W. Chester, and G. T. Kerr, J. Catal., 69 (1981) 89. 20 R. A. Dalla Betta and M. Boudart, Proc. 5th Int. Congr. Catal., 2 (1973) 96-1329. 21 A. W. Chester, J. Catal., 86 (1984) 16. 22 M. M. Otten, M. J. Clayton, and H. H. Lamb, J. Catal., 149 (1994) 211. 23 G. S. Lane, F. S. Modica, and J. T. Miller, J. Catal., 129 (1991) 145. 24 W. C. Buss and T. R. Hughes, UK Patent Application 2116450 (1983). 25 M. Vaarkamp, J. V. Grondelle, J. T. Miller, D. J. Sajkowski, F. S. Modica, G. S. Lane, B. C. Gates, and D. C. Koningsberger, Catal. Lett., 6 (1990) 369. 26. N. D. Triantafillou, J. T. Miller, and B. C. Gates, J. Catal., in press.

63 27. D. C. Koningsberger and R. Prins, R. (eds.), X-ray Absorption: Principles, Applications, Techniques of EXAFS, SEXAFS, and XANES, Wiley, New York, 1988. 28. B. C. Gates and D. C. Koningsberger, CHEMTECH, 22 (1992) 300. 29. Z. Zhang, F. A. Cavalcanti, and W. M. H. Sachtler, Catal. Lett., 12 (1992) 157. 30. S. B. Rice, J. Y. Koo, M. M. Disko, and M. M. J. Treacy, Ultramicroscopy, 34 (1990) 108. 31. M. Boudart, M. G. Samant, and R. Ryoo, Ultramicroscopy, 20 (1986) 125. 32. M. G. Samant and M. Boudart, J. Phys. Chem., 95 (1991) 4070. 33. M. Vaarkamp, F. S. Modica, J. T. Miller, and D. C. Koningsberger, J. Catal., 144 (1993) 611. 34. J. T. Miller, B. L. Meyers, F. S. Modica, G. S. Lane, M. Vaarkamp, and D. C. Koningsberger, J. Catal., 143 (1993) 395. 35. Oil Gas J., 190 (1992) 29. 36. D. Rotman, Chem. Week, 150 (1992) 8. 37. J. R. Bernard in L. V. C. Rees (ed.), Proc. 5th Int. Zeolite Conf., Heyden, London, 1986, p. 686. 38. T. R. Hughes, W. C. Buss, P. W. Tamm, and R. L. Jacobson, in Y. Murakami, A. Iijima, and J. W. Ward (eds.), New Developments in Zeolite Science and Technology, Elsevier, Amsterdam, 1986, p. 725. 39. G. Larson and G. L. Hailer, Catal. Lett., 3 (1989) 103. 40. P. W. Tamm, D. H. Mohr, and C. R. Wilson, in J. W. Ward (ed.), Catalysis 1987, Elsevier, Amsterdam, 1988, p. 335. 41. J. R. Bernard and J. Nury, U.S. Patent 4,104,320 to Elf France (1978). 42. R. J. Davis and E. G. Derouane, Nature (London), 349 (1991) 313. 43. E. Mielczarski, S. B. Hong, R. J. Davis, and M. E. Davis, J. Catal., 134 (1992) 359. 44. C. Besoukhanova, J. Guidot, D. Barthomeuf, M. Breysse, and J. R. Bernard, J. Chem. Soc. Faraday Trans., 77 (1981) 1595. 45. W.-H. Han, A. B. Kooh, and R. F. Hicks, Catal. Lett., 18 (1993) 193. 46. W.-H. Han, A. B. Kooh, and R. F. Hicks, Catal. Lett., 18 (1993) 219. 47. J. T. Miller, F. S. Modica, B. L. Meyers, and D. C. Koningsberger, Prepr., Div. Petrol. Chem., Am. Chem. Soc., 38 (1993) 825. 48. S. J. Tauster and J. J. Steger, J. Catal., 125 (1990) 387. 49. G. Larson and G. L. Hailer, Catal. Today, 15 (1992) 431. 50. C. Besoukhanova, J. Guidot, D. Barthomeuf, M. Breysse, and J. R. Bernard, J. Chem. Soc., Faraday. Trans., 77 (1981) 1595. 51. B. L. Mojet, M. J. Kappers, J. C. Muijsers, J. W. Niemantsverdriet, J. T. Miller, F. S. Modica, and D. C. Koningsberger, in J. Weitkamp, H. G. Karge, H. Pfeifer and W. H~lderich (eds.), Studies in Surface Science and Catalysis, Vol. 84B, Zeolites and Related Microporous Materials, State of the Art 1994, Elsevier, Amsterdam, 1994, p. 909.

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Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All fights reserved.

65

CATALYTIC DEACTIVATION D. L. Trimm

University of New South Wales, School of Chemical Engineering &Industrial Chemistry, Sydney NSW 2052 Australia Abstract

The hydroprocessing of heavy oils is associated with extensive catalyst deactivation mainly as a result of the deposition of coke and of metal sulphides. Deactivation is fast when the catalyst is first brought on line: subsequent loss of activity is much slower. The importance of various effects occurring during deactivation has been re-assessed. Initial deactivation is almost certainly due to adsorption of asphaltenes on acidic sites on the catalyst. Analysis of the catalyst shows that the initial deposits block smaller pores to cause up to a 50% loss of surface area. Subsequent deactivation is slower, and is associated with the deposition of relatively larger amounts of metal sulphides. The total amount of deposit is more than enough to block the pore volume of the original catalyst. Molybdenum sulphide is found to migrate through such deposits to the external surface and it is suggested that interactions between nickel, vanadium and molybdenum sulphides results in the formation of a deposit with significant catalytic activity. This activity would be expected slowly to decrease as further metal deposits block any porosity in the deposit itself. Introduction

The products obtained by refining and upgrading heavy crude oils are largely determined by market demand. The consumption of transport fuels is always such as to focus attention on the conversion of lighter [1] or heavier [2] feed stocks to appropriate distillation cuts. The processing of heavy residual oils into lighter oils is now widely practised, although the methods used may vary from refinery to refinery [3,4]. The specification of all petroleum products is dictated by the application and by environmental constraints. Increasing attention is focused on the latter, with- for example lower and lower limits on sulphur in the fuel oils being the norm. Essentially sulphur-free diesel fuel should be in demand in the near future, with subsequent legislation possibly being aimed at aromatics in the fuel. All of these factors have led to increasing demands on the refiner which, in turn, are reflected in more efficient petroleum processing. Higher demands are placed on processes designed to convert heavy residual oils, and the introduction of secondary processes designed mainly to meet environmental specifications is becoming a necessary expense. It would be desirable to produce on-specification fuels from the initial processing, but this is becoming increasingly difficult.

66 Various industrial processes have been developed to convert heavy crude oils into transport fuels [3,4]. Most of those in use are based on residual cracking or on hydroprocessing over cobalt-molybdenum, nickel-molybdenum or nickel-tungsten based catalysts [3]. Given the nature of the feed and the severity of the processing, it is not surprising that catalyst deactivation is a major problem. Although resid cracking is becoming increasingly popular, the bulk of heavy oil processing still involves hydrotreating to various degrees of severity As a result, attention is focused here mainly on an appraisal of catalyst deactivation during heavy end hydroprocessing.

Catalytic Hydroprocessing The nature of crude oils depends on their source. Initial separation into components is carried out by atmospheric and vacuum distillation. Heavy ends are particular boiling point cuts, which can include atmospheric gas oil (250-350~ atmospheric residues (350~ vacuum gas oil (350-550~ and vacuum residues (550~ The descriptions are based on boiling points and, within a particular distillation cut, various chemical species can be identified. These include saturated and unsaturated hydrocarbons, aromatic and polyaromatic hydrocarbons and inorganic atoms such as V, Ni, and S, which are associated with large organic molecules [5]. As a result of this complexity, the composition of the boiling cuts is otten described in terms of their content of oils, resins and asphaltenes [6,7,8], the relative amounts of which vary depending on the cut and the source of the crude [6] Of these species, asphaltenes are particularly important in the present context since they are known to be associated with heavy coke formation [7,8]. It is essential to recognise that asphaltenes are defined in terms of the fraction of the oil that precipitates where mixed with a lighter hydrocarbon such as pentane or hexane [7,9]. As a result, solubility is an important parameter, not least in the context of heavy oil processing. Thus for example, Komastsu et al [10] have produced a diagram that relates amounts of saturates, aromatics and asphaltenes with a tendency to form sludges (figure 1). The asphaltenes are only kept in solution by the solubilising power - mainly of aromatics. If hydroprocessing removes aromatics, asphaltenes will deposit. Dautzenberg and de Deken [ 11] have suggested that such processes could be important in hydroprocessing. They viewed the overall process as a competition between thermal and catalytic reactions. Over ca 430~ thermal reactions were suggested to predominate and the removal/inter-conversion of oils, resins and asphaltenes proceeded at about the same rate. At somewhat lower temperatures, preferential removal of oils and resins was suggested via a catalytic reaction. The consequent removal ofasphaltene solvent then led to precipitation of the asphaltenes and to coking. Similar arguments can be used to explain the results of Mochida et al [12] and ofNakata et al [13]. Part of the difficulty of understanding the catalytic hydroprocessing reaction mechanism arises from the nature of the asphaltenes [7,8,9]. It was believed that asphaltenes consisted of "sheets" of aromatic molecules, held together to form large entities [14]. However, recent work by Speight [8] has thrown doubt on these ideas with a much more open structure being suggested. Over 200 samples of asphaltenes were examined to find single, double and triple ring polynuclear aromatics together with thiophenic sulphur. Nitrogen and oxygen containing

67

60 50 SATURATES Figure 1. Stability diagram for sludge formation from asphaltenes/aromatics/saturates mixtures (adopted from ref. 10).

species typical of naturally occurring products were detected, and were suggested to be involved in intermolecular association. Whatever the composition of the component molecules, association into larger species is certain. Variations in molecular weight between ca 3000 [15] and 150,000 [16] have been reported, as have asphaltene sizes of 60-90A [16] - again considered to be an over-estimate [7,8]. In the context ofhydroprocessing it is obviously the size of the entity than actually reacts on the catalyst rather than the absolute size of the molecule or conglomerate that is important. Resins and asphaltenes are also of interest as a result of the fact that most of the heteroatoms reside in these fractions [7,16]. Beaton and Bertolacini [5] report thatabout equal amounts of V and Ni reside in the resins as in the asphaltenes, while sulphur is concentrated in resins. The chemical associations between the heteroatoms and organic molecules are far from certain, although general descriptions in terms of porphyrin like structures have been advanced [5]. Sulphur appears to be easily accessible and thiophenic sulphur has been identified [8]. Vanadium is more accessible than nickel. Catalytic hydroprocessing is designed to reduce the molecular size of molecules via hydrocracking (HC) and to remove unwanted heteroatoms by hydrodesulphurisation (HDS), hydrodenitrogenation (HDN) and hydrodemetallation (HDM).

68

;~ 100 >.F> ItO

<

Iii

> m

I< _1 ILl

,',"

0

TIME

ON LINE

Figure 2. Typical catalyst deactivation curve.

HC

R- R'+ H2 = RH + R'H

HDS

R S H + H2 = R H + H2S

HDN

RN + 2 H2= RH + NH3

HDM

R M + H2 = R H + H M X

Sulphur and nitrogen are removed as gases. Metals are deposited on the surface of the catalysts. This may be as a result of decreasing solubility of asphaltenes, as discussed above, or of chemical attack on the organic species associated with the heteroatom. The production of a molecule in which the inorganic species forms a larger fraction also leads to reduced solubility. Either associated with the chemical attack or subsequent to the initial attack, metals are converted to metal sulphides [ 17,18]. The operating conditions for catalytic hydrotreating depend on the particular heavy end and the severity of the transformation required. Temperatures of the order of 375-430~ and pressure in the range 500-2000 psig are not uncommon. Recycle of hydrogen is a normal feature of industrial operations. Given the tendency of heavy oils to coke and the deposition of metals on catalysts, it is not surprising that catalysts deactivate. A typical deactivation curve is shown in figure 2, the initial loss in activity occurring over only a short time. Subsequent slow deactivation then

69 occurs until final catastrophic deactivation. It is the processes associated with catalyst deactivation that form the main focus of this paper. Initial deactivation:

There is little doubt that the rapid initial deactivation is associated with heavy deposition of coke and the resulting loss of catalyst surface area (figure 3) [ 18,19,20]. About 50% of the surface area of a typical hydroprocessing catalyst can be lost [19,21], largely as a result of blockage of pores of small diameter [ 19, 21 ]. Macroporous volume is not greatly affected: 2-5 nm radius pore volume is greatly reduced [22]. The initial heavy deposits of coke almost certainly originate mainly from asphaltenes in the feed. Preferential adsorption of asphaltene fractions have been observed on cobalt and molybdenum oxides [23] and sulphides [24] as well as on sulphided nickel-molybdenum catalysts [25]. Adsorption seems to involve nitrogen atoms in the asphaltene linked to acidic sites on the catalysts. Basic nitrogen containing groups have been identified by Speight [8], and the initial catalyst deposits are known to be rich in nitrogen [19, 26]. The nitrogen bonding in the asphaltene is strong [7,19,20] and adsorption of the whole asphaltene entity via the nitrogen atom(s) is probable. The importance of catalyst acidity is emphasised by comparisons between initial coke formation on cobalt molybdate based catalysts and on sodium molybdate based systems [24]. In the latter case, the acidity is low, the steady state specific activity is not too dissimilar from cobalt molybdate and the initial deactivation is very much less. In similar studies, the extent of formation of coke deposits during the initial stages of reaction has been found to decrease with the acidity of a range of catalyst [28]. Once deposited, the coke was found to reduce the activity of catalysts for acid-catalysed reactions such as isomerisation [24] or hydrocracking [29]. The inference is clear- initial coking involves acidic sites on the catalyst which are deactivated by coke deposition. The initial coking causes several other undesirable effects. The loss of surface area resulting from filling smaller pore by coke certainly causes loss in activity. At the same time, larger pores in the catalyst become coated with adsorbed asphaltene, at least to the detriment of acid-catalysed coke formation. Such reactions appear not to include hydrodesulphurisation, which has been found not to be affected by initial coke deposition [18]. This presumably reflects the ease of removal of sulphur from the feedstock. The initial coking also has some interesting indicators for metals distribution in the catalyst, as can be illustrated using data reported by Tamm et al [20]. They found that nickel distribution in pellets at the inlet and the outlet of the reactor passed through a maximum across the fractional radius of the pellet. Near the inlet, the maximum deposition was some distance inside the pellet, as opposed to vanadium which was deposited predominantly at or near the external surface. Similar observations have been observed by others [21, 22]. Beaton and Bertolacini [5] have suggested that ca 50% of the Ni in a heavy oil resides in the asphaltene fraction. Given that asphaltenes absorb preferentially during the initial stages of deactivation, the use of a heavy Arabian residue containing 25 ppm Ni [20] would infer that

70

300

30

200

20 =

I

I v

r...)

100

N

0

10

0 TIMEONSTREAM

Figure 3. Surface area changes and coke formation during the initial stages of catalyst deactivation.

12.5 ppm Ni is deposited with the coke. lie initial deposit involves ca 25 wt% coke on the catalyst, from which it can be calculated that only ca 3 ppm Ni/gm catalyst can be expected. Comparison with the overall weight of nickel deposited during a run (0.02- 0.003 g cm "3 catalyst [20]) shows that metals deposition during the initial stages of deactivation is of little importance. There is some evidence of more extensive metal deposition during these initial stages of reaction [31 ] but, even so, nickels deposition cannot affect initial deactivation to any large extent. The same calculation may be made for the deposition of vanadium. Although the amounts deposited are significantly higher (ca 8 x Ni [19,20]), it is clear that most of the deposition of metals occurs after the initial deactivation period. Nonetheless, initial deactivation can be expected to affect subsequent deposition - largely as a result of the blockage of small pores and the deposition of asphaltenes on the surface of the large pores. The available surface area drops (figure 3) and the reactivity of the catalyst surface changes. Steady state deactivation Once the initial stages of deactivation are past, slow deactivation occurs over longer periods of time (figure 2). The amount of coke on the Catalyst appears to reduce, but this is more a function of the fact that coke deposition is reported per grain of catalyst. As metals

71 deposit on the solid, densification takes place and the percentage (weight for weight) coke drops. During steady state deactivation, the smaller pores are blocked (initial coke deposition) and further adsorption on acidic sites is unlikely. As a result, the accumulation of coke and metal salts must be considered relative to the catalyst surface area in the larger pores. It is known, however, that metals deposit as very long crystallises (some tens ofnanometers in length) originating from fixed nucleation sites [32]. As a result, consideration of metal deposits even on large pore surface area may not be accurate. It is also clear that the removal of vanadium and sulphur is much easier than the removal of nickel and nitrogen. Deposition profiles across catalyst pellets show that most vanadium is deposited near the pellet exterior, while the maximum for nickel lies some distance inside the pellet [ 19, 20]. Turning first to the deposition of nickel, the most commonly accepted mechanism is that a two stage process is involved. Using nickel etioporphyrin as an example, Wei et al have suggested a consecutive reaction involving a partly hydrogenated intermediate [33, 34, 35]. Ni R + H2 -~T-Ni RH + H2 -~ Ni deposit The difference in the deposition patterns observed between the inlet and the outlet of the reactor [20,36] can then be described in terms of conversion of the porphyrin to the hydrogenated species and then to the deposit. This allows the molecules to penetrate further into the pellet before hydrogenation and deposition occurs. Near the end of the reactor, the partly hydrogenated intermediate is more common in solution and deposition occurs closer to the edge of the pellet. Further support for the concept of a sequential reaction leading to metals deposition comes from an analysis of the deposition profiles of vanadium [ 17]. A kinetic analysis of a two step process for V removal was developed, in which the first step was suggested to involve hydrogenation of the V containing porphyrin, while the second produces deposits of vanadium sulphide on the surface. Increased levels of H2S in the gas stream inhibited HDS of the material but increased HDV. For vanadium, the second reaction leading to metal deposits, with levels of H2S expected in an industrial reactor, was found to be some 85 x faster than the first hydrogenation reaction [17]. Under these conditions, it is not surprising that the bulk of the deposition of vanadium occurs near the pellet surface. One further difference exists between HDS and HDM. Bridge [37] has shown, very clearly, that HDS is not limited by diffusion while HDM is. Using a nickel-molybdate based catalyst with a unimodal microporous size distribution, the demetalation of Arabian heavy atmospheric residuum was found to be affected by catalyst particle size, while HDS was not. As the diameter of the pore was decreased, the maximum in the metals deposition profile moved closer to the external surface of the pellet, again indicating diffusional limitations for HDM. Similar observations were reported by Toulhoat and Plumail [38], who found HDV and hydrodeasphaltene reactions to pass through a maximum as the pore size increases. Up to ca 20 nm pore diameter, mass transfer appeared to limit access to the pore system. Above this value, the effective surface area dropped away as a result of increasing pore diameter to give a reduction in the overall rate of HDM.

72

12[ z

Reactor inlet

/

08.j ~'

z

x

~~eactor d outlet

0.4......

o

0

Figure 4. ref. 37).

4'

0.2

o ~

~

~

. . . . . . . . .

0./+ 0.6 0.8 RELATIVETIMEONLINE

1.0

Vanadium deposits as a function of reactor length and time on line (adopted from

Despite the difficulties of determining the actual nature of the asphaltenes, Sughrue et al. [39] have used size exclusion chromatography to determine molecular sizes of molecules associated with vanadium. As would be expected, the size varies from feed to feed, but - in general - a small amount of vanadium is contained in molecules of ca 3 nm diameter, while most is contained in molecules with a Gaussian size distribution centred on 8-10 nm diameter species. Given the relative diameters of the molecules (ca 8-10 nm) and the pores (ca 20 nm) it is not surprising that vanadium does not penetrate far into the pellets. Assuming that the vanadium deposits consist mainly of vanadium sulphides, it is then possible to extend the arguments of Tamm et al [20] to calculate possible effects of deposition of coke and metals on pore closure. Accepting that 200 m2g"1 is a typical surface area for hydrotreating catalysts, it is known that ca 50% of the surface area is lost during the initial deactivation [ 19-21]. It is also known that V deposition occurs in the outermost part of the pellet [20,31,37,38]. As a result, the amount of V needed to provide a monolayer over the surface area available can be calculated to be between 0.034 and 0.079g cm "3 Vanadium deposits of up to 0.9 g cm 3 have been obtained from Arabian heavy atmospheric residue (figure 4 [20]), inferring ca 10-25 monolayers on the surface. Based on the crystallite dimensions of vanadium sulphide, this corresponds to a deposit some 3-8 nm in depth. Given that the deposit forms, together with coke, on both sides of a pore of diameter ca 20 nm, it is not surprising that the pores are closed off and that the catalyst loses almost all activity [20]. This observation raises some interesting questions with respect to steady state deactivation. The fact that so much vanadium is deposited, alongside similar deposits of nickel and coke infers that the original catalytic surface is not accessible throughout the run. Hence,

73 subsequent catalytic activity would appear to be due either to a non-catalytic reaction or to a reaction catalysed by the deposit itself. Both suggestions seem possible, but a catalytic reaction would seem to be more important. Sughrue et al [39] have shown the existence of that some thermal reactions leading to vanadium deposition is essential to explain the results observed: However these effects are, at most, responsible for only ca 5-10% of the total deposit and, as a result, must be regarded as of minor importance below about 430~ [ 11]. A catalytic reaction involving the deposit as the catalyst seems more likely. Hydrotreating reactions have been suggested to be favoured by Ni and V sulphides, albeit at a lower rate of reaction [35]. Welters et al [40] have shown that nickel sulphides in or on a zeolite are active hydroprocessing catalysts and vanadium sulphides have also been found to promote hydrotreating [27], as least to some extent. Much more interesting, however, is the observation by Simpson [36] that molybdenum is always observed on the surface of used catalysts, despite the presence of up to 85 wt% V and 20 wt% Ni on the catalyst. Evidently molybdenum sulphide can migrate readily through the deposit to the surface, a finding confirmed by at least two other studies [40, 41 ]. As a result, there may well be combinations of sulphided nickel, vanadium and molybdenum at the surface of the catalyst and available to promote further hydroprocessing. The amounts of molybdenum at the surface are only ca 10% of the original catalyst content [36], and catalytic activity presumably reflects combinations of vanadium, nickel and molybdenum containing sulphides. The influence of the original support would be small, but coke deposited during the initial deactivation and subsequently ordered [27] could well provide a suitable support [43]. Pore blocking by coke or by vanadium sulphides also restricts access of nickel salts. It has been suggested., that nickel compounds may penetrate further into the pellet as a result, at least in part, of their smaller size [20, 31]. This seems inherently unlikely in that a narrow boiling point cut can be expected to give a general molecular size around the same average. What seems more probable is that the low reactivity of nickel compounds [35] allows greater penetration both of the pellet and of deposits on the pellet. If this is the case, then nickel deposition, will also be affected by pore blocking by vanadium. As a result, subsequent removal of nickel should occur near the deposit exterior, as has been observed towards the end of the bed [35]. The resulting improvement in the catalytic activity of the deposit (as compared to the activity of vanadium sulphides) would be significant in view of possible interactions between nickel and molybdenum sulphides, and the role of nickel sulphide in facilitating the distribution of molybdenum sulphide across the surface [44]. If this explanation is correct, then one also has to consider the subsequent slow decline in activity (figure 2). One possible explanation involves the deposition of more vanadium, slowly decreasing the amount of molybdenum at the surface. More probably, the slow decline reflects slow loss of surface area by infilling of pores in the deposit, rather than in the catalyst. The initial deposition of carbon should create a porous deposit on the catalyst via the deposition of large crystallites of coke and of metal sulphides [35]. Slow infilling of these pores via metals deposition should lead to slow loss of surface area and of activity.

74 It is accepted that these concepts are speculative, and that examination of the catalytic activity of deposits from hydrotreating catalysts is required to prove the suggestions. Nonetheless, the proposals do explain many of the reported observations. Final catastrophic deactivation

It is well established that eventually the catalytic activity dies away rapidly (figure 2). The most probably explanation of this involves the migration of the wave of heavy metal sulphide deposits through the bed. Examination of spent catalyst removed from a trickle bed reactor shows clearly that the top layer of the bed is very heavily contaminated and that the contamination decreases down the bed. If the reaction is allowed to proceed to the point where heavy deposits spread throughout the bed, then catastrophic deactivation can be expected. Conclusions

Reappraisal of catalyst deactivation during heavy end hydrotreating suggests that initial deactivation involves the adsorption of asphaltenes on acidic sites in the catalyst. This results in the blocking of smaller pores and loss of ca 50% of surface area. Subsequent deposition of vanadium, nickel and coke is shown to be considerably in excess of that needed to fill the catalyst pore volume. Migration of molybdenum to the used catalyst surface is established and is suggested to result in reasonable catalytic activity via the formation of sulphides containing nickel, vanadium and molybdenum. Subsequent deactivation would result from infilling pores in the deposit, with final catastrophic deactivation resulting from the wave of heavy deposits of metal sulphides approaching the end of the catalyst bed. References

1. W.M.H. van Wechem & M.M.G Sendeu in Natural Gas Conversion II, H.E. Curry-Hyde & R.F. Howe (Eds.), Elsevier Amsterdam (1994) 43 2. H. Qabazard, R. Adarme & B.L. Crynes, Catalysts in Petroleum Refining 1989: Studies in Surface Science & Catalysis vol. 53, Elsevier Amsterdam (1989) 3. A. El Hariry, Ibid p 129 4. J.H. Gary & G.E. Handwerk, Petroleum Refining Technology & Economics 2nd Ed. Marcel Dekker NY (1984). 5. W.J. Beaton & R.J. Bertolacini, Catal. Revs. Sci. Eng., 31 (1991) 281. 6. M.M. Boduszynski, Preprints Div. of Petroleum Chem. ACS, 20 (1985) 1376. 7. J.G. Speight, The desulphurisation of heavy oils and residue, Dekker NY (1981). 8. J.G. Speight, Preprints Div. of Petroleum Chem. ACS, 34 (1989) 321. 9. J.G. Speight & R.J. Paneirov, Preprints Div. of Petroleum Chem., ACS, 28 (1983) 1319. 10. S. Komatsu, Y. Hori & S. Shimizu, Hydrocarbon Process, May (1985) 42. 11. F.M. Dautzenberg & J.C. de Deken, Preprints Div. of Petroleum Chem., ACS, 30 (1985)8.

75 12. I. Mochida, X.Z. Zhao, S. Sakanishi, H. Yamamoro, H. Takashuira & S. Vemura, Ind. Eng. Chem. Res., 28 (1989) 418. 13. S. Nakata, S. Shimizu, S. Asaoka, Y. Shiroto & Y. Fukui, Preprints Div. of Petroleum Chem. ACS, 32 (1987) 477. 14. J.P. Dieke & T.F. Yen, Anal Chem. 39 (1967) 1847. 15. J.G. Speight, Preprints Div. of Petroleum Chem. ACS 32 (1987) 413 16. J.G. Speight & S.E. Moschopedis in 'Chemistry of Asphaltenes", J.W. Bunger & N.C. Li (Eds), Adv. Chem. Ser. 195 (1981) 1. 17. J. Bartholdy & P.N. Hannerup in 'Catalyst Deactivation 1991", C.H. Bartholomew and J.B. Butt(Eds.), Elsevier Science Publishers Amsterdam (1991) 273. 18. P. Wirrel, P. Zeuthen & A.C. Jacobsen, Ibid p257. 19. D.S. Thakur & M.G. Thomas, Ind. Eng. Chem. Prod. Res. Div. 23 (1984) 349. 20. P.W. Tamm, H.F. Harsberger & A.H. Bridge, Ind. Eng. Chem. Proc. Des & Dev. 20 (1981)262. 21. M. Ternan & J.F. Kritz in 'Catalyst Deactivation (Eds.), Elsevier Amsterdam (1980) 283.

1980", B. Delmon & G.F. Froment

22. R.J. Bertolacini, L.C. Gutberiet, K.K. Robinson EPRI report AF - 1084 for Project 408-1 June 1979. 23. J.J. Sault, Preprints Div. of Petroleum Chem. ACS 23 (1978) 1462. 24. F. Melo Faus, P. Grange & B. Delmon, App. Catal. 11 (1984) 281. 25. D.R. Milburn, B.D. Adkins & B.H. Davis, Preprints Fuel Chem. Div. ACS 33 (1988) 380. 26. G. Furimsky, Ind. Eng. Chem. Prod. Res. Div. 17 (1978)329. 27. M. Absi Halabi, A. Stanislaus & D.L. Trimm, App. Catal. 72 (1991) 193. 28. J.F. Kritz & M. Ternan in 'Catalysis on the Energy Scene", S. Kaliaguine & A. Mahali (Eds.), Elsevier, Amsterdam (1984) 545. 29. A. Nishijima, L.T. Shiurada, Y. Yoshimuro, T. Sato & N. Matsubayashi, in 'Catalyst Deactivation 1987", B. Delmon & G. Froment (Eds.), Elsevier, Amsterdam (1987) 39. 30. C.L. Thomas, 'Catalytic Processes & Proven Catalysts", Academic NY (1970) 31. D.S. Thakur & M.G. Thomas, App. Catal 6 (1983) 283. 32. H. Toulhoat, J.C. Plumail & G. Martino, Preprints Div. of Petrol. Chem. ACS 85 Feb. (1985) 33. R. Agrawal & J. Wei, Ind. Eng. Chem. Process, Des. Dev. 23 (1984) 505, 515. 34. R.A. Ware & J. Wei, J. Catal. 93 (1985) 122, 135. 35. J. Wei in 't2atalysts Deactivation 1991", C.H. Bartholomew & J.B. Butt (Eds.), Elsevier Science Publ., Amsterdam (1991) 333.

76 36. H.D. Simpson, ibid p 265 37. A.G. Bridge in 'Catalysts in Petroleum Refining 1989", D.U Trimm et al (Eds.), Elsevier Science Publ., Amsterdam (1990) 363. 38. H. Toulhoat & J.C. Plumail in 'Catalysis in Petroleum Refining 1989", D.L. Trimm et al (Eds.), Elsevier Science Publ., Amsterdam (1990) 463. 39. E.L. Sughrue, R. Adarme, M.M. Johnson, C.J. Lord & N.D. Phillips in 'Catalyst Deactivation 1991", C.H. Bartholomew & J.B. Butt (Eds.), (1991) 281. 40. W.J.J. Welters, G. Vorbeck, H.W. Zandberger, J.W. de Haan, V.H.J. de Beer & R.A. Van Santen, J. Catal., 150 (1994) 155 41. T.H. Fleisch, B.L. Meyers, J.B. Hall & G.L. Ott, J. Catal., 86 (1984) 147. 42. A. Stanislaus & K. A1-Dolama, J. Catal., 101 (1986) 536. 43. O. Weisser & L. Landa, 'Sulphide Catalysts, Their Properties & Applications", Pergamon, Oxford (1973). 44. U.S. Ozkan, L. Zhang, S. Ni & E. Mociezuna, J. Catal. 148 (1994) 181.

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All fights reserved.

77

PREPARATION AND CATALYSIS OF HIGHLY DISPERSED METAL SULFIDE CATALYSTS F O R HYDRODESULFURIZATION Yasuaki Okamoto

Department of Chemical Engineering, Faculty of Engineering Science, Osaka University, Toyonaka, Osaka 560, Japan Abstract

Highly dispersed Mo, Co and Co-Mo composite sulfide catalysts were prepared by using Mo and Co carbonyls anchored on A1203 or zeolite. The dispersion, structure and chemical states were examined by using EXAFS and XPS techniques. It is demonstrated that highly dispersed Mo sulfides lead to the formation of highly dispersed and highly active Co-Mo sulfide catalysts for the HDS reaction of thiophene. It is suggestedthat the promotional effects of Co are generated by direct chemical bondings between the Mo and Co sulfide phases. I. Introduction

Hydrotreating of petroleum feedstocks have become more and more crucial not only for protecting environments but also for efficient utilization of natural resources. Extensive researches have been conducted to develop highly active catalysts for hydrotreating, in particular for hydrodesulfurization (HDS) [1-3]. Industrial HDS catalysts are usually based on Co(Ni)-Mo sulfides supported on alumina or silica-alumina. In order to develop more active HDS catalysts, many attempts have already been carried out up to now. An addition of further promoters such as P, B or F is reported to improve catalyst performances [4,5]. Fine controls of the pore size distributions of the catalyst supports are very successful. Some improvements of the catalyst preparation methods have been found to be considerably effective. In particular, the addition of chelating compounds such as nitrilotriacetic acid [6] into the impregnation solution is a promising technique. Developments of new kinds of support materials such as zeolite, [7] zirconia, titania and their mixed oxides [8] and novel catalytically active metals such as Rh and Ru [9] or new kinds of materials such as Mo nitride and carbide [10] are strong candidates for new types of liDS catalysts with high performances. One of the possible strategies for improving the catalyst performances may be significant improvements of the dispersion of catalytically active Co(Ni) and Mo sulfide species over the supports. The origin of catalytic synergies between Co(Ni) and Mo sulfides is still controversial now [3,11]. At present, mainly two synergy models are under arguments. A surface model of alumina supported Co-Mo binary catalysts is presented in Fig.1 [2]. Mo is present as microcrystalline MoS2 and Co as CogSs, Co 2+ in the alumina phase and Co species interacting with the MoS2 phase. The active species are proposed to be a so called CoMoS phase in which atomically dispersed Co sulfides are believed to be anchored on the edge surface of MoS2

78

'//

/

I

s co

\\

Mo

Figure 1. Structure model of a sulfided Co-Mo/A1203 catalyst.

L.

~=36A S

|

Distorted tetrahedral Co

e

Rapid prismatic Co

0

Mo

9 Figure 2. Top and side view of a squarepyramidal CoMoS structure.

Figure 3. Model of the Co-Mo active phase.

particles [2]. The local structures of the Co and Mo atoms in Fig. 2 proposed on the basis of XAFS studies are consistent with the formation of the CoMoS phase for carbon supported CoMo catalysts [12,13]. A 59Co solid NMR study of Co-Mo sulfide catalysts by Ledoux et al. [14], combined with high resolution electron microscopic observations [15] suggested a structure in Fig. 3 for the CoMoS phase. Promotional effects of Co or Ni are also construed in terms of a "remote control" theory proposed by Delmon and co-workers [3]. On the basis of the findings that catalytic synergies are generated even for physical mixtures of supported component sulfides, it is claimed in this theory that the catalytic,, activity of Mo sulfides is enhanced by spillover hydrogen originally generated on highly dispersed Co or Ni promoter sulfides in the proximity.

79 Table 1. Interaction modes between molybdenum and alumina deduced from the behavior of the surface hydroxyl groups as function of molybdenum concentration. Mo conc. 1013 cm2 <10 10-30 25 30 >30

interaction modes a predominant formation of Motet species simultaneous formations of both Motet and Mooet species; formation of multilayered molybdate species limit of Motet formation; maximum amount of Motet species, 17 x 1013 Mo atoms cm2 limit of a high dispersion of molybdenum formations of dispersed and bulklike molybdate species such as MoO3 and

A12(MoO4)3 a Motet and Moo~t, molybdate species in tetrahedral and octahedral coordinations, respectively.

Figure 4 schematically depicts the CoMoS and remote control models. It is rational to conjecture in both models that an increase in the dispersion of Mo sulfide species leads to an increase in the concentration of the CoMoS phase or in the number of the active sites on the Mo sulfide edge surface promoted by spiUover hydrogen, resulting in enhanced HDS activities. However, the knowledge about the preparation, structure and catalytic behaviors of highly dispersed Mo or Co sulfide species is lacking now despite indispensable information for the design of highly active HDS catalysts. It seems difficult to prepare highly dispersed Mo sulfides on AlzO3 by using conventional ammonium heptamolybdate (AHM) as a starting material because of specific Mo oxide-Al/O3 surface interactions [16-17] summarized in Table 1. Decreasing the Mo content to aim at a higher Mo dispersion enhances the interactions between the Mo oxide phases and the A1203 surface, which produces Mo species less susceptible to reduction and sulfidation, and thus resulting in a decrease of the fraction of catalytically active Mo species as corroborated by the XPS study of the sulfidation of MoO3/A1203 catalysts [19]. With the supports other than A1203, activated carbon is reported to provide highly dispersed metal sulfides [20]. Utilization of Mo complexes as precursors instead of AHM may be another possible route for the preparation of highly dispersed Mo sulfides. In the present study, molybdenum hexacarbonyl, Mo(CO)6, and cobalt carbonyls, Co(NO)(CO)3 and Coz(CO)8, were employed for the preparation of supported Mo and Co sulfides and composite Co-Mo sulfide catalysts.

MoSz

04

Co~_..S_,/~ + H2

MoS2

H2SO

/ /- / / / / / / / /

Co-Mo-S model

/~/-~l

Remote control model

Figure 4. Catalytic synergy generation in Co-Mo catalysts.

80 Table 2. Catalytic properties of sulfided Mo catalysts prepared from Mo(CO)6.

Zeolite

Mo-loading (wt%)

Ads-MoSx HY 9.4 LiY 12 NaY 12 KY 10 RbY 9.3 CsY 8.7 Imp-MoS2 NaY 4.8 TiO/ 6.1 A1203 6.1 SiO/ 6.1 a measured at 673K on a fixed bed flow reactor b proportion of cus-Mo sites

HDS activity a (%g-1 Mo-I)

(NO)2/Mo b

48 136 176 205 134 128

O.23 0.27 0.31 0.48 0.40 0.46

60 149 63 93

0.12 0.07 0.14 0.05

Vrinat et al. [21] examined the HDS activity for dibenzothiophene of the Mo sulfide catalysts prepared by using Mo(CO)6 encaged in zeolite. Relatively low HDS activities were obtained. This is considered to partly result from low dispersions of Mo sulfide species induced by complete decompositions of Mo(CO)6 at 600 K prior to sulfidation. In our previous study [22,23], Mo(CO)6 encaged in a series of alkali metal cation exchanged Y- or X-zeolites was directly sulfided at 673 K for the preparation of supported Mo sulfides. Table 2 shows the catalytic results of the Mo sulfides supported on zeolites for the HDS of thiophene. It was found that the Mo sulfide species thus prepared are highly active for the HDS as compared with the Mo sulfides prepared by a conventional impregnation method using AHM. The dispersion of the Mo sulfides was estimated on the basis of an NO adsorption capacity. In line with these observations, Laniecki and Zmierczak [24] showed that zeolite supported Mo sulfides synthesized in a similar way exhibit high activities for the HDS of thiophene and water-gas shift reactions. More recently, Vorbeck et al. [25] studied the distribution of Mo sulfides in sulfided Mo(CO)6/NaY using 129Xe-NMR techniques in combination with Xe adsorption and HREM/EDX measurements. Cobalt carbonyls were examined by Halbert et al. [26] to modify unsupported or AlzO3 supported Mo sulfide catalysts. Decoration of the Mo sulfide catalysts with Co2(CO)s, followed by sulfidation, was found to generate catalytic synergies between Co and Mo sulfides for HDS reactions. In the present study, the preparation and catalysis of highly dispersed Mo, Co and Co-Mo sulfide catalysts supported on AI/O3 or NaY zeolite were investigated to show the effects of the Mo sulfide dispersion on the catalyst performances and synergy generations. The mechanisms of the generation of the catalytic synergies between Co and Mo sulfides are discussed.

81

2. A!203 Supported Mo Sulfide Catalysts With A1203 supported Mo sulfide catalysts prepared from Mo(CO)6 (denoted as MoSx/AI203 here), it was revealed on the basis of EXAFS results that the dispersion of the Mo sulfide is very high as compared with that in conventional impregnation catalysts (MoSE/A1203). The loading amount of Mo on A1203 was limited to be 1.9 wt% Mo on a single adsorption procedure of Mo(CO)6 and subsequent sulfidation (MoSx/AI203(1.9)). MoSx/AI203(3.6) was prepared by introducing Mo(CO)6 onto MoSx/A1203(1.9), followed by a second sulfidation. The sulfidation was carried out at 373 K and subsequently at 673 K in a stream of H2S/H2 (0.1 V/V) throughout this study. XPS results indicated that Mo species are completely sulfided in MoSx/AI203 catalysts in spite of low Mo contents. The EXAFS results showed that the structure of the Mo sulfide species is a highly dispersed MoS2 like structure in MoS2/AI203(10) or MoSx/AI203(3.6). With MoSx/AI203(1.9), however, the formation of a Mo sulfido dimer species having $22 bridging ligands are suggested on the basis of a short MoMo bond distance (0.28 nm) and a Mo-Mo coordination number close to unity. The thermal stability of the Mo dimer complex was rather low and the structure of the surface complex was transformed into a MoS2 like one, accompanying a slight Mo agglomeration, on a prolonged sulfidation at 673 K, MoSx/AI203(1.9,2S). The Mo-Mo coordination numbers, N(Mo-Mo), were calculated to be 0.7, 2.2, 2.8 and 3.5 for MoSx/A1203(1.9), MoSx/AI203(1.9,2S), MoSx/AI203(3.6) and MoSJA1203(10), respectively. On the basis ofastructuralmodel of MoS2 [ 12], the numbers of Mo atoms constituting the sulfide clusters are estimated to be 2, 34, 4-5, and 7-8, respectively, for these catalysts. It is demonstrated that highly dispersed Mo sulfide species are prepared on A1203 using Mo(CO)6. The turnover frequencies per Mo atom ( T O F ) o f thiophene HDS and butadiene hydrogenation (HYD) are shown in Fig. 5 as a function ofN(Mo-Mo), a parameter of the dispersion of Mo sulfides. The TOF of the HDS is relatively invariant with the Mo sulfide dispersion when the dispersion is considerably high, indicating that the HDS activity of the edge sites of Mo sulfides is almost independent of the cluster size. This is consistent with the observations by Topsoe et al. [2]. On the other hand, the TOF of the HYD increases as the dispersion of the Mo sulfide increases. These findings may suggest that the active sites for the HDS and HYD reactions have different configurations and that the fraction of the active sites for the HYD greatly increases as the dispersion of the Mo sulfide increases. It is also plausible that the bridging $22 ligands in the Mo dimer species are responsible for the high HYD activity at the highest dispersion, since sulfido Mo dimer complexes have been reported to activate molecular hydrogen under mild reaction conditions [27].

3. Highly Dispersed Co-Mo/A!203 Catalysts Cobalt was added from a gas phase to the A1203 supported Mo sulfides by using CoffCO)8 or Co(NO)(CO)3, followed by a second sulfidation at 673 K. Figure 6 compares, as a function of the Co loading, the HDS activities of CoSx-MoSj A1203(10) prepared by the Co carbonyl addition into MoS2/AI203(10) and conventional Imp-Co-MoS2/Al2Offl0) catalyst prepared by a double impregnation method. The HDS activity Of CoSx- MoSJA1203 increases at a lower Co content, reaching a plateau activity on a further Co addition.

82

09

~ 4 T O

7O3 7 o 2

@

"~.

@

~'-.......

v t.l_

20 c O t'13 c'(D 03 O "O >,., I -10 "~ q03

7O

9

U..

o

6O N(Mo-Mo)

Figure 5. Dependencies of TOF of the HYD of butadiene (O, 473 K) and thiophene HDS (o, 673 K) upon the coordination number of the Mo-Mo bondings for Mo sulfides supported on A1203. Similar observations have been reported by Halbert et al. [26]. The maximum activity of Imp-Co-MoS2/Al203(10) exceeded the plateau activity of CoSx-MoS2/ A1203(10). This is probably due to an increased dispersion of Mo sulfides by the presence of Co [20]. The HDS activities of CoSx-MoSx/AI203 catalysts having 1.9 and 3.6 wt% Mo are presented in Fig. 7. In both catalyst systems, plateau activities are attained, as observed for CoSx-MoS2/AI203(10). The promotional ratio is defined here as the ratio of the activity of the Co-Mo composite catalyst to that of the host Mo sulfide catalyst. The promotional ratio is plotted in Fig. 8 against N(Mo-Mo) of the host Mo sulfides. It is evident that the promotional ratio increases as the dispersion of the Mo sulfide increases. It is demonstrated that highly dispersed Mo sulfides lead to the formation of highly dispersed and highly active Co-Mo sulfide catalysts for HDS. The optimum amount of Co defined as the smallest amount of Co required for the maximum or plateau HDS activity of the Co-Mo catalyst was found to increase as the dispersion of the Mo sulfide species. Figure 9 correlates the promotional ratio with the optimum amount of Co as expressed by the Co/Mo atomic ratio. A linear correlation in Fig. 9 may suggest that the Co species decorating the edge sites of the Mo sulfides are responsible for the HDS activity. The optimum Co/Mo atomic ratio reached unity at the highest dispersion of the Mo sulfide species. 4. Zeolite Supported Mo and Co Sulfide Catalysts Highly dispersed Mo sulfide species were found to be prepared by sulfiding Mo(CO)6 encaged in zeolite even at a high concentration of Mo (more than 10% Mo) [22-24]. The NO adsorption capacity suggested a significantly higher dispersion of the Mo sulfides in the zeolite systems than that in supported Mo sulfide catalysts prepared by a conventional impregnation method. The EXAFS results in Fig. 10 for MoSx/NaY also indicated a high dispersion of the Mo sulfide species (N(Mo-Mo) = ca. 1). In the case of MoSJNaY prepared from a MoO3/NaY impregnation catalyst, Mo-O bondings were found to appear even alter a sulfidation for 5 h at 673 K as well as Mo-S and Mo-Mo bondings, suggesting incomplete sulfidation of the Mo

83 z,O Imp Co - MoS2 / AI,_O3 "7

@

"7

,.~ 30

20

O

-

A W

@

_

Mo " 3.6wt%

E

O

E O

20 .~_

!

,...,

._

< 10

< r~

c~ 10c 1

CoSx / AhO3 I

1

0

t

( I

2

/

Mo " 1.9wt%

1

4

I

I

6

2

Co-Loading I wt%

Figure

1

I

4

Co-Loading / wt%

6.

HDS activities of CoSxand Imp-Co-MoS2/Al203 as a function of the Co content. The liDS activity of CoSx/A1203 prepared by using C02(CO)s is also shown for comparison.

MoS2/ml203

Figure

7.

HDS

activities of CoSx-

MoSx/A1203 catalysts having 1.9 and 3.6 wt% Mo against the Co loading.

10

10

~8

c~8 -.9 6

oi o 6

4

._~ 4

I-

O

E

E

O

O

~- 2 I

0

I

I

2

I

4 N(Mo-Mo)

Figure 8. Promotional ratio of the Cocatalyst for HDS as a function of N(Mo-Mo) of the host Mo sulfide species. O ; CoSx-MoSx(MoS2)/A1203 and

Mo/AI203

@ ; Imp-Co-MoS2/Al203

0

I

I

I

I

!

0,2

0,4

0,6

0,8

1,0

Co/Mo Atomic Ratio

Figure 9. Promotional ratio of the CoM o / A 1 2 0 3 catalyst for HDS as a function of the optimum Co content as expressed by the Co/Mo atomic ratio. O ; CoSx-MoSx(MoS2)/A1203 and @ ; Imp-Co-MoS2/A1203.

1,2

84

o

0

u_

:~

•

6

'

3

s

Distance / A

Figure 10. Fourier transforms (k3 X(k)) of the Mo K-edge EXAFS for MoSx/NaY and MoS2/NaY. oxides. This was corroborated by XPS measurements. Sulfiding the Mo oxide dimer species, (MOO3)2, fabricated in the supercage by mild oxidation of Mo(CO)6/NaY [28] was found by using EXAFS techniques to provide the Mo sulfide species which possess the identical structure and dispersion with those of the Mo sulfides prepared by the direct sulfidation of Mo(CO)6/NaY. The HDS activities of MoSx/NaY, MoS2/NaY and CoSx/NaY (prepared using CO(NO)(CO)3) are shown in Fig. 11 as a function of the number of the metal atoms per supercage (SC). With MoSx/NaY and CoSx/NaY, the HDS activities linearly increased up to the metal content of 2M (M = Mo or Co)/SC, suggesting the formation of uniform sulfide species in these catalysts. MoSx/NaY was found to show a much higher HDS activity than MoS2/NaY as reported previously [22,23]. It is remarkable that Co sulfide species in CoSx/NaY exhibit a considerably high HDS activity. It is demonstrated that highly dispersed Mo sulfide species are prepared even at a high Mo concentration (ca. 10 wt%)by using Mo(CO)6/NaY. The location of Mo sulfide species, inside or outside of the host zeolite, is always a difficult problem to determine. The high resolution electron microscopic observations for MoSx/NaY with 4 Mo/SC clearly demonstrated that the cage structure of the host zeolite is not destroyed and that no crystalline Mo sulfide species are observed outside of the zeolite. These results suggest that highly dispersed, thermally stabilized Mo sulfide species are prepared inside of the zeolite cages at high Mo loadings using Mo(CO)6 as a starting material.

85 0 120

200 -

loo

5

Co-loading / wt% 10 15 '

2Q1.5 '

t

16o-

1

_c

~. 80

z 12o~-///k.J

-60

--,~

o

t

i 0.5T

,.

......

0

1

2 3 4 5 Metal-aloms / supercage

~

4

6

Figure 11. HDS activities (623 K) of MoSx/NaY($), CoSx/NaY(O) and MoS2/ NaY (&) against the Mo or Co content as expressed by the atom number/supercage.

0 Co/Mo atomic ratio

Figure 12 HDS activity of CoSx-MoSx/ NaY (2 Mo/SC, sulfided for 1.5 h at 673K) as a function of the Co content. The activity ratio of HYD/HDS is also shown. I:! sulfided for 20h.

5. Zeolite Supported Co-Mo Sulfide Catalysts Zeolite supported Co-Mo composite catalysts were prepared by introducing Co(NO)(CO)3 or Mo(CO)6 into MoSx/NaY or CoSx/Nau respectively, followed by a subsequent sulfidation at 673 K. Figure 12 shows the HDS activity of the composite catalyst, CoSx-MoSx/NaY (Co was introduced after Mo, 2 Mo/SC) , as a function of the Co/Mo atomic ratio. It is revealed that the maximum activity is obtained around Co/Mo = 1. No activity decrease was observed even after prolonged sulfidation at 673 K. The HYD/HDS activity ratio decreased with increasing Co content and reached the ratio for CoSx/NaY at the composition where the maximum HDS activity is attained. In addition, the butene distribution in the butadiene HYD became identical with that of CoSx/NaY at Co/Mo = ca. 1. The HDS activity of MoSx-CoSx/NaY catalyst in which Mo was added to the pre-existing Co sulfide species was identical with that of CoSx-MoSx/NaY at the same composition. The HDS activity of MoSx/NaY, however, was remarkably decreased by the addition ofFe by using Fe(CO)5 (FeSx-MoSx/NaY). The k3-weighted Fourier transforms of the Mo K-edge EXAFS are shown in Fig. 13 for the zeolite supported Mo sulfide catalysts. With CoSx-MoSx/NaY, the Mo-Co bondings are obviously observed and the bond distance was calculated to be 0.282 nm. The bond length is close to that reported for CoMoS phases [ 13,20]. The Mo-Co bondings were also observed for MoSx-CoSx/NaY and CoSx-MoSx/NaY sulfided for 20 h. No Mo-Fe bondings were detected for FeSx-MoSx/NaY as shown in Fig. 13. The XPS binding energy of the Co2p3/2 level for C o S x - M o S x / N a Y was higher by 0.6 eV than that for C o S x / N a Y , suggesting the formation of Co-Mo binary sulfide species [29] in conformity with the EXAFS results in Fig. 13.

86

O3

v

L_ O

13: v ii

CoSx-MoSx/N,:qY

FeSx-MoSx/N~]Y

MoSx/NaY 0

1

2

3

4

5

6

Distance / A

Figure 13. k3-weighted Fourier transforms of the Mo K-edge EXAFS for Mo sulfide catalysts. Consequently, the above findings suggest that highly dispersed Mo sulfide species are decorated with Co and vice versa, possibly forming highly dispersed Co-Mo binary sulfides, in a limit Co2Mo2Sx cluster, at the maximum HDS activity and that the Co sites constitute an important part of the catalytic center for the butadiene hydrogenation and probably for the HDS of thiophene.

6. Generation of Catalytic Synergies Dicobalt octacarbonyl, Co2(CO)s, can be used to produce Co sulfides on the external surface of zeolite particles, since the diffusivity of Co2(CO)g into zeolite cavities is considerably small at low temperatures. Figure 14 schematically shows the model catalyst systems examined in the present study. In Fig. 14, for example, CoSx/MoSx/NaY denotes the catalyst in which Co sulfide was supported on MoSx/NaY using Co2(CO)8 dissolved in hexane, followed by an evacuation at room temperature and a subsequent sulfidation at 673 K. The XPS results for the catalysts were consistent with the distributions of Co and Mo sulfides envisaged in the models. The HDS activities of the composite catalysts possessing a composition of 2Co + 2Mo/SC are compared in Fig. 15 with the sum activities of the corresponding component CoSx/NaY and MoSx(MoS2)/NaY catalysts. Obviously, a catalytic synergy was observed only for CoSxMoSx/NaY system in which highly dispersed Co and Mo sulfides are combined, forming binary sulfides. Taking into account also the linear relation in Fig. 9 for CoSx-MoSx(MoSE)/AI203

87

H2S/H2

H2S/H2

H2S/H2

T

f

T

T

7

I

002(CO)8 Mo(CO)6 Co(NO)(CO)3 '

CoSx-MoSx/NaY

IH2S/H2

Mo(CO)6

CoSx/MoSx/NaY

Co(NO)(CO)3 /

CoSx-MoSdNaY

002(CO)8 !

CoSx/MoSdNaY

MoSx-CoSx/NaY

Figure 14. Schematic models for the Co-Mo binary sulfide catalysts supported on a NaY zeolite.

I

200 I..t3

o "!--

x

..o

>-. Z

6~ ]00 0

E > o <

~//////~ .

. . . . . . . .

.

.

.

.

.

-..-.:.~.-:.-.~ . . . . . . .

Co+Mo CoMo Co+MoCoMo Co+MoCoMo Co+Mo CoMo CoSx/MoSz/NaY CoSx/MoSx/NaY CoSx-MoSz/NaYCoSx-MoSx/NaY

Figure 15. HDS activities of the Co-Mo binary sulfide catalysts (2Mo + 2Co/SC) supported on NaY as compared with the sum activities of MoSx/NaY or MoS2/NaY (2Mo/SC) and CoSx/NaY (2Co/SC).

88 between the promotional ratio and the optimum amount of Co expressed as the Co/Mo atomic ratio, it is suggested that catalytic synergies between Co and Mo sulfides are generated only when Co and Mo sulfides are directly bonded to each other to form Co-S-Mo chemical bonds and that the Co sites play important roles in the HYD and HDS reactions. 7. Conclusions

In conclusion, it is demonstrated that highly dispersed Mo sulfides lead to the formation of highly dispersed and highly active Co-Mo sulfide catalysts for the HDS reactions of thiophene. The promotional effects of Co are considered to be generated by a direct decoration of the Mo sulfide phase with Co under the present preparation and reaction conditions. In addition to the above scientific findings, the industrial implications of the present study reside in the demonstration that one of the strategies to prepare highly active hydrotreating catalysts is the preparation of highly dispersed Mo sulfide phases decorated with Co. Further investigations are required to synthesize and stabilize highly dispersed Co-Mo sulfide species by employing preparation procedures easily accessible in chemical industries. Acknowledgments

We are grateful to Prof. N. Nomura and staff of the Photon Factory, National Laboratory for High Energy Physics, for assistance in measuring the XAFS spectra (Proposal Nos. 89-138, 90-150 and 93G-163). We also express our gratitude to Prof. O. Terasaki in Tohoku University for the HREM measurements of the catalyst. References 1. P.Grange, Catal. Rev.-Sci. Eng., 21 (1980) 34. 2. H.Topsoe, B.S.Clausen, N.-Y.Topsoe and E.Pederson, Ind.Eng.Chem. Fundam., 25 (1986) 25. 3. B.Delmon, Catal.Lett., 22 (1993) 1. 4. Ch.Papadopoulon, A.Lycourghiotis, P.Grange and B.Delmon, Appl. Catal., 38 (1988) 255. 5. H.K.Matralis, A.Lycourghiotis, P.Grange and B.Delmon, Appl. Catal., 38 (1988) 273. 6. J.A.R.van Veen, E.Gerkema, A.M.van der Kraan and A.Knoester, J. Chem. Soc. Chem. Commun., (1987) 1684. 7. I.E.Maxwell, Catal. Today, 1 (1987) 385. 8. M.Breysse, J.L.Portefaix and M.Vrinat, Catal. Today, 10 (1991) 489. 9. R.R.Chianelli, Catal. Rev.-Sci. Eng., 26 (1984) 361. 10. S.T.Oyama, Catal. Today, 15 (1992) 179. 11. R.Prins, V.H.J.de Beer and G.A.Somorjai, Catal. Rev.-Sci. Eng., 31 (1989) 1. 12. S.M.A.M.Bouwens, R.Prins, V.H.J.de Beer and D.C.Koningsberger, J. Phys. Chem., 94 (1990) 3711. 13. S.M.A.M.Bouwens, J.A.R. van Veen, D.C.Koningsberger, V.H.J.de Beer and R.Prins, J. PhyS. Chem., 95 (1991) 123. 14. M.J.Ledoux, O.Michaux, G.Agostini and P.Panissod, J. Catal., 96 (1985) 189. 15. M.J.Ledoux, G.Maire, S.Hantzer and O.Michaux, "Proc. 9th Intern. Congr. Catal." 1 (1988) 74.

89 16. W.K.Hall, Proc. of the Climax Fourth Intern. Conf. on the Chemistry and Uses of Molybdenum (H.F.Barry and P.C.H.Mitchell eds.), Climax Molybdenum Co., Ann Arbor, Michigan, 1982, p.224. 17. Y.Okamoto and T.Imanaka, J. PhyS. Chem., 92 (1988) 7102. 18. H.Knozinger, Proc. 9th Intern. Congr. Catal., 5 (1988) 20. 19. C.P.Li and D.M.Hercules, J. PhyS. Chem., 88 (1984) 456. 20. 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 and D.C.Koningsberger, J.Catal., 146 (1994) 375. 21. M.L.Vrinat, C.G.Gachet and L.de Moutgues, Catalysis by Zeolites (B.Imelik ed.), Elsevier, Amsterdam, 1980, p.219. 22. Y.Okamoto, A.Maezawa, H.Kane and T.Imanaka, Proc. 9th Intern. Congr. Catal., 1 (1988) 11. 23. Y.Okamoto, A.Maezawa, H.Kane and T.Imanaka, J.Molec.Catal., 52 (1989) 337. 24. M.Laniecki and W.Zmierczak, Zeolite, 11 (1991) 18. 25. C.Vorbeck, W.J.J.Welters, L.J.M.van de Ven, H.W.Zandbergen, J.W.de Haan, V.H.J.de Beer and R.A.van Santen, Zeolites and Related Microporous Materials: State of the Art (J.Weitkamp, H.G.Karge, H.Pfeifer and W.Holderich, eds.), Elsevier, Amsterdam, 1994, p. 1617. W.J.J.Welters, Thesis, Eindhoven University of Technology, 1994. 26. T.R.Halbert, T.C.Ho, E.I.Stiefel, R.R.Chianelli and M.Daage, J.Catal., 130 (1991) 116. 27. M.R.du Bois, Chem. Rev., 89 (1989) 1. 28. Y.Okamoto, Y.Kobayashi and T.Imanaka, Catal.Lett., 20 (1993) 49. 29. I.Alstrup, I.Chorkendorff, R.Candia, B.S.Clausen and H.Topsoe, J.Catal., 77 (1982) 397.

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Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All fights reserved.

91

NEW DEVELOPMENTS IN OLEFIN POLYMERIZATION WITH METALLOCENE CATALYSTS W. Kaminsky and A. Duch

Institute of Technical and Macromolecular Chemistry, University of Hamburg, Germany ABSTRACT Work in the application of metallocene-based catalysis to olefin polymers has become a research topic of growing interest in recent years. A great number of symmetric and chiral zirconocenes have been synthesizedto give totally different structures of isotactic, syndiotactic, atactic or block polymers. The isotactic sequence length of polypropylene is influenced by the nature of the ligands of the metallocene. New ring or bridge substituted metallocene/methylalumoxane catalysts for the olefin polymerization are described. 1. INTRODUCTION For more than a decade there has been a growing interest of research and publication that describes the amazingly high activity of metallocene-based catalyst systems to polymerize olefins [1-9]. The productivity of these metallocene systems now easily exceeds the range of 100 to 500 kg polyethylene (PE) or 10 to 30 kg isotactic polypropylene (PP) per gram of catalyst, typical of Ziegler-Natta catalyst systems presently in commercial use by a factor of 10 to 100 [10]. Furthermore, metallocene or single site catalysts have the capability to permit the user to better control polymer tacticity, molecular weight, and molecular weight distribution [ 11-16]. They give the opportunity to tailor polymer performance to specific applications, both in terms of processability requirements and end use performance. The application of this technology is in a rapid state of development in stiffness, higher impact strength and, in certain cases, higher clarity and lower density. For olefin copolymers, better control may be exerted on comonomer content, as well as on distribution statistics. Table 1 gives an overview about the efficiency of the zirconocene/alumoxane catalysts. The analogous titanium and hafnium compounds form active catalysts, too. Especially at higher temperatures the zirconium catalysts are more stable and active than the titanium or hafnium systems. The co-catalyst has a main influence. The most used co-catalyst is methylalumoxane (MAO). At lower temperature, fluorinated borane compounds activate the catalyst, too [17,18]. The structure of MAO is complicated. Sinn characterized MAO by different analytical methods and found it to be a cage with a formula of A116012(CH3)24 [ 19]. With metallocene catalysts, not only homopolymers such as polyethylene or polypropylene can be synthesized but also many kinds of copolymers and elastomers, copolymers of cyclic olefins, polyolefin covered metal powders and inorganic fillers, oligomeric optically active hydrocarbons [20-25]. In addition, metallocene complexes represent a new class of catalysts for the cyclopolymerization of 1,5- and 1,6-dienes [26]. The enantio-selective cyclopolymerization of 1,5-hexadiene yields an optically active polymer whose chirality derives from its main chain stereochemistry.

92 Table 1 Efficiency of the Zirconocene/Alumoxane Catalyst 1.

2.

3.

4.

5. 6. 7. 8. 9.

Ethylene

Homopolymerization high activity (40 x 106 g PE/g Zr x h) MJM, = 2, highly linear Ethylene Copolymerization random distribution, LLDPE co-monomers: propene, 1-butene, 1-octene EPDM Elastomers, Terpolymers ofEthene, Propene and Diene low transition metal concentration in the polymer, narrow molecular weight distribution Propene Polymerization to: a)Pure atactic PP b) Highly isotactic PP c) Syndiotactic PP d) Stereoblock PP e) Isoblock PP Cyclo-olefin Polymerization to Isotactic Polymers high melting point COC-Copolymerization of Ethene and Cyclic Olefins to transparent polymers high glass transition point Cyclopolymerization of 1,5-hexadienes Oligomerization to Optically Active Hydrocarbons Polymerization in the Presence of Filling Materials

2. SYNTHESIS OF SUBSTITUTED ZIRCONOCENES It is known that in propene polymerization the activity, the molecular weight, and the tacticity of the produced polymers differ by the substitution of the cyclopentadienyl ring and the substitution of the bridge. Some of these changes are dramatic and instructive. Changing the dimethylbridge into a diphenylbridge, increases the activity and the molecular weight by electronic effects. Benzyl ligands in the bridge are more bulky and could show a steric effect. Five benzyl or phenyl bridged zirconocenes are synthesized. In a first step benzylbromide is treated with magnesium, then with tetrachlorsilane and reacted with methylcyclopentadienyl sodium [27].

,,,

SiCI 4

I[~~2SiCI2 Synthesis of the Bridge

93

, , ~ CH3 =

Si

| 2Li

b-c., Synthesis of the Ligand The next step is the reaction with lithium and after this with zirconiumtetrachloride.

CH3

2

~

CH 3 No

,~CH 3 Si

,,,...-

H3

=,.,,..-

Si

Zr

.Ct "~Ct

CH 3

Synthesis of DibenzylsUyl bis(3-methylcyclo-pentadienyl)zirconium dichloride The produced dibenzylsilyl(3-methyl-cyclopentadienyl)zirconium dichloride 1 exists in racemic and meso forms. Figure 1 shows the ~H-NMR spectra of zirconocene 1 with different molar ratios of the racemic and meso-diastereomers.

94

r

m

"

m

'

'

i

'

""

r r

''

m

'

6.0

rac:meso

r

I

'

r-

m

,

m

~

5.0

=

I

'

r

~

r

t

m

l

6.1~

1"2

rac:meso

I

5.~

=

1:2

Figure 1.1H-NMR spectroscopic measured cyclopentadienyl protons of the racemic (r) and meso (m) diastereomers of dibenzylsilyl(3-methyl-cyclopentadienyl)zirconium dichloride.

.CI 9 Si

d

O>

'".. Si

"

CH~

rac-,meso-Diphenylsilylbis-(3-methylcyclopentadienyl) zirconium dichloride

o

Zr

. .Cl '

'

rac-,meso-Dibenzylsilylbis (2-methyl-4tert-pentyl cyclopentadienyl)zirconiumdichloride

Figure 2a. Synthesized cyclopentadienyl substituted zirconocenes. The chemical shifts could be assigned to the racemic and meso cyclopentadienyl protons by the fact, that only one signal is observed by an additional substitution in m-position of the cyclopentadienyl rings (ct-CsH3: r = 5.25 and 5.48 ppm, m = 5.08 and 5.68 ppm; 13-C5H3 : r 6.54, m = 6.43 ppm). By recrystallization the meso form can be partly removed. The structures of the analog produced zirconocenes are given in Figures 2a and 2b.

95

9.Ct

c,

Si

Zr

Si

Zr

"~Ct

rac-Dibenzylsilylbis-(indenyl) zirconiumdichloride

rac-Dibenzylsilyl-bis-(tetrahydroindenyl) zirconiumdichloride

Figure 2b. Synthesized zirconocenes with dibenzylsilyl bridges Table 2 Comparison of the polymerization of propene with different zirconocenes. Polymerization temperature: 30~ propene pressure: 2 bar; time: 1 h; solvent: 200 ml toluene; 300 mg; MAO; 6 x 10-6 mol. zirconocene. Activity (kg PP/mol. Zr x h x bar)

Catalyst

En(Ind)2ZrCl2 Me2Si(Ind) 2ZrC12 BzESi(Ind) 2ZrCl2 PhESi(Ind) 2ZrCI2 Me2Si(MeCp }2ZrC12 BzESi(MeCp) 2ZrCI2 PhzSi(MeCp) 2ZrC12

6 7 3 8 9 1 2

235 270 270 300 16 300 9 600 398

Molecular weight MTI

31 500 78 700 74 600 89 600 1 36 700 <106 <106

Melting point (~

136 148 145 146 147 145 141

3. PROPENE POLYMERIZATION

With the racemic mixture of zirconocenes in combination with MAO it is possible to obtain highly isotactic polypropylene. Non-chiral metallocenes as the meso forms produces only atactic polypropylene. The new synthesized zirconocenes (1), (2), (3) and for comparison some other compounds are used as catalysts for the propene polymerization under standard conditions. Table 2 shows that the silylbridged catalysts are more active than the ethylene bridged (6). The molecular weights of the bisindenyl compounds (3), (7), (8) are twice higher. The isotacticity is not much controlled by the ligands of the silyl bridge. This picture is different for the methyl-substituted cyclopentadienyl compounds. The methyl ligand in the bridge (9) shows the highest activity, but the produced polymer has the lowest molecular weight. The new compound (1) gives an excellent activity and a polypropylene with a high molecular weight. In comparison with compound (6), the activity is 40 times higher as well as the molecular weight, and the melting point is higher by 9~ For technical use, only the complicated separation of the racemic forms from the meso forms of compound (1)is a problem.

96 Table 3 Comparison of the polymerization of ethene with different zirconocenes. Polymerization temperature: 30 ~ ethene pressure: 2 bar; solvent: 200 ml toluene; 300-600 mg MAO; 6 x 10 -6 mol. zirconocene. Catalyst

Temp. (~

Bz2Si(MeCp)2ZrCI2 CpEZrCI2 Cp* 2 Z r C l 2 (EtMe4Cp) 2ZrCI2 (EtMe4Cp)EHfCI2

3 10 11 4 5

30 30 50 20 20

Activity (kg PE/mol. Zr x h x bar) 2 750 60 000 16 675 10 600 2 696

Molecular weight MB 350 000 620 000 539 000 1 609 000

4. ETHENE POLYMERIZATION

For the polymerization of ethene chiral, metallocenes are not necessary but high activity and linear polymers are wanted. Table 3 compares some metallocene catalysts for the polymerization of ethene. The activities for the substituted metallocene catalysts are lower than for the simple biscyclopentadienylzirconocene (10). In contrast to this, compound (5) shows a very high molecular weight of 1,6 millions. The activity of (EtMe4Cp)2ZrCI2/MAO for the ethene polymerizatzion increases with the temperature (Figure 3).

Activity 14,5 _ 14,0. 13,5 _ 13,0. 12,5 _ 12,0 _ 11,5 _ 11,0 _ 10,5 _

10,0

_

9,5_ 9,0

~b

2b

3b

4b

5b

6b

Temperature Figure 3. Activity of the polymerization of ethene with Catalyst (5) versus temperature; Activity in 103 kg PE/mol. Zr x h x bar.

97 There is a crack in the line at 30~ temperatures.

indicating a lower activation energy at higher

Polyethylenes synthesized by metallocene/MAO catalysts have a molecular weight distribution of MJM, = 2. The molecular weight can easily be lowered by increasing the temperature, increasing the metallocene concentration, or decreasing the ethene concentration. The narrow molecular weight distribution is characteristic for a single site catalyst. Nearly every zirconocene forms an active site of a cationic metallocene - and an anionic MAO compound or a complex of both. The nature of the active site would be clearer if more details were known about the structure of the alumoxane. 5. REFERENCES

1. H. Sinn and W. Kaminsky, Advances in Organometallic Chemistry, Vol. 18, Academic Press, New York, p. 99. 2. J.A. Ewen, J. Am. Chem. Soc., 106 (1984) 6355. 3. W. Kaminsky, H. Kfilper, H.H. Brintzinger and F.R.W.P. Wild, Angew. Chem. Int. Ed. Engl., 24 (1985) 507. 4. W. Kaminsky, R. Engehausen, K. Zoumis, W. Spaleck and J. Rohrmann, Makromol. Chem., 193 (1992) 1643. 5. W. Spaleck, F. Ktiber, A. Winter, J. Rohrmann, B. Bachmann, M. Antberg, V. Dolle and E.F. Paulus, Organometallics, 13 (1994) 954. 6. B. Rieger and J.C.W. Chien, Polym. Bull., 21 (1989) 159. 7. J. A. Ewen, L. Haspeslagh, M.J. Elder, J.L. Atwood, H. Zhang and H.N. Chong, in W. Kaminsky and H. Sinn (eds.), Olefin Polymerization, Springer Verlag, Berlin, 1988, p. 1536. 8. W. A. Herrmann, J. Rohrmann, E. Herdtweck, S. Spaleck and A. Winter, Angew. Chem., 101 (1989) 1536. 9. K. Soga and M. Kaminaka, Macromol. Rapid Commun., 15 (1994) 593. 10. N. Brockmeier, Proceeding, MetCon 94, 26-28 May, 1994 in Houston. 11. T. Mise, S. Miya, and H. Yamazaki, Chem. Lett., (1989) 1853. 12. W. Kaminsky, Catalysis Today, 20 (1994) 257. 13. M. Galimberti, G. Balbontin, I. Camurati, and G. Paganetto, Macromol. Rapid Commun., 15(1994)633. 14. Z. Guo, D.C. Swenson and R.F. Jordan, Organometallics, 13 (1994) 1424. 15. D. Fischer and R. Mtilhaupt, Macromol. Chem. Phys., 195 (1994) 1433. 16. G. M. Diamond, S. Rodewald and R.F. Jordan, Organometallics, 14 (1995) 5. 17. H. W. Turner, Europe Patent Appl. Nr. 0277004 to Exxon (1988). 18. M. Bochmann and S.J. Lancaster, Organometallics, 12 (1993)633. 19. H. Sinn, Symposium 40 Years Ziegler Catalysis, 1-3 Sept. 1993, Freiburg. 20. W. Kaminsky and R. Spiehl, Makromol. Chem., 190 (1989) 515. 21. S. Collins and W.M. Kelly, Macromolecules, 25 (1992) 233.

98 22. W. Kaminsky, A. Ahlers, O. Rabe and W. K6nig, in D. Enders, H. Gais and W. Keim (Eds.), Organic Synthesis via Organometallics, Vieweg, Braunschweig, 1993, p. 151. 23. N. Piccolrovazzi, P. Pino, G. Consiglio, A. Sironi and M. Motet, Organometallics, 9 (1990) 3098. 24. W. Kaminsky and S. Lenk, Macromol. Chem. Phys., 195 (1994) 2093. 25. A. Zambelli, A. Proto, P. Longo and P. Oliva, Macromol. Chem. Phys., 195 (1994) 2623. 26. G. W. Coates and R.M. Waymouth, J. Am. Chem. Soc., 199 (1993) 91. 27. H. Wiesenfeldt, A. Reinmuth, E. Barsties, K. Everth and H.H. Brintzinger, J. Organomet. Chem., 369 (1989) 359.

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All fights reserved.

99

NEW D E V E L O P M E N T S IN HYDROPROCESSING

J.W.M. Sonnemans Akzo Nobel Chemicals B.V., P.O. Box 247, 3800 A E Amersfoort, The Netherlands

I. ABSTRACT Changes in the demand for the various oil products combined with tighter product specifications has in the past and will continue to have in the future a strong impact on the role of hydroprocessing in refining. New technologies and, even more important, new catalysts have been developed to meet these changes. This paper will review the changes in product slates and product qualities of refined oil, both in the past and in the future, which influence hydrotreating and hydrocracking processes in the refinery. These changes in demands and specifications are generally met by applying and improving existing technologies. Catalyst development is crucial in this respect. A review will be given regarding new catalysts introduced into the market since 1990. New technologies introduced since 1985 are summarized. Four of the most interesting and successful new technologies are discussed in more detail. 2. I N T R O D U C T I O N The refinery industry is continuously faced with changes. These changes are in the quality of the crudes processed, the product slate and the required qualities of the products. To cope with these changes refineries must introduce new technologies from time to time or improve already practised technologies by revamping process units, by applying different, often improved, catalysts, or both. To be prepared for future changes and to improve the economics of the refinery, knowledge about available technologies and possible improvements of applied technologies, such as the use of advanced catalysts, is essential. This paper can be helpful to update that knowledge. First the paper reviews the changes in crude processed, product slate and product qualities. This is done for the period 1985 - 2000. Changes which occurred in the past for the various regions of the world may give indications of what might be expected in the near future. Next is a review of which technologies have been installed to address these changes followed by a summary of new hydroprocessing technologies, introduced during the last ten years. The four most successful new technologies are described in more detail. Catalyst developments are discussed next. Hydroprocessing catalysts introduced during the last five years are reviewed. The last chapter is a summary of the most important changes expected for the next five years.

100 3. F E E D AND P R O D U C T CHANGES

3.1 Feed changes The average quality of crudes processed worldwide was fairly constant the last ten years. However, important differences can be observed for the various geographic areas. In the USA the average crude changed considerably to a poorer quality, higher in density and sulfur as can be seen in figure 1. Lower quality of imported crudes, in particular from South America (24.8~ S = 2.20 wt%) and decreasing availability of the sweeter USA crudes (32.3~ S = 0.87 wt% ) influenced the average crude quality. It is expected that the sulfur content will further increase due to increasing imports from the Middle East (32.9~ S = 1.80 wt% ). The density will stay fairly constant at 31.7 [1]. The average crude quality in Western Europe has been rather constant over the last ten years (34.6 ~ S = 1.02 wt%). In the future we expect that Western Europe will face the changes as observed in the USA over the past ten years because of the decreasing availability of North Sea crudes (36.1 ~ S = 0.23 wt%, [2]). However, this trend will not start before the end of the century due to the still increasing availability of sweet crudes [3]. 1.8.

37, SULFUR

API 36

1.6. JAPAN

1.4.

JAPAN

35

/ "

-~~

1.2 33

J

USA

USA

32

0.8 0.6

31 85

Fig. 1.

87 YEAR89

91

93

85

87 YEAR 89

91

93

Changes in crude quality processed in the USA and in Japan during the last ten years.

Japan faced a trend opposite to what was experienced in the USA (figure 1). After 1987 the average crude processed became lighter and lower in sulfur, despite the fact that the percentage of crude imported from the Middle East increased from 68 % to 77 %. Other imported crudes were very sweet crudes, for instance those imported from Indonesia. The present low price difference between light and heavy crude [3] will enable Japan to process a light average crude quality in the near future, but on longer term Japan will also face an increase in sulfur and density.

3.2 Product slate changes The continuous shift in product slate from fuel oil to transportation fuels is a well-known trend of the last 10 years. It is interesting to compare this shift for the various geographic areas presented in figure 2. This figure is based on statistics annually supplied by BP Oil [4]. The gasoline fraction is constant in the USA, but increases in all other regions. Middistillate is increasing in most regions with the highest growth in Japan. For LDC (lesser

101

45 USA

40

w . EUROPE ,..1

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developed countries) we see a constant value of 36 %, lower than the value for OECDEurope, but higher than the value for North America. In Western Europe distillate use is high, due in part to the relatively high percentage of personal diesel cars. Relative fuel oil consumption decreases in all regions. The level of fuel oil consumption is quite different for the various geographic areas. Japan has now reached the level typical for Western Europe 10 years ago; Western Europe has not yet reached a level typical for North America 10 years ago. The use of oil products for other applications showed a slight increase of 2 % in most areas, on a level of about 18 %. Because of the wide variety of applications (LPG, solvents, lubricants, bitumen, refinery fuel) this hardly influenced refinery configuration.

What can be expected the next 5-l0 years ? Fuel oil will continue to decrease in all regions, but only to a minor degree in North 4O America, where fuel oil use is already quite *%~176 35 low. In Western Europe and Japan the 30 decrease will also be less than experienced during the last 10 years, because fuel oil has o 25 JAPAN "" ...... now reached a price level equal to coal, taking away incentive for power stations to -" ...... "~.-E~-R-O~R"" 15 shift from oil to coal. For the LD Countries USA a considerable decrease is expected to meet the increasing demand for transportation fuels. 83 85 87 89 91 93 YEAR Relative gasoline demand in North America will stay constant or may even Fig 2. Gasoline, mid-distillate and fuel oil decline slightly.. Depending on technology consumption as percentage of the total crude improvements for diesel cars and tax for various geographic areas, benefits for diesel oil Western Europe still may face a small increase. Gasoline consumption as percentage of oil consumption will increase further in all other regions. Mid-distillate use will increase in all geographic areas, absolute as well as relative (percentage of total crude). Developments in LD Countries are of extreme importance in 83

85

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102 this respect. In these countries not only the shift in product slate has to be addressed but, more importantly, a strong increase in total consumption will dominate the developments (figure 3). In particular the strong growth in Asia (about 5 % each year) will result in an annual shortage in transportation fuels in this area [5]. Mid-distillate can be either heating oil, diesel oil or jet fuel (kero). Europe faced a shift from heating oil to diesel oil allowing them to produce additional barrels of diesel oil without installing conversion Fig. 3 Oil consumption in 1985, 1993 and capacity. The percentage of jet fuel (kero) 2000 in various geographic areas, is at present 10 % in the USA, 6 % in Western Europe and 12 % in Asia Pacific [6]. This will increase worldwide because the expected strong growth in air-traffic of 5 % each year for the next 15 years. How these changes in product slates have been achieved in the past will be discussed in chapter 4.

3.3 Product quality changes Product qualities are strongly environmental driven. Main trends in product oil specification in the past and expected changes for the near future are presented in Table 1. Changes in gasoline specifications strongly influenced and will further influence the operation of FCC units, particularly in the USA. The upcoming further decrease in sulfur in gasoline will have an effect on hydroprocessing. At present about 30 % of the worldwide FCC capacity processes hydrotreated feeds. This percentage will increase. Moreover, in the future an important part of the FCC gasoline will be post-hydrotreated to meet gasoline sulfur specifications. The specifications for the other oil products have been met by installing hydrotreating and hydrocracking capacity and by applying more active catalysts. More information is presented in chapter 4.

4. TECHNOLOGY DEVELOPMENTS

4.1 Technologies installed to meet changes in feeds and products To address the changes in feed quality, in product slate and in product qualities, the refining industry has installed both well established and new technologies. Product slate changes were met by installing conversion capacity, either thermal or catalytic. Table 2 presents the relative installed conversion capacity for the various geographic areas.

103

Table 1 Changes in product oil specification Product Type

Gasoline

Diesel oil

1 9 9 0 - 1994

1995 - 2000

o Lead phase out o Lower RVP

o

Reformulated gasoline in USA o Lower sulfur content, 0.05 wt% in OECD countries o Slightly lower aromatics in OECD countries

o Lower sulfur content; 0.05 wt% in USA, 0.2 wt% in Japan and W. Europe o Lower aromatics in California and Scandinavia o

Heating oil

after 2000

Sulfur content below 100 ppm o Further decrease of sulfur content in all countries o Higher cetane number o Lower endpoint

Lower sulfur content in W. Europe (0.1 wt%) Lower sulfur content in W. Europe (1.5wt%)

Fuel oil

Table 2 Installed conversion capacities ~ early 1985, 1990 and 1995 in various geographic areas

Geographic area

North America OECD Europe Japan Middle East LDC

1985

47.9 % 21.2 % 11.2 % 14.7 % 18.1%

1990

50.6 26.8 19.3 22.7 20.9

% % % % %

1995

54.0 % 33.1% 18.9 % 19.8 % 20.6 %

~ Conversion capacity is total of thermal cracking, FCC and hydrocracking capacity divided by crude oil processing capacity [7]. The increase in conversion is due in part to installed new conversion capacity and in part to the closing down of refineries which lacked conversion capacity. On worldwide basis FCC is still the most widely applied conversion process. Each year about 300.000 bpsd capacity has been added the last ten years resulting in a present worldwide capacity of 12.3 million bpsd. Thermal conversion is growing with 170.000 bpsd to a present total of 6.9 million bpsd. For hydrocracking these figures are 140.000 bpsd and 3.2 million bpsd, respectively. This makes hydrocracking the relatively fastest growing conversion technology. The last years the installed

104 hydrocracking capacity reached a level of 200.000 bpsd each year. For distillate type feedstocks Unocal/UOP has been by far the most successful company in this field by licensing the majority of the added hydrocracking capacity. IFP and Chevron have been less successful licensors. Newcomers on this market segment, Texaco and the alliance between Mobil, Akzo Nobel and Kellogg (MAK), may change this picture for the next decade. The type of conversion units installed differ greatly for the various regions. This is dictated by the changes in product slate, but also influenced by imports and crude quality changes. In North America about 50 % of the added conversion capacity in between 1985 - 1995 was based on hydrocracking; FCC adding only 20 %. For Japan these figures were 19 % and 81 %, respectively. Japan did not install thermal cracking capacity, but thermal cracking remains to play an important role in North America. The most important trend in product qualities changes for all oil products is the continuous decrease in sulfur level resulting in a worldwide increase in hydrotreating capacity as is shown in table 3. Table 3 Hydrotreating capacity as percentage of total crude processed (hydrocracking and hydrofining included) [7] Geographic area

North America OECD Europe Japan Middle East LDC

1985

1990

1995

63.1% 43.4 % 69.5 % 32.1% 28.7 %

68.4 % 46.3 % 80.9 % 47.4 % 31.1%

73.0 % 55.1% 77.4 % 40.5 % 31.3 %

A wide variety of technologies have been applied to increase desulfurization. In resid desulfurization 43 % of the capacity coming on stream the last 10 years was licensed by Chevron and 42 % was licensed by Unocal (KNPC units included). Chevron is the leading process licensor of (vacuum) resid upgrading today. For lighter feeds Unocal has been successful with 7 units added after 1985, and 8 units in design or under construction. The large number of diesel hydrotreaters which came on stream the last 10 years apply a variety of technologies, licensed (e.g. by Exxon, Unocal, IFP) as well as non licensed. Non licensed designs were often based on know how supplied by catalyst manufacturers in cooperation with engineering companies. In the future tighter product specifications (table 1) will increase the degree of desulfurization in all geographic areas. In 1992 in Western Europe 3 1 % of the crude sulfur was removed as sulfur. About 8 % left the refineries as SO2 emissions. The rest remained in oil products [8]. The new sulfur specification for diesel oil will increase the sulfur removal as solid sulfur to 36 %. The worldwide increase in desulfurization together with the increase in sulfur level of average crude processed will result in a strong increase in solid sulfur coming on the market originating from crude oil as is shown in figure 4 [9]. This will upset the market for sulfur, unless the contribution of other sources will be reduced by at least 50 % over the next 15 years, or unless new applications for sulfur will be found.

105

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I

1985

I

I

1990

I

I

1995

I

I

2000

I

I

2005

YEAR

Fig. 4. Worldwide supply and demand for sulfur.

4.2 New hydroprocessing technologies Each year new hydroprocessing technologies are announced, mostly at well attended conferences such as the annual spring meeting of the National Petroleum Refiners Association. Table 4 summarizes new technologies announced after 1984. It is interesting to see that a number of new resid conversion technologies were announced in the eighties. Most of these technologies lack success due to: a. Less interest in resid upgrading mainly caused by the low profitability of the refining industry and abundant supply of sweet crudes b. Successful introduction of resid FCC technologies c. Selection of well proven technologies such as fixed bed processes from Chevron and Unocal and ebullated bed processes from HRI (now IFP/Texaco) and Lummus. Three new resid technologies were licensed to two or more other companies. The Hycon process was licensed to Agip and Showa Yokkaidi, but the demetallization part containing the bunker flow reactor will not be constructed. The VCC process was licensed twice, but it is still not certain whether the VCC units will be build. The VCC process might become an important resid technology in the next century, because of its flexibility, good operation experience in the Bottrop plant and good quality of the products. A very interesting new resid technology is the OCR process licensed by Chevron. See section 4.2.1. Two new technologies were recently announced for hydrocracking of VGO type feeds. The different concepts used in these technologies, ebullated bed licensed by Texaco and a lower pressure process resulting in lower investment costs licensed as MAK technology, may give these new technologies an edge over existing technologies. Moreover, the cooperation between a major oil company, with in house experience in hydrocracking, a leading catalyst supplier and a well experienced engineering company, could give the MAK technology a strong competitive basis. Most of the technologies announced in the nineties do have the production of high quality diesel oil as objective. Dewaxing, desulfurization and hydrogenation are offered to reach high quality diesel oils. The Akzo-Fina CFI process is unique in this respect, because it offers all

106 Table 4

New hydroprocessing technologies announced after 1984 Name (year of announcement)

Licensor

Feed

Process

Hycon ('85)

Shell

Resid

Conversion

R-HYC ('85)

Idemitsu

Resid

Conversion

Hyvahl-F ('87)

IFP-Total-Elf

Resid

Conversion

Resid hydroconversion technology ('88)

Cosmo Oil

Resid

Desulfurization and conversion

Akzo-Fina CFI ('88)

Akzo Nobel-Fina

Light and heavy distillate

Dewaxing, hydrotreating, hydrocracking

Veba Combi Cracking ('89)

Veba

Resid

D.O.D.D. ('89)

Exxon

References

Licensed to third parties

10

Yes Yes

12

Yes No

29-31

Yes

Conversion

14

Yes

Middistillate

Desulfurization

32

Yes

HDH ('90)

Intevep

Resid

Conversion

IFP deep HDS and HDA ('91)

IFP

Middistillate

Desulfurization and hydrogenation

16

Yes

Integrated HDS/HDA process ('92)

Litwin Engineers and Constructors

Middistillate

Desulfurization and hydrogenation

17

No

OCR ('92)

Chevron

Resid

Demetallization

28

Yes

Synsat ('92)

Lummus-Criterion

Middistillate

Desulfurization and hydrogenation

33-34

Yes

Process for improving diesel quality ('92)

Topsoe

Middistillate

Desulfurization and hydrogenation

18

No

MHUG ('93)

Synopec

Middistillate

Desulfurization and hydrogenation

19

No

T-star ('93)

Texaco

Light and heavy distillate

Hydrocracking, FCC pretreatment

20

Yes

Octgain ('93)

Mobil

Gasoline

Desulfurization

SMDH ('93)

Shell

Middistillate

Desulfurization and hydrogenation

Biodesulfurization ('93)

Energy BioSystems

Light and heavy distillate

Desulfurization

MAK ('94)

Mobil-Akzo Nobel-Kellogg

Light and heavy distillate

Hydrocracking

25

No

Two stage deep HDS ('94)

Nippon Oil

Middistillate

Desulfuization

24

Yes

Synsat ('94)

Lummus-Criterion

Gasoline

Desulfurization

26

No

MAK ('95)

Mobil-Akzo Nobel-Kellogg

LCO

Hydrocracking

27

No

No

No 22

No

No

107 these functions, including hydroconversion, in one process step. See section 4.2.2. From the many technologies offering desulfurization combined with hydrogenation resulting in diesel low in sulfur and aromatics, the Synsat process is the most successful. See section 4.2.4. Next to the Synsat and Akzo-Fina CFI technology, three new technologies have become available to desulfurize mid distillate type feedstocks. Exxon was very successful in licensing the DODD technology in the recent years. See section 4.2.3. On longer term biodesulfurization might be attractive for diesel oil and kero desulfurization, because of low hydrogen consumption and low investment. Gasoline desulfurization is not yet a main topic, but two technologies were recently announced to achieve gasoline desulfurization with no or limited octane loss, the Octgain and the Synsat process respectively. The lower octane loss at a given level of desulfurization will give the Octgain process an edge. However, other gasoline desulfurization technologies are expected to enter the market in the future. The continuous decrease of fuel and heating oil and still increasing installed FCC capacity might result in a problematic future for LCO. The expected decrease in final boiling point of FCC gasoline will further magnify this. The recently announced MAK-LCO process [27] is an interesting option in this respect compared to the already practiced LCO upgrading technology licensed by Unocal, now UOP. Below we will discuss in more detail the four most successful technologies introduced after 1985.

4 . 2 . 1 0 C R process [28] The main objective of OCR process is demetallization of residues. To decrease reactor volume and to make an effective use of the catalyst Chevron developed a technology in which oil and catalyst are fed to the reactor in counter-current flow (figure 5). The beaded catalyst is added once or several times a week to the top of the reactor and moves in plug flow to the bottom of the reactor, where the spent catalyst is removed. Chevron carried out extensive experiments to develop a safe and reliable system to feed fresh and to withdraw spent catalyst. The catalyst is added as a slurry in Fresh Catalyst Bin oil. To prevent clogging of the catalyst in the reactor, catalyst Catalyst ~ Product properties and oil flow are chosen Feed Vessel to RDS .... Reactor in such a way that the catalyst bed is slightly expanded. The slight motion of the catalyst prevents agglomeration, but is small enough OCR to secure plug flow of oil and Reactor catalyst. Catalysts used in the OCR process are ICR 138 and ICR 143. The OCR process can be applied in two ways: Low Pressure Catalyst Vess, 1. in new designs to decrease Feed in reactor volume and to improve flexibility to process different Spent Catalyst Bit crudes. Fig. 5. OCR reactor system.

108 2. as pretreat step in existing units to allow refiners to process heavier crudes.In ref.28 it is mentioned that the demet reactor of an existing unit can be revamped to an OCR unit, but we doubt whether this is a realistic option. Main advantage of the OCR technology is that heavier crudes can be processed (metals above 200 ppm). The technology has been installed in the Aichi refinery of the Idemitsu Kosan Company where it has been in operation from May 1992. Capacity of this unit is 50000 bpsd. Three other refineries apply (IKC, Hokkaido) or will apply (Mitsubishi, Mizushima and Chevron, Valero) the OCR technology. The commercial performance up till now is rather successful. Several problems faced after starting up of the first unit could be solved. The problems as faced in the first step of the Hycon process, a similar technology, in which process the clogging of catalyst particles resulted in long shut downs of the unit, have not been observed in the OCR process. The demet performance was satisfactory: on the other hand the longer expected cycle length has not been achieved. The present low price delta between light and heavy crude does not make it economically attractive to install the OCR technology. However, in the future we expect that this price delta will increase again and more OCR units may be build, if with further improvements the cycle length will be increased.

4.2.2 Akzo-Fina CFI process [29-31] The main objective of the Akzo-Fina CFI process is to improve cold flow properties (cloud point, pour point, cold filter plugging point) of hydrocarbon feedstocks. However, in addition to dewaxing activity, desulfurization, denitrogenation, hydroconversion or hydrogenation activity also can be incorporated giving this technology a high degree of flexibility. For instance, high sulfur gasoils can be processed to yield diesel oils low in sulfur with excellent cold flow properties. Light vacuum gasoils can be processed to yield a product which can be sent without further treatment or distillation to the diesel oil pool. Fina developed the technology which was practiced in their own refineries before they decided to license it in cooperation with Akzo Nobel to other companies. Akzo Nobel has successfully marketed the technology with eleven licenses sold today in Europe, North America and Japan. The process scheme is identical to typical hydrotreating units. Process conditions are similar to what is practiced in LGO or VGO hydrotreating or mild hydrocracking units. Catalysts applied in the process are only available under license. The first licensed Akzo-Fina CFI unit came on stream in Italy in November 1990 and was able to operate for nearly two years without any reactivation of the catalyst. A typical result is presented in table 5. This first unit was a grass root unit, but the technology can and has been also applied in existing hydrotreaters and units of similar design (e.g. dewaxing units). Examples of possible applications of the technology, also demonstrating the excellent regeneration of the catalyst, are presented in ref. 30 and ref. 31.

4.2.3 Dodd process [32] The objective of this process is the deep desulfurization of middistillate types of feedstocks. Exxon licensed hydrotreating units in the past, but decided that a new technology had to be developed for deep desulfurization of diesel oil. Main differences from the existing technology is the different desulfurization kinetics for deep desulfurization and the importance of a good

109

Table 5 Commercial performance of the Akzo-Fina CFI process

Crude Parameters

Low S SOR

High S SOR

Low S 18 MOS

0.872 0.49 25 28 22

0.888 1.40 25 27 7

0.850 0.53 27 31 25

394

406

427

3.8 19.7 76.5

2.5 13.4 84.1

4.2 20.9 74.9

0.871 0.005 - 15 -4 54

0.881 0.015 - 15 -4 31

0.875 0.015 - 9 -7 40

Feedstock properties Density Sulfur Pour point Cloud point 350~ -

g/ml wt% ~ ~ vol%

Feedstock distillation 90 vol%

~

Product yield (wt%) LPG Naphtha 180~ +

Product (180~ § properties Density Sulfur Pour point Cloud point 350~ - fraction

g/ml wt% ~ ~ vol%

liquid distribution in deep desulfurization. Substituted dibenzothiophenes are the most resistant sulfur compounds which remain in partly desulfurized diesel oil. Conversion of these compounds does follow different kinetics from the general 1.8 order of reaction observed for desulfurization of high sulfur diesel oils; at high degree of desulfurization the order of reaction decreases. For the proper design of deep diesel oil hydrotreaters detailed knowledge of the kinetics is required. Exxon claims that they have a large data base as a result of tests done with different feeds applied in a total of 400 different conditions. Detailed knowledge of sulfur compounds distribution and higher performing catalysts are mentioned as keys for minimizing costs. Also oil distribution is becoming a very critical aspect in deep desulfurization requiring special reactor internals technology to achieve good wetting of the catalyst and to avoid channelling. The Exxon process for diesel oil production with sulfur levels between 200-500 ppm was introduced in 1989. The process is particularly interesting for feeds containing cracked stocks. The low pressure results in an efficient use of hydrogen in diesel upgrading. Since its introduction 9 units have been licensed outside the Exxon group. Various catalysts are approved for this process. RT 601 is the most active one.

110 4.2.4 Synsat process [33, 34] The objective of the Synsat process is to convert feedstocks boiling between 200 and 360~ to high quality diesel oils low in sulfur and aromatics. The technology was developed in cooperation between Criterion Catalyst Company and ABB Lummus Crest. It finds its basics in the desulfurization and hydrogenation catalysts manufactured and sold by Criterion and the technology applied by Lummus in their Arosat process. A typical flow sheet of the Synsat process is presented in figure 6.

Vapor/liquid ] Separation/Recycle System

"'"'""['"'"'~~,1 I

Make

rv, Product Stripping

Fresh Feed Fig. 6. Flow scheme of the Synsat process. Main characteristic of the process is a two step approach. The first step achieves deep desulfurization which may be carried out in two beds. In this part of the process hydrogen and oil are processed in co-current flow. In the second step desulfurized oil is stripped by hydrogen to remove HzS and than is processed in counter current flow over a hydrogenation catalyst. The counter current flow with fresh hydrogen feed at the bottom of the hydrogenation reactor has the advantage of conditions which are most favorable for optimal hydrogenation from thermodynamic point of view. According to the licensor this saves considerable reactor volume compared with co-current flow operation which is only to be expected in case final aromatic content is close to the equilibrium level. The poorer wetting and the higher investment of counter-current flow operation compared to the widely applied co-current operation are apparently offset by the improved kinetics. The two steps of the process can be achieved in one reactor or two separate reactors. Important is that in between the two steps gas liquid separation takes place at high temperature and that the liquid formed after cooling down of the hot gas stream is recycled to the reactor. The process can be operated at moderate pressures, 5 0 - 70 bar at reactor outlet. Temperatures have to be low (below 320 ~ to achieve deep hydrogenation. Space velocities are low (below 0,5 hr -1) if low product aromatic contents of e.g. 5 vol% are required. The first step of the process is not different from a typical diesel hydrotreating unit. Five different catalyst types in staged loading can be used to obtain the required desulfurization and hydrogenation level. Catalysts are only available under license. At the spring NPRA meeting of 1995 it was announced that 13 companies have taken a license. This impressive success is for an important

111 part due to the possibility of a staged investment, at present for desulfurization and at a later date, when required, for hydrogenation. In ref. 34 it was mentioned that 20 - 25 % of the total investment can be delayed at the drawback of a 10 % higher total investment. Most of the Synsat licensees decided to install only the hydrodesulfurization part, which was not the case for Scanraf (paper AM-93-24) and Beta (paper AM 94-62) who are planning to produce Class 1 diesel for Sweden (10 ppm sulfur, 5 vol% aromatics, FBP 300 ~ Beta is in fact the first company who practiced the Synsat process. 5. C A T A L Y S T D E V E L O P M E N T S Every year many new catalysts become available for a wide variety of applications. Because it is impossible to discuss all these catalysts, we will limit the discussions to a summary of new grades introduced in the last five years and review only the most important developments for the various application segments, being resid upgrading, hydrocracking, FCC pretreatment, middistillate desulfurization and deep hydrogenation. Table 6 summarizes new catalysts introduced in the last 5 years and still being sold in 1995 [35]. Each year 15 - 20 new catalysts are introduced in the market. Akzo Nobel, Chevron and Criterion appeared to be the most active companies. Only a limited number of new catalysts have been introduced for ebullated bed service. Criterion and Grace dominate this market segment; both companies have introduced new catalysts. Criterion introduced two new catalysts, one which was developed in cooperation with Texaco, one of the process licensors. Criterion also introduced a new series of catalysts for fixed bed resid upgrading which has narrowed the gap in performance between Criterion and leading companies in this market segment, Akzo Nobel and CCIC. The many new catalysts commercialized by Chevron for resid upgrading has made it also a leading company in fixed bed resid upgrading. In the hydrocracking pretreat segment Criterion and Akzo Nobel dominate the market with licensed catalysts and open market catalysts (HC-H and C-411 of Criterion; HC-K, Ketjenfine 843 of Akzo Nobel). Both companies recently introduced new catalysts (DN-120 from Criterion, HC-P and Ketjenfine 846 from Akzo Nobel); Criterion and Akzo Nobel will probably continue to dominate this market segment the coming years. The developments in hydrocracking are very interesting. Many new catalysts have entered the market, new catalysts from technology licensors as Unocal/UOP and Chevron as well as open market catalysts. These catalysts target for selective naphtha production as well as for selective diesel or diesel/jet production. Despite the development of improved grades by the licensors Unocal/UOP and Chevron we notice an important shift in the market from licensed catalyst to open market catalyst. Criterion has been successful with sales to 9 units outside Shell hydrocrackers (ref. 28); Akzo Nobel supplied 5 units. In the eighties resid upgrading shifted from licensed catalysts to open market catalysts and we see now the same happening for hydrocracking in the nineties. FCC pretreatment is an area dominated by Akzo Nobel, which has introduced two new catalysts for this service, Ketjenfine 841 and Ketjenfine 901H. The last catalyst is sold very successfully, especially in Japan. Several other companies also have introduced new catalysts; notable is the less active role of Criterion in this market segment despite the fact that they are second supplier after Akzo Nobel.

112

Table 6 N e w c a t a l y s t s i n t r o d u c e d in the nineties

Application

Company

Catalyst codes

Resid Ebullated-bed

Akzo Nobel Criterion Grace Haldor Topsoe

KF 1302, KF 1303 HDS-2443, TEX-2710 GR-31, GR-250 TK-871

Resid Fixed-bed

Akzo Nobel CCIC Chevron Criterion Orient Unocal/UOP

KF 645, KFR 11, KFR 20 Actmax, CDS-R95 ICR-130-133, ICR-135, ICR-137 C- 117, RN/RC-400, RN/RC-410,RM-430 HOP-606 RF-1000, RF-1100

Hydrocracking Pretreatment

Acreon/IFP Akzo Nobel Criterion Unocal/UOP

HTH 548 KF 846 DN-120 HC-P

Hydrocracking

Akzo Nobel CCIC Chevron Criterion (incl. Zeolist) Unocal/UOP

KC 2300, KC 2610, KC 2710 NHC-10 ICR 136, ICR 208, ICR 120, ICR 139 Z-763, Z-603, Z-713, DN-801 HC-26, DHC-32, HC-33

Mild hydrocracking

Akzo Nobel Chevron Criterion Orient

KC 2602 ICR 140 DN-800 HOP-522

FCC pretreatment (incl. VGO upgrading)

Acreon/IFP Akzo Nobel Chevron Haldor Topsoe Leuna Unocal/UOP

HR 360 KF 901H, KF 647, KF 847, KF 841 ICR 403, ICR 134 TK-525 8209 HC-J, HC-DM

Distillate Desulfurization

Distillate hydrogenation

Acreon/IFP

HR 316C

Akzo Nobel CCIC Criterion Crosfield Exxon Haldor Topsoe Leuna Orient Unocal/UOP

KF 756 CDS-D21, CDS-LX3 C-448, DC- 130 520, 465 RT-601 TK-524, TK-554 8408 HOP-463, HOP-471 N-108

Akzo Nobel Criterion Haldor Topsoe

KF 852 C-624, Synsat 1-4 TK-907, TK-908

113 As expected many new catalysts have been commercialized for light feed hydrodesulfurization. Nearly all catalyst manufacturers are active in this segment. At the end of the eighties Akzo Nobel introduced Ketjenfine 752 which until recently was the most active diesel HDS catalyst with sales of over 15.000 tons. A few catalyst manufacturers recently introduced catalysts with a performance in a range similar to that of Ketjenfine 752. Early this year Akzo Nobel introduced Ketjenfine 756. This catalyst was developed in cooperation with Exxon and successfully used as RT 601 in several units in Exxon refineries and DODD units. Ketjenfine 756 is most probably the highest activity commercially applied diesel HDS catalyst today. Several new hydrogenation catalysts have been announced. Criterion and Haldor Topsoe are the most active companies in this segment. Sales are small at present, but may become important when more severe diesel aromatics or cetane number specifications are adopted. For one step hydrogenation Akzo Nobel's Ketjenfine 852 could be an attractive catalyst.

6. MOST I M P O R T A N T D E V E L O P M E N T OF THE COMING YEARS

In summary the most important developments are reviewed in this section. Crude oil supplied to the refineries will not change significantly in quality the next five years because of the wide availability of sweet crudes. The product slate will continue to shift from fuel oil to transportation fuels (outside the USA). Distillate and jet will be the products with the strongest growth [36]. All oil products will face tighter specifications; in particular sulfur levels will be decreased. Conversion has been increased worldwide to meet the higher demand of transportation fuels. FCC is still the conversion technology with the highest amount of barrels licensed each year. In relative terms hydrocracking is the fastest growing conversion technology. For the future we expect a slow shift from FCC to hydrocracking. A number of new technologies were announced the last ten years, in particular for resid upgrading and middistillate production. Most interesting of the new technologies are the OCR process licensed by Chevron for resid demetallization, the Akzo-Fina CFI process licensed by Akzo Nobel for dewaxing with hydrotreating/hydrocracking, the DODD process licensed by Exxon for deep middistillate desulfurization and the Synsat process licensed by ABB Lummus Crest for middistillate production low in sulfur and aromatics. Many new catalysts have entered the market particularly for resid upgrading, hydrocracking and middistillate deep desulfurization. The leading catalyst suppliers today, Criterion and Akzo Nobel, have been the most active companies with new catalysts for nearly all market segments. REFERENCES

1. 2. 3. 4. 5.

Oil and Gas Journal, Dec. 27, 1993, page 30. Oil and Gas Journal, May 23, 1994, page 43. P. Horsnell,The future for heavy crude oil, Oxford Energy Forum, Febr. 1995, page 10. BP Statistical Review of World Energy, June 1994. A.I. Spiers, B.F. Burke and J.T. Boepple, 7th Annual Conference on International Refining & Petrochemicals, May, 1994, Singapore

114 6. 7. 8. 9. 10.

11. 12. 13. 14. 15.

16.

17.

18. 19.

20. 21.

22. 23. 24. 25. 26. 27.

J.R. Green and C.R. Kennedy, Technology trends and Needs for the Petroleum Refining Industry, Catcon Conference, June 1994, Philadelphia, U.S.A. Oil and Gas Journal, Dec. 31, 1984; March 18, 1985; Dec. 25, 1989; March 26, 1990; Dec. 19, 1995. Concawe Review, vol. 3, Oct. 1994, number 2, page 8. G.H. van den Berg, Akzo, the largest independent supplier of hydroprocessing catalysts, Akzo Nobel Catalyst Symp., June 1994, page 1. W.C. Van Zijll Langhout, P.B. Kwant and G.J. Lambert, The development of Shell's Residue Hydroconversion Process, Symp. on Oil Refining, Sept. 30Oct. 4, 1985, La Plata, Argentina. H. Sue, M. Fujita, Oil and Gas Journal, May 26th, 1986, page 51. M.E. Morrison, G. Stevens, A. Billon, A. Hennico and J.P. Peries, Hyvahl-Solvahl; key processes for upgrading of residues, NPRA annual meeting, paper AM 94-25. H. Yamazaki, K. Tawara and T. Tomino, New resid hydroconversion technology boosts up mid-yield, NPRA annual meeting, paper AM 88-60. F.W. Wenzel and K. Niemann, A commercial route for bottom of the barrel upgrading, NPRA annual meeting, paper AM 89-14. R.B. Solari, J. Guitian, J. Krasuk and R. Marzin, HDH, hydrocracking as an alternative for high conversion of the bottom of the barrel, NPRA annual meeting, paper AM 90-38. J.P. Peries, A. Billion, A. Hennico and S. Kressman, IFP deep hydrodesulfurization and aromatics hydrogenation on straight run and pyrolysis middle distillates, NPRA annual meeting, paper AM 91-38. D. Eastwood, C. Tong and C.S. Yen, Integrated hydrotreater/aromatic saturator process for the production of oils with reduced aromatics and polynuclear aromatics, NPRA annual meeting, paper AM 92-63. P. SCgaard-Andersen, B.H. Cooper and P.N. Hannerup, Topsoe's process for improving diesel quality, NPRA annual meeting, paper AM 92-50. Y.L. Shi, J.W. Shi, X.W. Zhang, Y.H. Shi and D.D. Li, MHUG process for the production of low sulfur and low aromatic diesel fuel, NPRA annual meeting, paper AM 93-56. W.F. Johns, G. Clausen, G. Nongbri and H. Kaufman, Texaco T-Star process for ebullated bed hydrotreating/hydrocracking, NPRA annual meeting, paper AM 93-21. M.S. Sarli, D.L. Fletcher, T.L. Hilbert, G.G. Karsner, S.S. Shih and P. Xayariboun, OCTGAIN TM, a new unique gasoline desulfurization process, NPRA annual meeting, paper AM 94-39. J.P. van den Berg, J.P. Lucien, G. Germaine and G.L.B. Thielemans, Fuel Processing Techn., 35, 1993, page 119. D.J. Monticello, Biological desulfurization (BDS) of middle distillates, NPRA annual meeting, paper AM 93-14. M. Ushio and M. Hatagama, Development of two stage desulfurization process of diesel gasoil, Pre conference TO-CAT 2, Aug. 1994, Tokyo, Japan. M.G. Hunter, D.A. Pappal and C.L. Pesek, Moderate pressure hydrocracking: a profitable conversion alternative, NPRA annual meeting, paper AM 94-21. D. Edgar, C. Bosland and A. Moore, Fuel Reformulation, May/June 1994, page 32 D.A. Pappal, C.K. Lee, K.W. Goebel and M.Y. Asim, MAK-LCO: MAK light cycle

115 oil upgrading to premium products, NPRA annual meeting, paper AM 95-39. 28. B.E. Reynolds, R.W. Bachtel and K. Yagi, Chevron's on-stream catalyst replacement (OCRTM)provides enhenced flexibilities to residue hydrotreaters, NPRA annual meeting, paper AM 92-61. 29. H.W. Homan Free, T. Schockaert and J.M.W. Sonnemans, The Akzo-Fina cold flow improvement process, Fuel Processing Techn., 35, 1993, page 111. 30. H.W. Homan Free and J. Mertens, The Akzo-Fina CFI process to improve the quality of the diesel pool, Akzo Nobel Catalyst Symp., June 1994, page 77. 31. T. Schockaert and J.W.M. Sonnemans, The flexibility of the Akzo-Fina CFI process, Akzo Nobel Catalyst Symp., May 1991, page 187. 32. W.M. Gregory and D.S. McCaffrey, Technology for the future: deep desulfurization to meet ultralow sulfur levels in distillate product, 7 th Refinery Technology Meeting, Dec. 1993, Bombay, India. 33. A. Suchanek, Reduction of aromatics in diesel fuel, NPRA annual meeting, paper AM 90-21. 34. A. Suchanek and G. Hamilton, Update on diesel deep desulfurization/aromatics saturation by Synsat process, NPRA annual meeting, paper AM 92-19. 35. Oil and Gas Journal, Oct. l lth, 1990; Oct. 14th, 1991; Oct. 12th, 1992; Oct. 11th, 1993; Oct. 10th, 1994. 36. P. Trimmer, Current product demand trends. The implications for refinery operations, by product and volume, the 9th European Petroleum and Gas conference, Nov. 1994, Amsterdam, The Netherlands.

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Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All rights reserved.

117

OPTIMIZING HYDROTREATER CATALYST LOADINGS FOR THE UPGRADING OF ATMOSPHERIC RESIDUES Jesper Bartholdy and Barry H. Cooper Haldor Topsoe A/S, Nymollevej 55, DK-2800 Lyngby, Denmark ABSTRACT The present paper gives an overview of the effect of catalyst bed composition on performance and catalyst life in residual oil upgrading. Examples of gradings are given, where a proper grading can increase the catalyst life by more than 50%. The effect of having very large pores present in the catalyst bed has also been investigated. It was found that the mass transport in 0.5 mm holes in the catalyst pellet was diffusionally limited. INTRODUCTION The use of composite catalyst fillings is the industrial standard when treating resids in a modem refinery. The grading can be made both with respect to size, shape and catalyst properties in order to optimize the performance, obtain the desired product quality and the longest catalyst life. The optimization is often a careful balancing of catalyst deactivation and performance, and it is therefore important to understand the mechanisms of deactivation. For this reason, a brief introduction to the various aspects of catalyst deactivation will be given: what deactivates the catalysts and how is catalyst deactivation generally tied in with catalyst properties? EXPERIMENTAL The resid deactivation experiments were conducted in a 250 cc isothermally operated test unit. The catalyst bed was diluted with 0.2-0.3 mm glass beads in order to minimize the effect of liquid maldistribution. Products were collected from the unit on a daily basis, and important product properties were determined using the methodology described in a previous publication [ 1]. After the experiments, the catalyst bed was transferred into a glass tube. The catalyst bed was divided into sections representing a distinct part of the catalyst bed. Spent catalyst from each section was soxhlet extracted for 24 hours in xylene and dried for 4 hours at 180~ prior to microprobe analysis of Ni and V distribution using a Jerol JXA-3A electron microprobe and chemical analysis of spent catalyst composition. From the microprobe analysis, the metal penetration parameter Qv can be derived. This parameter is defined as follows: Qv = Average V concentration in pellet Max V concentration in pellet The experiments were conducted on a Kuwait atmospheric resid with characteristics as given in Table 1. The catalysts tested were commercial resid desulfurization catalysts in the shape of 1/32" cylindrical extrudates. Important properties of the catalysts have been listed in Table 2.

118 Table 1. Properties of Kuwait Atmospheric Resid Cut: S Ni+V CCR

650 ~ 4.3 wt% 82 ppm 10.5 wt%

Table 2. Properties of Catalysts Tested Catalyst name Carrier Mo wt% Ni wt%

Pore radius, ! Surface area, m2/cc V-penetration, Qv HDS act. HDM act. Metal capacity

TK-711

TK-751

TK-771

A1203

A1203

A1203

4.0 1.7 75

6.5 1.9 62

9.5 2.8 51

150 0.7 Moderate

190 0.4 Medium Medium Medium

225 0.2

High High

High Moderate Moderate

In one experiment, a few small ring-shaped parti-cles were placed on top of the reactor. Some of the rings were sealed in one end, others were open in both ends. After the experiment, where Ni and V were allowed to deposit, the particles were soxhlet extracted and prepared for microprobe analysis of V-distribution. The V-distribution was determined at two distances, 2 and 4 mm from the pellet end. Pellet dimensions and positioning of microprobe scans are given in Figure 1.

Figure 1. Dimensions of Ring-Shaped Catalyst Particle with one Sealed End for studying Diffusion in Large Pores.

119 Table 3. Catalyst Configurations Investigated.

TK-711 (HDM) vol% TK-751 (HDM/HDS) vol% .TK-77 ! (HDS) vol%

I:IDS 7 93

LOADING TYPE EIDM 38 62

Optimum 7 31 62

The experiments were conducted in constant HDS mode of operation, i.e. the temperature was adjusted to give 0.5 wt% S in the fuel oil. In the experiments relating to the influence of reactor loading on the performance, three different loading configurations have been tested. The composition of these is given in Table 3. CATALYST DEACTIVATION Resid desulfurization catalysts are deactivated simultaneously by both coke and metals. Which is the dominating factor depends on the type of operation and its severity as well as on the properties of the feed. In a recent publication [2], examples are given where more than 50% of the deactivation observed in upgrading &Middle East resids can be ascribed to catalyst coking rather than to metal deactivation. However, these findings are the exception; normally, coke deactivation is responsible for the least part of the deactivation. As a catalyst poison, coke deposition is different from metal deposition in several respects, e.g. the rate of contaminant build-up and the contaminant distribution over the surface. On exposing the catalyst to the feedstock, an initial coke layer is formed [3]. For catalysts operating on resid, the initial coke laydown corresponds to more than one monolayer of coke. Such quantities will inevitably cause deactivation but exactly how much is hard to say as the initial coke is formed in a few hours, before an activity measurement can be made. In the subsequent initial phase of a resid upgrading operation, catalyst deactivation is rapid. In the run plots given later, the initial deactivation could last 1-2 months depending on the loading configuration. This initial deactivation is not caused by catalyst coking alone but predominantly by metal deposition as described in a recent publication [4]. After the initial phase, coke is also laid down and causes deactivation of the catalyst. The amount and the rate of laydown are controlled predominantly by the operating temperature [2]. The demetallation of a resid causes a steady accumulation of metals on resid catalysts. Among the metals, Ni and V are the most abundant. These accumulated metals reduce the (HI)S, HDN and HDCCR) activity of the catalyst in two ways: a) by coveting the active surface, b) by completely blocking the access to the catalyst system during the last part of the run (pore mouth restriction) [5]. The HDM activity is affected in a different way by accumulated Ni and V on the catalyst. DemetaUation is autocatalytic in nature, i.e. deposited Ni and V have a significant HDM activity [6]. Deposited Ni and V do not deactivate the HDM activity by coverage of the active surface but rather by restricting the access to the catalyst panicle.

120

0.8

KuwaitAR 1/32"C

0.6

o/

./0 /

/

ArabianHeavyAR

I

1/32"C

.m

~ j ~ ~ , 0.2 i~"

'~TK-771 "%"

ArabianHeavyAR 1/16"C %

0.0 40

50

60

70

80

90

\

100

Pore Radius,A

Figure 2. Influence of Particle Size, Feedstock and Pore Size on Qv.

Ni + V on Catalyst

Figure 3. Deactivation of Different Catalyst Types, TK-711 (HDM,, Q,,=o.7), TK-751 (~DM/HDS, Q,,=0.4) and TK-771(HDS, Q,,=0.2)

Metal deposition in a catalyst pellet yields a characteristic deposition profile that is determined by the pellet and pore size of the catalyst particle, since the mass transport of metal bearing species is controlled by diffusion [7]. For convenience, many research groups have defined a simple parameter to characterize the deposition profile. In the present publication, this parameter is designated Qv and has been defined in the experimental section. High Qv means uniform deposition over the whole cross section of the catalyst pellet, whereas low Qv means that the deposition of metals mainly takes place in the outer layer of the catalyst particle. The value of Qv is predominantly controlled by the pore and particle size of the catalyst, but feedstock properties also have some impact as can be seen from Figure 2, showing the variation in Qv with pore size for two Middle East feeds: Arabian Heavy and Kuwait AR. For the Arabian Heavy, Qv vs. particle size is also included. The value Qv significantly influences the catalyst deactivation rate. This is illustrated in Figure 3. TK-711 in Figure 3 is a high Qv HDM catalyst, Qv~0.7. For catalysts with a high Qv, the contaminants are spread evenly across the pellet diameter, which gives a high storage capacity. However, the activity loss is high due to contamination of all sites in the particle. For low Qv catalysts, the inner part of the particle is left uncontaminated, which gives a low storage capacity but good activity retention. In Figure 3, the deactivation curve TK-771 is an example of the deactivation of a low Qv catalyst, Qv~0.2. During a resid hydroprocessing run, large amounts of metals and coke are accumulated. For catalysts situated in the top of the loading, the metal deposits not infrequently add up to 60-70 wt%, and in the bottom, 10-20 wt% Ni and V on a flesh catalyst basis are accumulated. These amounts can, if distributed evenly, cover the entire catalyst surface area with several monolayers. In spite of this, the catalyst is still active. This is due to a combination of things: only the outer part of the catalyst particles is heavily contaminated, and the particle center is protected from contamination; the metal deposits have some HDS activity and in particular HDM activity [6]; and the contaminants do not form a full monolayer but are rather deposited in random positions, ensuring that the active phase is accessible even at high contaminant loadings, i.e. poison deactivation (catalyst fouling) as suggested by J. Wei [8].

121 410 0 o

Max operating temperature o

=

,

#P

=po=, o

400

HDM,, ~ ~

L (1) Q.

F:

,

~,~,#"f OPTIMUM

390

"0

| ._N

380

tO

E

370

p~s ~

r

~~

HDS

Z

360 2

4

6

8

10

Months

Figure 4. Temperature Required for Producing 0.5 wt% Sulphur Fuel Oil. Influence of Loading Configuration. RESULTS AND DISCUSSION

The composition of reactor loadings for upgrading of atmospheric resid requires a careful balancing of metal capacity on the one side and adequate activity for the desired reactions on the other. These considerations can best be illustrated by examples of how reactor performance changes with catalyst configuration. Three loadings have been tested, the composition of each is given in Table 3. The loadings designated "HDS" and "HDM" are loadings with only two types of catalyst: TK-711 and TK-771 (HDM and HDS catalyst, respectively). The "HDM" loading uses a fairly large volume ofTK-711 on top of TK-771. In the "HDS" loading, the TK-771 layer is expanded at the expense ofTK-711. In the "optimum" loading, a layer ofTK-751 is introduced in between TK711 and TK-771. The performance of the three loadings is shown in Figure 4. The activity is given as the temperature normalized to maintain a constant product sulfur level, 0.5 wt%, in fuel oil. The "HDS loading" has the lowest SOR temperature of the three. This is to be expected since TK-771, the active HDS catalyst, occupies more than 90% of the reactor volume. The "HDS" catalyst system is rapidly deactivated since the pore system of the small pore (low Qv) HDS catalyst is readily plugged. This is because the layer of TK-711 is too small to offer protection for the subsequent TK-771 catalyst. The "HDM loading" has a higher SOR temperature because a large fraction of the catalyst bed consists of TK-711, a large pore and low surface area demetallation catalyst. The "HDM loading" deactivates faster than the other loadings, which reflects that the metal contaminants penetrate and deactivate the entire surface of the high Qv demetallation catalysts. On the other hand, the "HDM loading" has a high capacity for metals, and the catalyst following the demetallation catalyst is well protected. The run plot of the "HDM" loading, Figure 4, shows no sign of any development of pore mouth restrictions (upward swing towards EOR). However, the capacity improvements made by using a large volume of TK-711 do not yield a longer life, since the maximum operating temperature is met before the capacity for metals is used up. In the optimum loading, the amount of HDM catalyst is reduced, and an intermediate catalyst, TK 75 l, is used. By doing so, the metal capacity of the entire loading can be kept high, whilst the deactivation by fouling is controlled by reducing the overall aveage Qv. As can be seen from Figure 4, the optimum loading not only has the longest life but also gives a full utilization of reactor metal capacity, the latter indicated by the upward bend of the deactivation curve at EOR. /

122

E Q. Q.

"~

20

HDS ~

,"

Q. > §

10 9 .. . . . . .

Z

2

HDM .= ,,., ,.,,,.,., ,,..,.,.,,..

4

6

8

10

Figure 5. Metal Content in Fuel Oil. Influence of Loading Configuration. The product metal content for the three runs is shown in Figure 5. The product from the HDS loading has a high metal content from the start of the run.This metal content rapidly increases during the run. This reflects the fact that we are carrying out demetallation of the resid on a catalyst system with predominantly small pores, i.e. low Qv and low HDM. As metals build up on the catalyst, the pore size of the pore mouth is reduced, decreasing the rate of demetallation further due to increased diffusional limitations. For the HDM loading we observe that the product metal content shows an initial decrease as a result of temperature adjustments made to compensate for loss of liDS activity by Ni and V deposition. As these deposits are active for demetallation, the increased temperature gives increased HDM. The optimum loading represents a compromise between the "HDM" and "HDS" loading with a medium product metal content, increasing towards the end of the run as the pore size of particularly the TK-751 and TK-771 is reduced and di~sional limitations develop. Optimizing reactor loadings using the three catalyst types described above can only be pursued to a certain point. For further improvement, more active catalysts are needed. Possible improvements are: better HDM activity, higher metals capacity or a generally higher activity for HDS, HDN, HDCCR, etc. It is, however, important to realize that the improvement of one catalyst functionality (e.g. HDM) should not be made at the expense of another catalyst functionality (e.g. HDS), as a lower activity would often require that the operating temperature is increased which could accelerate the coke deactivation and thus shorten the run [2]. It is possible to develop more efficient HDM catalysts. The HDM reaction is diffusionally limited, and an increase of the effectiveness factor of HDM would give an increased HDM activity. An improvement of the effectiveness factor of HDM can be obtained in several ways: by reducing the particle size, widening the pore size [ 7] or introducing a network ofmacropores to facilitate easy access of the metal bearing species to the catalyst pore system. Reduction of particle size can be a successful route to more efficient HDM catalysts since this results in an increased Qv as is evident from Figure 2. However, pressure drop constraints in industrial units often limit this option considerably. Pore enlargements can also improve the metal diffusion and thus demetallation. However, for the largest pore catalysts, pore enlargement will result in a marginal increase of Qv (metal capacity and HDM activity, see Figure 2), but at the same time, the surface area is reduced and thus the activity for the reactions that are not diffusionally limited (HDS, HDN, etc.). If Qv is low, pore enlargements can have a net positive effect on metal capacity and HDM activity.

123

r

2mm

L_

from

entrance

to

E

~.~.~_ _ -

-

~

.o

f._ r

I~

4 m m from entrance

E: (3) o E: o

E ::3 .m

Outer surface

Inner surface

Figure 6. Profile of Deposited V on the Walls of a Ring-Shaped Particle. Profiles: 2 and 4 mm from the Ring Entrance. ( - - ) closed ring and (---) open ring. Improving the demetallation and metal capacity by introducing macropores to the catalyst is an intriguing option. The macropores could facilitate an easy access of large molecules to the catalyst and thus improve the effectiveness factor for HDM. However, our work has shown that the effectiveness factor for HDM (and thus Qv [9]) can only increase to a certain value (Qv ~ 0.7-0.8) independent of the pore size and the amount ofmacropores present because bulk diffusion also limits the mass transport. This finding has been illustrated by the following experiment. Tests were carried out on catalyst particles in the shape of small rings, as described in the experimental section. Some of the tings were sealed at one end (see Figure 1), giving in effect a particle with a bimodal pore system, where the center hole acts as one big "macropore". When such a particle is exposed to residual oil in a reactor, oil flows on the outside and metal compounds from the oil diffuse into and are deposited on the pellet walls. The inner parts of the particles, however, can mainly be reached through the center hole, through which the oil has to diffuse since one end is sealed. The "macropore" is 0.5 mm in diameter and therefore no diffusional limitation was expected in the "pore". However, from the microprobe analysis of the spent catalyst (Figure 6), it is clearly seen that less V is deposited from the inside than from the outside, decreasing from a distance of 2 to 4 mm from the entrance. In Figure 6, the dashed curve gives the deposition profiles obtained on tings with open ends, which shows almost the same deposition from the inside as for the outside, demonstrating that the center hole is accessible during the experiment. The fact that the amount of V decreases when the center hole is closed at the one end suggests that - for large pores - the mass transport is no longer controlled by hindered diffusion but rather by bulk diffusion. This gives an explanation of why For low Qv catalysts, controlled by hindered a Qv equal to unity can never be achieved.diffusion, introduction of macropores can give improvements, but for a large Qv HDM catalyst, macropore addition does not improve diffusion as the mass transport is limited by bulk di~sion in the pores, i.e. unaffected by the pore size of the catalyst.

124 355 0 01d Ilen~mtien

-.~

[

e

_

315

375

=..,..,.

1D II) ._N m

_.,

,-

"

"

"

-

9,p., "

,....

~

=., _

_

e

== o

Nowgellei'~tJli'l catalysts

365 Z

355 l

1l ( l !

20(10

3(IH

Run Hours

Figure 7. HDS Activity of New Generation Catalysts compared with that of Previous Generations. Temperature Required to give 0.5 wt% Sulphur Fuel Oil.

In recent years, better catalysts have been developed with a better activity for HDS, HDN, HDCCR, etc. To a less extent, the asphaltene removal and demetallation have been improved. The improvements in HDS activity are substantial as can be seen in Figure 7, comparing the activity of previous catalysts with that of new generation catalysts. Similar improve-ments have been achieved for HDN and HDCCR activities. A low product metal content can be achieved using the new generation high HDS activity catalyst types, utilizing the grading technology described previously. The higher HDS activity of the catalysts can justify the use of less HDS and more HDM catalyst, giving the same HDS performance as the case designated "optimum" but with a better HDM activity CONCLUSION Catalyst grading is an essential tool for tailoring catalyst loadings to treat atmospheric resids. Catalyst life and performance can be balanced to give the desired product quality, achieving maximum run length at the same time. The demetallation reaction is diffusionally limited and thus requires large pores to achieve unhindered access to the pore system. However, for catalysts with large pores and even macropores, where hindered diffusion should not limit the mass transport, diffusional limitations are seen as bulk diffusions become limiting. REFERENCES 1.

2. 3. 4. 5. 6. 7. 8. 9.

A. Nielsen, B.H. Cooper and A.C. Jacobsen, ACS Atlanta Meeting, 26,, 440 (1981). J. Bartholdy and B.H. Cooper, ACS Div. Petrol. Chem., 38, 2, 386 (1993). P. Wiwel, P. Zeuthen, A.C. Jacobsen, Stud. Surf. Sci., 68, 257 (1991). G. CJualda and S. Kasztelan, Stud. in Surface Science and Catalysis, 88, 145 (1994). P.N. Hannerup, A.C. Jacobsen, ACS Div. Petrol. Chem., 28, 3, 576 (1983). C. Takauchi, S. Asaoka, S. Nakata and Y. Shiroto, ACS Div. Petrol. Chem., 30, (1985).(NY) S. Kobayashi et al., Ind. Eng. Chem. Res., 26, 2241 (1987). J. Wei, Stud. Surf. Sci., 68, 333 (1991). J. Bartholdy, P.N. Hannerup, Stud. Surf. Sci., 68, 273 (1991).

96

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All rights reserved.

125

H Y D R O T R E A T M E N T OF RESIDUALS USING A SPECIAL NiMo-ALUMINA CATALYST. Alfredo Morales and Rodoifo Bruno Solari

INTEVEP S.A., Refining and Petrochemical Division, Apartado Postal 76343, Caracas 1070A, Venezuela. ABSTRACT This article presents experimental results obtained at bench scale using a special NiMo-Alumina catalyst developed by Intevep, S.A. to hydrotreat and demetallize deasphalted oils, heavy crude oils and residuals. Using this catalyst to process residuals, it is possible to reach 92% demetallization under stable operating cycles of more than six months with a feedstock containing less 600 ppm of metals (V+Ni), 17 % asphaltenes, 16% Conradson Carbon and API gravity increment higher than 4 degrees, under moderate operating conditions (1500 psi, 340~ LSHV = 1 h"1). The technical and economic feasibility of using this catalyst in a conventional refinery to process residuals, is discussed. I. INTRODUCTION All around the world there exists an installed capacity to process residue utilizing solvent deasphalting as a carbon rejection technology followed by hydrotreatment of the deasphalted oil (DAO) or the residual. Severity of the hydrotreating stage depends on the downstream use of the hydrotreated DAO or the residual, which can be used as feedstocks to catalytic cracking/hydrocracking or as components of low sulfur fuel. In refineries with existing deasphalting capacity and/or idle hydrotreating and catalytic cracking capacity, it is possible to implement low investment projects to increase the feed to the downstream conversion units. In fact, by increasing the yield of deasphalted oil and/or hydrotreating it, or by means of direct hydrodesulfurization of the residue, it is possible to feed a fraction of these streams to catalytic conversion units, such as catalytic crackers or hydrocrackers. However, the high content of metals, Conradson Carbon and nitrogen in these feedstocks could decrease the project profitability. These type of processes have been practiced on a commercial scale for approximately twenty years. As examples, we have the Corpus Christi Refinery in Texas [1, 2]. and BarrancabermejaRefineryin Colombia [3]. In most of the cases, an increase of the heavier fraction components in the conversion process feedstocks is limited by the catalytic system being used, which is usually affected by the additional amount of metals and Conradson Carbon in the feed. This restriction limits significantly any increase in the amount of deasphalted oil in the residue fraction that could be added to the refinery stream. This paper presents a catalyst that is able to incorporate a higher fraction of heavier feedstock to the conversion system. The NiMo-Alumina catalyst permits process scheme adaptations to achieve these objectives in an economic way, depending upon the particular refinery and its product slate to satisfy the market requirements.

126 Table 1 Catalyst properties of NiMo-alumina (INT-R1R)

Type Pellet size (inches)

NiMo-A! 1/20"

Surface area (m2/g)

140

ERD (g/cm 3)

4.6

Bulk density (g/cm 3)

0.6

Pellet length (mm)

4.5

Bulk Crush Strength (kg/cm 2)

7.8

2. EXPERIMENTAL The catalyst has been developed by Intevep, S.A. [4,5] and is being commercialized by AZKO NOBEL Chemie from Holland under the trademark INT-R 1 a. High molecular weight compounds are able to be processed with the catalyst due to its special pore distribution. Therefore, it can be used to treat feedstocks with a significant asphaltene content without losing its high desulfurization and demetallization activity. Also, the catalyst has a high metal retention capacity and allows for moderate conversion of the 510~ + fraction. INT-R1 a catalyst physical properties are shown in Table 1. The hydrotreatment was carried out in a fixed bed down flow reactor of approximately 5 l/day capacity. The oil feed and hydrogen were premixed before entering the reactor. The light gases were separated from the liquid product in a high pressure separator. The liquid, after been stripped with nitrogen, was analyzed. The operating conditions for the isothermal reactor were 1500 -1800 psig; space velocity of 1 v/v.h; H2(NPT)/feed of 1000 v/v; the temperature range was 360~ to 420 ~ Metal removal was followed by analyzing daily feed and product vanadium content. 3. RESULTS AND DISCUSSION The INT-R 1 a catalyst has been tested with extremely difficult feedstocks regarding their metal, Conradson Carbon and asphaltene content. Pentane and hexane deasphalted oils from extra heavy crude oils, and atmospheric residue of these crudes were among the feedstocks tested [6,7]. The INT-R1 R catalyst can be used in fixed bed hydrotreating units to improve the residue quality, from a high sulfur residual to a low sulfur fuel oil [6]. Tests were performed to demonstrate the technical feasibility of directly processing these residues, using atmospheric residue of Cerro Negro and Iranian Gach Saran crude oils as feedstocks. These residues are characterized by a higher content of asphaltenes and Conradson Carbon than the deasphalted oils [5,6]. Therefore, the effect of these two variables on the performance of the catalyst, can be evaluated.

127 Table 2. Feed properties

API gravity Sulfur (% wt) Nitrogen (ppm) Conradson carbon (% wt) Asphaltenes (% wt) Vanadium (ppm) Nickel (ppm) Viscosity (cSt) 100~ Distillation (% wt) IBP- 190~ 190- 270~ 270- 343~ 343 - 510~ 510oC+

Residue 350~ Gach Saran 12.7 2.6 4,690 10 4.2 191 62 71

0.0 0.0 1.0 41.0 58.0

Cerro Negro Crude oil 10.7 3,3 5,550 11.0 8.0 400 73 540

0.0 2.0 20 20 58.0

Residue 350~ Cerro Negro 4.6 4.1 7,305 17.9 16.9 548 117 3,270

0.0 0.0 2.0 26.5 71.3

Feed properties are shown in Table 2. In this case, more severe operating conditions are required and the pressure level is increased up to 1800 psi to avoid coke deposition on the catalyst surface. Gach Saran long residue have lower metals, asphaltenes, sulfur and nitrogen than Cerro Negro long residue indicating that this feedstock is a good example of a long residua easy to process. On the contrary, Cerro Negro residue is an example of an extremely difficult feedstock due to its chemical properties while the complete Cerro Negro crude oil was considered as an intermediate feedstock, with a Conradson carbon content (11%)a little higher than the Gach Saran residue but double the asphaltene(8. 5%) and metal (473 ppm) content. Table 3 summarizes INT-R1 R catalyst performance when processing each of these feedstocks. Although these results cannot be directly compared due to differences in the operating conditions, there is a trend that is worth to be mentioned. The HDM and HDS functions are reduced when processing residuals as compared to the DAOs, clearly showing the effect of the Conradson Carbon and asphaltene content over the HDS and HDM functions [6]. The same can be observed regarding the catalyst capacity to convert asphaltenes and Conradson Carbon which decreases as the concentration increases. However, the INT-R1 catalyst is able to keep HDS and HDM activity levels higher than 60% and 75%, respectively, even at the highest level of asphaltene (17%) and Conradson Carbon (18%) content, at moderate operating conditions (LSHV =1 h-l, T = 400~ This different catalyst behavior with diverse type of feedstocks indicates that operating conditions should be adjusted to each type of feedstock to fit the required severity according to the objectives set for the hydrotreating process.

128 Table 3. INT-R1R catalyst activity using different feedstocks. Gach Saran residue 350~ +

Cerro Negro crude oil

Cerro Negro residue 350~ +

Hydrodesulfurization (% wt)

77

61

63

Hydrodemetallation (% wt)

81

77

77

Hydrodenitrogenation (% wt)

32

Asphaltenes conv. C7 (% wt)

64

56

46

Conradson carbon conv. (% wt)

43

41

32

Fraction 510~ +conv. (% wt)

34

34

41

1

0.3

1

400

390

400

15

150

15

LHSV h -1 Temperature (~ Days of run

,oo

60 .I 40 -

NV.

20

i

I

5

I0

% ASPHALTENES

I

I

15

20

IN FEED

Figure 1. Catalytic activities as a function of asphaltene content Fig. 1 shows the effect of the asphaltene content over the catalytic activity for HDM, HDS and Carbon Conradson conversion. This effect being more severe in HDS function than in the HDM function, varying from a HDM/HDS selectivity ratio of 1.0 to approximately 1.5. Conradson carbon conversion capacity is also affected by the high asphaltene content. However, figure 1 also show that these functions tend to stabilize as the asphaltene content

129

80-

jT=380-390~

LHSV=O,3m3/m3.

h; P = 1 8 0 0

psig;Hj/FEED=IOOONm3/m3~.,

" 700

O

60

0

o

0

t)

50.4 % v

80 ~

.,,C

=

0

v

n

(1

0

'~

O

7"

C

0

0

70"

I

0

20

I

40

I

I

60

80

I

100

I

120

I

140

I

160

I

180

I

I

200

220

f D~21~ )

Figure 2. Long Term Test using Cerro Negro Heavy crude oil. increases even for values close to 20%. This confirms the catalyst stability in a long term run using a feedstock with high asphaltene content, as shown in Fig. 2, corresponding to a long term test (220 days) using Cerro Negro crude oil as feedstock. In this case the HDM function was constant along the test, reaching up to 50% metal content on the catalyst. At the start of run, high HDS was observed but atter about 50 days HDS was stabilized at approximately 61%. This behavior is probably due to coke deposition over the catalyst surface that tends to decrease its desulfurization capacity. This different catalyst behavior with diverse type of feedstocks indicates that operating conditions should be adjusted to each type of feedstock to fit the required severity according to the objectives set for the hydrotreating process. From Fig. 3, it can be predicted catalyst life cycle for operations with the two types of residues being studied. At high severity (0.3h-1 space velocity, 1800 psig) and using Gach Saran long residue, it is possible to reach a 14 months cycle at 75% demetallization while operating with Cerro Negro long residue this operating cycle is reduced to 7.5 months. These forecasted life cycles obtained from the experimental data indicate a profitable operation of INT-R1R catalyst, despite the difficulties derived from the type of feedstock used in the process. Catalyst life prediction has been done using the procedure proposed in ref. [8] based on shorter screening period for demetallization catalysts. In any commercial application, the process economy depends mainly on catalyst life cycle which in turn is a function of catalyst metal retention capacity. The experimental tests have shown that INT-R1 R catalyst is able to accept a metal content equivalent to 100% without losing significant catalytic activity. Nevertheless, these results guarantee a stable operation with

130 F-

~" 100" t.9 IM

~

80-

Z 0

Z

uJ 60U,I .J

N 4020-

0

0

i .......

2

I"

4

'I

6

I

8

I

10

I

12

I

14

LIFE CYCLE (MONTHS)

Figure 3. INT-R1R life cycle operating with Cerro Negro and Gach Saran long residues. catalyst cycles of more than 12 months when using lighter feedstocks at a moderate severity, which makes the commercial application of this catalyst economically attractive.

3.1. Refinery Applications Results obtained from the tests using atmospheric residue from heavy crudes show the feasibility of reaching long term operating cycles with INT-R1R catalyst. However, these types of feedstock are not commonly found in refineries. In conventional refineries processing lighter feedstocks, the long residue is always fed to a vacuum distillation unit. In these cases, it is possible to envisage three process schemes to increase the amount of feedstock to FCC using an existing hydrotreating unit designed to operate at 1200-1500 psi, filled with lNT-R1R catalyst. The first process scheme considers an existing deasphalting unit that could be adapted to increased DAO yield, either by increasing solvent molecular weight, or changing operating conditions. DAO is hydrotreated using INT-R1R catalyst in a fixed bed reactor to reduce Conradson carbon and metal content and to yield a feedstock stream that can be mixed with the virgin VGO and send to a FCC unit. This scheme would also increase the refinery operational flexibility to incorporate a larger amount of heavy crude oil in the refinery feedstock. Table 4 presents results obtained when processing a metal-rich vacuum residue of 7 ~ API. Deasphalting this residue with pentane produce a 74% of low asphaltene DAO, which is hydrotreated to obtain a 18~ product with low metal and sulfur content. This product is mixed with virgin VGO to obtain, from the long residue, 30% wt of additional feedstock to FCC that meets metal, nitrogen and Conradson carbon specifications. This operation scheme is economically attractive due to high DAO yield.

131 Table 4. Increase of FCC feedstock via deasphating.

API gravity Sulfur, %wt Nitrogen, %wt Conr.C., %wt Metals, ppm Asphaltenes, %wt Yield over reduced crude, %wt

VGO

Vacuum Residue

DAO-C5

DAO-HDT

VGODAO-HDT

19 2.5 0.04 0.50 61

7 4 0.3 20.7 430 10 39

10 3.9 0.24 12.1 86 0.5 29

18 0.62 0.18 5.03 7 0.1 30

18.7 0.2 0.086 1.99 2.3 0.03 91

The second process scheme considers that a fraction of vacuum residue is segregated to be mixed with virgin VGO, and treating the mixture through a hydrotreating bed of INT-R] n catalyst. A mixture of 75% by weight of virgin VGO and 25% of vacuum residue yields a 16~ API feedstock with 5.5% CC, 2.5% asphaltenes and 100 ppm of metals to be fed to the hydrotreating unit, as indicate in Table 5. The mixture, after being hydrotreated, is incorporated as additional feedstock to FCC. Taking the atmospheric residue as the base case, a 22% net increase in the FCC feedstock is achieved, reducing asphalt production to only 19% as compared to the original 39% produced in the vacuum distillation unit. Hydrotreatment of this mixture yields a 22 ~ API product that meets specifications as feedstock to a modem FCC unit. If this product quality is compared to that obtained in the previous scheme, it could be seen that the main difference is in the asphaltene content, making Table 5. Residual feedstock to FCC via hydrotreatment. VGO

Vacuum residue

VGO / residue 75/25

Hydrotreated product

API gravity

19

7

16

22

Sulfur (% wt)

2.5

4

2.87

0.7

Nitrogen (% wt)

0.04

0.3

0.105

0.080

Conradson Carbon (% wt)

0.50

20.7

5.5

2.60

Metals (ppm)

-

430

106

9

Asphaltenes (% wt)

-

10

2.5

1.05

61

39

81

82

Yield over reduced crude (%wt)

132 this feedstock more difficult to be processed. Also, the increment in the FCC feedstock is approximately 9% lower. These two main differences make the deasphalting plus hydrotreating scheme economically more favorable than the second scheme. However, it should be considered that in the first scheme the required investment is higher. In the second application, investment is low, being limited to minor cost changes to adapt the hydrotreating unit to process the VGO/residue mixture. In this case, operating costs will be affected only by additional hydrogen and catalyst consumption, due to the higher asphaltene content in the HDT feedstock. If a refinery has FCC idle capacity and an existing hydrotreating unit that is available to operate at 1200 psi, this second alternative could significantly increase the refinery net profit by increasing the amount of distillates and reducing production of high sulfur fuel.

The last scheme application considers to add INT-R 1R as a top bed of the an existing hydrotreating unit for protection of the HDS conventional catalyst. In this case virgin VGO could be mixed with HKGO and HKN obtained of cracking process to reach a feedstock containing between 3 to 7 ppm of metals, and be hydrotreated using a conventional HDT catalyst protected in the top of the reactor by INT-R 1 R catalyst. Commercial operation using this scheme has been evaluated in PDVSA affiliates improving the operating cycles of the conventional HDT catalyst. 4. CONCLUSIONS Experimental results show the high stability of INT-R1 R catalyst when processing feedstocks with high metal, Conradson Carbon, and asphaltene content. Operating cycles of a least six months were demonstrated at bench scale using feedstocks with 400 ppm metals, 8 to 10 wt% Conradson Carbon and 8 %wt asphaltenes. Expected INT-R1 R catalyst life with lighter feedstocks shows the feasibility of reaching a stable operation for more than one year with up to 100% metal retention on the catalyst. The utilization of INT-R1 R catalyst to produce additional FCC feedstock could be economically attractive to refineries running with idle capacity in FCC and hydrotreating units. With a low investment it is possible to incorporate up to 50% of vacuum residue as FCC feedstock by hydrotreating directly a segregated fraction of the vacuum residue. In refineries where it could be possible to modify existing deasphalting units, the additional FCC feedstock increment could be even larger, reducing asphalt production to only 10% of the atmospheric residue. These examples show a high economic potential in the utilization of the INT-R1 R catalyst allowing the refinery to incorporate additional feedstock to FCC. REFERENCES 1. E. Le Roi, J. Hutchings, G. Sikonia and R. M. Ponder, Demex process proves successful, NPRMeeting, San Antonio Texas, March 19-21 (1978). 2. P. Pemming, A.G. Vickers and B.R. Shah, The increasing importance of solvent extraction for heavy oil conversion. 47th. Midyear Refining Meeting, New York, American Petroleum Institute. May 11 (1982).

133 3. R. G. Zambrano, Experiencias en el procesamiento de residuos de vacio en el Complejo Industrial de Barrancabermeja. XLIX Arpel Experts Meeting, Rio de Janeiro, August (1982). 4. A. Morales et al., US Patent No. 4,642,179, granted to INTEVEP, S.A., (1986). 5. A. Morales, R. Galiasso, D. Huskey, R. Carrasquel, HDS and HDN Catalyst to Hydrotreat Heavy Crude Oils, XLIXArpel Experts Meeting, Ciudad de M6xico, May 1983. 6. B. Solari and A. Morales, Vision Tecnologica, Vol.2-1 (1994) 19. 7. H. Kum, J.J. Garcia, R. Galiasso, L. Caprioli, A. Morales, A. de Salazar, Hydrotreatment of heavy Crude oils and residues. Rev. T~c. INTEVEP, 5(1) (1985) 17. 8. L. Reyes, C. Zerpa and J. Krasuk, B. Delmon and G.F. Froment (eds.), Catalyst Deactivation 1994, Vol 88 (1994),85. ACKNOWLEDGMENT Permission to publish this paper, by PDVSA and INTEVEP S.A., is gratefully acknowledged

This Page Intentionally Left Blank

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All rights reserved.

135

RESIDUE HYDROPROCESSING: DEVELOPMENT OF A NEW HYDRODEMETALLATION (HDM) CATALYST Opinder K. Bhan 1 and Safa E. George 2

1Shell Development Company, Westhollow Technology Center, 3333 Hwy. 6 South, Houston, Texas 77082, USA 2Criterion Catalyst Company, 16825 Northchase Drive, Two GreenspointPlaza, Houston, Texas 77060, USA ABSTRACT Increased emphasis is currently being placed on development of improved hydrodemetallation (HDM) catalysts, that maximize metal removal without deactivating excessively, and thus effectively protect the hydrodesulfurization catalysts in fixed-bed multiple reactor systems. With the greater research emphasis in this area, HDM catalysts are emerging as increasingly distinct from conventional hydroprocessing catalysts. By a combination of optimized catalyst pore size and structure, metal impregnation procedures, and improved catalyst preparation techniques, we have been able to develop a new HDM catalyst (RM-430), which combines high metal removal activity with high metal deposition capacity. In this paper we will discuss the results of testing of RM-430 with various feed-stocks, and also some of the analytical techniques used to analyse the feeds, reaction products, and the resulting aged catalysts. INTRODUCTION Residue hydrodemetallation (HDM) is gaining increasing importance with the emergence of residue catalytic cracking (RCC) as a viable heavy feed conversion process. Reduction in feed metal content has a significant impact on RCC catalyst replacement costs, and hence, on the overall refinery economics. In addition, the general trend in the petroleum industry for the past two decades has been towards the use of heavier crudes with higher metal and other contaminant levels. This has resulted in residue catalysts being operated at more severe processing conditions, necessitating the use of newer generation catalysts with greater capacity to tolerate higher level of metals and other contaminants. Catalyst manufacturers have generally kept pace with these more stringent demands of refiners, and new generations of catalysts are commercialized regularly. In this paper, we will discuss the development of a new HDM catalyst (RM-430), which has been developed to provide high metal removal activity and metal deposition capacity for fixed-bed residue processing. This catalyst has distinct physical and chemical properties which have been a result of incorporating recent advances in powder preparation technology, and improved active metal impregnation techniques. RESIDUE HYDRODEMETALLATION CATALYST DEVELOPMENT HDM catalysts are designed with the objective of maximizing metal removal and providing large capacity for deposition of metals inside catalyst pores. Since organometaUic

136 Table 1. Properties of Selected Residue Feedstocks. Heavy Arabian

Kuwait

Canadian Bitumen

API Gravity Viscosity (cSt) at 100~ Nitrogen, wt% Sulfur, wt% Basic Nitrogen, ppmw C5 Asphaltenes, wt% Conradson Carbon, wt% (MCR) Metals, ppmw Ni V Fe

11.2 231.0 0.354 4.41 824 15.5

13.4 56.0 0.267 4.29 580 12.1

9.7 174.0 0.439 4.10 1099 18.7

14.8

11.5

13.7

33 106 7

19 54 8

64 149 7

1000 ~

63.2

52.8

56.8

compounds contained in petroleum crudes are predominantly present in large asphaltene structures [1,2], it is imperative that the catalyst pores exert a minimum resistance to the diffusion of metal-beating molecules into the catalyst pore structure. Ideally, metal molecules should deposit uniformly inside catalyst pores. However, due to restrictive diffusion and the intrinsic surface activity, metals may be more densely deposited at the catalyst pore inlet. This causes a loss in metal deposition capacity due to constriction in catalyst pore diameter, and thereby shortens catalyst effective life. Increasing the catalyst pore diameter reduces the diffusional resistances and increases the metal penetration into catalyst pores. Increase in pore size over a certain size is neither desirable, due to loss in other catalyst activities, nor practical due to the significant loss in catalyst strength associated with the increase in pore diameter. Thus, an optimum in pore diameter exists, where catalyst HDM activity is balanced with other activities and catalyst physical strength. We have conducted detailed studies in our laboratories to study the effect of catalyst physical and chemical characteristics on metal removal activity, metal deposition capacity, and deactivation rates. Considerable research and modelling effort [3-5] has resulted in the definition of an optimum pore size for HDM catalysts for maximizing metal penetration into the catalyst interior, while maintaining high catalyst activity and mechanical strength. Besides pore constriction and blockage due to metal deposition, hydroprocessing catalysts also deactivate due to coke deposition [6]. An ideal HDM catalyst should minimize coke make, maximize metal deposition, and maintain activity for long run durations. RM-430 catalyst was developed with theses point in mind. In this paper, we will discuss some of the methodology that went in the development of this catalyst. E X P E R I M E N T A L METHODS Feedstocks

Several feedstocks covering a range of metal and other contaminant levels were used to develop and test RM-430 catalyst using a wide range of process conditions. RM-430 catalyst

137 was extensively tested with Heavy Arabian, Canadian bitumen, Maya, Kuwait and other atmospheric and vacuum residue feeds of Middle Eastern origin. Properties of some of these feeds are given in Table 1. In addition to routine chemical and physical analyses specified in the table, selected feed samples were also analyzed by a high temperature Size-Exclusion Chromatographic/Inductively Coupled Plasma-Mass Spectrometric (SEC/ICP/MS)technique to analyze differences in molecular sizes between the various feeds [7]. This information allowed us to more accurately optimize catalyst pore diameters for maximizing metal removal activity and deposition capacity.

Micro-Reactor Testing Catalyst testing was conducted in down-flow micro-reactors with once through hydrogen. In some cases, testing for as long as a year was conducted to evaluate catalyst long-term activity. Precautions for maintaining normal plug-flow distribution through the reactor system were taken. At the end of each catalyst cycle, the aged catalysts were removed, categorized according to their position in the reactor, and solvent extracted and oven dried. The washed and dried catalyst samples were analyzed for metal and coke content and the distribution of deposited metals.

Catalyst Features RM-430 demetallation catalyst is designed for maximum metal deposition capacity and metal removal activity and is made from Group IVB metals highly dispersed over a high surface area support. RM-430 catalyst has an optimal pore size distribution with very large pore volume, which provides it with a sizeable capacity for deposition of metals. The increase in pore volume, without the concomitant loss in catalyst pellet strength or surface area, has been accomplished by the use of specialized support materials and distinctive support preparation techniques. Properties of RM-430 and two other commercial Criterion Catalyst Company residue catalysts used in our studies are given in Table 2.

Aged Catalyst Metal Deposition Profiles Aged catalyst samples from the pilot-plant testing were analyzed for deposited metal profiles using a scanning electron microscope (SEM/EDX) fitted with an energy dispersive Xray analyzer (JEOL 8600 Microprobe and Noran EDX Analyzer). The information provided from this analysis technique regarding the penetration of metal species into the catalyst interior proved invaluable in tailoring catalyst pore structure for maximizing metal deposition capacity. RESULTS AND DISCUSSION

Kinetic Testing Heavy Arabian and Kuwait atmospheric residue feeds were used for evaluating kinetic parameters, temperature response, and metal deposition profiles for RM-430 catalyst. The data from these experiments indicated a one and one-half order dependence for metal removal at the process conditions tested. The temperature response of RM-430 catalyst is shown in Figure 1. The activation energy for vanadium and nickel were: Vanadium-36.1 kcal/mol, Nickel-27.3 kcal/mol.

138 Table 2. Hydroprocessing Catalyst Properties. RM-430 RN-410 Alumina Alumina Group VII and/or VI Metals TL TL 1.6,2.5 1.3,1.6,2.5 0.54 0.65 3.5 4.8 0.87 0.67 150 155 TL:TRILOBE*

Cartier Active Metals Shape Size, mm Bulk Density, g/cc Crush Strength, lbs/mm Pore Volume, cc/g Surface Area,

m2/~

RN-400 Alumina TL 1.3 0.65 5.0 0.67 220

*Registered Trademark 2.0 H e a v y A r a b i a n Long Residue Feed

E, K c a l / g m o l e

I ~ oo

.=43=- Vanadium - - O ' - Nickel

36.1 27.3

1.5

0 U

n.-

1.0

-

0.5

-

._~

0 ~ o _J

0

1.47

I

I

I

I

I

I

I

I

I

1.48

1.49

1.5

1.51

1.52 1/TX 103

1.53

1.54

1.55

1.56

1.57

Figure 1. Effect of temperature on activity. Figure 2 shows a typical electron scanning micrograph profile of vanadium and nickel deposited on aged catalyst pellets from a test run conducted with Heavy Arabian atmospheric residue feed. From this analysis and tests conducted with other feeds, we deduce that RM-430 catalyst has an effectiveness for vanadium and nickel removal of over 90%, and a total metal deposition capacity, based on fresh catalyst weight of over 100 wt%. RM-430 catalyst is therefore ideally suited for services where overall metal load limits operation. The pore diameter of RM-430 has been maximized in order to maximize the effective diffusivity of metal-containing molecules into catalyst pores. In addition to pore size and structure, metal deposition profiles have been improved for RM-430 by modifying the catalyst preparation techniques, i.e., catalyst support type and impregnation chemistry. We have optimized the impregnation chemistry of this catalyst to result in both a high metal removal activity and a high metal deposition capacity. The resultant hydrodesulfurization activity is relatively reduced.

139 V Line Scan ,

,

,

,

Ni Line Scan ,

,

,

,

,

,

,

,

,

,

,

,

,

,

,

,

'

'

'

'

I

'

'

'

'

I

i

[

i

i

,

,

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'

'

,

,

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'

'

'

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'

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i/,,

i

i

l

ii

,,i

i

i

ii

[

i , , , i

Pellet Diameter

,,

I

I

,

Pellet Diameter

Figure 2. Typical line scan &vanadium and nickel on aged RM-430 catalyst. It has been generally observed for Middle Eastern feeds that the catalyst metal deposition profiles (SEM/EDX) for nickel are steeperthan for vanadium leading to the conclusion that either the diEusional resistance to nickel-containing molecules is greater compared to vanadium molecules (implying difference in molecular size) or the reactivity of the two species is different. The SEC/ICP-MS profiles of nickel and vanadium compounds present in Heavy Arabian and Kuwait atmospheric residues are shown in Figure 3. The relative concentration of the large sized molecular species (with lower retention times) is l~gher for nickel-containing species than for the vanadium-containing compounds. Moreover, the relative concentration of metalloporphyrins is higher for vanadium than for nickel species. These profiles are consistent with the SEM/EDX profile of the metal deposition on catalyst. Clearly, the smaller metalloporphyrin vanadium species are able to penetrate deeper into the catalyst interior Our work using SEC/ICP-MS has also indicated that in hydroprocessed liquid samples, the smaller molecules are consumed at relatively milder process conditions, whereas the large molecular species require more severe process conditions. Based on these results we deduce that the larger nickel molecules are relatively more diEusionally hindered versus the vanadium molecules. Effect of Temperature Hydrodemetallation tests were conducted with Kuwait atmospheric residue feed at operating conditions selected to simulate the guard reactor in a typical residue processing service. In one such test, the temperature was initially held constant at a typical residue startof-run (SOR) temperature condition for nearly three and one half months and then raised to a typical middle-of-run (MOR) temperature condition and maintained at that level for an additional equivalent length of time. In an another similar experimental run, SOR temperature was maintained for nearly two months and then the temperature was raised to a typical end-of-

140

Heavy Arabian Atmospheric Residue Nickel Concentration (1000 COUNT/SEC)

Vanadium Concentration (1000 COUNT/SEC)

70 60

"-"U

,o

40 30

j _

17

2'1 2'3 2'~ 2~

1'9

_

i

2'9 3~

i

7

19

i

i

i

i

21

23

25

27

/

29

31

Retention Volume (mL)

Retention Volume (mL)

Kuwait Atmospheric Residue Vanadium Concentration (1000 COUNT/SEC)

Nickel Concentration (1000 COUNT/SEC)

~

40 30 20-

/

10-

!

17

0

!

19

2=1 23

25

27

/

17

3

29

19

21

23

25

27

2"9 31

Retention Volume (mL)

Retention Volume (mL)

Canadian Bitumen Nickel Concentration (1000 COUNT/SEC) 22 20" 18 16 14 12 10 8 6 4 2 0 -2 16

Vanadium Concentratio n (1000 COUNT/SEC)

600 500 400300 200 100 0

_

I

I

18

I

I

I

20

I

I

22

I

I

I

24

I

I

I

26

I

I

28

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32

Retention Volume (mLs)

-100

I

34

36

16

~ ~ ~ ~ ~ ~ ~ ~ ~ w ~ ~ ~ ~ ~ ~ ~ i ~ . 18 20 22 24 26 28 30 32 34 36 Retention Volume (mLs)

Figure 3. SEC/ICP-MS profiles of various feeds.

141 run (EOR) temperature condition, and maintained at that level for an additional five months. In both the cases, the total run length was seven months. The total system pressure, the gas rate, and the liquid space velocity were all held constant in both the cases. Over the nearly three and onehalf months of processing with Kuwait atmospheric residue over Nickel + Vanadium 4 RM-430 at the SOR temperature a condition, the relative loss in the EOR rate constant for metal removal 2 ~lr ~ ,_ ._ SOR MOR was less than 5% (see Figure 4). The activity loss at the MOR 1 temperature

condition was initially

slightly higher, however, aider one month of operation, the activity stabilized and remained very steady for the two and one-half additional months. When these results were compared with the results derived from testing conducted at the EOR temperature condition, it was observed that although the HDM activity increased with the increased temperature, the HDM deactivation rate also increased,

0 4

activity was higher at the higher operating temperature, the relative gain in the reactivity of other activities could, in some instances, make operating at a higher temperature more attractive for refiners. The relative rate constant for sulfur removal is shown in Figure 4 as a function of catalyst age for both the cases. The HDS activity rate stabilized at a higher activity level at the higher temperature, however, at the same deactivation rate as at the lower

temperature operation. The increase in the HDS activity was commensurate with the activation energy of sulfur removal for this catalyst. Similar results were noted

~> o EOR

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0 7 6 5 4 1

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6

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8

Run Length in Months

Figure 4. Relative rate constant for RM-430 using Kuwait atmospheric feed.

142 for nitrogen, Conradson carbon removal and pitch conversion, where activity increased with the increased temperature and remained stable at the higher temperature conditions, without deactivating at a higher rate. It would thus depend upon the refiner's processing objectives as to whether they would operate the fixed bed hydroprocessing unit at a higher temperature and gain improvements in conversion activity (pitch, CCR, HDN and distillate make), or operate at a lower temperature, and thereby maximize the run length. By increasing the temperature severity towards the start of run, the refiner can gain an advantage in increased activity for all functions. The ultimate cycle length, however, would depend upon the type of feeds processed, the metal levels of the feeds processed, and the other feed properties and process conditions. Overall, our data demonstrates that RM-430 catalyst hydrodemetallation capacity is very stable for HDM and other reactivities at the normal residue guard reactor operating temperature conditions. Operating at a higher temperature increases demetallation activity, but also increases the demetallation deactivation rate. All other activities are higher and stable at the higher operating temperature. It is interesting to note that the hydrodesulfurization activity (and HDM activity to a lesser extent) of RM-430 catalyst increased significantly over the first month of operation and then decreased at a relatively low rate beyond this period for all the feeds tested in our test programs. This activity increase is related to the auto-catalytic activity of the deposited metals. RM-430 catalyst has been designed to take advantage of the autocatalytic activity of metal sulfides deposited from the feed during hydroprocessing. The pore structure and the impregnation metal chemistry of the support is such that the deposited metal sulfides catalyze hydrodesulfurization and demetallation reactions. In the beginning of the run cycle, metal sulfides are deposited on the fresh surface and are active as catalyst sites, thereby, enhancing activity significantly. After a certain level of metals is deposited, further deposition occurs on previously deposited metal species, thus reducing the net activity gain. The ultimate deactivation rate is a result of deactivation due to core poisoning and pore mouth plugging due to coke and metal deposition, and the activity improvement due to the creation of new active sites due to metal sulfide deposits. PERFORMANCE OF RM-430 WITH HDS CATALYSTS IN STACKED CATALYST COMBINATION Long-term tests were conducted using stacked-beds containing RM-430 catalyst and other Criterion Catalyst Company residue hydrodesulfurization catalysts. In one of the tests, one third of the reactor volume was RM-430 (reactor top). Two other HDM/HDS catalysts: RN-410 and RN-400 (see Table 1 for properties) were used in combination with RM-430. Our objective was to process RM-430 with a high metal content feed. A Canadian bitumen-derived feed was selected for this purpose. This run was conducted for seven months at a total hydrogen pressure of 2,000 psig. As shown in Figure 5, despite the high metal content of the feed and the relatively high content of deposited metals over RM-430, the catalyst stacked-bed maintained a high demetallation activity. A high degree of liDS and Conradson carbon residue (Micro-carbon residue, MCR) conversion was also achieved for relatively long periods of time. At the end of the seven months of testing, nearly 60% metals was deposited on the RM-430 catalyst. No

143 indication of metal break-through was observed. Analysis of the aged RM-430 catalyst revealed that coke content was 13.6 to 16.1 wt% (from reactor top to bottom). This level of coke is relatively low compared to what would be expected for this high a Conradson carbon feed service. As shown in Table 3, on an average, the aged catalyst lost less than 50% of its surface area (SA) and less than 40% of its median pore diameter (MPD) even after deposition of nearly 60% metals. This loss in surface area and pore diameter is inclusive of the loss associated with coke deposition. Due to the relatively low coke deposits on this catalyst, larger catalyst volume capacity was available for metal deposition, Table 3. Analysis of Aged RM-430 Catalyst Samples - Canadian Bitumen Feed.

Position in Reactor

Coke (wt%)

%SA Retained

%MPD Retained

Top

13.6

47

68

Middle

15.0

55

58

Bottom

16.1

55

58

Relative Reactivity of Feedstocks Comparison of relative reactivities of Kuwait, Arabian Heavy, and Canadian Bitumen atmospheric feeds over RM-430 catalyst at hydrogen pared pressures normally encountered in residue processing revealed some interesting results. As shown in Table 4, Kuwait residue was the most reactive for sulfur removal (reactivity is defined as rate constant at a constant temperature and space velocity condition and reported relative to Heavy Arabian, which was assigned a ranking of 100). Heavy Arabian residue was the least reactive. The nickel removal activity of Heavy Arabian and the Canadian Bitumen feed were nearly equal and that of the Kuwait feed nearly a factor of two higher. Vanadium removal activity of the Canadian bitumen was the lowest and that of Kuwait the highest (a factor of two higher versus Heavy Arabian). here was no correlation between the sulfur removal activity and the bulk sulfur content nor was any correlation between the total nickel and vanadium content and the HDM activity. The SEC/ICP-MS analysis of the feed and product samples from these test runs also did not indicate any clear trend. The SEC/ICP-MS chromatogram indicated relatively smaller nickel and vanadium containing species to be present in the Canadian bitumen feed. The Middle Eastern feeds contained relatively larger molecular species (see Figure 3). Analysis of the hydroprocessed species from these experiments showed that the concentration of smaller sized molecules was reduced to a higher degree than the larger molecules. This points out that for the Kuwait feed, the larger sized species may be more heat labile. At the reactor temperature conditions, the large molecules probably break into smaller species, whose net reactivity is more than the reactivity of the smaller molecular species present originally in the feed (bitumen in this case). This association points out that reactivity is more a function of several physical and chemical characteristics of crudes and that at the reactor process conditions, feeds go through a substantial molecular transformation. A more detailed analysis of feed and product properties is thus required for predicting reactivity. Correlating reactivity to bulk feed properties may not be sufficient.

144 Table 4. Relative Reactivities of Feeds. Heavy Arabian

Kuwait

Canadian Bitumen

HDS

100

138

123

HDNi

100

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97

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Figure 5 - Reactivities using Canadian bitumen feed. We have also observed that the relative reactivity of the vanadium molecules for Middle Eastern crudes is higher than the nickel-containing species. As stated earlier, the vanadiumcontaining compounds present as metalloporphyrins are more abundant than the nickel compounds, probably causing the difference in reactivity Nickel removal was also observed to be more dependent on hydrogenation activity of the catalyst, indicating a stronger hydrogenation pathway prior to the hydrogenolysis step for nickel-containing species. It is our understanding that analytical tools like SEC/ICP-MS and SEM/EDX and other

145 chromatographic and mass spectroscopy techniques can be used to gather greater insights into the functioning of hydroprocessing catalyst. It is essential that the nature of the feedstock be fully characterized and analyzed before selection of catalysts for a processing service [8,9]. Better understanding of catalytic hydroprocessing will ultimately lead to development of superior catalysts and more efficient catalytic processing. CONCLUSIONS A new HDM catalyst has been developed which maximizes HDM activity and metal deposition capacity without sacrificing catalyst strength. We have tested this catalyst under various process scenarios and have determined that this catalyst can maintain high HDM activity at relatively large metal loading levels, with stable activity for sulfur, nitrogen, and Conradson carbon removal. REFERENCES

1. J. P Dickie, M. N. Hailer, T. E Yen, J. of Colloidal and Interface Science, 29 (1969) 475. 2. J. G. Erdman, J. of Chem. and Eng. Data, 8 (1963) 252. 3. E M. Dautzenberg, J. van Klinken, K.M.A. Pronk, S. T. Sie, J. B. Wijffels, Chemical Eng. Science, (1978) 254. 4. W.C.V.Z. Langhout, Oil & Gas Journal, Dec 1 (1980) 120. 5. J. M. Oelderik, S. T. Sie, D. Bode, Applied Catalysis, 47 (1989) 1. 6. C. H. Bartholomew, Catalytic Hydroprocessing of Petroleum and Distillates, M. Co Oballa and S. S. Shih (editors), Marcel Dekker Inc., 1993. 7. A. A. Del Paggio, G. J. Kamla, A. R. Forster, M. A. Shepherd, ACS, Division of Petroleum Chemistry, Symposium on Resid Upgrading, Washington D.C.,A-ugust 26-31 (1990) 606. 8. A. J. Suchanek, Oil & Gas Journal, Dec 17 (1984) 115. 9. J. R Hohnholt, WK. Shifiett, A. J. Suchanek, Oil and Gas Journal, May 28, (1990) 72.

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Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All fights reserved.

147

C O M M E R C I A L EXPERIENCE IN VACUUM RESIDUE HYDRODESULFURIZATION

Hiroki Koyama, Eiichi Nagai, Hidenobu Torii, Hideaki Kumagai Mizushima Oil Refinery, Japan Energy Corporation, 2-1 Ushio-dori, Kurashiki-shi, Okayama 712, Japan ABSTRACT Mizushima Oil Refinery of Japan Energy Corporation has succeeded in a high conversion operation of vacuum residue in the residue h y d r ~ ' o n unit~luipped with the fixed bed reactors. To complete a sixmonth cycle operation, hot spot ~ c e , pressure-drop build-up, and catalyst deactivation had been the most imtmrtant subjects to be solved.The conmaercialoperation has demonstratedthat good liquid distribution, which is obtained by uniform catalyst loading,an appropriate ca_t__a_lystshape, and good liquid distributors, prevents hot spot occurrence. Dispersing solids throughout the reactorsby an appropriate catalyst combinationhas beeneffectiveto control the pressure-drop increase in the first bed which is ~ by plugg~. Activitytests of the ~ catalysts showed that the catalyst in the last bed was most deactivateddue to coke fouling. It has been demonsWatedthat controlling the conversion in each bed reduces coke deactivationin the last bed. 1. I N T R O D U C T I O N The Mizushima Oil Refinery of Japan Energy Corporation first implemented an operation of vacuum residue hydrodesulfurization in the conventional fixed bed reactor system in 1980. We have also conducted a high conversion operation to produce more middle distillates as well as lower the viscosity of the product fuel oil to save valuable gas oil which is used to adjust the viscosity. Vacuum residue hydrodesulfurization in fixed bed reactors involves the characteristic problems such as hot spot occurrence and pressure-drop build-up. There has been very little literature available discussing these problems based on commercial results. Jaffe analyzed hot spot phenomena in a gas phase fixed bed reactor mathematically, assuming an existence of the local flow disturbance region [1]. However, no cause of flow disturbance was discussed. To seek for appropriate solutions, we postulated causes of hot spot occurrence and pressure-drop build-up by conducting process data analysis, chemical analysis of the used catalysts, and cold flow model tests. This paper describes our solutions to these problems, which have been demonstrated in the commercial operations. Feed properties and operation conditions determine catalyst life in the residue hydrodesulfurization. In a high conversion operation of vacuum residue, catalyst deactivation due to coke is as important as the one due to metals. Though many researchers have worked on understanding and modelling deactivation of residue hydrodesulfurization catalysts, there has still been a controversy in a coke deactivation mechanism [2, 3]. Very few publications are available discussing an effect of a bed temperature profile on catalyst deactivation in large scale adiabatic commercial reactors. Most of the studies on deactivation of residue hydrodesulfurization catalysts have been done with small-scale isothermal reactors [2,3,4,5]. The activity tests of the used catalysts were conducted to study the catalyst deactivation in the commercial reactors. This paper also describes an effect of a bed temperature profile on coke deactivation, which was tested in the commercial reactors.

148 2. PROCESS DESCRIPTION 2.1. History The residue hydrodesulfurization unit, which is the first commercial plant designed by Gulf Company, was constructed at the Mizushima Oil Refinery in 1970. The unit was equipped with two parallel single reactors, which were designed to produce fuel oil of 1.0 wt% sulfur content from 27,760 BSD of Kuwait atmospheric residue in a six-month cycle operation. In 1980, the residue hydrodesulfurization unit was remodeled to process a mixture of 50% atmospheric residue and 50% vacuum residue of the Middle East crude oil. Because of high metal concentration and difficulty in hydrodesulfurization, a lower space velocity and a large volume of the demetallation catalyst were required. Therefore, a new reactor was added in each train to process 19,000 BSD of feedstock. Since 1981, we have conducted a high conversion operation versus a constant desulfurization operation. Figure 1 shows the variation in the components and API gravity of the feedstock after the remodeling. Since Mizushima Oil Refinery has aimed at the operation of heavy crude oil conversion, the residue hydrodesulfurization unit has been required to process vacuum residue as much as possible. The feed ratio of vacuum residue had been raised nearly to 100 % by 1985. The demetallation and desulfurization catalysts which we developed greatly contributed to increasing the ratio of vacuum residue beyond the design. 20000

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Feedstock 1.04 100,000 4.8 0.43 22 170

Product fuel oil 0.97 900 1.0 0.32 11 64

149 2.2. Current operation Table 1 shows typical properties of the feedstock and the product fuel oil at the middle of a run. The specification of the product sulfur content determines an operation cycle. For the high conversion operation, the reactor temperature is fast increased high enough to convert vacuum residue into low viscosity fuel oil. Then, the reactor temperature is increased to compensate for the gradual decrease in the catalyst activity of the conversion to the end of a run. 3. HOT SPOTS 3.1. Hot spot mechanism A hot spot is defined as an existence of high local temperature in a catalyst bed. We often experienced hot spot occurrence in the first bed of the first reactor during the middle of a run. The bed temperature profiles are measured with thermocouples in three parallel, vertical wells. We usually observed hot spots in the center between the middle and bottom levels of the bed. Figure 2 shows changes in the temperatures of the middle level of the first bed. Only the temperature close to the center increased with time relative to the other temperatures and finally became a hot spot. Although we observed the same behavior of the temperatures in the bottom level of the first bed, no temperatures increased in the second bed. This suggests that the high temperature region was limited. We assume that a hot spot occurs in a local region of low liquid flow, where heat is release by the increased reaction rate. A composite catalyst sample was taken from each of thirteen regions every meter along the bed depth after a hot spot occurred. Figure 3a and 3b show the radial distribution of coke and metal deposit on the catalyst in different bed depths, respectively. Coke and metals were uniformly distributed at the top of the bed. However, relative coke deposit increased in the center with bed depths, while relative metal deposit decreased in the center with bed depths. This suggests that liquid flow rate was low in the center of the lower part of the bed, where hot spots were observed, because low liquid flow accumulates less metals and raises a conversion to deposit more coke on the catalyst. We propose the following mechanism of the hot spot occurrence. For some reason, maldistribution occurs and forms a low liquid flow region in the center of the lower part of the bed at the beginning of a run. In this low liquid flow region, a decrease in the bed voidage due to coke deposit further decreases liquid flow with time. A gradual increase in heat release in this region increases the temperature and finally causes a hot spot. Therefore, we conclude that the initial maldistribution is a cause of a hot spot.

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3.2. Experiments to understand a cause of maldistribution We conducted cold flow model experiments in a air-water/glycerin system to investigate a cause of maldistribution in a catalyst bed. The apparatus used was a 30 cm I.D. acrylic column equipped with a liquid distributor at the top and a liquid collector with 33 compartments at the bottom. Bed depth can be varied by combining the pipes. Liquid distribution at a given depth of the bed was estimated by measuring the liquid flow from each compartment of the collector. We examined effects of gas and liquid velocity, liquid viscosity, particle shapes, and ways of catalyst loading on liquid distribution in the bed. An increase in liquid velocity or viscosity slightly improved liquid distribution. However, gas flow rate did not affect liquid distribution. Three different ways of catalyst loading were tested to examine an effect of the particle orientation. Scattering the particles uniformly over the bed with a special equipment maintained the bed surface fairly flat during catalyst loading. Dropping the particles onto the center of the bed through a tube formed a convex bed surface during catalyst loading, while dropping them along the inner wall formed a concave bed surface. After loading all the catalysts, the surface of the bed was flattened to eliminate an entrance effect in all cases. Figure 4 shows radial liquid distribution for three different ways of catalyst loading using the trilobe catalyst. Liquid distribution was uniform in case of the catalyst loading which maintained the bed surface flat. However, in case of the catalyst loading which formed the convex bed surface, liquid flow was faster near the wall. On the contrary, in case of the catalyst loading which formed the concave bed surface, liquid flow was faster in the center. This indicates that liquid flows down along the slope of the bed surface which is formed during catalyst loading. Assuming that catalyst particles lie in parallel with the bed surface formed during catalyst loading, we conclude that liquid tends to flow along the particle orientation. Figure 5 shows radial liquid distribution for the catalysts with different shape, which were loaded in the way of forming a convex bed surface. The trilobe catalyst showed the fastest liquid flow near the wall. However, it was surprising that the cylindrical particles produced as good liquid distribution as the spherical particles, which have no particle orientation. Therefore, we assume that the pleats along the shaped catalyst particle lead liquid to flow along the particle orientation. In the first bed, we have used a sock loading method instead of a dense loading method, because we prefer higher bed voidage to avoid a pressure-drop increase due to plugging with solids. We had also used the shaped catalysts in the first bed except the top of the bed to increase the bed voidage for the same reason. In the sock loading method, catalysts are loaded in the reactors with a flexible hose. Since this method is easy to drop the catalysts around the center of the bed, the catalysts flow toward the shell and form a descending slope during catalyst loading. Therefore, we conclude that such poor loading of the shaped catalysts causes

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the low liquid flow in the center of the lower part of the bed at the beginning. 3.3. Development of liquid distributors A good liquid distributor is necessary to prevent maldistribution. The existing liquid distributor consists of a tray and a number of short chimneys. Liquid is collected on the tray and flows onto the bed from the chimneys through the small holes on the side. A cold flow model test showed that liquid flow rate from the short chimneys was extremely sensitive to the level of the tray. It was also pointed out that liquid dispersion from the short chimneys was poor. Therefore, we developed a new liquid distributor, which improved the defects of the existing one. The new liquid distributor with tall chimneys can achieve uniform liquid distribution, even if the tray is declined. Each chimney also has a feature to well disperse liquid onto the bed. A cold flow model test also showed that an effect of a liquid distributor was limited in a certain depth of the bed, which varied with a catalyst size. However, we have expected that a good liquid distributor lower a chance of the maldistribution which is caused by non-uniform deposition of solids on the top of the bed. 3.4. Trials in the commercial unit and evaluation To prevent hot spot occurrence, we evaluated the methods which improved liquid distribution, in the commercial operations. Figure 6 compares changes in the temperature deviation of the first bed during the different runs. The axial number, which is an average of the standard deviation of the temperatures at each level, is defined as an index of a degree of maldistribution. A small number indicates good liquid distribution. We took the countermeasures successively run after run. Before the improvements, a hot spot occurred within two months. During the first trial run, an increase in mass flow rate in the reactors by recycling the product fuel oil delayed hot spot occurrence a little. In the second trial run, we modified the sock loading method. To avoid forming a slope of the bed surface during catalyst loading, we have scattered the catalysts over the bed instead of dropping them around the center of the bed. This showed a great improvement in liquid distribution. Finally, changing the shaped catalysts for the cylindrical catalysts and replacing the existing liquid distributors with the new ones which we developed maintained good liquid distribution throughout the third trial run. Radial distribution of coke and metals in the first bed also verified the improvement in liquid distribution. We have had no hot spot problem since then. Therefore, it has been demonstrated that good liquid distribution prevents hot spot occurrence.

152

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Time Base " + Product recycle ......... + Improved catalyst loading + Cylindrical catalysts & new liquid distributors Figure 6. Changes in the temperature deviation in the first bed Liquid distribution in trickle bed reactors has been mainly discussed from the aspect of flow channels between particles [6, 7]. However, since most of the commercial catalysts are extrudates, an effect of the particle orientation on liquid distribution is much more important than flow channel, which relates to mass flow rate and a particle size. Shaped catalysts have a higher volume activity than cylindrical catalysts when an effect of diffusion on the reaction rate is large [8]. Therefore, the shaped catalysts have been commonly used for hydrodemetallation of residue. However, since an effect of liquid distribution on the catalyst performance is important in large-scale commercial reactors, catalyst shape should be carefully selected to maximize the effectiveness of the catalyst usage in a commercial application. 4. PRESSURE-DROP BUILDUP We often observed a pressure-drop increase in the first bed. The solid line in Figure 7 shows a typical curve of a pressure-drop increase in the first bed of the commercial reactor. The pressure drop starts increasing half way into the run and rises exponentially with time. The unit has to be shut down before the pressure-drop reaches the limitation. When we first experienced a serious pressure-drop increase, we observed that the fines of iron sulfide plugged the voidage of the bed surface. The iron sulfide particles suspended in the feedstock are so fine that most of

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153 them go to the reactors through the feed mechanical filter. We increased the catalyst size and voidage at the top of the bed to increase the tolerance for plugging. However, as the ratio of vacuum residue increased, an increase in solid concentration made the pressure-drop in the first bed increase sharply again, as shown in Figure 7. We took the catalyst samples at various bed depths to analyze the solids on the catalysts. Detecting a large amount of solids beneath the large catalyst layer led us to the conclusion that the solids through the large catalyst layer deposit on the top of the small catalyst layer. The components of the solids were not only iron sulfide but also coke and other inorganic compounds. However, this result prompted us to disperse the inorganic solids throughout the reactors to control the pressure-drop increase in the first bed. The cold flow model test in an air/slurry system demonstrated that the solid deposition rate on the catalyst particles was proportional to their specific surface area (the ratio of the particle surface area to the volume), which varies with particle size and shape. This result suggests that distribution of the inorganic solid in the reactors can be controlled by combining the catalysts with different size and shape. The pressure-drop in the fresh catalyst bed is also a function of equivalent particle diameter and voidage, which also relate to particle size and shape. Therefore, we conclude that combining the catalysts with different size and shape in the first bed can control the rate of the pressure-drop increase with cumulative reactor feed. The dotted line in Figure 7 shows that an appropriate combination of the catalysts with different size and shape prevented the pressure-drop increase in the first bed. The catalyst activities were also designed properly to control coke deposition. About 30% of iron sulfide carried into the reactors accumulated on the first bed before the improvement. However, an appropriate catalyst combination has reduced the iron solid accumulation to half. This suggests that dispersing inorganic solids throughout the reactors is effective to control the pressure drop increase in the first bed. We have not observed any serious pressure-drop increase in the lower beds. 5. CATALYST DEACTIVATION 5.1. Catalyst deactivation in the commercial reactors The activity tests of the catalysts used in the commercial reactors were conducted in the bench-scale reactor. The aged catalyst samples were taken from the second bed through the fourth, where the hydrodesulfurization catalyst was packed. The aged catalysts were Soxhletextracted with toluene followed by drying. The activity tests were conducted for the fresh and aged catalysts with Arabian Heavy atmospheric residue at a temperature of 360 ~ and pressure of 12 MPa. A detail of the study on the catalyst deactivation in the commercial reactors will be discussed elsewhere [9]. Table 2 summarizes the chemical analysis and the relative activities of the typical catalyst samples. The coke content increased with the bed depths, while the metal content decreased. The activity tests show that the catalyst in the fourth bed was most deactivated. The activity tests of the aged and regenerated catalysts, using model compounds in a gas phase, were also conducted [9]. It suggests that the fourth bed catalyst was heavily deactivated by coke fouling. Table 2 Chemical analysis and activities of the catalysts used in the commercial reactors Samples Chemical analysis of the aged catalyst Ni + V g/100g-fresh catalyst Carbon g/100g-fresh catalyst Activity relative to the fresh catalyst

2nd bed

3rd bed

4th bed

21.2 12.1 27

15.2 15.7 28

6.8 22.1 21

154 The work by Tamm et al. showed that the activity of the aged catalyst in the exit of the reactor was higher than that in the entrance after the constant desulfurization operation [2]. However, Myers and Lee observed the lowest catalyst activity due to coke fouling in the third of three consecutive expanded reactors after the high conversion operation, where the conversion range of 1100+OF boiling material was 60 to 70% [5]. They operated the reactors isothermally. Therefore, we assume that an increase in the conversion of high boiling material increases coke fouling in the fourth bed, or the last bed. 5.2. Effect of a bed temperature profile on coke deactivation in the fourth bed It has been thought that coke is produced by the precipitation of large molecular hydrocarbons such as asphaltenes when their solubility in oil is lowered [10, 11]. An increase in the conversion of vacuum residue increases the aromaticy of the asphaltenes and decreases the aromaticy of the maltenes [12]. Consequently, the solubility of the asphaltenes in the maltenes decreases. However, an increase in the aromaticy of the asphaltenes may be controlled if we choose an appropriate operation condition where polymerization or condensation of the cracked asphaltenes is prevented by hydrogenation of the radical bonds. Absi-Halabi et al. point out that the asphaltenes partly have a responsibility for coke fouling of the catalyst subsequent to the initial rapid coke deactivation [ 11]. Therefore, we assume that controlling the conversion in each bed to maintain the solubility of the asphaltenes reduces coke fouling in the fourth bed. We tested an effect of a reactor temperature profile on the coke fouling in the fourth bed in the commercial operations. Figure 8 and 9 show the reactor temperature profiles at the middle of the runs and the corresponding catalyst activity curves in the third and fourth beds, respectively. The catalyst activity of each bed was calculated from the correlation between hydrogen consumption and a bed temperature rise. As show in Figure 8, an increase in the reactor inlet temperature decreased the temperature rise, or the conversion rate in the fourth bed. This indicates that more conversion took place in the upper beds. Figure 9 shows that the catalyst deactivation rate with metal accumulation in the fourth bed was higher than that in the third bed. This is consistent with the result of the activity tests shown in Table 2. Figure 9 also shows that the operation with a higher reactor inlet temperature maintained the catalyst activity in the fourth bed higher than the operation with a lower reactor inlet temperature. In case of the lower reactor inlet temperature operation, since the conversion rate in the fourth bed was higher, we assume that insufficient hydrogenation of the cracked asphaltenes caused severe coke fouling there. In case of the higher reactor inlet temperature operation, although the

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Metal on catalyst O 3rd bed (Lower reactor inlet temperature) 6 4th bed (Lower reactor inlet temperature) 9 3rd bed (Higher reactor inlet temperature) 9 4th bed (Higher reactor inlet temperature) Figure 9. Catalyst activity curve in each bed

155 catalyst activity in the second bed became lower because of some shift of the conversion to the upper beds, the total catalyst activity became higher. Therefore, we conclude that controlling the conversion in each bed by the bed temperature control can minimize the coke deactivation in the fourth bed and maximize the run average catalyst activity. It should be noted that solving a hot spot problem in the first bed has enabled to increase the reactor inlet temperature. 6. CONCLUSIONS Effective solutions to the problems of the vacuum residue hydrodesulfurization unit equipped with the fixed bed reactors, such as a hot spot, pressure-drop buildup, and catalyst deactivation by coke fouling, were discussed. Improving liquid distribution can prevent hot spot occurrence. Dispersing inorganic solids throughout the reactors can control a pressure-drop increase in the first bed. For a high conversion operation, controlling the conversion in each bed can minimize the coke deactivation in the fourth bed. REFERENCE 1. Jaffe, S. B., Ind. Eng. Chem. Proc. Des. Dev., Vol. 15, No. 3,410 (1976) 2. Tamm, P. W., Harnsberger, H. F., and Bridge, A. G., Ind. Eng. Chem. Proc. Des. Dev., Vol. 20, No. 2, 262 (1981) 3. Bartholdy, J. and Cooper, B. H., ACS Prepr. Div. Petrol. Chem., 205th National Meet., Denver, 386 (1993) 4. Johnson, B. G., Massoth, F. E., and Bartholdy, J., AIChE J., Vol. 32, No. 12, 1980 (1986) 5. Myers, T. E. and Lee, F. S., AIChE Symp. Series, Vol. 85, No. 273, 21 (1989) 6. Fox., R.O, Ind. Eng. Chem. Res., Vol. 26, No. 12, 2413 (1987) 7. Herskowitz, M. and Smith, J. M., AIChE J., Vol. 24, No.3, 439 (1978) 8. Cooper, B. H., Bonnis, B. B. L., and Moyse, B., OGJ, Dec. 8, 39 (1986) 9. Koyama, H., Nagai, E., and Kumagai, H., to be published in ACS Symp. Series (1995) 10. Wiehe, I. A., Ind. Eng. Chem. Res., Vol. 32, No. 11, 2447 (1993) 11. Absi-Halabi, M., Stanislaus, A., and Trimm, D. L., Applied Catalysis, Vol. 72, Elsevier, Amsterdam, 193 (1991) 12. Takatsuka, T., Wada, Y., Hirohama, S., and Fukui, Y., J. Chem. Eng. Japan, Vol. 22, No. 3, 298 (1986)

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Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All fights reserved.

157

C O M P A R I S O N OF OPERATIONAL MODES IN RESIDUE H Y D R O P R O C E S S I N G M. de Wind a, Y. Miyauchi b and K. Fujita b

aAkzo Nobel Chemicals B. V., P.O. Box 247, 3800AE Amersfoort, The Netherlands bNippon Ketjen Co., Ltd., Jatxm I. ABSTRACT In the past few years the residue hydrodesulfurization process has gone through a number of changes. Deeper desulfurization and more conversion to mid-distillate have become a primary target for several units. At the same time, heavier residues are being processed. To address these and other questions, Nippon Ketjen has developed a new series of resid catalysts, viz. the KFR series. The two most common modes of resid hydroprocessing applied on commercial scale, are illustrated with pilot plant test data and data from commercial units. 2. I N T R O D U C T I O N Since the late sixties, residue hydrodesulfurization plants have been constructed for the production of low sulfur fuel oil, specially in Japan. After the oil crisis, their role has changed significantly. Nowadays they are also used as: 1. Mild hydrocracker to produce middle distillates from heavy feedstocks at high temperatures. 2. Feed pretreater for Residue Fluid Catalytic Cracking (RFCC) units. Ketjenfine resid catalysts, the so-called KFR series, have been developed by Nippon Ketjen to meet these revised demands and have been successfully introduced into the resid hydroprocessing market. More than 6000 tons of these types of catalysts have been used already. In resid hydroprocessing, the following issues are of prime importance: i. Guard reactor fouling. ii. Dry sludge or sediment formation in product oils. iii. The mode of operation, that is constant hydrodesulfurization (HDS mode) or high temperature conversion (MHC mode). iv. The effect of feedstock on activity and stability. In this paper we discuss the effect of the mode of operation on catalyst deactivation and product properties in residue hydroprocessing. Other mentioned issues have been addressed in other publications [ 1-2]. 3. THE T W O O P E R A T I O N A L MODES Table 1 gives a comparison of two operational modes practiced in residue hydroprocessing: MHC and HDS. In the MHC mode of operation, usually heavy feedstock is processed at a relatively high liquid space velocity and temperature to produce maximum middle distillate. The desulfurized bottoms are fed to a vacuum tower to produce FCC feedstock, are used as a blending stock for

158 Table 1. Comparison between the MHC and HDS mode of operation. Mode of operation Type of feedstock

MHC AR, VR/AR, VR

HDS AR, AR/VR

Process objectives

Production of middle distillate Blending feed for RFCC or FCC Low sulfur fuel oil

Pretreatment of RFCC feed Blending feed for RFCC or FCC

Feedstock properties: Viscosity Density Sulfur CCR Ni + V Fe Process conditions' H2 partial pressure LHSV Hz/oil Temperature Catalyst life 1000~ conversion

cSt @ 50~

1000-27000

300-30000

g/ml @ 15~

0.97-1.01

0.95-0.98

wt% wt%

3.5-5.0 10-18 120-250 10-40

1.7-4.4 8-14 35-120 5-15

~

105-165 0.25-0.55 900-1300 400-410

90-145 0.17-0.45 700-1300 395-405

months wt%

3-6 max. 45-50

8-22 max. 45-50

ppm ppm bar h"l Nm3/m3

Product properties: Fraction Viscosity

cSt @ 50~

Density Sulfur CCR Ni + V

g/ml @ 15~ wt% wt% ppm

Desulfurized bottoms AR AR < 600 150-450 0.92-0.97

0.92-0.95

< 1.5 < 12 < 90

0.1-1.0 3-7 5-30

residue FCC, or are used as heavy fuel oil. Especially in this operation, the following problems may be encountered: i.

Guard reactor fouling becomes more severe, because hydrotreating of heavy feedstocks with a high iron content at a high temperature causes deposition of iron at the top of the guard catalyst bed. This results in pressure drop buildup, oil mal-distribution, hot spot formation and catalyst agglomeration [2-4].

ii. Sludge formation in product oils lilnits the maximum conversion attainable in commercial units. If the 538~ (1000~ exceeds 45-50 wt%, the unit can no longer be operated. The dry sludge

159 deposition, which occurs in flash drums, effluent heat exchangers and fractionators causes a drastic decrease in the heat transfer coefficient and buildup of pressure drop. Therefore, the unit should be operated below the critical conversion (45-50 wt%) at lower reactor temperature. The critical level depends on the characteristics of the feedstock to be treated and the type of catalyst system used but typically lies within the range mentioned [5]. iii. The MHC mode requires a more metal tolerant catalyst system because heavy feedstocks with high vanadium and nickel contents are hydrotreated at high temperature. Despite the use of improved catalysts, the catalyst life is much shorter than obtained with the HDS mode of operation. The cycle length could become as low as 3 to 6 months. In the HDS mode of operation, the cycle length is longer (typically one year). Guard reactor plugging and dry sludge formation only tend to occur towards end of run, when the temperature reaches similar levels as applied in MHC mode. Besides metal tolerance, coke formation on the tail end catalyst is a predominant mechanism of deactivation in the HDS mode. 4. PILOT PLANT TEST RESULTS We have done pilot plant tests under both MHC and HDS conditions, to show and explain the particularities of each mode. The MHC mode of operation is characterized by a reactor temperature of 410~ to produce maximum middle distillates using Iranian Heavy Atmospheric Residue (IH-AR) feed. A constant hydrodesulfurization test was done at 93% HDS with Arabian Medium Atmospheric Residue (AM-AR) feed.

Table 2. Process conditions for pilot plant test Operation mode Feedstock

MHC Constant conversion

HDS 93% HDS

IH-AR

AM-AR

Reactor temperature

~

410

Increase required for 93% HDS

PPH2

bar

130

135

Hz/oil

Nl/l

base + 200

base

LHSV

h"1

base + 0.12

base

Table 2 shows the process conditions applied in the pilot plant tests. Figure 1 gives the catalyst configurations. The two demetallization catalysts, KFR 11 and KFR 10 have a very high metal absorption capacity. KFR 11 is more metal tolerant than KFR 10. KFR 30 is a dual function catalyst for hydrodesulfurization and demetallization which is used as an intermediate stage in the catalyst system. KFR 50 is used as a downstream catalyst that has a high activity for desulfurization and is more coke resistant

160 KFR 10 (DEMET)

KFR 30 (HDS/HDM)

50 ~KFR

KFR 50 (HDS)

30

MHC AT 410~ PILOT PLANT

20

11 (DEMET) MHC AT 410"C COMMERCIAL

/

21

I

9 3 % HDS IN PILOT PLANT

55 I

0

50

100

C A T A L Y S T POSITION IN REACTOR, %

Figure 1. Catalyst configurations. The MHC mode of operation requires more demetallization catalyst in the reactor to obtain a more metal tolerant system. Table 3 shows the properties of the feedstocks used in the pilot plant tests. The Iranian Heavy AR for the MHC mode of operation contains 261 ppm of metals (vanadium + nickel), while the Arab Medium AR for the 93% HDS mode contains 87 ppm of metals. A summary of the test results is presented in Figures 2-7. Table 3. Properties of feedstocks Feedstock origin Operational mode

Iran Heavy MHC

Arab medium 93% HDS

Feedstock properties: Density @ 15~

g/ml

0.999

0.983

Viscosity @ 50~ Sulfur Nitrogen Vanadium Nickel CCR

cSt wt% wt% ppm ppm wt%

14500 3.53 0.52 201 60 15.5

1381 4.19 0.23 66 21 11.8

n-C7 insoluble

wt%

5.5

4.5

GC Distillation: IBP

~176

160/320

295/563

720+~

wt%

97.4

93.5

IO00+~

wt%

65.0

63.4

161 420 MHC MODE

400 A HDS MODE 380

360

I

0

100

I

200 DAYS ON STREAM

300

Figure 2. WABT program for pilot plant testing. 2.0

MHC MODE/~

1.5

22~

/

1.0

U_ ....1

0.5

B

ir

HDS MODE A "'~"

I

0.0

100

I

200 DAYS ON STREAM

300

Figure 3. Sulfur in total liquid product. 120

E __.1

LU I-'-O (23 O 12L

r

100 M 80 60 40 HDS MODE

20 w I

100

9 I

200 DAYS ON STREAM

Figure 4. Metals in total liquid product

I

300

162 30 ,,'k

A

.~.

/k

MHC

MODE

o~

O-

20

HDS

(.9

MODE

.-I-

z In O .-J

10

LU m

>..

O

I

I

I

I

O

100

200

300

DAYS

ON STREAM

Figure 5. Yields of naphtha and gas oil. The most remarkable observations are: - The middle distillate yield is almost constant during the Mild Hydrocracking mode of operation, except at the end of the run. - The temperature stability of the KFR catalyst system is extremely good for the constant lIDS mode, in the range of 392-400~ WABT (Figure 2). 100 _

~k-_.

..... , 80

...........

o~ -

z O

60

>

" ~ ~ ~ .

.

.

.

.

.

.

.

.

.

.

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

-

r nnun

40

...........

HDS

9........................ .. ................................. -.~. .-. . ~ . . - .,-,. . . . , ,

.

.

.

"

......

.

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

"~

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

9

" "--.,~o

.... H t E i E i ~ i ~

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

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

9 ""'~~

1000*F 9C O N V

:..:..~..

~ .......

'"~ .......

~t ...............................................................................................................

"<:'~

.........

Z

O o

_

20

...........

~

9. . . . . . . . . . . .

fi- .............

.....

.--..-..-:-.:..~.:.-..-.-..-.-..:.-....t.

720~

0

o_

...........

ca ...........

-o ..............

a..

........

....... n ....... J . . . . . . . . . . . . . . . ..................................................... m .........

"--.o...

n -.-.. . . . . . .

CONV

5~

1~0 DAYS ON STREAM

Figure 6. Performance of the pilot plant MHC operation.

1~0

"'\.. x

......

163 100 :~

__,

H~ D-S__ __ __,~ .

~-

. . . . . .

~

- __ - : ~

. . . .

C~

80

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

" ~ . ~ . ~ S -

o

HDCCR ~

:

:

:

.....:::::::::::::::::::::::::::::::::::::::: ..........~ .....-,,...........

60 -

(/) rr uJ

>

.....

HDM

~

..... ~

z" 0

~

zx

~

HDN

~

n

40

..~---"

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

~-

I. ~ ~ ~ -

""

""

~

~

~

~

~

1000*F CONV

"~ ...........................................................................................................

Z

0 o

f

-

20

720"F CONV

"

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

_

/ I

.

~ ............................................... ! ...................................

' ' /

DAYS ON STREAM

Figure 7. Performance of the pilot plant HDS operation. From Figures 6 and 7, showing the removal of several impurities in each mode of operation, we can conclude the following: - The removal rate of asphaltenes (N-C7 insoluble) goes down as the time on stream increases. This is probably because the pore size of the demetallization catalysts is reduced by metal deposition, resulting in a lower rate of diffusion of asphaltenes into the catalyst pores. - The mechanism for removal of Conradson Carbon Residue (CCR) appears to be the same as for sulfur removal. This may be concluded from the similar behavior in deactivation. 93% sulfur removal corresponded to 62-65% removal of CCR in both modes of operation although the feedstock properties were quite different. - In MHC mode, the conversion of the 382+~ (720~ and the 538+~ (1000~ fraction stays constant except at end of the run. - The catalyst life at 93% HDS mode is about 280 days. At end of run, the weighed average bed temperature (WABT) reaches 398-399~ The critical conversion of the 538+~ (1000~ fraction for sludge formation is around 50%. The amount of dry sludge in the 385+~ fraction from the last part of the test was in the range of 0.1-0.2 wt%. In this range the plant operation becomes critical due to sludge deposition in the reactor effluent heat exchangers. Product yields and properties at end of run are summarized in Table 4. Commercial experience at EOR in the MHC mode of operation is also included. We observe that: There is a good agreement between pilot test results and commercial data. - The desulfurized bottoms from the HDS mode meet the usual quality targets for residue FCC feedstock. The acceptable quality of residue FCC feedstock depends on the design of the catcracker. -

164 Table 4. Product yields and properties at end of run Mode of operation

MHC Pilot plant

MHC Commercial unit

HDS Pilot plant

Hydrogen

wt%

-1.2

-1.5

H2S

wt%

3.05

4.18

NH3

wt~

0.23

0.18

C~-C4

wt%

2.08

1.76

wt.

vol.

vol.

wt.

Naphtha

%

2.6

3.5

3.4

3.6

vol. 4.7

Gas oil

%

20.1

23.1

21.3

13.4

15.6

Bottoms Total

% %

73.1 95.8

76.3 102.9

77.4 102.1

78.4 95.4

83.2 103.5

Properties of desulfurized bottoms: Density @ 15~ Viscosity @50~

g/ml

0.9576

0.9506

0.9207

cSt

647

416

124

CCR

wt%

9.9

-

4.1

n-C7 insoluble Sulfur

wt% wt%

4.0 0.81

1.0

2.6 0.31

Nitrogen

wt%

0.39

-

0.10

Vanadium

ppm

54

65

9

ppm Nm3/m3

27 136

27 138

6 166

Nickel 1-t2 Consumption Cut points (TBP): DSN

~

160

191

DSGO

~

382

343

5. C A T A L Y S T DEACTIVATION The kinetics for liquid volume conversion and the removal of sulfur, and metals are described as follows: 1

(1-1)

1 1 Cp(,_, ) C f ( , . , ) ) L H S V '~ = A(t) e -gIRt

with:

(1)

Reaction order Concentration of sulfur and metals in total liquid product, wt% and ppm respectively. Cf = Concentration of sulfur and metals in feedstock LHSV = Liquid hourly space velocity, hr"~ n

--

Cp

=

165 a

A(t) E R

T

= = = = =

Power of LHSV Frequency factor (rate constant) which is a measure for the catalyst activity. A is a function of the time on stream (or metals on catalyst). Activation energy Gas constant Weighed average bed temperature, K

The relative volume activity (RVA) versus the rate constant at SOR is defined as follows: RVA

=

A~ x A(SOR)

(2)

100

The SOR is defined as 10% of the run length after the first day of oil-in. Using these formulas we obtained Figures 9, 10 and 11, which show the relative volume activity of sulfur, metals removal and the 720~ conversion versus metals on catalyst. Commercial data under MHC conditions are included in these figures The metals on catalyst depend on the time on stream as shown in Figure 8. These graphs indicate the following: More demetallization catalyst leads to a more metal tolerant catalyst system (compare Figure 1 and 9). - The commercial MHC run as shown in Figure 9 was terminated when RVA HDS was 24. This corresponded to 29 wt% of metals on catalyst. A hydrotreating run is usually terminated when RVA HDS reaches the range of 40 to 20, independent of the mode of operation. This is because the HDS deactivation is strongly enhanced below RVA HDS of 40. No deactivation of the 382+~ conversion was observed in both the commercial and pilot plant MHC runs till RVA HDS reached 20 (metal on catalyst < 45 wt%). - The 93% HDS mode showed some deactivation of the 382+~ conversion. -

-

50

40

MHC MODE

30

d O

20

HDS MODE

10

I

100

200 DAYS ON STREAM

Figure 8. Metal buildup on catalyst during HDS and MHC operation

I

3O0

166

100

MHC

PILOT

50 40 HDS PILOT

30 20

10

MHC COMMERCI

I

i

i

20

10

0

i

30

METALS

ON CATALYST,

40

50

wt%

Figure 9. Stability of the desulfurization.

100 >-

9

~-~ i-0

50

tu

40

i---

.~

L

30 20

HDS PILOT

10

I 10

0

~

o

MHC

I. 20 METALS

PILOT

I 30 ON CATALYST,

I 40

50

wt%

Figure l 0. Stability of the demetallization. 150 i

> i

/x

c.) /x LU

> I--

MHC

PILOT

A

100 * M~C

COMMERCIAL

L.U rr

8O

z 12) ~

7O

-

HDS PILO

rr"

> z 12) 0

t30

5O

m

I

10

I

I

20 METALS

I

30 ON CATALYST,

Figure 1 l. Stability of the 720~ conversion.

40 wt%

50

167 In the following section we will discuss why the catalyst system is less metal tolerant in the 93% HDS mode and why more deactivation in conversion is observed in the HDS mode than in the MHC mode, although the HDS mode requires lower temperatures during operation. 6. DEACTIVATION BY COKE AND METALS The metals and coke (toluene insoluble) on spent catalyst (expressed on flesh catalyst basis) from the pilot plant runs and the commercial MHC run are presented as a function of the catalyst position in the reactor in Figures 12 and 13. These two graphs indicate the following: - The demetaUization catalysts (KFR 11 and 10) have a much higher metal absorption capacity than the other catalysts (KFR 30 and 50). In the MHC mode, the coke deposition on the catalysts is almost constant from the reactor top to the reactor bottom. This means that coke deactivation has the same significance for all catalyst types and is not such a dominant factor. Deactivation by metal deposition is dominant for each catalyst. Therefore demetallization catalysts with a high metal tolerance (higher metal absorption capacity) are essential for the catalyst life in the MHC mode. - In the 93% HDS mode, the coke level on the upper and middle bed catalysts (KFR 10 and 30) is similar as found in the MHC operation. Coke on the downstream catalyst is significantly higher. -

The average coke levels on KFR 50 spent catalyst (downstream catalyst) from several pilot plant tests and commercial units are plotted as a function of the hydrodesulfurization percentage (HDS%) at the end of the run (EOR) in Figure 14. Clearly, deep desulfurization, that is above 8590%, results in much higher coke formation on the catalyst. Figure 15 shows the coke on KFR 50 from the downstream section as a function of WABT at EOR in a 93-94% HDS mode. There is a strong correlation between coke formation and EOR WABT. In Figure 16, we show the effect of time on stream on coke formation for the 93% HDS mode. 100

o~

I".

80

I.-:

r

>-o z 12)

60

4O

L

~

"''" . . . . . . . ~ . . . . . . . MHC MODE

....I

ua

2o

t 0

, os oo , 20 CATALYST

40 POSITION

60 IN R E A C T O R ,

Figure 12. Metals on catalyst as function of reactor position.

80 %

100

168 50

o~

40

I---(.f) >...

30

.ml

cO z O ILl O cO

HDS P

9 ..........or.--- /

I

L

~

MHC P I ~ / N

20 "~"

m

-I

B- . . . . . . . . . MHC COMMERCIAL

11 ........................

10

0

I 20

H5

I 40

CATALYST

I 60

POSITION

I 80

IN R E A C T O R ,

100

%

Figure 13 Coke on catalyst as function of reactor position. 40

O

PILOT TEST

9

COMMER

30 tar) >.. -...I

I----

C) 20

-(2)

co z O LIJ O CO

9 10

IO

--

KFR 50 0

I

30

I

40

50

I

60

I

I

70

80

I

90

1 O0

HDS AT LOB, %

Figure 14. Average coke level on KFR-50. 40

o~ "~ I--(.f) >-

30

m

.._J

~

20

0 I2) u.l 0 0

10

KFR 50 I

380

I

I

400 WABT AT EOR,"C

Figure 15. Increase of coke on KFR-50 with WABT.

420

169 40

"~

30

i-.->...J

I-.< O

20

Z

O I.U

~" O O

10 KFR 50 I

I

100

200 DAYS

I

300

ON STREAM

Figure 16. Increase of coke on KFR-50 with days on stream.

Figures 14, 15 and 16 make clear that the cycle length, EOR temperature and degree of desulfurization determine coke make on the KFR 50 catalyst. It is obvious that this effect is more predominant in HDS mode than in MHC mode. 7. CONCLUSIONS In mild hydrocracking of atmospheric residue for maximum middle distillates production, the metal tolerance of the demetallization catalyst is the most important factor determining the catalyst life. The deactivation by metals is the prevailing mechanism in this mode of operation. In pretreatment of residual FCC feedstock, downstream catalysts with high coke resistance are essential to obtain sufficient catalyst life. The improved metal tolerance for upstream catalysts and improved coke resistance for downstream catalysts are the main success factors of the KFR catalyst system. REFERENCES

1. Y. Miyauchi c.s., Nippon Ketjen Seminar 1992, "Correlation between catalyst performance in

2. 3. 4. 5.

laboratory tests and commercial unitsfor resid hydrotreating with the KFR catalyst system '" Tokyo. F. L. Plantenga, Akzo Catalysts Symposium 1991, "Akzo Chemicals'guard bed technology", Amsterdam. Y. Miyauchi c.s., JPI-Petrotech 13(1) (1990) 44. K. Fujita c.s., First Tokyo Conference on Advanced Catalytic Science and Technology, "New guard catalystfor descaling and de-iron", 1990. S. Saito, AIChE Annual Meeting, "Experience in operating high conversion residualHDS process with ABC catalyst", San Francisco, 1984.

This Page Intentionally Left Blank

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All fights reserved.

171

MINA ABDULLA REFINERY EXPERIENCE WITH A T M O S P H E R I C RESIDUE DESULFURIZATION (ARDS) A. A I - N a s s e r , S. R. C h a u d h u r i , a n d S. B h a t t a c h a r y a

Mina A bdulla Refinery, Kuwait National Petroleum Co., Mina Abdulla, Kuwait.

ABSTRACT Mina Abdulla Refinery was modernized in the late Eighties. The Crude Processing capacity was expanded and a number of downstream processing units were installed. Based on the design philosophy of minimum Fuel Oil production from the Refinery, the new processing scheme included a Hydrocracker and a Delayed Coking unit with Vac. Gas Oil and Vac. Residue as feed stocks respectively. Light and Middle Distillates Hydrotreaters along with Atmospheric Residue Desulfurization (ARDS) Unit were also incorporated in the refinery process scheme.. Three different catalyst systems have been utilized in the ARDS Unit, namely single HDS, dual HDS/HDM and multiple HDM/HDS/HDS-HDN catalyst combinations. This paper presents the performance and an analysis of the three systems with particular reference to catalyst life. The Refinery experience will be useful in selecting improved catalyst type and combinations to optimize run length and products. INTRODUCTION ARDS unit works as the springboard in the new scheme of Mina AbdullaRefinery operation. Primarily a desulfurization unit, ARDS also reduces the metals, asphaltenes and nitrogen in the products, thereby, ensuring proper quality of feed for downstream conversion units. As an additional benefit, ARDS is also a mild hydrocracking process, partially upgrading high sulfur atmospheric residue to low boiling products like naphtha and distillate. Choice of the catalyst system and severity of operation, therefore depends on the following major factors : - ability of the catalyst system to meet the target product qualities during the run length. - ability of a catalyst system to ensure a steady economically viable run length. - stable by-product distillate.

ARDS UNIT G E N E R A L DESCRIPTION ARDS unit has two trains, each having a Guard Chamber and three main reactors with a common fractionation section. A simplified flow scheme is shown in figure (1). Other than the Off-Gases, the three liquid products are Naphtha (C5- 375~ TBP), Distillate (375-680~ TBP) and Low Sulfur Fuel Oil, that is LSFO (680~ + TBP). Feed stock

The quality of Feed to the ARDS unit generally remains comparable to the specification used in unit design as shown in Table 1.

172 M A K [ UI' Ft..,,

I'UiI(;F

GAS

v----

NAPHTHA

RESID

GAS

FEED

MAKEUP

OIL

HYDROGEI~

JRGE FUE L OII

GAS

Figure 1. Process flow diagram of the ARD unit. Table 1. Properties of ARD feedstock.

Gravity Sulfur Con. Carbon Metals (Ni & V) Sodium Nitrogen Distillation recovery @ 680~

~ wt. % wt. % ppm wt. ppm wt. ppm wt. vol. %

Design

Typical

12.5 4.5 12.2 max. 88 max. 3 2800

13.5 4.3 12.0 75 2.0 2500

6

6

Unit Processing Objectives Processing objectives of ARDS unit may be categorized as below: 1. Feed Preparation for Down Stream Units - Reduce Nitrogen in VGO to specification requirement.

conform to Hydrocracker and FCC units feed

- Control sulfur in Vacuum Rerun Unit feed in order to obtain a desired coke quality as well as to be within metallurgical limits of the unit.

173

2. Fuel Oil Blending - Produce LSFO of typically 0.5 % wt. sulfur for use as a blending component in the F.O. 3. Product Slate improvement In addition to the above, mild hydrocracking associated with the process leads to production of Distillate and Naphtha thus helping upgradation / optimization of refinery product slate.

Reactors Arrangement Each ARDS train has four reactors holding about 28,000 CFT of catalyst. First one is a Guard Chamber loaded with about 7 % of total catalyst. The other three reactors hold about 3 1 % each of total catalyst. Quench Hydrogen is available at the inlet/outlet of the Guard Chamber and outlet of reactors 2 & 3. Quench is utilized for adjusting the reactors inlet temperatures to maintain desired temperature profiles in the catalyst beds. The Guard Chamber can be bypassed in case of high pressure drop due to scale or coke accumulation to enable continued operation. About 20% ( max. ) of the total reactor system pressure drop is admissible across the Guard Chamber.

Unit Operating Variables & Modalities The ARD trains are generally operated at design capacity of 33,000 BPD each. The Recycle Gas (85% H2) rates are maintained as high as the system permits - usually giving a Gas : Oil ratio between 4500 - 5000 SCF/BBL at the reactor system outlet. The reactors system pressure is maintained at the design level between 1800 - 2000 psig. There is facility for water injection in the charge heater coils to improve catalyst performance. Generally, a fiat temperature profile is maintained at the outlet of the three main reactors, with GC outlet remaining ~ 20 - 30~ lower. Maximum exotherm allowed in the three main reactors is around 50~ GC exotherm is generally limited to 25~ This is to ensure a more controlled and uniform deactivation of the catalyst system. Some catalyst systems are characterized by higher pressure drop in the reactors than the other. To elaborate a little, pressure drop in the reactors is a function of shape (i.e. cylindrical, spherical, trilobe or quadralobe) of catalyst and its size. It has been observed that pressure drop can be brought within operating limits by varying Gas / Oil ratio or changing oil viscosity. However, reduction in pressure drop by Gas/Oil ratio adjustment is mostly at the cost of expected life of the catalyst. Further discussion on this is followed in the next paragraph. C A T A L Y S T SYSTEMS USED AT MINA ABDULLA R E F I N E R Y Three distinct catalyst combinations have been used at Mina Abdulla ARDS unit. A : Mono Catalyst - HDS B : Dual Catalyst - HDM + HDS C : Multiple Catalyst - HDM + HDS + l I D S / H D N General specifications of the catalysts are given in Table 2.

174 Table 2. Catalyst Specifications.

Catalyst System Catalyst type

Combination %

Active metal

Shape Size

B

A

C

HDS

HDM

HDS

HDM

HDS

HDN

100

52

48

32

52

16

CoMo

CoMo

NiMo

Mo

NiMo

NiMo

Tri-/ Quadralobes 1/20" 1/4" 1/8"

Quadra lobes 1/20" 1/4" 1/8"

Quadra lobes / Cylindrical 1/20", 1/32" 1/5" 1/10"

All the three systems employed graded bed in order to tackle the reactor pressure drop, especially at the front end where the operating temperatures are lower and catalyst encounters the virgin stock. Most of the catalysts used are tri-lobes or quadra-lobes which are known to ensure better wetting of the catalyst surface with consequent improvement in liquid flow pattern and sustenance of catalyst activity. On the other hand, cylindrical shaped catalysts though results in lesser void fraction has poorer wetting characteristics. Catalyst size selection is also of utmost importance. More catalyst can be packed in available volume with lesser catalyst size. However, reaction section hardware may pose a limitation. A minimum size of 1/20" seems to be optimum for Mina Abdulla Refinery reactor system. Last but not the least, criterion for good performance of a catalyst system is its crush strength. A refinery catalyst unit is liable to encounter upsets during its run length. A quality that a refiner looks for in the catalyst is its ability to sustain such upsets without any adverse impact on its performance. PERFORMANCE OF THE CATALYST SYSTEMS

Catalyst System A: HDS During initial stages of operation the catalyst was very active i.e., SOR WABT was about 25~ lower than predicted. The rise in WABT in first four months was continuous @ about 6~ Subsequently, the catalyst activity declined sharply(WABT rise 10-12 ~ The pressure drop across the main reactors remained low (35-55 psi) and steady throughout the run length.

175 The naphtha and distillate makes were significantly high. The catalyst behavior was explainable in view of the operating requirement i.e., primarily LSFO sulfur. Since the catalyst was solely HDS type, a high degree of desulfurization was obtained at lower temperature at SOR. On the other hand, the feed metals poisoned the HDS catalyst very fast, initially at the front end and then the subsequent beds. Once this happened, the WABT had to be raised steeply to attain target LSFO Sulfur. Due to the high fouling rate, the catalyst cycle length was lower than expected.

Catalyst System B: HDM / HDS The catalyst showed high rate of deactivation (9~ WABT ) during first 2 months or so followed by a relatively steady period of 5-6 months (4~ WABT). The EOR was characterized by steep temperature rise. After six months of operation, the GC was by-passed in both trains. Pressure drop across the main reactors remained almost steady throughout the run (55-75 psi). Presumably the effect of catalyst fouling was countered by the increase in reactors temperatures from SOR to EOR. The behavior of HDM/HDS combination was as expected. As the amount of liDS catalyst was much less than that in Catalyst System-A and desulfurization activity of front end catalyst (HDM) declined fast, WABT had to be raised progressively right from the SOR to obtain target LSFO Sulfur. At mid-of-run, both HDM and HDS catalysts were at their optimum temperature levels and a rather steady activity (& LSFO Sulfur) was observed for some time. This was expected as the HDM catalyst in the front end of the reactor system protects the HDS catalyst from getting fouled up by metal deposition. At EOR, the scavenger HDM catalyst was exhausted and metals in feed started poisoning the HDS catalyst. This phase was characterized by a steep rise in WABT to maintain LSFO Sulfur. The drift of exotherm from Reactor-1 to Reactor-2, as observed, was due to saturation of HDM catalyst and consequent shiit of relatively more virgin feed (& reaction) to the next catalyst bed. The pressure drop characteristics of this batch remained stable throughout the run which allowed comfortable Gas/Oil ratio. This did contribute to the overall relatively good performance of this catalyst in terms of cycle length.

Catalyst System C: HDM/HDS/HDS- HDN An initial period of stabilization (about two months) was experienced with this catalyst system, when high (8~ catalyst deactivation rate was observed. After this, a period of low catalyst deactivation rate followed (4-5~ month WABT). Regarding the reactors exotherms, the same trend as with Catalyst System-B was observed. First main reactor showed about 50% of total delta temperature which drifted to the next reactor at EOR. As expected, the behavior of Catalyst System-C was similar to that of Catalyst System-B in view of similar HDM catalyst guard at the front end. However, the reactors pressure drop was much higher from the SOR itself, enforcing a lower Gas/Oil ratio that resulted in enhanced catalyst fouling and deactivation. The higher pressure drop was related to the characteristics (i.e. size and shape) of the catalyst. The products yield was, as expected, similar to that of Catalyst System-B.

176 Table 3. Comparative Performance of Catalyst Systems. ATTRIBUTES Effective run length Effective catalyst life Operating LSFO Sulfur Demetallation Denitrification Naphtha Dist. Yield VGO Make-up Hydro[en Rx. Pressure drop

UNIT

CAT-A

CAT-B

CAT-C

Months BBL/Ib %wt. %wt. % wt. % Vol. % Vol. SCFB psi.

8.8 8.1 0.65(0.5-0.9) 50 35 16-33.2 43.5 1005 45

10.5 11 0.58(0.53-0.73) 70 40 20-25 40-45 1015-1125 (65)*

10 9.5 0.65(0.5%0.$2) 82 39 19-30 46 1015-1050

(80)*

* GC was bypassed partially or totally

CATALYST PERFORMANCE ANALYSIS The performance of the three catalyst systems are summarized in Table 3. The three systems operated with the same feed and nearly the same target Sulfur in LSFO. Their most significant differences in performance were as follows: In Catalyst System-A, the metals in the feed poisoned the catalyst very fast imposing higher severity on the remaining part which deactivated as fast. This was apparent in the progressive and continuous increase of WABT throughout the cycle length (Fig.2). Though the WABT, to start with, was lower than catalyst systems B & C, rate of deactivation was very high resulting in shorter run length. Other relevant operating and performance parameters are shown in Fig. 3 & 4. However, this system generated more distillate towards EOR which again could be attributed to the characteristics of the catalyst system. As was expected, demetallation/denitrification of the feed stock was substantially lower than that for Catalyst Systems B and C. Even though product pattern with this system was superior, short catalyst cycle length and consequently limited On-Stream Factor (OSF) weighed heavily against it. Short catalyst cycle length was operationally uneconomical due to frequent shut downs. Catalyst System-B was a significant improvement in terms of cycle length. The Demet catalyst at front end of the reactor system acted as metals scavenger and protected the activity of HDS catalyst down below. In general, the run was also characterized by almost constant distillate production from SOR uptill EOR situation. It is worth mentioning here that this Catalyst System-B tolerated a prolonged shut down without any appreciable deterioration in performance. Catalyst System-C incorporated Denitrification catalyst intended to lower Nitrogen in VGO (Hydrocracker Feed stock). The advantages over Catalyst System-B, if any, was not clear. Higher Distillate stability was expected but could not be substantiated. Distillate yield was comparable with that of Catalyst System-B.

177

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t

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100

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DAYS ON OIL CAT. SYSTEM A

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CAT. SYSTEM B

9

CAT. SYSTEM C

Figure 2. Days on oil vs. Performance and operating parameters.

178

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.

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#

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179

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100

150

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DAYS ON OIL =

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Figure 4. Days on oil vs. product quality and distillate yield.

180 OPTIMUM CATALYST BLEND To safeguard against metals poisoning, some section of reactors volume is necessarily allocated to the Demet catalyst. The clue to optimum catalyst combination, therefore, is in developing improved catalyst and a balance among various types, i.e. a. Search for higher metal removal capacity HDM catalyst to leave more reactor volume to HDS catalyst. b. Search for improved HDS catalyst that gives: i. Superior activity and stability. ii. higher tolerance to metals deposition. c. Improve physical characteristics of catalysts to maximize Compacted Bulk Density (CBD) within mechanical limitation of the reactor system. CONCLUSIONS 1. Mono catalyst system, i.e. only HDS, by itself is a poor performer for the feed stock processed in ARDS unit. 2. A dual catalyst system for ARDS units processing heavy stock, i.e. combination of HDM and HDS catalysts gives a more stable and longer catalyst cycle length, more suited to refiner's over-all operational plan. 3. Advantage of a multiple system, i.e. combination ofHDM, HDS and HDN combination catalyst, over that of a dual-catalyst system is not conclusive. However, HDN catalyst is expected to impart color stability to naphtha and distillate products. In short operating experience with different catalyst systems helps in selecting improved catalyst combinations to optimize run length, product yields and qualities.

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All rights reserved.

181

C O S M O RESID HYDROCONVERSION CATALYST: CATALYST C O M B I N A T I O N TECHNOLOGY

Yasuo Yamamoto, Yoshihiro Mizutani, Yukio Shibata, Yasushi Kitou and Hatsutaro Yamazaki Cosmo Research Institute, 1134-2 Gongendo, Satte, Saitama 340-01, Japan ABSTRACT For the hydroconversion process of residual oil such as atmospheric residue and vacuum residue, Cosmo Oil Co. has succeeded in developing a combination system with Cosmo CF catalyst. The catalyst system was designed with investigations of the performance of each catalyst and the combined catalyst effect. The catalyst system properly combined with the pretreatment catalyst suitably designed produced less degradation of the activity at high temperature ranges. This stable catalyst system still showed high degradation after reaching a point of a very high metal deposition, called the metal breakpoint. Accordingly, the deactivation behavior of each catalyst of the catalyst system was investigated in several long-term bench plant tests. The relationship between the hydrodemetallization (HDM) catalyst behavior and the total deactivation was cleared. In addition, the metal capacity of the catalyst was suggested, and the breakpoint could be estimated. Finally, a comparison of performance between the bench plant test and the commercial operation showed good agreement. Catalyst combination technology for op'tunizing a suitable catalyst system was built upon. 1. I N T R O D U C T I O N Recently, the processing of heavy crude oil has been important, while the demand in Japan for middle distillates has been increasing steadily, demand for residual fuel oil has been declining. Therefore, oil refiners have been interested in converting residues to middle distillate products. Cosmo Oil Co. has been developing the new residual hydroconversion technology since the 1970s. This technology is aimed at converting residues to more middle distillates by utilizing a residual hydrodesulfurization unit (R-HDS unit). As a result, we have succeeded in developing a catalyst system with less degradation of activity in a high temperature range, as well as simulation technology for the system. The developed catalyst system has been successfully used in commercial units through several demonstration runs at the Cosmo Chiba refinery from 1983 to 1987, using an R-HDS unit from UNOCAL [1,2,3], and at the Cosmo Sakaide refinery from 1988 to 1990 using an R-HDS unit from GULF [4]. This catalyst system consists of three types ofCosmo CF catalysts in combinations, which each have individual functions, such as a high metal uptake capacity, and high desulfurization activity, high conversion. Furthermore, the developed technology simulates the performance of the catalyst system, such as deactivation rates, catalyst life, optimum operation guides, and product yields.

182 Before the demonstration operation in commercial units, extensive basic research on catalyst combination technology was carried out. Therefore, we present a catalyst system and combination technology deriving from the following items: (1) (2) (3) (4) (5)

Characteristics of each catalyst Relationship between catalyst properties and performance. Performance behavior of catalysts used in combination. Activities and deactivation behavior of the catalyst system. Deactivation of each catalyst used in combination.

2. EXPERIMENTAL APPARATUS AND METHOD For the experiments in this study, three types of reactor system were used. One had some separate micro-scale reactors, and it was used to measure HDS and HDM activity at constant operational conditions. The volume of each micro-reactor was 70 cc. The two other systems were bench-scale plants which had two or three reactors used in series for performance tests of catalysts used in combination and for which intermediates from each reactor could be sampled and analyzed. The volume of each reactor in the bench plant was 1000 cc. The two-reactor bench plant was used for the investigation of the activity of catalysts used in combination, and the three-reactor bench plant was used for the investigation of the deactivation behavior of catalysts during long-term operation. Presulfiding conditions consisted of a presulfiding source of light gas oil which included 5 wt% carbon disulfide, an LHSV of 1.0 h1, a maximum temperature of 370~ a ratio of feed oil to hydrogen gas of 500 Nm3/m3, and a presulfiding time of 20 hours. 3. CHARACTERISTICS OF COSMO CF CATALYSTS Our main purpose for developing residual hydroconversion catalyst is the upgrading of petroleum residue, decomposition of asphaltenic components and hydrocracking of hydrocarbons to obtain useful middle distillates from petroleum residue. Through extensive studies on HDS catalysts, hydroconversion was determined to be entirely dependent on reaction temperature [2]. On the other hand, coking and metal deposition onto catalyst were reported to occur under such high temperatures as to decrease catalyst activity and shorten catalyst life [5,6,7, 8]. Therefore, an excellent catalyst system, which had high activity and high metal tolerance needed to be developed. Series of Cosmo CF catalysts were developed under such a philosophy [9]. Their typical properties are shown in Table 1. They consist of three types of catalyst: 500 series, 3 series, and 2 series. They are used in combination to have optimum performance. CF-2 series catalyst was designed to show good hydroconversion and hydrodesulfurization (I-IDS), and to be loaded into the latter part of reactors. CF-3 series catalyst showed good hydroconversion and hydrodemetallization (HDM) in addition to hydro-desulfurization. CF-500 series catalyst was designed to have excellent hydrodemetallization and metal uptake capacity. Their performance features are shown in Table 1. 4. RELATIONSFIIP BETWEEN CATALYST' PROPERTIES AND PERFORMANCE There are many factors which control catalyst performance. A pore structure, such as pore diameter and pore volume, is one of the important factors which strongly influence HDS activity and deactivation rates of the catalyst. While developing Cosmo CF catalyst, several test catalysts were investigated for their performance.

183 Table 1. Features of Cosmo CF catalysts i

Carrier Active Metal Size CBD, g/cc SA, m2/g MPD Metal capacity Demetallization Desulfurization

CF-500 Alumina Groups VI&VIII Metal 1/20" 0.54 120-280 La~e High High Low

CF-3 ~ +-~-0.62-0.65 220 Medium Medium Medium Medium

CF-2 ~-1/16"or 1/20" 0.68-0.72 230-250 Small Low Low High

The catalyst which has a larger pore diameter tends to show a lower deactivation rate, as well as lower HDS activity. Figure 1 shows one example of the results of residual hydrodesulfurization experiments testing three kinds of catalysts which have different pore diameters. The micro-reactors were operated under the same conditions, as shown in Figure 1. Catalyst A, Catalyst B and Catalyst C were the test catalysts which have the same properties with different pore diameters ( C > B > A ). The activity and deactivation rate of each catalyst were shown to depend strongly on pore diameter, as shown in Figure 1.

Figure 1. Effect of MPD. on HDS activity and deactivation rate. Pore diameter: Catalyst C > Catalyst B > Catalyst A HDS activity: Catalyst A > Catalyst B > Catalyst C Deactivation rate: Catalyst A > Catalyst B > Catalyst C

184

-~ O1

2

e,..) o 1.5

",o ~ , \ ", ,, ~ ,

" [] CAT-C/CAT-A o CAT-B/CAT-A *- CAT___-A/CAT-A"

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~\\'r5~..M,,,

.m

ec

1

...................................... "-9.....~.---:-,~.~-~.~.~2-.,.43 1

1

J

Sul,ur level to catalyst A, wt% Figure 2. Example of Synergistic Effect of Catalyst B and C against Catalyst A (CAT-A: HDS Catalyst A; CAT-B: PretreatrnentCatalyst B; CAT-C: PretreatmentCatalyst C) 5. BEHAVIORS OF CATALYSTS USED IN COMBINATION

For the R-HDS units, catalystmight be used in combination. This catalyst combination system should be optimized with catalyst selection and a combination ratio against the operation and feed conditions. The performance of the catalyst system can be estimated by the HDS activity and deactivation rate of each catalyst. Estimation of HDS activity is complicated because HDS activity increases with the use of catalysts in combination [4]. Figure 2 shows the HDS activity of liDS catalyst combined with pretreatment catalyst. The lIDS activity of Catalyst A, expressed as the reaction rate constant Ks, increased with the severity of pretreatment Catalysts B and C, expressed as a decrease in the sulfur levels of catalyst A. The same catalyst, used as HDS catalyst for pretreatment, caused no increase in HDS activity when set in the latter part of the reactor. Different pretreatment catalysts caused some increases in HDS activity. This was called the combination effect [4] on HDS activity, also said to be the synergistic effect. Figure 3 shows the relationship of a catalyst combination ratio with its HDS activity and deactivation rate. The HDS activity level is expressed as a product sulfur level on a typical operation condition. With an increase of the ratio of pretreatment catalysts, the HDS activity tended to decrease, and the product sulfur level increased. The HDS activity of the catalyst system was calculated with the activity levels of each catalyst and the synergistic effect of the HDS activity on the catalyst combinations. Then, the deactivation rate decreased with increases in the ratio of pretreatment catalysts. The deactivation rate of the catalyst system was calculated in a weighted manner for each volume ratio. An optimal catalyst system should have maximum lIDS activity and a minimum deactivation rate under the given operating condition. The best catalyst combination can be estimated with the required sulfur level. A series of catalyst systems were able to copi~ with the dotted area shown in Figure 3 and meet the required activity levels. In these catalyst systems, the catalyst system which had lower deactivation rate would be selected. This catalyst system would have high stability in a high temperature range.

185

Figure 3. Relationship of catalyst combinations with product sulfur level and deactivation rate. 6. ACTIVITIES AND DEACTIVATION BEHAVIORS OF A CATALYST SYSTEM An optimal catalyst combination system can be estimated through the selection of the catalysts and their combination ratio versus the operation conditions, as shown before. Before commercial application of such an optimum catalyst system, the catalyst system should be tested with a bench plant for a long period under conditions equivalent to the commercial conditions anticipated in order to obtain detailed data with commercial applicability. Figure 4 is an example of a long-term bench plant test for a catalyst combination system. Several ten days after the start-of-run, the catalyst system showed stable deactivation. During a stable deactivation period, the catalyst deactivation rate is constant. If the operation mode was a constant product sulfur mode, the temperature-increase-rate of reaction (TIR) was constant and small. Then, after the stable deactivation period, the catalyst system showed a higher deactivation period, in which the TIR became constantly larger than that during the stable deactivation period. The point at which the deactivation rate changes is called a breakpoint. 7. BEHAVIORS OF EACH CATALYST WHEN USED IN COMBINATION A basic concept of the catalyst combination system is that pretreatment catalyst, such as HDM catalyst, loaded in the upper section of a reactor removes the deactivation components included in feed residual oils, and it protects the desulfurization catalyst and the hydroconversion catalyst loaded in the latter section of reactor. Several bench plants which had multi-reactor systems and intermediate product sampling systems were operated to investigate the deactivation behavior and metal accumulation of each catalyst. The required temperature and the metal accumulation can be calculated through the analysis of each intermediate product for sulfur levels and metal levels.

186

Figure 4. Example of catalyst deactivation in long-term bench plant tests. Figure 5 shows the results of the long-term bench plant test of Catalyst System A, investigating the behavior of each catalyst in combination as it is usesd. The guard reactor included the HDM catalyst. Reactor A and Reactor B included the desulfiafization catalyst and the hydroconversion catalyst. With increases in the days on stream, catalyst activities decreased and amount of the metal that ~ t f l a t e d on catalysts (MOC) increased. Alter about 275 days on stream, the HDM catalyst in the guard reactor started deactivating faster. Then, the TIR of the total catalyst system became larger. At this break point, the MOC of the guard reactor was calculated to be about 43 wt%.

Figure 5. Comparison of catalyst deactivation and metal accumulation for each catalyst in the long-trem bench plant tests. (the guard reactor includes HDM catalyst. Reactor A and Reactor B include HDS catalyst and HC catalyst).

187

LHSV 0.50h-1 FEED Boscan Crude WABT 395 deg. C

r

O

o~ -1

O

O O O O O

t" n

-2

-3 )-)

-4

\,

-5 0

I 10

i 20

I 30

J I 40

I 50

60

MOC, wt%

Figure 6. Results of metal capacity measurement of HDM catalyst with Boscan crude oil. Figure 6 shows the result of the metal capacity measurement of HDM catalyst with Boscan crude oil. Catalyst deactivation was measured with an accelerated method. The HDM catalyst shown in Figure 6 was the catalyst used as pretreatment catalyst in Catalyst System A. HDS activity, expressed as a reaction rate constant Ks, decreased with increasing MOC. After the point at about 43 wt% of MOC was reached, catalyst activity decreased faster. This point is the breakpoint of the catalyst and the MOC at this point was suggested as the metal capacity. This MOC level was nearly equal to those shown in Figure 5. Cosmo CF-500 series catalyst has been developing for a higher metal capacity. Figure 7 shows a comparison between the long-term bench plant test results and the commercial operation results using the same catalyst system. Their operation conditions were different for LHSV, feed base, and so on. Here, the required temperature was used to normalize operation conditions and compare activity. The MOC was used to normalize the historical severity. Both deactivation behaviors on the bench plant test and the commercial operation showed fair agreement. The small disagreement in breakpoint might have been caused by the severe operations at the end-of-run in the commercial operation, according to the one-year-operation typically found in Japan. A normalized reaction temperature, such as the required temperature, is very effective for finding out the catalyst activity level under conditions different from the base data produced from the bench plant test. During commercial operation, the constant product sulfur mode operation would be run with changing feed sources and charge rates, which would be different from those tested in the bench plant operation; however with the required temperature and the breakpoint estimation, the catalyst life could be estimated fairly. 8. CONCLUSION We have carried out the basic research on catalyst combination technology. The results reported on this paper are summarized as follows: 1) The HDS activity and deactivation rates of each catalyst were mainly dependent on the pore diameters of the catalysts.

188

Figure 7. Comparison of the results of the bench plant test and the commercial operation 2) The synergistic effect on HDS activity was clear, and the performance of a catalyst system could be estimated. 3) The catalyst system could be optimized on the basis of the estimated HDS activity and deactivation rates for each catalysts. 4) The rapid deactivation of catalyst was related to the break-point of the first pretreatment catalyst, and could be predicted by the metal capacity measurement for an accelerated deactivation method with Boscan crude oil. 5) The bench plant test and the commercial operation were compared and showed good agreement for MOC and required temperature. Another basic study on simulation technology for life estimations has been carried out to achieve the commercial use of developed catalyst. Cosmo Oil Co. has succeeded in commercializing the developed catalyst and catalyst system through this basic research and demonstration operation experience. REFERENCES

1. C. Sera, H. Yamazaki and T. Tomino, Akzo Catalysts Symposium, May (1988) H-13. 2. H. Yamazaki, K. Tawara and T. Tomino, NPRA Annual Meeting, March (1988) AM-88-60. 3. C. Sera, K. Usui and H. Yamazaki, Studies in Surface and Catalysis, 44 (1989) 291. 4. H. Yamazaki, T. Tomino, Y. Yamamoto, M. Yumoto and Y. Mizutani, 2nd Joint Saudi-Japan Workshop on Recent Developments in Selected Petroleum Refining and Petrochemical Processes, December (1992). 5. D. Alvarez, R. C_~asso and P. Andreu, J. Japan Petrol. Inst., 22, No. 4, (1979) 234. 6. O. Togari, H. Takahasi and M. Nakamura, J. Japan Petrol. Inst., 23, No. 4 (1980) 256. 7. H. Nomura, Y. Sekido and Y. Ohguti, J. Japan Petrol. Inst., 23, No. 5 (1980) 321. 8. E.C. Sanford and R.P. Kirchen, Oil & Gas J., Dec. 19 (1988) 35. 9. Research. Assoc. Residual. "Nichel Molybedenum Alumina Catalysts of Specified pore distribution used for combined hydrodesulfurization and hydrocaracking of Heavy Oil" US Patent No.4 732 886. (1988).

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All fights reserved.

189

I N F L U E N C E OF CATALYST P O R E SIZE ON ASPHALTENES C O N V E R S I O N AND C O K E - L I K E SEDIMENTS F O R M A T I O N DURING C A T A L Y T I C H Y D R O C R A C K I N G OF K U W A I T VACUUM RESIDUES

A. Stanislaus, M. Absi-Halabi and Zahida Khan

Petroleum Technology Department, Kuwait Institute for Scientific Research, P. O. Box 24885, 13109 Safat, Kuwait ABSTRACT A critical factor that limits the maximum attainable conversion of heavy residues to lighter cuts in commercial residue hydroprocessing units is coke-like sediments formation. Suppression of sediments formation is highly desirable to increase distillate yields. As part of a research program on the factors which influence sludge or sediments formation during hydroprocessing of Kuwait vacuum residue for high conversion, we have investigated the relation between catalyst pore size, asphaltene conversion and coke-like sediments formation. Five Ni-Mo/7-AI203 catalysts with different unimodal and bimodal pore size distribution were used in the study. A unimodal pore catalyst with maximum pore volume in medium size mesopore range (100 - 200 A dia) showed the highest activity for the overall conversion of the residual oil to distillates. However, a relatively larger percentage of sediments was also observed for this catalyst. Catalysts with a large proportion of macropores, particularly in the 800-3000 A range produced little or no sediments, but showed poor activity for asphaltene cracking and overall conversion of residual oil to distillates. Molecular weight of the residual asphaltenes in the product increased with decreasing catalyst pore size. The concentrations sulfur and vanadium in the residual asphaltenes were found to be higher for catalysts having only small and meso-pores without macropores. The results have been explained on the basis of the importance of the ratio of feed molecular size to catalyst pore size in determining the diffusion and reaction rates in residue hydroconversion. INTRODUCTION Deep conversion of heavy petroleum oils and residues to lighter cuts by catalytic hydrocracking is becoming increasingly important in modem petroleum refining due to increasing market demand for cleaner transportation fuels with very low sulfur levels (1,2). One of the problems faced by the refiners in residue hydrocracking for high conversion is the formation of coke-like sediments (3-6). The coke or sediment usually deposit on the reactor and down stream vessels as well as on the catalyst surface and cause both operability and catalyst deactivation problems (6,7). In addition, the product stability is also affected. The problem becomes particularly more important at high temperatures when the conversion of residues to distillates is high. In order to minimize these' problems, the refiners are often forced to operate at low conversion levels (ca. 50%). Despite its importance as a critical factor limiting the maximum conversion attainable in commercial residue hydroprocessing units, the problem of sediment

190 formation has not received much attention and the mechanism of its formation is not fully understood. Suppression of sediment formation is highly desirable to increase distillate yields in catalytic residue hydrocracking. Factors influencing the formation of coke-like sediments during catalytic hydrocracking of Kuwait vacuum residues have been the subject of some investigations in this laboratory (8,9). The macromolecules of the heavy feedstocks such as asphaltenes are generally believed to contribute significantly to sediment formation and coke deposition (10,11,12). Consequently, in our studies particular attention was paid to the fate of asphaltenes during heavy residue hydroprocessing. In a previous paper we reported the effect of operating temperature on asphaltene conversion and coke-like sediments formation (8). In the present work, we have studied the effect of catalyst pore structure on asphaltenes conversion and sediment formation during hydroprocessing of Kuwait vacuum residues. Five Ni-Mo/y-A1203 catalysts with different unimodal and bimodal pore size distribution were used in the study. The nature of changes that take place in the asphaltenic and non-asphaltenic fractions of Kuwait vacuum residue during hydroconversion over catalysts of different pore size were examined by separating the asphaltenes from the liquid products and analyzing them by various techniques. The primary objective of the study was to understand the role of catalyst pore size on asphaltene conversion and coke-like sediments formation and to determine desirable pore size for minimizing the formation of undesirable sediments and improving various conversions during hydroprocessing of Kuwait vacuum residue. EXPERIMENTAL Five Ni-Mo/y-A1203 catalysts with wide variations in pore size distribution were used in the present study. They were all in the form of extrudates. The catalysts were characterized according to standard procedures. A mercury porosimeter (Quantachrome - Model-Autoscan 60) was used to determine pore size distribution. A Quantasorb adsorption unit was used for BET surface area measurements. The chemical composition and physical dimensions of various catalysts used in the present study were not appreciably different. The pore volume distribution curves for the five catalysts are shown in Figure 1. Table 1 summarizes the meso (30-500 A diameter) and macro (>500 A dia) pore size distribution for the five catalysts. Activity tests were conducted in a fixed bed reactor using Kuwait vacuum residue as feedstock (API gravity = 6.8; S =5.2 wt%; N = 0.44 wt%; V = 94 ppm; Ni = 26 ppm; asphaltenes = 9.2 wt%; CCR = 19.2 wt%). A J0 ml sample of the catalyst, diluted with an equal amount of carborundum, was charged into a tubular reactor. Thermocouples inserted into a thermowell at the center of the catalyst bed were used to monitor the reactor temperature at various points. After loading the catalyst the system was purged with nitrogen, and the temperature was increased to 150~ gradually. Then, the system was purged with hydrogen and pressurized to 120 bars. Under these conditions, the presulfiding feed (recycle gas oil) was fed and presulfiding was carried out using standard procedures (8).

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P Q R S T

Total Pore Volume (ml/g) 0.53 0.60 0.73 0.75 0.69

Meso-pore Distribution (%) 30-100 A 38 4 7 55 38

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When presulfiding was completed, the feed (Kuwait vacuum residue) was injected at 100 ml/h and the conditions were adjusted to desired operating temperature, pressure, hydrogen flow and LHSV. Testing was carried out under the following conditions: pressure, 120 bar, LHSV, 2h1; H2/oil, 1000 ml/ml/h; temperature, 440~ After 6 hours of operation under the set conditions, liquid product samples were collected every 48 h for various tests. Feed and product samples were analyzed using standard procedures. Molecular weight distribution were determined by gel permeation chromatography (Waters Associates). Sediment content in the liquid products was estimated by filtration through a glass fiber Whatman GF/A (1.6 ~tm porosity) filter at 100~

192

Figure 2. Influence of Catalyst Pore Size on Asphaltene Conversion, Distillate Yield and Sediment Formation. RESULTS AND DISCUSSION 3.1. Effect Catalyst Pore Size on Asphaltene Cracking, Distillate Yield and Sediments Formation.

The percentages of asphaltenes conversion on different catalysts with widely varying pore size distribution are presented in Fig. 2a. Interestingly catalyst P with maximum (60%) pore volume in the meso-pore range (100 - 200 A dia) shows the highest activity. Catalyst R that contains 34% of pore volume in 100-200 A pores ranks next. The activities of the large unimodal pore catalyst Q (that contains a major proportion of its pore volume in 800 - 3000 A pores) and the narrow pore bimodal catalyst S that contains a large proportion of narrow pores (<100 A dia) are relatively low for asphaltenes cracking. Figure 2b shows the yield of distillates (boiling below 524~ for various catalysts. It can be seen that catalyst (P) shows a substantially higher activity than the others for the conversion of Kuwait vacuum residue to distillates. As observed in the case of asphaltene conversion, catalyst R ranks second for conversion of residues to distillates. Compared with the bimodal pore catalyst S the meso-pore unimodal catalyst P and catalyst R show respectively, about 8% and 4% higher conversion. It appears that catalysts having a large amount of 100-200A diameter pores are more active in promoting asphaltene conversion and residue cracking. To confirm this point we conducted an additional run with catalyst T prepared from catalyst S by mild hydrothermal treatment. This catalyst contained a higher amount of 100-200 A pores than catalyst S (Table 1) and showed a substantially higher activity for conversion of residue to distillates (Fig. 2b). Mesopores in the range 100 - 200 A, thus, play an important role in the overall conversion of residual oils to distillates.

193 With a view to gain information on the influence of catalyst pore size on the formation of coke-like sediments during hydroconversion of Kuwait vacuum residue, the toluene insoluble materials formed in the hydrocracked product over various catalysts were estimated using a standard procedure. The results are presented in Figure 2c. A relatively high percentage of toluene insoluble sediments is noticed for catalyst P. In general the extent of sediment formation over different catalysts is in the order of their ranking for residue conversion to distillates. 3.2. Characteristics of Residual Asphaltenes

The nature of changes in the characteristics of asphaltenes in Kuwait vacuum residue during hydroprocessing over catalysts of different pore size were examined by separating the asphaltenes in the liquid products from various runs and analyzing them for the H, C, S, V, and Ni contents. The molecular weight distributions of asphaltenic fractions of feed and products were also determined by gel permeation chromatography. The purpose of these tests was to have a better understanding of the relation between catalyst pore size, asphaltene cracking and sediments formation. The S/C and V/C ratios in residual asphaltenes in the products from different runs are plotted in Figs. 3a and 3b. The lowest sulfur and vanadium concentrations in the asphaltenes are noticed for the run conducted with the monomodal macropore catalyst Q. The catalyst R which contains a large proportion of both macro and meso pores ranks next for the removal of sulfur from asphaltenes. Although the conventional bimodal catalyst (S) with a large proportion (> 50%) of micro pores and about 20% macropores shows good activity for removal of sulfur from asphaltenes similar to catalysts R its activity for vanadium removal is poor.

Figure 3. S/C and V/C ratio in residual asphaltenes for different catalysts.

194 The highest amount of sulfur and vanadium in residual asphaltenes is found for the unimodal meso-pore (100-200 A) catalyst (P), although its activity for the overall asphaltenes conversion is higher than the other catalysts. The molecular size of the residual aspahltene that contains high concentrations of sulfur and vanadium is probably too large to allow its diffusion into the narrow meso-pores predominantly present in this catalyst. To illustrate this in a better way it would be useful to discuss the molecular weight and size of the petroleum asphaltenes. Data on molecular weight of asphaltenes reported in literature have varied considerably, depending on the measurement technique. For example, earlier studies based on ultra-centrifugation (13,14) have shown molecular weights as high as 300,000. On the other hand, viscosity and vapor pressure osmometry (VPO) and gel permeation chromatography (GPC) methods (15) have yielded significantly lower values, typically in the 2000-8000 range. Based on NMR spectroscopic measurements molecular weights in the range 600-1000 have been calculated for condensed aromatic sheets with alkyl and alicyclic substituents. The difference between the NMR values and other measurements have been accounted for by proposing that C-C bonds and sulfur bridges (16) link several condensed polycyclic aromatic sheets to yield macromolecules of repeating structure. This has been confirmed by the work of Asoaka (17). The 2000-8000 range molecular weight measured by GPC and VPO would correspond to a stacking of four to six sheets. Molecular weights in the 40,000 range or more reflect association of particles into micelles (18). To what extent the lower molecular weight components aggregate in the resid fraction is uncertain. With regard to the size of asphaltenes, molecular radius ranging from 20-150 A have been reported in literature (12,19,20). The information available in literature thus indicate the existence of species with varying molecular size distribution in petroleum asphaltenes. The diffusion of the asphaltene molecules into the pores of the catalyst to reach the active catalytic site within the pore structure is an important requirement for the reaction. Catalytic hydrotreating reactions involving large molecular clusters in petroleum residues are diffusion limited. The ratio of molecular size to pore size is important in determining the reaction rate in residue hydrotreating, especially in asphaltene conversion. In the present studies it is noticed that catalyst (P) with pore maximum in 100 - 200 A diameter range, is able to crack a large proportion of the total asphaltenes present in the feed. This implies that a major portion of the asphaltenic species have sufficiently lower molecular size for diffusion and reaction within the catalyst's pores. The remaining portion of the asphaltenes, probably having larger molecular dimensions is unable to diffuse into the pores and consequently are not attacked by the catalyst sites. This is further confirmed by the higher molecular weight distributions of the residual asphaltenes for this catalyst (Fig. 4). The large molecular weight and high concentrations of sulfur and metals in the residual asphaltenes indicate that the catalyst is not able to attack and remove the heteroatoms from the large size asphaltene molecules. In the case of the catalysts with a high percentage of large pores the sulfur and vanadium concentrations of the residual asphaltene are significantly low. The molecular weights are also substantially low.

195 The results clearly indicate that pores larger than 200 A diameter, especially in the 8003000 A range are important for cracking a portion of large size asphaltene molecules present in the residual oil. However, the presence of large pores alone in the catalyst appears to be ~r ~6 , O - ,~-~.~t- ""...

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Figure 4. Molecular weight distribution of asphaltenes in hydrotreated products for different catalysts (Note: Low elution volumes indicate higher molecular sizes). inadequate for the over all performance. It is noticed that the catalyst P with maximum amount of mesopores in the 100 - 200 A diameter range possesses very high activity for various conversions, including the cracking of asphaltenes of smaller molecular size. However, a relatively high percentage of toluene insoluble sediments is noticed for this catalyst (Fig. 2C). The exact mechanism for the formation of these sediments is not known. One proposition is that the sediments are simply asphaltenes or asphaltene fragments precipitating as a result of the disturbance of the ratio of resins to asphaltenes during the reaction (7). The residual oil can be considered as a colloidal system consisting of oils, resins and asphaltenes. The asphaltenes remain dispersed in the less polar oil medium due to the presence of resins (12,15). The micelles of the dispersed phase contain asphaltenes and resins. The asphaltenic core of the micelles absorbs high molecular aromatic hydrocarbons from the resin fraction which absorbs further hydrocarbons, until the periphery of the micelles contains hydrocarbons having a hydrogen content that approximately corresponds to the hydrogen content of the resin dispersing agent. The micelles are in a state of equilibrium with respect to the surrounding oil phase. The solubilizing and dispersing power of resins is controlled by their degree of aromaticity. In other words, the adsorption equilibrium will be disturbed and the

196 solubility of asphaltenes will change if the nature of the resins is modified by reactions during hydroprocessing. It is believed that the sediments formed during deep conversion of heavy petroleum residues are simply asphaltenes or asphaltene fragments precipitating as a result of changes in the properties of the resin phase. The reactivities toward catalytic hydrocracking of the three major components of the heavy oil are in the order. Oils > Resins > Asphaltenes Catalyst (P) possesses a high activity for various conversions. Since the catalyst contains predominantly meso-pores and contains negligible amount of larger pores, it is possible that the rate of cracking of resins and oils in the feedstock occurs at a faster rate than that of the large molecular size asphaltenes. As a result, the ratio of resins to asphaltenes in the product will decrease. Consequently, the asphaltenes may become incompatible in the oil fraction and precipitate out as sediments. The sediment formation may probably be reduced if the catalyst contains certain amount of macro-pores in addition to the meso-pores. In catalysts containing predominantly macro-pores with insignificant amount of meso- and micro-pores, (e.g. catalyst Q), sediment formation is very low. Such catalysts show the highest activity for removal of vanadium. However, the activity for hydroconversion to lighter products as well as for sulfur removal is minimum for the macro-pore catalyst. A catalyst containing predominantly meso-pores together with some micro-and macro-pores in appropriate proportions may be expected to show a reasonably high activity for various conversions, including asphaltenes cracking, without the problem of sediment formation. ACKNOWLEDGEMENT The authors thank the members of the H-Oil task force at KNPC for their helpful suggestions and remarks during the course of this work. The authors also gratefully acknowledge the financial support and encouragement provided by KNPC and KISR managements. REFERENCES 1. 2. 3. 4.

G. Heinrich, M. Valais, M. Passol, and B. Chapotel. Thirteenth World Petroleum Congress, Paper No. 18 (1), 1991. I.E. Maxwell; J. E. Naber; and K. P. de Jang, Appl. Catal. A: General, 113 (1994) 153. S. Kamatsu, Y. Hori and S. Shimizu. Hydrocarbon Processing, May, 1985, p. 42.

5.

I. Mochida, X. Z. Zhao, K. Sakanishi, S. Yamamoto; H. Takashima and S. Vemura, Ind. Eng. Chem, Res. 28 (1989) 418. J.F. Kriz and M. Ternan. Stud. Surf. Sci. Catal. 73 (1992) 31.

6.

W.I. Beaton and R. J. Bertolacini, Catal. Rev. Sci. Eng, 33 (1991) 281.

7. 8.

M. Absi-Halabi, A. Stanislaus and D. L. Trimm, Appl. Catal., 72 (1991) 193. M. Absi-Halabi, A. Stanislaus, F. Owaysi; Z. Khan and S. Diab, Stud. Surf. Sci. Catal, 53 (1990)201.

197 9.

A. Stanislaus, M. Absi-Halabi, F. Owaysi and Z. H. Khan. Effects of temperature and pressure on Catalytic hydroprocessing of Kuwait Vacuum Residues. KISR Publication No. 2754 (1988). 10. T. Takatsuka, Y. Wada, S. Hirohama and Y. Fukui, J. Chem. Eng. Japan, 22 (1989) 298. 11. I.A. Wiehe, Ind. Eng. Chem. Res, 32 (1993) 2447. 12. J. G. Speight, Upgrading of Heavy Oils and Residue: Nature of the problem. "Catalysis in the Energy Scene" Elsevier, 1984, pp. 515-527. 13. R. S. Winford, J. Inst. Petroleum, 49 (1963) 215. 14. S. Wales and V. Waarden, ACS Div. Petrol. Chem. Preprints, 9 (1964) B-21.

In

15. J. G. Speight, ACS Div. Petrol. Chem. Preprints 32, (1987) 413. 16. J. G. Speight and S. E. Moschopedis. In "Chemistry of Asphaltenes (Edited by J. W. Bunger and N. C. Li), Advances in Chemistry Series, 195 (1981) 1. 17. S. Asaoka, S. Nakata, Y. Shiroto and C. Takeuchi, Ind. Eng. Chem. Process Design and Dev, 22 (1983) 242. 18. J.P. Dickie and T. F. Yen, Anal. Chem, 39 (1967) 1847. 19. R.J. Quan, R. A. Ware, C. W. Hung and J. Wei. Advances in Chem. Eng, 14 (1988) 95. 20. E. W. Baltus and J. L. Anderson, Chem. Eng. Sci, 38 (1983) 1959.

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Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 1996 Elsevier Science B.V.

199

ORIGIN OF THE L O W REACTIVITY OF ANILINE AND H O M O L O G S IN HYDRODENITROGENATION. M. Cailant', K. A. Holder b, P. Grange" and B. Delmon" a Unit~ de Catalyse et Chimie des Mat~riaux Divis~s, Universit~ Cathofique de Louvain, 2/17 Place Croix du Sud, 1348 Louvain-la-Neuve, Belgium. b BP Oil, M&S Technology Development Unit, Chertsey Road, Sunbury-on-Thames, Middlesex TW16 7LN, United Kingdom. ABSTRACT This contribution deals with the origin of the low reactivity of aniline and its homologs in hydrodenitrogenation of petroleum fractions. This low reactivity is very surprising, because aniline, in the absence of other nitrogen containing molecules, reacts readily. We performed a series of experiments where two model molecules in the feed were reacted competitively. The experiments were carried out with an industrial sulphided phosphorus-containing NiMo/q,-AI203 catalyst. The influence of the H2S pressure was also studied. The low reactivity of aniline in mixtures is due to the inhibition of the first step in the main pathway to HDN, namely the hydrogenation of the benzenic ring, by both basic (pyridine and indoline) and non-basic (indole and pyrrole) heterocyclic molecules. Compounds containing nitrogen in the cycle adsorb much more strongly on the catalyst than aniline. These results are discussed together with results published in literature. 1. INTRODUCTION The industrial interest in hydrodenitrogenation (HDN) arises from the growing necessity of using heavy crudes as the source for transportation fuels. Heavy crudes are characterised by a high nitrogen content [ 1]. Hydrodenitrogenation became recently still more critical because of new fuel specifications (e.g. low aromatic content and deep hydrodesulphurization for diesel fuel). The complex hydrorefining processes to be used necessitate a deep nitrogen removal as a prerequisite. The main (but not exclusive) reason is that nitrogen blocks the acidic sites necessary for mild hydrocracking. The emphasis is increasingly laid on aromatic amines. These compounds exist as such in the petroleum crudes. They also constitute stable intermediate products formed by the partial reaction of polycyclic nitrogen compounds [2]. Aniline and its homologs are extremely refractory to HDN. For instance, Toulhoat and Kessas [3] and Kasztelan et al. [4] reported that the alkyl-aniline content of a coker gas oil subjected to hydrotreatment increased after the treatment. This can be explained by the fact that the alkyl anilines are produced from the polycyclic compounds, but have a low reactivity in the reaction conditions. This result is very surprising since, when alone in the feed, aromatic amines are highly reactive. The object of the present work was to elucidate the origin of the low reactivity of aniline and homologs when treating an industrial feedstock. We studied the HDN reaction pathway of aniline and we considered two important industrial parameters: the effect of the H2S partial pressure and the competition between aniline and an important category of nitrogen compounds present in petroleum cmdes, namely the

200 heterocyclic nitrogen compounds. The heterocyclic compounds tested in the frame of our experiments were basic (pyridine and indoline) as well as non-basic (indole and pyrrole).

aniline

0

NH2

cyclohexylamine

cyclohexene

cyclohexane

NH2

benzene Figure 1: Aniline reaction pathway Two pathways are reported in the literature for aniline HDN (Figure 1). In pathway A, the hydrogenation of the aromatic ring precedes the breaking of the C-N bond. The second pathway (B) is the direct hydrogenolysis of the aromatic amine. The occurrence of pathway B can be proven by the presence of benzene in the reaction products. It has indeed beea shown by several authors [5-8] that, in the conditions of hydrotreating, the hydrogenation of benzene was negligible with respect to the hydrogenation of aniline. Different results in the literature concern the relative importance of pathway A versus pathway B. Pathway B (direct hydrogenolysis) was reported as the exclusive pathway on an oxide CoMo-A1203 catalyst [9]. In the studies of Moreau et al. [7,10], the occurrence of both pathways was considered. The nature of the promoter appeared to be determinant: pathway A (hydrogenation) was predominant on a sulphided NiMo-A1203 catalyst while pathway B (direct hydrogenolysis) was the main pathway on a CoMo-A1203 catalyst. Finiels [11], Oliv6 [8], Geneste [5] and Schulz [2,12] tested in similar conditions aniline compounds on sulphided NiW-A1203 and NiMo-A1203. Their results showed that pathway A (hydrogenation) was the predominant pathway. There is, as yet, no report in the literature on the competition between aromatic amines and non-basic nitrogen heterocycles. In contrast, several authors investigated the competition between aromatic amines and basic nitrogen heterocycles. Perot et al. [13-17] reacted opropylaniline and 6-methylquinoline over a sulphided NiMo-A1203 catalyst and found that the conversion of o-propylaniline was much lower in the presence of 6-methylquinoline than when reacted alone. A similar result was obtained when 2,6-diethylaniline was reacted in competition with 1,2,3,4-tetrahydroquinoline: the latter compound strongly inhibited the reactivity of 2,6diethylaniline. The conversion and product distribution of 1,2,3,4-tetrahydroquinoline was not influenced by the presence of 2,6-diethylaniline. These results were explained by the difference of gas-phase proton affinity between the two compounds, 1,2,3,4-tetrahydroquinoline being more basic than 2,6-diethylaniline and thus more strongly adsorbed. It was concluded that the unshared electron pair of the nitrogen atom is involved in the adsorption on the active sites. The inhibition of aniline (alkyl-substituted or not) by quinoline or 1,2,3,4-tetrahydroquinoline has also been reported by Moreau et al. [ 18-20], Cocchetto and Satterfield [21] and Toulhoat and Kessat [3]. The influence of the H2S partial pressure on the reactivity of aromatic amines has not received much attention in the literature. The only indication we found comes from a study of Yan et al. [22]. They noticed that the addition of H2S slightly decreased the conversion of oethylaniline.

201

2.

E X P E R I M E N T A L

The catalytic tests were performed in a bench-size continuous-flow reactor as described elsewhere [23, 24]. The catalyst tested was a commercial NiMoP catalyst supported on 'talumina (composition 2.9 wt% Ni, 12.6 wt% Mo, 2.9 wt% P). We selected a catalyst containing phosphorus because the most active HDN catalyst presently available on the market contains this additive. The catalyst was in the oxide state (NiO, MOO3) when introduced in the reactor. It was pretreated in situ according to a procedure which ensured an optimal catalyst sulphidation. The catalyst was first heated to 423 K under Ar and left at this temperature for half an hour. The activation gas - - a H2S(15 vol%)/H2 m i x t u r e - was introduced afterwards and the temperature raised, first up to 573 K where it was maintained for half an hour, then up to 673 K for one hour. Between each step of the pretreatment procedure, the heating rate was 0.17 Ks -1. The total gas flow rate was 1.67 10-6 m3s-1 during the whole process. At the end of the pretreatment, the catalyst was maintained under the H2S-H2 atmosphere and the temperature was lowered to 573 K before starting the reaction. The reaction conditions were: - weight of catalyst: - temperature: - total pressure: - hydrocarbon feed flow rate to the reactor: - H2 flow rate to the reactor:

8 10-4 kg 573 K 5 106pa 8.33 10-9 m3s-1 at STP 8.33 10-6 m3s-1 at STP.

The conversion levels were evaluated at steady state catalyst activity: sample analysis gave constant results after 10 to 15 hours on line. The reported results correspond to data collected after this time. The reacting gas phase resulted from the mixing of pure H2 and a hydrocarbon feed. The hydrocarbon feed contained the nitrogen model compounds, CS2 as H2S precursor and a hydrocarbon solvent (n-heptane) which is inert in the reaction conditions. We report in Table 1 the composition of the hydrocarbon feed (and the corresponding gas phase) which we take as standard feed. To study the influence of the H2S partial pressure, we varied the CS2 concentration of the hydrocarbon feed between 0 and 0.62 M. This gave rise to a H2S partial pressure in the gas phase comprised between 0 and 132 kPa. We also investigated the effect of the H2 partial pressure by testing the standard feed with a mixture of H2-Ar (50-50 vol%) instead of pure H2: the H2 partial pressure was thus 2110 instead of 4220 kPa. The reaction samples were analysed by temperature programmed gas chromatography using a Hewlett Packard instrument (model 428) equipped with a 25 m capillary DB-5 column and a FID detector. The concentration of reactants and products were calculated using n-heptane as internal standard. Table 1 Composition of the standard model feed. COMPOSITION OF THE HYDROCARBON FEED COMPOSITION OF THE REACTING GAS PHASE

molar conc. (M)

% weight

0.0246 0.0504 0.2167 6.6857

0.33 0.85 2.38 94.45

aniline

indole CS2

n-heptane

partial pressure (kPa) H2 H2S n-heptane methane indole aniline

4219.8 45.6 703.8 22.8 5.3 2.6

202 3. RESULTS

We report first the effect of indole on aniline reactivity (Table 2). Indole strongly inhibited the conversion of aniline: aniline conversion decreased from 90% when reacted alone to less than 20% when reacted with indole. The two products of aniline HDN were cyclohexene and cyclohexane. No benzene was detected. This result confirms that, on sulphided NiMo catalysts, the HDN of aromatic amines proceeds essentially through the hydrogenation of the aromatic ring and the subsequent formation of cyclohexylamine (path A, Figure 1). Cyclohexylamine must be very reactive since it was not found in the reaction products. This result shows that aniline reactivity is controlled by the rate of aniline hydrogenation to cyclohexylamine. In order to further prove the validity of this interpretation, we studied the reactivity of benzene and cyclohexylamine and we investigated the effect of the hydrogen partial pressure on aniline reactivity. Cyclohexylamine and benzene were reacted in the same conditions as reported earlier, alone and in the presence of indole (in order to simulate a possible inhibiting effect). Cyclohexylamine was found to be completely converted either in the presence or the absence of indole in the reacting gas phase. Concerning benzene, no conversion was observed. The results concerning the influence of the H2 partial pressure are reported on Figure 2. A first order relation was observed. Table 2 Aniline reactivity. aniline reacted in competition with indole

aniline reacted alone

19

91

9 5

85 4

% aniline conversion p r o d u c t distribution:

% cyclohexane % cyclohexene

% aniline conversion

Reaction condition: T= 573K Ptot = 5 MPa PH2S = 45.6 kPa Paniline = 2.6 kPa Pindole = 5.3 kPa

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2

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H2 concentration in the reacting gas phase (MPa)

Figure 2: Effect of H2 on aniline reactivity

203 The inhibition exerted by indole on aniline is further illustrated in Figure 3. The experiment consisted of a three-stage reaction: 1: reaction of aniline alone, 2: reaction of the mixture aniline-indole, 3: reaction of aniline alone. This experiment clearly indicated a drop of aniline reactivity when adding indole in the reacting gas phase (stage 1 to stage 2). When removing indole from the gas phase (stage 2 to stage 3), the aniline conversion was restored to its previous level. Following our results on the couple aniline-indole, we extended our study to other nitrogen heterocycles: indoline, pyrrole and pyridine [23-25]. No difference was found in the intensity of the inhibition exerted by the heterocycles on aniline reactivity (Figure 4). In all the experiments, the reactivity and the product distribution of the nitrogen heterocycles were not modified by the presence of aniline.

o

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stage I 9 stage 2 9 stage 3 9 P aniline = 2.6 k P a P aniline = 2.6 k P a P aniline = 2.6 k P a P indole = 5.3 k P a 100

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0

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Figure 3: Inhibition of aniline reactivity upon indole addition The last result concerns the effect of the H2S partial pressure on aniline reactivity. H2S was found to inhibit aniline reactivity (Figure 5): aniline conversion dropped from 47% in the absence of H2S in the reacting gas phase to 15% at a H2S partial pressure of 132 kPa. We investigated the effect of the H2S partial pressure on aniline alone in order to evaluate whether the effect of H2S was modified by the inhibition exerted by indole. Aniline conversion also decreased with the increase of the H2S partial pressure. The same result was observed when dimethyldisulphide (instead of CS2) was used as H2S precursor [26].

204 R e a c t i o n condition: T=573K Ptot = 5 M P a o tl9 - A 9 ~lPa

% aniline conversion 100

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aniline in competition with"

75

A indole-indoline

50

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25

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1

2 3 4 5 6 heterocyclic compound concentration in the reacting gas phase (kPa)

Figure 4: Effect of nitrogen heterocyclic compounds on aniline reactivity

% aniline conversion 50 40

R e a c t i o n condition: T= 573 K Ptot = 5 M P a P H 2 = 4.2 M P a aniline = 2.6 k P a indole = 5.3 k P a

30 20 O 10 I

0

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50 100 150 H2S concentration in the reacting gas phase (kPa)

Figure 5: Effect of H2S on aniline reactivity

205 4. DISCUSSION The products of aniline are cyclohexane and cyclohexene. As benzene cannot be converted to cyclohexane, the absence of benzene in the product distribution indicates that the HDN of aniline occurs through the hydrogenation of aniline to cyclohexylamine (Figure 1, path A). This result is in agreement with the literature as far as nickel promoted MoS2 catalysts are concerned [2,5,7,8,10-12]. Cyclohexylamine cannot be detected in the reaction products because, as we showed directly, cyclohexylamine is very reactive (conversion 100 %) even in the presence of indole. This high reactivity of cyclohexylamine is in agreement with literature [5,11,27,28]. This result implies that the hydrogenation of aniline to cyclohexylamine is the rate limiting step of the reaction. Our results concerning the effect of the H2 partial pressure conf'Lrrn this interpretation. As a consequence, we can ascribe the inhibition exerted by H2S and the heterocyclic nitrogen compounds on aniline reactivity to the poisoning of the hydrogenation function of the catalyst. It is well known that H2S inhibits the hydrogenation function of hydrotreating catalysts. This effect has been reported in various studies dealing with the reactivity of nitrogen compounds (quinoline [18,22,29-35], 1,2,3,4-tetrahydroquinoline [36-38], 7,8benzoquinoline [39], 5,6-benzoquinoline [40], pyridine [41-44], piperidine [45] and indole [46]). In these studies, the catalysts tested were sulphided NiMo-~,A1203, CoMo-~,A1203 or NiW-~,A1203. An inhibiting effect of H2S has also been reported for hydrogenation reactions different from those involved in HDN reactions [47-52]. Concerning the inhibiting effect of nitrogen heterocycles on the hydrogenation of aniline to cyclohexylamine, we confirmed our interpretation by verifying that indole strongly inhibits a typical hydrogenation reaction: the reduction of naphthalene to tetraline [23]. Other authors [34, 53-57] have also reported that nitrogen heterocycles inhibit the hydrogenation function of hydrotreating catalysts. Our results show that the nitrogen heterocycles have a higher adsorption coefficient on the hydrogenation sites than aniline. But our experiments do not allow a differentiation between the adsorption coefficients of the various heterocyclic compounds. The higher adsorption coefficient on the hydrogenation sites of pyridine in comparison with aniline is in agreement with the results of Nagai et al. [53]. These authors correlated the adsorption constant of nitrogen compounds (on the hydrogenation sites of a NiMo-~,A1203 catalyst) with their gasphase basicity. The nitrogen compounds studied were acridine, quinoline, pyridine, v-picoline and aniline. A linear correlation was found. The authors concluded that the nitrogen compounds adsorbed on Br0nsted acid sites, which were supposed to be OH groups or SH groups adjacent to an anion vacancy on the surface of the sulphided catalyst [58]. They inferred that the hydrogenation sites involve these Br0nsted acidic sites. The difference of gas-phase basicity can explain the inhibiting effect exerted by pyridine on aniline. The same factor was invoked by Perot [13] to explain the inhibiting effect exerted by 1,2,3,4-tetrahydroquinoline on 2,6diethylaniline. In the case of indole, a similar explanation can be proposed since indole is readily hydrogenated to indoline, a basic compound. A second factor which obviously plays a role in the adsorption of molecules on the hydrogenation sites is the n electron density. This factor was considered to be predominant in the study of Moreau et al. [10]. The role of the n electrons can explain the fact that, in the study of Nagai et al [53], cyclohexylamine and piperidine did not fit the linear correlation between the adsorption coefficient and the gas phase basicity: the adsorption coefficient of these saturated nitrogen compounds was about twice lower than expected on the basis of the above mentioned correlation. In our experiments, the high adsorption coefficient of the pyrrolic ring can be related to the very high n character of this heterocycle [59].

206 5. C O N C L U S I O N Aniline reactivity is strongly inhibited by the heterocyclic nitrogen compounds. The heterocyclic compounds tested in the frame of our experiments were basic (pyridine and indoline) as well as non-basic (indole and pyrrole). This inhibition is reversible, namely disappears as the inhibiting substance is removed. A second factor, coming in addition, is the H2S partial pressure: aniline reactivity is inhibited by H2S. The consequence is that, although aniline is among the most reactive nitrogen compounds when reacted alone, it becomes difficult to decompose in the presence of other nitrogen compounds. This conclusion very likely also applies to real feeds which contain a large amount of sulphur containing molecules (source of H2S). These two factors explain why aniline compounds which are formed during the hydrotreatment of industrial feedstocks, are very stable. There is no doubt that this is the major origin of the low degree of hydrodenitrogenation in many industrial feeds. ACKNOWLEDGMENT We gratefully acknowledge BP Oil, Research Centre Sunbury, UK, for supporting this work. REFERENCES

1. 2. 3. 4. 5. 6. 7. 8. 9. 10. 11. 12. 13. 14. 15. 16. 17. 18.

J.G. Speight, in Chemical Industries, vol 41 : Fuel Science and Technology Handbook, Marcel Dekker Inc., New York, 1990, p. 71. H. Schulz, M. Schon and N. M. Rahman, in L. Cerveny (Editor), Studies in Surface Science and Catalysis, vol 27 : Catalytic Hydrogenation, Elsevier, Amsterdam, 1986, p. 201. H. Toulhoat and R. Kessas, Revue de rlnstitut Franqais du P6trole, 41 (1986) 511. S. Kasztelan, T. des Courti~res and M. Breysse, Catal. Today, 10 (1991) 433. P. Geneste, C. Moulinas and J. L. Oliv6, J. Catal., 105 (1987) 254. Y. Liu, F. E. Massoth and J. Shabtai, Bull. Soc. Chim. Belg., 93 (1984) 627. C. Moreau, J. Joffre, C. Saenz and P. Geneste, J. Catal., 122 (1990) 448. J.L. Oliv6, S. Biyoko, C. Moulinas and P. Geneste, Appl. Catal., 19 (1985) 165. A.K. Aboul-Gheit and I. K. Abdou, J. Inst. Pet., 59 (1973) 188. C. Moreau, C. Aubert, R. Durand, N. Zmimita, and P. Geneste, Catal. Today, 4 (1988) 117. A. Finiels, P. Geneste, C. Moulinas and J. L. Oliv6, Appl. Catal., 22 (1986) 257. W. B/Srhinger and H. Schulz, Bull. Soc. Chim. Belg., 91 (1991) 831. G. Perot, Catal. Today, 10 (1991) 447. G. Perot, S. Brunet, C. Canaff, and H. Toulhoat., Bull. Soc. Chim. Belg., 96 (1987) 865. N. Gnofam, L. Vivier, S. Brunet, J. L. Lemberton and G. Perot, Catalysis Letter, 2 (1989) 81. L. Vivier and G. Perot, in Proceedings of the 12th Ibero-American Symposium on Catalysis, vol 2, Rio de Janeiro, 1990, p. 582. L. Vivier, S. Kasztelan and G. Perot, Bull. Soc. Chim. Belg., 100 (1991) 801. C. Moreau, L. Bekakra, A. Messalhi, J. L. Oliv6 and P. Geneste, in M. L. Occelli and R. G. Anthony (Editors), Studies in Surface Science and Catalysis, vo150 : Hydrotreating Catalysts, Elsevier, Amsterdam, 1989, p. 107.

207 19. C. Moreau, L. Bekakra, R. Durand, N. Zmimita and P. Geneste, in M.L. Occelli and R.G. Anthony (Editor), Studies in Surface Science and Catalysis, vol 50: Hydrotreating Catalysts, Elsevier, Amsterdam, 1989, p. 115. 20. C. Moreau, L. Bekakra, P. Geneste, J. L. Oliv6, J. C. Duchet, M. J. Tilliette and J. Grimblot, Bull. Soc. Chim. Belg., 100 (1991) 841. 21. J.F. Cocchetto and C. N. Satterfield, Ind. Eng. Chem. Process Des. Dev., 20 (1981) 49. 22. J.-W. Yan, T. Wakatsuki, T. Obara, and M. Yamada, Sekivu Gakkaishi, 32 (1989) 129. 23. M. Callant, P. Grange, K. A. Holder and B. Delmon, Bull. Soc. Chim. Belg., 91 (1991) 823. 24. M. Callant, PhD Thesis, Universit6 Catholique de Louvain, Louvain-la-Neuve, Belgium, 1993. 25. M. Callant, K. A. Holder, P. Grange and B. Delmon, in preparation. 26. M. CaUant, K. A. Holder, P. Grange and B. Delmon, accepted for publication in J. Mol. Catal. 27. E.W. Stem, J. Catal., 57 (1979) 390. 28. S. Eijsbouts, C. Sudhakar, V. H. J. de Beer and R. Prins, J. Catal., 127 (1991) 605. 29. S.H. Yang and C. N. Satterfield, J. Catal., 81 (1983) 168. 30. S.H. Yang and C. N. Satterfield, Ind. Eng. Chem. Process Des. Dev., 23 (1984) 20. 31. C.N. Satterfield and S. Giiltekin, Ind. Eng. Chem. Process Des. Dev., 20 (1981) 62. 32. C.N. Satterfield, C. M. Smith and M. Ingalis, Ind. Eng. Chem. Process Des. Dev., 24 (1985) 1000. 33. S. Giiltekin, M. Khaleeq and M. A. A1-Saleh, Ind. Eng. Chem. Res., 28 (1989) 729. 34. M.V. Bhinde, S. Shih, R. Zawadzki, J. R. Katzer and H. Kwart, in H. F. Barry and P. C. H. Mitchell (Editors), Chem. Uses Molybdenum, Proc. Int. Conf., 3rd, Climax Molybdenum Co., Ann Arbor, 1979, p. 184. 35. S. S. Shih, J. R. Katzer, H. Kwart and A. B. Stiles, A. C. S. Div. Pet. Chem., 22 (1977) 919. 36. A. Olalde and G. Perot, Appl. Catal., 13 (1985) 373. 37. L. Vivier, P. D'Araujo, S. Kasztelan and G. Perot, Bull. Soc. Chim. Belg., 100 (1991) 807. 38. S. Brunet and G. Perot, React. Kinet. Catal. Lett., 29 (1985) 15. 39. K. Malakani, P. Magnoux and G. Perot, Appl. Catal., 30 (1987) 371. 40. J. Shabtai, G. J. C. Yeh, C. Russel and A. G. Oblad, Ind. Eng. Chem. Res., 28 (1989) 139. 41. F. Goudriaan, H. Gierman, and J. C. Vlugter, J. Inst. Pet., 59 (1973) 40. 42. R.T. Hanlon, Energy Fuel, 1 (1989) 424. 43. C.N. Satterfield, M. Modell, and J. A. Wilkens, Ind. Eng. Chem. Process Des. Dev., 19 (1980) 154. 44. M. Cerny, Coll. Czech. Chem. Commun., 47 (1982) 1465. 45. M. Cerny, Coll. Czech. Chem. Commun., 47 (1982) 928. 46. F.E. Massoth, K. Balusami, and J. Shabtai, J. Catal., 122 (1990) 256. 47. G. Perot, S. Brunet, and N. Hamze, in M. J. Philips and M. Ternan (Editors), Proc. 9th Int. Congress Catal., Calgary 1988, vol 1, The Chemical Institute of Canada, Ottawa, 1988, p. 19. 48. Lee and Butt, J. Catal., 49 (1977) 320. 49. C.N. Satterfield and G. W. Roberts, AICh J., 14 (1968) 159. 50. A.V. Sapre and B. C. Gates, Ind. Eng. Chem. Process Des. Dev., 21 (1982) 86. 51. R.J.H. Voorhoeve and J. C. M. Stuiver, J. Catal., 23 (1971) 228. 52. S. Giiltekin, S. A. Ali and C. N. Satterfield, Ind. Eng. Chem. Process Des. Dev., 23 (1984) 181. 53. M. Nagai, T. Sato, and A. Aiba, J. Catal., 97 (1986) 52. 54. V. Moravek, J.-C. Duchet and D. Comet, Appl. Catal., 66 (1990) 257. 55. F.W. Kirsch, H. Shallt and H. Heinemann, Ind. Eng. Chem., 51 (1959) 1379.

208 56. 57. 58. 59.

M. Nagai and T. Kabe, J. Catal., 81 (1983) 440. F.E. Massoth and J. Miciukiewicz, J. Catal., 101 (1986) 505. F.E. Massoth and C. L. Kibby, J. Catal., 47 (1977) 300. P. Sykes, in A Guidebook to Mechanism in Organic Chemistry, Longman, London and New York, 1981, p. 161.

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All fights reserved.

209

DEEP HDS OF MIDDLE DISTILLATES USING A HIGH LOADING CoMo CATALYST S. Mignard a, S. Kasztelana, M. Dorbon b, A. Billon a and P. Sarrazin c ab~:stitut Frangais du Pdtrole, 1 & 4 av. de Bois Prdau, 92506 Rueil-Malmaison, France stitut Frangais du Pdtrole, CEDI, BP3, 69390 Vernaison, France Cprocatalyse, 212 av. Paul Doumer, 92506 Rueil-Malmaison, France ABSTRACT The effect of the nature of the support material (),-alumina or alumina-based proprietary support) and of the metal loading on the catalytic properties of CoMo catalysts have been studied. For a conventional 3wt% CoO and 14wt% MoO 3 catalyst, the use of the proprietary support has led to a 50% increase of the toluene hydrogenation activity. From Transmission Electronic Microscopy experiments, no major morphological differences between catalysts have been found suggesting an increase in the intrinsic catalytic activity rather than an increase in the number of sites. Measurements of the catalytic properties of a conventional industrial catalyst and the new high loading CoMo catalyst manufactured with the proprietary support have been carried out with natural feedstocks. The new high loading catalyst exhibits a gain in iso-conversion temperature of 8~ In addition, for the same hydrodesulfurization level, this new catalyst has not exhibited an increase of the hydrogen consumption. 1. INTRODUCTION The reduction of sulfur content of middle distillates down to very low levels (500 ppm or less for diesel fuel) will be imposed nearly world-wide in the near future. In order to reach these targets without large capital expenses, very high performance hydrodesulfurization (HDS) catalysts are needed [1, 2]. IFP and Procatalyse have worked extensively to improve the HDS activity of CoMo catalysts. Improvements have been possible due to constant efforts to isolate the main parameters which determine the catalyst performances. CoMo HDS catalysts have been extensively studied and it is well known that the active species is molybdenum sulfide and that cobalt is a promotor [3-8]. Proper design of the support and adjustment of molybdenum and cobalt composition are very clearly the key points for HDS catalyst design. In the early 80's, Bachelier et al. [9] demonstrated that catalytic efficiency depends on the Mo loading of the catalyst. We have studied the effect of molybdenum loading on HDS activity and found an optimum metal loading of about 6 wt% for a selected y-alumina support (Figure 1). Such a behaviour has been rationalized by a change in the molybdenum sulfide particle size [9]. With the addition of cobalt, the activity per molybdenum atom increases at low cobalt content and then reaches a plateau as shown in Figure 2. Such a result confirms earlier works done by Bachelier et al. on NiMo catalysts [10, 12]. Today, the preferred interpretation is that cobalt atoms decorate the molybdenum sulfide particles. This hypothesis was predicted by a geometrical model [8] and confirmed experimentally [12].

210

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Co/Mo atomic ratio Figure 2. Influence of the Co/Mo atomic ratio on HDS activity at constant Mo loading of CoMo catalysts

Higher activity catalyst can be achieved by increasing the metal content up to the limit of the support capacity, although the molybdenum efficiency decreases. Consequently, we have worked on the different steps of a catalyst preparation (cartier selection and shaping, Co/Mo ratio, molybdenum and cobalt introduction methods, promotor, thermal and hydrothermal treatments...) and examined the activity of the resulting catalyst at each step. In this work, we report a comparative study of 3%COO and 14%MOO3 catalysts made with either a ~,-alumina support or a proprietary Procatalyse support. We have compared their toluene hydrogenating activity and characterized the sulfided catalysts by Transmission Electronic Microscopy (TEM). We have also compared two industrial catalysts from Procatalyse : a 3wt% CoO and 14wt% MoO3 on a conventional y-alumina support (HR306C) and a high loading catalyst (4wt% and 18wt% in CoO and MOO3) on the proprietary support (HR316C). Their HDS and aromatic hydrogenation performances have been compared under industrial conditions on either a straight run gas-oil (SRGO) or a mixture of SRGO and light cycle oil (LCO). The results clearly show the gain in activity obtained with this new high loading CoMo catalyst. 2. EXPERIMENTAL

2.1. Catalyst preparation The catalysts studied in this work have been prepared by using either a conventional yalumina (SBET=240 m2/g, pore volume=0.5 cm3/g) or a proprietary support both supplied by Procatalyse. Molybdenum or cobalt and molybdenum have been introduced by wet impregnation of the extrudates by using aqueous solutions of Co(NO3) 2 and Mo7024(NH4)6. Then, the catalyst have been dried over night and calcined under air (7.5 vol% H20 ) at 500~ for 4

211 hours. The catalysts with a CoMo content of 3wt% of CoO and 14wt% of MoO 3 have been prepared with the y-alumina support (CoMo-A) and with the proprietary support (CoMo-B). In addition, two industrial catalysts have been studied. The CoMo-C catalyst is the commercial HR306C (3wt% of CoO, 14wt% of MoO3 on ,/-alumina) and the CoMo-D catalyst is the new industrial high metal loading catalyst HR316C (4wt%CoO, 18wt%MoO 3 on a proprietary support).

2.2. Toluene hydrogenation test The catalyst activities for toluene hydrogenation have been measured in a fixed bed reactor at a pressure and a temperature close to those of liDS industrial unit conditions: 6MPa and 350~ In order to maintain the catalyst in the sulfide state, H2S (ex DMDS) is continuously added. 2.3. TEM experiments Before TEM experiments, the samples were sulfided with an H2S/H 2 gas mixture (15/85vol/vol) with a flow 41/h at atmospheric pressure. The temperature has been increased from room temperature up to 400~ at a rate of 5~ and maintained at 400~ for 2 hours. The H2S/H 2 mixture was then replaced by helium and the samples cooled to room temperature. The reactor cell was isolated then transferred in a glove-bag under helium. Three or four extrudates were then crushed to a powder under ethanol and the powder was deposited onto a coated toper-grid. Aiter wetting, the grid was introduced into the TEM prechamber. It is one more time wetted under vacuum in the prechamber for 5 minutes before exposure to the electron beam. The instrument used was a JEOL 2010 with a LaB6-filament operating at 200kV with an objective aperture of 30~m.

2.4. Gas-oil hydrotreating tests The tests have been carded-out with continuous flow, once-through, pilot unit. The reactor (1 liter) was filled with 600 cm3 of catalysts between two beds of alumina balls. The feed was mixed with pure hydrogen and preheated before enterdng the reactor. Potentially dissolved H2S was removed by submitting the product to a caustic soda wash. Before a test run and after each change of operating conditions or of feedstock, the unit was stabilised for at least 48 hours. Test runs were carried out during at least 12 hours and only the analyses of the products collected during the tests run have been taken into account in order to estimate the performances of a catalyst. The sulfur, nitrogen and aromatics contents of the feeds and products were obtained by X Ray Fluorescence, Lumazote and 13C M R respectively. Two feeds have been used, a Middle East medium straight run gas-oil (SRGO) and a mixture (50/50vol) of the same SRGO and an FCC light cycle oil (LCO). The analysis of the SRGO, LCO and the mixture are reported in Table 1. As it can be seen in Table 1, the aromatics content of the LCO is very high which is confirmed by the high specific gravity and poor cetane number. Because sulfur and nitrogen are essentially in aromatic rings, sulfur and nitrogen contents of the LCO are also very high. Industrially, it is very common to treat mixtures of SRGO and LCO. HDS activities have been determined on the SRGO/LCO mixture by varying the temperature at 6MPa total pressure. Hydrogenation activity has been determined with the SRGO feed by varying the total pressure at a temperature of 326~ Before use, the catalysts were sulfided in-situ with SRGO spiked with dimethyldisulfide (DMDS).

212 Table 1 Analysis of SRGO, LCO and their mixture (50/50vol) Specific gravity Sulphur (wt%) Nitrogen (wt ppm) Cetane number (-) Aromatics (wt %) ASTM distillation (~ IBP 5% vol. 10% vol. 50% vol. 90% vol. 95% vol.

SRGO 0.853 1.49 100 55 31

LCO 0.941 2.80 570 21 83

SRGO + LCO 0.897 2.16 334 38 57

219 241 255 302 352 363

210 241 242 269 339 360

213 240 248 285 346 362

3. RESULTS AND DISCUSSION

3.1. Comparison of supports The catalysts CoMo-A and CoMo-B have been prepared in the lab with a conventional CoMo loading i.e. 3wt% and 14wt% respectively in CoO and MoO 3 and tested in toluene hydrogenation test. The relative toluene hydrogenation activity for the CoMo-B catalyst is 150% of the value obtained for the CoMo-A catalyst (Table 2). Thus, the proprietary support provides a 50% gain in hydrogenating activity. This could be due either to an increase of the number of active sites or an increase in the specific activity of each site or both. Table 2 Relative toluene hydrogenation activity of 3%COO-14%MOO 3 catalysts (a.u.) Catalyst CoMo-A CoMo-B

Support y-alumina proprietary

Hydrogenation activity 100 150

To determine the influence of the support on the morphology of the active phase, some TEM investigations have been performed on the sulfided catalysts. About 200 particles have been counted for each sample and the particle size distribution and the number of slabs per particle have been determined. The distribution of the number of slabs per CoMo particle is reported in Figure 3. As can be seen, the use of the proprietary support modifies the distribution of the number of slabs. More mono-slab particles and less double-slabs particles (respectively 70/25%) are observed with CoMo-B than with CoMo-A (respectively 54/38%). The distribution by length of particles is reported in Figure 4 for the two CoMo catalysts. The use of the proprietary support appears to lead to an increase in the number of particles above 50! at the expense of the particles below 10!.

213

Figure 3. Distribution of the number of slabs per particle for the CoMo-A and CoMo-B catalysts

Figure 4. Particle size distribution by length of particles for the CoMo-A and CoMo-B catalysts

Table 2 shows that the use of the new proprietary support leads to a better catalytic activity in toluene hydrogenation. At the same time, the TEM investigations show that the CoMo phase is better dispersed with the proprietary support since the number of mono-slabs particles is larger. This fact goes in the way of a higher number of active sites and so, a higher activity. On the other hand, the length of the particles tends to increase with the use of the new support. That means that the number of molybdenum atoms per particle is larger in this case and considering that the active sites are located on the edges of the particules, it means that there is a smaller number of active sites per molybdenum in the catalyst. So, two antagonist effects are observed. It would be necessary to perform more detailled experiments, particularly by using catalysts having the same dispersion. This would be probably easier at lower CoMo content. Nevertheless, if there is an increase of the number of active sites, it is not enough to explain such a gain in activity. So, the gain in activity would be more probably due to a higher activity per active sites.

3.2. Effect of high metal loading The increased metal loading produces an increase in catalytic activity. To evaluate the catalytic performances on various feedstocks, catalysts prepared by Procatalyse at the industrial scale have been used. In order to obtain a very active CoMo catalyst, Procatalyse has decided not only to use a new proprietary support but also to use a high CoMo loading. We have compared a CoMo catalyst prepared with a ~/-alumina support and a classical CoMo content (CoMo-C) and a new high loading CoMo catalyst based on the proprietary support (CoMo-D). The catalytic activities have been measured with either the pure SRGO feed or with the SRGO/LCO mixture.

214

Hydrodesulfurization The HDS activities of the CoMo-C and CoMo-D catalysts have been measured with the SRGO/LCO mixture. The HDS conversions from 320~ to 350~ at constant LHSV and total pressure are reported in Figure 5. It can be seen that, whatever the reaction temperature, the new high loading CoMo-D catalyst is more active than the conventional catalyst. It is useful to compare catalysts in term of iso-conversion temperature. The Apparent Activation Energy has been calculated by using the Arrhenius law and for the two catalysts. A value of about 25kcal.mol- 1 has been obtained. The gap in reaction temperature has been calculated for a 500 wt ppm sulfur content in the effluent (the future specification) with a sulfur content of the feed equal to 2.16wt% which correspond to an HDS conversion of 97.7%. We have determined a reaction temperature of 351 ~ for the conventional CoMo catalyst and 343~ for the new high loading CoMo catalyst. That means that to obtain the same conversion, the new CoMo catalyst can be operated at a temperature 8~ lower than that of the conventional catalyst. This gain in activity is due both to the use of the new proprietary support and the increase of the CoMo loading.

Hydrogen consumption The two main causes of hydrogen consumption during hydrodesulfurisation processes are the hydrogenation of aromatic hydrocarbons and the removal of sulphur as H2S. With cracked middle distillates, another main cause is the hydrogenation of olefins. Other causes, that can be neglected compared to those previously mentioned are hydrodenitrogenation (the amount of nitrogen is very small compared to the amount of sulphur, even in cracked products), and hydrocracking reactions very low under hydrodesulfurization conditions. The hydrogen consumption due to lIDS and due to olefin hydrogenation is about the same. Therefore, for a given hydrodesulfurization level and for a given feedstock, the difference of hydrogen consumption between two catalysts is only due to differences in aromatic hydrogenation. Figure 6 shows the comparison of aromatic hydrogenation between the two catalysts carried out on the same pilot unit, at the same operating conditions and with the same feedstocks. The precision of the percentage of the aromatic hydrocarbon measurement is about 5%. Therefore, the percentages of aromatic hydrocarbons in products treated on the two catalysts are the same. It means that the difference in hydrogen consumption between the two catalysts could only be due to the difference in hydrodesulfurization. Therefore, for a given hydrodesulfurization, the hydrogen consumption on one or the other of the two catalysts is the same. The results presented on Figure 6 obtained on SRGO have been confirmed on conversion gas-oils and on mixture of cracked gas-oils with straight-run ones (50/50vol).

Catalyst stability Tests on several feedstocks were performed with the CoMo-C and CoMo-D catalysts. These tests were carried out for over 1000 hours, i.e. about one month and an half. During testing, the reactor temperature (the main parameter that causes the catalyst ageing)was ramped from 325~ to 360~ with an average temperature higher than 340~ At the end of the tests, the operating conditions were returned to the initial values in order to estimate catalyst deactivation.

215

Figure 5. HDS conversion versus reaction temperature for CoMo-C and CoMo-D catalysts with a 50/50 SRGO/LCO

Figure 6. Comparison of aromatic hydrogenation versus total pressure for CoMo-C and CoMo-D catalysts

The deactivation is measured by the increase of temperature that is necessary to obtain the same HDS performance for the final point as for the initial point. It has been shown that the new high loading CoMo catalyst exhibits the same level of deactivation as the conventional CoMo. The maximum deactivation obtained with a conversion feedstock atter high temperature operation for more than 1,000 hours is about 1~ 4. CONCLUSION Use of the new proprietary support leads to a better toluene hydrogenation activity with a sulfided standard 3wt% CoO and 14wt% MoO 3 catalytic phase. The TEM experiments would show that the use of the new proprietary support leads to an increase of the intrinsic activity of each site. The use of a new support and a higher CoMo loading has been taken into account to design a new HDS catalyst. Procatalyse introduced this new high activity CoMo catalyst (I-H~I6C) in 1994. This catalyst exhibits, in HDS of a SRGO, a gain in iso-conversion temperature over of 8~ (iso-volumic activity). The activity and stability of this new catalyst have since been proven in commercial units processing straight-run or cracked gas-oil. The improvement in hydrodesulfurization activity was obtain without increasing aromatic hydrogenation or hydrogen consumption. In view of the tight hydrogen availability in most of refineries, this is another main advantage. ACKNOWLEDGEMENT We gratefully acknowledge Mrs A.K. Araya and E. Merlen for the TEM experiments and Mrs C. Guitton for the catalysts preparation and model molecule test.

216 REFERENCES

.

.

4. 5.

.

9. 10. 11. 12. 13.

D.C. McCulloch, M.D. Edgar, and J.T. Pistorius, "Higher Severity HDT Needed for Low-Sulfur Diesel Fuels" Oil & Gas Journal, April 13 (1987) 33-3 8 R.M. Nash, "Process Conditions and Catalysis for Low-Aromatics Diesel Studied" Oil & Gas Journal, May 29 (1989) 47-56 B. Delmon, Studies in Surface Science and Catalysis, 53 (1989) 1 V.H.J. DeBeer, G.A. Somorjai and R. Prins, Catal. Rev-Sci. Eng. 31 (1989) 1 H. Topsoe, B.S. Clausen, N.Y. Topsoe and P. Zeuthen, Studies in Surface Science and Catalysis, 53 (1989) 77 M.L. Vrinat, The kinetics of the hydrodesulfurization process, Appl. Catal., 6 (1983) 137 B.C. Gates, J.R. Katzer and G.C.A. Schult, Chemistry of Catalytic Processes, page 407, McGraw-Hill Book Co., New York, 1979 Le Page J.F. Applied Heterogeneous Catalysis. Editions Technip, Paris, 1987 J. Bachelier, M.J. Tilliette, J.C. Duchet and D. Comet, J. Catal., 76 (1982) 300 S. Kasztelan, H. Toulhoat, J. Grimblot and J.P. Bonnelle, Appl. Catal., 12 (1984) 127 R.R. Chianelli, A.F. Ruppert, S.K. Behal, B.H. Kear, A. Wold and R.J. Kershaw, J. Catal., 92 (1985) 56 J. Bachelier, J.C. Duchet and D. Comet, J. Catal., 87 (1984) 283 J. Bachelier, M.J. Tilliette, J.C. Duchet and D. Comet, J. Catal., 87 (1984) 292

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All fights reserved.

217

E N V I R O N M E N T A L L Y F R I E N D L Y DIESEL FUELS P R O D U C E D F R O M M I D D L E D I S T I L L A T E S G E N E R A T E D BY C O N V E R S I O N P R O C E S S E S R. Zamfirache and I. Blidisel

Research Institute for Petroleum Processing and Petrochemistry, B-dul Repubficii hr. 291,4, 2000 Ploiesti, Romania ABSTRACT Some diesel fuels specifications already in use or proposed for the near future are concerning with a 500 ppm sulphur content and a total aromatics content limited to 35 vol. % or even less. A new process aimed to meet the proposed diesel fuels specifications has been developed by Research Institute for Petroleum Processing and Petrochemistry Ploiesti. The pilot plant studies have been carried out to establish the best solution for revamping existing gas oil hydrotreating plants to reach both deep hydrodesulphurization (lIDS) and aromatics content reduction. It was found that a two stage hydrotreating process at medium pressure is the best approach for middle distillates higher in sulphur and aromatics content such as blends of straight run gas oil with thermally cracked gas oil. In the first stage the sulphur removal to very low levels is achieved combining the high HDS activity of a new type of promoted NiMo catalyst with variation of the process conditions. Aromatics hydrogenation is much more challenging than sulphur removal as requires additional hydrotreating capacity. A moderate pressure hydrogenation stage using a specific noble metal catalyst appears to be sufficient to reach a deep aromatics reduction if a feed desulphurization less than 150 ppm is performed. Tests of aromatics saturation at moderate pressure performed with a special high nickel containing catalyst have also been reported. 1. I N T R O D U C T I O N The Romanian refining industry will soon be facing a challenge to satisfy the demand of greater amount of high quality fuels in accordance with stringent international fuels specifications. New directive proposed by the European Community Commission [ECC] calls for cutting sulphur levels in all gas oils to 0.05 wt % by October 1996 to regulate sulphur dioxide emission from diesel engines. The assumption that a lower aromatic fuel reduces engine combustion temperature, thereby reducing nitrogen oxide formation, results in a limitation of total aromatics content to 3 5 vol. % in USA or even less such as a 5 vol. % limit in Sweden. Some of the new and proposed environmentally friendly diesel fuels specification are summarized in Table 1. Although gas oils obtained from the atmospheric distillate still remain the main source of diesel fuels, in order to cope with the increased consumption of naphtha and middle distillates almost all refineries in Romania use conversion processes such as fluid catalytic cracking on vacuum distillates and coking or visbreaking on residue. These processes generate middle distillates with higher olefins, diolefins, sulphur, nitrogen and aromatics content compared to gas oil obtained from an atmospheric distillation unit

218 Table 1. Present and proposed diesel fuel specifications

Country

Romania EEC USA California Sweden - Class 1 - Class 2 - Class 3 Japan

Max. Sulphur wt %

Max. ...............B.o.!!!.ng.r.a.nge ............... Aromatics IBP FBP Min. vol. % ~ ~ Cetane No

0.5 0.05 0.05 0.05 0.001 0.005 0.05 0.2 0.05

180

360

45

35 10

170 170

350 350

40 -

5 20 -

180 180 180

300 320 330

50 50 -

Valid from

1980 1996 1993 1993 1991 1991 1991 1992 1997

Since 1968 ICERP Ploiesti has been developed hydrofining technologies and catalyst systems for treating various petroleum cuts including blends of distillates with cracked feedstocks. The commercial hydrofining units in domestic refineries as well as those licensed in different countries are characterized by high service factors. In order to establish the best solution for revamping the existing gas oil hydrotreating units a new process to produce diesel fuels at the sulphur and aromatics contents specification by hydrotreating blended feed of cracked gas oil (CKGO) with straight run gas oil (SRGO) has been developed. In this paper, consideration will be given to the type of catalysts required, to the process conditions and to the possibilities of process implementation to revamp the existing gas oil hydrotreating units. 2. D E E P H Y D R O D E S U P H U R I Z A T I O N

OF GAS OIL BLENDS

2.1. C o n v e n t i o n a l a p p r o a c h

There are some different approaches to extend gas oil hydrotreating technology so as to increase sulphur compounds conversion to reach the new specification of 500 ppm. At this lower level the most difficult sulphur-containing molecules in gas oil have to be removed. An extensive study [ 1] concerning the relative reaction rates of various kinds of sulphur compounds on a CoMo catalyst shows that at 375~ and a hydrogen partial pressure of 34 bar the hardest to react sulphur species are dibenzothiophenes which are controlling the HDS rate. The rate of removal sulphur-containing molecules increases as follows:

219

R < S

S

S

--R< s

Increasing in reaction temperature is the first way to take into consideration. A 30-50~ over the operating temperature is necessary to reach the 500 ppm sulphur but such an increase could not be acceptable because of the colour degradation and of the shorter catalyst cycle length [2,3]. A hydrogen partial pressure increase, which is favorable for sulphur compounds conversion is limited due to mechanical constraints of maximum total pressure of existing hydrotreating reactor. However, a substantial increase of hydrogen partial pressure could be obtained by purification of make up hydrogen or recycled hydrogen. Another approach to reach 500 ppm sulphur in product is to adjust the plant capacity working at lower liquid hourly space velocity [LHSV]. If such a reduction is not acceptable, the volume of catalyst will have to be increased by addition a new HDS reactor. 2.2. Performance of the new type of HDS catalysts.

It is well-known that a new generation of hydrotreating catalysts prepared with a silica promoted alumina support has been developed and are in use in a number of commercial hydrotreating units. Improved and more flexible operation should be possible especially in thermally cracked feedstocks with these catalysts having a higher HDS activity and resistance to carbon deposition than conventional CoMo or NiMo catalysts. Model 23 R-16 (NiMo) is a new HDS promoted catalyst developed by ICERP to be used in desulphurization units in order to reach the new diesel fuel specifications. The catalyst is based on a new type of alumina obtained by an original preparing method which offers a correct interaction degree between metal and its support. The acidic property and the pore size were improved by the addition ofpromoteurs such as silica and phosphorus (P2Os).Some of the properties of new 23 R-16 (NiMo) promoted catalyst in comparison with the standard hydrofining catalyst are listed in Table 2. Table 2. Catalyst properties. Catalyst type NiO, wt % MOO3, wt %

Standard catalyst 5.21

23 R-16 (NiMo) 5.34

13.12

15.87

SiO2, wt %

-

3.78

P205, wt %

-

2.13

Na20, wt % Sp. surface area, m2/g Total pore volume, cm3/g Radius average, A Crushing strength, kgf, g Attrition strength, %

0.047 241 0.50 41 5.5 97.2

0.03 262 0.62 48 6.5 98.5

220 A comparison between the new 23 R-16 hydrotreating catalyst and standard catalyst is presented in figure 1. The results exhibit a higher HDS activity at the same process conditions.

,I, ~ 1000

-"-

- A- *-

- ~~~..._

LEGEND Standard 23 R-16

: -

-

500 "~ r.~

L Feedstock: SRGO+CKGO (60/40 vol%) r Hydrogen pressure: 50 bar .~LHSV: I 1,5h "l

100 340

350

360 Reaction temperature, ~

370

380

Figure 1. HDS activity of 23 R-16 vs. Standard catalyst

2.3. Pilot plant data The proposed objective of liDS tests was to reduce sulphur content in blended feed of 20 to 40% of CKGO with SRGO in order to obtain diesel fuel at the new specification of 0.05 wt % level. The properties of typical SRGO + CKGO blends are presented in Table 3 and The operating conditions and the test results are summarized in Table 4. Table 3. Properties of typical SRGO + CKGO blends. Feed

SRGO

SRGO+CKGO

SRGO+CKGO

80/20

60/40

CKGO

Density (15~

0.842

0.852

0.870

0.8841

Sulphur, wt %

1.3

1.28

1.24

1.16

Nitrogen, ppm

170

185

201

241

Aromatics, vol. %

25.8

31.7

38.5

55

Cetane index

53

50

47

39

IBP

186

189

213

215

Distillation (~ 10%

232

228

238

235

50%

298

289

292

284

90%

346

340

358

354

FBP

360

358

398

396

221 Table 4. Deep HDS of gas oil blends. SRGO + CKGO

Feed

SRGO + CKGO

80/20

60/40

Operating conditions: - Catalyst -

23 R- 16 (NiMo)

HE part. pressure, bar

- Temperature, ~ - LHSV, h1

23 R- 16 (NiMo)

60

60

360

360

1.0

- HE Consumption, m3/m3

1.0

42

60

360

480

42

48

29.5

36.5

51

49

Product quality: - Sulphur, ppm -

Nitrogen, ppm

- Aromatics, vol. % -

Cetane index

1500

1500

1000 -

1000-

500

500

J I

320

I

I

340

I

I

360

I

I

380

Reaction temperature, ~

Figure 2. Sulphur in product vs. reaction temp.

0.0

f

J I

0.5

I

I

I

1.0 LHSV, h-1

Figure 3. Sulphur in product vs. LHSV

The results in Table 4 indicate that a temperature of 360~ is sufficient to reach the proposal objective of 500 ppm sulphur. Figures 2 and 3 show corelations between operating conditions such as reaction temperature as well as LHSV and sulphur content in product. As can be seen, a low sulphur level is relatively accessible by using a good hydrotreating promoted catalyst such as 23 R-16 (NiMo) and by varying the process conditions. It should be noted that the maximum limit of 60 bar imposed by mechanical constraints of existing hydrotreating reactor was chosen to reach the minimum possible pressured needed in second stage aromatics hydrogenation reactor. By using such a moderate pressure the sulphur content

222 can be reduced in the first stage to a level that does not affect too much the performance of aromatics hydrogenation in case of a sensitive to poisoning noble metal based catalyst will be used in the second stage. 3. HYDROGENATION OF GAS OIL BLENDS

3.1. Conventional approach The reduction of the aromatics level in SRGO to reach 10 vol. % can be attained with a moderate hydrogen partial pressure of 60 bar and a NiMo catalyst. Hydrogenation of gas oil blends containing cracked feedstocks required a higher severity involving higher hydrogen partial pressure and lower LHSV. As well known, the hardest compounds to hydrogenation are the monoaromatics which are controlling the dearomatization rate. Increasing in reaction temperature is limited due to the thermodynamic limitation of aromatics hydrogenation [4].In order to attain a 10 % level at aromatics content in blended feeds of 20 to 40% CKGO with SRGO by using a moderate pressure of 60 bar the method of choice is to get the proper hydrogenation catalyst.

3.2. Hydrogenation catalysts The effect of catalysts based on noble metals on the aromatics hydrogenation have been well documented in literature. These catalysts are very sensitive to poisoning by very small amounts of sulphur compounds in the feedstocks [6 ]. As the sulphur level tolerable by such catalysts ranges from 1.5 ppm to 600 ppm, a deep HDS has to be performed. Model 1-6 is a new hydrodearomatization (HDA) catalyst developed by ICERP to be used in aromatics hydrogenation of gas oil blends. The catalyst has been obtained by a highly improved NiO dispersion on the promoted alumina support having a bimodale pore distribution with a total pore volume of minimum 45 cm3/g. 1-6 has a good HDA activity under rather moderate hydrotreating conditions i.e. 60 bar total pressure and a remarkable sulphur resistance.

3.3. Pilot plant data The proposed objective was to reach 10 vol. % aromatics level in product by hydrogenation of the desulphurised blended feed of 20 to 40 % CKGO with SRGO. In order to compare the HDA activity of I-6 catalyst and commercial HDA noble metal (Pt) catalyst the pilot tests have been perform on a deep desulphurised feed containing 150 ppm sulphur for both catalysts. The operating conditions and test results are summarized in Table 5. As can be seen, a 60-70% degree of aromatics saturation in the desulphurised blended feed containing up to 40 vol. % CKGO has been obtained at 350~ for both catalysts, that is sufficient to reach the proposed objective of 10 % aromatics content in product. The tests also confirm the nitrogen removal and the Cetane index improvement. The effect of reaction temperature on the aromatics saturation is shown in figure 4.

4. IMPLEMENTATION OF HDS / HDA PROCESS Depending on the proposed objective - deep HDS or combination deep HDS / aromatics saturation - the developed technology can be applied as a single or two stage process. The revamping offers the refinery a much more lower cost route to meet the new gas oil

223 Table 5. Hydrogenation of gas oil blends. SRGO+CKGO

SRGO+CKGO

Feed

60/40

80/20

Noble metal (Pt)

Catalysts Operating conditions - H2 pressure, bar - Temperature, ~ - LHSV, h

60 380 1.0

HDA, % Product quality - Sulphur, ppm - Nitrogen, ppm - Cetane index

Noble metal (Pt)

I-6 60 350 1.0

1-6

60 380 1.0

60 350 1.0

66.0

64.8

72.7

71.5

20 <2 57

17 <2 55

21 <2 55

19 <2 54

100 Legend

90

-A-.-

1-6 Noble metal

80 70 60 50 40

Feed: SR ulphur: 150 ppm Hydrogen pressure: 60 bar LHSV: 1 h"l

30 20 280

300

320

340

360

380

400

Reaction temperature, ~ Figure 4. HDA activity of 1-6 and noble metal catalysts

specifications than a grassroots unit. Thus a single stage deep HDS process can be implemented for revamping the existing gas oil hydrofining unit to produce 500 ppm sulphur diesel fuels by hydrotreating blended feed up to 40 % of CKGO + SRGO.

224 The results reported in this work indicate that a 500 ppm sulphur level in product is accessible by using combination of 23 R-16 (NiMo) hydrotreating catalyst developed by ICERP and suitable operating conditions. To reach the proposed aromatics content specification of 10 vol. % a new reactor and separation system should be added in series to existing hydrotreating unit. A proper hydrogenation catalyst such as sulphur resistant 1-6 developed by ICERP can be used to attain a 10 vol. % aromatics content in desulphurised blended feed up to 40 % CKGO with SRGO. The two stage process has the advantage to combine the moderate pressure in deep HDS stage with the same moderate pressure in aromatics saturation stage since the second stage can be operated essentially at low sulphur content. It should be mentioned that the moderate to severe proposed reduction of aromatics content in diesel fuels will have a major impact on the hydrogen availability on the refinery. 5. CONCLUSION The revamping of existing hydrotreating units to reach the proposed diesel fuels specifications of less than 0.05 wt % sulphur is achievable at moderate pressure combining the high HDS activity of the new generation of hydrotreating catalysts such as 23-R-16 NiMo promoted catalyst developed by ICERP and suitable operating conditions especially lower LHSV. To remove the critical production capacity constraints concerning with the lower LHSV on cracked feed hydrotreatement, the co-processing of blended feed up to 40 % CKGO with SRGO has been considered. Aromatics reduction up to 10 vol. % is much more challenging than sulphur removal. If and when such a significant reduction is required a new reactor and separation system has to be added and integrated with the deep HDS unit. A proper sulphur resistant HDA catalyst such as Model 1-6 developed by ICERP can be used to attain a 10 % level of aromatics content in desulphurised blended feed up to 40 % CKGO with SRGO. A two stage process approach at moderate pressure has the advantage to combine the deep HDS and aromatics saturation and, therefore, to meet the process objectives with minimum investment cost. REFERENCES

1. A. Nastasi, Inginefia Prelucr~fii Hidrocarbufilor, vol. 2, p. 297, Ed. Tehnic~, Bucuresti 1992. 2. S Kasztelan., N. Marchal and S. Kresmann, AICHE Spring Meeting, Houston, Texas, 1993 3. A Suchanek., Oil and Gas J., May 7, 1900, p. 109 4. I. Barbier, Advances in Catalysis, 37 (1990) 279. 5. N Milam, AICHE Spring Meeting, April 1991 6. A.Cavanaugh and W. M Gregory, Hydrocarbon Technology International, 2 (1994) 109.

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All rights reserved.

225

FACTORS INFLUENCING THE PERFORMANCE OF NAPHTHA HYDRODESULFURIZATION CATALYSTS

Jamal A. Anabtawi, Syed A. Aii, M. Abdui Bari Siddiqui and S. M. Javaid Zaidi

The Research Institute, King Fahd University of Petroleum andMinerals, Dhahran 31261, Saudi Arabia Abstract The performance of the two CoMo catalysts was evaluated for hydrodesulfurization of two Saudi naphthas in a bench-scale unit under Industrial conditions. Various types of sulfur compounds in the naphtha feed stocks and products were determined by chemical analysis and gas chromatography. The results show that the product sulfur decreased with increasing temperature down to a minimum, followed by an increase at higher temperatures. This increase was attributed to the occurrence of H2Salkene recombination reaction leading to the formation of mercaptans. The effect of increasing the space velocity was to slightly increase product sulfur, while the hydrogen gas rate had an insignificant effect. The product sulfur for the two catalysts was below 1 ppm in the temperature range 280-350~ However, one catalyst was selected based on its superior HDS performance attributed to higher concentrations of molybdenum, cobalt and phosphorus oxides. 1. INTRODUCTION Bimetallic reforming catalysts (Re/Pt >1) are more susceptible to sulfur poisoning. Sulfur can cause shorter cycle length, rapid increase in reactor temperature to maintain octane level and faster catalyst deactivation. Selection of a hydrodesulfurization (HDS)catalyst is dependent on several factors, including feedstock properties, product specifications, catalyst characteristics, and process economics. Pilot plant tests, carried out for comparative performance of competitive catalysts, are complicated due to experimental unit design and procedures, high initial activity, and slow catalyst deactivation. The main challenge is the trace sulfur analysis (below 1 ppm) which is influenced by the presence of dissolved H2S in the product. This paper describes the challenges faced in the evaluation of liDS catalysts under refinery commercial conditions, and illustrates the factors influencing performance data. 2. EXPERIMENTAL PROCEDURES

2.1 Catalysts, Feedstocks and Products Analysis The two cobalt-molybdenum catalysts were characterized for their physical properties: surface area (ASTM D-3663), pore volume, and average pore diameter (ASTMD-4222), mechanical properties: crushing strength (ASTM D-4179), and attrition loss (ASTM D-4058), chemical composition, as well as carbon and sulfur content by coulometry.

226 The feedstocks (straight-run naphtha (SRN) and a blend of SRN and hydrocracked naphtha) and hydrotreated products were analysed by ASTM methods for density, carbon, hydrogen, hydrocarbon and boiling point distribution. Total sulfur was determined by ASTM D-4045 method, mercaptan sulfur by the potentiometric method (ASTM D-3227 and UOP212), disulfides by the UOP-202 method, polysulfides by polarography [1], and elemental sulfur by the UOP-286 method. The Perkin-Elmer gas chromatograph (Model 8700), equipped with a flame photometric detector (GC/FPD) and a DB-1 fused silica capillary column (30 m x 0.53 mm), was used for identification of individual sulfur compounds [2-6]. The sensitivity of the GC/FPD technique was maximized by optimizing the gas flow rates and temperature programming as presented elsewhere [ 1].

2.2 Catalysts Activity Tests The HDS activity tests were conducted in a fully automated bench-scale unit designed for unattended operation. The catalysts were calcined ex-situ at 425~ for two hours and then diluted with 2-3 mm size alpha-alumina to give 46 cm long catalyst bed in a 2.7 cm internal diameter trickle-bed reactor. They were presulfided in-situ using a 1 wt. percent carbon disulfide diluted in SRN (5 h1 and hydrogen gas rate of 150 Nm3/m3 at 200~ for three hours, at 250~ for one hour, followed by 16 hours at 300~ Pure hydrogen gas was used in oncethrough mode and the pressure was maintained at 28.2 bar during all tests. The temperature profile inside the catalyst bed was monitored by three thermocouples and controlled by five-zone furnace. The performance activity tests were carried out for 48 hours after the steady state had been reached. The liquid product stream from the reactor was condensed at 10~ in a gas-liquid separator, and was immediately stripped of dissolved hydrogen sulfide by extraction with acidified cadmium chloride solution prior to analysis. The gaseous products were analyzed periodically by refinery gas analyzer. The experiments were designed to investigate the effects of temperature (220-350~ space velocity (10 and 13 hl), hydrogen gas rate (67 and 80 Nm3/m3) with two feedstock types on the performance of the catalysts. 3. RESULTS AND DISCUSSION

3.1 Catalyst Characterization The data shown in Table 1 indicate that catalyst B has higher concentrations of molybdenum and cobalt oxides compared to catalyst A, and contains about 1.8% phosphorous oxide, which acts as promoter for sulfur and nitrogen removal [7]. Catalyst B showed higher crushing strength, lower density, more abrasion resistance, and larger pore volume and surface area. Although these results indicate the superiority of catalyst B, performance with actual feeds and under commercial operating conditions is usually the criteria for catalyst selection.

3.2 Feedstocks Analyses Analysis of the feedstocks, shown in Table 2, indicates that the density, refractive index and carbon content of the blend naphtha are higher than those of SRN. This was due to higher aromatics and naphthenes, lower alkanes and higher final boiling point. To study the effect of catalyst type on sulfur compounds selectivity, the feedstock and products were characterized

227

Table 1 Catalyst Characterization Results Catalyst A Property

fresh

Physical properties Form Extrudate, Average length (mm) 7.2 Bulk density (kg/m3) 740 BET surface area (m2/g) 193 Pore volume (ml/g 0.45 Mean pore radius 47.3 Attrition loss (%) 1.65 Crushing strength (N/mm) 10.6 Chemical analysis (wt%, dry base) Molybdenum oxide 15.00 Cobalt oxide 4.00 Aluminium oxide 74.70 Sodium oxide 0.08 Silicon oxide 0.15 Phosphorous oxide 0.07 Carbon 0 Sulfur 0 Loss on ignition at 1000~ (wt% wet base) 4.6

Catalyst B

spent

fresh

spent

9.7

Extrudate, 4.9 700 233 0.54 46.3 0.70 24.3

1/ 16" 4.2 -161 O.35 43.5 -22.0

3.67 5.97

19.50 4.60 68.00 0.13 0.46 1.83 0 0

3.07 7.25

1/16" 6.4 148 0.37 48.6

5.30

Table 2 Analytical Results of Naphtha Feedstocks Property

Physical Properties Density (kg/m3) Refractive index

Chemical Analysis Alkanes (wt.%) Alkenes (wt.%) Naphthenes (wt.%) Aromatics (wt.%) Carbon (wt.%) Hydrogen (wt.%) Total sulfur (ppm) Mercaptans (ppm) Disulfides (ppm) Sulfides + thiophenes (ppm) Elemental sulfur and H2S Boiling Range (~ IBP 50% FBP

SRN

Blend naphtha

740.5 1.4194

749.5 1.4257

69.3 0.1 18.8 12.0 85.8 15.6 897 500 10 387 trace

58.6 0.1 28.3 13.0 87.7 14.7 534 297 5 232 trace

66 128 187

65 128 194

228 by chemical analysis for mercaptans, sulfides, disulfides, thiophenes, polysulfides and hydrogen sulfide. The results presented in Table 2, indicate that the SRN contains mercaptans (55.7%), thiophenes and sulfides (43.1%), disulfides (1.1%) and traces of polysulfides, hydrogen sulfide and elemental sulfur. Further analysis by GC/FPD showed the presence of fifty-two sulfur compounds, seventeen of which were identified by matching the retention times. The DB-1 column, which separates sulfur compounds and elutes them in an increasing order of their boiling points, was used to estimate the boiling points of the unidentified peaks. Twenty-four sulfur compounds were recognized by matching the boiling point. Table 3 shows the forty-two sulfur compounds identified include ten sulfides, thirteen disulfides, thirteen mercaptans, and sixteen thiophenes. Analysis of commercial product naphtha identified four compounds namely methyl mercaptan, isopropyl mercaptan, n-butyl mercaptan and 2,3,4,5-tetramethyl thiophene. The presence of this thiophene was expected because alkyl thiophenes, are difficult to desulfurize. However the presence of mercaptans indicates the occurrence of recombination reactions between alkenes and H2S [8] as follows: k2 k3 Mercaptans " > Alkenes > Alkanes (1) + H2 kl + H2S Pure component studies indicate the rate of mercaptan formation is sufficiently rapid at hydrotreating conditions compared to the saturation step which lead to alkane [8]. The exothermic reversible reaction, which shifts to the left at higher hydrogen sulfide partial pressure, is also dependent on temperature, feedstock type, total sulfur, partial pressure of hydrogen and alkenes, space velocity and catalyst type. Furthermore the size of the reactor affect the balance between the kinetic sulfur removal and alkene saturation [9].

3.3 Comparison of Bench-Scale Unit with Commercial Hydrotreater. Performance of the bench-scale unit was compared to a commercial hydrotreater using blend naphtha and catalyst A under similar operating conditions (320~ 10 h~, and 80 Nm3/m3). The products from the refinery and bench-scale units showed that the total sulfur was 0.3 and 0.78 ppm, respectively. The higher sulfur content in the bench-scale unit could be attributed to lower superficial mass velocity, deviation from plug-flow, and poor catalyst utilization. Results also show that the mercaptan content was about 0.3 ppm in both products suggesting the occurrence of H2S recombination reaction takes place in the reactor outlet piping and heat exchangers systems as the fluids cool down. Hydrogen sulfide has a strong retarding effect on HDS, due to inhibition caused by adsorption on catalytic sites in competition with the sulfur compounds. This lIDS inhibiting effect decreases with rising temperature and is more pronounced in benchscale units. On the other hand, definite partial pressure of H2S is required to maintain the activity of liDS catalysts [10-11]. 3.4 Effect of Temperature The HDS performance data of the two catalysts are presented in the Table 4. The total product sulfur is shown as mercaptan and other sulfur types for catalyst A and B with SRN and blend naphtha. Figure 1 shows that the blend naphtha is easier to desulfurize and catalyst B is more active. The total sulfur for catalyst A decreased from 72 ppm at 220~ to a minimum of 0.69 ppm at 300~ However, upon further increase in the temperature up to 350~ the total sulfur increased as a result of mercaptan formation. Figure 2 shows the total and mercaptan sulfur distribution as a function of temperature for catalyst A with blend naphtha.

229 Table 3 Sulfur Compounds in Naphtha Feedstocks Identified by GC/FPD Method Serial no.

Compound name

Retention time(min.)

Matched peak #

Boiling point (~

1.48 3.00 4.07 5.04 5.42 7.76 9.11 10.29 13.18 13.64 16.88 20.30 20.75 22.84 25.25 34.18 38.18

1 3 4 5 6 7 8 10 14 15 19 24 25 28 31 44 52

36 56 64.2 67.5 84 92 100 112.4 115.4 126 132.5 135.5 142 151 176 188

1.51 12.57 15.60 16.15 18.51 19.75 22.11 23.38 24.37 25.25 25.79 25.79 27.01 28.05 28.85 31 59 32.65 33 57 34 85 35 25 35.65 35.89 36.48 37.18 38.13

2 13 17 18 21 23 27 29 30 31 32 33 35 36 37 41 42 43 45 46 47 48 49 50 51

6 109.7 120.7 122.5 130 132 139.5 144 149 152 153 154 156 159 161 169 172.7 175 179 179.5 181 181 183 185 186

Compounds identified by matching retention times .

2. 3. 4. 5. 6. 7. 8. 9. 10. 11 12. 13 14. 15 16 17

Hydrogen s Ethyl mercaptan Isopropyl mercaptan tert-Butyl mercaptan n-Propyl mercaptan see-Butyl mercaptan Di-ethyl sulfide n-Butyl mercaptan 2 Methyl thiophene 3 Methyl thiophene n-Amyl mercaptan 2 Ethyl thiophene 2,5 Di-methyl thiophene Di-n-propyl sulfide n-Hexyl mercaptan n-Heptyl mercaptan Di-n-butyl sulfide

Compounds identified by matching boifingpoints .

2. 3 4. 5. 6. 7. 8. 9. 10. 11. 12. 13. 14. 15. 16. 17. 18. 19. 20. 21. 22. 23. 24. 25

Methyl mercaptan Methyl disulfide Di-isopropyl sulfide n-Butyl methyl sulfide Ethyl methyl disulfide Isopropyl propyl sulfide Methyl 2 methyl butyl sulfide 3,4 Di-methyl thiophene tert-Butyl sulfide 2 Isopropyl thiophene Ethyl disulfide 2 Hydroxy ethyl mercaptan 3 Isopropyl thiophene Cyclohexyl mercaptan 2 Ethyl 3 methyl thiophene 3 tert-butyl thiophene 2,3,4 Tri-methyl thiophene 1,3 Di-thiacyclopentane 2 benzothiozole thiol 2 Methyl 5 propyl thiophene 2 n-Butyl thiophene 2,5 Di-ethyl thiophene 3 n-Butyl thiophene 3,4 Di-ethyl thiophene 2~3~4~5~Tetra-methyl thiophene

230

Table 4 Comparison of Catalyst Performance Data for SRN and Blend Naphtha Operating conditions Temperature ~

Space velocity (h- 1)

Gas rate

Catalyst A Total sulfur

(Nm3/m3) ppm

220 250 280 300 320 350

10 10 10 10 10 10

67 67 67 67 67 67

72 95 1 06 0.69 076 089

280 300 320 350 250

10 10 10 10 13

80 80 80 80 67

0.91 0.76 0.78 0.80

250 300 320 350

13 13 13 13

67 67 67 67

300 320 350

13 13 13

220 250 280 300 320 350

Catalyst B

Mercaptan sulfur ppm %

Blend Naphtha 4 5.6 0.67 0.27 0.25 0.43 0.51

7.1 25.5 36.2 56.6 57.3

0.20 0.25 0.47 0.53 . . . .

22.0 32.9 60.3 60.3

1.33 0.94 0.76 0.72

0.1 0.24 0.25 0.53

7.5 25.5 32.9 73.6

80 80 80

1.01 0.69 0.75

19.8 29.0 92.0

10 10 10 10 10 10

67 67 67 67 67 67

184 2.6 0.79 0.52 O.37 0.87

0.20 0.20 0.69 SRN 8.7 0.80 0.28 0.20 0.33 0.73

4.7 3.1 35.4 38.5 89.2 83.9

280 300 320 350

10 10 10 10

80 80 80 80

0.81 0.32 0.23 0.92

0.33 0.13 0.13 0.76

40.7 40.6 56.5 82.6

250 280 300 320 350

13 13 13 13 13

67 67 67 67 67

4.1

2.24 0.53 0.43 0.26 0.65

5.5 36.1 82.7 92.9 83.3

1.47 0.52 0.28 0.78

Total sulfur ppm

Mercaptan sulfur ppm %

5.8 4.7 0.52 0.32 0.54 0.75 0.34 . . 0.28 . . 7.3

. .

7.7 1.00 0.16 0.21 0.33 0.60

13.3 21.3 30.8 65.6 61.1 80.0

0.25 . . 0.28 . . . 1.36

73.5

.

100 18.6

0.40 0.25 0.32 0.62

0.36 0.20 0.30 0.60

90.0 80.0 93.8 96.8

16.1 0.36 0.38 0.42 0.69

1.1 0.15 0.27 0.27 0.64

6.8 41.7 71.0 64.3 92.8

0.38

0.20

52.6

0.40

0.26

65.0

24.5 0.62 0.37 0.66 0.79

2.13 0.27 0.21 0.43 0.64

8.7 43.5 56.8 65.2 81.0

231

180

e Catalyst A-Blend 9 Catalyst B -Blend

16o

9 Catalyst A - S R N 9 Catalyst B - S R N

140 1=. 120

r,~

8o

o [,.

60 40

9

200

!

9

|

220

9

240

|

9

260

T

9

280

9

300

9

320

|

~

340

|

9

360

380

Temperature, Deg. C

Figure 1. Effect of temperature on catalysts performance.

1.4.

e

Total Sulfur

o

M~v, aptan

1.2,

E gg gg

1.0,

%....r

t: 0.8" r,/3 0.6.

:3 "m @ t,.

"~

0.4,

0.2

0,0

~

250

i

275

9

i

300

9

i

9

325

i

350

9

375

Temperature, Deg. C

Figure 2. Effect of temperature on sulfur distribution for blend naphtha with catalyst Ao The mercaptans in the feedstock, which was about 56 percent of the total sulfur, was reduced to about six percent at 220~ and then increased to about 36 percent at 300~ and to about 57 percent at 350~ as shown in Table 4. These data confirm the occurrence of mercaptan forming reaction along with HDS of other types of sulfur. Similar results were observed by Sekhar and Rahimi while hydrotreating naphtha derived from coal liquids and heavy oil [12]. Our data show that the maximum temperature of operation for catalyst A is 300~ above which recombination reactions become dominant.

232

n

L H S V = 10

9

L H S V = 13

40

B t2

30

d 9--

20

O

0

9

240

9 2;0

"2;o

" 300

9

320

|

340

360

Temperature, Deg. C Figure 3. Effect of space velocity on product sulfur for SRN with catalyst A.

Table 4 also shows that although the products obtained with catalyst B contain lower total sulfur than with catalyst A, the mercaptan sulfur content was significantly higher. The mercaptans were about 13 percent of the total sulfur at 220~ and increased with temperature to 80 percent at 350~ It can be inferred that catalyst B has higher HDS activity, but it also favors the formation of mercaptans by H2S-alkene recombination reactions. However, a maximum of 300~ was also observed for optimum catalyst performance. The data indicate that the optimum operating temperature is dependent on the feed composition and not the catalyst used (both catalysts were of Co-Mo type). The GC/FPD results indicate that the product obtained with catalyst B at 250~ contained twelve sulfur compounds: three mercaptans, two sulfides, and seven thiophenes. At temperatures above 280~ thiophenes were removed almost completely, but mercaptans were still present suggesting the occurrence of recombination reactions. Therefore the hydrotreater has to be operated at the lowest temperature possible to minimize the alkene production leading to mercaptans formation [9]. 3.5 Effect of Space Velocity and Gas Rate The effect of space velocity on HDS was studied at 10 and 13 h~. For SRN and catalyst A at 250~ the total sulfur increased from 26 to 41 ppm as the space velocity increased from 10 to 13 h ~ as shown in Figure 3. Meanwhile, the mercaptans increased slightly from 0.8 to 2.2 ppm as presented in Table 4. The effect of hydrogen gas rate on HDS was also investigated at 67 and 80 Nm3/m3. The data, presented in Table 4, indicate that the effect of gas rate on the desulfurization of blend naphtha was insignificant within the range of operating conditions used. It was found that the effect of hydrogen gas rate is more pronounced at lower space velocity and that higher hydrogen gas rate increases desulfurization and suppresses H2S-alkene recombination reactions.

233

1.4

9 SRN 9 Blend

1.2

~

1.0

~

12

0.8

r,~

0.6

0 ~

0.4

0.2 0.0 25O

9

|

9

9

9

|

9

Temperature, Deg. C

Figure 4. Performance of catalyst A with SRN and blend naphtha.

3.6 Effect of Feedstock

The effectiveness of catalysts A and B to desulfurize SRN and blend naphtha was investigated and the results are shown in Table 4. Figure 4, which shows the performance of catalyst A, illustrates that it is easier to desulfurize SRN than blend naphtha. The results also confirmed higher HDS performance with blend naphtha than SRN with both catalysts. This could be due to the refractive material in the hydrocracked fraction of the blend naphtha. With blend naphtha and catalyst A the minimum total sulfur of 0.69 ppm was obtained at 320~ while with SRN the minimum was 0.37 ppm at 300~ Above these temperatures, the occurrence of H2S-alkene recombination reactions increased the total sulfur. Nickelmolybdenum catalysts are known to reduce recombination reactions by hydrogenating alkenes. Higher temperatures and very active catalysts can cause cracking at the reactor outlet allowing alkenes production[ 13]. 3.7 Comparison of Catalyst Performance Data

Performance of the two catalysts was compared and the results are presented in Table 4 and in Figure 1. These data indicate a superiority in HDS activity of catalyst B (up to 3.3 percent better than catalyst A for blend naphtha and 1.1 percent for SRN under refinery operating conditions). It should be noted that catalyst B has higher concentrations of molybdenum and cobalt oxides compared to catalyst A, and contains about 1.8% phosphorous oxide. However, due to the operating conditions used, significant difference in activity could not be observed. Stanulonis and Pedersen [7], who investigated the effect of promoters on the importance of hydrotreating catalysts, reported that presence of phosphorus may produce a variation in acidity. This variation, similar to that caused by Co in Mo/alumina, would improve the desulfurization and denitrogenation activity of the catalyst. Catalyst B as a result of its

234 acidity exhibited higher hydrocracking and olefin production thus promoting recombination reaction at higher temperatures. 4. CONCLUSIONS This study investigated the factors affecting the performance of naphtha HDS catalysts in bench-scale reactor. Although naphtha HDS is routinely carried out at the refineries, the results confirmed the occurrence of H2S-alkene recombination reactions by conducting detailed analysis of sulfur compounds. Product sulfur decreased to a minimum, after which increasing temperature enhanced mercaptan formation. This minimum was found to be a function of sulfur concentration and type of the feed. On the basis of characterization and performance data, catalyst B was found to be better due to higher oxide loading and promoting effect of phosphorus. However, the undesirable mercaptan formation could not be eliminated. Therefore, hydrogenation of alkenes should be considered while designing newer hydrotreating catalysts. 5. ACKNOWLEDGMENTS This work is a part ofKFUPM/R/Project No. 21101 sponsored by a refinery in Saudi Arabia. The authors wish to acknowledge the support of the Research Institute of King Fahd University of Petroleum and Minerals. 6. REFERENCES 1. 2. 3. 4. 5. 6. 7. 8. 9. 10. 11. 12. 13.

J.A. Anabtawi, K. Alam, M.A. Ali, S.A. Ali and M.A. Siddiqui, Presented at the First Intl. Conf. on Chem. and its Applications, Dec. 7-9, (1993), Doha, Qatar. S.O. Farwell and C.J. Barinaga, J. Chromatographic Sci. 24 (1986) 483. K.J. Hyver and D. Diubaldo, HP-GC Application Brief, (1986). J.F. McGaughey and S.K. Gangwal, Anal. Chem., 52 (1980) 2079. F.M. Ali, H. Perzanowski and S. Koreisk, Fuel Sci. Tech. Intl., 9 (1991) 397. R.S. Hutte, N.G. Johansen and M.F. Legier, J. High Resolution Chrom., 13 (1990) 421. J.J. Stanulonis and L.A. Pederson, Proc. Symposium on Novel Methods of Metal and Heteroatom Removal, Houston, March 23-28, (1980) 255. D.P. Satchell and B.L.Crynes, Oil and Gas J., Dec. 1, (1975) 123. NPRA Questions and Answers, Oil and Gas J., 82, 14 (1984). M.L. Vrinat, Applied Catalysis, 6 (1983) 137. S.C. Schuman and H. Shalit, Catalysis Reviews, 4 (1970) 245. M.V.C. Sekhar and P.M. Rahimi, Adv. Hydrotreating Catalysts, (1989) 251. NPRA Questions and Answers, Oil and Gas J., 87, 9 (1989).

Catalysts in PetroleumRefining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All rights reserved.

235

HYDROCRACKING OF PARAFFIINIC HYDROCARBONS OVER HYBRID CATALYSTS CONTAINING H-ZSM-5 ZEOLITE AND SUPPORTED HYDROGENATION CATALYST I. Nakamura and K. Fujimoto

Department of Applied Chemistry, Faculty of Engineering, The University of Tokyo, 7-3-1, Hongo, Bunkyo-ku, Tokyo 113 Japan ABSTRACT A hybrid catalyst, which was prepared by physical mixing of a H-ZSM-5 and Pd/SiO2, showed an excellent activity for the hydrocracking of n-paraffins. In the case of n-dodeeane hydrocracking, the hybrid catalyst gave high selectivity for hydrocarbon fragments in the middle rang (C8-C9) and for the isomer of dodecane in a hydrogen atmosphere, in spite of its high conversion level. In the n-heptane cracking, the hybrid catalyst gave only isomerized heptane and propane and equimolar amount of ibutane 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 polymerization and cracking. The simple products for tile H2-hybrid system should be formed through no other reaction path than the cracking reaction on H-ZSM-5. I. INTRODUCTION Hydrocracking of petroleum heavy hydrocarbons have been practised extensively commercially 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 alkane is as follows: (1) the dehydrogenation of alkane to alkene on the supported metal; (2) proton addition to the alkene to form carbenium ion on the acidic component; (3) 13-scission ofsugiid the carbenium ion to form smaller carbenium ion and alkene on the acid component; (4) hydrogenation of the cracked alkene to alkane on the metal [1]. On the other band, it was proposed that acid catalysed reactions such as skeletal isomerization of pamifim [2], disproportionation [3], dehydration of alcohols or cumene cracking over metal supported acid catalysts were promoted by spillover hydrogen (proton) on the acid catalysts. Hydrogen spillover phenomenon from noble metal to other component at

236 room temperature has been reported in many cases [4-5]. In the present work, hydrocracking at low reaction temperature was studied using hybrid catalysts containing H-ZSM-5 zeolite and a supported noble metal from the standpoint of hydrogen spillover.

2. E X P E R I M E N T A L 2.1

Catalyst preparation

Pd/ZSM-5 (0.5wt%) catalyst was prepared using a commercially available ZSM-5 (Toso, HSZ-840NHA) with silic'a/alumina ratio of 44. Pd was introduced by the method of ion-exchange with aqueous solution of tetra ammine palladium chloride. The ion-exchange was carried out at 373 K for 6 h with 0.1wt % Pd(NH3)4CI2 aqueous solution under stirring, the supported Pd/H-ZSM-5 was washed by water until no chloride ion was detected. Oxide-supported palladium was prepared by impregnating a commercial available SiO2 (Aerosil 380, BET specific surface area 380 m2/g) with PdCl~ from its aqueous hydr~xzhrolic 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 4 weight parts of the H-ZSM-5 with one weight part of Pd/SiO2 (2.5wt%) and pressure molding the mixture to granules to 20/40 mesh. Catalysts were activated in air at 723 K for 2 h and reduced in flow hydrogen at 673 K for lh, before use.

2.2

Reaction apparatus and procedure The hydrocracking of n-paraffins was conducted with a continuous Ilow type fixed bed reaction apparatus under pressurized conditions. The reactor was a stainless steel tube with an inner dimneter of 6 mm. The feed material which had been deeply desulfurized was fed by a liquid pump. Prcxtucts were analyzed by a capillary gas chromatography.

3. RESULTS AND DISCUSSION 3.1. Reaction of n-Cl 2112 Figure 1 shows the changes of catalytic activities of a variety of catalysts containing Pd/SiO2 and/or H-ZSM-5 as a function of reaction time. The catalytic activity of H-ZSM-5 was unaffected by the atmospheres and decreased quickly. Pd/SiO2 showed little activity for both dehydrogenation and cracking of n-C12H26. On the oilier hand, the catalytic activity of a hybrid catalyst comprising Pd/SiO2 and H-ZSM-5 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.

237 It is well known that the supported platinum show a high catalytic activity for the dehydrogenation of alkane whereas the supported palladium does not. The rcsults shown in Figure 1 suggest that the dehydrogenation activity of supported metal is not essential for the appearance of the alkane hydrocracking activity. The essential point is that the hydrogen-activating component is contacted with acidic catalyst. The present authors have pointed out that the skeletal isomerization of lower alkane is effectively promoted by the hybrid catalyst composed of Pd/SiO2 or Pt/SiO2 and H-ZSM-5 and that the hydrogen spillover is the key step of isomerization reaction. In the present case also, the hydrogen migration from Pd/SiO2 to H-ZSM-5 should be essential for the high and stable catalytic activity. 100 90

-O-Pd-Iiybrid cat. in i-12a)

80 70

--O-Pd-hybrid cat. ill N2a)

60

-13-II-ZSM-5 il~ I12 b)

5o :,- 40

--II-II-ZSM-5 i~ N2 b)

0

9

0

tO 30 20

-/X- l'd/SiO 2 ill 1-12b)

10 0 1 ~A

0

,A

/k

/1~

I 2 3 Time on stream (h)

Figure 1. Hydrocraking of dodecane with Pd/SiO2-H-ZSM-5 hybrid catalyst. 503 K, I 0 MPa, tl2/n-C 12=9, apd/SiO2:It-ZSM-5=I' I, W/F=2.4 g h tool -I , bW/F=I.2 g h tool-1

Figures 2 shows the carbon number distribution in tile products obtained by hydrocracking of n-Cl2H26. The carbon number distribution of H-ZSM-5 catalyzed reaction was not affected by the atmosphere. In the absence of hydrogen, the carbon-number distribution in the hybrid catalyst containing Pd/SiO2 system was very similar to that in H-ZSM-5 system. However, the hybrid catalyst containing H-ZSM-5 and Pd/SiO2 showed high selectivity for hydrocarbon fragments in file middle range (C8-C9) and for the isomers of C12H26 under hydrogen atmosphere, in spite of its high conversion level. It can be said that secondary cracking of cracked fragment is prevented ill H2-Pd-hybrid catalyst system.

238 3O

_o_Pd-hybrid ill 112 t

25

__e._Pd-hybrid in N 2 "d 20

E~

_u_H-ZSM-5 i~ I12

.r.-~

tJ

--m-II-ZSM-5 ii~ N 2

10 I

r.,o

0 0

2 4 6 8 10 12 Carbon Ilunlber of products

14

Figure 2. Carbon number distribution in products of hydrocrackillg of n-dodecane. The same experiment as Figure I (I h time on stream).

3.2.

Reaction of n-C7H1 6

In tile case of n-dodecane hydrocracking, the product distribution is affected by secondary cracking reaction of the middle range hydrocarbon formed in primary cracking. Therefore hydrocracking of n-heptane, which should give only C3 mid C4 hydrocarbons in acid catalyzed cracking was studied. In Figs. 3 to 5 show catalytic activity as a function of reaction time, carbon-number distribution of products and the composition of C4 products. As was the case of n-dodecane cracking, 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 (little n-Chill0 was formed), whereas the products on H-ZSM-5 alone distributed from C3 to C9 and C4 products contained all kind of paraffins and olefins. The high and stable catalytic activity of Pd-hybrid catalyst in hydrogen atmosphere should be attributed the proton (H+so) formed in the spillover phenomenon from Pd to H-ZSM-5. The wide product distribution for H-ZSM-5 system should be attributed to the reaction path comprising polymerization and cracking. The simple products for the H2-hybrid system should be formed through no other reaction path than the cracking reaction on H-ZSM-5.

239

100

- O - l'd-hybtid

90 80

1!2

''}

--O--I'd-hybrid N 2 a)

~. 70 -El- 11-ZSM-5 112 b)

60 "~ 50

g

40

r,.)

30

- I I - i i-ZSM-5 N 2 t0

:-0----0

0

0 ~ 0 ~ 0

- -:

_/~_ Pd/SiO 2 tl 2 t,) --ilk Pd/SiO 2 N 2 t,)

IO o

0

I

2

3

4

T i m e on Stream (hr) Figure 3. Hydrocrakillg of tl-Ileptalle with Pd/SiO2-1t-ZSM-5 llybrid catalyst. 503 K, 10 MPa, 112/n-C7=9, apd/SiO2:It-ZSM-5=I" 1, W/I:=2.4 g h tool -I, bW/F= 1.2 g h tool-I

60 _ o _ P d - h y b r i d in I-I2 50

-

__.__Pd-hybrid in N 2 -6 40

_ ~ H - Z S M - 5 in H 2 30

- m - H - Z S M - 5 in N 2

20 O'9

0

nil

0

roll

2

K,J

4

~

~

6

8

w

iiw

10

~

~

12

roll

14

Carbon number of products Figure 4. Carbon ilulnber distfibutioll in products of lmydrocrackilig of n-heptane. Tile same experinlelit as Figure 3 (I h tiine on strealn).

240

Figure5. Distribution of C4 hydrocarbon formed in the hyclrocracking of n-C 7 The same experiment as Figure 3.

In hydrocracking of normal paraffin with metal supported acid catalyst, tile iso/nornral ratios in the paraffinic products gencrally exceed the thermodynanlic equilibfiuln. 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. Especially the fact that isobutane is the sole C4 product suggest that the hydmcmcking on the hybrid ca "talyst proceeds through the reaction path shown in Figure 6. C-C-C-C-C-C-C 11~,

+II' ~,

-II2

-I-

C-C-C-C-C-C-C ~

r

C

l l-~so

I1~,.-,

I1~)

c--c-c-c-c

~ slow

C-C=C + C-C-C-C

,~ C-C=C +

"

slow

C

fnst ~

-C-C

C-~--C

+

C-C-C

C-C-C

+

C-C-C

r

Figure 6. llydrocracking model of n-heptane with IM-hybrid catalyst.

241 It has been suggested that formation of multibranched isomers from the feed and cracking are consecutive reactions [6]. Cracking of a normal paraffin must thus proceed thin)ugh the stage of formation of monobranched isomers, dibranched isomers and finally cracked product as in Figure 6, because the high energy barrier for B-scission of monobranched carbenium ion. The isobutylene, which is one of a pair of the primary cracked product, will be hydrogenated to isobutane in the presence ol hydrogen and palladium catalyst. As will be discussed later, the hydride (H-so) as a counter anion of proton (H+so), which is formed in hydrogen spillover process, stabilizes the propyl-carbenium ion to give propane. Thus oligomerization of the cracked fragments and consecutive cracking reaction is prevented in the HaTPd-hybrid system. 3.3.

Effects of h y d r o g e n s p i i l o v e r a n d r e a c t i o n m o d e l

Experimental results and discussion shown above suggest that not olfly cracking activity but also cracking pattern were affected by synergistic effect of hydrogen addition trod supported metal catalyst. 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 on to zeolite surface as a proton and a hydride, where proton promotes the acid catalyzed reaction such as skeletal isomerization, on the other hand, hydride ion stabilizes intermediate carbenium ion to prevent oligomerization and cracking k~ improve selectivity for isomerization:'-). If the hydride ion is not reacted with carbenium ion, the carbenium ion will leave from the acid site as olefin while leaving proton on the acid site. The olefin will be polymerized to higher hydrocarbons and then be cracked on the acid site. Same phenomenon should occur in this system. Hydrogen gas is dissociated on the palladium on SiOz and spills over onto the H-ZSM-5. Hydrogen transfer between the particles as in the case of the hybrid catalyst is a well-known phenomenon. The spillover hydrogen presunmbly exist on the zeolite

II z

r162 tl )x~

I1'

-

I-I-

I!"

Figure 7. Hydrogen spillover model on Pd-Ilybrid catalyst.

242 surface as protons and hydride. Tile proton promotes cracking reaction even at low reaction temperature. The hydride generated simultaneously stabilizes intermediate carbenium ion to prevent over-cracking and promote isomerization of alkane. The model of hydrogen spillover in the hybrid catalyst is shown in Figure 7.

4. C O N C L U S I O N A hybrid catalyst, which was prepared by physical mixing of a H-ZSM-5 and Pd/SiO2. showed an excellent activity for the hydrocracking of n-paraffins. Hydrogen gas is dissociated on the palladium on SiO2 and spills over onto the H-ZSM-5. Tile spillover hydrogen presumably exist on the zeolite surface as protons 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 case of n-dodccane hydrocracking, the hybrid catalyst gave high selectivity for hydrocartxm fragments in the middle rang (C8-C9) and for the isomer of dodecane, in spite of its high conversion level. 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.

5. A C K N O W L E D G E M E N T

This work has been carried out as a research project of tile Japan Petroleum Institute commissioned by the Petroleum Energy Center with the subsidy of the Ministry of International Trade and Industry.

REFERENCES

1. B.S. Greensfelder, H.t{. Voge and G.M. Good, Ind. Eng. Chem., 41 (1949) 2573. 2. K. Fujimoto, K. Maeda and K. Aimoto, Applied Catal., 91 (1992) 81. 3. I. Nakamura, R. lwamoto and A. l-ino, in "New Aspects of Spiliover Effect in Catalysis" (T. Inui, K. Fujimoto, T. Uchijima and M. Masai, eds.), Elsevier, Amsterdam, 1993, pp. 77-84. 4. S.Khoobiar, J. Phys. Chem., 68 (1964) 411. 5. A.J. Robell, E.V. Ballou and M. Boudart, J. Phys. Chem., 68 (1964) 2748. 6. M. Steijns, G. Froment, P. Jacobs, J. Uytterhoeven and J. Weitkamp, Ind. Eng. Chem. Prod. Res. De v., 20 (1981) 654.

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All fights reserved.

243

E F F E C T OF PRESULFIDING ON THE ACTIVITY AND DEACTIVATION OF H Y D R O T R E A T I N G CATALYSTS IN PROCESSING KUWAIT VACUUM RESIDUE M. Absi Halabi', A. Stanislaus', A. Qamra b and S. Chopra b

aPetroleum Technology Department, Kuwait Institute for Scienti)qc Research, P. O. Box: 2 4885, 13109 Safat, Kuwait. bTechnical Services Department, Shuaiba Refinery, Kuwait National Petroleum Co., Kuwait ABSTRACT Presulfiding of hydrotreating catalysts plays an important role in creating the essential surface requirements for optimum activity. It transforms the oxides of molybdenum to MoSx crystallites which are the primary active species in this category of catalysts. In this paper, the results of our investigations of the effects of presulfiding on various reactions taking place in residues hydroprocessing are reported. The changes that the catalyst undergoes as a result of pretreatrnent is also discussed. It has been observed that presulfiding has practically no effect on sulfur and nitrogen removal. In contrast, some improvements were observed for vanadium and nickel removal, asphaltenes reduction, and conversion to distillates. A detailed investigation of the properties of the catalysts with and without presulfiding revealed that the catalysts undergo early deactivation by coke deposition; however, presulfiding reduces the extent of this early deactivation. The unsulfided catalysts were observed to lose over 60% of their initial surface area due to coke deposition and the accessibility to the inner portions of the catalyst pellets appears to be restricted. INTRODUCTION Activation of hydroprocessing catalysts through presulfiding has been widely practiced by the petroleum refining industry in processing distillate cuts. It has been observed that such pretreatment improves the activity and reduces the deactivation rate of the catalyst (1-4). Factors influencing the presulfiding process has been thoroughly investigated (5-10). In addition, techniques for ex-situ presulfidation of hydroprocessing catalysts have been developed (11-13). Studies on the influence of presulfiding on the performance of residual oil hydroprocessing catalysts are relatively scarce in comparison with those related to distillate hydrotreating (14-16). Information available in the literature on the subject are focused on hydrodesulfurization and are often incomplete or conflicting. Furthermore, no reports have been cited on the effect of presulfiding on other hydroprocessing reactions such as hydrodenitrogenation, hydrodemetallation, or hydrocracking, despite the industrial importance of these reactions. The scarcity of the studies and the inconsistencies of the results can largely be attributed to the wide variations in the properties of the feedstocks and catalysts used in the studies, the complexity of the reactions taking place during residue hydroprocessing, and the deactivation mechanisms of the catalyst due to both coke and foulant metals deposition.

244 Table 1. Characteristics of Catalysts Samples Used in Presulfiding Studies

NiMo/AI203 CoMo/Ai203

Test Parameter

Chemical Composition (wt%) (dry basis) MoO3 NiO CoO Physical Properties Average Particle Diameter (mm) Surface Area (m2/g) Pore Volume (H20 Adsorption) (cm3/~)

13.20 4.03 -

13.71 3.4

0.96 312.0

1.01 270.8 0.69

0.70

At Kuwait, a number of residue upgrading processes are operated. The processes use typical hydrotreating catalysts which are brought in contact with the residue feedstock without formal presulfiding. In the present work, a research study was undertaken to assess the effects of presulfiding on various reactions taking place during the hydroprocessing of Kuwait vacuum residue. The study also included a detailed investigation of the effect of pretreatment on catalyst deactivation during the early stages of residue hydrotreating. EXPERIMENTAL Two commercial residue hydrotreating catalysts, NiMo/AI203 and a CoMo/ml203, were used in this study. The properties of the catalysts are summarized in Table 1. Gas oil and Kuwait vacuum residue were used as feedstock for performance evaluation. Detailed analyses of these petroleum fractions are presented in Table 2. Table 2. Physico-Chemical Properties of Residue Feedstock and Gas Oil Used for Presulfiding a Property Density @ 15~ API Gravity Total Sulfur Total Nitrogen C. C.R. Kin-Viscosity @ IO0~

Test Method IP-190 D-1250 Columax

IP-13 IP-71

Unit g/ml API wt% wt% wt% cSt

Vacuum Residue 0.9955 10.6 4.8 0.41 16.9

840 5.81

@ 5o~ Ash Content Metal in Ash Ni V Asphaltenes

Gas Oil 0.8687 31.3 2.08 0.02

IP-4

wt%

0.02

ICAP ICAP IP 143

ppm ppm wt%

36 79 8.4

a Dashes indicate that the test is not applicable or below detection limit.

245 Table 3. Summary of Experimental Conditions of the Test Runs*. Test Run No.

Pretreatment

Feedstock

Duration

R01 R02 R03 R04 R05

None Presulfided None Presulfided Presulfided

Vacuum Vacuum Vacuum Vacuum Gas Oil

10 days 10 days 10 days 10 days 10 days

R06

Presulfided

R07

Presulfided

R08

Soaked in gas oil for 2h. Soaked in gas oil for 2h.

R09

resid resid resid resid

Vacuum resid

Remarks

Test temperature was 380 ~ Test run was terminated after pretreatment

6 h. after presulfiding Test run was terminated after pretreatment

Vacuum resid

6 h. after presulfiding

*Test runs R03 and R04 are carried out using CoMo/AI203.All other test runs are with NiMo/A1203. The performance of the catalysts was studied in a fixed bed reactor testing unit. A 50 ml sample of the catalyst diluted with an equal volume of carborundum was used in the test. In a typical test run in which the catalyst is presulfided, the diluted catalyst is loaded in the reactor and the unit is pressurized with H2 to a pressure of 30 bar. The reactor is heated to 200~ and the presulfiding feed, which consists of gas oil spiked with 5% DMDS, is introduced at a rate of 100 ml/h. The reactor temperature is then raised to 250~ gradually in 2 hours. These conditions are maintained for 8 hours, then the temperature is raised again to 350~ gradually in 8 hours, and maintained under these conditions for an additional 8 hours. The residue feed is then introduced and the unit is brought to the operating conditions of the test, namely, P = 120 bar; LHSV = 2.0; T = 425~ HjOil = 1000 V/V. For the test runs in which the catalyst is tested without presulfiding, the residue feedstock is introduced at 200~ alter the unit is pressurized to 120 bar and the reactor is brought to the operating conditions of the run in 8 hours. The operating conditions and the durations of the test runs conducted in this study are summarized in Table 3. The product is sampled every 12 h for sulfur determination and every 48 h for complete product analysis. Sulfur content is determined using an Oxford Model 2000 Sulfur Analyzer. Selected samples of feed and product are subjected to detailed analysis including total nitrogen, boiling range, asphaltenes and metals content. The spent catalysts were analyzed for Ni, V, C and S. Surface area and distribution profiles of metals in the catalyst pellets are determined according to standard procedures. RESULTS Effect of Presulfiding on Catalyst Performance. The first group of test runs (Test Runs R01-R04) involved comparative evaluation of the effect of presulfiding on hydroprocessing

246

Figure 1. Comparison of the effect of presulfiding on NiMo and CoMo catalyst performance towards various hydroprocessing reactions. reactions. Each of these test runs was carried out for a total duration of 10 days. The reactions that were monitored included hydrodesulfurization (HDS), hydro-denitrogenation (HDN), vanadium (HDV) and nickel (HDN) removal, asphaltenes reduction (HDA), and conversion (HDC) to distillates (524 ~ minus products). Typical deactivation curves were observed for all reactions. The catalyst bed showed high activity at the initial stages of the test run, then the activity exponentially decreased to an equilibrium activity during the first 60 hours on stream. After attaining equilibrium activity, the performance of the catalyst bed remained practically unchanged till the end of the run. Figure 1 shows a comparison of the activities of the presulfided and untreated catalysts toward various reactions. The data indicated are for the catalyst activity after stabilization. It is observed that presulfiding has practically no effect on catalyst activity or deactivation rate for the HDS and HDN reactions. The differences between the efficiencies of the unsulfided and presulfided catalysts are about 2-4%, which is of the same order of magnitude as the experimental error of the analytical procedures used in determining these elements. On the other hand, the results for demetallation, i.e., V and Ni removal, and asphaltenes reduction for both catalysts revealed definite improvements. The improvements for V and Ni removal are on the order of 8-10% and 15-20%, respectively. Similarly, the results presented in Figure 1 show that the presulfided catalyst is more active than the unsulfided catalyst by around 5-10% towards asphaltenes reduction. For hydrocracking, the NiMo/AI203 catalyst exhibited around 15% higher conversion for the presulfided catalyst over the untreated catalyst. The results demonstrate that presulfiding has an overall positive effect on catalyst performance in processing residues. The improvements are specifically in areas of significant interest in residue upgrading namely, conversion to distillates and asphaltenes reduction. The

247 Table 4. Properties of the Spent Catalysts of Pilot Plant Test Runs Under Different Presulfiding Conditions Test Run Code

R01 R02 R03 R04 R05 R06 R07 R08 R09

NiO wt% 4.3 4.1 1.4 1.3 3.9 3.6 3.7 2.8 4.7

V2Os, wt% 8.7 7.7 8.2 7.1 0.1 0.01 0.66 0.01 0.6

C wt% 17.5 18.1 17.1 17.9 8.3 7.1 17.4 13.4 17.0

S wt% 4.8 6.0 6.1 7.1 4.1 4.2 4.1 0.67 3.9

Surface area (m2/g) 123 135 138 122 161 278 124 175 85

failure of the catalyst to show any improvement in hydrodesulfurization is attributed to the conditions of the reaction, particularly the temperature and the properties of the catalyst. Thus at 425~ thermal cracking of C-S bonds is anticipated to be the dominating route for desulfurization of the residue feedstock. As a result, the catalyst role would be marginal and any improvement in catalyst performance would not be reflected by deeper desulfurization. In addition, the observation that the catalysts rapidly deactivate (as discussed in more detail below) by coking and pore plugging when the heavy feedstock was introduced suggests that the active surface of the catalyst is masked from the reactants. Therefore, the rate of catalytic desulfurization would be significantly reduced due to diffusion limitation effects. The improvements in conversion to distillates, asphaltenes reduction, and metals removal observed for the presulfided catalysts, particularly for the NiMo catalysts, is a significant result of this study. The close connection between these three types of reactions is in line with previous results (17) indicating that the metals are associated with asphaltenes. Any increase in the rate of asphaltenes cracking would also result in increases in metals removal and conversion to distillates. The observation that the NiMo/AI203 catalyst is more active than the CoMo/AI203 catalyst is in line with the prevailing views that NiMo catalysts are better hydrogenation and hydrocracking catalysts. The spent catalysts from the Test Runs R01-R04 were examined to assess the effect of presulfiding on catalyst deactivation. Table 4 includes the results of these analyses. The percentages of C, NiO, and VzO5 are nearly the same for both the presulfided and unsulfided catalysts. The surface area of all catalysts decreased substantially to around 35-40% of that of the fresh catalyst. Both chemical and physico-chemical properties revealed that the catalysts were deactivated to nearly the same extent irrespective of the pretreatment conditions. However, the distribution of foulant metals in the spent catalyst pellets revealed clear differences between the presulfided and untreated catalysts. Figure 2 shows vanadium and nickel distribution for the spent NiMo/AI203 and CoMo/A1203 catalysts of the unsulfided and presulfided test runs. For the unsulfided catalysts, both Ni and V are observed to be more highly concentrated on the outer edges of the pellet.

248

R01, NiMolUnsulfided

R03, CoMolUnsulfided

..

R02, NiMolPresulfided

! . . .

I . . . . , ,

R04, CoMolPresulfided

Figure 2. Vanadium and nickel distribution profile across pellets of the spent catalysts of test runs R0 l-R04. The y-axis is the relative concentration of metal.

In contrast, the two metals show higher penetration into the pellet for the presulfided catalysts. This provides clear indication that diffusion of feedstock molecules within the catalyst pellets is somehow more restricted for the unsulfided catalyst in comparison with the presulfided one. The exact reasons for the improvements observed regarding the hydroprocessing reactions are not fully understood. Sulfiding transforms the oxides of molybdenum to MoSx crystallites in highly dispersed form (5,10,18,19). The MoSx crystallites are present as thin hexagonal shaped slabs and the promoter atoms (Co or Ni) are located on the edges of the slabs forming a highly active phase known as Co-Mo-S or Ni-Mo-S phase (19,20). Although it is now well established that the sulfur vacancies at the edges and corners of MoSx slabs are the active sites for the catalytic reactions taking place during hydrotreatment, less is known about the processes occurring during the build-up of the active phase from the oxidic precursor of the catalyst. Furthermore, the influence of coke deposited on the surface of the catalyst during the sulfidation process with spiked gas oils appears to have received little or no attention. It is likely that passivation of the highly active acidic oxidic sites by presulfiding reduces the coke forming tendency. In addition, the small amount of coke deposited on the catalyst during the sulfiding process may also have a passivating effect. This coke apparently has no

249 detrimental effect on the catalyst surface area, but appears to have a beneficial passivating effect. Consequently, the presulfided catalysts provide better access of the large asphaltene molecules to the internal pores, thus, enhancing the catalytic route for conversion. This is further confirmed by higher penetration and even distribution of foulant metals within the interior of the catalyst pellets in the case of the presulfided catalysts. Effect of Pretreatment on Early Catalyst Deactivation. To understand further the mechanism through which presulfiding affect the performance of residue hydroprocessing catalysts, special test runs were conducted using the NiMo/A1203 catalyst. The test runs were aimed to assess the effect of presulfiding on early catalyst deactivation during the processing of residues and gas oil. To assess the effect of pretreatment on early coke formation, two sets of runs, R06/ R08 and R07/ R09, were conducted. In the first set, the runs were terminated immediately after pretreatment without introducing vacuum residue, while for the second set, the runs were terminated after vacuum residue was introduced for 6 hours only. The catalysts in runs R06 and R07 were presulfided as described in the Experimental Section, while those in runs R08 and R09 were simply soaked in recycled gas oil for 2 hours, simulating a practice adopted by the industry. A comparison of the carbon percentages of the spent catalysts (Table 4) of test runs R01 and R02 with those of test runs R07 and R09 shows that all four catalyst samples have nearly the same carbon content. This clearly demonstrates that almost all of the coke on the spent catalyst is deposited during the first few hours of the run. Furthermore, the surface area for both presulfided and unsulfided catalysts was significantly reduced during the early hours of the run. The data in Table 4 also shows that conventional presulfiding results in lower coke deposition compared with soaking in gas oil (Run R06 vs. R08). Furthermore, the conventionally presulfided catalyst retained most of its surface area. It is also interesting to observe that when the catalyst is exposed to residue, the percentage of coke deposited is unaffected by pretreatment (Run R07 vs. R09); however, the surface area of the presulfided catalyst is significantly higher than that of the unsulfided catalyst. This may be partly due to differences in the nature of coke deposited and partly to the pore plugging by the foulant metals. The effect of presulfiding on the fouling of the catalyst by metal deposition during the early hours of the run was investigated by studying the distribution profile of the low levels of vanadium in the spent catalysts of test runs R07 and R09. The penetration of vanadium in the presulfided catalyst of Run R07 was around 50% of the radius of the pellet, whereas that for the unsulfided catalyst (Run R09) was limited to around 25% of the radius, despite equal carbon content. This lends additional support to our earlier argument that presulfiding passivates the highly active sites on the surface of the catalyst, permitting deeper diffusion of the feed into the catalyst pellet. The observation that vacuum residue causes significant coke deposition on catalysts was further ascertained by conducting a special test run (Run R05) which is similar to Run R02 except straight run gas oil was used instead of vacuum residue as feedstock at 380 ~ The

250 properties of the spent catalyst of this test run when compared with that of R02 (Table 4) shows that the percentage of carbon is over 50% lower and the surface area is significantly higher. Our results on the effects of coking on the catalyst and its performance are in agreement with the generally accepted view on the deactivation of residue HDS catalysts, in that the main cause of deactivation is coke and metal deposits near the pore mouths (21). These deposits lead to constricting, but not completely blocking the pores of the catalyst. However, the results of the present work highlight the importance of initial coking on catalyst deactivation.

Conclusion The present study provides evidence that presulfiding by conventional procedures has an overall positive effect on catalyst performance in processing residues. On the other hand, soaking the catalyst in gas oil leads to partial sulfidation, but is not an adequate pretreatment. The study also revealed that the catalyst rapidly loses most of its surface area during the early hours of the run by initial coking. Presulfided catalysts were slightly better than the untreated catalyst in maintaining the surface. Furthermore, presulfided catalysts showed improved foulant metals distribution throughout the catalyst pellet. These conclusions are specific to the catalyst and feed used. Further work on the effects of various operating conditions on initial coke formation, its nature, and its role in catalyst deactivation is in progress.

References 1. J. S. Jepsen and H. F. Rase, Ind. Eng. Chem. Prod. Res. Dev. 20 (1981) 467. 2. J. Laine, K. C. Pratt, and D. L. Trimm, Ind. Eng. Chem. Prod. Res. Dev. 18 (1979) 329. 3. H. Gissy, R. Bartsch, and C. Tanielian, J. of Catalysis 65 (1980) 150. 4. B. Scheffer, E. M. van Oers, P. Arnoldy, V.H.J. de Beer, and J. A. Moulijn, Applied Catalysis 25(1986) 303. 5. H. Hallie, Oil and Gas Journal, Dec. 20 (1982) 69. 6. F. C. Riddick, Jr. and B. Peralta, Hydrodesulfurization of Oil Feedstock with Presulfided Catalyst, US Patent No. 4 213 850 (1980). 7. D. R. Herrington and A. P. Schwerko, Hydrotreating Process Utilizing Elemental Sulfur for Presulfiding the Catalyst, US Patent No. 4 177 136 (1979). 8. P. Clements and O. L. Davies, Activation of Hydrotreating Catalysts, U S Patent No. 3 915 894 (1975). 9. R. Prada Silvy, P. Grange, F. Delanny, and B. Delmon, Applied Catalysis 46 (1989) 113. 10. J. van Gestel, J. Leglise and J. C. Duchet, Journal of Catalysis 145 (1994) 429. 11. G. Berrebi, Process of Presulfurizing Catalysts for Hydrocarbons Treatment. US Patent No. 4 530 917 (1985).

251 12. M. de Wind, J. J. L. Heinerman, S. L. Lee, F. L. Platenga, C. C. Johnson, D. C. Woodward, Oil & Gas J., Feb. 24 (1992) 49. 13. B. J. Young, Catalyst Composition and Sulfiding Method. US Patent No. 3 563 912 (1971). 14. S. J. Yanik, A. A. Montagna and J. A. Frayer, Method for Presulfiding Hydrodesulfurization Catalysts, US Patent No. 4 111 796 (1978). 15. T. Takatsuka, H. Nitta, S. Kodama, and T. Yokoyama, Preprints Symposium on Advances in Petroleum Processing, American Chemical Society, (1979) 730. 16. W. L. Brunn, J. A. Frayer, J. A. Paraskos, and S. J. Yanik, Residual Oil Hydrodesulfurization Process by Catalyst Pretreatment and Ammonia Addition, US Patent No. 3 859 204 (1975). 17. A. Stanislaus, M. Absi-Halabi, F. Owaysi, and Z. Hameed, 'Effect of temperature and pressure on the hydroprocessing of Kuwait vacuum residue." Kuwait Institute for Scientific Research, Report No. KISR2754, Kuwait (1988). 18. R. Prada Silvy, F. Delannay; P. Grange; and B. Delmon, Polyhedron 5(1/2) (1986)195. 19. E. Payen, S. Kasztelan, S. Housenbay, R. Szymanski, and J. Grimbolt, J. Phys. Chem. 93 (1989) 6501. 20. H. Topsoe and B. S. Clausen, Applied Catalysis 25 (1986) 273. 21. M. Absi-Halabi, A. Stanislaus, and D. L. Trimm. Applied Catalysis 72 (1991) 193.

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Catalysts in Petroleum Refining and Peu'ochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All fights reserved.

253

CONTINUOUS DEVELOPMENTS OF CATALYST OFF-SITE REGENERATION AND PRESULFIDING P. Dufresne a, F. Valeri a and Dr. S. Abotteen b

a Eurecat SA - Quai Jean-Jaurbs 07800 La Voulte-sur-Rhdne, France bAl-Bilad Catalyst Co., PO Box 10174, Jubail Industrial City, Saudi Arabia

ABSTRACT This paper presents the insight of ex-situ regeneration and ex-situ presulfiding of hydroprocessing catalysts. The ex-situ regeneration of used hydroprocessing catalysts offers a better performance recovery because of the precise temperature control. During the oxidation step, carbon and sulfur removals are optimized to get a maximum surface area and oxygen chemisorption capacity recovery, this last parameter being related to active sulfide phase dispersion. The ex-situ presulfided catalysts having an oxysulfide form of active metals can be converted into a sulfided form during the activation stage in a hydroprocessing reactor, providing a quick and convenient start-up of the commercial units. INTRODUCTION Off-site regeneration of hydroprocessing catalysts has been widely accepted over the last ten years by the petroleum refining industry. A large proportion of hydroprocessing catalysts (hydrotreating, hydrocracking) containing noble metals (Pd, Pt) or non-noble metals (Mo, W, Co, Ni) on inorganic oxides such as alumina, silica-alumina or zeolites are currently regenerated off-site (1). This technique is preferred above the conventional in-situ technique for a number of reasons, including safety, time savings and better activity recovery. The high quality achieved by the modern technologies of off-site regeneration often allows to perform more than one cycle with the same catalyst batch. In addition to this economic incentive for regeneration, more and more stringent environmental regulations for disposal of spent catalysts encourage their reuse. The primary objective of oxidative regeneration is coke removal. However, this cannot be done without the oxidation of metal sulfides to their corresponding oxides. The structures of the oxide and sulfide forms of these metals are completely different, but conversion from one form to the other is quite reversible. Laboratory scale regeneration has been performed on different commercial spent catalysts which were in use in various types of units. This property, which allows the catalysts to be regenerable under oxidizing conditions, has of course been confirmed by industrial results. On the other hand, the structural identity of fresh and regenerated catalyst has often been proven (2), as well as the potential loss of physical and catalytic properties under severe regeneration conditions (3,4).

254 Table 1. Characteristics of Spent Catalysts.

Type Metals amount Unit feed C wt% S wt% LOI wt%

Cat A

Cat B

Cat C

Cat D

CoMo medium AGO* 5.7 6.7 12.4

NiMo medium AGO* 7.5 7.5 16.3

NiMo high VGO * 15.8 9.3 19.8

CoMo medium AGO* 10.9 10.4 16.6

* AGO: Atmospheric Gas Oil, VGO: Vacuum Gas Oil

EXPERIMENTAL The spent catalysts used in these studies are commercial catalysts unloaded from different types of hydroprocessing units. The characteristics of the spent catalysts are summarized in Table 1. Lab scale catalyst regeneration is performed in a ventilated muffle furnace, using 50g of catalyst dispersed as a monolayer on a plate. This plate is introduced directly into the furnace which is preheated at the desired temperature for 4 hours. For regeneration temperatures higher than 300~ the furnace is first held one hour at 300~ the temperature is then raised at a rate of 10~ to the desired regeneration temperature. Carbon and sulfur contents are measured in a LECO CS 125 analyzer. LOI is the weight loss at 500~ Surface area is measured in a MICROMERITICS Flowsorb 2300 (BET 1 point). Dynamic Oxygen Chemisorption analysis is performed as follows: fresh or regenerated catalyst is presulfided by impregnation with a polysulfide to introduce a sulfur amount equivalent to 100% of the stochiometry (MoS2, Ni 3 $2, Co9S8). Then lg of the presulfided catalyst is introduced into the glass cell of the chemisorption apparatus (GIRA chromatograph), activated with H 2 at 350~ and cooled under a helium flow at 60~ where pulses of oxygen are mixed to the flow of helium and analyzed by a thermal conductivity detector. Thermogravimetric and differential thermal analysis are performed with a Setaram TG DTA 92-12, using 50mg of sample, a 5~ heating rate under air. Carbon and sulfur removal It is performed in a ventilated muffle furnace, as described above. Catalysts A and B were regenerated at increasing temperatures and the remaining amounts of carbon and sulfur measured at each step. In Figure 1, the carbon and sulfur removal of both Cat A (CoMo) and Cat B (NiMo) are presented. It can be clearly seen that for both types of catalyst the carbon removal occurs in a single, swit~ step whereas the sulfur removal clearly shows two separate domains of removal.

255

[wt%] 10

Carbon Cat A m

Sulfur Cat A D Carbon Cat B Sulfur Cat B 0

0

0

200

400

600

800

Regeneration temperature [~ Figure 1. Carbon and sulfur removal for Cat A and B versus regeneration temperature. It appears that sulfur removal begins between 150 and 200~ with a maximum between 200 and 250~ Subsequently, the remaining sulfur is very slowly removed at increasing temperatures. Carbon combustion takes place between 300 and 350~ and is completed at 450~ Thermogravimetric and Differential Thermal Analysis has been performed on Cat D. The TG and DTA profiles in Fig 2 show three different steps. The first one is the evaporation of hydrocarbons up to 200 ~ with a moderate endotherm. The second step is the oxidation reaction of metal sulfides to oxides (most of the Mo sulfide, and part of the Co sulfide), starting around 200-250 ~ The third step around 350-450 ~ is strongly exothermic, due to carbon burn-off as well as the remaining of sulfides oxidation. The carbon burn-off reaction finishes around 500 ~ in this experiment performed on a dynamic mode at the heating-up rate of 5 ~

256 -G c~'/~.:~'

i~E~rr ~

I

;~:-;'o,,~

-

'

'

"

'~

'

~

'

" -'

-'

t:3

0 : O,

300

-0.50

DTG

:C,O

~

-25 -50

: .

9

vl

_

200 !

~

~

250 .

.

..

300 .t

350 ~

"00 9

~50 ,

500 ~.,

550 I

E,O0 !

TEP'.PER~TURE (C.) 700 750 BOO_30

650 I

-.

."

,,

_ I

,

9

Figure 2. Thermogravimetric and Differential Thermal Analysis of Cat D under air flow and a heating rate of 5 ~ Diffusion limitations The oxidation reactions of carbon and sulfur on hydroprocessing catalysts seem to be kinetically controlled by oxygen diffusion inside the catalyst porosity. Figure 3 shows the carbon and sulfur removal for Cat C which contains a very high amount of nickel and molybdenum, and an appreciable load of carbon. It is clear that the sulfur elimination occurs at higher temperatures than for the other catalysts and is simultaneous to carbon combustion. A tentative explanation of this phenomenon would be that the diffusion of oxygen in the microporosity is limited by coke deposit which needs to be at least partly removed to allow complete sulfur oxidation. Quality assurance by SA and DOC The quality of the regenerated catalyst was studied by means of Surface Area (SA) and Dynamic Oxygen Chemisorption (DOC). DOC has been proven previously to be an elegant technique for the evaluation of hydrotreating catalysts. Hydrodesulfurization activity has been correlated with the amount of oxygen chemisorbed at low temperature (5, 6).

257 [wt%]

[% relative]

20

100 i

18

90

16 80 14 12

70

Carbon

10

60

Sulfur []

50 40 SA DOC

30 0

"

0

200 400 600 Regeneration temperature [~

Figure 3. Carbon and sulfur removal for Cat C (NiMo) versus regeneration temperature.

800

20 400

9 ,

,

,,

I

~

,

,

[] I

,

,

,

I

,

500 600 700 Regeneration temperature [*C]

,

,

800

Figure 4. Relative surface area and DOC compared to fresh Cat C (NiMo).

This relationship between DOC and activity can be explained by the greater affinity of oxygen for the edge sites of the MoS2 layered structure, whereas these sites have been indicated as being responsible for HDS activity (7). So it is clear that DOC must not be used directly to compare different types of catalyst, and it will not replace catalytic activity measurements. However, DOC gives a good indication about the dispersion state of the active phase. The results presented here below show that the DOC response is very sensitive to regeneration conditions for a given catalyst, and that the technique can be used as a quality control. The Surface Area (SA) of the regenerated catalyst and the values of Dynamic Oxygen Chemisorption (DOC) of the presulfided catalyst were established as a function of temperature. Figures 4 shows the relative SA and DOC of a regenerated Catalyst D, respectively, compared to fresh. The sintering of the alumina support occurs at temperatures higher than 620-650~ which is typical for such high SA alumina carriers. However, the loss of DOC is already very severe at temperatures around 500~ This indicates that active phase dispersion could be affected already at this low temperature. X-ray diffraction was used as a complementary technique and evidenced the formation of a crystallized phase for the regeneration at higher temperatures. The peaks can be attributed to a mixed oxide of molybdenum and nickel with the following formula: x (NiO), (MOO3), z (H20). So the metals partially sinter during regeneration into a bulky phase, which is no longer active. The appearance of this nickel - molybdate type phase, expressed as the area of

258 Table 2 Results of industrial regenerations on commercial CoMo HDS catalyst of 1.3 mm diameter extrudate. Catalyst 1st Cycle - Used - Regenerated 2nd Cycle - Used - Regenerated 3rd Cycle - Used - Regenerated

LOI

C

S

(%)

(%)

(%)

SA

DOC

20.0 -

8.5 0.3

10.7 0.6

173

6.3

16.9 -

10.9 0.3

10.4 0.6

166

17.4 -

11.9 0.1

9.3 0.4

160

BCS

L

CBD

(mm)

(g/cc)

1.0

3.1 3.1

0.79

5.9

1.5

2.5 2.6

0.80

5.3

1.2

2.4 2.4

0.80

(m2@) (cc/g)(MPa)

the primary peak, has been linearly correlated to the loss of active sites, expressed by the relative loss of DOC for regenerated catalyst compared to fresh.

Industrial regeneration results The industrial oxidative regeneration was performed under the optimized condition determined by the laboratory regeneration study. Table 2 summarizes the results of three cycle use of the same catalyst for an LGO HDS application, made possible by the successful regenerations. The loss of SA may have resulted not only from the regenerations, but also from the hydrotreating operation itself (metal poisoning, thermal shock etc.). For an industrial catalyst, physical properties such as strength and loading density are important as well as C/S contents, SA and DOC. Although some length reduction can be observed from the first cycle to the second cycle, other properties like the bulk crushing strength (BCS) were within acceptable level. The relative DOC figures are 100, 94 and 84% for the 1st, 2nd and 3rd cycle respectively, suggesting that in case the 4th cycle might be critical for high performance requirement. PRESULFIDING

Introduction The ex-situ presulfided hydroprocessing catalysts processed by Sulficat | Process have been successfully used over 30,000 tons worldwide in petroleum refining industry since 1986. The main advantage over a conventional in-situ sulfiding is to allow a refinery a quick and efficient start-up without requiring addition of a sulfur compound. Conventional hydroprocessing catalysts contain oxidic active metal forms of Mo, W, Co and Ni, which are converted into their active sulfided forms by sulfiding in an hydroprocessing reactor (in-situ sulfiding). Since the sulfiding reactions are rather highly exothermic, careful

259 attention should be paid to the addition of a sulfur agent, resulting in a long procedure for insitu sulfiding. The ex-situ presulfided catalyst has a metal form of stable intermediate oxysulfide with the sulfur amount sufficient enough to be converted to the working sulfided form. Therefore atter loading the catalyst into a reactor, the activation can be done with hydrogen gas without requiring any sulfur compound addition. The development and application of ex-situ presulfided catalysts have been discussed in previous papers (8-12). The reactions which take place during the activation of the catalyst will first be presented and then followed by an example of industrial application.

Experimental The presulfided catalyst used in this study was industrially prepared of a commercial CoMo type HDS catalyst by the Sulficat| Process, impregnation with an organic polysulfide compound followed by fixation at an elevated temperature. The amount of sulfur introduced on the catalyst has been calculated from the theoretical stoechiometry necessary to convert all the molybdenum and cobalt into MoS 2 and Co9S 8. XPS: The X-ray Photoelectron Spectroscopy (XPS) was performed, with a FISONS Instruments ESCALAB 200R, using the A1Kot ray at 487 eV, to characterize the presulfided catalyst and the catalyst aider activation in comparison with the starting oxidic catalyst and the sulfided catalyst. The activation of the presulfided catalyst was made under H 2 at 20 bar by raising the temperature from 20 ~ to 350 ~ at 240 ~ heating-up rate and then keeping it at 350 ~ for 1 hr, with a simulation of gas recycling by injecting 200 ppm H2S from 200 ~ The sulfided catalyst was prepared by treating the oxidic catalyst under 15% H2S/85% H 2 at 350~ for 1 hr. HP-TPR: The High Pressure Temperature Programmed Reduction (HP-TPR) with Mass Spectrometer (MS) from FISONS Instruments was performed to observe the activation reactions of the presulfided catalyst under H 2 at 20 bars and 240 ~ heating-up rate.

Results Figure 5 shows the HP-TPR profiles. Sulfur in the presulfided catalyst begins to react with H 2 at 150 ~ to yield H2S, which immediately reacts with the metal oxysulfides to form the sulfides and H20. The activation reactions appear to complete by 300 ~ It should be noted that some H2S slipped out, in this case approximately 7% of the stoechiometry. Figure 6 shows the XPS spectra ofMo 3d, Co 2p and S 2p. The binding energies of the presulfided catalyst are different from those of the oxidic catalyst, but also from those of the sulfided catalyst. The calculated proportions of Mo oxidation number calculated by peak decomposition were 62% Mo 6+ and 38% Mo 4+, although the binding energy of the Mo 4+ is shifted by about 1 eV from that of Mo 4+ in MoS 2. The binding energy of the S corresponds to that of a polysulfide. The cobalt spectrum indicates that Co is more sulfided than Mo. By the o Mo 4 + of MoS 2 and 20~ o activation of the presulfided catalyst, the Mo changed to 80~ Mo 6+, of which proportion is equivalent to that of the sulfided catalyst, 78% Mo 4+ and 22%

260 (n

_

i

u

E ,

R

03

"

O

"

3:

"

E 03

t'N

"

0

gO

100 150 200 T e m p e r a t u r e (~

250

300

Figure 5. HP-TPR profiles of presulfided catalyst under H 2 at 20 bar.

Molybdenum 3d XPS spectra

Cobalt 2p 3/2 XPS spectra

Pr Su Ac H2

Su H2S -

/

~~..,.. /

223

225

227 229 231 233 235 Binding Energy (eV)

PrSuAcH2 ~

P

r

S

u

Ox 237 239 774 776 778 780 782 784 786 788 79,0 792 Binding Energy (eV)

Sulfur 2p 1/2 XPS spectra

]

~ H_~2_2s

159

161

163 165 167 Binding Energy (eV)

169

171

Figure 6. XPS spectra of Mo 3d, Co 2p and S 2p regions. (Ox: Oxidic catalyst; PrSuAcH2: Activated catalyst of PrSu under 350~ CPrSu: Presulfided catalyst; SuH2S" Sulfided catalyst of Ox under 15% H2S 85% H 2 at 350 ~

261

(~

400

4

3O0

3

o

O

2~

200

tO t~

"r"

E'

s

I-..- 100 9

,

o

9

!

s

9

,

9

,

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:

I

~s

i

I

t

20

l

l

2s

i

I

30

i

'

~s

0

4O

Time (hour)

Figure. 7. Activation of presulfided catalyst in industrial start-up for Distillate Hydrotreater in gas-liquid mixed phase with gas recycle. Mo 6+. Cobalt became also essentially sulfided and S is characteristic of a sulfide S2". We may conclude that the presulfided catalyst after the activation results in giving similar active metal forms to those in the catalyst sulfided by a mixture of H2S and H2.

Industrial application example An example in Figure 7 presents the gas-liquid mixed phase activation for the presulfided catalyst (60 t) in a 120,000 BPSD Distillate Hydrotreating unit. The activation reactions started at 140 ~ with exotherm of delta T of 39 ~ at the maximum, which subsided within 1 hr. As H2S was emitted during the activation, the effluent gas was recycled. The unit achieved a quick and successful start-up saving the time by 24 hr compared to the conventional in-situ presulfiding. ACKNOWLEDGMENTS The TPR experiments and the XPS spectra have been obtained by the laboratory directed by Mr. Breysse at IRC, Institut de Recherches sur la Catalyse, Villeurbanne (France). The authors are grateful to them as well as to N. Brahma, J. Darcissac and other Eurecat people for their contribution to this paper. REFERENCES 1. J.H. Wilson, AIChE Summer Meeting, San Diego, August (1990) 2. Y. Yoshimura, E. Furimsky, T. Sato, H. Shimada, N. Matsubayashi and A. NishijimaProc.9th Int. Cong. Catal., Calgary (Canada), M.J.Phillips and M.Ternan(Editors), p.136 (1988) 3. A.V. Ramaswamy, L.D.Sharma, A.Singh, M.L.Singhal and S.Sivasanker- Appl.Catal., 13 (1985) 311

262 4. 5. 6. 7. 8.

S.M. Yui, NPRA Annual Meeting, San Antonio, Texas, AM 91-60, (1991) S.J. Tauster and K.L. Riley - J.Catal. 67 (1981) 250 S.J. Tauster and K.L. Riley- J.Catal. 70 (1981) 230 H. Topsoe, B.S.Clausen, R.Candia, C.Wivel and S.Morup - J. Catal. 68 (1981) 433 J.H. Wilson and G. Berrebi, "Off-Site Presulfiding of Hydroprocessing Catalyst", Am. Chem. Soc., Toronto Meeting, 1988. 9. P. Dufresne, G. Berrebi and J.H. Wilson, "A New Way to Start-up Hydrocrackers", Am. Chem Soc., San Francisco Meeting, 1989. 10. J. H. Wilson, "Innovations in Off-Site Catalyst Services", Am. Inst. Chem. Eng. Spring National Meeting, 1991. 11. M. de Wind, J..J.L. Heinerman, S.L. Lee, F.L. Plantenga, C.C. Johnson and D.C. Woodward, Oil and Gas Journal, February 24 (1992). 12. S. R. Murff, E.A. Carlisle, P. Dufresne and H. Rabehasaina, "The Sulficat Presulfided Catalyst Experience", Am. Chem. Soc. Denver Meeting, 1993.

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All fights reserved.

263

THE PRODUCTION OF LARGE POLYCYCLIC AROMATIC HYDROCARBONS DURING CATALYTIC HYDROCRACKING John C. Fetzer

Chevron Research and Technology Company, P.O. Box 1627, Richmond, California 94802 U.S.A.

ABSTRACT Modem analytical techniques, including HPLC with diode-array UV detection and spectrofluorometry, have been used to identify the large polycyclicaromatic hydrocarbons (PAHs) produced in catalytic hydrocracking. Several reaction pathways have been inferred from these structures. New simpler analytical methods can then be used to monitor PAH production. I. INTRODUCTION Polycyclic aromatic hydrocarbons (PAHs) are produced as side-products during the catalytic hydrocracking of processed petroleum to produce lighter products such as the motor fuels and lubricating oils. These P AHs have been implicated in a variety of process and product problems, including catalyst fouling, process pipe plugging, and product oxidation and coloration. As part of a continuing program of investigating the chemical changes that occur during petroleum processing, efforts in this laboratory have led to a better understanding of the PAH species produced and their production routes. This work will be summarized in a comprehensive manner that describes the various analytical Procedures and results. Several specific reaction pathways were found and will be described [ 1-5]. The analysis of PAHs were done on two types of samples, hydrocracked oils and extracts from process deposits. High-performance liquid chromatography with diode-array detection (HPLC-DAD) can be used to analyze the PAHs of up to 12 rings. A DAD collects the complete UV absorbance spectrum of the chromatographic eluents through the simultaneous monitoring of a large number of wavelengths. For PAHs, the UV absorbance spectrum is a fingerprint that reflects the number and arrangement of the aromatic rings. Each PAH, therefore, inherently possesses a characteristic (fingerprint) spectrum [6,7]. Isomers, even those of many rings and very similar structures, differ greatly in their spectral patterns because the n electrons in each experience a different environment based on the isomer's shape and n electron delocalization. Alkyl substitution on a PAH results in a red shift of the spectrum, but the same overall pattern of maxima and minima is seen. Therefore, even though hydrocracked oils are chemically very complex, the P AHs are relatively easy to determine by HPLC-DAD because the saturate hydrocarbons do not respond, there are only trace amounts of heteroatom containing species, and the PAHs can be differentiated by their aromatic classes.

264 Normal-phase HPLC, using an amino-bonded phase, was used for determination of the PAHs of up to 7 rings. This type of separation results in elution by the number of rt bonds. A special reversed-phase octadecyl column was used for PAHs of 7 through 12 rings. This HPLC packing, Vydac 201TP5, is well known for its orderly structure and separates the PAHs by their overall shapes. It has been compared to the liquid-crystal phases used in gas chromatography. It provides the best isomer specific separation of PAHs. For PAHs of more than 12 rings, HPLC-DAD methods cannot work due to the strong affinity of these PAHs to the HPLC packings. Other analytical tools based on mass and fluorescence spectrometries must be used. 2. E X P E R I M E N T A L The HPLC was a Perkin-Elmer 410 quaternary-solvent system with a Waters 991 DAD. A model compound library was used for PAH identification. Standard PAHs were obtained either through synthesis or from a wide variety of sources. Fluorescence spectra were collected with a Perkin-Elmer MPF-6. All solvents were Burdick and Jackson HPLC grade, and used as received. With normal-phase HPLC, oil samples were analyzed as is by simple dilution in nhexane. A Du Pont Zorbax amino-bonded phase column, 25 cm x 0.46 cm ID, was used, with n-hexane and dichloromethane as solvents. For reversed-phase HPLC, Vydac 201TP5 columns were used (25 cm x 0.46 cm ID for analytical scale and 25 cm x 1 cm ID for preparative scale). Samples for reversed-phase HPLC were fractionated in order to remove the saturated hydrocarbons which can interfere with the separation mechanism. The samples dissolved in n-hexane were passed Baker silica solid-phase extraction cartridges. The PAH fraction was then collected by eluting with a 1:1 mixture of dichloromethane and methanol. Acetonitrile and dichloromethane were used in the HPLC gradient. Process deposits were exhaustively Soxhlet extracted with n-hexane to remove residual oil, then with dichloromethane to remove the PAHs. The efficiency of extraction was determined by a fluorescence spectrum of the extract. 3. RESULTS AND DISCUSSION The HPLC analysis ofhydrocracked oils and the dichloromethane extracts of process deposits revealed a number of PAHs, but not the large variety expected. Under normal operating conditions, only one, two, or three PAHs were found for each ring number (alkylation is disregarded, only the core aromatic structures of the molecules are considered). These were (up to 10 tings, with ring number in parentheses): benzene (1),naphthalene (2), phenanthrene (3), pyrene (4), benzo[e]pyrene (5), benzo[ghi]perylene (6), coronene (7), dibenzo[e,ghi]perylene (7), benzo[a]coronene (8), benzo[pqr]naphtho[8,1,2-bcd]perylene (8), p h e n a n t h r o [ 5 , 4 , 3 , 2 - e f g h i ] p e r y l e n e (8), n a p h t h o [ 8 , 1 , 2 - a b c ] c o r o n e n e (9), dibenzo[ij,rst]naphtho[2,1,8-defg]pentaphene (9), ovalene (10), and benzo[rst] dinaphtho[2,1,8-defg:2',l',8'-ijkl]pentaphene (10). The structures are shown in Figure 1. When the number of aromatic carbons, rather than the number of tings, is considered, in many cases only one isomer was found. Phenanthro[5,4,3,2-efghi]perylene and naphtho[8,1,2abc]coronene were newly discovered PAHs.

265

Figure 1. Naphthalene zigzag path of 1-ring additions. These PAHs could be arranged so that a series of successive two- and four-carbon additions would result in all these species being produced through a buildup of rings. The additions are not random, but are determined by two simple structural features. If a PAH had a three-sided bay (concave) region (such as phenanthrene or benzo[e]pyrene), a two carbon addition closed that area (e.g., going from phenanthrene to pyrene). This is conceptually similar to a Diels-Alder reaction. If there is no bay region in the PAH structure, a four-carbon addition occurs at the carbon-carbon bond of highest n electron localization (pyrene going to benzo[e]pyrene). This series of two- and four-carbon addition as a production route is identical to those PAHs found theoretically by Stein to be the energetically most favored [8]. This sequence has been called the "naphthalene zigzag" because several of the species are fused naphthalene structures arranged in a zigzag pattern [9].

266 At one point in this 1-ring buildup there are three possible isomers (Figure 2). The 7-ring dibenzo[e,ghi]perylene has three nonequivalent bay regions. A two-carbon addition to each of these will produce three different 8-ring PAHs. Two of these were previously known before this work,benzo[a]coronene and benzo[pqr]naphtho[8,1,2-bcd]perylene. The third was not, but a unknown compound was found which almost co-eluted in the HPLC with benzo[a]coronene. This was isolated and shown to be the third isomer [5]. This can be taken as a tacit proof of the 1-ring buildup path. Under higher-temperature operation, some other PAHs of 7 and 9 rings were found. The 7-ring PAHs were the isomers dibenzo[cd,lm]perylene and naphtho[8,1,2-bcd]perylene,

),

Benzo [ghi] Perylene

jr

/ Dibenzo [e,ghi] Perylene

Benzo [pqr] Naptho [8,1,2-bcd] Perylene

Benzo [a] Coronene

I '"

Phenanthro [5,4,3,2-efghi] Perylene

I~

Naptho [8,1,2-abc] Coronene

Figure 2. Detailed portion of 1-ring buildup showing 8-ring isomerization.

267 with a molecular weight of 326 daltons. The former usually found at about 100 times greater abundance. The 9-ring PAHs were the isomeric pair dinaphtho[2,1,8,7-defg:2',l',8',7'opqr]pentacene and dinaphtho[2,1,8,7-defg:2',l',8',7'-ijkl]pentaphene, with a molecular weight of 400 daltons. The former compound is usually produced at three to five times the amount of the latter PAH. Additionally, naphthenic (fused cycloalkyl) substituted species of the two 7-ring PAHs were found, and upon isolation mass spectrometry showed these had molecular weights of 366 and 406 daltons. A reaction that explains the occurrence of this second set of PAHs is the condensation of two pyrenes through the formation of a bridging ring (the Scholl condensation of pyrene). This reaction yields the two 9-ring PAHs, and subsequent hydrogenation and cracking gives the other PAHs. This reaction path is shown in Figure 3. The residue from dichloromethane extraction of hydrocracker deposits is a reddish powder. This material was found to be sparingly soluble in 1,2,4-trichlorobenzene (TCB). A mass spectrum of the TCB extract showed a major ion at 596 daltons, with peaks at 610, 624, and 638 daltons due to alkylation of this first species. Smaller peaks were seen at 620,

+H2 2

"-

+

~,

M W = 4O0

--t-

'---

MW = 406

-.I-

MW = 366

-t-

MW = 326

Figure 3. Scholl condensation reaction of pyrene, and subsequent hydrogenation and cracking to smaller PAHs.

268 644, and 694 daltons. The fluorescence spectra of the extract matched a known 15-ring PAH, benzo[ 1,2,3-abc:4,5,6-a' ,b',c']dicoronene (trivially known as dicoronylene) [10]. This molecule is the product of a Scholl condensation of two coronene molecules. The species at 694 daltons is due to the condensation of a coronene and an ovalene molecule, and is the largest PAH ever reported. It has a fluorescence excitation maximum at 545 nm, and is a purple colored compound. It was a newly discovered PAH, and is the largest ever reported. These structures of these reactions are shown in Figure 4. The other two masses at 620 and 644 daltons are due to PAHs produced when the two-carbon addition at bay regions occurs for dicoronylene. These two compounds [ 11 ] have been synthesized, and a similar fluorescence spectrum has been seen in those samples showing the 620 daltons species. The Scholl reactions of coronene and ovalene, and the subsequent products from two carbon addition are shown in Figure 5. 4. R O U T I N E ANALYSES Dicoronylene is the predominant molecule that causes process pipe plugging [2,3]. Its low solubilities in process oils are only on the order of a few parts-per-million, so saturation is reached very early in the process run. This PAH is so large that it is not volatile enough for gas chromatographic analysis and too absorbing for HPLC analysis. The low solubility requires a very sensitive analysis. Synchronous-scanning fluorescence (SSF) relies on the inherent spectra behavior of some PAHs [4]. In many PAHs (those with a strong 0 to 0 transition) the highest wavelength excitation band is only a few nanometers lower than the lowest-wavelength emission band. These two bands are also the most intense in each of their respective spectra. This difference

Coronene SchSII Condensation Dicoronylene

Coronylenovalene

Ovalene Figure 4. Scholl condensation of coronene and ovalene.

269

"-,2

I

Figure 5. The two large PAHs resulting from two-carbon additions to dicoronylene. in wavelength is known as the Stokes' shift. An SSF spectrum is collected by scanning simultaneously through both excitation and emission monochromators, with the emission monochromator offset from the excitation monochromator by the Stokes' shift. For a single PAH, only the spectral region immediately around these two bands has a response. At lower wavelengths there is excitation but not emission. At higher wavelengths there is emission but not excitation. A mixture shows discreet peaks for each PAH. The PAHs produced by the 1-ring additions have been found to not have strong 0 to 0 transitions, so they do not fluoresce with small Stokes' shifts [ 12]. They therefore do not interfere with any PAHs that do. This is as long as the solutions are dilute enough to avoid self-adsorption, which would cause the emitted light from the compounds of interest to be absorbed by the high concentrations of other molecules at these same wavelengths. The small number of PAHs produced by Scholl condensation (of pyrene, coronene, and ovalene) all have very strong 0 to 0 transitions, and so all have small Stokes' shifts and strong responses in SSF. All these structures differ significantly enough, however, so that the range of wavelengths of SSF spectral response for each PAH are different.

270 An SSF spectrum of a standard mixture of several of the PAHs found in a typical hydrocracker oil is shown in Figure 6, while that of an oil sample is shown in Figure 7. Quantitation is accomplished by generating a calibration curve from known concentrations of dicoronylene or by spiking samples through standard additions.

4 3

1

m

I 440

I

I

I

I

I

460

480

500

520

540

~

9

560

(nm)

Figure 6. Synchronous-scanning fluorescence spectrum of a mixture of standard PAHs.

m

I 440

I 460

I 480

I 500

, 520

I

I

540

560

(nm)

Figure 7. Synchronous-scanning fluorescence spectrum of a hydrocracker oil sample.

271 The HPLC methods cannot be used to routinely monitor the other PAHs because the instrumentation is rather sophisticated and data interpretation and analysis requires a high level of expertise. Instead simple UV absorbance methods are used to measure the absorbance at wavelengths characteristic of certain key PAHs. These include 305 nm for coronene, 335 nm for pyrene, 345 nm for dibenzo[cd,lm]perylene, and 495 nm for dinaphtho[2,1,8,7-defg:2',l'8',7'-opqr]pentacene. Tracking of each of these absorbances from process startup allows process engineers to follow the buildup of the PAHs and gain an idea of the chemistry going on within the process. 5. CONCLUSIONS The use of modern analytical methods has led to the determination of the PAHs which are produced in catalytic hydrocrackers. A variety of HPLC-DAD, fluorescence, and UV absorbance methods were developed to determine the occurrence of the PAHs. These PAHs result from a small number of reactions. These are either a new ring forming through twoor four-carbon addition or the condensation of pyrene, coronene, or ovalene. The latter reactions result in very large PAHs which cause process problems because of their low solubilities. Their production rates (and eventual precipitation in the process streams) can be monitored through the use of UV absorbance and fluorescence spectrometries. A synchronous-scanning fluorescence method was developed to monitor the production of dicoronylene during process operation. The results of these analyses can then be used to determine process performance.

REFERENCES 1. 2.

R.F. Sullivan, M. M. Boduszynski, and J. C. Fetzer, Energy Fuels, 3 (1989), 603. J.C. Fetzer and D. G. Lammel, "Hydrocracking Process With Polynuclear Aromatic Dimer Foulant Adsorption," U.S. Patent 5,190,6330 (1993). 3. J . C . Fetzer, J. M. Rosenbaum, R. W. Bachtel, D. R. Cash, and D. G. Lammel, "Hydrocracking Process With Polycyclic Aromatic Dimer Removal," U.S. Patent 5,232,577 (1993). 4. J . C . Fetzer, Polycyclic Arom. Cmpds., 4 (1994), 19. 5. J.C. Fetzer and W. R. Biggs, Polycyclic Arom. Cmpds., 5 (1994), 195. 6. E. Clar, "The Aromatic Sextet," Wiley-Interscience, New York (1972). 7. J.C. Fetzer and W. R. B iggs, J. Chromatogr., 642 (1993), 319. 8. S.E. Stein, J. Phys. Chem., 82 (1986), 566. 9. E. Clar and O. Kuhn, Justus Liebigs Ann. Chem., 601 (1956), 181. 10. L. Boente, Brenstoff Chemie, 36 (1955), 210. 11. M. Zander and W. Friedrichsen, Chem. Zeitung., 115 (1991), 360. 12. W.E. Acree, S. A. Tucker, and J. C. Fetzer, Polycyclic Arom. Cmpds., 2 (1991), 75.

This Page Intentionally Left Blank

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All rights reserved.

FOULING

MECHANISMS

AND EFFECT OF PROCESS

273

CONDITIONS

ON

DEPOSIT FORMATION IN H-OIL EQUIPMENT

Maurice A. Bannayan,a Harald K. Lemke, b and W. Kirk Stephensonb

aHusky Oil, Lock Box 1710, Lloydminster, Sask., Canada $9 V 1M6 bNalco/Exxon Energy Chemicals, RPC Chemicals, POB 87, Sugar Land, Texas, U.S.A. 77487-0087 ABSTRACT The H-Oil process is a high pressure, high temperature hydrocracking process, which uses an ebullated bed of catalyst to convert lower value heavy oils into upgraded higher value products. Deposit formation in the equipment downstream of the H-Oil reactor and high sediment accumulation in heavy fuel oil product streams are confining factors in current attempts to maximize H-Oil unit conversion. Analyses of deposit and stream samples from a commercial H-Oil unit indicate that several mechanisms influence fouling. The rejection of "vanadium- and nickel sulfides" from the catalyst and the precipitation of polycyclic aromatics appear to contribute to fouling in the reactor, reactor outlet line, and high pressure separator. "Asphaltene precipitation" is the prevalent fouling mechanism in the H-Oil vacuum tower. The primary objective of the current study is to investigate the mechanisms which lead to deposit formation in the reactor recycle cup and vacuum tower of a commercial H-Oil unit. 1. INTRODUCTION In many refineries thermal cracking processes are used to convert residues into lighter products. Low value petroleum coke is a product from the more severe cracking processes. The H-Oil process made it possible to convert the asphaltenic carbonizable portion of the residue to higher value liquid products rather than coke. In the H-Oil process an ebullated bed of catalyst is used to convert lower value heavy oil into upgraded higher value products in the presence of hydrogen. The ebullated bed reactor is an expanded bed of catalyst maintained in constant motion by the upward flow of liquid. The reactor behaves as a well mixed continuously stirred tank reactor. The catalyst activity in the reactor is maintained at a constant level by the daily addition of fresh catalyst and withdrawal of an equivalent amount of catalyst from the reactor. A flow diagram of the H-Oil unit is shown in Figure 1. H-Oil units have been operated successfully at residue conversion levels above 60%. However, at the higher conversion levels there is increased fouling, sedimentation problems in the operating units.

274 (ON(~

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MAKE-UP HYDROGCN

~--I.

LEAN AMINE

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,~1

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.

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Figure 1. Simplified H-Oil flow diagram. These fouling phenomena are believed to be due to the precipitation of asphaltenes from hydrocracked, effluent streams. Fragmentation reactions decrease the solubility power of effluent maltenes and the solubility of asphaltene micelles, thus facilitating precipitation [ 1]. The concentration of polyaromatic hydrocarbons (PAHs) might also contribute to fouling. This could happen when the formation rate ofPAHs (through dehydrocyclization)is greater than the hydrocracking rate of the process stream. In particular, PAH concentration occurs with fixed bed catalysts when the small pore size limits the access of large PAHs to the catalytic sites [2]. Phase separations due to supercritical conditions in hydroprocessing equipment may also facilitate deposit formation [3]. Also of interest are the influences of certain operating and design factors on deposit formation and prevention. These include the use of middle distillates (e.g., FCC light cycle oil), fresh feed recycling, proper separator quenches, downstream vessels with conical bottoms, and antifoulant programs. 2. EXPERIMENTAL An atmospheric residue was hydrocracked in a commercial H-Oil unit at about 440~ and 18,000 kPa. Figure 1 shows the processing sequence. Analyses were performed on deposits from the recycle cup, reactor outlet line, and vacuum tower trays; and on samples of the

275 atmospheric residue (i.e., H-Oil feed), vacuum tower charge (i.e., cracked atmospheric residue) and residue, and two vacuum tower side streams. The loss of ignition was determined by grinding the deposit to 200 mesh, drying it overnight in a programmable muffle furnace (Fisher), and burning the sample at 800 ~ for 3 hr. For X-ray determinations, the samples were ground and dried prior to analyses with a KEVEX Analyst 770. The amount of resins and asphaltenes was determined with an Iatroscan (Mark 5) instrument. The GC/MS system is from Waters (Waters/Extrel-ELQ 400), E.I. mode, scan rate 500 amu/s (range 33-600 amu.).LC/UVsystem is from Waters (model 600E)and uses a synchrom column (100 mm-4.6 mm). The flocculationpoint apparatus consists of a transmittance detector and an automatic titrator. The atmospheric residues are diluted in an aromatic solvent and titrated with n-pentane. The onset of the flocculation is indicated by a decrease in transmittance. NMR data were recorded with a Varian Unity 300 (13C-NMRdoped with Cr(AcAc)3, 2s delay, gated decoupling, 1000 repetitions). The microscope used to observe sedimentation coke formation is equipped with a video camera so that images may be recorded as the samples are being observed. Figure 6 shows two miscroscopic pictures of HOil atmospheric and vacuum bottoms. It is noticeable that the amount of sediments in the fractionator bottoms is higher than the vacuum bottoms. Hot stage microscopy [4] is now becoming a very reliable method of observing the coking process for heavy hydrocarbon samples at elevated temperature and pressure. A number of phenomena may be observed, of which the formation of mesophase is of most interest. 3. RESULTS AND DISCUSSION

3.1. Fouling in the H-Oil Reactor Recycle Cup Deposit formation in the reactor recycle cup is undesired due to potential plugging of catalyst distribution, which could lead to an uneven distribution of hydrogen in the ebullated catalyst bed. Analyses of recycle cup deposit and reactor effluent samples provide insight as to the mechanism(s) that contributes to recycle cup fouling.

3.1.1. Recycle Cup Deposit Analyses Loss ON IGNITION. The loss on ignition value of 53% (by weight) indicates that about half of the deposit is composed of inorganic material. X-RAY. X-ray analysis data, summarized in Table 1, shows that the deposit contains 20.8 % nickel (Ni) plus vanadium(V), and 8.8 % sulfur (S). Interestingly, the data reveals only 2.5% of each of the primary catalyst components aluminum (A1) and molybdenum (Mo); hence, it is unlikely that carryover of the catalyst contributes significantly to the accumulation of V and Ni. Furthermore, it appears unlikely that major amounts of V and Ni derive from the reactor metallurgy, as the deposit contains only 1.7% iron (Fe). The rejection of"vanadium-and nickel-sulfide"[5] from the catalyst surface may account for the high amounts of V, Ni, and S. CHROMATOGRAPHY. Materials extracted from the recycle cup deposit using a proprietary technique provide insight as to the organic molecules that contribute to fouling. GC/MS and LC/UV spectroscopy were used to analyze the extracts.

276 Table 1. Elemental analysis of recycle cup deposit (weight %).

Recycle Cup 12.0 8.8 2.5 2.5 1.7 8.8

V Ni AI Mo Fe S

Relative

t

[

~

Deposit a

Abundan

. . . . .

1OO

-A-L

.

200

300

400

500

600

.

700

.

.

flOO

900

.

.

1000

1100

i

1200

!200

1400

Scan Number Relative Abundance

100

~oo

Deposit b

~oo

,oo

=_oo ~oo

Too

Boo

900

~ooo ~1oo 12oo ~3oo : , o o Scan Number

Figure 2. GC/MS n-alkane distribution of extracts from (a) reactor recycle cup and (b) reactor outlet line.

Figure 2a shows the n-alkane distribution of the extract, as determined by GC/MS. In contrast to the normal n-alkane distribution observed for the H-Oil feed, the extract exhibits a bimodal distribution, with abundance maxima at heptadecane (C17)and hexacosane (C26) molecular weights. Figure 2b shows similar data for an extract from a reactor outlet line deposit. The n-alkane distribution corresponding to the outlet line deposit has an even more pronounced bimodal pattern than that from the recycle cup deposit. The abundance maxima correspond to shorter chain lengths. The bimodal distribution of the n-alkanes may reflect a preferential cracking behavior o f t he catalyst. (Planned investigations will help evaluate whether such "fingerprints" support the idea of a phase separation [3,6] in a hydrocracker).

277

Relative Abundance

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

t ~

_

Effluent

~..-.v'~%~'~ ~ - ~ ' H - O i l __~.,~"

Feed

Retention Time

Figure 3. LC/UV spectra at 254 nm for H-Oil feed and effluent. A comparison of LC/UV spectra for (uncracked) H-Oil feed versus (cracked)H-Oil effluent shown in Figure 3 reveals that small- and large-ring polycyclic aromatics concentrate in the latter. One particular four ring polycyclic aromatic compound is highly enriched. The formation of polycyclic aromatics during hydrocracking can occur at high temperatures where dehydrogenation reactions are favored [1 ]. 3.1.2. Reactor Effluent Analyses

FLOCCULATION POINT. Figure 4 shows that the flocculation point of a reactor effluent sample (cracked atmospheric bottom) is much lower than that of the reactor feed. Generally, the flocculation point decreases as the amount of solids in the sample increases. This has been confirmed by the Shell Hot Filtration Test IP/375/ASTM 4870 procedure.

H-Oil Feed

3.5

r c c~

2.5

E

2

C L_

p-

1.5

- Effluent 0.5

o:1 2.14

6.22

t0.3

14.311

10.44

22.54

26.62

mL OF PRECIPITANT Figure 4. Flocculation curves for H-Oil feed and effluent.

2K).?

278

Relative Abundance Effluent

It

\, H-Oil \

/

20,000

u

/

Feed

800

u

\

'\

Relative Mol. Weight

Figure 5. GPC spectra of H-Oil feed and effluent.

1H-NMR AND 13C-NMR.To investigate the flocculation point decrease on a molecular level, 1H-NMR and 13C-NMR techniques[3] were employed. NMR spectra do not indicate significant molecular differences between the streams. (NMR investigations are in process to access cracking-induced molecular alterations of the asphaltenes.) CHROMATOGRAPHY. On the basis of latroscan measurements, cracking appears to decrease the resin/asphaltene ratio. This result confirms the observed destabilization of the stream. Figure 5 displays gel permeation chromatography (GPC) apparent molecular size distribution data for both streams. In contrast to reported results[3], GPC data indicate that the components of the highest molecular weight fraction crack quantitatively, lowering the maximum apparent molecular weight from 23,500 u to 4500 u. Molecules of apparent molecular weight 800 u are most abundant in the H-Oil feed. The hydrocracked effluent has a bimodal molecular weight distribution with abundance maxima at 800 u and 520 u. The bimodal distribution indicates that the applied H-Oil catalyst preferentially cracks molecules with an apparent molecular weight of about 650 u rather than 800 u, the most abundant apparent molecular weight of the H-Oil feed.

3.2. Vacuum Tower Fouling A fouling-induced pressure increase in the effluent tower leads to decreased profits due to a lower HVGO recovery rate. The prevention of such fouling is therefore of considerable interest.

279

Figure 6. Microscopic pictures (500X) of >5 m solids for (a) Vacuum tower charge and (b) vacuum tower bottom. The fouling rate of the tower correlates with the SHFT and microscopically determined solids content in the vacuum tower charge. Figure 1 shows the areas where liquid and solid samples were taken starting in the feed to the reactor, the bottom of the hot low pressure separators, the feed to the atmospheric tower, the feed to the vacuum tower, the vacuum bottoms and the wash oil section. Visual inspection of the tower revealed that most of the fouling occurs on the trays slightly above the vacuum tower inlet and not at the shed decks of the conical reactor bottoms. Microscopic measurements are consistent with the finding of a clean reactor bottom by revealing that the vacuum bottom contains fewer particulates than the charge to the vacuum tower (see Figures 6a and 6b). 3.2.1. Stream Analyses 13C-NA/[R. Table 2 shows 13C-NMR data [7] for liquid samples from the bottoms of the atmospheric and vacuum towers, and for side streams from and above the most seriously fouled vacuum tower trays. The data show the vacuum bottom to contain 12 % more aromatics than the atmospheric bottom, while the amount of paraffins is 6 % lower. This indicates that the vacuum residue has a greater "solubilizing power" than the atmospheric tower residue, which possibly accounts for the decreased amount of microscopic particulates in the vacuum bottom. In contrast, the side stream from the severely fouled trays contains only 25 % aromatics and 47 % aliphatics, suggesting a relatively low "solubilizing power" for the local reflux.

280 Table 2. Group distribution of vacuum tower streams determined by 13C-NMR [7]. Tray (above foul) Tray (foul) Atm. Residue Vacuum Residue

Aromatics

Aliphatics

Naphthenes

31 25 32 44

49 47 43 37

20 28 25 19

3.2.2. Vacuum Tower Deposit Analyses GC/MS analyses give evidence that some heavy ends material entering the vacuum tower rises through entrainment and impacts fouling on the upper trays. X-ray and combustion analyses reveal the deposit to be primarily organic: the LOI value is 96 %; S and N contents are low at 3 . 1 % and 1.8 %, respectively; V and Ni values are only 0.9 % and 0.5 %; and the Fe content is negligible. These data are typical for deposits caused by asphaltene precipitation. Tower deposition appears stratified, indicating localized fouling. One possible explanation is the preferential precipitation of the heavy ends asphaltenes facilitated by the low "solubilizing power" of the local reflux. The reduced fouling effects observed for the higher trays support this hypothesis. Side streams from the less fouled trays contain considerably more aromatics and naphthenes than those drawn from the more fouled trays. 4. CONCLUSIONS H-Oil cracking appears to significantly increase the readiness of an atmospheric residue to precipitate asphaltenes. The increased readiness can be explained on a molecular and colloidal level. It appears likely that reactor fouling is related to the rejection of "vanadium- and nickelsulfides" from the catalyst surface. The relatively high aromaticity of the organic extract from the reactor deposit (versus the H-Oil feed) may be due to excessive dehydrocyclization at the expense of naphthene and naphtheno-aromatic hydrocracking reactions. The high reaction temperatures employed in high conversion H-Oil units favor such reactions. The occurrence of an effluent molecular weight distribution with two major peaks supports the idea of insufficient hydrocracking of molecules around 800 u. Reactor deposit analyses do not reflect compositions consistent with asphaltene deposition: the V, Ni, and S values are high., and the LOI value is low. The current authors doubt that the increased readiness of the effluent to precipitate asphaltenes is of relevance at reactor conditions. In contrast, asphaltene deposition seems to be a major contributor to vacuum tower fouling in the H-Oil unit. REFERENCES 1. I. Mochida, X. Zhoa, and K Sakanishi, Ind. Eng. Chem. Res., 29 (1990) 2324. 2. R. F. Sullivan, M. M. Boduszynski, and J. Fetzer, Energy and Fuels, 3 (1989) 603.

281 3. M. Ternan, P. M. Rahimi, D. M. Clugston, and H. D. Dettman, Energy and Fuels, 8 (1994) 518. 4. P.L. Sears, "Hot Stage Microscopy," Division Report ERL93-03 (CF), ANCMET, Energy Mines and Resources Canada, (1993). 5. S. Asaoka, S Nakata, Y. Shiroto, and C. Takeuchi, in Metal Complexes in Fossil Fuels (ACS Symposium Series 344), American Chemical Society, Washington, D.C., 275, (1987). 6. J. M. Shaw, R. P. Gaikwad, and D. A. Stowe, Fuel, 67(11) (1988) 1554. 7. L. G. Galya and D. C. Young, American Chemical Society, Div. of Petr. Chem, 28(4) (1983) 1316.

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Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All rights reserved.

283

BED EXPANSION AND PRODUCT SLATE P R E D I C T I O N S OF H-OIL P R O C E S S VIA NEURAL N E T W O R K M O D E L L I N G E. K. T. K a m a, M. M. AI-Mashan b and H. Dashti a

apetroleum Technology Department, Kuwait Institute for Scientific Research, P.O.Box 24885, 13109, Safat, KUWAIT bKuwait National Petroleum Company, Shuaiba Refinery, KUWAIT ABSTRACT The H-Oil process is a very complex unit used to upgrade vacuum residues to lighter product by catalytic hydrotreating and hydrocracking at reasonably high pressure and temperature. It is an important operation at the Shuaiba Refinery of Kuwait National Petroleum Company (KNPC). Despite considerable research conducted by its licenser and other organizations, the technology is still not well understood. However, this can be enhanced through the development of reliable models which can predict the hydrodynamics and product slate due to changes in process or operation parameters. This contribution demonstrates how neural network can be used as an operational support tool for H-Oil process by providing the insight into the multiphase bed behaviour and predicting a complete picture of the product slate. The ebullated-bed behaviour is characterised by a hydraulic model expressed in term of ebullated rate, catalyst loading, catalyst addition rate, and gas and liquid feed rates. The reaction model is used to predict the product slate covering naphtha, kerosene, light diesel, fractionator bottoms and 975~ + conversion. 1. I N T R O D U C T I O N The H-Oil units operated in Shuaiba Refinery, KNPC since 1968 are single stage processes which hydrotreat and hydrocrack vacuum residues in an ebullated-bed reactor (EBR) in which hydrogen, heavy oil feed stock and catalyst particles are brought together in constant contact. The feed is constituted of a mixture of very large molecule hydrocarbons, metals, sulphur and many other undesirable materials. The catalysts are suspended in the heavy oil and are ebullated by a constant flow of liquid and gas from the bottom of the reactor. This creates an homogeneous environment for to hydrotreat the undesirable materials and hydrocrack the heavy feed stock. Lighter and more environmental friendly products are separated and taken to a separation and purification train. The unconverted heavy oils are recirculated back to the reactor with a small amount of diluent to improve the efficiency of overall conversion and selectivity [1 ]. Due to the process complexity and inadequate understanding of the technology, no reliable models are readily available and capable to postulate the physico-chemical phenomena of the H-Oil process covering the reaction kinetics, heat and mass transfers and ebullation hydrodynamics of catalyst, oil and hydrogen gas. In an attempt to capitalize the benefits from the previous joint efforts between KNPC and KISR through the PET-47 project [2], it will be advantageous to utilize the captured plant operation and laboratory research data

284 to develop models to provide insights to the hydrodynamic behaviours and to predict the product slate specific to the Shuaiba Refinery H-Oil units. To model this reactor, consideration of hydrodynamics, reaction kinetics, mass transfer and heat transfer properties of the system is essential. Although a number of studies have been made to determine these parameters [3-5], the findings are neither comprehensive nor coherent. Moreover, other parameters such as pressure drop, flow regimes, bubble characteristics, catalyst stratification and attrition and interphase mixing complicate the system further. It is important to recognize that any H-Oil process models must be capable of multiple inputs and outputs, and flexible in model updating due to the complexity of the unit. All these require extensive experimental investigation to develop reliable correlations which can then be coupled to the deterministic model equations describing the hydrodynamics, heat and mass transfer of the entire unit. Similar procedures in stochastic model development are also expected in terms of time and efforts. The best way is to use artificial neural networks (ANN) as the modelling tool because the neural network models are initially trained with historical plant, laboratory or published data to give a preliminary set of weights which can be progressively refined as new information becomes available. Since the simulated results are subjected to match with the data used in the training to give the best fit, this leads to improve accuracy and extend the range over which the models were originally intended. 1.1 Neural network modelling

Neural networks are computing system composed of many simple computational elements (neurons) locally interacting across very narrow bandwidth channels (connections), which process information by their dynamic response to external inputs. Computations are collectively performed by the entire network with knowledge represented as distributed patterns of activity over all processing elements. A schematic representation of a general neural network architecture is shown in Figure 1. The circles, squares and rectangles representing the neurons are arranged in three layers identified as input, hidden and output respectively. A bias node, encircled by an ellipse, is used to stabilize the evaluation process. The lines provide the connections between neurons. The input layer obtains external information in the form of input patterns from which predictions will be based. The information flows from input to output layers through the hidden layer. The input to each neuron in the hidden layer is a weighted sum of the outputs of all the input neurons because each neuron is weighted with respect to information (signal strength). The outputs are passed onto the neurons in the output layer and are also the weighted sum of all the hidden layer outputs. The collective operations result in a high degree of parallelism which enable the network model to solve complex problems rapidly. It is well equipped to solve complicated, ill-structured problems with multiple input-output nonlinear structures. Some models have been developed for complex process simulation such as Fluid Catalytic Crackers [6,7], and fractionators [8]. 2. NEURAL NETWORK MODELS The EBR is one of the novel multiphase catalytic reactors which can be regarded as intermediate to slurry and fixed-bed operations. It is usually operated co-currently where the

285

Figure 1. A schematic of a three-layer ANN with 8 input, 7 hidden and 1 output nodes.

catalyst particles are fluidized mainly by liquid, and the gas flows in the form of discrete bubbles. 2.1. Ebullated-bed expansion model

Ebullated-bed reactor, Figure 2, is the main power house for the H-Oil process. To suspend the catalyst particles, which are lager than those normally used in fluidized-bed contacting units, an internal recycle of liquid by an ebullation pump is needed. The mixing and mixture provide excellent heat transfer properties. The ability for catalyst replacement during processing is particularly important in hydrodemetallation and hydrodesulphurization of heavy residues because of the rapid catalyst deactivation. The advantages of employing an EBR can be found in several areas: 1. well mixed between the different phases, 2. thermal stable and near isothermal operation, 3. continuous catalyst replacement without interrupting the operation, and 4. constant pressure drop. These are the essential conditions for a catalytic reactor to provide consistent activity, reactivity, stability and selectivity and they are direct or indirectly affected by the catalyst bed

286 expansion. The level of interface between multiphase regime and freeboard region is believed to account for the extend of catalytic hydrotreating and hydrocracking reactions and that of the thermal cracking. It is important to maintain this interface at a desired level for optimum conversion and safely operation. The interface is controlled by the ebullation rate, feed temperature, catalyst loading, process pressure and feed flowrates. An increase in the bed expansion can effectively increase the residence time for hydrotreating and hydrocracking reactions. However, this will greatly affect the heat generation, as well as the separation processes between the phases. Consequently, the reaction severity, the gas/liquid product distribution and gas/liquid/solid entrainment in ebullation can result instability in process operation and violation in product specifications. Furthermore, the entrainment of gas or solid can affect the ebullation pump differently, but both eventually will lead to the bed slump [9] and even process shut down. The hydrodynamic interactions are difficult to be described quantitatively because of lacking information in the particle velocity, catalyst stratification, bubble size, phase holdups and feed and product properties. Although some correlations have been developed [10-13], most are applicable to aqueous and non-porous particle systems or under cold flow conditions. Generally, they cannot be applied directly to commercial operations. Several expressions correlating bed expansions with parameters such as liquid and gas flowrates, liquid density and viscosity, catalyst loading, catalyst size (length and diameter), and catalyst skeletal density have been developed for the KNPC Shuaiba H-Oil Units [14]. An ANN model as shown in Figure 1 is developed to simplify the above development but still covering all the parameters used previously. The neural network architecture consists three neural layers. The eight specified input nodes to predict the expanded bed height as the output node through nine hidden nodes. All input and output values obtained from the KNPC H-Oil Process Operation Manual [14], are normalized to synchronize their magnitudes to minimize the number of training events and errors. The initial number of hidden nodes can be determined by the following expression, H = 2 * 4I + o

(1)

where H, ! and O are the number of hidden, input and output nodes respectively. The reliability of the network model can be assessed by comparing the predictions (recalls) with measurements expressed in terms of the absolute relative deviation (ARD%) and absolute average deviation (AAD%) as,

(ARD%)i =

(~[Yexpt. -Ypred j)i J O * 100% [2 (Yexpt)j]i J

AAD% = 1 ~ ( A R D % ) i

n i--1

i

= lton

(2)

(3)

where y represents the concerned parameter, i is the ith data set, j is the jth output node, and n is the total number of learning data sets (patterns).

287 Figure 3 shows the comparison of the normalized bed height from the H-Oil reactor data and ANN model predicted values after two millions training events. The maximum ARD% is 13.8% with an AAD% of 1.92% for the 85 sets of input data employed. However, if only data with bed height values below the allowable upper level are considered, the max. ARD% and AAD% reduce to 10.7% and 1.43% respectively. It is clearly demonstrated the predicted results from the ANN ebullated-bed expansion model are very close to the literature values. This model by no mean limits its applications just to predict the interface level. It can be extended to cover heat generation in terms of exotherms, spread temperature and/or catalyst average temperature (CAT) from data recorded in the technical report [15]. 2.2. H-Oil product slate model

Due to its flexible nature, the same ANN architecture can be applied in which the input and output parameters based on the H-Oil units are given in Table 1. There are twenty input nodes, ten from each train (A and B). They cover the fresh and recycle liquid feed rates, feed API, makeup and recycle gas rates, hydrogen partial pressure, reactor pressure, exotherm, reactor average temperature and catalyst addition rate. The model also consists nine hidden and six output nodes. However, the number of nodes in each layer or the inclusion of which process parameters can be modified as occasion arises, such as changes in process requirement, model refinement or data availability. This is a multiple input and output problem where there are 20 input and 5 output nodes. The simulated results of the product slate covering hydrogen consumption, naphtha, light kerosene, light diesel, fractionator bottoms and 975~ + conversion are given in Table 2. The AAD%'s for the five set of test run data used to train the model are ranging between 0.292 to 0.56 % which results the mean AAD% less than 0.5 %. The maximum ARD% which appears in the light diesel of the fourth test run data set, is 1.843 %. The reliability of this model is clearly shown. The model is then applied to predict the output patterns from some recent shift and start-up data, and the results are shown in Table 3. The mean AAD%'s of the shift and start-up are far too high because of the difference in the process status, liquid feed used and human factors in determining some process values. Although these have not been covered in the input patterns, they can be overcome by including either the additional features in data sets such as feed properties, or additional plant data. After re-training the model with both the test run and shift data, the AAD% improvement in data prediction is tremendous, the trained shift data from 134.7% to 1.6% and the untrained start-up data from 137.2% to 9.64%. They can outweigh the minimal drop in AAD% of the test runs. Further improvement can be realized by training all the test, shift and start-up runs data together, and the resulting AAD%'s are 0.43, 0.82 and 1.64% respectively. The usefulness of the of artificial neural networks as a modelling tool is apparent. A more general H-Oil product slate model can be developed by including the feed and catalyst properties. It can also easily be adapted to model the other aspects of the H-Oil process such the hydrotreating and hydrocracking reaction kinetics or coke lay down tendency in the separation units with the appropriate input and output patterns.

288

Figure 2. A typical ebullated-bed reactor.

1.0 0.9 ,s=

0.8

i

:IZ

Allowable upp '

0.7J l

0.6' i

Z

0.5

i

0.4 0.4

0.5

0.6

0.7

0.8

0.9

Normalized Bed Height Published values

. Predicted values

Figure 3. Comparison of measured and predicted bed heights.

1.0

289

Table 1 Input and Output Parameters ANN Layer

Process Parameters

Engineering Units

Operation Range (Normalized)

Input

Liquid Feed Flowrate 'A' Feed API 'A' Gas Oil Recycle Rate 'A' Make-up Hydrogen Flowrate 'A' Recycle Gas Flowrate 'A' Reactor Average Temperature 'A' Reactor Pressure 'A' Actual Exotherm 'A' Catalyst Additional Rate 'A' Hydrogen Partial Pressure 'A'

BPD oAPI BPD MMSCFD MMSCFD OF

O- 1.00 0 - 0.90 O- 1.00 0 - 0.90 0 - 0.90 O- 1.00 O- 1.00 0 - 0.90 0-0.50 0 - 1.00

Liquid Feed Flowrate 'B' Feed API 'B' Gas Oil Recycle Rate 'B' Make-up Hydrogen Flowrate 'B' Recycle Gas Flowrate 'B' Reactor Average Temperature 'B' Reactor Pressure 'B' Actual Exotherm 'B' Catalyst Additional Rate 'B' Hydrogen Partial Pressure 'B'

BPD oAPI BPD MMSCFD MMSCFD OF

Hidden

9 Nodes

Output

Hydrogen Consumption 'A' Naphtha Light Kerosene Light Diesel Fractionator Bottoms 975~ + Conversion

PSIG OF LBS/BBL PSIG

PSIG OF LBS/BBL PSIG

SCF/BBL BPD BPD BPD BPD vol %

0 - 1.00 0 - 0.90 0 - 1.00 0 - 0.90 0 - 0.90 0 - 1.00 0 - 1.00 0 - 0.90 0-0.50 0 - 1.00

0.00 0.60 0.40 0.100.100.60 -

0.85 0.90 0.90 1.00 1.00 0.85

290 Table 2 Comparison of Values from Test Run and the ANN Model. Test Run

Product Slate

Plant Data

Prediction

ARD%

AAD%

Run 1

Hydrogen Consumption (SCF/BBL) Naphtha (BPD) Kerosene (BPD) Light Diesel (BPD) Factionator Bottoms (BPD) 975+F Conversion (vol %)

798.4490 3570.000 7440.300 4136.900 1596.000 48.715

803.551 3573.522 7481.493 4148.421 1612.584 48.925

0.634 0.099 0.554 0.278 1.039 0.427

0.505

Hydrogen Consumption (SCF/BBL) Naphtha (BPD) Kerosene (BPD) Light Diesel (BPD) Factionator Bottoms (BPD) 975+F Conversion (vol %)

801.975 3499.800 7200.000 4100.310 2500.000 51.500

799.311 3507.144 7240.554 4100.664 2475.152 51.498

0.332 0.209 0.563 0.016 0.994 0.003

0.353

Hydrogen Consumption (SCF/BBL) Naphtha (BPD) Kerosene (BPD) Light Diesel (BPD) Factionator Bottoms (BPD) 975+F Conversion (vol %)

744.000 3150.000 4200.300 8069.620 3420.000 50.499

747.552 3149.526 4229.154 8069.292 3428.344 50.678

0.477 0.015 0.687 0.004 0.244 0.356

0.297

Hydrogen Consumption (SCF/BBL) Naphtha (BPD) Kerosene (BPD) Light Diesel (BPD) Factionator Bottoms (BPD) 975+F Conversion (vol %)

669.035 4095.000 6419.700 1110.280 7239.200 56.050

672.319 4100.862 6453.720 1130.739 7241.720 56.229

0.491 0.143 0.530 1.843 0.035 0.320

0.560

Hydrogen Consumption (SCF/BBL) Naphtha (BPD) Kerosene (BPD) Light Diesel (BPD) Factionator Bottoms (BPD) 975+F Conversion (vol %)

642.005 3850.200 8420.400 4200.040 4500.000 60.301

644.473 3858.054 8433.999 4230.593 4499.904 60.464

0.384 0.204 0.162 0.727 0.002 0.272

0.292

Run 2

Run 3

Run 4

Run 5

Mean AAD% = 0.401 Maximum ARD% = 1.843

291

Table 3 AAD% Comparison of the ANN Model before and after Re-training.

Particular

AAD% from Model trained with only Test Run Data Test Run

RUN RUN RUN RUN RUN SHIFT SHIFT SHIFT SHIFT SHIFT SHIFT

1 2 3 4 5

Shift

0.505 0.353 0.297 0.560 0.292

A B C D E F

Test Run

Shift

57.088 191.730 196.175 93.306 201.963 67.953

0 335 2 968 1 895 0 703 2 829 0916 72.566 137.613 202.798

0.401

Start-up

0.948 0.525 0.284 0.841 0.795

START-UP o~ START-UP [3 START-UP Y Mean AAD %

Start-up

AAD% from Model re-trained with both Test Run and Shift Data

134.702

137.155

14.792 12.240 6.035 0.672

1.608

9.640

3. CONCLUSIONS The complexity of H-Oil operation devices mainly on the wide range of mixed heavy feed stock and employs the not well understood ebullation technology. This restricts any attempts in modelling the H-Oil process mechanistically. One alternative is to use artificial neural networks which is very suitable for any complex process or not well defined physical phenomenon because of its parallel processing capability. Hence it can handle truly multipleinput and -output problems. Once an ANN architecture is set, it can be applied to the areas of interest as demonstrated in H-Oil process modelling applications. From the ebullated-bed expansion model, the maximum ARD and mean AAD%'s are 10.7 and 1.43%'s respectively, while 1.962 and 0.759%'s are achieved by the H-Oil product slate model. Both models adopt the same 3-layer ANN architecture using back-propagation but with corresponding input and output patterns as well as the adequate number of hidden nodes. It does not require to redevelop equations or re-program the software due to new applications. However, it will require some efforts to design an attractive user-interface to make the tool more user-friendly.

292 Furthermore, it can be modified to predict the preferable process parameters according to a specified product slate by an inversion procedure which is currently being investigated. The applications are not limited to the examples illustrated. It can handle other complicated situations or refining processes such as fluid catalytic cracking units and hydrocrackers. 4. NOTATION AAD% ARD% H I

i J n O Y

absolute average deviation % as defined in Eqn. (3) absolute relative deviation % as defined in Eqn. (2) number of hidden nodes number of input nodes the ith data set the jth output node total number of training data sets number of output nodes output parameter

REFERENCES

1. 2.

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

14. 15.

M. Embaby, Studies in surface science and catalysis - catalysts in petroleum refining, D.L. Trimm, S. Akashah, M. Absi-Halabi and A. Bishara (eds.), 53 (1989) 165. M. Absi-Halabi, E.K.T. Kam, A. Stanislaus, S.Y. Diab and F. Owaysi, H-Oil process and catalyst evaluation and development, Final Report, KISR 4014, Kuwait Institute for Science Research, Kuwait, 1992. Y.T. Shah, Gas-Liquid-Solid Reactor Design, McGraw Hill Book Co., New York, 1979. V.T. Sinha, M.S. Butensky and D. Hyman, Ind. Eng. Chem., Proc. Des. Devel., 25 (1986) 321. L.S. Fan, Gas-Liquid-Solid Fluidization Engineering. Butterworths, Boston, 1990. V. Venkatasubramanian and K. Chan., AIChEJ, 35 (1989) 1993. M.J. Bagajewicz and V. Manousiouthakis, AIChEJ, 38 (1992) 1769. C. McGreavy, M.L Lu, X.Z. Wang and E.K.T. Kam, Chem. Eng. Sci., 49 (1994) 4717. A. Li. and D. Lin, AIChE 1981 Annual Meeting, New Orleans, USA. (1981) Paper n 2d. J.M. Begovich and J.S. Watson, Fluidization, J.F. Richardson and D.L. Keairns (eds.), Cambridge University Press, England, 1978. Y. Kato, K. Uchida, K. Kago and S. Morooka, Powder Tech., 28 (1981) 173. W.D. Deckwer and A. Schumpe, German Chem. Eng., 77 (1984) 168. E.K.T. Kam, E. Alper, S. A1-Safadi, F. Abu-Seedo, M. Absi-Halabi and M. Sabri, Cold flow studies on catalyst properties and hydrodynamic characteristics in ebullated-bed model reactor, Technical Report KISR 4169, Kuwait Institute for Scientific Research, Kuwait, 1992. KNPC, H-Oil Operation Manual, Kuwait Nation Petroleum Company, Shuaiba Refinery, Kuwait, 1985. E.K.T. Kam, S. A1-Jadi, A.R. Meziou, M. Sabri and F. Abu-Seedo, H-Oil process monitoring II. Further data analysis and correlations development, Technical Report KISR 4105, Kuwait Institute for Scientific Research, Kuwait, 1992.

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All fights reserved.

293

R E N E W E D A T T E N T I O N TO THE EUREKA PROCESS: T H E R M A L C R A C K I N G P R O C E S S AND RELATED T E C H N O L O G I E S F O R RESIDUAL OIL U P G R A D I N G T. Takatsuka a, R. Watari a and H. Hayakawa b

a Chiyoda Corporation, Tsurumichuo 2-12-1, Tsurumi-ku, Yokohama 221, Japan b Fuji Oil Co. L TD, Ohtemachi, 1-2-3, Chiyoda-ku, Tokyo 100, Japan ABSTRACT The EUREKA process is a commercially proven thermal cracking process of semi-batch reactors system to produce valuable cracked oil and aromatic petroleum pitch from heavy residual material. The first unit was built in Sodegaura Refinery of Fuji Oil, Japan in 1975. The second unit was built in Nanjing, P. R. China in 1988. Both the units are now operated effectively for heavy oil upgrading. In this process, reaction proceeds at lowered hydrocarbon partial pressure by injecting steam into the reactor, keeping petroleum pitch in a homogeneous liquid state. A higher cracked oil yield is obtained by this reaction than that of Delayed Cokers. The cracked heavy oil aider hydrotreating is used as FCC or hydrocracker feedstock, feedstock for olefin production and blending stock for low sulfur fuel oil. The cracked light oil is blended with LGO and hydrodesulfurized to produce diesel fuel. The cracked naphtha is fed to catalytic reformer aRer hydrotreating. The accumulated experiences and further developments improved the latest design of the EUREKA process to a great extent. The conversion of the residual feed in the furnace is increased from ca. 25% to 35 % on total conversion of 70%. That enabled a wide selection of feedstocks as well as a reduction of utility consumption and improved operability. The residual product of the processes is not coke but pitch. Pitch is very easy and smooth to withdraw from the reactor, and it makes the plant site clean and compact compared to Coking plants, because pitch is handled as liquid in the process and product solid pitch is none dusty. The pitch is utilized as boiler fuel or gasification feedstock having an excellent quality. It has high heat of combustion and good characteristics for burning, comparable to fuel oil. It is also precious alternative source of coking coal for steel industry. The design tools developed for EUREKA process are very sophisticated, a) Practical Model of Thermal Cracking of Residual Oil and b) Tubular Fouling Model for Residue Cracking Furnaces made our designs very sure and reliable. 1. P R O C E S S F L O W Figure 1 shows a simplified flow diagram of the Eureka process. The feed, typically vacuum residue, is fed to the preheater and then enters the bottom of the fractionator, where it is mixed with the recycle oil. The mixture is pumped up to the charge heater and fed through an automatically operated switching valve to the reactor system which consists of a pair of reactors operating alternately. Since the switching valve operates typically every one and a half to two hours in this process, the cycle time is of the order of three to four hours. In the reactor, thermal cracking reaction takes place in the presence of injected superheated steam. The function of this steam is to strip the cracked products out of the reactor and supply a part of heat required for cracking reaction.

294 Fractionator ~ ~

Reactor Charge Heater

~

I

~

] ] ~ ~

~._'

.

'

] [ Light Oil

~5~ ~ ff Vacuum Preheater Residue

o

r

j ou, ~U~er Water [ Stripping Unit

Steam Super Heater Steam ~ 1

~Stabilizer ~" CrackedLight Oil ,.

~

,I

] Cracked Gas I[ Fuel Gas ' ] ] Sweetening [ -'(~[Unit ]

[ Pitch Flakerc~ ~

~ ~Stripped _ _ Water .

Cracked Heavy Od Petroleum Pitch

Figure 1. EUREKA Process Flow Diagram

At the end of the reaction, the bottom product is quenched and then allowed to blow down to a buffer drum. The pitch is sent to the pitch flaker by pump where liquid pitch is cooled and solidified to flakes. The cracked products - oil and gas from the top of the reactor and steam enter the lower section of the fractionator, where a small amount of entrained residue is washed and removed. The upper section is an ordinary fractionator, where the heavier fraction of cracked oil is drawn as a side stream. The lighter fraction of cracked oil is obtained from the overhead drum as normal practice. 2. DISTINGUISHED CHARACTERISTICS AND PROCESS IMPROVEMENTS

The notable aspects of the Eureka Process are: a) Cracked residue is not solid coke but pitch which is continuously processed as liquid in the plant. Refineries are free from troublesome and tough operations of solid coke handling. b) The yield of liquid product is advantageous over other coking processes. Both the characteristics are brought about by having lower hydrocarbon partial pressure by injecting steam into the reactor. During thermal cracking in a vessel-type reactor, cracked oil products are flushed out of the reaction system in accordance with its pressure. Heavy cracked oil remains in the reactor and is cracked further to light cracked oil with more amount of residue yield under the condition of higher hydrocarbon partial pressure, while more reaction time is required. Lower hydrocarbon partial pressure strips cracked oil out of the reactor and suppresses overcracking of liquid product and results in more liquid and less residue yields in a short residence time. When residual feedstock is cracked, polycondensation takes place at the same time. The polycondensed product is dissolved stably in the liquid phase of the product when the conversion is low. But compatibility decreases as conversion increases, and polycondensed

295 material as a coke precursor is easier to be separated, because the polycondensed product becomes more aromatic while the matrix portion of the cracked oil becomes more paraffinic. Having a narrow residence time distribution of reactant in the reactor as well as a lower hydrocarbon partial pressure to strip paraffinic cracked oil out of liquid phase in the reactor, is also important point of a view to keep liquid phase homogeneous in the reactor. The accumulated experiences and further development improved the latest design of the EUREKA process. There are two major modifications. a) The conversion of the residual feed is increased in the furnace from ca. 25% to 35%. b) A residence time of residue in the reactor or cycle time of a reactor swing is cut to three from four hours. The above modifications were realized after detailed investigations into basic process conditions with short and long term trial operations. The typical operating condition and product yields are shown in Table 1. The credits of the modifications in 105 ton/hr feed case are summarized as follows: 1) Saving of steam consumption by 5 to 10 ton/hr 2) Saving of fuel gas consumption by 2 to 3 MMkcal/hr 3) Reduction of pitch yield to 96 to 98 % 4) Increasing of light cracked oil yield to 103 to 106 %

Table 1. Improvement of the EUREKA Process

Feedstock Sp. Gr. CCR 538 de~.C+

d15/4 wt% wt%

1.011 17.2 87.0

Improved

Original

3

4

deg.C

495

489

deg.C

439

434

deg.C -

685 half

630 base

wt% wt% wt% wt%

5.2 36.0 32.2 26.6

4.9 33.9 34.2 27.1

Operating Conditions Operation Cycle Furnace Outlet Temp. Reactor Max. Temp. S.H.S Temp. Steam Rate

hr

Yields C4- Gas CLO (C5 to 370 deg.C) CHO (370 deg.C+) Pitch

296 The modifications also enabled a wide selection of feedstocks even from Chinese residue such as Shengli with relatively paraffinic properties which had been hard to process because of a poor operability in the reactor. Higher reaction temperature in the furnace is favorable for cracking against polycondensation of residue, because activation energy of cracking is higher than polycondensation. A residence time distribution is also so improved with more residence time in the furnace of plug flow and less in the reactor that unfavorable reaction of polycondensation is suppressed. A prediction model of fouling rate of the furnace was developed to decide optimal decoking cycles of the furnace. It is a very effective tool, of course, for the design of the new plant with three hours cycle of a reactor swing. 3. COMPARISON WITH DELAYED COKER

The performance of the improved EUREKA Process with three hours cycle of reactor swing is compared to a Delayed Coker in Table 2. The gas and pitch yields of the Eureka Process are less than that of Delayed Coker, which leads to the higher production of valuable liquid product. This difference is caused by the reaction condition, especially the hydrocarbon partial pressure in the reactor and residence time. In the coker drum, the hydrocarbon partial pressure is higher, which allows the heavy oil produced to remain longer in the drum, resulting in the production of additional coke and gas. It should be noted again for the EUREKA Process that refineries are free from troublesome and tough operations of solid coke handling experienced in Delayed Coker. Table 2. Comparison of the EUREKA Process to Delayed Coker Feedstock Sp. Gr. CCR 538 deg.C+

d15/4 wt% wt%

1.030 22.4 89.0 ....... ~ . ~ . ~

Operating Conditions Operation Cycle Furnace Outlet Temp. Reactor Max. Temp. Pressure. Yields C4- Gas CLO (C5 to 370 deg.C) CHO (370 deg.C+) Pitch Coke

........... P ~ ! . ~ . r . . ~ . g . . C . . ~ . ~ . ~

3

24

deg.C

495

500

deg.C Kg/cm2G

437 0.3

435 1.4

wt% wt% wt% wt% wt%

5.3 33.6 28.4 32.7 -

10.4 39.3 16.3

hr

34.0

.....

297 Table 3. Characteristics of Various Fuels

Volatile Matter Heat of Combustion H.G.Index Required Frame of Boiler

wt% kcal/kg

EUREKA Pitch

Delayed Coke

Fluid C o k e

Coal

40 - 50 8800-9200

6 - 14 7800-8600

4 - 10 7700-8000

150 - 170

50 - 100

15 - 20

7 - 50 3000-6000 50 - 70

small

larse

larse

larse

4. EUREKA PITCH AND ITS APPLICATIONS a) Fuel Pitch

EUREKA pitch is quality pulverized fuel as its characteristics shown in Table 3. The conventional boiler designed for fuel oil can be used for pitch burning as it is. Several Japanese manufacturers had employed EUREKA pitch as boiler fuel. 1) It is easy to burn because of its high content of volatile matter. 2) It is easily crashed to make pulverized fuel. 3) It is very safe against autoignition in storage, because the volatile matter contained in the pitch is very stable at ambient temperature. 4) It has a very high heat of combustion. 5) It has a small amount of ash. 6) It is free from troublesome material like shot coke in Delayed Coker. However, it is necessary to care that the pitch has relatively high content of sulfur, nitrogen, vanadium and nickel which are condensed from residual oil of feedstock to the EUREKA Process. b) A l t e r n a t i v e of C o k i n g Coal for Steel Industry

So far, EUREKA pitch is mainly consumed as an alternative of coking coal for steel industry in Japan. Most of EUREKA pitch produced in Sodegaura(Fuji Oil Company, Japan) and Nanjing(P. R. China) refinery are exported to Japanese steel industry. It is approximately 370,000 ton/year (70% of total production). The production of cokes blended with the EUREKA pitch from Sodegaura refinery is accumulated to 70 million ton to date. It is economical to use non-coking coal, which is cheep and abundant, for metallurgical coke supply to steel manufacturer. EUREKA pitch is consumed to increase mechanical strength of metallurgical coke from non-coking coal. Coking coal is able to be replaced by 30 % of EUREKA pitch and 70 % of non coking coal. When the steel industry recognizes a merit of employing EUREKA pitch, a very large market of EUREKA pitch is created along with a growth of the industry. It results in the saving of the imported coking coal.

298 370~ ~538~ C6 k16/~ )k k k 2

538~ C5 k15* ~ ~ CI

k12 ~ C 2

Heptane Soluble

150~ ~-370~ C7 ~ k 2 8

k23 ~ C 3

Heptane Insoluble Toluene Soluble

180~ & GAS C8 ~ k34

Toluene Indoluble Quinoline Soluble

~C4 Quinoline Insoluble

Figure 2. Reaction Mechanism Model. 5. RELATED TECHNOLOGIES

The design tools developed for the EUREKA process are very refined. They can be also applied to any process designs of thermal cracking other than EUREKA. a) A Practical Model of Thermal Cracking of Residual Oil

It is very important to have a mathematical reaction model, not only to analyze or understand the reaction mechanism but also to utilize the experimental results for designing a reactor or for operating it. The reaction of residue thermal cracking is fast in the early stage of reaction, but slows later. In the early stage of reaction there exist many molecules with long alkyl chains or with a structure represented by bibenzyl type bonding. They have a high reaction rate and result in a high yield of cracked oil. Polymerization and condensation of residue proceeds at the same time. Molecules without an easy-cracking portion have a very slow reaction rate and yield very small amounts of cracked oil. In the final reaction stage, polycondensation of residue and gas production, such as from dehydrogenation, is dominant and little cracked oil is produced. The reaction of residue cracking can be predicted as follows with the lumping model shown as in Figure 2, which is commonly used in the petroleum refining reaction. The effects of hydrocarbon partial pressure and residence time distribution are easily incorporated into the reaction model, when the reaction kinetics are mathematically described. Figures 3 and 4 show the examples of simulated results by use of the model. They clearly show that the effect of hydrocarbon partial pressure on the product distribution is larger than generally recognized and low pressure is preferable to obtain high liquid yield and that residence time distribution should be controlled as narrow as possible to reduce coke precursor(Q 1) in the residual component. Figure 5 shows the simulated results of reaction in the one cycle of the EUREKA operation.

299

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40

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RESIDUEIQS) 20RESIDUE(QI) I-----] TEMP:425*C SP 180"C

lO 0 Porg lS.SkPo 36.0 kPo lOl.SkPo 445.8kPe To 523"C 478"C 425*(:: 330"C HYDROCARBON PARTIAL PRESSURE

O~

1

-Y

50

lO0

t

I

150 200

l

250

SOFTENING POINT (*C)

Figure 4. Effect of Residence Time Distribution.

Figure 3. Effect of Hydrocarbon Partial Pressure on Product.

b) A Tubular Fouling Model for Residue Cracking Furnaces Tubular fouling by coke deposition on its inner wall has restricted the development of a high-conversion furnace for residual oil. Such a furnace is desirable from the viewpoint of economics and operability. The proposed tubular fouling model as shown in Figure 6 well represents the complex phenomena of tubular fouling of a residue conversion furnace by modelling a sedimentation of coke precursor and its reaction into coking material in and out the boundary film.

4OO

60

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~

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Figure 5 Simulation of Commercial Operation

lO0

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TIME (rain)

200

300

I

I

i 7000 , hr 50O

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S e p a r a t i o n of coke p r e c u r s o r in bulk flow Diffusion of coke precursor into boundary film F o r m a t i o n of caking m a t e r i a l Cx in b o u n d a r y film Diffusion of caking m a t e r i a l out of boundary film

Figure 6. Tubular Fouling Mechanism.

A

/ / _

//

-"~ /

~

o

II

..%,,.s/ f t

5000 hr - 2.0 12_ 0-.-'-hr" " " ' " ..-.----~__,-.,_...

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I

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I

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t

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TEME- 1.5 w rr" ZD /2" / '-:;:~ (/3 #// / / PRESSURe%.. - 1.0 u') LI.J 121 13_ U ......... / SKIN TEMP. DATA "'%

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TUBE SKI~/ f " " 0-'---- o TEMR/ p.. I--'- ~----"'--

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0 12_ w Eb

r - - - ~ -5I-0- '0- 0' r ' - - - h I

1

1

I

I

IN

I

OUT

TUBE LENGTH

Figure 7. Simulation of Tubular Fouling in a Commercial Furnace

Figure 7 shows analytical results of a commercial Eureka furnace. No direct measurements of coke thickness were available, but the increase of skin temperature and Dp against operation time were well predicted by the simulation model. In the improvement of the EUREKA process a very high furnace conversion is aimed at. From the conventional concept of furnace design it may seem to be beyond a practical scope, but studies of furnace design with the tubular fouling simulator indicated the realization of a new type of furnace. 6. CONCLUSIONS 1. The EUREKA process is a commercially proven thermal cracking process. 2. The liquid yield is advantageous over other caking processes. 3. Cracked residue is not solid coke but pitch. It makes the operation site clean and compact compared to other Caking processes, because pitch is handled as liquid in the process. 4. Refineries are free from troublesome and tough operations of solid coke handling. 5. The accumulated experiences and further developments enabled a wide selection of feedstocks as well as a reduction of utility consumption. 6. The pitch is utilized as boiler fuel or gasification feedstock having an excellent quality comparable to fuel oil. 7. It is also a precious alternative source of caking coal for steel industry.

301 8. The design tools developed for the EUREKA process are very sophisticated and made our designs very sure and reliable. REFERENCES 1. R. Watari, et al, "The Development of the New Eureka Process", NPRA Annual Meeting, San Antonio, March, 1987 2. T. Takatsuka, R.Kajiyama, H.Hasimoto, I.Matsuo and S.Miwa, "A Practical Model of Thermal Cracking of Residual Oil", J. Chem. Eng. Japan, 22, (3) (1989) 304. 3. T. Takatsuka, R.Kajiyama, H.Hasimoto, I.Matsuo and T.Hanazawa, "A Tubular Fouling Model of Residue Cracking Furnaces", J. Chem. Eng. Japan, 22, (2) (1989) 149. 4. I. Matsuo et al, "The Development of the New Eureka Process", International Congress of Pacific Basin Societies, Hawaii, December, 1989 5. RAROP, "Heavy Oil Processing Handbook", p9, 1991

This Page Intentionally Left Blank

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All fights reserved.

303

NEW C A T A L Y T I C T E C H N O L O G Y F O R FCC GASOLINE SULFUR R E D U C T I O N W I T H O U T YIELD PENALTY Ulrich Alkemade a and Timothy J. Dougan

b

a Manager Catalyst Evaluation, GRACE Davison FCC Europe, Worms, Germany b Manager of Marketing Services GRACE Davison FCC, Baltimore, USA GRACE GmbH, In der Hollerhecke 1, Postfach 1445, D-6 7545, Germany 1. A B S T R A C T Sulfur content limitations in gasoline are becoming increasingly stringent as environmental legislation is enacted throughout the world. There is a strong incentive to reduce the sulfur content of FCC gasoline in the most cost effective manner since FCC gasoline is the main contributor to the gasoline pool. This paper introduces emerging new GRACE Davison technolgy that enables a reduction in the FCC gasoline sulfur content by up to 30%. This technolgy is selective to sulfur species in the middle of the gasoline boiling range and converts those species to H2S, while preserving most of the base catalyst selectivities. This new technology is contrasted to existing options for reducing FCC gasoline sulfur content. 2. I N T R O D U C T I O N As in the US, member nations of the European Community will introduce further specifications for transportation fuels over the next few years. Besides other components, the sulfur content of transportation fuels and gasoline in particular will be limited.FCC gasoline can contribute up to 90% of the sulfur in the gasoline pool. The parameters thatcontrol the sulfur levels in gasoline have been described by various authors in the past. The main determinant of sulfur levels in the FCC gasoline is the feedstock. Researchers found that the reactions that converted the feed sulfur compounds in the FCCU were kinetically controlled and were dominated more by catalyst contact time than by catalyst-to-oil ratio [ 1]. Furthermore, it was shown that for most feedstocks, as the H2S yield increased for a given feedstock, the sulfur in the gasoline increased. They also found that mild hydrotreating of the feed tended to remove H2S precursors, so that the H2S yield was decreased, as well as the sulfur in the gasoline, but that the sulfur in the heavy liquid products remained constant. From a commercial point of view it is important that hydrotreating or desulfurisation technologies involve high capital investments [2]. A technically easy option to reduce the sulfur content of FCC gasoline is to lower the endpoint, since a large fraction of the sulfur is concentrated in the heavy end. A decrease in the end-point from 221 ~ to 17 I~ will, for a given example, lower the sulfur content of the gasoline by 42%. Commercially, however, this option is not at all attractive since the gasoline

a corresponding author

304 yield decreases by about 15 wt% for the same drop in the end-point depending on the distillation curve [3, 4]. Looking at the various options described above, the development of technology that enables the refiner to desulfurise products within the FCC process catalytically appears very attractive. 3. EXPERIMENTAL Catalysts were tested in the Davison Circulating Riser, DCR, in the adiabatic mode with a riser outlet temperature of 521 ~ This pilot unit gives a very good simulation of the refinery scale FCCU operation [5]. Complete sulfur balances were also performed around the individual runs. H2S in the light gas fraction was determined by GC analysis using a Chromosorb 107 (Alltech) column at 90~ Feed sulfur as well as the total liquid sulfur (gasoline, LCO, HCO) was determined by LECO sulfur analysis. Sulfur on the coke was determined by SO 2 analysis of the flue gas from the regenerator. Gasoline sulfur analysis was determined by GC using an atomic emission detector, as described below. Gasoline compositions were determined using a gas chromatograph equipped with an atomic emission detector (GC-AED). The separation of the individual sulfur compounds was accomplished using a column coated with a specially developed liquid phase, SPB-1 SULFUR (Supelco Inc., Bellefonte, PA 16823-0048). With this column, retention times were found to be proportional to boiling points over the range studied. The GC-AED is a unique GC detector system; as compounds elute from the GC column, they are broken down into constituent atoms and excited to a higher energy state by a microwave induced plasma. These atoms then relax to a lower energy state, and emit light characteristic of the atom. This light is filtered using a diffraction grating and detected by a photo diode array positioned along the focal plane, where wavelength and intensity are measured. When the detector is adjusted to the emission wavelength of sulfur (181 nm), only those compounds which contain sulfur are detected. In this way the GC-AED acts as a "filter", removing standard hydrocarbon peaks from the chromatogram [6]. Individual peaks in the sulfur GC-AED chromatogram shown in Figure 1 were assigned by either gas chromatography-mass spectrometry (GC-MS) or by retention time conformation using prepared standards. For GC-MS identification, compounds were first separated and detected using the GC-AED, monitoring both the carbon (191 nm) and sulfur (181 nm) emission lines. The separation was then duplicated on the GC-MS using the GC-AED carbon trace as a "fingerprint" guide to zero in on the retention window where the sulfur compounds could be found. Electron ionisation GC-MS followed by significant background subtraction was then used to obtain the mass spectra. Comparison of these spectra to a mass spectral library of pure compounds yielded the identity of these sulfur compounds. Assignments were confirmed by retention time on the GC-AED whenever possible. The sulfur chromatogram shows in order of increasing retention time (boiling point): mercaptans, thiophene, methylthiophenes, tetrahydrothiophene, ethylthiophenes, propylthiophenes, butylthiophenes, and benzothiophene. The last eluting peak identified in the chromatogram is benzothiophene, which boils at 218~ very near the end-point for standard 221 ~ gasoline.

305

Mercapt.ans

ThiopheMel nehyllhiophenes jTelrahydrothiophene

I EthyIthiophenes ] | Propylthiophenes

I /

13enzothtophene

I Buiylthiophenes i

f 2M'J

0

~p

.

9.

.

.

.

.

.

.

.

""

R

R---SH

Figure 1. Sulfur GCAED Chromatogram.

The properties of the feedstocks used for the GSR study are given in Table 1. A low- and mid-sulfur feed was used, besides the different sulfur levels the feeds also showed different crackabilities which was expected from the different K factors. 4. RESULTS AND DISCUSSION Table 2 shows the properties of the catalysts tested aiter a 4 hours, 1500~ 100% steam deactivation. The catalysts were chosen to give a wide range of catalytic properties. In the simplest terms, the USY/MATRIX and the USY-G are octane oriented catalysts with low hydrogen transfer activity, whereas the REY and the REUSY are maximum gasoline yield oriented catalysts with high hydrogen transfer activity. The gasoline composition from the low hydrogen transfer systems is more olefinic, whereas, in the high hydrogen transfer case the olefins are converted to isoparaftins and aromatics. Each of the four catalysts was tested in the DCR under adiabatic conditions with a reactor outlet temperature of 521~ For each catalyst a range of conversions was obtained by varying the cat/oil ratio via feed preheat temperature, so that conversion effects could be studied. Figure 2 shows the sulfur balance data for the DCR study as a percentage of feed sulfur versus conversion. Material balances ranged from 99-102% for the study. Most of the feed sulfur stays either in the heavy liquid fraction, LCO/HCO, or is converted to H2S. About 3% of the feed sulfur is deposited with the coke, with about 4% converted into the gasoline range.

306 Table 1. Feedstock Properties. Feedstock Name Gravity Aniline Pt. Sulfur Total Nitrogen Basic Nitrogen Conradson Carbon

N V Fe Na SIMDIST Distillation: Vol %, Temp. o C ibp 10 50 90 fbp 'K' Factor

kg/m 3 C~ wt% wt% wt% wt% ppm ppm ppm ppm

Feed A Low S 887.4 82 0.474 0.07 0.02 0.16

Feed B Mid S 894.1 82 1.05

0.14 0.10 0.4

0.32 0.68 9.1 2.9

170 252 368 491 574 11.68

181 266 380 503 610 11.59

0.23

Table 2. Catalyst Properties. Hydrothermal Steam *

MAT Zeolite Area Matrix Area Unit Cell Re203 * 4 hrs, 816 ~

wt. % m2/gm m2/gm A wt. %

REY

REUSY

USY-G

62 38 24 24.49 4.73

72 131 23 24.30 2.79

71 196 25 24.24 -

USY/ MATRIX 64 81 55 24.19 -

100 % stm., 0 psig.

Figures 3 and 4 show the total sulfur in gasoline concentration for the low sulfur feed for standard T90 = 193~ gasoline and cut gasoline at T90 = 149~ The effect of the lower T90 is to reduce the gasoline sulfur content from this feed by about 40%, with a 15 wt% decrease in the gasoline yield at the same time. This figure also shows that the standard T90 = 193~ gasoline has both a conversion and a catalyst-type dependence. In contrast, the gasoline cut shows a weaker dependence on conversion, but is still dependent on catalyst type. The catalysts which have the lowest hydrogen transfer activity yield more sulfur in the gasoline than the catalysts with higher hydrogen transfer activity.

307

Figure 3. For Different Catalyst Types versus Conversion (DCR data, 521 ~ Reactor, Low S Feed, 0.5% S).

The effect of reactor temperature variation is shown in Figure 5. Decreasing riser outlet temperatures showed somewhat reduced sulfur levels in the gasoline for a given conversion.The effect of the GSR technology on a USY/Matrix catalyst using a mid- and lowsulfur feed is shown in Figures 6-9. The sulfur level of the full range gasoline is reduced over the whole conversion range for both feeds and the gasoline sulfur levels of the base catalyst and the GSR catalyst decrease roughly parallel with increasing conversion. The GSR technology shows an effect on sulfur compounds over the whole boiling range, which would result in a dependence on gasoline end-point as shown previously. The effect of the GSR technology on a REUSY catalyst is shown in Figures 10-11 for the low sulfur feed. Here the same applies as for the USY/Matrix catalyst, however, the absolute sulfur levels are somewhat lower than in the USY/Matrix case. This clearly demonstrates the importance of hydrogen transfer reactions for GSR sulfur reduction especially for the smaller substituted thiophenes. In this case a relatively small end-point correction should have a significant effect.

308

Figure 4 Sulfur Gasoline Product Yields at Constant Conversion = 72 wt.% Low S Feed, 0.5% S.

Figure 5 . Sulfur in Full Range and Cut Gasoline For Different Reactor Temperatures versus Conversion DCR data. Low S Feed, 0 5% S. USYNATRIX Catalyst.

Figure 6 . Sulhr in Full Range and Cut Gasoline for USYMATRIX GSR Technology versus Conversion DCR data, 521°C Reactor, Low S Feed, 0.5% S.

Figure 7. Sulfur Gasoline Product Yields at Constant Conversion = 70 wt.% Low S Feed, 0.5% s.

Most selectivities of the base catalysts are not affected by the GSR technology, except for approximately 8% higher coke yields. This problem can be overcome in many cases by proper choice of coke selective base catalysts. The sulhr removed from the liquid products is converted into HzS.

Figure 8. Sulfir in Full Range and Gasoline Cut for USY/MAIlUXGSR Technology Versus Conversion; DCR Data, 521 C Reactor, Mid S Feed 1 .O% S.

Figure 9. Suffir Gasoline Product Yields at Constant Conversion = 69 wt.% Mid S Feed, 1.0% S.

Figure 10. Suffir in Full Range Gasoline versus Conversion; DCR data, 0.5% S feed.

Figure 1 1. Sulfur Gasoline Product Yields at Constant Conversion = 75 wt.% Low S Feed, 0.5% S.

The effect ofthe GSR technology on the sulfur content of the L C 0 is illustrated in Figure 12. This shows a decrease of approximately 10% sulfur content over the whole conversion

range. For this experiment the USYfMatrix system was used on the low sulfur feed. Unlike the gasoline, the LC0 shows an increase in the sulfur level with conversion, which is most likely due to a concentration of stable aromatic sulfur compounds in the LCO.

310

0,68 9

,~uSu GSR Tr S Y / M A T R I X .]. I~,..

~0,64

. ........... II

..,,--_~~;~ ~

.,..~

I i ' t ..* 9 ,

....,,-

~0,60 I1~ ,""'~

~0,56 0,52

,.. t " ' ~ ..... /

...................'..................~.................................!.................................. 60

62

64

66 68 conversion, wt. %

70

72

Figure 12. Sulfur % in LCO versus Conversion; DCR Data, 521 C Reactor, Low S Feed, 0.5 % S. 5. SULFUR REDUCTION MECHANISM The various sulfur compounds present in the FCC feedstock show different behaviour, which is due to different chemical compositions. For example, most mercaptanes decompose under typical FCC conditions into hydrocarbons and hydrogensulfide as shown in Figure 13, while the thiophenic compounds are relatively stable in the FCC unit. The lower molecular weight thiophenic compounds end up in the gasoline and LCO, while the higher molecular weight compound remain in the HCO and coke [7]. The sulfur in the coke is subsequently transferred into the regenerator and oxidised to SOx. The goal of a sulfur reduction technology is of course to release all of the sulfur as hydrogen sulfide and to keep the sulfur in the coke to a low level in order to keep the SOx level in the regenerator flue gas to a minimum. In Figure 13 (part 2) a mechanism is proposed which suggests that thiophenic compounds can decompose into hydrocarbons and hydrogensulfide. This reaction involves tetrahydrothiophene, which is known to decompose under FCC conditions. Another way to reduce sulfur in liquid products is to decompose thiophene precursor molecules i.e. to decompose compounds that can form thiophenic compounds under FCC conditions. These reactions which reduce the sulfur content of liquid products from the FCC process are catalysed by the new GSR technology. This was substantiated by the experimental data presented in the previous section. 6. CONCLUSIONS The data presented in this paper show that the GSR technology effectively reduces the sulfur content of liquid FCC products. The sulfur compounds are converted into hydrocarbons and hydrogensulfide. With the exception of a somewhat increased coke yield, the selectivities of the base catalyst are not affected by the GSR technology. The higher coke yield can be controlled by proper choice of coke selective base catalysts in most cases. The data also showed that the GSR technology can be combined with all commercially available base catalyst systems.

311

1)

R - C . -cI-h- sI-h

H,,

/H

~

R - c . : c.2+ H2s

H

2)

H / ~ s / ~ H + 2 H2------t~ i ~

3)

Inhibition of Thiophene Formation

H

H ~

H +H2S

Figure 13. Decomposition Route for Mercaptans.

ACKNOWLEDGEMENTS The GSR technology has been invented and developed by R.F. Wormsbecher et al at the W.R. GRACE & Co. Research Center at Columbia Maryland, USA. Also all of the presented data were generated at this location. REFERENCES 1. E.G. Wollaston, W.L. Forsythe, I.A. Vasalos; Oil & Gas Journal August 2, (1971) 64. 2. G.P. Huling, J.P. McKinney, T.C. Readal; Oil & Gas Journal May 19, (1975) 73. 3. GRACE Davison FCC Technology Conference; Athens, September (1994) 27-30. 4. R.F. Wormsbecher, D.S. Chin, R.R. Gatte, T.G. Albro, R.H. Harding; NPRAAnnual Meeting New Orleans, March 1992. 5. G.W. Young, G.D. Weatherbee; A.I.Ch.E. Annual Meeting, San Francisco, November 1989. 6. T.G. Albro, P.A. Dreifuss, R.F. Wormsbecher; J. High Resolution Chromatography, 16 (1993) 13-17. 7. R.H. Harding, R.R. Gatte, J.A. Whitecavage, R.F. Wormsbecher; ACS Symposium Series 552 (1994) 286.

This Page Intentionally Left Blank

Catalysts in PetroleumRefining and PetrochemicalIndustries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All rights reserved.

313

THE INFLUENCE OF FEEDSTOCKS AND CATALYST F O R M U L A T I O N ON THE D E A C T I V A T I O N OF FCC CATALYSTS

R. Hughes 1, G. Hutchings 2, C.L. Koon 1, B. McGhee 3 and C.E. Snape 3

l Department of Chemical Engineering, University of Salford, M5 4W~, U.K. 2 Leverhulme Centre for Innovative Catalysis, Department of Chemistry, University of Liverpool, P.O. Box 147, Liverpool L69 3BX, U.K. Department of Pure and Applied Chemistry, University of Strathclyde, Glasgow G1 1XL, U.K. ABSTRACT The effect of quinoline and phenanthrene additions to a n-hexadecane feedstock has been determined for a model four-component FCC catalyst by means of a MAT reactor with analysis of all products and characterisation of the coke produced. Both additions lead to an overall loss in conversion; quinoline is considered to act as a poison while phenanthrene participates strongly in coke formation and the resultant coke becomes more aromatic in nature. The cracking propensity and associated coke formation have been measured for a series of FCC catalysts with differing compositions. Increasing amounts of zeolite in a matrix lead to increasing extents of conversion but with little effect on the extent of coke production. However, a pure zeolite gave a very high coke content. I. I N T R O D U C T I O N The catalytic upgrading of petroleum fractions in fluid bed/riser reactors (FCC) is a major refinery operation. Because of its industrial significance, this process has been the subject of much study, yet there remains a lack of detailed knowledge concerning the mechanisms of product formation and the deposition of coke during FCC operations. In addition, there have been relatively few studies concerning the role of poisons and feedstock additives in this complex process (1). Deactivation of catalysts, particularly by coke deposition (the main means of reversible FCC catalyst deactivation) has been the subject of intensive study over the past 50 years (2-4). Initially, the loss of activity was correlated with the time on stream, but it is now generally accepted that a more appropriate approach to understanding the effect of deactivation by coke is to relate deactivation to the deposited coke concentration (5). Furthermore, few studies on the effect of catalyst formulation on both the product distribution and coke formation have appeared in the open literature. In recent years, attention has tended to be focused on coke deposition in zeolites (6, 7) in order to characterise the coke formed. In one specific study Groten et al (8) carded out a study of coke formation using zeolite USHY with n-hexane as reactant, but in this case, as in others (6, 7) it was necessary to deposit excessive amounts of coke (> 5%) to enable characterisation of the coke deposits to be achieved. However, if demineralisation of the catalyst is used to concentrate the coke as in the present work the inherently quantitative single pulse excitation (SPE) ~3C NMR procedure may be used to characterise coke deposits on FCC catalysts at realistic levels of ca 1% by weight.

314 In this paper we present our results on a study of the deactivation and characterisation of FCC catalysts, together with product yields at realistic coke levels (0.5 to 1.0%), that are typically found on FCC catalysts during industrial operation. In particular, the effect of quinoline and phenanthrene as additives to the n-hexadecane feedstock has been studied at two concentration levels and the relative roles of these additives as catalyst poison and coke inducer are discussed. A further aspect investigated is the influence of catalyst formulation. Pure zeolites are seldom used as FCC catalysts: instead, catalysts comprise a number of components, which apart from the zeolite, may include matrix, binder and clay. In the present work, catalyst formulations ranging from 100% matrix to 100% zeolite have been examined and the influence of the various catalyst compositions on product distribution and coke formation is assessed. 2. EXPERIMENTAL A number of FCC catalysts was used in the present study. For comparison of the effects of quinoline and phenanthrene additions to the n-hexadecane feedstock a model catalyst of composition, zeolite US-Y (30%), silica binder (25%), Kaolin (25%) and pseudo Boehmite matrix (20%) was used. Quinoline and phenanthrene additions to the n-hexadecane amounted to 1% and 10%. The catalysts used to assess the effect of composition on product yields varied from a basic matrix material through a variety of zeolitic catalysts containing 20% zeolite and 15% silica binder, the remainder being clay, to a pure zeolite catalyst. Data for all the catalysts used are presented in Table 1. In addition for the 13CNMR analysis a sample of coked refinery catalyst obtained from a unit processing heavy feedstock was obtained. The coke level on this catalyst was 0.9% and 30g. of this catalyst was demineralised by standard HF treatment to produce a 250 mg sample of coke concentrate containing 65% carbon. Reactivity and coking were determined using a standard MAT test reactor (9). The products from the MAT reactor were analysed by gas chromatography (GC)and peak identifications for the liquid products were made with the aid of GC-MS used in conjunction with the concentration of the aromatic species by open column chromatography on alumina. To give a clear indication of the boiling point distribution of the products, the peaks in the chromatograms have been grouped using successive n-alkanes, although specific quantification of individual isomers was also obtained. This procedure could not be used as precisely for the n-hexadecane/quinoline mixtures because of the overlap of the quinoline and product peaks close to C12. Coke levels were measured by combustion and by the weight gain of the catalyst; both methods gave good agreement. Mass spectrometric (MS) analysis was conducted on the deactivated catalysts from the MAT reactors using a Vacuum Generators instrument in which the probe was heated from ambient to 500~ at a rate of 200~ minl and spectra over the mass range 50 to 600 were recorded every 5 s. Spectra were recorded in both electron impact (El) and chemical ionisation (CI, with ammonia) modes. A number of deactivated samples have also been analysed after extraction in chloroform to remove physically-trapped molecular species. Solid statel3C NMR analysis of the coke concentrate was carried out using a Brake MS4 100 spectrometer. The single pulse excitation procedures described elsewhere (10,11), were used to derive carbon aromaticity and the proportion of bridgehead aromatic carbon.

315 In this paper we present our results on a study of the deactivation and characterisation of FCC catalysts, together with product yields at realistic coke levels (0.5 to 1.0%), that are typically found on FCC catalysts during industrial operation. In particular, the effect of quinoline and phenanthrene as additives to the n-hexadecane feedstock has been studied at two concentration levels and the relative roles of these additives as catalyst poison and coke inducer are discussed. A further aspect investigated is the influence of catalyst formulation. Pure zeolites are seldom used as FCC catalysts: instead, catalysts comprise a number of components, which apart from the zeolite, may include matrix, binder and clay. In the present work, catalyst formulations ranging from 100% matrix to 100% zeolite have been examined and the influence of the various catalyst compositions on product distribution and coke formation is assessed. 2. EXPERIMENTAL A number of FCC catalysts was used in the present study. For comparison of the effects of quinoline and phenanthrene additions to the n-hexadecane feedstock a model catalyst of composition, zeolite US-Y (30%), silica binder (25%), Kaolin (25%) and pseudo Boehmite matrix (20%) was used. Quinoline and phenanthrene additions to the n-hexadecane amounted to 1% and 10%. The catalysts used to assess the effect of composition on product yields varied from a basic matrix material through a variety of zeolitic catalysts containing 20% zeolite and 15% silica binder, the remainder being clay, to a pure zeolite catalyst. Data for all the catalysts used are presented in Table 1. In addition for the ~3CNMR analysis a sample of coked refinery catalyst obtained from a unit processing heavy feedstock was obtained. The coke level on this catalyst was 0.9% and 30g. of this catalyst was demineralised by standard HF treatment to produce a 250 mg sample of coke concentrate containing 65% carbon. Reactivity and coking were determined using a standard MAT test reactor (9). The products from the MAT reactor were analysed by gas chromatography (GC)and peak identifications for the liquid products were made with the aid of GC-MS used in conjunction with the concentration of the aromatic species by open colunm chromatography on alumina. To give a clear indication of the boiling point distribution of the products, the peaks in the chromatograms have been grouped using successive n-alkanes, although specific quantification of individual isomers was also obtained. This procedure could not be used as precisely for the n-hexadecane/quinoline mixtures because of the overlap of the quinoline and product peaks close to C~2. Coke levels were measured by combustion and by the weight gain of the catalyst; both methods gave good agreement. Mass spectrometric (MS) analysis was conducted on the deactivated catalysts from the MAT reactors using a Vacuum Generators instrument in which the probe was heated from ambient to 500~ at a rate of 200 ~ min1 and spectra over the mass range 50 to 600 were recorded every 5 s. Spectra were recorded in both electron impact (EI) and chemical ionisation (CI, with ammonia) modes. A number of deactivated samples have also been analysed after extraction in chloroform to remove physically-trapped molecular species. Solid statel3C NMR analysis of the coke concentrate was carried out using a Brake MS4 100 spectrometer. The single pulse excitation procedures described elsewhere [ 10,11 ], were used to derive carbon aromaticity and the proportion of bridgehead aromatic carbon.

316 TABLE 1 Physical Properties of Catalysts Catalyst

Type

Surface Area m2.g -1

Alumina Wt.%

Rare Earth Oxides Wt.%

Unit Cell Size A~

MAT16

4 Component

298

37.5

BPM1

Matrix + Clay/Silicabinder

102

45.6

0.0

Z-A2

Zeolite + Clay/Silica binder

125

24.1

0.6

24.26

Z-A4

Zeolite + Clay/Silica binder

134

24.3

1.1

24.28

Z-A6

Zeolite + Clay/Silica binder

143

24.2

2.7

24.33

LZY1

Zeolite Only

24.53

26.0

3 RESULTS AND DISCUSSION 3.1 Influence of additions to n-hexadecane feed.

Using the standard model four component catalyst, experiments were carded out in the MAT reactor for a n-hexadecane feed containing 1% and 10% of quinoline and phenanthrene additions. The results obtained are presented in Figs. 1 and 2 respectively in terms of a normalised yield, defined as the weight of product divided by the weight of injected feed. In all cases a feed rate of 2.7 ml/min was used with a catalyst charge of 4 g in the reactor. The temperature of operation was 530~ Analyses of the gaseous product were made for the C1 to C5 range, while the liquid product distribution was examined in the C5 to C15 range. Liquid products were characterised using GC-MS and a range of aromatic compounds were identified in which the concentrations of alkybenzenes are greater than those of alkylindans and naphthalenes while polynuclear aromatic compounds (PACs) were only minor constituents. The prominent group of constituents eluting between n-pentadecane and n-hexadecane are mixtures of alkenes, alkylbenzenes and naphthalenes. Phenanthrene addition had no significant effect on the overall liquid product distribution. The product distribution shows a maximum for C3, C4 and C5 products for n-hexadecane and for both additives. In general, the effect of increased additive is to decrease the extent of individual product formation. This effect is most marked for quinoline where even the addition of

317

4.5

N-HEXADECANE

F/77/] I WT~. QUINOLINE

I.-

3.5 C~ ._.I t.~

I0 WT% QUINOLINE

3

,.,,.,,

>" 2.5 E3 l.t.l N 2 ._I

~E 1.5 r'Y 0 Z 1

0.5

Cl C2 C3 C4 C5 C6 C7 C8 C9C10CllC12C13C14C1~0KE

CARBON NUMBER FIGURE 1. EFFECT OF QUINOLINE ADDITIONS ON YIELD OF N-HEXADECANE FEED.

N-HEXADECANE

F/7~ I WT% PHENANTHRENE

I--

):

4.

I0 WT% PHENANTHRENE

...J I.i >'3 C3 bJ N ,==,. _.J ~;2 rY O z

I

I

I

I

I

1

I

I

h h _ LtJ l

I

I

I

I

I

I

C1 C2 C3 C4 C5 C6 C7 C8 C9C10CllC12C13C14C15C0KE

CARBON NUMBER FIGURE 2. EFFECT OF PHENANTHRENEADDITIONSON YIELD OF N-HEXADECANE FEED.

318 1% of this compound causes a considerable reduction in product yield. At the 10% level the effect is even more marked and is considerably stronger than the effect of phenanthrene. Thus, addition of 10% quinoline caused a 30 fold reduction in C5 products, whereas the addition of the same amount of phenanthrene reduced the C5 product yield to 40% of the pure n-hexadecane feed. Coke levels were surprising constant for all these experiments, averaging about 0.7% by weight, but the 10% phenanthrene gave a value for about 1.0%, as might be expected, due to its aromaticity. From these results, the quinoline appears to act as a severe catalyst poison. However, while hydrocarbon products are drastically reduced, coke levels remain relatively unaffected and are comparable to those of the pure n-hexadecane, suggesting that quinoline acts as a coke inducer as well as a catalyst poison. An important factor in commercial operation is the relative amounts of alkene produced, relative to alkanes. Alkene/alkane ratios for the C1 to C5 range are presented in Fig. 3 for nhexadecane and for 1% and 10% additions of quinoline and phenanthrene to the n-hexadecane feedstock. In all cases the ratio was greater than unity, with 1% addition of additives having relatively little effect on this ratio. However, at 10% addition, phenanthrene enhanced this ratio, whilst quinoline showed a corresponding decrease. Thus, although these additives diminished the individual yields of the gaseous components, with a marked reduction in the case of quinoline, small concentrations had little effect on the alkene/alkane ratio. Coke deposits were studied using mass spectra obtained from the probe E1 and CI analyses of the deactivated catalysts arising from the various feed streams. Alkane and alkene fragments were observed to dominate the individual mass spectra (particularly, m/z 57, 71 and 55, 69, respectively, in the E1 mode). Although alkylaromatics were evident for the catalyst from the tests with n-hexadecane and the n-hexadecane/phenanthrene mixture PACs are only present in trace quantities. Quinoline addition gave rise to much less intense ions from the deactivated catalyst due to its lower carbon content and the reduced sensitivitymade it difficult to observe the aromatic fragments. Indeed, the most intense peak was from quinoline itself(m/z 129 El, 130 CI). Phenanthrene addition would not appear to significantly increase the amounts of aromatic fragments evolved from the deactivated catalyst. These are primarily alkylbenzenes as observed for 100% n-hexadecane. However, leaving the catalyst at reaction temperature for 15 min. gave rise to a significant increase in the abundance of the aromatics fragments with naphthalenes (m/z 128) evolving in much larger quantities. Chloroform extraction appeared to reduce the concentrations of aromatic fragments observed indicating that the actual coke forms is highly aliphatic in character with alkene groups accounting for most of the sp2 hybridised carbons. All the above experiments were based on a n-hexadecane feedstock. In order to characterise coke deposits using ~3CNMR, a catalyst deactivated from the processing of a heavy oil feedstock was used. The carbon skeletal parameters obtained are summarised in Table 2. The present technique uses the single pulse excitation (SPE) analysis and a comparison with the more conventional cross polarisation (CP) technique, shows that CP significantly underestimates the carbon aromacticity (0.92 compared with 0.96). The fact that over 80% of the carbon has been observed by SPE indicates that the procedure is quantitively reliable for catalyst cokes and that graphitic layers are not present in significant amounts. If present, their paramagnetism would have detuned the probe resulting in little of the carbon being observable. As for the aromaticity, CP also

319

2.5 N~ADECANE

I PHENANTHRENE

2

QUINOLINE 1.5

.5 < 0.5

0 0

1

10

WT% OF POISONIN N-HEXADECANE. FIGURE 3. ALKENE/ALKANE RATIO FOR MAT16 CATALYST WITH QUlNOLINE AND PHENANTHRENE. 14 i

12-

:::

+

LZY 1

F7777A Z-A6

I.-

~:io-

BPM1 ILl >a N ...J .< ]E

QC

8

6-

4.

0 z

tkl ,I

I

--

l

I

I

I

i

I

I

I

!

C1 C2 C3 C4 C5 C6 C7 C8 C9C10CllC12C13C14.C15COKE

CARBON NUMBER FIGURE 4. PRODUCTYIELD OF LZY1, Z-A6 AND BPM1.

320 TABLE 2 13C~ Results for Coke FCC Concentrate. ~3C T~ (aromatic): 0.5 and 10 s (two components of similar proportions)

SPE

CP

Carbon aromaticity:

0.96

0.91

Quaternary aromatic C

0.72

0.51

(Cqa/Car) CH3/aliphatic C: 0.75 Fraction of bridgehead aromatic C: 0.65 (.'. highly condensed).

grossly underestimates the fraction of quaternary aromatic carbon. From the value of 70% derived by SPE (Table 2), it is estimated that bridgehead aromatic carbons account for ca 65% of the total aromatic carbon. The only assumption needed is that each aliphatic carbon is bound to one aromatic carbon which is not unreasonable in view of the fact that arylmethyl groups account for 75% of the aliphatic carbon (Table 2). The aromatic structure is dearly highly condensed corresponding and the proportion of bridgehead aromatic carbon corresponds to 15-20 fiased aromatic tings.

3.2 Effect of Catalyst Formulation As indicated above MAT experiments were made to assess the influence of catalyst composition for a number of materials with zeolite contents ranging from 0% zeolite (matrix only) through various rare earth additions Z-A2, Z-A4, Z-A6 to 100% zeolite (LZY1). The product yield for BPM1 (matrix), Z-A6 and LZY1 are illustrated in Figure 4. As expected the matrix material BPM1 gave the lowest overall product yield, while the zeolite LZY1 gave the highest gas product yield, but the 20% zeolite catalyst Z-A6 gave the highest yield for the liquid products range. The most remarkable of Fig. 4 is the extremely large amount of coke obtained using the zeolite LZY1, which produced approximately 12-15 times as much coke as the other catalysts including the MAT 16 catalyst. A plot of the alkene and alkane yields and the alkeneYalkane ratio forthese three catalysts and the MAT 16 catalyst (the model for component material) is given in Fig. 5. Again the 100% zeolite catalyst LZY1 produces by far the greatest yield ofalkane whereas all the other materials produce more alkene than alkane and thus producing values of the alkene/alkane ratio in excess of unity. On the basis of product yield the catalyst Z-A6 is seen to be superior. Fig. 6 shows the effect of rare earth additions on product yield. Increase of rare earth content (together with the associated increase in surface area), results in a significant increase in product yields.

321

1

~30 I

1"8 !

1

o__16;

t

414 re,

i

i

,- 2 r--,

!

I.--

!

'-20

z

iT

w12 i Z

9

.~

i i

''

-15 i

J

4

\0.8

~

F10

rh L.I.J N ...J

ps

r"F O Z

-~

Z

Ii

'j " ov 6 9 j 4044

i t

I

i

0.2+i

I

,

~

BPM I

Z-A6

MAT 16

&LKENE/ALKANE RATIO I

LZYI

ALKENE, WT%.

ALKANE, WT%.

FIGURE 5. ALKENE/ALKANE RATIOS FOR LZYI, Z-A6, BPMI AND MAT 1 6.

14

i i

(1.1)

i.-

d

_J Lo >rm L0 N _J 4 rF O Z

(2.7)

Z-A CATALYSTS:

12~

,4-

I

-I-

! iI

ALKANE,

WT%.

ALKENE,

WT%.

/ /,"

/ I

_

// ,,,

RARE EARTH OXIDES, (WT,o) ~ iN BRACKETS.

~

/

m--

(0.6)//./ / /

( 2-

0 100

I()5 --110

~115

120

125

130

155

I~0

_J LIJ >-.

145

SURFACE AREA, m 2GI FIGURE 6. EFFECT OF SURFACE AREA AND RARE EARTH CONTENT ON ALKENE AND ALKANE YIELD.

322 4. CONCLUSIONS An experimental study has shown that the addition of quinoline and phenanthrene to a nhexadecane feedstock in MAT experiments leads to a loss in overall conversion. Characterisation of the coke from this feedstock, indicates that the initial coke formed is highly aliphatic in nature. Quinoline acts primarily as a catalyst poison but also favours coke formation. Solid state 13CNMR was used to characterise the coke formed from a heavy oil feedstock on demineralisation of the deactivated catalyst. The coke was now observed to be aromatic and highly condensed and it was possible to achieve this characterisation at realistic coke levels of ca. 1% without employment of large coke deposits as hitherto. An examination of catalyst formulation on product yield for a number of catalysts of various zeolitic content has shown that the most effective catalyst is of intermediate zeolite content. A catalyst containing 100% zeolite results in a very large amount of coke deposition. REFERENCES .

2. 3. 4. 5. .

7. 8. 9. 10. 11.

J.R. Kittrell, P.S. Tam and J.W. Eldridge, Hydrocarbon Processing 64, No. 8 (1985) 63. J.S. Butt, Catalyst Deactivation, Adv. Chem. Series 109 (1972) 259. R. Hughes, Deactivation of Catalysts, Academic Press, London (1984). E.H. Wolf and F.Alfani, Cat. Rev. Sci. Eng. 24 (1982) 329 and references therein. G.F. Froment in "Progress in Catalyst Deactivation". (J.L. Figueiredo, Ed). NATO Adv. Study Inst. Series-E54, Nijhoff, The Hague, 1982. M. Guisnet and P. Magroux, Appl. Catal. 54 (1989) 1. J. Biswas and I.E. Maxwell, Appl. Catal., 63 (1990). W.A. Groten, B.W. Wojciechowski and B.K. Hunter, J. Catal. 125 (1990) 311. R.W. Mott, Oil and Gas Journal, Jan 26th (1987) 73. G.D. Love, R.V. Law and C.E. Snape, Energy and Fuels, 7 (1993) 639. M.M. Maroto-Valer, G.D. Love and C.E. Snape, Fuel (1994), In Press.

ACKNOWLEDGEMENTS We thank the SERC (UK) for financial support of this work and the SERC Mass Spectrometry service at the University of Swansea for analysis of deactivated samples. We also acknowledge the generous assistance of Dr. N. Gudde at BP Oil and of Crosfield Chemicals for provision of catalyst samples and data on these.

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All fights reserved.

323

RESID FCC OPERATING REGIMES AND CATALYST SELECTION P. O'Connor a, S.J. Yanik b

aAkzo Nobel Catalysts, The Netherlands bAkzo Nobel Catalysts, USA 1. ABSTRACT There are some very clear differences in the operation and catalyst requirements of various commercial Resid FCC (RFCC) units. In this paper, the differences between activity-limited and delta-coke-limited RFCC operations are elucidated and the related catalyst performance requirements and catalyst selection methods are discussed. The effect of the catalyst-to-oil ratio on conversion and on catalyst site utilization and poisoning plays a key role in the transition of an RFCC unit from a catalyst-activity-limited regime to a cat-to-oil-limited regime. For the activity-limited operation the catalyst resistance to poisons with the given feedstock will be the most important selection criterion. For the delta-coke-limited operation, a reduction of the commercial delta coke of the catalyst is crucial. Commercial delta coke consists of various components, which are discussed in this paper along with methods for their evaluation. In both cases the use of realistic catalyst deactivation methods and feedstock will be essential in order to arrive at the correct catalyst choice. 2. RESID FCC AND OPERATING REGIMES The processability of resid in FCC and the role of the catalyst have been extensively discussed in the literature[I-4]. Depending on the feedstock, feed pretreatment, unit design, and operating philosophy, the priority of the various catalyst performance characteristics may differ considerably [4,5]. An interesting example is the comparison of the operation of FCC units with and without heat removal. Resid FCC units with heat removal are of'ten limited by the activity of the catalyst and consequently the (bottoms) conversion which can be obtained, while Resid FCC units without heat removal are mostly limited by the delta coke of the catalyst and hence resid intake or feed quality (e.g., feed concarbon residue content). Clearly, improvements in FCC catalyst metal resistance and activity retention and in coke selectivity will allow the refiner to increase (bottoms) conversion and increase the intake of lower-valued residual feedstocks. On the other hand, the RFCC operating constraints will in general have a bigger impact on the profitability of the unit than incremental yield improvements. It is worth noting here that the absence of a regenerator catalyst cooler does not automatically entail a delta-coke-constrained operation, while on the other hand if heat removal facilities are present, the unit operation can still be constrained by delta coke, for instance if the objective is to increase the resid content of the feed.

324 In this paper we address the differences between activity and delta-coke-limited Resid FCC and discuss the resulting appropriate operating regimes, related catalyst performance criteria and catalyst selection methods. 3. ACTIVITY-LIMITED AND DELTA-COKE-LIMITED RFCC 3.1. RFCC constraints

We can distinguish two "generic" types of RFCC applications [4]. For the first type the conversion of the resid-type feedstock is limited by the activity of the catalyst or by the volume of light gas produced. For this type of operation we require a catalyst which has a good activity retention even in the presence of metals, with good bottoms cracking and gas selectivity. For the second type, the critical success factor will be the ability to circumvent the limitation in coke production and/or the maximum regenerator temperature constraint. Obviously, this type of operation will require a catalyst which excels in coke selectivity. We will discuss the two generic types in more detail, making use of some simple causal loop diagrams with the conventions as shown in Figure 1.

A

B

C

DIRECTION OF CHANGE

Figure 1. Casual loop diagram conventions. 3.2. Activity-limited RFCC conversion

The case where RFCC conversion is limited by activity is quite simple and is illustrated in Figure 2. If the objectives of the operation are to increase conversion and increase resid intake, the options are as follows: 9 Increase catalyst addition; 9 Increase the activity of the fresh catalyst; 9 Increase the catalyst's resistance to deactivation by poisons (V, Na). Obviously, the conversion in a commercial unit is not only a function of the catalyst activity (reaction rate, KR), but also of the catalyst-to-oil ratio (CTO) and the effective contact time in the reactor (t). The simplified FCC kinetics assuming second-order cracking are summarized as follows: Conv = KR X

CTO x t

100 -Conv

where Cony is conversion, KR is reactor rate (activity) and t is reaction time

325

RESID \ . . . . . . i INTAKEI %

.,., _

,,

#'

Objective Hi Resid

Consequence Lo Conversion

Lo Addition

Lo Conversion

SolLtdon Hi Addition Hi Fresh Activity Hi V, Na, Resistance Hi Fresh Activity Hi V, Na Resistance

Figure 2. Activity-limited RFCC conversion. In a heat-balanced operation at constant reactor temperature, activity, delta coke, and CTO are related in the way shown in Figure 3. Consequently an increase in catalyst activity will have a direct positive effect on conversion on one hand, but will also have a negative effect because of the increase in delta coke and hence reduction in CTO. 3.3. Delta-coke-limited RFCC conversion

As mentioned in the previous section, the RFCC operation can become delta-coke-limited by a constraint on the regenerator temperature or the air blower capacity. If the objectives of the operation are to increase resid intake or conversion, the options then will be: 9 Reduce delta coke of the catalyst; 9 Improve selectivity of catalyst to dehydrogenation (Ni); 9 Increase CTO by reducing catalyst activity. Clearly the last option mentioned is the most controversial one, as it implies that an optimum catalyst activity can be found which maximizes the conversion of a certain operation. Indeed we have experienced several RFCC operations where this appears to be the case. In Figure 4 we have designated the RFCC operation where activity dominates the unit (bottoms) conversion as the Activity Regime and the operation where CTO dominates the unit (bottoms) conversion as the Cat-to-Oil (CTO) Regime. It should be noted that the CTO regime can start before the maximum delta coke (regenerator temperature) constraint is reached (Figure 5). What we have then is no longer regenerator-temperature-limited RFCC, but CTO-limited RFCC.

326

. - _ ~..

:

-

,1 I'I'

Figure 3. Effect of activity on delta coke.

CAT TO OIL REGIME

" "

J I

I

I

I

"

I

DELTA COKE

Figure 4. RFCC operating regimes.

I Z 0 0 s

(0 0 II-0

" "

rn

/

IL R EG I

M

~

Z

-

I

l MAX. RGT

DELTA COKE '

HEAT

'

MA x . R G T

REMOVAL

Figure 5. CTO regime and maximum delta coke.

327

UNIT A Z

UNIT B

74

O i n,' LLI > Z O O

72

7O s 68

/

r

/,

p/ 1/ I

I

s

EQUILIBRIUM

I

I

CATALYST

I

I

I

ACTIVITY

Figure 6. Commercial RFCC operating regimes. Two examples of commercial cases are shown in Figure 6. The two regimes can be encountered in a single unit depending on the (equilibrium) catalyst activity, as in unit A. Obviously, this makes it very difficult to decide which catalyst to select if a unit is operating in the transition zone between the two regimes.

3.4. Possible explanations for the changing regimes It is well known that too high a catalyst activity can lead to overcracking and excessive gas and coke formation. The decrease in (bottoms) conversion is, however, a relatively new phenomenon and seems to be related to the processing of heavier feedstocks. One possible explanation is the occurrence of concarbon residue (CCR) poisoning. When poor-quality residual feedstocks are processed, part of the increasing CCR will be deposited quite instantaneously on the catalyst flowing into the reactor riser. The lower the CTO, the higher this delta coke by CCR will be, resulting in a larger drop of the initial activity of the catalyst in the riser (Figure 7). This effect can be further aggravated by the fact that the fast deposit of CCR delta coke will tend to result in pore mouth blocking and plugging (Figure 8). RFCC operations at higher CTO ratios will result in a dilution of the reversible (regenerable) catalyst poisons like coke and nitrogen. A second factor which needs to be considered is the fact that in practice catalyst activity in RFCC is o~en boosted by increasing the zeolite activity and/or stability (metal traps). While the additional zeolite sites will contribute to more VGO cracking activity and an increased delta coke, they may not do that much as far as cracking the large hydrocarbon molecules is concerned. In fact an increase in zeolite activity because of the higher delta coke, and hence lower CTO, may result in a drop in the concentration of accessible "matrix" sites: Accessible sites per oil weight

=

Accessible site per catalyst weight

x CTO

(1)

328

CATALYST ACTIVITY

CONCARBON

POISONING CCR/CTO I

TIME IN RISER I

I

I

I

Figure 7. Activity poisoning by feed CCR coke.

Y

Pore M o u t h Plugging

Active Surface P o i s o n i n g

S m a l l d r o p in surface activity Big c h a n g e in ( p o r e ) s e l e c t i v i t y

Large drop in surface activity S m a l l drop in ( p o r e ) s e l e c t i v i t y

Figure 8. Fouling profiles in FCC. The presence of poisons like nitrogen or coke will make the situation even worse as the poisons and the large hydrocarbons will compete to occupy the most accessible catalyst. This is illustrated in Figure 9 as a "supply and demand model." Another result is that the effect of poisoning will be greater at lower CTO ratios [5,6]. DEMAND HYDRO

SUPPLY

C A R BO N S ,

SITES I

POISONS

I

I

I

I

I

(V,Ni,Na,N..)

I

I

I

I

I

I

l

I

I

I

l

I

I

I

I

I

I

I

I

I

HIGH

LOW

AC C ESSIB

ILITY

Figure 9. Supply and demand model of cracking.

329

I.--

////// / /S//

z 0 o (n

:S o v. 1 I0 m < l.J uJ 0 0

//

,""

/

J

/ ~ 0

.

-

____.

:

HIGH.METALS

LOW METALS

t

=

r

[

2

4

6

8

10

CCR, WT%

Figure 10. CCR effect on bottoms conversion. We can support the foregoing by evaluating two catalysts differing in active site accessibility. The delta in bottoms conversion increases with higher feed CCR, higher metal levels, and lower CTO ratios. The catalyst with the lowest number of accessible sites is most sensitive to coke and metal poisoning (Figure 10). From equation (1) there are two possible solutions to the problem: 1.Increase the accessibility of the active sites, and thus the number of accessible sites per catalyst weight; 2.Increase the CTO, and hence reduce the delta coke of the catalyst. Both options will be discussed in the following sections. 4. CATALYST SELECTION FOR ACTIVITY-LIMITED RFCC Catalyst activity, bottoms cracking, and gas selectivity will be essential for the activity of RFCC operation. As catalyst screening by pore volume metals impregnation and steaming can give misleading results [5,7], the more true-to-life cyclic deactivation method should be used. Considering the importance of the active site accessibility as discussed in the foregoing section and references [5,6], the selection of the proper feedstock will also be crucial for testing the catalysts [2,5]. Recently, a lot of attention has been given to the development of more vanadium-resistant catalyst and vanadium traps. We have found that the accessibility of these traps can be even more important than the quantity and/or strength of the trapping sites. To evaluate the effect of accessibility, we investigated the zeolite retention (in % micropore volume)of the catalyst given in table 1. The catalysts were impregnated with 5000 ppm V by the traditional Mitchel pore volume impregnation method and by the cyclic deactivation method. With the pore volume method (PV) the vanadium is distributed homogeneously over the catalyst. With the cyclic deactivation method (CD), the vanadium profile over the particle is as in commercial practice.

330 Table 1. List of catalysts investigated for zeolite retention.

Zeolite

Metal trap

Accessibility

A- 1 A-2 A-3

low R E 2 O 3 - Y low R E 2 O 3 - Y low R E 2 0 3 - Y

matrix - 1 matrix - 2 matrix- 3

base base base

B-1 B-2 B-3

low R E 2 0 3 - Y low R E 2 O 3 - Y high RE203 - Y

dedicated trap dedicated trap dedicated trap

base base base

C-1 C-2 C-3 C-4

low R E 2 O low R E 2 O low R E 2 O high R E 2 0 3

3 3 3 -

Y Y Y Y

matrixmatrixmatrixmatrix-

3 3 3 3

high high high high

Figure 11 shows that a high-accessibility system will give the best zeolite protection when evaluated by the realistic cyclic deactivation method. This has been confirmed in commercial operations (Figure 12). The FCC catalyst ability to rapidly deactivate the deposited metals will be an important factor in resid cracking.

MORE ACCESSIBLE 80

TRAPS

C-4

9

60

/c-z

A-3 O 40

BASE

O A-2

MORE

(STRONGER)

TRAPS

0 A-1

20

I

I

I

40

I

I

$0

% Y RETENTION,

5000 PPM V

BY PVMITCHELMETHOD

Figure 11. Methods for testing vanadium traps.

I 80

331 13 G.

rn

I+

"-s,,.

3000-5000

PPM

VANADIUM

12

ID (/)

~ 11 0 I-p. 0 m 10

9

62

I

I

I

I

64

66

68

70

CONVERSION, A

BASE

--*--

"~ 72

WT%

IMPROVED

ACCESSIBILITY

Figure 12. Commercial vanadium resistance. Vanadium interacts with nickel in a manner which inhibits the deactivation behavior of nickel. Metals-resistant catalysts must therefore be evaluated in the presence of both nickel and vanadium. Also, the mobility of vanadium is reduced in the presence of nickel. In general, cyclic deactivation will be the preferred deactivation method in order to simulate the actual metal distribution and interactions on the catalyst and the correct metal age distribution. Furthermore, the presence of SOx during the regeneration stage seems to be essential as the SOx in the regenerator flue gas competes with vanadium oxide in the reaction with certain compounds to nonmobile vanadate species. There is only a limited amount of information on the deactivation mechanisms and rates of vanadium and nickel migration. The formation of metal silicates and/or aluminates has been proposed, as they seem to form more easily by reduction and oxidation cycles. Rajagopalan et al. [8] confirm that methods involving cyclic redox aging of metals in the presence of sulfur are needed for screening metals-tolerant catalyst. They propose a cyclic test (the cyclic propylene steam method), which addresses the redox aging of the metal, but not the nonuniform laydown and age distribution of metals on the catalyst. We feel that it is critical to also simulate the metals profile over the catalyst, because of the diffusion-limited progressive shell penetration of the metal deposits in real FCC conditions. Catalysts with a more accessible metal-trapping function will perform better under these conditions. Recently, the application of a nickel-tolerant shell-coated FCC catalyst has been suggested[9]. The idea is to introduce an inert silica-rich surface shell coating. As the large molecules carrying nickel and vanadium will only penetrate the outer shell of the catalyst particle, the nickel which is then deposited in this silica-rich region will be poorly dispersed and the absence of an alumina surface to activate the nickel will result in low dehydrogenation activity. A potential drawback of this approach is that the larger hydrocarbons cannot penetrate the catalyst deep enough to reach the active cracking sites and are hence not effectively converted.

332 The target should be to limit the dehydrogenation activity of the nickel without upsetting the conversion of large hydrocarbons in this very important outer shell of the catalyst. In addition, the application of separate vanadium-trapping additives will be less effective, as has been demonstrated in the past [ 10]. 5. CATALYST SELECTION FOR DELTA-COKE-LIMITED RFCC 5.1. T y p e s o f delta coke

Commercial delta coke consists of several components [4,11,12], namely: reaction or catalytic coke feed conradson carbon residue (CCR) coke adsorbed hydrocarbons, which in the case of extended contact time will be converted to soaking time coke [ 11 ] - hydrocarbons trapped in the catalyst by poor blocking etc. - hydrocarbons entrained in the interstitial spaces -

As reported by Ho [12], the types of delta coke formed in Resid FCC can be classified based on the length of time needed for their formation. CCR coke will form nearly instantaneously at the inlet of the reactor and is therefore also called "entrance coke." The second type of coke is formed by the adsorption of highly aromatic and basic materials on even weakly acidic surfaces; this process also occurs quite rapidly. Finally, reaction or catalytic coke will form in what is clearly the slowest coke formation process. Consequently, as illustrated in Figure 13, the relative importance of the nonreaction delta coke components will increase with operations with a short contact time operations. In order to correctly evaluate the delta coke of a catalyst, we need to distinguish between reaction and nonreaction delta coke. In what follows we will use the terms "hard" and "soft" delta coke. "Hard" delta coke is the delta coke measured after a long period of ideal stripping. "Hard" Delta Coke

=

Reaction Coke + Feed CCR Coke

"Soft" delta coke is the difference between total delta coke and "hard" delta coke: "Soft" Delta Coke

Adsorbed Hydrocarbons + Trapped Hydrocarbons + Hydrocarbons entrained in interstitial spaces

I< ~

O z

A

O Ill

0 0

CONTACT

TIME

(SEC.)

Figure 13. Hard coke and soft coke versus contact time.

NS

333

Table 2. Recent improvements in FCC catalysts. Hard delta coke (relative)

Catalyst type REY zeolite

1970s

100

USY zeolite

1980s

75

Modified USY's + metal traps

1990

55

State-of-the-art

1995

40-45

5.2. Reduction of hard delta coke In general the main emphasis and progress in the development of low-delta-coke Resid FCC catalysts has been in the reduction of reaction coke [4,5,10]. Table 2 gives an impression of the improvements which have been obtained in recent years. According to several researchers [ 1,12,13], a reduction in the fraction of the feed CCR which is converted to CCR delta coke is possible by an increase in the feed-catalyst-reactor mix temperature (Figure 14). Ho [ 12] shows that specific coke yield (*) increases slightly with temperature when cracking VGO (CCR = 0.27 wt%), whereas the specific coke drops significantly as temperature increases when cracking a Taching Resid (CCR = 7.0 wt%). Clearly a different, thermal-cracking type of mechanism is involved. Recent research by Moore et al. [14] shows that CCR coke varies significantly with the composition of the crude. Regression of the data from this study shows that if "additive" or CCR coke is considered to be proportional to the measured CCR, the coefficient can vary from 0.58 to 1.0, depending on the crude source. It seems logical to assume that the fraction of CCR converted to coke should also vary with the catalyst used. 7O

s "6 uJ

0

i s

J

OSO

s

0

s

""

...

-"

~ S

~

-~

.

-"

""

9

7 f

~ lg

J

s

""

INCREASED

~'-

.."

I

R E A C T O

O3O

o ~ A R

E A S T

B R E N T

A R A B

I FCC

L I G H T

I FEED

RESlD

I FAC

TO

R

Figure 14. Coke from feed Concarbon residue.

* Defined as: coke yield x conversion/100 - conversion

R

T E M

P~

334 As far as catalyst design is concerned, results seem to indicate that the specific coke caused by CCR will be higher for zeolite cracking than for matrix-type cracking. The foregoing adds support to our earlier statement that it is essential to test catalysts with a representative resid feedstock in order to obtain a realistic assessment of the delta coke of the catalyst. 5.3. Reduction of soft delta coke The hydrocarbons which are entrained or adsorbed by the FCC catalyst and are not stripped off before the catalyst enters the regenerator will clearly contribute to the total delta coke. Fast and effective stripping of the catalyst will therefore be essential in order to minimize the sott delta coke (Figure 15).

We have devebped a strippWiQ tcst w M cm be pedmned dwbg the stripping stage of the mbdy&&advation in a cydic d u c h t h unit [5,7J. F i 16 gives an crumple of a stripOriag amre. The initial 'da delta &is crtarlatod by subwetiag the "bardaamlytk time (nonideal stripping). d m fiom the total wke rffa r short

Figure 16. SoIt delta coke and stripping rate.

335 The "hard" delta coke is defined as the delta coke alter a long period of intense stripping. A (first-order) stripping rate can also be defined. Our investigations show that catalyst composition and architecture can have significant effects on the initial quantity of adsorbed hydrocarbons, i.e., soft delta coke, as well as on the stripping rate. The initial soft delta coke increases with zeolite content and the proportion of small-pore matrix systems. This roughly corresponds to the empirical observation that soft delta coke tends to increase with a higher surface area on the deactivated/equilibrated catalyst (Figure 17). 1 1 A

I..-

509 ~ 0

o8

i._o7 006

100

~

i

i

i

110

120

130

140

SURFACE

i , 150

160

170

(m21g)

AREA

Figure 17. Effect of Surface area on soft coke. Note that the "free" pore size distribution of coke catalyst leaving the riser will be different from the regenerated equilibrium, due to selective coking of the smallest mesopores. Stripping rates are remarkably constant, except for higher-accessibility catalyst systems where a doubling of the stripping rate can be observed (Figure 18).

2.5 STANDARD O w I-

IMPROVED

ACCESS.

2

iY z

DOUBLING

E 1.5 a,,

STRIPPPING

OF RATE

I-

...i w

1

.,

Ix

A

A

I

I

I

I

I

I

I

I

J-1

J-2

J-3

J-4

J-5

J-6

J-7

J-8

CATALYST

Figure 18. Effect of Accessibility on stripping.

336

STANDARD Cs

~

J~

~ I M~P R O1 / ~ I ~ V ED ~~,,~J ACCESS.

S M A L L PORE SA I :REDUCED HYDROCARBON ADSORPTION II : FASTER STRIPPING

Figure 19. Catalyst effect on delta coke. From the foregoing we can expect that a resid catalyst based on a moderate zeolite content in a more highly accessible large-pore matrix system will have a double benefit for the reduction of sott delta coke, because the quantity of adsorbed hydrocarbons will be lower and the stripping rate will be higher (Figure 19). In summary, the quantity of sott coke seems to increase with the surface area in the smallpore range (zeolite and matrix), while the stripping rate is determined inversely by the accessibility of the catalyst sites and increases with larger and nonconstrained pore systems. We can conclude that for delta coke limited RFCC catalyst selection it will be essential to assess the diferences in all the factors contributing to commercial delta coke. 5. REFERENCES

1. J.L. Mauleon and J.B. Sigaud, Characterization and Selection of Heavy Feeds for Upgrading through FCC. 23rd WPC Houston, John Wiley & Sons Ltd,1987. 2. F.H.H. Khouw, M.J.R.C. Nieskens, M.J.H. Borley, and K.H.W. Roebschlaeger, The Shell Residue FCC Process: Commercial Experiences and Future Developments. NPRA Annual Meeting, 25-27th March 1990, paper AM-9D-42, 1990. 3. M.M. Mitchell Jr., J.F. Hoffman, and H.F. Moore, in FCC Science and Technology, J.S. Majee and M.M. Mitchell Jr. (Edts), Studies in Surface Science and Catalysis, Elsevier Science Publishing Co., Amsterdam. Vol. 76 (1993) 293. 4. P. O'Connor, A.W. Gevers, A.P. Humphries, L.A. Gerritsen, and P.H. Desai, in Fluid Catalytic Cracking II, M.L. Occelli (Edt), ACS Symposium Series No. 452, 1991, p. 318. 5. P. O'Connor, A.C. Pouwels, and J.R. Wilcox. "Evaluation of Resid FCC Catalysts." Symposium on Catalytic Cracking of Heavy Oils, 1992 AIChE Annual Meeting, 1-6 November 1992, paper 242E. 6. P. O'Connor and A.P. Humphries, American Chem. Soc. Div. Petr. Chem. Preprints, 38(3)(1993)598. 7. L.A. Gerritsen, H.N.J. Wijngaards, J. Verwoert, and P. O'Connor, Catalysis Today, 11 (1991)61.

337 8. K.R. Rajagopalan, W.C. Cheng, W. Suarez, and C.C. Wear. "Resid FCC Catalyst Technology: Today and Future." 1993 NPRA Annual Meeting, paper AM-93-53, March 1993. 9. D.M. Stockwell, W.M. Jaglowski, and G.S. Koemer. Symposium on Catalytic Cracking of Heavy Oils. 1992 AIChE Annual Meeting, paper 242C, 1-6 November 1992. 10. P. O'Connor, L.A. Gerritsen, J.R. Pearce, P.H. Desai, A.P. Humphries, and S.J. Yanik. "Catalyst Development in Resid FCC." 1991 Akzo Catalysts Symposium, June 1991, Scheveningen. 11. S.J. Yanik, P. O'Connor, D.H. Abner, and M.C. Friedrich. "FCC Catalyst Pore Architecture and Performance." 1991 AIChE Annual Meeting, 18-20 November 1992. 12. T.C. Ho, "Study of Coke Formation in Resid Catalytic Cracking." Ind. Eng. Chem. Res. 31 (1992) 2281. 13. Hydrocarbon Processing, September 1987, pg. 166. 14. H.F. Moore, T.L. Goolsby, S.L. Mago, E. Chao, and M.M. Mitchell Jr. "Catalytic Cracking of Residual Fractions." Symposium on Catalytic Cracking of Heavy Oils, 1992 AIChE Annual Meeting, 1-6 November.

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Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All rights reserved.

339

NOVEL FCC CATALYST SYSTEMS FOR RESID PROCESSING Ulrich A i k e m a d e

a

and Steve Paloumbis b

"Manager Catalyst Evaluation, GRACE Davison FCC Europe b Director Sales & Marketing, GRACE Davison FCC Europe GRACE GmbH, In der Hollerhecke 1, Postfach 1445, D-67545, Germany 1. ABSTRACT Changing economic scenarios and available processing options often compel a refiner to pursue resid processing. Due to the varied properties of resid feeds, the refiner must carefully consider the choice of available FCC catalyst technology. This paper reviews novel matrix and zeolite technologies for resid processing applications to obtain better coke selectivity, gas selectivity and bottoms upgrading. Commercial experience and mechanism of separate particle vanadium traps to control vanadium deactivation is also reviewed. 2. INTRODUCTION A recent "GRACE" survey of the European refining industry showed that over 40 percent of European refiners add various amounts of resid to their FCC unit feeds. The term "Resid" covers a broad range of feeds boiling above 350~ such as Long Resid or Atmospheric Tower Bottoms, Short Resid or Vacuum Tower Bottoms, Heavy Vacuum Gas Oil, Aromatic Extracts and Deasphalted Oil. Such heavy feeds differ from gas oil primarily by their much higher boiling range (only partly vaporized at 560~ and their higher content in polynuclear naphthenes and aromatics, resins, asphaltenes, contaminant metals (V, Ni, Fe, Cu), sulphur, nitrogen, and Conradson carbon. Most resid feeds contain molecules having carbon numbers above 3 5 and size between 10 and 25A depending on aromaticity and degree of branching. Vacuum resids in particular are known to contain molecules with molecular weights between 1000 and 100,000 and size up to 150A. The characteristics of the resid portion can vary widely as shown in Table 1. Some resids derived from paraffinic/sweet crudes are readily cracked in an FCCU with minimum coke penalty. However, most resids contain high levels of CCR, sulphur, nitrogen and metals, and require in addition to adjusted processing conditions an optimized catalyst matrix design. Nickel has considerable dehydrogenation activity, which can be reduced by specially designed Selective Active Matrices made with reactive aluminas that react with nickel thus rendering it inactive for dehydrogeneration reactions.

a Corresponding author

340 Table 1. Typical Range of Properties for Resid Components in FCC Feeds. Nickel Vanadium Sulfur CCR Specific Gravity

ppm ppm wt.% wt.%

0.5-50 0.5 - 150 0.1-3.5 0.5-15 0.84-1.0

Vanadium, while not the only contributor to fluid cracking catalyst (FCC) deactivation, frequently dictates the amount of fresh catalyst added to the FCC unit to maintain activity. Improvements have been made to both zeolites and matrices to minimize the effect of vanadium [ 1]. Another method of protecting the catalyst from vanadium deactivation is to use traps that prevent the vanadium from contacting the catalyst in the first place. Vanadium traps have frequently shown more promise in laboratory testing than has been realized commercially[2,3]. Sulfur, present in commercial operations, has been known to interfere with previous traps' ability to capture vanadium. Recently it has been shown vanadium traps can be designed to perform successfully under commercial conditions. This paper discusses newly developed GRACE Davison catalyst technologies that are designed to crack heavy feedstocks more selectively. 3. CATALYST DESIGN FOR HEAVY FEEDS The typical effects of adding heavy aromatic and metals contaminated resids to the normal VGO feedstock at constant riser outlet temperature are: 9 Reduced conversion due to lower cat/oil ratio resulting from higher delta coke (higher regenerator temperatures). 9 Higher dry gas yields due to feed quality and higher temperature of the regenerated catalyst at the bottom of the riser. 9 Lower gasoline yield due to loss in conversion and increased feed aromaticity. 9 Higher cycle oil yields due to loss in conversion. 9 Possible increase in gasoline octane primarily due to feed aromaticity and more thermal cracking reactions. The loss in conversion is also partly caused by lower "effective catalyst activity" in the riser as a result of increased coke blockage of the catalyst pores with coke and higher vanadium and hydrothermal deactivation of the catalyst. The negative effects of resid processing on FCC yields can be reduced by adjusting the FCC process conditions (lower feed preheat, increased catalyst make-up, increased steam dispersion and stripping) and by the use of FCC catalyst formulations more suitable to such applications. 3.1 Zeolite Selection

It is generally accepted that the most suitable zeolites for resid processing as well as maximum octane-barrel applications are of the RE-USY type. The rare-earth exchange/ stabilization is tailored to match the FCC unit's LPG/Gasoline quality requirements. For the

341 lowest possible zeolitic coke selectivity, the rare-earth exchange of the zeolite should be such as to lead to an equilibrated unit cell size in the range of 24.27 - 24.30 A. High concentrations of ultrastable zeolite are necessary (more than 30 %) in order to provide sufficient activity under resid processing conditions (high metals, high regenerator temperatures).

MICROSTRUCTURE Faujasite-Type Zeolite

C~176

' ',r,:

! / ! (~)~J~ ~

~

6 Microns

/

P~ Binder '

MACROSTRUCTURE

65 Microns (avg.)

Figure 1. FCC Catalyst Components for Heavy Oil/Resid Processing.

3.2 Matrix Selection

Most resid catalysts have medium to high activity matrices with a high percentage of large pores. (Figure 1). The selection of the appropriate amount and quality of matrix activity (acid site strength and density), pore volume, and pore size distribution of the matrix are key criteria for resid catalyst selection. The selected matrix formulation depends heavily on feedstock characteristics such as aromaticity, concarbon, nitrogen and metals. Furthermore, the selection of catalyst has to take into account the optimal Z/M ratio for low dry gas and coke selectivity as well as a low SA/K number[l].

3.3 New Matrix Designs for Resid Processing

The degree of selective cracking of heavy hydrocarbons to useful liquid products determines the profitability of processing residue or generally heavy feeds.As far as the matrix is concerned, this implies a matrix that cracks the bottoms with minimum coke and gas penalties. The most desirable matrix for such application is one that not only has intrinsic bottoms upgrading ability but at the same time provides resistance to nickel and vanadium as well as eliminates hydrocarbon diffusion limitations by customized pore structures. GRACE Davison has developed new matrix technologies utilizing special Structured Reactive Aluminas (SRA). These SRA components are chemically reacted with the proprietary GRACE aluminasol active binder system leading to Selective Active Matrix (SAM) systems with unique properties (Table 2). SPECTRA-400 series and the new ULTIMA-400 catalyst family utilize these matrix systems.

342 Table 2. ,,SAM" matrices: Produced by the chemical reaction of specially formulated Structured Reactive Alumina (SRA) with the GRACE Davison Alumina-Sol binder system. Al.(OH)b(H20)c + aLxOy(OH)z

(A1)d(O)c(OH)f+ H20 T. time

AI-Sol

SRA

SAM

Depending on type of SRA component, different SAM matrices can be formulated SAM-XYZ

~ X YZ

= type of SRA = amount of SRA

Example: SPECTRA-447 ULTIMA-447

~ ~

SAM - 110 SAM - 210

3.4 SAM Matrices Crack Resid With Lower Delta Coke

The ability of an FCC unit to process profitably a heavy feed will mainly depend on the delta coke that results from the feedstock/catalyst combination (Figures 2 and 3). The four types of coke contributing to the overall coke burned in the regenerator have been well described. i) Catalytic coke produced by the cracking reactions on the acid sites of the zeolite and matrix. ii) Contaminant coke produced by the dehydrogenation reactions of metals (Ni, V, Fe, Cu) on the catalyst. iii) Cat/Oil or occluded coke resulting from carryover of hydrocarbons in the catalyst pores and incomplete stripping. iv) Feed residue coke, well correlated with feed Conradson Carbon Residue (CCR). The delta coke strong dependence on feed is illustrated graphically in Fig. 3. Other parameters affecting delta coke are listed in Table 3. Table 3. Parameters Affecting Delta Coke other than Feed Quality. ~,

1. 2. 3. 4.

Reactor T, P Contact time in the reactor Dispersion of feed nozzles Catalyst Design

343 Constant Riser Outlet Constant Coke Operation (Unit at Max Blower Capacity)

~

,00

~

Regen T 0.80 Unit Conversion

0.50 0.30 0.10

C/O Delta Coke, wt.% Increasing Resid content

/

Feed Residue Coke Catalytic Coke

~

Cat/Oil Coke

ontaminant Coke

Decreasing Feed Quality ~ B ~ " Increasing: S.G, Con Carbon, Metals, S, N. Increasing Resid Content

A/P ratio, Enal~nt-'--'-

1)

CatalyticCoke Decreases due to lower effective activity

1. Lower C/O severity (Higher Regen T)

2)

ContaminantCoke (Metals)increases

2. Less effective activity due to metals contamination, coke blockage of pores and higher nitrogen.

3)

Cat/Oilor coke is the same or shows slight increase

4)

Feedresidue Coke (Con Carbon) increases

Lower Conversion by:

Figure 2. Conversion Dependence on Delta Coke.

Figure 3. Conversion Dependence on Delta Coke.

Catalytic Coke can be best reduced by selection ofRE-USY zeolites with an equilibrated UCS in the range 24.27-24.30 A.The selection of Z/M activity ratio will depend on feed composition and process conditions (Figure 4). An active but Coke Selective Matrix (SAM) is better suited for reduction of catalytic coke. Contaminant Coke The use of a low matrix surface area to lower the dispersion of nickel and therefore its dehydrogenation activity is a possible option which, however, is associated with poor intrinsic bottoms upgrading capability. GRACE Davison's Selective Active Matrices (SAM) made with Structured Reactive Aluminas are designed to react with nickel thus rendering it inactive for dehydrogenation reactions. Vanadium "fixation" on the SAM matrix also avoids destruction of the zeolite by hydrolysis of the SiO2/Al203framework by Vanadic acid (HsVO4) and inhibits the formation of Rare-Earth Vanadates which result in cleavage of the RE-O-RE stabilizing bridges in the sodalite cages. The SAM matrix also minimizes the formation of the low melting Na20-V205 eutectic with the zeolite leading to loss of crystallinity. SAM + Nickel(AldOr Porphyrin SAM + Vanadium(AldOo(OH)~ (V-R) + 02 Porphyrin

; (Nickel Aluminate) *SAM) +CO2+H20

(Aluminium)*SAM)+CO2+H20 Vanadate

344

1.4 -

~

Pilot Plant Data 930 F Reactor Temperature

1.3

~

75 Sec. Contact Time

1.2

~1.1 1.o ~ ~. 0.9 ~0.8

N

\ "

"

"n

n, 0.7

,~

0.6 0.5 0.4

ParaffinicFeed 75 LV% Conversion

o.o

I

I

,'.o

Z/M Ratio

Figure 4. Effect of Z/M Activity Ratio on Catalyst Delta Coke.

GRACE Catalysts incorporating SAM matrix technology have exhibited commercially high activity (67-73 MAT) and low coke and gas selectivities with very high levels of Ni+V (8000 -12000 ppm) Occluded or Cat/Oil Coke This coke results from carryover of hydrocarbons adsorbed in the catalyst pores and by incomplete stripping in the stripper. It can be reduced by shifting the pore size distribution to higher values by: i) Increased zeolite mesoporosity: 20 200A pores (use of hydrothermally produced USY zeolite is preferred since it ensures better zeolite mesoporosity) ii) Increased matrix meso- and macropores: >200 A.

The pore size distribution of the catalyst matrix is important for the catalytic performance. The optimal matrix pore size distribution will depend on a balance of mesopores and macropores depending on feedstock quality and reactor conditions (e.g. conventional vs. short contact time riser operation). SAM-technology catalysts (SPECTRA, RESIDCAT, ULTIMA) exhibit different pore size distributions that are matched to various types of feedstock and unit conditions. Figure 5 exhibits typical pore size distribution of SPECTRA944, SPECTRA-444 and ULTIMA-444 catalysts. Since the only differentiating characteristic of these three catalysts is the matrix formulation, the pore size distribution variation is characteristic of the different matrix design: Zeolite % RE203 Matrix

SPECTRA-944 RE-USY SPECTRA-444 RE-USY ULTIMA-444 RE-USY

1.0 1.0 1.0

Modified M-Sol SAM-110 SAM-210

In Table 4 metals free selectivities of an ULTIMA catalyst are given. 4. EVALUATION OF METALLATED CATALYSTS The selectivity improvements of the SAM-200 containing ULTIMA catalysts are especially pronounced when the catalyst is metallated to simulate the equilibrium catalyst conditions in a high metals environment arising from processing of heavy feeds in the FCC Unit. Table 5 summarizes Riser Pilot Plant (DCR) results of a competitive Resid Catalyst versus ULTIMA-445 after Cyclic Metals Impregnation of the catalysts to 5000 ppm Ni+V. The dramatic improvement in the bottoms to coke relationship in the high metals environment is the result of the selective bottoms cracking of the SAM-200 matrix.

345 I~,s-

,1

-~ 0.4

~ B.2

.-7:!0.1

Pore Diameter,

/~ SPECTRA-944

(A)

'~ SPECTRA-444

X ULTIMA-444

(SAM-110)

(SAM-210)

Figure 5. Influence of Selective Active Matrices on Pore Size Distribution (atter AM- 1500 Steaming).

Table 4. ULTIMA Converts Slurry to Useful Products in the FCC Riser* Through High Matrix Activity but Low Matrix Coke.

Catalyst

Competitor A

ULTIMA-443

Fresh Catalyst Activity 76 76 Equilibrium Unit Cell Size 24.27 24.27 Conversion wt.%ff 75.7 76.7 Hydrogen wt.%ff 0.04 0.02 C~+C2 wt.%ff 2.3 2.2 C3+C4 wt. %ff 17.8 18.2 C5 + Naphtha wt.%ff 52.5 53.3 LCO wt.%ff 15.3 15.1 Slurry wt.%ff 9.0 8.2 Coke wt.%ff 3.0 3.0 *Davison Circulating Riser, Reactor Temperature 52 I~ Regenerator Full Bum, Feed Pre-Heat varied, Countrymark feed, 0.9003 g/cc @ 15~ 0.3 wt.% S, 0.53 wt.% ConCarb., 90% Pt. 530~ Metalsfree, steam equilibrated catalysts.

346 Table 5. ULTIMA Shows Better Bottoms Upgrading in the Riser as the Nickel and Vanadium Increases*.

Catalyst

Competitor E

ULTIMA-445

Nickel ppm 2000 2000 Vanadium ppm 3000 3000 Conversion wt.%ff 64.0 70.7 Hydrogen wt.%ff 0.28 0.18 C~+C2 wt.%ff 2.6 2.4 C3+C4 wt.%ff 11.3 14.6 Cs+Naphtha wt.%ff 46.8 50.5 LCO wt.%ff 19.8 17.7 Slurry wt.%ff 16.2 11.6 Coke wt.%ff 3.0 3.0 *Catalysts: CPS (Mettallated and Cyclic) steaming. Test Conditions: Davison Circulating Riser, Reactor Temperature 521~ C, Full Bum Regenerator, Countrymark feed. A large amount of data generated by cracking highly aromatic and metals contaminated FCC feedstocks with ULTIMA catalysts versus a wide range of resid catalysts has shown that the SAM-200 matrix is particulary suitable for selective upgrading of these most difficult feeds. Such an example is shown in Table 6 where a refinery extremely aromatic FCC feed with high Sulfur and Concarbon was cracked with metallated catalysts (5000 ppm Ni+V). The results interpolated at constant coke show the dramatic improvements possible with the new GRACE technology when heavy feeds are processed in the FCC Unit. SAM technology catalysts are already in commercial use and field results confirm what has been consistently observed in a multitude of riser pilot plant and MAT evaluations. 5. VERY HIGH VANADIUM CONTAMINATION The previous examples showed that moderately high metals levels on catalyst are handeled very well by the new matrix systems. For extremely high vanadium levels on catalysts (>6000 ppm), a new material called RV4+ has been developed by GRACE Davison and has been tested successfully in several commercial FCC units. Vanadium reductions on equilibrium catalyst as high as 23.4% were observed with as little as 4.3% material in inventory. RV4+'s affinity for vanadium was as high as six times that of fluid cracking catalyst. Improvements in equilibrium catalyst microactivity were observed that are directly related to higher zeolite surface area, a sure sign that the effects of vanadium were being mitigated. One refiner was able to reduce fresh catalyst additions by 20% and still maintain activity. No sulfur interference was observed during the commercial trials. Refiners can elect to take advantage of this technology in several ways. The most obvious is to process lower cost, higher metals feed or increase the amount ofresid fed to the unit. Another option is to reduce fresh catalyst additions. Cost savings range from hundreds of thousands of dollars per year to several million depending on feed rate and #/BBL usage

347 Table 6. ULTIMA Catalysts show dramatic yield improvement with highly aromatic and metals contaminated feed*.

Catalyst Competitor A ULTIMA-443 Vanadium ppm 3000 3000 Nickel ppm 2000 2000 Conversion wt.%ff 55.2 59.7 Hydrogen wt. %ff 0.80 0.61 C1+C2 wt.~ 2.5 2.5 C3+C4 wt.%ff 8.8 10.4 C5+ Naphtha wt.%ff 37.1 40.2 LCO wt.%ff 20.5 19.8 Slurry wt.%ff 24.3 20.5 Coke wt.%ff 6.0 6.0 *Catalysts: Cyclic Metals Impregnation and steaming (CPS). Test Conditions: MAT, fixed-bed, 527~ Reactor Temperature, 30 s Contact Time, Aromatic feed, 0.948 g/cc @ 15~ R.I. @ 20~ 1.6501, 3.2 wt.% S, 3.3 wt.% Con. Carbon.

rates. Spent catalyst disposal costs would be decreased as well. Laboratory results of future RV technology showing even greater promise than that tested commercially are also presented.

5.1 Deactivation Mechanism All crude oils contain metals, the most common of which is vanadium. Vanadium is usually associated with organo-metallic compounds found in the higher boiling range fractions. Distillation concentrates the vanadium in the fractions frequently sent to the FCC unit. Vanadium quantitatively deposits on the catalyst, destroys the zeolite and contributes to increased coke and hydrogen yields. Many other factors such as inherent catalyst stability, regenerator conditions, and average catalyst age also play a role in determining the activity of FCC catalyst. However, the dominant role of vanadium is demonstrated by plotting equilibrium microactivity versus vanadium level for the entire industry[6].

5.2 Historical Traps One common type of vanadium trap contains a basic species to react with and neutralize the acidic vanadium compounds. The vanadic acid can react with the basic component of the trap according to the general reaction scheme: 2MeO + 2VO(OH)3

Me2V207+3H20

Compounds that have been proposed to react by this mechanism include barium titanate, calcium titanate, calcium carbonate, strontium titanate and magnesium oxides[8,9,10]. All these basic compounds should theoretically react with vanadic acid and bind it in the trap and have proved effective in laboratory evaluations. However, sulfur competition negatively affects the performance of these traps in commercial units[8,11 ].

348 Sulfur oxides in the FCC regenerator flue gas can react with these alkaline earth metals to form sulfates. On the basis of thermodynamic data, the formation of calcium and barium sulfates is favored over the formation ofvanadates at typical regenerator conditions[11,12]. The other trap materials may or may not be affected by sulfur competition, depending on the SOx concentration and regenerator conditions. In any case, the effect of sulfur competition can not be overlooked when designing effective vanadium traps.

5.3 Integral vs Dual Particle Approach A vanadium trap can either be integral to the catalyst particle or contained in a separate particle. GRACE Davison employs both technologies. Each has advantages and disadvantages and neither has emerged vastly superior to the other in testing to date. Integral traps are closer to the zeolite and may provide better protection in units with low vanadium mobility such as those in partial burn or with low steam partial pressure. However, incorporating the trap in the catalyst particle can change the selectivity of the catalyst and its physical characteristics. Dual particle or separate traps such as RV4+ must have attrition and fluidization properties similar to FCC catalyst. Their advantages are that they do not change the selectivity of the base catalyst and theoretically have a higher capacity for vanadium capture. Performance evaluation of dual particle traps is usually simpler. They can often be isolated from equilibrium catalyst and analyzed for vanadium capture. Confirmation of preferential pick up on integral traps tends to be a bit more qualitative. A disadvantage may be that they are more dependent on vanadium mobility than integral traps.

5.4 Vanadium Mobility Since the effectiveness of a separate particle vanadium trap such as RV4+ depends on the ability of the vanadium to migrate from the catalyst to the trap, a number of laboratory experiments and commercial evaluations were designed to measure vanadium mobility. Vanadium mobility can be discussed in terms of intraparticle mobility, interparticle mobility from the catalyst to the trap, and interparticle mobility from the trap to the catalyst (irreversibility). These three areas are discussed below[6].

5.5 Intraparticle Mobility Time Of Flight Secondary Ion Mass Spectrometry (TOF SIMS) analyses of Ecat and RV4+ from a commercial trial. Show that while the vanadium concentration may be higher on the surface of the particles, vanadium is found throughout the RV4+ particle, not only on the outer surface. The SIMS scan also shows that vanadium is found throughout the catalyst particle as well. This shows that over time, there is intraparticle mobility of vanadium in both catalyst and RV4+ particles[5].

5.6 Interparticle Mobility Fresh RV4+ blended with equilibrium catalyst (90wt.% catalyst/10wt.% RV4+, 50wt.% catalyst/50wt.% trap, and 10wt.% catalyst/90wt.% trap) was steamed by the Cyclic Propylene Steaming (CPS) procedure[6]. During this short steaming time (20 hrs), the RV4+ removed vanadium from the Ecat. This is clear evidence that not only does the vanadium trap pick up metals from the incoming feed, but the trap can also remove "old" mobile metals directly from the Ecat by interparticle migration[6].

349 Table 7. Vanadium Removal by Trap Improves MAT Activity*. Ecat

Conversion C/O

wt.%

55 3.8

Ecat/10% Fresh RV4+ 55 3.1

H2

wt.%

0.21

0.19

Total C1 + C2

wt.~

1.3

1.2

C3='s Total

wt.% wt.%

3.1 3.7

3.1 3.6

C4= Total

wt.% wt.%

3.8 6.8

4.2 6.7

Gasoline

wt.%

39.8

41.0

LCO Bottoms

wt.% wt.%

24.5 20.5

25.6 19.4

Coke

wt.%

3.2

2.4

* 1300~ CPS, 90/10 wt.% Blends

Microactivity testing of the 90% Ecat/10%RV4+ sample compared to a 100% Ecat sample steamed by CPS was also performed. Results in Table 7 show a dramatic improvement in yields and activity. Interparticle mobility is proven by electron microprobe scans of cyclic metal impregnated (CMI)[6] Residcat| 767Z4+ which incorporates RV4+ technology. Since the catalyst and the RV4+ were simultaneously exposed to the metals during the CMI procedure, the rate of deposition of vanadium on the catalyst and trap surfaces should be similar. However, the catalyst particles, contain virtually no detectable vanadium. In contrast, the RV4+ particles containing the Active Trap Component are high in vanadium. This is another indication of particle to particle vanadium mobility[6]. 5.7 Irreversibility

A trap was then blended with 90wt% fresh catalyst and steamed by cyclic propylene steam (CPS) for 20 hours. After steaming, the catalyst and trap were density separated and analyzed for vanadium. Results are presented in Table 8. As shown in the table, less than 6% of the vanadium migrates back to the catalyst. This represents an insignificant amount of the total vanadium transferred. Additionally, since the vanadium on the catalyst may migrate back to the trap over time, the degree of reversibility may actually decrease with time.

350 Table 8 Low V Mobility from Trap to Catalyst

Before CPS

Vanadium (ppm) After CPS

Impregnated Trap 11,350 Fresh Catalyst 50 % Vanadium Migration is less than 6% 1400~ CPS, 90/10 wt.% Blend

10,680 80

5.8 Measuring Performance The ultimate measurement of trap performance is if microactivity increases at constant fresh catalyst additions and metals levels or if the improved stability provides the flexibility to reduce additions or process higher vanadium containing feed. From an evaluation standpoint, it helps to have additional methods of determining success. Dual particle traps can frequently be separated from equilibrium catalyst if their densities are slightly different. The two fractions can then be analyzed for vanadium. If the trap is preferentially picking up vanadium, then it confirms that the technology is working even if there is too little trap in the inventory to improve the microactivity or if another variable is at work reducing microactivity. We have found the ratio of vanadium on the two fractions to be an effective means of confirming trap performance. We refer to this ratio as the Pick-up Factor (PUF) and express it as follows: Pick-up Factor (PUF) =

Vanadium on Trap. ppm Vanadium on Ecat, ppm

Another useful comparison is the amount of vanadium "removed" from the equilibrium catalyst. This is somewhat of a misnomer because it represents not only vanadium that has migrated from the equilibrium catalyst to the trap but vanadium that has deposited directly on the trap. Had the trap not been there, all of the vanadium would have deposited on the equilibrium catalyst so it is in essence the amount of vanadium removed. Mathematically it is expressed as: % V Removed =

(V on Trap ppm) (wt% of Trap in Inv) (Vanadium on Total Blend, ppm)

6. C O M M E R C I A L RESULTS Seven commercial trials have been conducted using RV4+ technology. A wide range of base catalysts, vanadium levels, unit designs and unit operations, including a partial burn operation, were studied. Table 9 summarizes the key results.

351 Table 9. Residcat RV4+ Technology Commercial Results. Trial A B C D E F G

%RV4+ in Inventory 43 2.2 45 3.6 37 4.6 56

% Vanadium 23.4 5.7 15.2 7.9 14.5 12.9 22.3

Pick-up Removed 6.8 2.7 3.8 2.6 4.4 2.9 4.8

Vanadium Associated Factor with RV4+,ppm 7,5O0 13,400 16,400 5,900 12,000 7,200 13,900

The wt.% vanadium removed varied from approximately 5-25% and correlated well with the amount of trap in inventory (Figure 6). In all cases, the targeted amount of RV4+ in inventory was 5%. While much of our laboratory work was done with 10% blends, a 5% blend was chosen for the commercial trials to minimize possible dilution effects. Several units did not attain the 5% level due to previously scheduled turnarounds. In two of the cases where the targeted level was achieved, Trials A and G, vanadium removal exceeded 20%. Interestingly, the partial burn operation, Trial F, was not that much lower than the full burn operations. The amount of RV4+ in inventory is a function of time. It stands to reason that the percentage vanadium removed would also vary with time. Figure 7 illustrates this relationship. In general, for the same number of days on the trap, the unit with the greatest % trap in inventory provided the highest vanadium removal. Taking a closer look at Trial G, the refiner's objective was to run higher metals feed without increasing fresh catalyst additions. Figure 8 tracks vanadium as a function of time. Shortly after the introduction of RV4+ to the unit, the vanadium level increased by over 1,500 ppm. Normally this type of increase would have significantly reduced activity. Instead, microactivity remained relatively stable. This had the effect of redefining the MAT versus vanadium deactivation curve for the unit (Figure 9). The shift to the right or to higher metals levels can be attributed to increased zeolite activity retention. In this case, also the same percentage of zeolite surface area is being retained at 1,500 ppm higher vanadium with constant catalyst additions. This clearly shows that the trap is protecting the zeolite from deactivation[7]. Sulfur competition has been the Achilles' heel of other technologies used to trap vanadium. While RV4+ technology picks up some sulfur, it does not appear to hinder its performance. In fact, its propensity to pick up sulfur diminishes rapidly as its ability to capture vanadium increases, suggesting that the rare earth vanadates formed are more stable than rare earth sulfates. This can be seen in Figure 10. Also evident in Figure 10 is the high amount of vanadium on RV4+, approximately 11,000 ppm. The highest level achieved was in excess of

352

Figure 6. Percent Vanadium Removed From E-Cat Commercial Summary.

Figure 8. RV4+ Improves Vanadium Tolerance Vanadium vs Time-Trial G.

Figure 7.Percent Vanadium Removed From E-Cat Commercial Results.

Figure 9. V4+ Technology Improves Activity Retention-Microactivity vs. Vanadium - Trial G.

16,000 ppm. The theoretical saturation point is several times greater than this. Given more time in the unit and more favorable conditions for vanadium mobility, the trap should continue to pick up vanadium. Figure 11 confirms the trap's ability to capture vanadium long atter the trial ended. However, there may be some factors at play limiting the amount of vanadium the trap can capture. Vanadium level on equilibrium catalyst, average catalyst/trap age, regenerator internals, steam partial pressure, and the amount of excess oxygen in the regenerator are just a few variables that come to mind which may influence trap performance. More commercial data is needed to sort out their respective roles. Figure 11 also shows good unit retention of the trap by the fact that the decay curve appears normal.

353 1.4 1.2

18,000

\ \

_

-15,ooo .~

~'5

5

Q.

?.4

/"

~'0.8 "0

m r

>

g~

0.6

3

0.4 0.2 0

I 20

I 40

I 60

I 80

I 100

2 120

~:2 N

1 0 , 6130193

-12,000 ~

J"

r

§

J//

9,000 >

J

6,000 > Trial

3,000

nded

8130193

11 I15/93

# of Days on Residcat (RV4+)

Figure 10. Vanadium and Sulfateon RV4+ Trial E.

Figure 11. RV4+ Continues Capturing Vanadium After Additions are stopped -Trial C.

It is well known that the MAT versus vanadium deactivation curve is different for different catalysts. Where a refiner is operating on this curve will influence the response to trapping technology. What may not be so obvious is that different unit designs have different deactivation curves and that the mode of regenerator operation also influences the MAT versus vanadium relationship. 7. CONCLUSIONS GRACE Davison has developed "Selective Active Matrix" catalysts (SAM-technology) based on structured Reactive Aluminas on Alumina-Sol binder. These matrices provide unique properties in cracking heavier and metals contaminated FCC feeds with minimum coke and gas yields. The choice of the specific SAM matrix formulation for any application will depend on feed quality and metals contamination as well as unit riser configuration. SPECTRA-400 series employs the novel SAM-100 matrix technology which isbest suited for heavy/resid feeds with low to intermediate aromatics/paraffins ratio and high metals content (V/Ni ratio greater than 2). ULTIMA-400 series employs the new SAM-200 matrix technology which is best suited for heavy/resid feeds with intermediate to high aromatics to paraffins ratio and high metals content (V/Ni ratio less than 2). Both these catalysts are extremely suitable for "Short Contact Time" riser designs where high activity is desired via the combination of high concentration o f l ~ - U S Y zeolite with Selective Active Matrices. Excellent commercial results with SAM matrix catalysts recently obtained in several "Short Contact Riser" FCC units in Europe have confirmed the advantages of this technology[5]. For extremely high vanadium application, RV4+ vanadium trapping technology has commercially demonstrated the ability to reduce vanadium on equilibrium catalyst by more

354 than 20% in a variety of units. Reducing vanadium loading leads to higher microactivity and improved zeolite surface area retention, confirming that RV4+ technology protects zeolites from vanadium deactivation. Sulfur competition, which prevented some previous traps from working commercially, was not a factor. RV4+ technology can save refiners up to several million dollars a year in catalyst costs or allow the option of processing higher vanadium feeds. REFERENCES

Paloumbis et. al, Davison Catalgram - European Edition 1/93 2. R.N. Cimbalo, Oil & Gas Journal 70 (20), 112 (1972). 3 C.C. Wear, Davison Catalagram No. 75, 4, 1987 4. B.B. Agrawel and F.B. Gulati, Petr. Hydrocarbons, 6, 193, (1972) 5 GRACE Davison FCC Technology Conference, Athens, Sept. 27-30, 1994 6. T.J. Dougan, U. Alkemade, B. Lakhanpal, L.T. Book, NPRA Annual Meeting March 2022, 1994 7. T.J. Dougan, U. Alkemade, B. Lakhanpal, L.T. Book, Oil & Gas Journal, September 26, (1994) 81 8. J. Scherzer, Octane Enhancing FCC Catalysts, Marcel Dekker Inc., New York, 1990 9. K.R. Rajagopalan, W.C. Cheng, W. Suarez, C.C. Wear, NPRA Annual Meeting 1993 10. D.J. Rawlence, K. Gosling, L.H. Staal, A.P. Chapple, Preparation of Catalysts V.G. Poncelet, P.A. Jacobs, P. Grange, B. Delmon Ed, Elsevier Science Publishers, Amsterdam 1991,407-419 11. H.L. Occelli Ed., Fluid Cracking: Concepts in Catalyst Design, ACS Washington DC 1991, Vol. 452 12. Lange's Handbook of Chemistry 13th Ed; J.A. Dean Mc Graw Hill, New York 1985 1

S.

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All rights reserved.

355

PROBING INTERNAL STRUCTURES OF FCC CATALYST PARTICLES: FROM PARALLEL BUNDLES TO FRACTALS R. Mann a and U.A. EI-Nafaty b~

a. Department of Chemical Engineering, UMIST, Manchester M60 1QD, UK b. Department of Chemical Engineering, KFUPM, Dhahran 31261, Saudi Arabia. ABSTRACT The phenomena of transport and reaction in the Reactor-Regenerator Cycle of FCC processes rely on the large surface area available in the porous catalyst particles. Most of this area resides in randomly interconnected sub-micron cavities within the particles. A good understanding of configuration of the pore space is essential for reliable modeling of cracking reactions as well as the catalyst deactivation and regeneration. Earlier approaches have utilized mercury porosimetry in conjunction with the Washburn equation to represent internal voids as parallel bundles of cylindrical pores. But the failure of these models to account for important geometric parameters, such as tortuosity, connectivity and morphology expected of the complex porous framework, renders them highly unrealistic. A number of alternative structural configurations have been developed incorporating various degrees of randomness to provide more realistic visualization of the chaotically oriented cavities. Beginning with intersecting versions of the parallel bundles, work has progressed through corrugated cylindrical pores, towards 2-D and 3-D stochastic pore networks and more recently to fractals. In this paper, a qualitative overview of hierarchical developments in pore space representation and quantification will be presented. Emphasis is given to stochastic pore networks and fractal geometrical concepts. 1. INTRODUCTION The fluid catalytic cracking (FCC) process has been one of the most important cornerstones of petroleum refining and at present accounts for nearly 30% of world gasoline production[I]. The heart of the process lies in application of high surface area (>200 m2/g) cracking catalyst particles (30 - 801am) composed of amorphous alumina matrix embedded with small (1-3~m) zeolite crystallytes. Although the FCC technology is now more than fiPty years old, the fast deactivation of the catalyst continues to pose a significant drawback to the overall economy of the process. A mounting body of experimental evidence has pointed to the internal pore structure of the catalyst as being the single most significant aspect affecting the process kinetics both in terms of reactivity and selectivity[2,3]. Not surprisingly, considerable research effort is geared towards investigating and correlating cracking performance with the structural configuration of the pore space within the catalyst particles. Table 1 gives a list of variables that must be incorporated for adequate assessment and representation of void spaces within catalyst particles. Although a comprehensive model that simultaneouly accounts for all these these parameters is yet to be developed, the rapid advancement in computing technology coupled with high quality image processing and characterization techniques, has made possible,

* Corresponding Author

356 Table 1. Variable and random parameters in pore structure modeling 1. 2. 3. 4. 5. 6. 7. 8.

Euclidean dimension(variable). Fractal dimension(variable). Pore length(random). Pore diameter (random). Topology (Pore connectivity) (random). Pore surface morphology (random). Pore cross-sectional shape (random). Tortuosity (variable).

the relaxation of many of the simplifying assumptions made in earlier models as well as developement of more realistic pore structural models. 2. THE CLASSICAL PARALLEL BUNDLE Mercury porosimetry and low temperature gas adsorption(LTGA), are two laboratory techniques commonly used as a means of probing void space within porous particles. The classical methods of analyzing the resultant data (penetration/retraction and adsorption/ desorption curves for the former and latter respectively), represent the void space as a bundle of straight parallel cylindrical pores (Figure 1). However, SEM studies of FCC particles, show, in common with most porous materials, that the pore spaces are an evident entangled mass of widely varying sizes. Pores are thus expected to be randomly jumbled together, but to be interconnected thoroughly amongst one another. Figure 2 shows how an FCC particle of about 70 lam in diameter appears when viewed by an SEM. The parallel bundle of nonintersecting tubes is hence a perfectly unrealistic representation of the complex realties of the porous particle depicted in Figure 2. Subsequent efforts to improve upon it have tended to incorporate either a non-cylindrical assumption or an element of interconnection and intersection[4]. An early modification of the parallel bundle model was described by Androutsopolus and Mann [5] who presented a version of the model termed "series pore model" in which each pore is subdivided into sub-segments of varying sizes and length with the diameters randomly distributed according to some statistical distribution. Although the segment sizes obey the same distribution function, no two pores in the network are identical. Mann and Thompson [6] have also applied the parallel bundle model concept in a modified form to study deactivation kinetics in a supported zeolite cracking catalyst. In their model, the zeolite micropores were assumed to be adjoint to the matrix micro- and meso-pores. Although these improvements to the parallel bundle description encompass variability in pore size and structure, the models are obviously too regular to adequately represent the entangled mass of interconnected pores shown in the SEM figure. The attendant gross over simplification risks serious distortion of the intraparticle transport processes. The greatest challenge in pore space representation is thus to incorporate the elements of randomness and chaos implicit in porous catalyst particles in such a way as to retain both structural realism and tractable quantitative treatment.

357

Figure 2. SEM view of a FCC catalyst particle

358

Figure 3. A simple 30x30 2-D stochastic Network 3. STOCHASTIC PORE NETWORK MODELS A stochastic pore network (SPN) in one in which simple pore segments form interconnecting networks within which pores can be either randomly or partly randomly distributed. Figure 3 shows a typical 30x30 2-D SPN composed of 1860 pores of equal length in which the pore radii obey a uniform distribution in the range 10 A to 4400 A. Such pore networks are meant to provide a more realistic basis for deducing pore structure and hence modeling different kinds of processes within catalysts Although SPNs were originally developed for mercury porosimetry [7], their application in characterization of catalyst pellets [8], low-temperature gas adsorption[9], and diffusion, reaction and coke laydown in FCC catalysts [6] has been clearly demonstrated. A distinctive feature of these models is the capability to account for hysteresis and entrapment characteristic of mercury porosimetry and adsorption/desorption isotherms[10]. Figures 4, 5, and 6 show respectively, predicted LTGA isotherm, mercury porosimetry intrusion/extrusion curves and accompanying mercury entrapment for the network shown in Figure 3. In this example, the predicted mercury entrapment is 45%. Although the network in question incorporates some element of randomness and connectivity, one shortcoming of the network is the uniform length allocated to the pores. Mann et al.[4] have incorporated additional structural variations to the network to include random pattern re-ordering of the pore junctions giving irregular node placements. Another possibility is to allocate lengths to pores of different diameters giving a sub-ordered corrugated feature to the pores. It is thus possible to have a simple regular or irregular network, or a sub-ordered corrugated regular or irregular network. Figure 7 shows the various structural developments of the simple square network.

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Figure 4. Predicted adsorption/desorption isotherm for the network depicted in Figure 3. The rules for pore size allocation, length variation and sub-ordering or reconstruction in 2-D networks, can be extended to three dimensions. A sample visualization of a 3-D 10xl0xl0 network which comprises some 3300 pores is depicted in Figure 8. The irregular network(b) is obtained by a relatively small random relocation of nodes of its equivalent counterpart(a). The use of 3-D model is undoubtedly more realistic than its 2-D network and with current rapid advancement in computing technology, it would soon be possible to use assemblies of millions of pores to not only deduce pore space quantitatively but also to predict, more accurately, several transport and reaction kinetics in FCC process 4. FRACTALS The use of fractal geometry, both deterministic and non-deterministic i.e. (stochastic), to model natural processes has become an intensive research area in recent years. This has extended to include characterization and analysis of the configuration of void spaces within porous materials. Qualitative geometrical analysis have shown a wide variety of natural and synthetic materials ranging from rocks, trees and clouds to charcoal, quartz and aluminas, to posses fractal properties [11,12]. Very often, common natural processes involving diffusion and reaction are found to obey power laws which for most of the time have been described within the domain of Euclidean space and hence restricted to integer powers. Table 2. gives a comparison of the Euclidean and fractal geometries. On the other hand, it is observed that a large number of heterogeneous reactions follow fractional-order kinetics under different process conditions [13]. But most classical transport theories, valid for Euclidean structures, fail when applied to transport processes

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362 in complex and disordered media. The advantage of the fractal geometry lies in the ability to relate properties and processes (both static and dynamic) in terms of non-integer power laws. Fractal geometrical concepts hence offer a potential tool for modeling catalysis and transport in porous media on a more fundamental and realistic basis. Table 2. Comparison of Euclidean and fractal geometries.[ 14]. EUCLIDEAN Traditional (> 2000 yr.) Based of characteristic size or scale Suits man-made structures Described by formulae Integer dimensions

FRACTAL Modem (<20 yr.) No specific size or scale Suitable for natural shapes Described by (recursive) algorithms Non-integer dimensions

4.1. Creation of fractal pores.

A fractal (line, curve, surface or structure), is constructed by an iterative procedure using a 'generator' applied to an 'initiator' (also referred to as 'attractor'). In general, a generator can be any chosen shape which when applied to the initiator, affects a change to it. The resulting fractal will however be deterministic or stochastic depending on whether the generator is fixed or randomly modified with each successive iteration. A typical example of a deterministic fractal that offers reasonable representation of a porous particle is the Menger Sponge (Figure 9). This was constructed by successive removal ofa cuboid (the generator) from the centre of a cubic initiator. The pores in this structure have square cross-sections as a result of the

Figure 9. The menger sponge : a potential fractal representation of pore space within FCC particles. ( fractal dimension = 2.7268) [15]

363

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Figure 10. A hierarchy of stochastic fractal pore development: a. Straight cylindrical constant diameter, b.Straight variable diameter, c. Tourtous variable diameter, d. Tortous non-circular cross-section. generator geometry. Mann and Wasilewski[16], have developed a stochastic fractal pore by successive random modification of the radius, centre line and cross-section of a cylindrical pore as shown in figure 10. With proper manipulation of the variables in the generating algorithm, it would be possible to extend this to construct realistic porous structures that closely represent FCC catalyst particles. It would then be straight forward to model complex transport and reaction processes using partinent fractal concepts. 5. CONCLUSION Pore structure modeling has been significantly refined over the years. Application of stochastic pore networks for quantitative deduction of pore space configuration and modeling transport and reaction processes in porous particles, is fairly well developed. The problem of mapping mercury porosimetry and LTGA in fractal geometry however, is still at a development stage~ The most critical step in this respect rests on successful development of suitable cubic or spherical initiators with appropriate generators. Pore space can then be generated by controlling the randomized modification of the generator as well as the recursion levels in the iterative algorithms to map experimentally measured pore structure and morphology. It is expected that calculation of FCC catalyst pore structure by a combination of SEM images,

364 3-D SPNs, and stochastic fractal geometry, would enable us to dispense with indirect laboratory measurements which require expensive instruments. ACKNOWLEDGEMENTS. The authors wish to thank King Fahad University of Petroleum and Minerals, for partial support and MCC UMIST for scanning facilities. REFERENCES

F. A. Zenz, Encyc. Chem. Tech., R.E. Kirk and D.F. Othmer(Eds.), 3rd Ed., 10 (1980) 548. 2. A.Wheeler, Adv. Catalysis, 3 (1951) 250. 3. J.J. Carberry, Chem. Eng. Sc., 17 (1962) 675. 4. R. Mann, J.J. Almeida, and M.N. Mugerwa., Chem. Eng. Sci. 41 (1986) 2663 5. G.P. Androutsopoulus., and R. Mann., Chem. Eng. Sc., 33 (1978) 673 6. R. Mann, and G. Thomson, Chem. Eng. Sc., 42 (1987) 555. 7. G.P. Androutsopoulus, and R. Mann, Chem. Eng. Sc., 34 (1979) 1203 8. R. Mann, and H. Golshan, Chem. Eng. Commun., 12 (1981) 377. 9 R. Mann, and G. Thomson, Ads. Sc. & Tech., Rodriquez, E.A., and Nijhoff, K. (Eds.) (1989) 63 10. R. Mann, and H. N. S. Yousef, Adsorption Sc. & Tech., 8(4), (1991) 196, 11. D. Avnir, and D. Farin, Nature, 303, (1984) 261 12. C. E. Krohn, and A. H. Thomson, Phys. Rev., B33, (1986) 6366. 13. B. Sapoval, In "Fractals and Disordered System.", A. Bunde and S. Havlin (Eds.), Springer-Verlag, Berlin Heidelberg, (1991) 207. 14 R.F. Voss, In "The science offractal images", Peitgen, H.-O., and Saupe, D. (Eds.), Springer-Verlag, New York Inc., (1988) 21. 15. B. B. Mandelbrot, Thefractal geometry of nature, Freeman, N.Y., (1983) 16. R. Mann, and M. C. Wasilewski, Trans. I. Chem. E. 68A (199o) 177.

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All fights reserved.

365

D E V E L O P M E N T OF M I C R O SCALE ACTIVITY TEST STRATEGY F O R F C C PROCESS ECONOMICS ENHANCEMENT O. H. J. Muhammad

Petroleum Technology Department, Kuwait Institute for Scientific Research, P.O.Box 24885, 13109 Safat, Kuwait. ABSTRACT The fluid catalytic cracking (FCC) unit is a complex but important refining process which has great potentials in enhancing process benefits. Any improvement, whether in design, operation, control or catalyst will result in a substantial reward in process economics and hence FCC research interests are worldwide[l]. There are several ways to achieve these goals. In this contribution, strategies in using the micro scale activity test (MSAT)rig to optimize the catalyst selectivity and equilibrium catalyst system due to catalyst deactivation are developed. Since only the coking effects on catalysts are investigated in the present study, the cracking of cumene over supported zeolite catalyst is employed in order to avoid complications from other deactivation agents such as metals and sulphur. It is found that the catalyst selectivity and the equilibrium catalyst properties are highly interactive. I. I N T R O D U C T I O N The FCC process is used to convert straight-run atmospheric gas oil, certain atmospheric residues and heavy stocks recovered from other refining operations into high octane and low sulphur gasoline, light gas oils and olefin-rich light gases over zeolite type of catalysts[2]. Zeolites are a class of crystalline alumino silicates. These materials have high surface area to enhanced catalytic activity. Their matrix structures act as "molecular sieves" by allowing smaller size molecules to get through but excluding those larger ones and hence adding the "selectivity" capability. These extraordinary properties of zeolites make them the most important cracking catalysts[3,4]. However, due to the continuous degrading of crude oil quality and a general shift in petroleum demand towards heavy distillates, there is a need for a more effective but flexible operation in all the refining processes in general and the FCC units in particular, to maintain profit and abide by the increasingly stringent environmental regulations. There are number of ways to maximize the profits and the principal benefits of any refinery conversion process can be realized through: 1. 2. 3. 4.

to increase product value through improved selectivity[5] to increase unit processing capacity to be able to crack heavy, and high metal and high sulphur content feeds economically[6] to have a suitable process control system capable of tighter control but without violating product specifications[7-8] 5. to reduce catalyst consumption[9] 6. to extend the process time-on-stream 7. to employ additives[10]

366 The benefit estimation of producing more valuable products can be easily determined by the delta between the products unit price. For example, a 30,000 BPD FCC unit producing 4,200 BPD heavy naphtha, for a 1% increase in this product from the less expensive light naphtha at a delta price of $2/bbl[11 ] can increase the revenue by $30,000/year. In a recent publication[9], a saving of $131,000/year in catalyst expenditure has been reported for a small 6,000 BPD FCC unit due to the reduction of 0.11b catalyst/bbl feed processed. This is a continuous research program aiming to develop an experimental strategy using a micro scale activity test (MSAT) equipment to improve FCC operation economics through better understanding of catalysts and their interactions with the process hardware and operation conditions. This is different than the commonly used micro activity test (MAT) for the evaluation of the FCC catalyst, in terms of the residence time and hydrodynamics. In the present study, an operation strategy to optimize FCC unit operation through better selectivity and to minimize catalyst consumption due to catalyst coking are investigated using a MSAT equipment. The cracking of cumene over a commercial zeolite catalyst is used as a model reaction to avoid the obscurity from other deactivations due to sulphur, vanadium or nickel. Cumene (isopropyl benzene) has been used to study the activity of numerous cracking catalysts in the past years. Many kinetic studies of cumene cracking have suggested the reaction taking place at a single site via a carbonium-ion mechanism and the products are mainly benzene and propylene[ 12], while propylene has been suggested as the coke precursor[13,14]. Furthermore, the nature of the disproportionation of cumene can represent the cracking of typical commercial FCC feeds such as gas-oils via the dealkylation of branched aromatics. 2. EXPERIMENTAL W O R K

2.1 Catalyst and Feed Specifications The catalyst used is a commercial catalyst known as the super-D manufactured by Crosfield Chemicals Ltd., UK. It is in the form of particulate spheroid with an average diameter of 81 microns and consists of 15-18% ion exchanged Re sodium Y-zeolites on a support silica-alumina matrix. Heat treatment of catalyst particles at 150oc for 48 hours is undertaken before cracking reaction commenced. The isopropyl benzene (cumene)has the purity higher than 99.5% which was supplied by Fissons Scientific Apparatus.

2.2 Experimental Apparatus Figure 1 shows the schematic diagram of the experimental apparatus. The catalytic cracking of cumene was carried out in a bench scale fluidized bed reactor made of quartz glass. The reactor could hold 5 to 10 grams of catalyst. Nitrogen gas was used to purge the reactor before and after the cracking experiments. The feed was pumped to the waste-line initially and it would not be admitted to the reactor before a steady flow was achieved. This was determined by obtaining a constant weight of cumene discharged from the waste-line at several fixed intervals. The feed line and the reactor were heated electrically to the required temperature which was then maintained by a temperature controller to the nearest +1 oc. The reactor effluent, which included the products and the unconverted cumene, was condensed and collected periodically for analysis. The experimental parameters and their ranges are: 9 Reaction temperature 475 - 525 oc.

367

9 9 9 9

Catalyst loading Feed flowrates Catalyst to Feed Ratio Time on stream

-

50.1 1:1 1-

10 gm. - 5 gm/min. - 100:1 gm-cat:gm-feed/min. 120 min.

2.3 Experimental Procedures and Analysis Before undertaking the cracking experiments, several 'blank' runs at different temperatures were carried out with no catalyst in the reactor to ensure only negligible conversion from thermal cracking in this apparatus. The maximum conversion found was about 2% which was in good agreement with the previous works undertaken in both fixed and fluidized-bed reactor with similar type of zeolite catalysts[ 15-17]. In addition, the mass balance of cumene over the entire reaction system was undertaken to ensure no leakage as well as to form a base case for the determination of cracking conversion in the later experiments. The liquid effluents were analysed by a Pye Unicam gas liquid chromatograph (GLC) containing a silicon OV17 stationary phase separation column under a specified temperature program. The chromatograms were quantified in terms of wt% by a Trivector Trilab 2000 integrator. A typical chromatogram of a known standard is given in Figure 2 which did not show any base-line drilt. The separation of all the species obtained from the product stream was excellent. The deactivated catalyst recovered from the reactor alter each run was analysed for its coke content using a LECO CS244 carbon/sulphur analyzer. The total surface area of the fresh and spent catalysts were measured using a Quantasorb Sortometer in the Catalyst Characterization Laboratories at Kuwait Institute for Scientific Research. The catalyst pore structures were also examined through a scanning electron microscope and images of the fresh and spent catalyst. 3. M O D E L F O R M A T I O N

3.1 Cumene Conversion Cumene conversion was determined from an overall mass balance over the cracking reaction system from both the normal and 'blank' runs.

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369

3.2 Benzene Selectivity Since benzene is the primary product from the cracking of cumene over super-D zeolite catalyst, it is assumed to be the desirable in the present development. The selectivity, S, can be defined as the moles of benzene produced per mole of cumene converted, ie.

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where C F is the catalyst to feed ratio. 4. RESULTS AND DISCUSSIONS

4.1 Cumene Cracking The cumene conversion at 3 levels catalyst to feed ratio, CF: 5, 50 and 100:1 are shown in Figure 3 in which the conversions are plotted against the time on stream. The general characteristics of catalytic activity (conversion) are exhibited in which a very rapid decline in activity occurs initially and then it is followed by a more graduated loss in activity. Since one of the products - propylene was the coke precursor and the benzene selectivity was high[14], the loss in surface area due to coking should be at minimum. This is reflected from the normalized catalyst surface area (%) against time on stream plots as shown in Figure 4. However, the severity in the activity loss also depends on C F. In the case of high catalyst to feed ratio, (C F = 100), the coke which is deposited on the catalyst surface will be spread out over the total amount of catalyst used. Hence, the level of coke content found in the catalyst at high C F is much less than that of lower C v values as shown in Figure 5. On the other hand, when catalyst to feed ratio is at 5, the conversion and loss in surface area are the lowest while the coke content is highest despite of the high benzene selectivity.

4.2 Benzene Selectivity The benzene selectivity expressed as a function of time on stream is given in Figure 6 for two values of C v. It is interested to note that the characteristics from both selectivity curves (C v - 5 and 50) are similar. In either case, there is an initial drop in selectivity. This is followed a period of constant selectivity. Another drop in selectivity is observed thereafter.

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50

60

.".............................

ratio on benzene

selectivity.

l l l l l l l J l

70

80

90

372 As clearly indicated in Figure 6, the benzene selectivity can reach 100% when the conversion of cumene is less than 5 % at C F = 5, while the maximum benzene selectivity at C F = 50 can only achieve 92%. However, this does not provide a clear picture of which catalyst to feed ration to be chosen from, neither the optimal conversion level, nor the status of the equilibrium catalyst. To effectively utilize the information obtained from the experimental work, it is advantageous to re-evaluate the same data based on the normalized benzene selectivity as described by Equation (5). A more useful trends can be seen as given in Table 1, in which the benzene production and normalized benzene selectivity are listed with respect to the cumene conversion for the same two values of C F as in the Figure 6. From columns 2 and 4, the maximum benzene production are clearly shown for C F = 5 and 50 respectively. The best production at low C F is when the cumene conversion is at 60%, while at high CF, this is at 75% conversion. These conversion levels are typical in the FCC unit operation. On its face value, the cumene conversion and benzene selectivity seem to be better at C F = 50 than 5. This is only because more catalyst has been used. In fact, it is 10 times more. The normalized benzene selectivities given in columns 3 and 5 with respect to C F = 5 and 50 provide clearer indication. The physical structures of the catalyst pores and zeolites obtained from electronic scanning microscope analysis are also important. However, they are not reported here since this is out of the scope of the present contribution and can be found elsewhere[ 15]. To achieve the optimal catalyst effectiveness and obtain the maximum product selectivity, the preferable operation conditions, in the present case, are at a catalyst to feed ratio of 5 while the conversion level is at 60%.

Table 1. Effect of C F on benzene production. Cumene

Conversion (%)

2.5 5.0 15.0 20.0 35.0 45.0 55.0 60.0 65.0 70.0 75.0 80.0 85.0

Cv = 5

C6H 6 Production

0.025 0.050 0.143 0.180 0.310 0.394 0.476 0.492 0.455 0.448 ............ ............ ............

C v = 50

C6H 6 Produced

C6H 6

C6H 6 Produced

(g/g-cat/min)

Production

(g/g-cat/rain)

0.329

0.0075

0.413 0.444

0.0083 0.0089

0.504 0.510 0.448 0.442

0.0108 0.0120 0.0090 0.0088

0.005 0.010 0.029 0.036 0.062 0.079 0.095 0.098 0.091 0.089

373 Furthermore, the catalyst status at this conversion level can be obtained from Figures 3 to 5 in terms of the catalyst coke content and surface area. Hence the status of the equilibrium catalyst used in the commercial operation should be adjusted by: 9 the degree of catalyst regeneration- temperature, residence time and gas composition 9 the space velocity, ie. the riser velocity 9 the catalyst to feed ratio 5. CONCLUSIONS Experimental work in a micro scale activity test equipment was performed to derive a testing strategy for the optimization of FCC unit operation due to coke deposition on catalyst. The cracking of cumene over a super-D zeolite catalyst was chosen as the model reaction because the nature of this reaction can represent the cracking of typical commercial FCC feeds such as gas-oils via the dealkylation of branched aromatics. In addition, this reaction can also eliminate any obscurity in catalyst deactivation from other contaminants. From this work, the analysis developed clearly shows how to optimize the catalyst to feed ratio to minimize catalyst inventory but maximize the feed conversion as well as the product selectivity. Although at the present stage, this cannot be directly applied to the commercial operation without further pilot plant substantiation. Since this is a continuous research program, further work will be undertaken using the typical FCC feed containing metal and sulphur fouling precursors to refine the developed strategy. This will be, finally adopted to run in a MAT for FCC catalyst evaluation. NOTATIONS CF

Mo,M~ P(ti) Pb(ti) S(t) Sd(ti) Si(t) ti Xi(t) Xb(ti) Xc(ti) Xd(ti)

catalyst to feed ratio, gm-cat/gm-feed/min molecular weights for cumene and benzene respectively, gm/mole liquid product mass from normal run at time t, gm liquid product mass from 'blank' run at time ti, gm normalized total benzene selectivity at any time t benzene selectivity at time t i total benzene selectivity at any time t time on stream between two sampling intervals, min total cumene conversion at any time t mass fraction of benzene in liquid at time t i mass fraction of unconverted cumene at time t i mass fraction of cumene conversion at time t i

ACKNOWLEDGMENT The author likes to thank Dr. R. Mann in Chemical Engineering Department at UMIST, UK, Dr. E. K. T. Kam in Petroleum Technology Department (PTD) at KISR, and Ms. K. AIDolama and Mr. D. Behzad in the Catalyst Characterization Lab., PTD at KISR.

374 REFERENCES

1. A. Stanislaus and H. Qabazard, Advances in fluid catalytic cracking: A review, Technical Report, Kuwait Institute for Scientific Research, Kuwait, KISR 4356, (1994). 2. P. B. Venuto and E. T. Habib, Cat. Rev. - Sci. Eng., 18 (1978) 1. 3. D. E. W. Vaughan, The Chem. Soc. (London), 33 (1979) 294. 4. K. Rajagopalan and E.T. Habib Jr., Hydrocarbon Processing, Sept. (1992) 43. 5. C.F. LeRoy, M.J. Hanshaw, S.M. Fischer, T. Malik and R.R. Koolman, Oil & Gas J., June 3 (1991) 90. 6. P. O'Conner, L.A. Gerritsen, J.R. Oearce, P.H. Desal, S Yanik. and A. Humphries, Hydrocarbon Processing, Nov. (1991) 76. 7. J.C. Brice and K.V. Kirkorian, Hydrocarbon Processing, May (1983) 83. 8. P. Grosdidire, A. Mason, A. Aitolahti, P. Heanonen and V. Vanhamaki, Computers Chem Eng., 17 (1993) 165. 9. R.F. Wong,, Hydrocarbon Processing, Nov. (1993) 59 10. A.S. Krishna, C.R. Hsieh, A.R. English, T.A. Pecoraro and C.W. Kuehler, Hydrocarbon Processing, Nov. (1991) 59. 11. International Crude Oil and Product Prices, July (1994) 174. 12.M. Viner and B. Wojciechowski, Canad. J. Chem. Eng., 62 (1984) 870. 13. T.E. Corrigan, H.F. Garrer, H.F. Rase and R.S. Kirk, Chem. Eng. Prog., 49 (1953) 603. 14. O.H.J. Muhammad, PhD. Thesis, University of Manchester Institute of Science and Technology, England, (1992). 15. I. Moore, PhD. Thesis, University of Manchester, (1983). 16. G. Thomson, MSc. Thesis, University of Manchester, (1984). 17. G. Thomson, PhD. Thesis, University of Manchester, (1986).

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All rights reserved.

375

P A R T I A L O X I D A T I O N OF C2-C 4 ALKANES INTO O X Y G E N A T E S BY A C T I V E O X Y G E N G E N E R A T E D E L E C T R O C H E M I C A L L Y ON G O L D T H R O U G H Y T T R I A - S T A B I L I Z E D ZIRCONIA

K. Takehira a, K. Sato b, S. Hamakawa a, T. Hayakawa a and T. Tsunoda a a National Institute of Materials and Chemical Research, Tsukuba Research Center, AIST, 1-1, HigashL Tsukuba, Ibaraki 305, Japan b Institute of Materials Science, University of Tsukuba, Tennodai, Tsukuba, lbaraki 305, Japan

Abstract. A

cell

system using yttria-stabilized zirconia (YSZ) as a solid electrolyte, e.g. Au IYSZlAg (Au: anode and Ag: cathode), was applied for the oxidation of C2-C4 alkanes, where the cell system worked as an 'oxygen pump' by transporting oxide ions from the anode to the cathode through the YSZ bulk. Despite the inertness of the gold anode for dissociative activation of molecular oxygen, partial oxidation of alkanes to the corresponding oxygenated compounds, i.e. ethane to acetaldehyde, propane to acrylaldehyde, and n-butane to acetaldehyde, acrylaldehyde and methyl ethyl ketone, was observed, together with complete oxidation over gold anode under oxygen pumping; the formation rate of the oxygenated products increased with increasing oxygen pumping through YSZ at 723-748K. No production of the oxygenated product was observed in alkane oxidation under open circuit conditions, i.e. no oxygen pumping. It is considered that oxide ions can be partially reoxidized at the triple phase boundary of Au-YSZ-atmosphere on the anode surface to form an active oxygen species, which can oxidize alkanes to the oxygenated compounds.

I . Introduction. Yttria-stabilized zirconia (YSZ) is a well known solid electrolyte that can transport oxide ions through its lattice via anion vacancies. A cell system using YSZ, e.g. P(O2), M I YSZI M', P(O2)' (M and M'; electrodes) can serve as an 'oxygen pump' [1], by which the oxygen flux transferred across the YSZ can be controlled by the electric potential externally applied between the two electrodes. The application of electrochemical cells using YSZ for the oxidation of ethylene [2], ethylbenzene [3], and methane [4] on metal or metal oxide catalysts has been demonstrated. However, the characteristics of the surface oxygen generated electrochemically through the YSZ have not been well established, since all the cell systems previously studied have been equipped with metal or metal oxide catalysts that can dissociatively activate molecular oxygen and therefore

376 catalyze the oxidation reaction by themselves. In the cell system, P(O2), MIYSZIM', P(O2),' molecular oxygen is reduced on the cathode surface and incorporated into YSZ as oxide ion (O2-), transferred across the YSZ,

and then reoxidized to molecular oxygen by 4-electron transfer on

the anode surface. In the oxidation chemistry, an activation of molecular oxygen can be generally achieved by its reduction, i.e. active species are reduced or disociated oxygen species, such as

O2-, O22-, O- or O [5]. It is therefore expected that the active oxygen species can be produced on the anode surface during oxygen pumping through the cell. We found that propene was partially oxidized to acrylaldehyde over an inert Au anode film under the oxygen pumping [6]: this suggests that oxygen species transported through YSZ onto Au surface are active for the selective oxidation of propene into acrylaldehyde. In this paper, we report the partial oxidation of C2-C4 alkanes into oxygenated compounds. Gold films alone cannot catalyze the oxidation by the surface oxygen supplied directly from the gas phase [7], and we can therefore observe the behaviour of lattice oxygen species from the YSZ.

2. Experimental. A schematic diagram of the reactor is shown in Fig. 1. An electrochemical reactor was constructed from an 8 mol% yttria-stabilized zirconia (YSZ) disk 32 mm in diameter and 1 mm in thickness (density, 5.4 g/cm3; Nippon Kagaku Togyo Co.). Thin compact films of Ag and Au were prepared as the cathode and the anode, respectively, on each face of the disk by vacuum evaporation. As the cathode metal, Ag was used because of its high ability in reductive activation of molecluar oxygen [8]. An Ag cathode film (1.5-2.0 ~m thick) and an Au anode film (0.5-1.0 ~m thick) was deposited from a molybdenum and a tungsten boat, respectively, onto the YSZ disk under 8x 10.6 Torr for 5 min. The electrodes of the disk were connected with an inert gold wire (Au anode) and a platinum wire (Ag cathode) to a dc power supply which controlled the oxygen transfer flux [ 1] from the Ag cathode to the Au anode across the YSZ by changing the applied voltage. The disk composed of AulYSZIAg was mounted between two vertical alumina tubes (A1203, 99.5 %; Nippon Kagaku Togyo Co.) and sealed by low-melting glass (Adhesion temperature, 763 K; Shibata-Hario Co.). The reactor thus obtained was placed in the electrical furnace to control the temperature of the electrochemical cell system. Electric current was measured to estimate the transportation rate of oxide ions across the YSZ by using Faraday's law, where reduction and oxidation between 02 and 0 2- on each of the two electrodes takes place via 4-electron transfer. In the cathode room, oxygen gas (0.6 g/h) was passed and an oxygen pressure of 101.3 kPa was maintained during the oxidation reaction. A gaseous mixture of alkane (5 %) nitrogen (5 %) and helium (90 %) was passed (1.2-1.4 g/h) over the Au anode at 450-475~ for testing the activity of oxygen species generated on the Au anode film. The products in the effluent gas were determined by gas chromatography using a thermal

377

........~ ,

I9

10

Au wire uartz tube umina tube

-

0

E

v=" c.._o

Ethane-N2-He

E

|

o

- -

dlllll1114-. ,~.,,-~y AO

~,. . . . .

",'~,

SZ (1 mm)

.c:

m -o

~ , ,,, ,,, , , . , ~ . j - - A g film (1.5-2.0 pm)

Ethane-O2-N2-He

lass seals ,,i,-

O O~

/lllliiii

''

Ptwire

':':

9

10

20

30

i

40

50

0 2 flux (lam01/h)

Figure 1. Reactor configuration

Figure 2. Ethane oxidation over AulYSZIAg.

conductivity detector and nitrogen as an internal standard. PEG 6000 and Molecular Sieve 13X were used as columns for analyzing the organic compounds and inorganic gas, respectively. 3 . Results and Discussion. In a blank test where nitrogen gas, instead of the reaction gas mixture, was passed over the Au anode at 723K, it was confirmed that the oxygen pumping was well controlled by the electric potential, i.e. the amount of oxygen gas evolved in the anode room (measured by the gas chromatography) coincided well with the value calculated from the electric current across the YSZ. 3-1.

Ethane Oxidation. When the ethane-nitrogen-helium mixture (5:5:90) was passed at a rate of 1.2 g/h over the Au

anode at 475 ~

no oxidation of ethane occurred under open circuit conditions, i.e. when no

oxygen was pumped through the YSZ. The oxidation took place when oxygen was pumped under closed-circuit conditions, and an increase in the oxygen pumping resulted in an increase in the rate of ethane oxidation [9]. No evolution of oxygen gas in the anode room was observed in this case. The rate of acetaldehyde formation is plotted as a function of the oxygen flux (Figure 2). The applied potentials varied from zero to 1.0 V in these experiments. At an 02 flux of 49.4 ~tmol/h, the formation rate of acetaldehyde and CO2 was 8.9 and 21.9 ~mol/h, respectively, and

378 the 02 consumption rate of 47.2 I.tmol/h was calculated by assuming water as the other oxidation product. The mass balance calculated between oxygen evolution and consumption was good and no other product was observed except a trace amount of ethene. Selectivities of acetaldehyde and CO2 were 45 and 55%, respectively, based on converted ethane (< 1%). The addition of oxygen (60 ml/h) to the mixed-gas flow in the anode space did not affect the rate of acetaldehyde formation even when the potential was varied (Figure 2), while that of CO2 formation was substantially enhanced. There was no detectable acetaldehyde formation under open circuit conditions in spite of the presence of 02 on the Au anode. The results obtained above clearly indicate no activity of an Au surface for the oxidation of ethane to acetaldehyde by surface oxygen supplied directly from the gas phase and, furthermore, no influence of applied potential on the acetaldehyde-formation activity of the surface oxygen from the gas phase. It is likely that acetaldehyde formation from ethane is performed only by the 'active oxygen species' generated electrochemically on the Au anode surface through the YSZ. It has been reported that ethane was oxidized to ethene and acetaldehyde over Mo, V, B, P and Nb oxide catalysts where N20 was used as the oxidant [10]; 02 gave a low selectivity of oxygenated products [ 11]. In the N20 oxidation of ethane over the Mo/SiO2 catalyst, O- ions, derived from N20, react with ethane by hydrogen-atom abstraction. The resulting ethyl radical reacts with surface Mo=O to form a surface ethoxide, which may either decompose to ethene or react further with the surface OH- to form acetaldehyde [ 10b]. It is likely that O- ions can be formed over the present Au anode surface during the 4-electron oxidation of 0 2- to 02 at the boundary between Au and ZrO2 (YSZ) and oxidatively dehydrogenate ethane. The ethyl radical thus formed can react with the surface oxygen to form a surface ethoxide, which may react with the surface OHof ZrO2 to form acetaldehyde. The surface oxygen over the Au anode is rather weakly bound to the anode surface compared to the Mo=O, and the ethoxide, therefore, may be more stable and react in preference with the surface OH- to form selectively acetaldehyde. When ethene (in place of ethane)-nitrogen-helium mixture (5:5:90) was passed at the same conditions, CO (10 ~tmol/h) was formed together with acetaldehyde (18 ~mol/h) and CO2 (45 ~tmol/h) at 50 l.tmol/h of 02 flux and the rate of each formation again increased with increasing oxygen flux. No oxidation of ethene occurred when no oxygen was pumped through the YSZ. When oxygen (60 ml/h) was added in the mixed-gas flow in the anode space, all the rates of CO, CO2 and acetaldehyde formations were greatly enhanced, even under open circuit conditions. Each value of CO, CO2 and acetaldehyde formation was 320, 210 and 130 ~tmol/h, respectively, and was not affected even when the applied voltage increased. This observation suggests that ethene can be oxidized by surface oxygen supplied directly from the gas phase, and therefore, ethane is oxidized not via ethene but directly to acetaldehyde via a surface ethoxide intermediate as mentioned previously [9].

379

Propane and n-Butane Oxidation.

3-2. 100r

14

t _~ 75 -

CO ~ "~ ~,.

o

E

,

CH3CH3 3~

"5

~,n

~._o "~50 -E

C 1.-13CH

"s~ I

,,'O ~

r"~

~.o_ .2"~E

CH2=CH2~~.'~

~

rr 251

E

CI..12=CHCHO

CH3CHO

*

"1

CH~CH~ 0

20

40

60

80

0 2 flux (pmol/h) Figure 3. Propane oxidation over AulYSZIAg.

Figure 4. Oxidation mechanism over AulYSZIAg.

When the propane-nitrogen-helium mixture (5:5:90) was passed at a rate of 1.4 g/h over the Au anode at 450~

no oxygenated product was formed under open circuit conditions when no

oxygen was pumped through the YSZ (Figure 3). Even under no oxygen pumping, however, propane was converted to several fragments, i.e. methane, ethane, ethene and propene. When the electric current, i.e. rate of oxygen pumping, was increased, the rate of methane, ethane and ethene formation first decreased and then increased, while that of propene formation monotonously increased. Acrylaldehyde was formed by closing the circuit, i.e. under oxygen pumping conditions, and the rate of its formation increased on increasing the electric current. No evolution of molecular oxygen was observed in this case, i.e. all oxygen species pumped through the YSZ were consumed by propane oxidation on the Au anode surface. The applied potential was varied from zero to 2.4 V in these experiments. We previously reported that there is no activity at an Au surface for the partial oxidation of propene to acrylaldehyde by surface oxygen supplied directly from the gas phase and, furthermore, that there is no substantial influence of applied potential on the acrylaldehydeformation activity of the surface oxygen from the gas phase [6]. Acrylaldehyde formation from propane appears to be performed only by the 'active species' generated electrochemically on the Au anode surface through the YSZ. It is likely that the Au surface catalyzed a thermal cracking of

380 propane to methane and ethane, which was slightly suppressed and then accelerated with increasing oxygen pumping. Both ethane and propane may be dehydrogenated to ethene and propene under the oxygen pumping conditions. The alkane and the alkene thus formed can be oxidized to acetaldehyde and acrylaldehyde over the Au anode. n-Butane oxidation gave methane, ethane, ethene, propane, propene, 1,3-butadiene, acetaldehdye, acrylaldehyde and methyl ethyl ketone. The last three oxygenated compounds were not observed under open circuit conditions, while they appeared upon closing the circuit and increased with increasing the oxygen flux, even though the amount of these oxygenated products was small compared with that of the hydrocarbons. A substantial amount of methane was produced in the C3 and C4 alkane oxidations; its amount was almost comparable to those of C2 and C3 hydrocarbons in the C3 and C4 alkane oxidations, respectively, suggesting that thermal cracking preferentially occurred on the Au surface. It is likely that alkanes can be directly oxygenated to the carbonyl compounds by the 'active oxygen species' generated electrochemically at the phase boundary between Au and ZrO2 (YSZ) through the YSZ because of the inactivity of the Au surface for the dissociative activation of adsorbed oxygen supplied directly from the gas phase.

3-3.

Active Oxygen Species.

In the present AulYSZIAg cell system, molecular oxygen is reduced to oxide ions by 4electron transfer on the Ag cathode, incorporated into YSZ, transferred across the YSZ, and then reoxidized again to molecular oxygen by 4-electron transfer on the Au anode. Therefore, it is likely that some active oxygen species can be formed on the Au anode surface during the reoxidation of oxide ions. In this electrochemical cell, oxygen evolution at the Au anode occurs through the triple phase boundary of Au-YSZ-atmosphere. High electric current under a high externally applied potential suggests the presence of both sufficient triple phase boundary and micro pores in the Au anode through which oxygen gas can permeate (Figure 4). Au has no ability for the dissociative activation of molecular oxygen and primarily works as an electron carrier. When oxide ions evolve through the triple phase boundary, the Au anode receives electrons from the oxide ions to produce a reoxidized oxygen species. Thus, an active oxygen species can be formed as a dissociated or a partially reduced form, i.e. O, 0-, O22-, or O2", from the oxide ions, in which the most plausible active species is O or O-. When an alkane is oxidized to an oxygenated compound by the active oxygen species, the alkane must diffuse from the anode space to the triple phase boundary through the micro pores. The active oxygen species, O or O-, formed at the boundary dehydrogenates the alkane to a surface alkoxide intermediate which can be converted to produce the oxygenated product. Then, the oxygenated product formed diffuses through the micro pores to the anode space.

381 It was reported that binary system of gold and transition-metal oxide showed catalytic activity for oxidation of carbon monoxide and hydrocarbons with gaseous oxygen [12] at low temperatures. Micro particles of gold (d < 10 nm) were supported on the transition metal and the periphery of the gold particle, i.e. boundary between the gold and the transition metal, is most likely the active sites. The transition metals (V, Cr, Mn, Fe, Co, Ni, Cu, etc.) also have their own catalytic activity for oxygen activation in the form of metal oxide and so a synergestic effect is observed in the boundary layer, i.e. gold activates the carbon monoxide while the transition metal activates adsorbed oxygen from the gas phase in the oxidation of carbon monoxide. A totally different mechanism from this must work in the present oxidation, because the surface oxygen supplied directly from the gas phase did not show any activity for partial oxidation. The active oxygen species must be evolved from the lattice oxygen in YSZ through the triple phase boundary by its electrochemical reduction. In the electrochemical conversion of hydrocarbons the NEMCA (non-faradaic electrochemical modification of catalytic activity) effect has been reported frequently over metal anodes [13] and rarely over metal oxide anodes [14]. The NEMCA effect is known to promote the rate of oxidation and, to the knowledges of the authors, such enhancement in catalytic activity is generally observed over the metal anodes which have original catalytic activity, e.g. Pt, Pd, Rh and Ag, and is also observed as a non-linear function of the electric current. In the present study, we observed an almost linear increase of activity with increase in the electric current. Lacking a reference electrode, it is beyond the scope of this work to elaborate on the work function of the anode material. However, it is likely that the contribution of the NEMCA effect is neglisible and the electrochemically generated ooxygen species operates in the partial oxidation of alkanes. 4. Conclusion.

The oxidation of C2-C4 alkanes has been studied by using a cell system, AulYSZIAg (Au and Ag were anode and cathode, respectively), where the cell system transports oxide ions from the cathode to the anode through the YSZ bulk. Partial oxidation of alkanes to the corresponding oxygenated compounds, i.e. ethane to acetaldehyde, propane to acrylaldehyde, and n-butane to acetaldehyde, acrylaldehyde and methyl ethyl ketone, was observed over gold anode under oxygen pumping and the amount of these products increased with an increase in the rate of oxygen pumping. No production of the oxygenated product was observed in alkane oxidation under open circuit conditions, i.e. no oxygen pumping. It should be noted that the oxygenates formed by the reaction of alkane with the oxygen species transported through the YSZ in spite of the inertness of gold anode in dissociative activation of molecular oxygen. It is likely that oxide ions can be partially reoxidized at the triple phase boundary of Au-YSZ-atmosphere on the anode surface to form an active oxygen species, which can oxidize alkanes to the oxygenated compounds.

382 References

1

E.C. Subbarao, 'Solid Electrolytes and Their Applications,' Plenum Press, New York,

2

M. Stoukides and C. G. Vayenas, J. Catal., 70 (1981 ) 137.

3

J.N. Michaelis and C. G. Vayenas, J. Catal., 85 (1984) 477.

4

D. Eng and M. Stoukides, Catal. Rev. -Sci. Eng., 33 (1991) 375.

5

'Proceedings of the International Symposium on Dioxygen Activation and

1980.

Homogeneous Catalytic Oxidation' 1987, Tsukuba Elsevier, Rausanne (1988). 6

T. Hayakawa, T. Tsunoda, H. Orita, T. Kameyama, H. Takahashi, K. Takehira and K. Fukuda, J. Chem. Soc., Chem. Commun., 1986, 961; T. Tsunoda, T. Hayakawa, T. Kameyama, K. Fukuda and K. Takehira, J. Chem Soc., Faraday Trans., 91 (1995) 1111.

7

G.I. Golodets, 'Heterogeneous Catalytic Reactions Involving Molecular Oxygen,' Elsevier, New York, 1983.

8

T. Arakawa, A. Saito, and J. Shiokawa, Bull. Chem. Soc. Jpn., 55 (1982) 2273.

9

T. Hayakawa, K. Sato, T. Tsunoda, K. Suzuki, M. Shimizu and K. Takehira, J. Chem. Soc., Chem. Commun., 1994, 1743.

10

a) M. Iwamoto, T. Taga, and S. Kagawa, Chem. Lett., 1982, 1469; E. Iwamatsu, K.-I. Aika, and T. Onishi, Bull. Chem. Soc. Jpn., 59 (1986) 1665; K.-I. Aika, M. Isobe, K. Kido, T. Moriyama, T. Onishi, J. Chem. Soc., Faraday Trans. 1, 83 (1987) 3139; A. Erdohelyi and F. Solymosi, Appl. Catal., 39 (1988) L11; A. Erdohelyi and F. Solymosi, J. Catal., 1990, 123, 31; A. Erdohelyi and F. Solymosi, J. Catal., 129 (1991) 497; A. Erdohelyi, F. Mate, and F. Solymosi, J. Catal., 135 (1992) 563. b) L. Mendelovici and J. H. Lunsford, J. Catal., 94 (1985) 37.

11

E.M. Thorsteinson, T. P. Wilson, F. G. Young, and P. H. Kasai, J. Catal., 52 (1978) 116; A. Argent and P. G. Harrison, J. Chem. Soc., Chem. Commun., 1986, 1058; Y. Murakami, K. Ohtsuka, Y. Wada, and A. Morikawa, Bull. Chem. Soc. Jpn., 63 (1990) 340; S. T. Oyama, A. M. Middlebrook, and G. A. Somorjai, J. Phys. Chem., 1990, 94, 5029; O. Desponds, R. L. Keiski, and G. A. Somorjai, Catal. Lett., 19 (1993) 17.

12

M. Haruta, N. Yamada, T. Kobayashi, and S. Iijima, J. Catal., 1989, 115, 301.

13

C.G. Vayenas, S. Bebelis, and S. Ladas, Nature (London), 343 (1990) 625; C. G. Vayenas, S. Bebelis, I. V/Yentekakis, and H. G. Linz, Catal. Today, 11 (1992) 303.

14

S. Bebelis, I. V. Yentekakis, S. Neophytides, P. Tsikaras, H. Karasali, and C. G. Vayenas, Proc. 3rd. Int. Symp. Solid Oxide Fuel Cells, May, 1993, ed. S. C. Singhal and H. Iwahara, PV 93-4, p. 926, The Electrochemical Society, Inc., Pennington, NJ (1993); A. G. Andersen, T. Hayakawa, K. Suzuki, M. Shimizu, and K. Takehira, Catal. Lett., 27 (1994) 221.

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All rights reserved.

383

THE E F F E C T S OF GAS COMPOSITION AND PROCESS CONDITIONS ON THE

OXIDATIVE COUPLING OF METHANE OVER Li/MgO CATALYST Saeed M. S. Al-Zahrani a and Lance L. Lobban b

a King Saud University, Chemical Eng. Dept. P.O.Box 800. Riyadh 11421, Saudi Arabia b University of Oklahoma, Chemical Eng. Dept. Norman, OK 73019, U.S.A. ABSTRACT The effects of CO2, steam, C2H6, and C2H4 partial pressures on oxidative methane coupling over a Li/MgO catalyst were studied under low conversion conditions in a fixed bed catalytic reactor. CH4:O2 ratios from 0.5 to 35 and temperatures from 973 K to 1073 K were used in the experiments and modeling. Results indicate CO2 has a poisoning effect on both COx and C2 formation rates, while selectivity to the C2 hydrocarbon was not significantly affected. Rate expressions rigorously derived from proposed mechanisms were obtained. The rate expressions agree well with the measured rates and predict well the high conversion experimental results under various reaction conditions. Results also indicated that addition of H20 to the feed enhanced the deactivation rate. The deactivation rate increased with increasing PH20 in the feed, and/or with increasing temperature. The deactivation rate was decreased by adding small amounts of CO2 to the reaction mixture. Results indicate that both CH4 conversion and C2 selectivity decrease with increasing C2H6 and C2H4 partial pressures in the feed. INTRODUCTION The direct conversion of methane into petrochemical feedstocks and or liquid transportation fuels is one of the most technologically important and challenging problems facing industrialized nations in the latest century.[1]. In the past thirteen years, the methane coupling has attracted the attention of many researchers. In spite of this effort, many aspects of the catalytic process are not yet well understood. The aim of most investigations has been to develop a catalyst which is able to convert methane selectively to higher hydrocarbons, preferably ethylene. Examples of promising catalyst system which have been studied to date are Li/MgO [2], Mn203/SiO 2 [3,4], PbO/AI20 3 [5], Sm20 3 [6], Sr/La20 3 [7], Mn/Na2WO4/SiO 2 [8]. Ross and co-workers [9,10] have explored the influence of CO2 on the oxidative coupling of methane over the Li/MgO catalyst. They found that carbon dioxide in the gas phase lowers both the methane conversion and the yield of ethane/ethylene products. They also found that carbon dioxide significantly improves the stability of the catalyst against deactivation. Based on the observations of surface species from FTIRS and transient experiments, In addition, most of the observations and experimental results reported to date cover a limited range of methane to oxygen feed ratios. There is a need to study the reaction over a wide range of methane to oxygen ratios and to quantify the effects of carbon dioxide on the reaction rates.

384 We have investigated the influence of CO2 reaction over Li/MgO over a wide range of CH4 and 02 flow rates. We have also quantified the effects using rate expressions derived from plausible mechanisms. The effects of carbon dioxide are significant over a wide range of methane to oxygen ratios. Our results also suggest a limited reaction between carbon dioxide and methane to produce carbon monoxide. The proposed rate expressions rigorously developed from mechanisms, predict well the reaction rates, selectivity and conversions of the methane coupling reaction over Li/MgO catalyst. Ekstrom [11] added 13C2H6 and 13C2H4 to th~me_thaneolx~enmi~ure^over Sm20 3 the total catalyst. These authors found that the oxidation of C2H6 ana t,2rt 4 m t s t ~ t ~ u CO x formation and decreased the unlabled C2 product. The conversion of added ethane and ethylene were 41 and 39 % respectively. At the same isotope concentrations, 15% of ethane and 39 % of ethylene were converted to COx, compared to 6 % of CH4 present at 10 times the concentration. The greater reactivity of the C2 products is clearly evident. The selectivity of ethane conversion to COx (versus dehydrogenation to C2H4) is comparable to the selectivity of methane conversion to COx . Except for this isotopic study there are no published investigations of the effects of C2H6 and C2H4 on methane conversion. In this paper we present the results of our study of the effects of ethane and ethylene on methane conversion. These studies were carried out to increase our understanding of the origin and mechanisms of the formation of the COx products because of their fundamental importance in this reaction. Varying flow rates of ethane and ethylene were mixed methane-oxygen mixtures prior to the reactor. Korf et al. [ 12] reported that the addition of steam to the reactor feed was deleterious to the stability of the Li/MgO catalyst at T=1073 K. They found that the lithium migrates from the catalyst, travelling in the direction of the gas stream. This caused a decrease in the activity of the catalyst. Kaminsky et al. [13] also reported that the addition of steam to the feed increases the deactivation rate of Li/MgO catalyst. Matsukata et al. [ 14] have observed that at T=998 K, steam accelerates the loss of the lithium from the surface of Li2CO3/MgO and LiNO3/MgO, but enhances the activity of the catalyst without decrease in the C2 selectivity. Kimble and Kolts [ 15] found that the addition of steam to the gas feed increased the C 2 yield in the presence of LiNO3/Mg(OH)2 at T=973 K. Recently Chang et al. [16] used LiNO3/Mg(NO3)2.6H20 at T=873 K and reported that the presence of steam is essential to the formation of the coupling product at low space velocities, but their data shows a very rapid deactivation of the catalyst. The above disagreement, besides the importance of studying the effect of steam on the reaction in addition to the limited number of studies, prompted our effort to elucidate the role of steam in the reaction with the ultimate goal of improving the catalyst and optimizing the reactor operating conditions. We have investigated the influence of adding H20 to the gas feed Li/MgO at different flow rates, temperatures and along with CO 2. In an attempt to obtain further evidence for the reaction pathways found for Li/MgO catalyst in a plug flow reactor, the effects of varying the residence time x were studied. EXPERIMENTAL The 7.0% Li/MgO catalyst used in these experiments and the properties of this type of catalyst have been described elsewhere [ 17]. Methane (99.97%), oxygen (99.99%), carbon

385

Computer Interface Module

Mass Spectrometer

ReactorTemp. Controller

'i'--1 ~ ] "

Gas Chrorn. i o ooo~i

Ioob l

Iiiii

Heating Tape

i I

t

I

CH4

o, 002 ( ~ .

I ;_, IL. -rLl3,

He

_

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~'~[

Furnace

bier

Mass Flow Controllers

'" zWater Condenser

Vent

Water Bath ( Feed Control Section )

( Reaction Section )

( Analysis Section )

Figure 1. Plug flow reactor experimental setup. dioxide (99.8%), ethane (99.9 %), and ethylene (99.9 %) were obtained from Linde, and all gases were used without additional purification. The reactant gases were mixed and then fed to the reactor. The experimental setup is shown in Figure 1. Reactions were carried out in a simple tubular quartz reactor (Figure 2). A 7 mm id quartz tube tapered to a 2 mm id quartz tube to remove the reaction gases from the reaction zone as quickly as possible to minimize gas phase reactions. Quartz wool was used as support for the catalyst bed. A thermocouple positioned against the outside wall of the reactor and a temperature controller (Omega CN1201) were used to control the reactor temperature. The actual temperature of the catalyst bed was calibrated in a separate experiment using a second thermocouple positioned in the center of the catalyst bed. Blank reactions with the catalyst bed filled with quartz chips of the same amount and mesh size as the actual catalyst showed no reaction at temperatures up to 1073 K. Gas flows were controlled by electronic mass flow controllers (Porter Instruments) which could also be monitored using the computer data acquisition system. A gas chromatograph (CARLE $400 AGC) was used for on line analysis and was the major source of experimental data. In addition, a Quadrupole Mass Spectrometer (BALZERS QMG420) was used to observe the transient behavior of the system. The bulk analysis of the catalysts was carried out using a Varian atomic absorption (AA) spectrometer to determine the lithium content of the catalyst. The BET surface areas of the catalysts were measured by a Micrometrics Flowsorb II 2300~

386

7_mmid

~

, ReactionGasMixture ~'"""~'~""'~':"i

~ ~ ! II ~

Heating Elements

Furnace fi~/'/.l~ I~

Tubular / 2 mmid /

-t~[~" ~ : ~

/

~

Q~"

w~176

Products

Figure 2. Tubular quartz reactor. The steady state reaction data were obtained under low conversion conditions over lg of 20 to 40 mesh Li/MgO catalyst; in few experiments (particulady at high CH4:O 2 ratio and 1073 K), oxygen conversion exceeded 10%. The total pressure was maintained at 1 atm with a total feed gas flow of 1000 SCCM (standard cubic centimeters per minute) for all the experiments. The partial pressure range of methane was 0.015 to 0.3 atm, while the oxygen range was 0.005 to 0.03 atm. By varying the methane and oxygen flow rates, the methane to oxygen ratio was varied from 0.5 to 35. The carbon dioxide partial pressure range was 0 to 0.005 atm. Helium was used as the balance gas. Ethane and ethylene were added to the methane -oxygen mixture at partial pressures ranging from 0 to 0.02 atm and from 0 to 0.01 atm, respectively. Additional experiments were carded out to increase reactor residence time further. The steam experiments were carried out by flowing helium through a water bubbler maintained at constant temperature in a water bath. The water flow rate was controlled by manipulating the helium flow rate and the temperature of the bath. The tube carrying the helium-water mixture was heated to 353 K to prevent any further condensation before the reaction zone. The water was trapped in a water condenser after the reactor. RESULTS AND DISCUSSION Effects of CO 2

In all experiments, the major products were ethane and carbon dioxide. Under some conditions, ethylene and carbon monoxide were also observed. In the following, R1 is the C 1 products (CO2 and CO) formation rate, and R2 is the C2 products (C2H6 and C2H4) formation rate. The methane conversion is defined as (R I+2R2)/CH 4 in feed. The selectivity to C2 products is defined as 2R2/(Rl+2R2), while the C 2 yield is defined as the product of conversion and selectivity. Our experimental results indicate that methane does react with carbon dioxide to produce carbon monoxide and either hydrogen or water under reaction conditions, but if oxygen is present, most of the carbon monoxide will be further oxidized to

387 6.0 ~.

0.60 0.50 "~.

q

~l\ \ "~~0.40 ~ X,

5.0 ~ ~k,

PcH4/P~ ~ model values ooooo exp. T=973 *K + + + + + exp.: T=1023 :K

X

r..4.0 ~ ~ ~ "

~ .0.30

~,3.0 (_)

"~ 0.20

~2.0

i~0.10

1.0

0.00 0.000

0.001

0.002 0.003 0.004 Pco2 (arm.)

0.005

0.006

Figure 3. Experimental and predicted methane conversion rate at three different temperatures vs. carbon dioxide partial pressure (PcH4 = 0.2 atm. and PO2 = 0.01 atm. in the feed).

PCH4/P~ ~ model values ooooo exp., T=973 *K +++++ exp., T=1023 *K exp., T=1073

o.o -~.............. ~ , , t ,

0.000

0.001

*K

.................. t .........

0.002 0.003 Pco2

(arm.)0.004

0.005

0.006

Figure 4. Experimental and predicted methane conversion (%) at three different temperatures vs. carbon dioxide partial pressure (PcH4 - 0.2 atm. and P02 = 0.01 atm. in the feed).

carbon dioxide. Though the net result of the above reactions is the oxidation of methane to carbon dioxide, the additional reaction steps suggested by the experiments are helpful in quantifying the observed effects of carbon dioxide on the coupling reactions.

RCH4,R2,R1, and Methane Fractional Conversion For each methane to oxygen ratio, reaction with no CO2 in the feed was first carried out, and then carbon dioxide was introduced at different partial pressures into the reactant stream while holding the CH4 and 02 partial pressures constant. Figures 3 shows the methane conversion rate (RCH 4 = R1 + 2R2)versus CO2 partial pressure CH4/O 2 =20. Results are shown at three different temperatures. At 973 K, even a low partial pressure of CO2 has almost eliminated the methane conversion. But at 1073 K, methane conversion is observed even for the highest partial pressures of carbon dioxide introduced. Figure 4 shows the corresponding methane fractional conversions. The value of CO2 partial pressure indicated in these figures is the average of the feed and effluent partial pressures. C2 formation rate versus CO2 partial pressure at methane to oxygen ratios of 20, is shown in Figure 5. Clearly carbon dioxide has a strong poisoning effect on the formation of C2 products. The effect is very strong at low partial pressure of CO2 (in a range from 0 to 0.002 atm), while at higher partial pressures of CO2 the reaction is less sensitive to increasing CO2. Similar behavior occurs at all methane to oxygen ratios. The results also show that when the reaction temperature is increased, the effect of CO2 on the reaction rates becomes more pronounced. The effect of CO2 on R1 is qualitatively similar to its effect on R2. The results at a methane to oxygen ratios of 20 is shown in Figure 6.

388 0.25

0.05 pcn4/po2=20

,_0.20 .,..; a (J E .,~.0.15

~

PCH4/Po2=20

- model values ooooo exp., T=973 *K +++++ exp., T=1023 *K ,, ., = *,K

"\

\

~...,0.04

\

.~9

\ \

E~ ,,,~,0.03

model values ooooo exp., T=973 *K +++++ exp., T=1023 *K " " " " " exp., T=1073 *K

~ "

._~ E _eO.lO o

~0.02

~J

~o.o5

0.00

EO.01

0.000

0.001

0.002

0.003

0.004

0.005

Pco2 (atm.)

o

0.006

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

0.000

0.001

, .........

0.002

0.003

.........

0.004

0.005

0.006

Pco2 (arm.)

Figure 6. Experimental and predicted carbon oxides formation at three different temperatures vs. carbon diooxide partial pressure (PcH4 = 0.2 arm. and P O2 = 0.01 atm. in the feed).

Figure 5. Experimental and predicted hydrocarbon formation rate at three different temperatures vs. carbon dioxide partial pressure (PcH4=0.2 atm. and P O2 = 0.01 atm. in the feed).

100.0

-

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4-

+

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+ zx ....

:---~. . . . . .

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

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60.0

Pax4=.0825, Po2=.0175 P~4=.0825, P~=.0175 Pert4=.200, Po2=.010 Pc~4=.200, Po~=.010 Pert4=.175, Po2=.005 P~4=.175, Po2=.005

atm. arm. arm. arm. arm. atm.

Go 40.0

20.0

0.0

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0.000

0.001

0.002

0.003

0.004

Pco2 (arm.)

0.005

0.006

Figure 7. Experimental and predicted hydrocarbon selectivity (%) vs. carbon dioxide partial pressure at PCH4 / PO2 = 4.7, 20 and 35 (in the feed).

Selectivity to C2 Products When feed CO2 partial pressure is increased, the selectivity at different temperatures and at different methane to oxygen ratios remains almost constant, except at very low methane to oxygen ratios. This behavior suggests a common methane activation site for both carbon oxide and C2 product formation which is poisoned by CO2 Figure 7 shows the experimental selectivities at methane to oxygen ratios of 4.7, 20 and 35. The conversion and selectivity data indicate that C2 selectivities stayed constant when CO 2 was added to the feed. When there was no CO 2 added to the feed, increase in methane conversion was always accompanied by decrease in the C2 selectivity.

389

Mechanism and Rate Expressions Lehmann and Baerns [18] have reported different reaction rate expressions based on a number of mechanisms to predict the hydrocarbon formation rate and carbon oxide formation rate in terms of PCH4 and PO2. None of the earlier studies included the dependence of R2 and R1 on PCO2. Ross and co-workers [19] developed expressions for RCH4 as a function of methane, oxygen and carbon dioxide partial pressures, but separate expressions for R1 and R 2 were not reported. The adsorption of CO2 and 02 was confirmed in our lab using FTIR by Bhumkar and Lobban [20]. The reaction of methane with CO2 was also confirmed in our lab. The rate limiting step was assumed to be the reaction between adsorbed or gas phase methane and adsorbed oxygen to form methyl radicals as given by Cant et al. [21]. The enthalpy of adsorption of methane on Li/MgO is relatively low (-30 kJ/mole) [22], thus resembling physical adsorption. This possibly also explains why Bhumkar and Lobban [20] in our lab have not been able to detect adsorbed methane species with Fourier Transform Infrared Spectroscopy. The enthalpies of adsorption of CO 2 and 0 2 are -185 kJ/mole and -62 kJ/mole [22]. The methyl radicals couple to form C2H 6 or further oxidize to CO and CO2. Further oxidation of ethane was assumed negligible due to the very short residence times in our reactor. However, at longer residence times, further oxidation of ethane and/or ethylene becomes important. The parameters (adsorption equilibrium constants, rate constants, and coefficient ~) were estimated using the experimental data with the nonlinear regression software package SAS (Statistical Analysis Software). The Proc Model (with Marquardt-Levenberg method) and Fit procedures in SAS were used for this purpose[23]. Model discrimination was carried out using the results of the parameter estimation and the goodness of fit (using F-value, r 2 , and Ttest outcomes). In some cases, unrealistic parameter estimation (e.g., negative activation energies) eliminated a mechanism. Among those mechanisms for which all estimated parameters were realistic, the measures of goodness of fit were used for discrimination. According to our experimental results and the discussions above, the following mechanism is the one considered to be the most practical and successful. It is based on the following assumptions: 1- There is only one type of active site on the Li/MgO catalyst; 2- CH4,O2, and CO 2 adsorb competitively on the active site of the catalyst. 3- Langmuir isotherm rules can be used to describe adsorption of all three components. 4- The reaction between adsorbed 02 and adsorbed CH4 to produce the CH 3" is the rate limiting step. 5- The CH 3. is an intermediate product at quasi-steady state, therefore, dPcH 3-/dt = 0. 6- The C2H6 is formed through the gas phase CH3" coupling.

390 Based on these assumptions, the mechanism is expressed as follows:

Ko 02 + c~ +--+ O2-cy

(1)

Km ~---+ CH4-o

(2)

Kc CO2 + o <----~ CO2-o

(3)

CH4 + •

kr --~ CH3" + HO2"

CH4-o + O2-o

(4)

kl CH4-o + CO2-o ~

CO, CO2, H20

(5)

k2 ~

CO, CO2, H20

(6)

CH3" + 89

k3 ~

2CH3"

(7)

C2H6

The expressions for the reaction rates and conversion were also derived: Methane consumption rate: --RcH 4 =

K mpC-H4 ( k r K o P 0 2 + k, KcPco 2 ) (1 + K 0 P02 -at"Kc 1)(',02 -}- Km P('H4 )2

(8)

Hydrocarbons formation rate: dl)c 2H6 _-- k 2~Po ~x

R2

16k3

dt

F(1

l

8k3kr

k22Po2 x

1/2 12

gol)o2gmP(,H4

(1+ gml)(TH4 "~- K01)02 4;- g(,Pc02)2)

-1

(9)

Carbon oxides formation rates:

R, =

_]_

dt

=

4k 3

(1 + k22P02 x (l+K,,,PcH4KoPo2+K<,P~,02) 2 )

- 1

(10)

k, KcK,,Pco2PcH4 . . (l+KoPo2 + K c Pc02 + K,,,Pcm )2

The Arrhenius form of the parameters are used, i.e., ki=Ai e-Ei/RT, Ki=Ai e~'i/RT, and the exponent ~ were estimated using the nonlinear regression software package SAS (Statistical Analysis Software). The Proc Model (with Marquardt-Levenberg method) and Fit Procedures in SAS were used for this purpose. The results are shown in Table 1 and Table 2.

391 Table 1. Predicted Kinetic Parameters Parameter

Predicted Values

Ar Er A1 E1 A32 E32 Ac ~c Am ~m Ao ~o . ~,

mmole/min./gm kJ/mole mmole/min./gm kJ/mole gm.min.atmX/mmole

1.282 1.992 8.830 1.564 3.940 19.44 2.185 1.323 2.777 62.65 2.753 30.92 4.75

atm -1 kJ/mole atm- 1 kJ/mole atm -1 kJ/mole

x x x x x

101 u 102 106 102 10-4

x 10-4 x 102 x 10-3

Table 2. Statistical Analysis Result (r 2 is coefficient of determination, SSE is Error Sum of Square and MES is Error Mean Square) # of Data

r2

SSE

MES

R2

166

0.965

2.64x10 -3

1.64x10-5

R1

166

0.827

6.86xl 0 -4

4.33x 10-6

The model describes the methane coupling reaction over Li/MgO quite well. The predicted reaction rates, conversions, and selectivity are plotted with experimental results in the previous figures. Rates and selectivities are also calculated over a wide range of conditions. According to the model, the decrease in the methane conversion rate with increasing CO2 partial pressure is due to the strong adsorption of CO2, reflected in the large values of K c . Effects of C2H 6 and C2H 4

Ethane at partial pressure ranging from 0 to 0.02 atm was added to the methane-oxygen mixture. Different methane-oxygen ratios were used in these experiments. Figure 8 shows the methane conversion and C 2 selectivity versus C2H 6 partial pressures at CH4:O 2 = 4.7 and T = 1023 K. Both methane conversion and C 2 selectivity decreased with increasing PC2H6 in the feed. The decrease in methane conversion is probably due to competition between CH 4 and C2H 6 for the active sites of the catalyst. Ethylene was added to the reaction mixture at partial pressures ranging from 0 to 0.01 atm. with different methane to oxygen ratios (i.e. CH4:O 2 = 4.7, 16.5 and 25)and at T =1023K. Figures 9 shows the methane conversion and C 2 selectivity versus C2H 4 partial pressure at CH4/O 2 = 4.7, respectively. Both methane conversion and C 2 selectivity decreased

392 100

10

,,-----,, co sd. (%) ,,----,, 4 Cony. (%) ~

~

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Figure 8. Methane fractional conversion and C2 selectivity versus PC2H6 in the feed over 1 gm of 7% LifMgO catalyst at CH4/O 2 = 4.7 and T = 1023 K.

10

100 ,,--~ "---"

9 o O., Sel. (%) @4 Conv. (%)

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0

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Q

.004

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PC2H4 (atm.)

Figure 9. Methane fractional conversion and C2 selectivity versus PC2H4 in the feed over 1 gm of 7% Li/MgO catalyst at CH4/O 2 = 4.7 and T = 1023 K. remarkably with increasing PC2H4 in the feed. The drop in methane conversion is due to a competition between CH 4 and C2H 4 for the same active site. The decrease in C 2 selectivity is due to the high reactivity of C2H 4.

393 8

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~176176176

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% ~ D~o

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9

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9

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T i ~ (hr)

Figure 10. Change in CH4 and 0 2 conversions and C2 selectivity as a function of time on stream over 1 gm of 7% Li/MgO catalyst at T= 1073 K, PCH4 = 0.0825 atm., PO2 = 0.0175 atm., and PH20 = 0.003 atm.

Effects of Steam Experiments with different water flow rates, temperature and a long with CO 2 were carried out. In all the experiments, the deactivation of the catalyst took place. The higher the feed partial pressure of water and/or the temperature the faster the deactivation rate. Figure 10 shows that when the partial pressure of water in the feed was 0.03 atm., the methane conversion dropped from 6.8% to 0.4% during 30 hours of operation. Within this period the 0 2 conversion dropped from 15.5% to 1% while the C 2 selectivity varied between 85%-90%. The data showed that the steam is clearly very detrimental to the stability of the catalyst. It is hypothesized that volatile LiOH was formed and evaporated from the Li/MgO catalyst. After 30 hours, when the H20 flow was stopped, the activity did not rise. This effect is probably due to the loss of the incorporated lithium as will be discussed later. It is believed that at the higher temperature the lithium gradually evaporates from the surface of the catalyst. No measurable reaction was found at T = 873 K and PH20 = 0.03 atm while the PCH4/Po2 varied between 1 and 10 with a total flow rate of 1000 SCCM. The catalyst deactivation rate decreased when CO 2 was introduced into the system with a very law partial pressure. These results indicate that the carbonate species were formed instead of the volatile species. The carbonate species are more stable but less catalytically active.

394 After each steam experiment, it was noticed that the reactor walls had deteriorated. Potentially this is from the volatile lithium diffusing into the quartz to form an inert lithium silicate phase. This lithium depletion also probably contributes to catalyst deactivation. The lithium content and the surface area of the fresh and used catalysts are listed in Table 3. Table 3. Li content and surface area for the fresh and used catalysts in steam experiments.

Time (hr)

T (K)

PH20 in the feed (atm.)

PCO in the feed~atm.)

Li (wt%)

Surface Area (m2/g)

3.0 30 1073 0.03 0.0 0.54 34 1073 0.003 0.0 0.78 44 1023 0.003 0.0 1.01 17.5 1073 0.03 0.001 1.84 Conditions: lg of 7% Li/MgO, PCH4=0.0825 atm., PO2=0.0175 atm., total flow= and Ptota I = latm. * After calcination.

Fresh catalyst* Used catalyst

0.64 0.34 0.55 0.40 0.41 1000 SCCM,

Experimental and Simulation of a Plug Flow Reactor The data used for parameter estimation were from experiments with very short gas residence times and generally low conversions. In order to check the validity of the derived expressions, we also carried out experiments at higher residence times. The proposed model reaction rate expressions {eq.8, 9 and 10} were used to predict the integral reactor behavior under the following assumptions: 1) Reactor operation is plug flow and adiabatic. 2) Inter- and intraparticle heat and mass transfer gradients are negligible. 3) The p and Cp of the gas phase remain constant independent of conversion. The resulting pseudo-homogeneous model is described by the following mass and energy balances; RTp, ]-' dP~ (11) d r = -r~

dCoCT) p, dr

= (-AHR,)R' + (-AHR=)R2

(12)

where: i = CH4, COx, 02 and C2H6; x = catalyst volume/total gas volumetric flow rate. Ps = catalyst density ; p = density of the gas mixture; Cp = specific heat of the gas mixture ; Pi = partial pressure of component i; AHR1, AHR2 = heats of reaction for the COx and C2 production reactions. Figure 11 shows good agreement between the predicted methane conversion of the model developed in this work from low conversion data and the experimental results at higher conversion, x was varied experimentally by varying the mass of the catalyst within the reactor and/or varying the total gas volumetric flow rate. The maximum experimental x of 0.6 see is about 12 times the value used in earlier experiments. Figure 12 shows the predicted and the experimental values of C2 selectivity at different residence times. The C2 selectivity decreased

395

ooooo

T~=1073

*K

,,~,~,,A T~=1023 *K PCH,~f = 0 . 0 8 2 5 arm. PO~ = 0.0175 arm.

o..,~_~

"80

60. 8.0 o O

"6 U3

:s

G

4O

u

4.0 Tf = 1 0 2 3 *K PCH,, = 0.0825 P02~ = 0 . 0 1 7 5

o.o

, N , , , , , , I , , , , .... , | , , .... , , , | , , , , , , , , , i , , , , , , , , , i , , , i , , i , , l , w l l ,

o.o,

o.lo

0.20

0.30

*

o4o

o.5o

o.~o

"~.7o

(s~o)

Figure 11. Comparison of model predictions for methane fractional conversion (%) under integral reactor conditions with the experimental data at two different temperatures (Pc ~ = 0.0825 atm. and P O2 = 0.0175 atm. in the feed).

o

.

o.oo

.

.

'"6;~b ..... ~3b'

.

.

.

.

.

arm. at.m.

.

. . . . . .0. . .3.0. . . . . . . . .0. . .4.0. . . . . . . . .0. . 5. .0. . . . . . 0 . 6 0

.

0.70

(s~c)

Figure 12. Comparison of model predictions under integral reactor conditions with the experimental data at T = 1023 K (Pc_H4 = 0.0825 arm. and PO2 = 0.0175 arm. in the feed).

with increasing the residence time. We expect the C2 selectivity that to continue decreasing for higher residence times due to further oxidation of the C2 product, which also support the idea of rapid separations of C2 product. Further investigations are currently underway in our lab to determine the effects of operating variables such as dilution ratio, reactor length and distributed oxygen feed on methane conversion and C2 selectivity. CONCLUSIONS Under steady state reaction conditions, the effects of CO2 on the methane coupling reaction over Li/MgO catalyst were quantitatively determined. Poisoning effects of CO2 on carbon oxide formation rate, C2 formation rate, and methane conversion were observed for all methane to oxygen ratios and all temperatures. However, C2 selectivity is relatively unaffected by CO2 partial pressure. The mechanism described here accounts for important elementary steps, especially the effects of carbon dioxide. Under the low conversion conditions used in this study, further oxidation of C2 products to CO and CO2 is assumed negligible. These reactions will become more important at high conversions. Rate expressions derived from the mechanism match well the experimental conversions and selectivities. The addition of ethane or ethylene caused a significant decrease in methane conversion and in C 2 selectivity. The ethane and ethylene probably compete with methane for the active sites on the catalyst. The reactivity of ethylene and ethane is higher than that of methane. The high reactivity of the C 2 products, in addition to decrease in CH 4 conversion when either C2H 6 or C2H 4 is added to the feed highlights one of the major problems facing the commercial development of this reaction. Maintaining high C2 selectivity without decrease in methane conversion might be possible through rapid separation of C 2 products from the reaction zone.

396 The addition of H20 to the gas feed is very detrimental to the stability of the catalyst. The higher the partial pressure of the steam, the faster the deactivation rate. Also, the higher the temperature, the faster the deactivation rate. In the presence of steam, the deactivation of the catalyst can be retarded if a low concentration of CO 2 is added to the reaction mixture. The steam results also show that the lithium is present in two forms i.e., Li+O - and inactive Li2CO 3 . A reactor-reaction model has been presented. The reaction model predicted well the result of integral experiments which were conducted using the same catalyst under different operating conditions. The methane conversion increased with increasing residence time while the C 2 selectivity decreased with increasing residence time from x = 0.05 to 0.6 see. REFERENCES

1. 2. 3. 4. 5. 6. 7. 8.

J.H. Lunsford, Catal. Today, 6(3) (1990) 235. T. Ito, J.X. Wang, C.H. Lin and J.H. Lunsford, J. Am. Chem. Soc., 107 (1985) 5062. J.A. Sofranko, J.J. Leonard and C.A. Jones, J.Cat., 103 (1987) 302. C.A. Jones, J.J. Leonard andJ. A. Sofranko, J.Cat., 103 (1987) 311. W. Byton and M. Baerns, Appl. Catal., 18 (1986) 199. K. Otsuka, K. Jinno and A. Morikawa, J.Cat., 100 (1986) 353. J. M. DeBoy and R. F. Hicks, J. Cat., 113 (1988) 527. Z. Jiang, H. Gong, L. Feng and H. Wang, Symposium on Methane and Alkane Conversion Chemistry, Am. Chem. Soc. Meeting, San Diego, CA, 1994, 39, 2, 255. 9. S.J. Korf, J.A. Roos, N.A. de Bruijn, J.G. van Ommen and J.R.H. Ross, J. Chem. Soc., Chem. Commun. (1987) 1433. 10. S.J. Korf, J.A. Roos, L.J. Veltman, J.G. van Ommen and J.R.H. Ross, Appl. Catal., 56 (1989) 119 11. A. Ekstrom, Methane Conversion by Oxidative Processes, edited by Wolf, E., (1992) 99. 12. S. J. Korf, J. A Roos, J.G. De Bruijn and J.G. van Ommen, Appl. Catal., 1990, 131. 13. M. P. Kaminsky, G. A. Huff, M. J. Spangler and T. P. Kobylinski, Symposium on Natural Gas Upgrading II, Ame. Chem. Soc., 37-1 (1992) 89. 14. M. Matsukata, E. Okanari, K. Komori, E. Matsuda, E. Kikuchi and Y. Morita, Symposium on Methane Activation, Conversion, and Utilization, Int. Chem. Cong.of Pacific Basin Soc. (1989) 58. 15. J. B. Kimble and J. H. Kolts, Energy Progress, 6-4 (1986) 226. 16. Y. F. Chang, G. A. Somorjai and H. Heinemann, J. Cat., 141 (1993) 713. 17. W. Tung and L.L. Lobban, Ind. Eng. Chem. Res., 31 (1992) 1621. 18. L. Lehmann and M. Baerns, J. Cat., 135 (1992) 467. 19. J.A. Roos, S.J. Korf, R.H.J. Veehof, J.G. Van Ommen and J.R.H. Ross, Appl. Catal., 52 (1989) 131. 20. S.C. Bhumkar and L.L. Lobban, Ind. Eng. Chem. Res., 31 (1992) 1856. 21. N.W. Cant, C.A. Lukey, P.F. Nelson and J. Tyler, J. Chem. Soc., Chem. Commun. 1988, 766. 22. S.J. Korf, Ph.D. Dissertation, University of Twente, 1990. 23. SAS-SAT Users Guides., SAS Inst. Co., 1990, Version 6, 4th ed.

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All rights reserved.

397

STUDY ON THE ACTIVE SITE STRUCTURE OF MgO CATALYSTS FOR OXIDATIVE COUPLING OF METHANE

Ken-ichi Aika and Takashi Karasuda

Department of Environmental Chemistry and Engineering, Interdisciplinary Graduate School of Science & Engineering, Tokyo Institute of Technology, 4259 Nagatsuta, Midori-ku, Yokohama 226, JAPAN ABSTRACT A general reaction mechanism of oxidative coupling of methane is proposed. According to this mechanism, a kinetic expression is related with both the catalyst morphology and the active site concentration. The active site of MgO catalyst is proposed to be the electron deficient surface oxygen ions such as O- relating to the surface defects which appear at the high temperature (above 973K). TPD, D.C. conductivity, and oxygen exchange measurements were carried out to confirm the defect model. 1. INTRODUCTION Oxidative coupling of methane (OCM) is a useful way to produce ethylene which can be turned to gasoline through the oligomerization. This is one of the future processes which convert natural gas to gasoline. 1.1. General mechanism of OCM reaction The following mechanism has been generally accepted. Activated surface oxygens abstract hydrogen from methane producing a methyl radical, which is either oxidized further to CO2 or reacted with another methyl radical producing ethane. The reaction scheme is thus shown in Eqs. 1 to 4, where O* means an active oxygen and xO means a less active oxygen which can not react with CH4 but can react with a methyl radical. If two steady state treatments are carried out for the concentration of the active surface oxygen and that of the methyl radical, the reaction rates of C2 formation (RC2)and Ci formation (RC0 are given as Eqs. 5 and 6. These kinetic expressions have been successfully adapted to kinetic data of OCM reaction on Na/MgO [1,2] and Li/NiTiO3 [3 ].

1.2. Relation between C 2 selectivity and morphology Eqs. 5 and 6 give a relative rate of C2 formation against Cl formation (RCz/RCI) (see Eq. 7). If kt, k2, or ks is a reaction with the surface, but k4 occurs in the gas phase apart from the surface, RC2/RCI is rewritten as Eq.8. RC2/RCt decreases with an increase of the surface area. Iwamatsu and Aika [2,4] have explained once why the active catalyst gives high C2 selectivity when the surface area is low. Yamagata et. al. [5] have also found that C2 selectivity is high when the catalyst has open space in the catalyst grain. These phenomena are explained in Fig. 1.

398

C2H6 C02

CH4

"CH3

(A) Open space

CH4

(B) Closed space

Fig.1 Relation between C2 selectivity and surface morphology

1.3. Relation between C2 selectivity and active site concentration Marcelin et. al. [6] have proposed almost the same kinetic model where k l is assumed to be greater than k2 and PO2 is assumed to be constant. The rate ratio R c 2 / R c I is thus simplified as Eq.9. They tried to apply this model for the mixed oxides catalyst such as PbO2-MgO or SnO2-MgO. Two kinds of active oxygen (k2' and k2") and two kinds of less active oxygen (k3' and k3") are assumed. The total activity depends upon the surface concentration of one component (Xl). In this case Rc2/Rc1 is rewritten as Eq. 10, which has a maximum at certain value of Xl (Xl- 0.01). This model successfully explained why a small amount of Pb (Xl- 0.01) was necessary in their results. If PbO2 covers all the surface like Fig.2-A, the produced methyl radical is oxidized easily to CO2. If concentration of PbO2 is limited like Fig.2-B, the produced methyl radical is not oxidized but dimerizes to give C2 compounds. If we can have a small concentration of very active site on MgO or Li/MgO like Fig.2-B, we can obtain a high C2 selectivity. Such an active site might be O- on these system if we have no PbO.

399 02+site

2k~ ; 2 0 *

CH 4 + O * 9CH 3

(3)

k4 ;C2H 6

1((

-dt - k4PcH3z - " i ~

m

(2)

k3 ; C O 2

de2

Re2-

RCI

k2 ~ . CH 3 + OH (a)

+ XO

2.CH 3

(1)

de, dt

-

-1

I+8CA+B

(4)

)2

(5)

x,2 4C1((1+8c A + )B 0.5 ) k3Pc"3 Pch ~-

~1

(6)

k4

where, A = kiP02 ,B = k2PcH4, C - b 2 p x "3 02 Pc,.

1[

0~

AB

1~

8klk2k4PozPcH4

1[{

0.5

]

-11

]

(7)

If heterogeneous-homogeneous reaction; kl, k2, k3 oc Specific surf. area, k4 independent from specefic surf. area ; 2

Rc,` Rcl

1[ -

4

const 1 + ~

I(

Ssp

/~

-11

j

(8)

If k 1) ) k 2,

,

Rc,` 1 [{. Rc, - 4 1 ( 1 + 8 c

,

8k4 PCH4k2

k32

)0.s-ll]

k4

where k4 - Po,` x ,Po,` - const" r

~

rt

If two sites exit (mixed oxide) k2 - xlk2 +(1- x~)k2", k3 - xlk3 +(1- xl)k3 , m - k2'/k2",n -- ka'/k3

1 If

P~

Rc~

8k4'PcH4[( m - 1)x1 + 1]k2' ! 0.5

II 1+ 4 [~

[ ( n - 1)x 1 + 112k3'2

/}

~.

]

-ll

]

(10)

400

1.4. Active site generation in MgO system Lattice defects can be generated on MgO if the surface is activated with UV light. The electron trapped in the defect reacts with N 2 0 to give O-, which immediately reacts with CH4 to give methyl radicals even at room temperature [7]. However, the defect formed with the UV is destroyed above 473K. On the other hand, MgO can bear defects at high temperature above 873K [8]. However, the real active structure is not known well at the reaction condition of high temperature, because any spectroscopy is not applicable to determine the active site structure at the high temperature [9]. Here we intended to study the DC conductivity, oxygen isotope exchange, and TPD techniques which give the active site information at the high temperature [10]. Two kinds of magnesia catalysts, i.e., MgO and Li/MgO were studied. The defect notations follow after Kroger [ 11].

C2H6

C02

CH4

"CH3

(A) All surface active oxygen

Fig.2

CH4

9CH3

l

9CH3

(B) Restricted number of active oxygen

Relation between C2 selectivity and concentration of surface active oxygen (bold circle)

2. EXPERIMENTAL A slurry containing high purity grade magnesium oxide (Soekawa Chemicals, purity min. 99.96%) and LiNO3 (Kanto Chemical, CICA Reagent) was prepared. The slurry was dried at 423K for 48h and then powdered in an agate mortar. The content of lithium in MgO was one mole percent. The powder was pressed into 2 cm diameter pellets at 200 kg/cm 2 followed by a sintering treatment in air at 1273K for 24h. TPD experiments were performed using an NEVA NAG 110 mass spectrometer with a glass-made vacuum system. Prior to every TPD measurement the sample was heated up at 1173K for lh under vacuum (10 -6 torr) to remove the adsorbed impurities. The typical water vapor adsorption pressure was around 20 torr and the heating rate was 10 K/min. The temperature programmer used in our experiments was Chino, model KP. Four electrodes method was used to measure the surface D.C. conductivity of the samples. The electrical contacts between the leading wires and the pellet surface were made by gold sputtering decomposition followed by high purity gold paste addition. The

401 conductivity was measured by a laboratory made electronic device. A constant (10V) potential was applied at the pellet end; the maximum sensitivity of our apparatus was 10-12A. Four high purity gold electrodes (~ = 0.8 mm, Tanaka company) were drawn out from the upper part of the quartz made sample chamber; the external parts of the electrodes were insulated by quartz tubes. The electrodes configuration are depicted in Fig.3a. The sample chamber used for D.C. conductivity measurements under controlled atmosphere is shown in Fig.3 b. An isotopic exchange reaction between 1802 and the oxide catalysts was done in a closed circulation system containing 0.2 g of the catalysts. Isotopic concentration was monitored by an NEVA NAG 110 mass spectrometer.

I ~

i i1

-

2

'

~

3

~

6

5

SAMPLE INLET ~

SOURCE

' i

,1 | I

(a) Y"

~7

1 OUTLET

Fig.3

(b)

a)Electrodes configuration for magnesium oxide surface D.C. conductivity measurement; where AU and I are the potential drop(V) and the current(A), respectively. b)The sample chamber for conductivity measurements under controlled atmosphere: 1.Thermocouple; 2.Quartz shield for the thermocouple; 3.Gold leading electrodes (tl~-0.8mm), 4.Quartz measurement chamber; 5.Ground connected metallic shield; 6.Furnace; 7.Sample

3. R E S U L T S

AND D I S C U S S I O N

3.1. TPD study Small amount of H 2 0 was absorbed on the catalysts. Desorption temperatures of H 2 0 from the two catalysts were different each other depending on the surface chemical nature. On the other hand, H2 evolution was observed from 973K. TPD spectra of MgO catalyst is shown in Fig.4. H2 evolution was assumed to come from the OH- anions which

402 are connected to the cation vacancy. After the H2 evolution, two O- are assumed to be left. The reaction scheme, Eqs. 11 to 13, are proposed, where OO~ means O-. 1.6 0 2 (69.3kPa) 1.5

o H20

/

1.4

0

/

"~ 1.3 ~ "-" 1.2

0 2 (32.7kPa)

600

800 1000 Temperature/K

1200

0.9 673

I

i

i

773

873

973

i

1073

i

1173

Temperature/K Fig.4 TPD spectra of H20 and H2 desorbing from MgO

Fig.5 Relative D.C. conductivity of MgO with 02 against that of 0 2 free

3.2. DC conductivity and oxygen effect The D.C. conductivity measurements are used to identify the most abundant and the most mobile charged defects on MgO. In this study, the effect of oxygen and water on the D.C. conductivity of MgO catalysts were measured. These gases are the main components in the OCM reaction. The D.C. conductivities of the sample with oxygen gaseous atmosphere are divided by that without oxygen but with helium, and the values, (o/o0)23, are listed as a function of the temperature in Fig.5. Data show that oxygen increases the conductivity. The positive effect of oxygen suggests the p-type conductance as is written as Eq. 14. 1%Li/MgO, the selective OCM catalyst, shows a clear p-type nature especially at the higher temperature at which the OCM reaction occurs. If Li20 is included in the MgO matrix, two Li + ions are located at Mg 2§ sites forming an oxygen ion vacancy (Eq.15). The oxygen vacancy can accept neutral oxygen (02) forming two holes (Eq. 16) or two Oanion radicals (Eq.17). In the case of Li/MgO, the strong oxygen effect on the D.C. conductivity is considered to be due to two effects: one by the pure MgO structure (Eq. 14) and one by the Li + impurity effect (Eq.16). The activation energy of D.C. conductivity of 1%Li/MgO sample was 26 kcal/mol below 773K and 42 kcal/mol above 773K. The same conductivity study on MgO single crystals has given two activation energies: 25 kcal/mol at the low temperature and 55 kcal/mol at the high temperature [8]. Dubois and Cameron [9] have pointed out the importance of p-type conductivity. A p,type oxide will generate chemisorbed anionic oxygen in the presence of gas phase oxygen (Eq. 14,16,17). The hole on an oxygen anion is identical to O- (Eq. 18).

403 MgO + H 2 0 ---, MgO(H20) --~ Mg(OH)2

(11)

MgO(H20) ~ MgO + H20

(12)

2(OH)o" --~2Oo~ + H2

(13)

VMg" + V O~ 0.5 0 2 --*VMg" + 2h~ + Oo x Li20 ~ 2LiMg' + O o x + V O~176 Li20 + 0.502--~ LiMg' + 2Oo 'x + 2h ~

(14) (15) (16) (17)

--* 2LiMg' + 2 0 0 ~ +11"---'00 ~ (or 02- + h ~ - ) VO" + O o x + H20 --, (OH)o~

3.3.

Isotopic

(18)

Hi~ O o x ---, 2(08)0"

(19)

exchange

The time course of concentration of 180 2 was measured when it contacted with MgO catalyst. The results are shown in Fig.6. As is seen in this figure, 160 in gas phase increases rapidly in the first 100 min and slowly after that. Thus, we proposed the two stages exchange model, the fast exchange with active surface oxygen in the first stage and the slow exchange with lattice oxygen major in the later period. The model is shown in Fig.7. Using the two-stage exchange model, the rates(R1 and R2 in Fig.7) and the amounts(N(1-Ctl)/Ctl and N(1-t~2)/t~2 in Fig.7) of both active surface oxygens and exchangeable lattice oxygens in 24 h are calculated at every temperature. A relative amount of active surface oxygens per surface area was high on Li/MgO, and low on MgO as is shown in Table. 1. This tendency was consistent with the order of the activity for OCM reaction. We suppose that the oxygen exchange occurs through surface O- evolved at the high temperature.

180 atom 50 t

~ 160 atom

'

~

'

4t 460'

Time / min

'

's6o

Fig.6 Time course of 180 and 160 concentration in gaseous oxygen during the exchange reaction with MgO at 973K

404 Table. 1 Relative ratio of the surface active oxygen number* against the surface oxygen number (Oact) and relative ratio of exchangeable oxygen number in 24h** against total lattice oxygen number (O) as a function of temperature. 673K

773K

873K

Oact

0.22

0.2

0

0.006

0.017 0.037

1.0

973K

1073K

1.4

2.1

0.042

0.079

3.7

8.8

1173K

Surface Area m2/g

MgO

23.5 m2/g

Oact Li/MgO O

x

-

-

35 0.55 m2/g

x

-

1•

-4 8.6x10 -3 0.018

0.14

x: Isotopic exchange was not observed. *:N(1-al)/o~l in Fig.7, **:N(1-ct2)/ct2 in Fig.7 180.,

Gas phase

JV]F]VIF MgO

Active surface phase

Less active bulk phase

160

Fig.7 Two stage exchange model Actiive phase ( 1): l+ctl R1

fel-f

t = I n ........

al N Less active phase (2): l+ct2 R2

fe2-f t = l n ........

ct2

N

where data within 100min was used.

fel-f0

where data after 100min was used.

fe2-fel

R: Oxygen exchange rate between gaseous 02 and MgO (mol/min) N : Amount of oxygen atom in gaseous phase (mol) a : 160/18 0 in gaseous phase at equilibrium t: time f0, f, fe: 160 content in gaseous 02 at time 0, t ,and equilibrium (el=90min e2=24h) respectively

405 3.4. R e l e v a n c e to O C M reaction

As the result of these experiments, formation of O-, which was an active site for the OCM reaction, was assumed to relate with the DC conductivity, isotopic exchange, and TPD spectra. The schematic feature is shown in Fig.8. Balint and Aika [12] have pointed out the importance of water which occupies the vacancy (Eq. 19) and restores it when it is released (Eq. 19).

C2H6 <

2"CH3

2

12(~ I ,t(~

/

2Lir~gO~

~

-H~

.: + o~l-, ~ - .,~ o "Eq.1 9 ~

~,~

CH4

\

2Oo 1""'~~ Eq2 '~irr~er hole

,

, ~q.~

I

"

+Oo +Ool

l+20z

iVo+O~l

Fig.8 Active site generation on MgO system and relevance to the OCM reaction

CONCLUSIONS We briefly reviewed the general mechanism of OCM catalysis and pointed out the importance of morphology (physical factor) in addition to the active site generation (chemical factor). At higher temperature, the active site was proposed to be generated, which was shown by D.C. conductivity, TPD, and isotopic exchange activity experiments. This work has been carried out as a research project of the Japan Petroleum Institute commissioned by the Petroleum Energy Center with the subsidy of Ministry of International Trade and Industry. REFERENCES

1.E. Iwamatsu, T. Moriyama, N. Takasaki, K. Aika, J. Catal., 113 (1988) 25. 2.E. Iwamatsu and K. Aika, J. Catal., 117 (1989) 416.

406 3.E.E. Micro, J. M. Santamaria, E. E. Wolf, J. Catal., 124 (1990) 465. 4.K. Aika, N. Fujimoto, M. Kobayashi, E. Iwamatsu, J. Catal., 127 (1991) 1. 5.N. Yamagata, Y. Abe, K. Igarashi, T. Ishikawa, M. Sahara, S. Okazaki, Chem. Lett. (1990) 1893. 6.S.K. Agarwal, R. A. Migone, G. Marcelin, J. Catal., 123 (1990) 228. 7.K. Aika and J. H. Lunsford, J. Phys. Chem., 81 (1977) 1391. 8.H. Katherin, F. Freund and J. Nagy, J. Phys. Chem. Solids, 45 (1984) 1155. 9.J. -L. Dubois and C. J. Cameron, Appl. Catal., 67 (1990) 49. 10.Ioan Balint and Ken-ichi Aika, in "Natural Gas Conversion ", H. E. Curry-Hyde and R.F. Howe (eds.), Elsevier Science B.V., Amsterdam, 1994, pp.177. 11.F.A. Kroger, The Chemistry of Imperfect Crystals. North Holland, Amsterdam (1964). 12.I.Balint and K.Aika, J.Chem. Soc., Farad. Trans., (1995) in press.

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All rights reserved.

407

VARIOUS CHARACTERISTICS OF SUPPORTED CoPc ON A1203 , SiO 2 AND SiO2-AI203 AS SELECTIVE CATALYSTS IN THE OXIDATIVE DEHYDROGENATION OF CYCLOHEXENE.

Salah A. Hassan, Salwa A. Sadek, Samy M. Faramawy, Mohamed A. Mekewi

Department of Chemistry, Faculty of Science, Ain Shams University, Cairo, Egypt ABSTRACT Various characteristics of CoPe supported on AI203 , SiO2 and SiO2-AI203 of varying SiO2 mole ratio were investigated. Different pore systems were found to exist depending on the support permitting entrance of CoPe molecules (SiO2 and silica-rich samples) or not (A1203 and alumina-rich samples). The major fraction of CoPe appeared as combined with the support surface was much less in silica and silica-rich samples. From ehemisorption stoichiometries, CoPe molecules are proposed to lie fiat on alumina and alumina-rich surfaces and inclined-edge oriented on silica and silica-rich surfaces. The studied catalysts exhibited good seleetivities in the oxidative dehydrogenation of cyclohexene toward benzene. A mechanism was suggested where the formation of an intermediate involving active oxygen species was considered the rate determining step. 1. INTRODUCTION Cobalt phthalocyanine (CoPc) supported on different carriers, e.g., A1203, SiO2 and SiO2AI203 of varying SiO2 ratios (especially those resembling the zeolitic matrices) represent catalyst systems of special interest, usable in widely different reactions related to many chemical industries [1-8]. Physicochemical characteristics of these catalyst systems have attracted the attention of several authors, in particular their adsorptive characteristics towards 02, N2, CO, NH 3 [9-11]. Inspection of literature reveals a serious lack in studies on interaction between supported complex phase and the surface of the above mentioned carriers as a controlling factor to the catalytic behavior. The present study was undertaken to correlate such interactive parameters derived from dissolution of free non-interacted CoPc complex and chemisorption of oxygen with the catalytic efficiencies in the oxidative dehydrogenation of cyclohexene. 2. EXPERIMENTAL

2.1. Preparation Pure Cobalt Phthalocyanine (CoPc) was prepared according to a method described elsewhere [ 12]. 7-A1203, SiO 2 and SiO2-AI203 of different mole% SiO 2 ranging from 32.6 to 97.1 were prepared by conventional methods [13, 14] and calcined at 600 o c for 4 hrs. The exact mole% with respect to both SiO 2 and A1203 were estimated according to [15] as expressed in terms of mole% SiO 2. Supported catalyst samples of different complex loadings, viz.,-~0.2- 10% w/w were prepared by impregnation of chosen support in n-butylamine solution of CoPc of appropriate concentration under N 2 atmosphere. The excess solvent was removed

408 by vacuum distillation. The solid catalysts were thermally pretreated at 200 oc for 10 hrs in a stream of pure nitrogen. The following designations were used for the different materials under study :CoPc = CP, AI203 = A, SiO2= S, SiO2-AI203 of 32.6 mole% SiO 2 = 32.6 SA, 0.6 wt% CoPc/SiO2-AI203 of 32.6 mole% SiO 2 = CP 0.6/32.6 SA and similar abbreviations were used for the other supported catalyst samples.

2.2. Characterization Surface texture was studied by the aid of adsorption-desorption isotherms o f N 2 at196 o c using a conventional volumetric apparatus. Specific surface areas and pore volumes were calculated from the isotherms by applying the BET equation [16]. The porosity was detected through the corrected modeless method [ 17]. Interaction characteristics were studied through dissolution of free non-interacted fraction of CoPc from the support surface and through chemisorption of 0 2 at 300 ~ The dissolution was performed at 25 oc in a mixture of 1:1 DMF-CHCI 3 and the extracted fraction was determined spectrophotometrically at ~=661 nm. For chemisorption study, the adsorption isotherm (first isotherm) was obtained in each case by plotting the amount of 0 2 adsorbed at 300 oc against the 0 2 equilibrium pressure, P~q, up to 250 mm Hg. The resorption isotherm (second isotherm) was measured after outgassing for one hour at the same temperature. The total chemisorption values were taken as the difference between the first and the second isotherms. To minimize adsorption on the support and maximize adsorption on the supported complex, the total chemisorption values on pure support were subtracted from the total chemisorption producing the so called net adsorption values (an) on the supported complex as expressed in m mol 0 2 g-1 CoPc. These values were used to calculate the degrees of dispersion (D), the chemisorption stoichiometries and specific surface areas of supported CoPc phase [ 18,19].

2.3. Catalytic activity The Catalytic activity was examined in the oxidative dehydrogenation (OXD) of cyclohe• using a pulse- chromatographic technique [20] (GC of Perkin-Elmer model 910 equipped with FID) in the temperature range, 280 - 350 oc. A stainless steel micro reactor (10 cm length x 0.3 cm id) packed with 0.1 g sample of the catalyst was placed in the injection part of the chromatograph. The catalyst was first activated in a stream ofN 2 (40 ml min -1) for 12 h at 350 oc. Pure oxygen gas was passed over the catalyst at fixed flow rate of 15 ml min-1 for 1 hr while pure nitrogen gas was flowing at a rate of 35 ml min-1. To achieve steady state of the catalyst, several pulses (2gl) of purified cyclohexene (PROLABO, 98%) were injected in this stream till two identical conversions were obtained. The effluent from reaction zone was carried to a stainless steel column (2.7 m length x 0.3 cm id) packed with 5% bentone 34 and 5% diisodecylphthalate on chromosorb p (80 - 100 mesh). The temperatures of the column and the detector were 70 and 250 oc, respectively. The products obtained were: cracked products (CP), methylcyclopentene (MCP), benzene (B) and toluene (T). In separate experiments, the effect of changing the concentration of peroxide in cyclohe• feed stock was tested by passing a flow of pure oxygen for 2 and 6 hrs. The concentration of produced peroxide was 0.26 and 0.8 g mol L-1, respectively as determined iodometricaly [21 ].

409 Table 1 Nitrogen adsorption data of parent supports and different supported catalyst samples Sample g-alumina 32.6 SA 52.3 SA 97.1 SA Pure silica

BET-C constant 130 77 64 59 132

SBE T m2 g-1 167.1 283.4 340.6 371.2 330.7

Vp .. ml g-1 0.260 0.458 0.602 0.707 0.700

rpCp Ao 31 32 36 38 40

St m 2 g-1 165.2 283.4 340.3 370.1 330.4

CP CP CP CP CP

0.6/A 0.6/32.6 SA 0.6/52.3 SA 0.6/97.1 SA 0.6/S

77 46 45 77 69

141.7 188.7 242.4 283.4 316.4

0.252 0.330 0.427 0.514 0.590

36 35 35 36 37

141.3 190.4 242.1 283.2 318.1

CP CP CP CP CP CP CP

0.22/97.1 0.40/97.1 0.60/97.1 0.80/97.1 1.20/97.1 2.00/97.1 10.4/97.1

69 88 77 151 142 67 96

316.4 329.4 283.4 288.5 307.4 324.8 301.8

0.660 0.676 0.514 0.555 0.625 0.676 0.606

42 41 36 39 41 42 40

317.4 330.3 283.2 292.2 310.3 325.1 302.1

SA SA SA SA SA SA SA

3. RESULTS AND DISCUSSION

3.1. Texture of the catalyst investigated The adsorption-desorption isotherms of nitrogen at -196 oc obtained on all the catalysts under investigation were mainly of Type IV of Brunauer's classification [16], exhibiting hysteresis loops closed at P/Po ranging between 0.25 and 0.55. The adsorption data are summarized in Table 1, including BET-C constant, specific surface area(SBET) , total pore volume (Vp), estimated from the saturation values of the adsorption isotherms and average pore radius (rcp), assuming cylindrical pore model for which superscript (cp) was used. The results indicate that ],-alumina support has a relatively small specific surface area, total pore volume and consequently small average pore radius as compared with silica support. Measurable increase in all these parameters is observed by increasing silica content from 32.6 SA to 97.1 SA. Upon loading 0.6% w/w CoPe on the various supports under study, a decrease in specific surface area is generally observed relative to those areas of the corresponding pure supports. However, some pore widening may occur in alumina and alumina-rich supported samples and some narrowing in silica and silica-rich supported samples. Supporting CoPe with different wt% (from --- 0.2 up to - 10% w/w) on one and the same support (97.1 SA) indicated that no effect on surface parameters was observed up to 0.4% w/w. However, for samples of 0.6 and 0.8 % w/w CoPe some decrease is observed in pore dimensions and surface areas as compared with those samples of lower and higher loadings. This should be linked with possible physical entrance of some CoPe molecules (of an average radius, 11 -12 A~ [22])in the

410 produced pore system and with the possible mode of orientation and dispersion of the supported molecules on the surface. For pore structure analysis, the t-curves of Sing et al. [23] on non porous alumina was used for )'-alumina support, alumina-rich support (32.6 SA) and their corresponding supported catalyst samples. However, the t-curve on non-porous hydroxylated silica of Sing et al. [24] was used as a reference one for silica support, silica-rich supports (52.3 SA, 97.1 SA) and their supported catalyst samples. Vl-t plots were constructed, where V 1 is the volume of N 2 adsorbed (ml g-l) and t is the statistical thickness (A). From the slope of the straight line obtained passing through the origin, the specific surface area, St, was calculated. The reasonable agreement between SBE T and St is the main criterion for the correct choice of the t-curve used in the analysis (Table 1). The analysis of the produced pores in each case was carried out using a computer program written in FORTRAN IV language [25] for an IBM computer assuming a cylindrical pore model (cp). The pore volume distribution curves shown in Fig. 1, representing the distribution of pore volumes (Av/Ar) as a function of the most frequent hydraulic core radii (rcCP, A ), indicate that silica and silica-rich support (97.1 SA) exhibit more fraction of more wider pores. Some fraction of micropores seems to exist in ),-alumina support (rcCP= 11 A) and similarly in 32.6 SA and 52.3 SA (rcCP= 14 A). Upon supporting CoPcwith constant 0.6% w/w on the studied supports, the distribution curves indicate a general decrease in the most frequent hydraulic core radii as compared with those of pure supports. For y-alumina support, no detectable change in the pore dimensions could be observed. The variation in CoPc content on 97.1 support does not affect appreciably the most frequent hydraulic core radii nor their fractions. In conclusion, one may suggest that, in silica and silica-rich supported samples, the pore system permits preferential entrance of some CoPc molecules, affecting sensibly the pore dimensions. In alumina and alumina-rich supported samples, the pore system, with its dimensions less than those of CoPc molecules, does not permit free entrance of complex molecules. Here, CoPc molecules may be forced to occupy positions in the defective structure of alumina. All these findings should be reflected on the mode of surface dispersion of the supported CoPc molecules. 3.2. Estimation of fraction of active species by dissolution method Dissolution technique was applied previously in our laboratories to elucidate the nature of metal-support interaction in the metal and metal oxide systems [ 18,26]. The fraction of free CoPc (wt%), i.e., extracted in 1:1 CHCI3-DMF solution is given in Table 2. It is clear that, in all cases, the major fraction of CoPc is combined with the support. The combined fraction may be assigned to some chemical interaction or diffusion into various sites of the support. For the catalyst samples of the same complex loading, viz., 0.6% w/w, the fraction of free non interacted CoPc shows its lowest level on pure alumina and increases gradually by increasing SiO 2 content. This means that the extent of interaction decreases in silica and silica-rich supported samples. Moreover, by increasing the complex loading, the fraction of free non interacted CoPc decreases gradually. This may be linked with the formation of clusters less susceptible to dissolution in the mentioned solvent mixture.

411

Pure

s

8R.8

5a.8

HA

_ 07" lJuA~

HA

'

\ ]

,

i

f

o.e/A 0.06

o.m/aa.s

o.e/en.a

liA

liA

o.e/eT.t

liA

f

o.m/li

-

/x

" -

~

0.03

- - ~ l i

-

I

o.oo

o.aa/oT.x

.

.

.4,,I

. I

liA

o.e/oT.x

_

.

.

t

.

I

1~ I

I

x.aloT.x

liA

I

J

tliA Ilia

liA

-~

o.o9

0.06 0.03 0.00

0

0

20

I 30

3JO

I0

0

0

10

20

I 30

~ 10

20

t 30

r.(z')

Figure 1. Pore volume distribution curves for parent supports and different supported catalyst samples. Table 2 Interaction characteristics as estimated from dissolution and oxygen chemisorption data for different supported catalyst samples. Catalyst sample CP CP CP CP CP

0.6/A 0.6/32.6 SA 0.6/52.3 SA 0.6/97.1 SA 0.6/S

CP CP CP CP CP CP CP

0.22/97.1 0.40/97.1 0.60/97.1 0.80/97.1 1.20/97.1 2.00/97.1 10.4/97.1

SA SA SA SA SA SA SA

wt% free CoPc 0.12 0.16 0.26 0.34 0.33

an mmolO2/gCoPc 13.76 31.99 26.04 10.05 2.98

SO2 A/0 2 285 166 258 781 2943

Stoich. 0.77 0.44 0.69 2.10 7.90

SCoPc m2/g 1410 5734 2976 377 30

0.27 0.15 0.34 0.11 0.09 0.05 0.04

11.06 8.93 10.05 5.02 4.20 1.56 0.65

2141 1532 781 1193 1012 1727 735

D 1.09 4.43 2.29 0.29 0.02

5.80 4.20 2.10 3.20 2.70 4.70 2.00

151 168 377 124 121 26 26

0.12 0.13 0.29 0.09 0.09 0.02 0.02

412 3.3. Surface dispersion of supported CoPc molecules from chemisorption data

It was suggested that 0 2 can be used as a selective adsorbate for the determination of degree of dispersion of metal phthalocyanines on the support surface [27,28]. Such adsorption was found to occur on the central metal ion as the active center [28,29]. The obtained net adsorption values (a n, in molecule 02 g-1 sample) were first used to calculate the surface area available for one 02 molecule in each case adopting the values of SBE T (m2/g sample) as follows: So2 = SBET/an (A2/O2 molecule)

(1)

Since monomeric CoPc is considered as a plane quadratic molecule of an area 123 A 2 [28] (i.e., one Co ion is present per 123 A2), and assuming equal distribution of(100), (01~ and (001) of the crystallite surfaces, one can conclude that one Co ion is situated per 369 AZarea of the surface. From this assumption and the available area for one 0 2 molecule, the number of 0 2 molecules adsorbed per one Co ion (i.e., stoichiometry) could be calculated as : Stoichiometry = S02/369

(2)

Knowing the stoichiometry of adsorption, the specific surface area of supported CoPc could be calculated as : SCoPc = a n.NA.tYo2/St~176

(3)

Where, NA is the Avogadro" s number and ~O2 is the cross-sectional area of oxygen molecule, being taken as 13.1 A 2 [30]. The degree of dispersion (D) could be detected from the calculated specific surface area (SCoPc) according to the following equation 9 SCoPc = (NA. O'CoPc/M).D

(4)

Where, t~CoPc is the cross-sectional area of CoPc taken as 123 A2, and M is the molecular weight ofCoPc (M = 571.47). The characteristic chemisorption data obtained for the various supported catalyst samples are summarized in Table 2. It is evident that, the O2-net adsorption (an) has in general lower magnitude for catalyst samples supported on pure silica and silica-rich support and higher magnitudes for catalyst samples supported on alumina and alumina-rich supports. However, these adsorption values decrease markedly as the CoPc content increases on one and the same support (97.1 SA); being related most probably to the mode of surface complex dispersion. The specific surface area of supported CoPc shows unrealistic values in the case of supported samples on alumina and alumina-rich supports. Consequently, the calculated degrees of CoPc dispersion are relatively high. On the other hand, the specific surface area of CoPc supported on silica and silica-rich support show much lower values with much lower degrees of dispersion. These parameters decrease markedly as the complex loading increases up to 10.4% w/w. It is of interest to notice that, the stoichiometry of adsorption of oxygen on samples of supported CoPc on alumina and alumina-rich supports is almost close to 1. However, in samples of CoPc supported on silica and silica-rich supports, the stoichiometry becomes >2. These findings indicate clearly that supported CoPc molecules interact with the different support surfaces under study along widely different ways.

413 Based on the above data, one may suggest that, CoPc molecules seem to lie flat on the surface of alumina and alumina-rich supports. This may arise most likely from the interaction of Co 2+ with AI3+ centers (with possibility of N(of CoPc)- weaker acid sites (of alumina) interaction [31,32]) resulting in a stoichiometry value- 1. In this case, Co ion in molecular dispersed CoPc is accessible for 0 2 from one side. However, the higher stoichiometries obtained in the case of samples on silica and silica-rich support may reflect that CoPc molecules are oriented on the support surface. This implies that CoPc molecule is accessible for 0 2 from different sides (both sides of central Co ion and exposed peripheral N" s). Such increased accessibility is favored by inclined edge interactions mainly between nitrogens of the macro-cycle of Pc and the hydroxyl groups of the support surface (Br~nsted interaction [32]). One cannot exclude the possibility of penetration of small 0 2 molecules (of diameter 2.05 A) between stacks of oriented CoPc molecules [28]). Again, the low degrees of dispersion of CoPc on these supports, i.e., silica and silicarich may reflect the presence of Copc as aggregates on the support surface. CoPc molecules on the surface become more crowded by increasing the loading to 2.0 and 10.4 % w/w CoPc.

3.4. Catalytic Activity In the catalytic oxidative dehydrogenation of cyclohexene (OXD), both the flow of molecular oxygen in the system [33] and the presence of peroxides [34] in cyclohexene feed stock were found to be of prime importance as operational conditions. (a) Role of molecular oxygen In order to elucidate the role of molecular oxygen on the behavior and mechanism of the dehydrogenation process, the reaction was first studied on CP 0.6/97.1 SA catalyst at temperatures >300 ~ in absence of oxygen. In this case, only N 2 was used as a carrier gas at a flow rate of 50 ml min-1. The obtained chromatogram revealed that only the pre-injected cyclohexene was produced at longer retention time where the broad peak obtained ensured its strong adsorption on the catalyst surface. The conversion has been observed only when several 0 2 pulses were introduced. These observations could indicate that an oxygen flow must be used in the reactant stream in order to prevent possible deactivation of the solid catalyst caused by excessive adsorption of cyclohexene. Here, the reaction may be expected to proceed at lower temperatures as it is facilitated by the pulses of 0 2 applied. The O2/N 2 ratios were varied as follows; 5/45, 10/40, 15/35, 20/30, and 25/25. It was noticed that with ratios below 15/35, the retention times of the products were markedly long. The obtained peaks were relatively broad and asymmetric which might lead to considerable error in peak area measurements. However, O2/N 2 flow rates above 15/35 were not recommended as they could affect the sensitivity of the flame ionization detector (FID) and the activity of the stationary phase. Therefore, O2/N 2 flow rate of 15/35 could be used as the optimum ratio giving almost reasonable results (the total flow rate being constant at 50 ml minl). (b) Role of peroxide Fig. 2a reveals that the conversion of cyclohexene at 300 ~ on CP 0.22/97.1 SA catalyst increases markedly with the increase of peroxide concentration. The product distribution

414

1)

'

t ~ ~b"

I

i

,

oo o ..,,.,i m

6o

o

4o

K

2 o [ ~

.o

,

I

,

J

,

40

o 0.0

0.2

0.4

Conc.

of

peroxide

0.6

0

(l[.mol/L)

Figure 2. a) Effect of peroxide concentration on % conversion of OXD of cyclohexene on CP0.2/97.1SA catalyst at 300~ using O2/N 2 flow rate of 15/35 ml min-1, b) Effect of peroxide concentration on product distribution of OXD of cyclohexene normalized to % conversion (g). [Benzene (B), Cracked Product (CP), Methycyclopentene (MCP)] normalized to % conversion (T) as a function of peroxide concentration is shown in Fig. 2b. It is obvious that the active oxygen produced from the decomposition of the peroxide affects the different pathways of the reaction in different manners. It enhances the oxidative dehydrogenation (OXD) pathway at the expense of both cracking and isomerization ones. (c) Role of reaction temperature The oxidative dehydrogenation of cyclohexene was carried out at temperatures ranging between 280 and 350 oc; the range being limited by the probable change in CoPc structure (including fragmentation) at t > 350 oc and by maximum adsorptivity of cyclohexene at t < 280 oc. The obtained data over the various catalysts under study indicated that benzene was generally the main reaction product, its concentration increases by increasing the reaction temperature. Both the cracked products (mainly methane) as well as the isomerization product (namely, methyl cyclopentene) were obtained only in small amounts being also increased by temperature.

415 O0

.~/

,

40

!-~ RO

A w

-

10

o

...

,

..8 i

,

~ e n P e

1~. l.,-,m.dlnl

,

1'.o I

,

,.. i

,

,,'.~

,

.'o

{ w ] w }

Figure 3. Effect of different CoPc loadings/97.1 SA on selectivity toward benzene at 330 oc. (d) Role of CoPc loading The effect of CoPc loading on 97.1 SA support (the peroxide activator content was selected at 0.26 g mol L-1) is illustrated in Fig. 3, where the catalyst sample containing 0.6 % w/w CoPc is shown to posses the maximum activity and selectivity toward benzene. Both cracked and isomerization products were found to suffer a considerable decrease in samples of 2.0% and 10.4% w/w CoPc which may confirm that both reactions are functions of the support (i.e. mainly of its acid sites). Considering the catalyst samples containing a constant 0.6% w/w CoPe over different supports (peroxide activator content being 0.8 g mol L -1) one may conclude that: i. The activity of samples supported on silica-alumina combinations seems to be much more than those on pure silica or pure alumina. This may be related to the produced pore systems affecting indirectly the distribution and dispersion of supported CoPc molecules. ii. The most active and selective samples are those ones supported on 52.3 and 97.1% SiO 2 of larger surface areas. (e) Suggested mechanism A reaction mechanism for the OXD of cyclohexane over Cu(II)-Y Zeolite, has been proposed in which the adsorption of oxygen was found to be the rate determining step [35]. A similar trend was observed for the OXD of cyclohexene over Co(II)-Y Zeolite [33], as a high adsorbed oxygen reactivity and benzene selectivity were found in both reactions. Five different reaction mechanisms, distinguished by the rate determining step, were proposed [36]. These mechanisms may possibly explain our experimental observations.

416 9 Mechanism 1 : Oxygen adsorption is the rate-determining step. 9 Mechanism 2 (Langmuir-Hinshelwood mechanism I) : The surface reaction of adsorbed oxygen and adsorbed cyclohexene on separate active sites of the catalyst is the rate determining step. 9 Mechanism 3 (Langmuir-Hinshelwood mechanism II) : The surface reaction of adsorbed oxygen and adsorbed cyclohexene on the same active site is the rate determining step. 9 Mechanism 4 (Rideal mechanism I) : The reaction of adsorbed cyclohexene with oxygen in the vapour phase is the rate determining step. 9 Mechanism 5 (Rideal mechanism II) : The reaction of adsorbed oxygen with cyclohexene in the vapour phase is the rate determining step. According to our experimental observations given above, adsorption of molecular 0 2 should be excluded as a rate determining step since it became evident from our chemisorption study that this step proceeds in a high rate leading to rapid saturation on almost all the catalysts under study. The strong adsorption of cyclohexene on the catalyst surface which seems most probably to displace the pre-adsorbed molecular 0 2 ensures the idea that this oxygen should adsorb only weakly on the catalyst surface. It is to be recalled that cyclohexene hydroperoxide is decomposed into active oxygen species (O*) and cyclohexene [34]. It can be thus considered that both of these products may be adsorbed on the same catalyst site favoring the formation of an active intermediate of the formula C6H10 -Cat-O*. This step seems to take place in some complicated manner as evidenced by the activation parameters (AEav.--13 kcal/mol, ASav.--59 e.u.). The higher conversions and selectivities toward benzene, i.e., more shifted toward OXD pathway may be a result of the increase in the concentration of the active intermediate indicating that the step of its formation is most probably the rate determining step. This runs in good harmony with the suggested scheme of the mechanism 9 (i) (ii)

O2(molec. ) + Cat. C6H10 + Cat.

(iii)

C6H9OOH

(iv)

C6H 10+O*+Cat-

(v)

C6H 10-Cat'O*

02 (ads.)-Cat. C6H 10( ads.)-Catrapid slow 1/2 02 ~ ~ rapid

C6H10 + 20* C6H10-Cat.-O*adsorbed on the same active site C6H 6 + Cat. + 2H20

Possible rapid side reactions may also be indicated as follows :i) Isomerization C6H10 -Cat- (acid sites)

02 >

CH3C5H 7 + Cat.

ii) Cracking C6H10 "Cat. (acid sites) C6H 6 + CH 4

13/2 02

1/2 02

> CH 4 + 5CO 2 + 3H20 >

C6H5CH 3 + H20

417 both being dependent mainly of the support acid sites and are taking place mostly in presence of molecular oxygen. 4. CONCLUSION From the obtained results one may conclude the following : * Silica and silica-rich supported samples, and alumina and alumina-rich supported samples exhibit widely different pore systems. * Major fraction of supported CoPc appear to be combined with the support surface. The formation of surface aggregates or clusters is favored in highly loaded samples. * Interaction goes along widely different ways in the catalyst systems studied; CoPc molecules lie flat on the surface of alumina and alumina-rich supports and edge oriented on the surface in silica and silica-rich samples. 9 All studied catalysts exhibit good selectivities in the OXD of cyclohexene toward benzene, particularly samples of 0.6% w/w CoPc on 52.3 and 97.1 SA combinations. The flow of molecular oxygen and the presence of peroxides in cyclohexene feed stock are of prime importance as operational conditions. A mechanism is suggested where the step of formation of an active intermediate of the type C6H10 -Cat'O* is most probably the rate determining step. REFERENCES

1. K. Tsuii, M. Imaizumi, A. Oyoshi, I. Mochida, H. Fujitsu and K. Takeshita, Inorg. Chem. 21 (2) (1982) 721. 2. I. Mochida, A.Yasutake, H. Fujitsu and K. Takeshita, J. Phys. Chem., 86 (1982) 3468. 3. H. Daud and S.A. Barisenkova, Deposited Doc., VINITI, 11 (1983) 5899. 4. H. Diegruber, P.J. Plath and G. Schulz-Ekloff, J. Mol. Catal., 24 (1984) 115. 5. N. Herron, G.D. Stucky and C.A. Tolman, J. Chem. Soc., Chem. Commun., (1986) 1521. 6. G. Schulz-Ekloff, D.Wohrle, V. Iliev, E. Ignatzek and A. Andreev, Stud. Surf. Sci., 46 (1989) 315. 7. R.F. Parton, L. Uytterhoeven and P.A. Jacobs, Stud. Surf. Sci., 59 (1991) 395. 8. T. Buck, D. Wohrle, G. Schulz-Ekloff and A. Andreev, J. Mol. Catal., 70 (1991) 259. 9. T.G. Boisova and B.V. Romanovskii, Vest. Mosk. Univ., Ser. 2: Kim. 18 (6) (1977) 732. 10.B.V. Romanovskii, R.E. Mardaleishvili, V. Yu. Zakharov and O.M. Zakharova, Vest. Mosk. Univ., Kim. 133 (5) (1978) 524. 11.Z. Weide, Z. Ruiyun, Y. Xinghai and W. Yue, Yingyong Huaxue, 10 (4) (1993) 39. 12.H. Junge and H. Bruenemann, (BASF A.-G.) Ger. Often. DE3, 106, 541 (C1. CO 91347106), 21 Oct. (1982). 13.D. Basmadjian, G.N. Fulford and B.I. Parsons, J. Catal. 1 (1962) 547.

418 14. J.H. De Beor, Faraday Discussion (1971) 52. 15. A.I. Vogel, Quantitative Inorganic Analysis (1977). 16.S. Brunauer, P.H. Emmett and E.J. Teller, J. Amer. Chem. Soc., 60 (1938) 309. 17.R.Sh. Mikhail, S. Brunauer and E.E. Bodor, J. Colloid Interface Sci., 26 (1968) 45. 18. S.A. Hassan, M.A. Mekewi, F.A. Shebl and S.A. Sadek, J. Mater. Sci., 26 (1991) 3712. 19. S.A. Hassan, M. Abdel-Khalik and H.A. Hassan, J. Catal., 52 (1978) 261. 20.A.K. Aboul-Gheit, A.M. EI-Fadly, S. Faramawy, S.M. Abdel-Hamid and M. AbdelKhalik, Erdol unfKohle Erdgas Ptrochimie, 40 (1987) 315. 21. C.D. Wagner, R.H. Smith and E.D. Peters, Anal. Chem., 19 (1974)976. 22.J.H. Zagal, M. Paez, J. Stum and S.U. Zanartu, J. Electroanal. Chem. 181 (1984) 295. 23.D.A. Ryne and K.S.W. Sing, Chem. Ind., (1969) 918. 24.J.D. Carruthers, P.A. Cutting, R.E. Day, M.R. Harris, S.A. Mitchell and K.S.W. Sing, Chem. Ind., (1968) 1772. 25.R.Sh. Mikhail, S.A. Selim and A. Goned, Egypt J. Chem., 18 (1975) 957. 26. S.A. Hassan, F.H. Khalil and F.G. E1-Gamal, J. Catal., 44 (5) (1976). 27.J.P. Contour, P. Lefant and A.k. Vijh, J. Catal., 29 (8) (1973). 28.F. Steinbach and H. Schmidt, J. Catal., 39 (1975) 190. 29.F. Steinbach and M. Zobel, Z.Phys. Chem., 87 (1973) 142. 30.S.J. Gregg and K.S.W. Sing, Adsorption Surface Area and Prosity, Acad. Press, Landon, New York (1967). 31.F. Campadelli, F. Cariati, P. Carniti, F. Marazzoni and V. Rgaini, J. Catal., 44(1976) 167. 32. G. Mercati and F. Marazzoni, Inorg. Chim. Acta, 25 (1977) L 115. 33.E.P. Garcia, M.R. De Goldwasser, C.F. Parra and O. Lead, J. Applied Catalysis, 50 (1989) 55. 34.R.K. Srivastava and R.D. Sreivastava, J. Catal., 39 (1975)317. 35.I. Mochida, J. Tetsuji, K. Akio and S. Tetsuro, J. Catal., 36 (1975) 361. 36.K.J. Laidler, "Catalysis" (P.H. Emmett, Ed.), Vol.1, Chaps. 3,4 and 5. Reinhold, New York, (1954); P.G. Ashmore, "Catalysis and Inhibition of Chemical Reactions" Chap. 7, Butterworth, London (1963); T. Kell, "Kinetics in Catalytic Reactions" (Catalytic Engineering, Vol. 1), p. 129, Chijinshokan, Tokyo (1969).

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All fights reserved.

DEHYDROGENATION OF PROPANE COMPARATIVE CHARACTERIZATION CATALYSTS

419

OVER CHROMIA/ALUMINA: A STUDY OF FRESH AND SPENT

A t a u r R a h m a n a and M o t a h e r u d d i n A h m e d b

apetroleum and Gas Technology Division, bEnergy Resources Division Research Institute, King Fahd University of Petroleum & Minerals, Dhahran 31261, Saudi Arabia. ABSTRACT

Chromia/alumina catalyst with 5% chromia loading was prepared and used in propane dehydrogenation reaction. The fresh and spent catalysts were then analyzed by a number of techniques to obtain a variety of relevant data. Distributions of Cr over the alumina support granules were measured by the Particle Induced X-ray Emission technique using a scanning proton microbeam (microbeam-PIXE). Thermogravimetric (TG) analysis together with Differential Thermal Analysis (DTA) were performed to measure transformation characteristics due to calcination. Temperature Programmed Reduction (TPR) was employed for the fresh and spent catalysts as well as bulk CrO3 to deduce metal-support interaction. X-ray Photoelectron Spectroscopy (XPS) was used to measure the oxidation states of the chromium ions. The data on the fresh and spent catalysts were compared to evaluate the properties of the catalysts. 1. I N T R O D U C T I O N Supported chromia catalysts have a wide range of applications such as hydrogenation and dehydrogenation reactions of hydrocarbons, the dehydrocyclization of paraffins, dehydroisomerization of paraffins, olefins, and naphthenes, and the polymerization of olefins [1-3]. In order to improve the activity and selectivity, characterization of some critical parameters for both fresh and spent catalysts is necessary. Particle Induced X-ray Emission (PIXE) technique has been shown to have great potentials for catalytic research [4,5]. The impact of high energy protons upon a sample causes the emission of characteristic X-rays which can be used for elemental analysis. The uniform distribution of elements across a catalyst surface is an important factor for catalyst design. PIXE technique with scanning microbeam can be used to measure elemental distributions with a spatial resolution of the order of micrometers. From X-ray Photoelectron Spectroscopy (XPS) measurements, the valencies of metal ions on thin surface layer (about 50*) of supported oxide as well as the metal support interaction can be deduced. The change in mass of a substance as a function of temperature can be measured using the Thermogravimetric (TG) technique. The amount of heat evolved or absorbed and the temperature at which these changes occur within the material can be estimated by Differential Thermal Analysis (DTA). Thus, by combining the results of TG and DTA, it is possible to deduce the transformation phenomena that occur when a catalyst is heated [6]. Temperature

420 Programmed Reduction (TPR) technique is a very useful procedure for investigating interactions between a supported metal oxide and a catalyst surface. To date, little information is available on the comparative study of fresh and spent chromia/alumina catalysts in propane dehydrogenation. While our prime objective is not to study propane dehydrogenation reaction in itself, we would like to report in this paper, physical characteristics evaluated for both fresh and spent catalysts in propane dehydrogenation employing multiple characterization techniques. 2. EXPERIMENTAL

2.1 Catalysts Preparation and Catalytic test Catalyst samples with 5 wt% chromia were prepared by impregnating a commercial gamma-alumina with an aqueous solution of CrO3 by wet impregnation technique. The support was first heated at 500 ~ in air for 16 hours in a furnace. Aqueous solution of CrO3 was prepared with a prescribed amount of CrO3 to yield 5% chromia/alumina sample. The solution was allowed to be soaked in the support for 2 hours and then the excess water was removed using a rotary evaporator under vacuum at 80 ~ The samples were calcined at 300 ~ 500 ~ and 800 ~ respectively to see the effect of calcination. The dehydrogenation of propane was carried out in a fixed bed tubular reactor using 2g catalyst previously calcined at 500 ~ The reaction was conducted at atmospheric pressure and at 600 ~ using a gaseous mixture of 50 mol% propane in nitrogen at a total flow rate of 60 ml/min as a feed. Prior to the run, the catalyst was preheated in a 30 ml/min flow of nitrogen upto 600 ~ it was held at that temperature for 1 hour before propane was introduced. Reactant and effluent reaction products were analyzed using an on-line gas chromatograph.

2.2 Characterization Techniques The scanning nuclear microprobe facility on the tandetron accelerator of King Fahd University of Petroleum & Minerals (KFUPM) [7, 8] was used for the measurements of Cr distributions over the alumina support particles using the PIXE technique. Targets of cylindrical disc shape were formed from the prepared powder samples by embedding them in epoxy resin, drying and polishing to have a fiat surface. To avoid any charge build-up during proton irradiation, the surface in addition was coated with a thin carbon layer. A 2.5 MeV proton microbeam of about 5 ~tm spatial resolution was employed to scan the sample to produce chromium distribution maps on alumina support. An X-ray energy spectrum was also acquired at the same time to measure the relative A1 and Cr concentrations. The XPS spectra reported in the present work were obtained using a PHI 5300 XPS system from Perkin-Elmer equipped with a dual Mg/A1 anode and using unmononchromatized Mg K a radiation (1253.6 eV). Thermal analysis of the samples were studied on a Netzsch simultaneous thermal analyser, STA 429, from ambient temperature up to 1000 ~ at a heating rate of 10 ~ in a dynamic air atmosphere (150 ml/min) with alumina as a reference material. One hundred mg of sample was placed in an alumina crucible. The same weight of aluminium oxide (A1203), which undergoes no thermal change in the temperature range of the experiment, was placed in an identical crucible as a reference sample. The temperature of the sample was measured by thermocouples of platinum and of platinum plus 10% rhodium. The parameters recorded simultaneously were temperature (T), change in weight (TG) and difference in temperature between sample and reference (DTA). All

421 temperature p r o g r a m m e d analyses were performed using an automated catalyst characterization unit using 10% H2 in argon as a reducing gas mixture. The sample was first treated under a 30 ml/min of pure argon flow, while heating from 25~ to 500~ at 10~ ramp. It was then held at 500~ for 5 minutes, then cooled to 50~ under the same flow conditions. In the reduction step, 10%H2 in argon was flown over catalyst at 30 ml/min while ramping from 50~ to 800~ at 20~ It was then held at 800~ for 30 minutes. Five pulses consisting of 100 microliters of pure argon were injected into the carrier stream (10%H2 in argon). Both flows were 30 ml/min. After the reduction step, the catalyst was cooled down from 800~ to 25~ under 30 ml/min of argon flow. 3. RESULTS AND DISCUSSION

3.1 Catalytictest Figure 1 shows conversion of propane as well as selectivity to propene with time on stream. A maximum propane conversion of 68% was achieved. Conversion was found to decrease with time and it dropped to 52% in 150 min. This decrease was caused by coke deposition due to cracking of propane. Very rapid methanation was observed in first few minutes of the reaction time. At 68% conversion, the selectivity towards propene was 44%. It declined in a similar pattern to that of conversion.

70

-70

60

-60

g 50

-50

g 40

-40 .,>"

v

>

=~

30

-30

o (I)

1.~ 2 0

-20

10

-10

0

0

i

i

60

120

0 180

Time on Stream (min) Figure 1. Propane dehydrogenation over 5% chromia/alumina catalyst at 600~ and at atmospheric pressure.

422 3.2 Microbeam-PIXE

From the PIXE energy spectra, Cr/A1 atomic ratios were calculated using a quantitative X-ray analysis software and are presented in Table 1. The PIXE ratios are somewhat higher than the XPS values for all three samples. This is because PIXE analyses deeper into the sample than XPS which is only a surface analysis. Figure 2 shows the spatial distribution maps of Cr over an area of 540 ~tm x 540 ~tm on the target surface for 5% chromia/alumina calcined (fresh) and spent catalyst samples. In both samples, Cr was found to be distributed over the alumina base. However, there appears to be an increase in Cr/A1 atomic ratios for the calcined sample compared to either the uncalcined or the spent sample. It seems that Cr is released to the surface from inside the alumina pores due to the heating process. The ratio decreases again after the catalyst go through the dehydrogenation reaction. It is believed that due to coke formation, surface area of chromia in spent catalyst decreases, resulting in low Cr/A1 atomic ratio. Gorriz et at [2] also observed decrease in surface area caused by coke deposition at low chromia loading during propane dehydrogenation. This behavior of the Cr/A1 data is similar to those obtained with XPS measurements reported in Table 1.

Figure 2. Microbeam-PIXE maps of chromium distribution over 540 ~tm x 540 ~tm area of the 5% chromia/alumina fresh and spent catalyst used in dehydrogenation of propane. A 2.5 MeV proton microbeam of 4 ~tm spatial resolution was used to scan the area. Darker shades on the gray scale indicate higher concentrations.

423 3.3 X P S

The binding energies of Cr 2p peaks as well as the corresponding Cr/A1 Atomic Concentration (AC) values obtained from XPS measurements are shown in Table 1. Chromia remains in Cr 6+ oxidation state in both the calcined and uncalcined samples. Following propane dehydrogenation, the binding energy value of Cr 2p3/2 peak is nearly identical to that previously ascribed to Cr 3+ oxidation state [6], suggesting that Cr 6+ is reduced to Cr 3+. During propane dehydrogenation, the surface of the sample is appreciably covered by coke as judged by the considerable increase in the intensity of the carbon 1s peak. Carbon is believed to be deposited on Cr although the participation of the alumina carrier to coke formation is not surprising [2]. This coke deposit probably reduces the signal intensity from the underlying chromium thus causing a reduction in the Cr AC values. XPS studies on calcined samples show a sharp increase of Cr/A1 ratio at calcination temperature upto 500 ~ for 5% Cr/alumina sample, while the ratio remains unchanged at higher calcination temperature. This is shown in Figure 3. Table 1. XPS data for 5% chromia/alumina samples. Sample

Binding Energy Cr2p 3/2 (eV) 6+

Uncalcined Calcined at 500~ Spent (used) in propane dehydrogenation a) XPS and PIXE atomic ratios

(Cr/A1)2p AC a) x 103 XPS PIXE

3+

579.8 579.9 -

577.1

83.8 99.5 85.5

90.5 111.7 103.0

c) o

o 120• 100

0

.,.,.,,. ,.I...

,,,,, 8 0 0

"~: 6 0 0

40< 0

200

I

0

I

I

I

200 400 600 800 Calcination Temperature ( C )

1000

Figure 3. Effect of calcination temperature on chromia/alumina atomic ratio obtained from XPS measurements.

424

3.4 ThermalAnalysis The thermal analysis results of bulk CrO3, 5% Cr/A1 fresh and spent samples are summarized in Table 2. Most significant weight loss is observed at an endothermic peak of 490 ~ corresponding to a thermal decomposition of Cr 6+ to Cr 3+. It is believed that the main phase transformation occurs at about 500 ~ When Cr is supported on alumina, different observation is made. The fresh sample shows only one endothermic peak at 170 ~ correspondingto moisture loss from the support alumina. No other peak was observed at higher temperature indicating that a metal support interaction has occurred in the case of supported catalyst. For the sample used in propane dehydrogenation, a small endothermic peak at 140 ~ due to moisture loss from the catalyst is observed. The exothermic peak at 470 ~ is due to coke formed during dehydrogenation reaction. Table 2 Thermal analysis results of chromia/alumina samples. Sample

Bulk CrO3

5% Cr/A1 (Calcined at

Temperature

Thermal

Tmax of DTA

Weight

Total weight

Range (~

Effect *

Peak

Loss (%) Loss (%)

20- 300 300- 380 380 - 450 450 - 1000

Endo Exo Exo Endo

200 340 400 490

0.5 2.5 3.0 7.0

13.0

20 - 200

Endo

170

1.9

1.9

140 470

1.9 7.6

9.5

500 ~ 20 - 200 Endo 5% Cr/A1 300 - 650 Exo (Spent) * Thermal effect: exothermic or endothermic 3.5 TPR

As shown in Figure 4, the TPR of bulk CrO3 consisted of reduction peaks at 280 ~ 462 ~ and 585 ~ Unlike the thermal decomposition of Cr 6+ to Cr 3+ at about 500 ~ (XPS and TG), the hydrogen reduction of Cr 6+ to Cr 5+ occurs at about 280 ~ Peak at 462 ~ can be assigned to the reduction of Cr 5+ to Cr 3+. The reduction peak at 585 ~ corresponds the reduction of Cr 3+ to either Cr 2+ or to the metallic state. The example of the TPR of supported 5 wt% chromia on alumina in Figure 4 shows the marked effect of the support in broadening the profile to a different temperature. No other reduction peaks are observed suggesting that in the supported catalyst, chromium species formed are difficult to be further reduced compared to the unsupported chromium oxide, which is an indication of metal support interaction. When the sample is used in propane dehydrogenation, the catalyst is reduced from Cr 6+ to Cr 3+ during reaction (Table 1). TPR profile shows only one peak at 555 ~ which is comparable to the peak at 585 ~ of the bulk CrO3, suggesting reduction of Cr 3+ to Cr 2+ or metallic state.

425

555

A vd z

(c)

329

o ==,=, IX, ::3

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

z

o

~L ,

280 ...

(3 z

(b)

462

14,1

(3 0D~ r~

>..

-T-

O 0

I

i

I

i

200

400

600

800

1000

TEMPERATURE ( C )

Figure 4. TPR of (a) unsupported (bulk) CrO3, (b) 5% chromia/ alumina calcined at 500 ~ and (c) 5% chromia/alumina spent sample in propane dehydrogenation.

4. CONCLUSION The present study demonstrates the usefulness of combined characterization techniques in the study of heterogeneous catalysts. Chromium is found well distributed throughout the support in both uncalcined, calcined and spent catalysts. Due to the coke formation in propane dehydrogenation reaction, the surface area of chromia decreases resulting in lower Cr/A1 atomic ratios. The number of active sites are believed to be reduced due to coke deposition. The XPS results indicate that chromia is entirely in Cr 6+ oxidation state in the case with both uncalcined and calcined samples. It further indicates that a peak due to Cr 3+ oxidation level appears after the catalyst is used in propane dehydrgenation. TG results agree well with the fact that the main phase transformation of Cr 6+ compounds occurs at about 500 ~ resulting in reduction to Cr 3+. Both TG and TPR results demonstrate the relative ease with which bulk CrO 3 can be reduced compared to the supported chromia catalysts due to metal support interaction. 5. ACKNOWLEDGEMENT The authors wish to acknowledge the support of the Research Institute of the King Fahd University of Petroleum and Minerals. The microbeam-PIXE part of this work was carried out at Energy Research Laboratory of the Research Institute.

426 6. REFERENCES

1. S.D. Rossi, G. Ferraris, S. Fremiotti, E. Garrone, G. Ghiotti, M.C. Campa and V. Indovina, J. Catal. 148 (1994) 36. 2. O.F. Gorriz, V.C. Corberan, and J.L.G. Fierro, Ind. Eng. Chem. Res. 31 (1992) 2670. 3. S.D. Rossi, G. Ferraris, S. Fremiotti, V. Indovina and A. Cimino, Appl. Catal., 106 (1993) 125. 4. J.A. Cairns and J.A. Cookson, Nucl. Instr. and Meth. 168 (1980) 511. 5. J.A. Cookson, "Applications of High Energy Ion Microbeams", (G.W.Grime and F.Watt, Eds.). p. 294. Adam Hilger Ltd., Bristol, UK, 1987. 6. A. Rahman, M.H. Mohamed, M. Ahmed and A.M. Aitani, Appl. Catal., 121, No. 2 (1995) 203. 7. M. Ahmed, J. Nickel, A.B. Hallak, R.E. Abdel-aal, A. Coban, H.A. A1-Juwair and M.A. Aldaous, Nucl. Instr. and Meth. B82 (1993) 584. 8. M. Ahmed, A. Rahman, J. Nickel and M.A. Garwan, "Micro-PIXE measurement of Ni distribution over supported nickel oxide catalysts", Thirteenth International Conference on the Application of Accelerators in Research & Industry, Denton, Texas, Nov. 7-10, 1994.

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All fights reserved.

D E A C T I V A T I O N M E C H A N I S M S OF A C H R O M I A - A L U M I N A COKE DEPOSITION

427

C A T A L Y S T BY

F. Mandani a, E. K. T. Kam b and R. Hughes c

aDepartment of Chemical Engineering, College of Technological Studies', P.O.Box 105, 44000 Sabah Salem, KUWAIT. bpetroleum Technology Department, Kuwait Institute for ,Scientific Research, P.O.Box 24885, 13109 Safat, KUWAIT. CDepartment of Chemical Engineering, University of Salford, M5 4WT, ENGLAND. ABSTRACT Catalyst deactivation is a normal phenomenon in catalytic processes and it comes in many forms - coking, poisoning, aging or sintering. In the case of coking, highly unsaturated, heavy hydrocarbons are adsorbed onto the catalyst active surface and condense leading to coke deposition. In this study, the mechanisms of coke formation as a side reaction from the dehydrogenation of 1-butene were investigated. The physical modifications in pore volume and surface area show that pore-blocking cause the rapid initial loss in catalyst activity while a slower active site coverage results gradual deactivation there after. However, the characteristics of the coke deposition cannot be described satisfactorily by either parallel or series fouling alone and the combined fouling mechanism is more appropriate. Moreover, the contribution to coke deposition from each of the individual mechanism changes with temperature. A simple model is developed to simulate such coking phenomenon and the results are compared well with the experiments.

1. I N T R O D U C T I O N Catalysts are normally deactivated once they are put on stream. Since this is an important industrial problem, numerous research works have been undertaken to investigate this phenomenon [1-4]. Coke formation is believed to be caused by the highly unsaturated species of high molecular weight hydrocarbons which can be reactants, products or both [5-6] which are adsorbed onto the catalyst active surface; further condensation reactions from the adsorbed materials lead to the deposition of coke. The coking precursors originate from reactions taking place over the catalyst and are not impurities so coking, unlike poisoning, cannot be minimized by purifying the feedstock or using a guard-bed. The consequences of catalyst coking are a reduction in product yields, lowering of product quality, modification of product slate; it even leads to process shut down. For example, a 1% increase in the time-onstream from a hydrocracking unit processing 30,000 bpd before performing the customary catalyst regeneration cycle which usually takes place after 18 months on stream, will give an

428 extra w e e k o f process run-time. There is a very high incentive to minimize coke formation in any o f the catalytic conversion processes. The dehydrogenation of 1-butene over a chromia-alumina catalyst is selected as a model reaction system to study the fouling mechanisms and their respective fouling precursors. The reaction and deactivation schemes can be taken as:

Hydrocarbon

Reactants

............ (Combined ral Hydrogen + Hydrocarbon

(Parallel C o k i n g ) ' .........

lel & S e r i e s Coking).. ........::.-..i ......

P r o d u c t s ................................................................................................... :::.::!~:' [Series C o k i n g )

where the hydrocarbon products include trans-2-butene, cis-2-butene and 1,3-butadiene. The reaction has been examined over a range of temperatures, several catalyst sizes and at different concentration of reactants.

Table 1 Catalyst properties (Harshaw) Particular

Data or Information

Catalyst Support Catalyst Shape Particle Diameter Particle Length Bulk Density Crushing Strength Surface Area Pore Volume Cr20 3 Content Crushed Particle Size Ranges : 7 - 10 mesh 1 8 - 22 mesh 25 - 30 mesh 40 - 60 mesh 70 - 85 mesh

Alumina Cylindrical 4 x 10"3 m 4 x 10-3 m 1.15 x 10 -3 kg / m 3 9.5 kg 2.95 x 104 m 2 / k g 3.35 x 10 -4 m 3 / k g 19% 1.68 - 2.83 x 10 -3 m 0.77- 0.92 x 10 -3 m 0.55 - 0.68 x 10 -3 m 0.25 - 0.37 x 10 -3 m 0.17 - 0.19 x 10 -3 m

429 2. EXPERIMENTS The experimental investigations discussed here are focused in two areas - coking and regeneration. Although the dehydrogenation experiments have been carried out in conjunction with coking, these will not be reported here but can be found elsewhere [7,8]. In either case, the catalyst used was 19% chromia-alumina catalyst which was in the form of small sized, cylindrical particles in 4 mm diameter and length. These particles were then crushed and sieved into five different sizes for experimentation. Their properties are given in Table 1. The coking investigation was undertaken in a stainless steel reactor, 2.2 x 10-1 m in length with 8 x 10-3 m inside diameter, which was mounted vertically. The reaction temperature was maintained by an electric furnace surrounding the reactor tube. The catalyst bed was 3 x 10-3 m long and situated 5 x 10-2 m from the reactor outlet. The thermocouple used was made of chromel-alumel wire of 5 x 10-4 m diameter which was placed in the center and 1.5 x 10-3 m from the bottom of the catalyst bed. The reactor can be operated in either differential or integral mode. To coke the catalyst, 1- butene was introduced from the top of the reactor together with nitrogen. The partial pressure of 1-butene ranged between 5-25 kPa. Coking was conducted isothermally at set temperatures from 798-873 K. The runs were terminated at different times (i.e. 300, 1200, 2400, 3600, 4800s). Catalyst characterization was also made using mercury porosimeter and sorptometer measurements to determine the pore size distribution and surface areas of fresh and coked catalysts. This was then used to assist the determination of coke deposition mechanisms.

2.1 Catalyst Regeneration The regeneration was carried out using a microbalance rig based on the thermogravimatric technique as shown in Figure 1. The air was supplied by mixing oxygen and nitrogen gases to the desired compositions before admitting to the reactor chamber which temperature was maintained by an electric furnace. The regeneration kinetics were determined based on the reduction in weight from the burn-off of the coke deposition which was recorded with process time at fixed intervals.

2.2 Experimental Results The study of coke formation kinetics as a side reaction during the dehydrogenation of 1-butene was carried out in two stages: coking and then regeneration. This is a complex problem because of the wide variety of reactions which are possible to form coke. The major sources of coke precursors are the reactant (1-butene), the primary product from hydrogenation (1,3-butadiene), and/or two other product species from isomerisation (transbutene and cis-butene). To overcome this, a series of experiments have been designed to eliminate and rank the precursors systematically. According to the findings, the isomerisation products comparing to that of dehydrogenation were negligible for the full ranges of temperatures and partial pressures employed in this study. Hence, it is reasonable to lump all products as one isomer species [9-11 ]. Moreover, the particle size has a considerable effect on coking and the results show a large increase in coke content found in the smaller size particles.

430 In the following analysis, the data used are obtained from the experimental conditions which are given in Table 2. 2.2.1 Coke formation on catalyst pore size distribution and surface area

The losses in the total pore volume and surface area due to coke deposition are shown in Table 3. Both the losses increase as the coke content increases. However, when the effect is taken in terms of unit coke wt% deposited, the loss in the active surface area at low coke content is comparatively more than that at higher coke content, while there is a maximum loss in pore volume observed at a certain coke deposition. Figure 2 depicts the change in pore size distribution of fresh and spent catalysts. The coke content in the spent catalysts range from 3 to 8.8 wt%. The fresh catalyst represented by the 0% coke curve has a considerably wide range of macropores, 100 - 1000 nm. The loss in pore volume due to coking in this range of micropores is minimal. In contrast, there is a significant drop in pore volume in the mesopore region ranging between 4.5 - 15 nm. Hence, from this analysis, it is shown that coke has a significant pore blocking effect on the mesopores, compared to the larger macropores which are affected only slightly. This pore-blocking of the smaller mesopores is more pronounced than the loss of the active sites and the severity diminishes when the coke level reaches a 7 wt% level. Any further increase in the coke level results in a continuous reduction in both total pore volume and active surface area.

Table 2 Coking experiment conditions Particular Catalyst weight Flow rate Temperature

Value 8 x l 0-4 kg 3.3 x 10-4 m3/s 873 K

Particular

Value

1-butene, PB Process time

2 kPa 7200 s

Table 3 Losses in pore volume and surface area due to coke deposition Coke Content

Loss

in

Total

Pore

Volume

Loss

in

Active

Surface

Area

[ wt% ]

[%]

per unit Coke Content

[%]

per unit Coke Content

0 (fresh) 3.0 5.0 6.7 8.8

0.00 13.00 24.00 44.00 50.00

Not applicable 4.33 4.80 6.57 5.68

0.00 22.03 35.59 47.15 61.02

Not applicable 7.34 7.12 7.04 6.93

431

Figure 1. Schematic diagram of the regeneration experimental rig.

Figure 2. Effect of coke deposition on pore size distribution.

432

2.2.2 Deactivation by coking The decrease in the dehydrogenation is caused by the coking of the catalyst. The coking precursors can be the reactants and/or the products. It is advantageous to examine the initial reaction and coke deposition rates. Figure 3 shows the effect of process temperature (xaxis) and reactant concentration (y-axis) on the initial dehydrogenation rate of 1-butene (zaxis) over the chromia-alumina catalyst. The initial reaction rate increases with the temperature. However, a maximum rate is observed at a particular partial pressure of the 1butene at one temperature and this observation applies to the entire temperature range. This is a classic example of the surface reaction controlling kinetics [3]. If the products are the coking precursor, it is logical to expect similar characteristics to be exhibited in the initial coke content curves, or the inverse if the reactant were the precursor. The effect of temperature and 1butene partial pressure on the initial coke content is given in Figure 4. The coke decreases with temperature as well as the 1-butene partial pressure monotonically. This indicates the fouling precursor is not a single species, but a combination of all the hydrocarbons present in the effluent stream. 3. PARAMETERS DETERMINATION AND MODEL FITTING The determination of more comprehensive coking mechanisms and rate equations requires simultaneous treatment of all experimental data to enable all the relevant parameters related to coking to be considered. After analysing the experimental data, numerical values of the rate and adsorption equilibrium constants were determined by statistical tests, and models were rejected if a negative constant was estimated at more than one temperature. It was found that the hyperbolic type of decay, as described in Equation (1), gives the best fit from the 9 models tested because it gives the least error from the sum of squares analysis [8], ~)c(t, T) =

1 [1 + ~c(T) Cc(t)]

(1)

where C c is the catalyst coke content, ~c is the deactivation function relating to coke content, (xc is the deactivation coefficient, t is the process time in s, and T is the process temperature in K. In a previous work [8], neither parallel or series coking mechanism was found to be satisfactory because the predominant fouling mechanism changes with temperature. Since the dehydrogenation reaction and coking formation takes place on the same type of active sites, a combined parallel and series mechanism is assumed in which ~c can be expressed in terms of process temperature, process time and the concentration of the respective foulents, as ~c(t,T) = kcB 7cB(T) PB(t) + kcD 7cD(T) PD(t) (2) where kcB and kcD are the rate constants for coking reactions due to feed and product respectively, ?cB and ?cD are the thermal factors for coking kinetics, and PB and PD are the concentrations of feed and product respectively.

433

Figure 3. Effect of PB and T on initial reaction rate.

Figure 4. Effect of P B and T on initial coke content.

434 3.1 Parameters Estimation and Results

To undertake the parameters estimation of the rate constants, deactivation coefficient and coking thermal factors, a combination linear and non-linear multiple parameters regression techniques were applied. The form of deactivation coefficient can be expressed as: (tc(T) =

4.12 x 104 T

- 34.8

(3)

and the coking thermal factors are, 7cB (T) = exp( )-01.7 T

(4a)

and '/cD (T) :

-65 exp(-f)

(4b)

for parallel and series coking respectively. The changes in feed and product with process time can be expressed as: PB(t) : 0.0711 exp(-x)

(5a)

and PD(t) = 0.000137 exp(-x)

(5b)

The coke content at different process temperature and time can be determined by putting Equations (3) to (5b) into (2) to solve for ~c which is then substituted in Equation (1). The simulated results are compared well with the experimental data at PB = 10 kPa as shown in Figure 5. Similar comparisons are also found for the other PB values.

4 CONCLUSIONS The mechanisms of coke formation as a side reaction from the dehydrogenation of 1butene over a 19% chromia-alumina catalyst were investigated over a range of 1-butene partial pressure and process temperature. The physical modifications in pore volume and surface area in the catalyst show that the pore-blocking occurs first which causes the rapid initial loss in catalyst activity. Subsequently, a slower active site coverage prevails and results in gradual deactivation afterward. The characteristics of the coke deposition is better described as the combined parallel and series fouling mechanism since the contribution in coke deposition from each of the individual mechanism changes with process temperature. This is very important when the optimal temperature policy is employed to compensate the loss in product yields due to catalyst deactivation by raising the process temperature. A simple model was also developed to simulate the coking phenomenon and the results compare well with experimental data. The model can be easily coupled into reactor design algorithms to improve the design of catalytic reactors which undergo similar catalyst deactivation.

435

Coke C o n t e n

[%]

12,000 ~

/

~ 8 4 0

8,000 Process T i m e [s]- ~ Experimental Data Simulated Values

9 798K

9 823K

800

T e m p e r a t u r e [K]

8-~8 K

A 873K

Figure 5. Comparison of experimental and simulated values of coke content at PB = 10kPa

NOTATIONS Cc

kcB, kcD L PB, PD

R

Rp t T Vp

catalyst coke content, wt% reaction rate constants for parallel and series coking respectively, s-1 active site concentration concentration of 1-butene and 1,3-butadiene respectively, kPa ideal gas law constant, kcal/kg-mol/K pore radius, nm process time, s process temperature, K total pore volume, m3/kg

Greek Letters c~c ~'cB and 7cD ~c "c

deactivation coefficient as defined in Equation (3) thermal factor for coking kinetics deactivation function as defined in Equations (1) and (2) dimensionless process time, ratio of process time to maximum process time

436 REFERENCES

.

3. .

5. .

7.

.

10. 11.

J. B. Butt, The Progress in Catalyst Deactivation, Proceedings of the NATO Advanced Study Institute on Catalyst Deactivation, Portugal, 1992. R. Hughes, Deactivation of Catalyst; Academic Press: New York, 1984. G. F. Froment and K. B. Bischoff, Chemical Reactor Analysis and Design, 2nd Ed., John Wiley: New York, 1990. A. S. Krishna, Catal. Rev.- Sci. Eng., 32 (1991) 279. A. G. Gayubo, J. M. Arandes, A. T. Aguayo, M. Olazar and JBilbao, Ind. Eng. Chem. Res., 32 (1993) 588. F. Garcia-Ochoa and A. Santos, Ind. Eng. Chem. Res., 32 (1993) 2626. F. Mandani and R. Hughes, Studies in surface science and catalyst - Catalyst deactivation, B. Delmon and G.F. Froment (eds.), 88 (1994) 507. F. Mandani, Kinetic and Deactivation studies during catalytic dehydrogenation, PhD Thesis, University of Salford, England, 1991. Y. Amenomiya and R. J. Cvetanovi'e, Canad. J. Chem., 40 (1962) 2130. S. Carra and L. Forni, Ind. Eng. Chem. Proc. Des. Dev., 4 (1965) 281. H. A. McVeigh, PhD Thesis, University of Deleware, USA, 1972.

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All fights reserved.

437

INVESTIGATION OF SYNTHESIS GAS PRODUCTION F R O M METHANE BY PARTIAL OXIDATION OVER SELECTED STEAM R E F O R M I N G C O M M E R C I A L CATALYSTS

H. AI-Qahtani Chemical Engineering Department, University of Bahrain, Isa town, P. O. Box 32038, State of Bahrain. I. ABSTRACT The production of synthesis gas (CO, H2) from methane by partial oxidation is investigated over commercial steam reforming catalyst at several flow rates, temperatures, and at different methane/oxygen ratios (R). Optimum synthesis gas selectivity and yield achieved are 70% and 60%, respectively at methane/oxygen ratio close to 2 and at flow rates of 500 cm3/min. An initial temperature (665 ~ is necessary to initiate the reaction and then the reaction is stabilized at 883 ~ The effect of methane/oxygen ratios and residence time are effective in determining the synthesis gas selectivity and yield. 2. INTRODUCTION Steam reforming is the principle process for carbon monoxide and hydrogen production. Steam reforming process is applied for several industrial applications to provide the necessary amount of the synthesis gas. Those industries such as oil refineries, iron and steel manufacturing, methanol and ammonia synthesis, and other several petrochemical industries. The future demand for synthesis gas utilization will increase especially when methanol is used as a combustible fuel in large scale and when compact fuel-cells is used in wider applications. One of the major alternatives methods for the production synthesis gas is the partial oxidation of fuel oil and coal gasification. However, capital costs for the partial oxidation of fuel oil and coal gasification are approximately 1.5 and 2 times higher, respectively, than that for steam reforming of natural gas [ 1]. Studies investigating the direct conversion of methane into methanol, formaldehyde, ethane, and ethylene found that these compounds could not be produced commercially due to the limitation on yield and selectivity of the desired products [2]. It is economically more viable to convert methane into synthesis gases and then to the final product [3]. A large amount of research on methane oxidative coupling has been conducted in recent years. The main setback of direct coupling is the high selectivity and yield of unfavoured products (CO2, and H20), and hence, the limited of C 2 yield [4]. Recently, active studies have been conducted investigating the possibility of oxidizing methane to synthesis gas catalytically at lower temperatures. Studies of methane to CO and H 2 over Ni/AI203 were reported. The formation of CO and H 2 rather than CO 2 and H20 were achieved at high synthesis gas selectivity (90%) and yield (95%) [5].

438 Chouddhury and co-worker[6] oxidized methane at high temperatures ranging from 300900~ over Ni/CaO. High methane conversion (90%) and high synthesis gas selectivity (92%) were found when the reaction took place over reduced Ni catalyst [6]. Schmidt et al. [7], studied the catalytic partial oxidation of CH 4 in air and pure 0 2 at atmospheric pressure over Pt and Rh coated monoliths. High selectivity for H 2 and CO (90's%) were achieved at 950~ over Rh catalyst when pure 0 2 was used; with air, the selectivity's were 70% and 40% over Rh and Pt, respectively. The production of synthesis gas from methane oxidation was also studied over Fe catalyst in fuel cell using solid electrolyte (YSZ) at 850-950~ at atmospheric pressure [8]. The anodic electrode was Fe and the cathode that was exposed to air was Pt. Reduced iron was more active than oxidized iron for synthesis gas formation. The maximum CO selectivity and yield were nearly 100% and 73%, respectively. Carbon deposition was reported at high methane to oxygen ration. The scope of the present study is the investigation of partial oxidation of methane over commercial steam reforming catalyst. Thus, the main purpose of using this type of catalyst is not to compare between the synthesis gas selectivity and yield of steam reforming to partial oxidation reactions over this type of catalyst, but to investigate the performance of partial oxidation reaction over commercial steam reforming catalyst. Satisfactory performance over the given catalyst is expected to provide information needed to develop commercial catalysts for partial oxidation. The reason for choosing this type of catalyst is due to the similarity between steam reforming and partial oxidation with respect to their operating conditions and type of species involved and produced during the reactions. 3. EXPERIMENTAL The system consisting of a tubular reactor, furnace, gas cylinders, flow meters, temperature controller, gas chromatography, and bubble meter is shown in figure 1. All flow rates measurements are monitored by the bubble meter. The reactor is a stainless steel tube with ID. = 2.0 cm and L. = 9.0 cm where 5 g of the catalyst is loaded in the tube (Figure 2). The catalyst used for this study is a commercial steam reforming type brought from the Gulf Petrochemical Industries (GPIC), the only petrochemical plant in the state of Bahrain. The catalyst consists 20% Ni and the rest is magnesium oxide mixed with a ceramic material. All the gases are premixed at room temperature, 25 ~ before entering the reactor. 4. RESULTS AND DISCUSSION Three sets of experiments have been conducted. The first set is examining the influence of methane/oxygen ratios on the performance of the catalyst; the second set is studying the effect of temperature on the synthesis gas formation; and the third set is investigating the influence of residence time on synthesis gas selectivity and yield. The experimental data are shown in tables 1 and 2. Selectivity, yield and conversion are defined according to the following: Selectivity o f H 2 = [rate of H2/2 (rate of CH 4 in - rate of CH4out)] Selectivity of CO = [rate of CO/(rate of CH 4 in - rate of CH 4 out)] Yield o f H 2 = [rate of H2/2 (rate of CH 4 in)]

(1) (2) (3)

439

2

2]

LIJ 1

(s) Figure 1. Schematic diagram of the tubular reactor system. (1: gas cylinder; 2: rotometer; 3:

reactor; 4: furnace; 5: temperature controller; 6: gas chromatograph; 7: bubble meter) in 3

out

Figure 2. Schematic diagram of the reactor. (1:furnace; 2: reactor; 3: thermocouple) Yield of CO Conversion (%X)

= =

[rate of CO / rate of CH4 in] [(rate of CH 4 in - rate of CH 4 out)/rate of CH 4 in]

(4) (5)

In the first set of experiments, the inlet flow rate is fixed at 500 cm3/min, and temperature at 883~ It is observed that the outlet flow rate is usually higher than the inlet by 100 to 150 cm3/min. As shown in table 1 and figures 3, 4, and 5, the rates o f H 2 and CO increased with the increase in the methane/oxygen ratios (R). It may be seen from the given figures that the hydrogen rate reached to a maximum at methane to oxygen ratio around 2. Therefore, most of the methane enters are converted to hydrogen and CO at that given R. At low methane to oxygen ratios (R < 2), the hydrogen yield

440 Table 1. Influence of methane/oxygen ratio on catalyst performance.

Methane/Oxygen Ratio ( 10 -3 mol/min.)

R=0.715

R=I.15

R=2.06

R=3.61

n (O2)in

11.930

9.510

6.668

4.440

n (CH4)in

8.530

10.950

13.780

16.020

n (CO)out

1.360

5.733

8.100

8.880

n (H2)out

2.730

11.739

16.380

18.325

n (CH4)out

0.191

0.730

2.730

4.095

%SH2

16.37

57.43

74.11

76.83

%Sco

16.31

56.09

74.12

74.46

%YH2

16.00

53.60

59.43

57.19

%Yco

15.94

52.35

59.43

55.43

%XCH 4

97.76

93.33

80.19

74.44

I

20

I

I

I

I

I f

15-

A

/

E "6 'o E

H2

/

/

o==

v

[]

/

10-

,-

[]

CO

5 -

0

I

0.5

1.5

I

I

I

i

2

2.5

3

3.5

ratio (CH4102)

Figure 3. Variation of H 2 and CO rates with methane to oxygen ratios at 500 cm3/min, and 883~

441 I

80

I

I

I

I

I o

CO

70-

D

60-

H2

> ..,= 0 0

50-

m

0

4030-

m

2010 0.5

I

I

I

I

I

I

1

1.5

2

2.5

3

3.5

ratio

(CH4102)

Figure 4. Variation of CO and H 2 selectivities at several methane/oxygen ratios at 500 cm3/min, and 883~

I

80

I

I

I

I

I

70-

m

H2 =,,.t

60-

m

co

50403020-

m

10 0.5

I

I

I

I

I

I

1

1.5

2

2.5

3

3.5

ratio

(CH4102)

Figure 5. Variation ofH 2 and CO yields at several methane/oxygen ratios at 500 cm3/min, and 883~

442 Table 2. Influence of inlet flow rate on catalyst performance.

Flow Rate . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .

Q=250 (cm3/min.)

Q=500cm3/min.)

Q=750 cm3/min.

n (O2)in

3.580

7.160

10.740

n (CH4)in

6.649

13.30

19.950

n (CO)out

5.650

7.490

5.670

n (H2)out

8.664

16.130

17.980

n (CH4)out

0.001

2.087

6.957

%SH2

65.15

71.93

69.19

%Sco

84.97

66.80

43.64

%YH2

65.15

60.64

45.54

%Yco

84.97

56.32

28.42

%XCH A

99.97

84.31

65.13

(10-3 mol/min.)

reduced due to the reaction of the excess oxygen available in the system with the hydrogen and, therefore, more carbon dioxide and water are observed at lower R values. At high methane to oxygen ratios, carbon deposition and C2+ are detected. This indicates that the limitation of oxygen species caused the free carbon formation. In the second set of experiments, temperatures are varied ( 400, 500, 600, 700, 800~ at constant inlet flow rate 500 cm3/min, and at a value of R about 1.86. All the given temperatures are reported from a thermocouple attached to the catalyst inside the reactor. At low temperatures (400 to 600~ formation of synthesis gas is insignificant. However, at about 665~ pulses of explosion occurs initially and then temperature increases rapidly above 800~ and the amounts of CO, H 2 increase significantly. At 700 and 800~ no pulses of explosion are observed but the temperature increases till it is stabilized at 883~ Therefore, heating of the reaction is needed only to initiate the reaction and then reaction is sustained by the exothermic heat of reaction. The explosion behavior that occurs at temperature of about 665~ is due to the sensitivity of the reaction to the variation of the temperatures. At temperature of 665~ the interaction between 0 2 and CH 4 over the catalyst surface is more likely to follow an explosion mechanism due to the types of intermediates that are dominated at this condition. In the third set of experiments, inlet flow rates are varied and temperature is held constant at temperature 883~ and at methane to oxygen ration 1.86. As shown in table 2 and figures 6, 7, 8, and 9, CO and H 2 rates increase then decreased slightly. Also selectivity and yield decrease at high and low flow rates. Methane conversion also decreased with the increase in the flow rate. At low flow rate ( < 400 cm3/min.), carbon deposition is detected. At high flow rate, lower CO and H 2 yields are recorded. Therefore, flow rate is an important parameter controlling the selectivity of synthesis gas.

443 100

I

I

I

I

I

I

I

I

I

I

I

I

1

1.5

2.5

3

3.5

9080x 7060504030 0.5

2 ratio

(CH4102)

Figure 6. Variation of methane conversion at several ratios of methane/ oxygen at 500 cm3/min, and 883 ~

I

18

I

I

I

I f

J

16-

H2

J

m

J 14-

m

J

A r o . .

E

12-

m

J

o

E 10-

m

8

-

6

-

200

m

CO m

o-"

I

I

I

I

I

300

400

500

600

700

Q (cm31mi

800

n)

Figure 7. Variation of CO and H 2 rates at several inlet flow rates at ratio = 1.9 and 883 ~

444 I

90

I

I

I

I

8070-

>, > .e..* O O ~)

u)

f

B--"'-

""e...

-'-~

H2

6050CO 4O 30m

2010

i 200

300

i 400

I 500

I 600

I

800

700

O(cm31min)

Figure 8. Variation of CO and H 2 selectivities at several inlet flow rates at ratio = 1.9 and 883 ~ I

90

I

I

I

I

m

80-

-D

706050-

H2-

4030-

CO

20

i 200

300

I 400

I 500

I 600

I 700

800

O (cm31mln)

Figure 9. Variation of CO and H 2 yields at several inlet flow rates at ratio = 1.9 and 883 ~

445 In the fourth set of experiments, different ratios and flow rates are examined in the absence of catalyst (homogenous). The rates of hydrogen and carbon monoxide are very low where their selectivity and yield are not exceeding 3% to 5%. This set of experiments indicates that the role of catalyst is significant to improve the synthesis gas production. It is believed that methane and oxygen are adsorbed dissociatively and then interact on the surface during the steam reforming and partial oxidation reactions over Ni, Ir, Pd, Re, and Pt [9-14]. The mechanism is summarized according to the following scheme : CH4(g ) + S O2(g ) + S H20(g ) + S

--> --> -->

C(ads) + 4H(ads) 20(ads) O(ads) + H2(gas)

The formation of CO, H2, carbon, H20, CO 2 may be expressed according to the above mechanism. Thus, at high ratios of R, adsorbed oxygen will be the limiting reactant and thus carbon deposition is achieved according to the following reaction: nC(ads)

+

mO(ads)

-->

mCO(ads)

+ (n-m) C(ads)

At low ratios of R, adsorbed oxygen sites are high and carbon sites on the surface are relatively low with the result that, adsorbed oxygen species may interact with adsorbed hydrogen to form water and with one carbon species adsorbed on the surface to form carbon dioxide. yC(ads) 2H(ads)

+ +

zO(ads) O(ads)

---> xCO(ads) --> H20(ads)

+

vCO2(ads )

Maximum synthesis gas selectivity and yield are about 70% and 60%, respectively, although those values are considered much lower than those achieved over Ni, Ir, Re, and others. 4. CONCLUSION Hydrogen and carbon monoxide production from partial oxidation of methane over commercial steam reforming catalyst is influenced by the methane to oxygen ratios and by the gas mixture flow rates. Both the selectivity and yield of synthesis gas are maximized at R about 2 and decrease at higher and lower ratios of methane to oxygen. H20 and CO 2 are formed at low ratios and carbon deposition is detected at high ratios. No heat is required to assist the reaction, however, initial heating is necessary to bring the reaction above the explosion temperature. Optimum selectivity and yield to synthesis gas are achieved at mixture flow rate of around 500 cm3/min, and methane to oxygen ratio of about 2.0. REFERENCE

1. T. Czuppon and J. Buridas, Hydrocarbon Process, 58 (1979) 197. 2. D. Eng and M. Stoukides, Catal. Rev.-Sci. Eng., 33 (1991) 375. 3. J. Lee and S. Oyama, Catal. Rev.-Sci., 30 (1988) 249. 4. A. Amenomiya and G. Sanger, Catal. Rev.-Sci. Eng., 32(3) (1990) 163.

446 5. D. Dissanyake, M. Rosynek, K. Kharas and J. Lunsford, J. Catal., 132 (1991) 117. 6. V. Chouddury, A. Rajput and B. Prabhakr, Catalysis Letters, 15 (1992) 363. 7. D. Hickman and L. Schmidt, J. Catal., 136 (1992) 300. 8. H. Alqhtani, D. Eng, and M. Stoukides, J. Electrochem. Soc., Vol. 140, 1993. 9. P. Munster, H. Grabe and Ber Bunseges, Phy. Chem., 84 (1980) 1068. 10. C. Cullis, T. Newell and D. Trimm, J. Chem. Soc. Faraday Trans., 68 (1972) 1406. 11.A. Frannet and G. Lienard, J. Chim. Phys. Physicochim. Biol., 68 (1971) 1526. 12. C. Coekelbergs, J. Delannois, A. Frannet and G. Lienard, J. Chim. Phys. Physicochim. Biol., (1964) 1167. 13.N. Meshenko, V. Veselov, F. Shub and M. Temldn, Kinet. Katal., 18 (1977) 962.

Catalysts in PetroleumRefining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All fights reserved.

AROMATIZATION CATALYSTS

OF

BUTANE

OVER

447

MODIFIED

MFI-TYPE

ZEOLITE

Tatsuaki Yashima a, Shigeyuki Ejiri b , Koichi Kato a, Mohmand M. Ishaq a'*, Makiko Tanigawa b Takayuki Komatsu a and Seitaro Namba b

Department of Chemistry, TokyoInstitute of Technology, 2-12-1, Ookayama, Meguro-ku, Tokyo 152 Japan b Department of Materials, The Nishi-Tokyo University, Uenohara-cho, Kitatsurugun, Yamanashi 409-01 Japan ABSTRACT The aromatization of butane on zinc modified HZSM-5 and gallium-and/or copper-modified HZSM-5 was studied. The activity, selectivity and thermal stability of the Zinc-modified catalysts prepared by various methods were discussed. The zinc loaded on HZSM-5 by the impregnation showed the highest activity. However, at a reaction temperature higher than 873K this catalyst lost part of the zinc from the zeolite surface. On the other hand, the zinc loaded into the zeolite framework showed relatively low activity and selectivity to aromatics donation, but it showed relatively high thermal stability. The gallium loaded on copper partially ion-exchanged HZSM-5 by impregnation showed high selectivity for aromatics donation. It is concluded that in this catalyst, gallium promotes dehydrogenation including the initial conversion of butane and the reaction step from C6-C9 olefin to aromatics. The copper ion mainly controls the acidity of the HZSM-5 to depress the cracking of the butane and C6-C9 olefins. 1. INTRODUCTION The associated gas is mainly composed of C1-C4 paraffins. Recently, the methane and ethane in the associated gas have been used effectively by petrochemical industries as a raw material However, propane and butane included about 10 vol % [ 1] in the associated gas are used only for fuels. Therefore, it is expected that propane and butane will be converted to liquid hydrocarbons, such as aromatics, for the effective total utilization of the associated gas. The reformation of lower paraffins to aromatics has been studied for about 20 yr by using zeolite catalysts. Recently, an excellent review was published of lower alkane transformation to aromatics on ZSM-5 zeolites [2]. From the studies of the mechanism of this reaction, it has been suggested that the bifunctional catalysts, having solid acidity and dehydrogenation activity, can effectively promote the aromatization of lower paraffins[3-6]. It has been reported that ZSM-5 and ZSM-11 are excellent solid acid catalysts [7] and the transition metals [8], Ga [9] and Zn [9] show high dehydrogenation activity in this reaction. In the case ofbifunctional * Present address: Department of Chemistry, University ofPeshawar, Peshawar, Pakistan

448 catalysts, a suitable balance of activity between the solid acid site and the dehydrogenation site is very important to accelerate the reaction effectively. In this study on the aromatization of butane, we want to control the acidity of solid acid sites on HZSM-5 and to improve the thermal stability of Zn supported on HZSM-5. We will discuss on the activity and selectivity of Ga and Cu supported on HZSM-5 and also on the effect of Zn supporting method on the stability. 2. EXPERIMENTAL DESIGN 2.1 C a t a l y s t ZSM-5, Ga-Silicate, and Zn-Silicate were synthesized hydrothermally. The protontypes of these zeolites were prepared by ammonium ion exchange followed by the deammoniation at 773K in He stream. The Zn loaded HZSM-5 catalysts were prepared by the impregnation and atom-planting methods using Zn nitrate solution and Zn chloride vapor, respectively. The copper-loaded HZSM-5 catalysts were prepared by ione• using copper acetate solution, and Ga loaded HZSM-5 catalysts were prepared by impregnation using Ga nitrate solution. The divalent cation loaded Ga-silicates were prepared by ion-exchange using corresponding metal acetate solution. Ga and Cu loaded HZSM-5 were prepared by the ionexchange of Cu cation using Cu acetate solution first, followed by the impregnation of Ga using Ga nitrate solution. All catalysts used in this work are as follows: - Zn-loaded ZSM-5 catalysts: Zn loaded on HZSM-5 prepared by impregnation, Si/Al=35, Si/Zn=24: Zn(Imp) Zn loaded on HZSM-5 prepared by atom-planting, Si / A1=38, Si/Zn=34: Zn(A-P) - HZn-silicate prepared by hydrothermal synthesis, Si/Zn=63 9Zn-Sil - Ga-loaded ZSM-5 catalysts: Ga loaded on HZSM-5 prepared by impregnation, Si/A1=22, Si/Ga = 110: Ga(Imp) Ga loaded on Cu partially ion-exchanged HZSM-5 prepared by impregnation, Si/Al=22, Si/AI=I 10: Ga(Imp)Cu(Ex) - HGa-silicate prepared by hydrothermal synthesis, Si/Ga=26, 34: Ga-Sil(26), Ga-Sil (34)

- Cu loaded on Ga-Sil(34) prepared by ion-exchange: Cu(Ex)Ga-Sil - Alkaline earth metal cation loaded on Ga-Sil(26) prepared by ion-exchange: Me(Ex)GaSil

- Cu loaded on HZSM-5 prepared by ionexchange: Cu(Ex) 2.2. A p p a r a t u s and P r o c e d u r e

The conversion of butane was carded out in a fixed-bed type reactor with a continuous flow system at atmospheric pressure. The reaction mixture was analyzed by gas chromatography.

449 3. RESULTS AND DISCUSSION 3.1. Zn loaded

ZSM-5

catalysts

The conversion of butane on Zn-Sil, Zn(Imp) and Zn(A-P) was studied. As shown in Figure 1, on all catalysts, the conversion of butane increased with W/F at 823K, while the selectivity to aromatics increased only slightly with W/F. The catalytic activity and the selectivity of Zn(Imp) were the highest in these catalysts. These results, suggest that the dehydrogenation activity of Zn loaded on ZSM-5 surface is higher than that of Zn loaded in the zeolite framework of ZSM-5.

ioo

100

0 80

-

Zn (Imp)

80 O t~

QJ

60

-

60~

Zn(Imp)

or)

O

o

9~,.=I ffl

E

40

-

40

o t~

Zn (A-P) Zn-Sll

O

O

r,.) 2o

-

20 "~ rd

rd3

r. 0

i 5

i

I

i

10

15

20

0

W / F / g-h.mol-1 Figure 1. Effect of W/F on Zn loaded catalysts at 823K.

Figure 2 shows the effect of reaction temperature on three kinds of catalysts. At a lower reaction temperature, Zn(Imp) showed much higher catalytic activity and selectivity to aromatics than two other kinds of Zn loaded catalysts. At a higher reaction temperature, the conversions of butane over Zn-Sil and Zn(A-P) increased dramatically. On the other hand, the selectivity to aromatics of all Zn loaded catalysts increased gradually with reaction temperature, and reached their maximum at 873K. At 923K, the selectivity decreased slightly. These results suggest that a part of the Zn loaded on HZSM-5 may exit the catalyst system, because Zn metal has a relatively low melting point (692K) and boiling point (1203K).

450 i00

I00

I

80

rj

Zn(

_

80 0

o _

:

<

60

60

L

0

~U v,,,,l

0 "m' 40

I Zn(A-P)II_ 4O E O /

=

J

Zn-Sil

i

O

U

O

O

20

.~ 9I,,,,l

-

U

0

I I I ! 723 823 923 Reaction t e m p e r a t u r e / K

0

Figure 2. Effect of reaction temperature on Zn loaded catalysts at 20 g.h.mol 4

100

lOO ~ .

0 ~,

a0 .s

80

6o -

Zn-Sll

0 =

::~ _ "

0 rd3

>

-

[

"

6O ~~

-

__.z~

E

Zn(A-P)

~

O t~

_

40

40

1-

20

U

0

C) 20T 0

,

I 2

~

I i I I I LI 6 8 Process time / h

i

I 10

o

T

Figure 3. Activity and selectivity changes in Zn loaded catalysts with process time at 823K and 20 g.h.mol l.

451 Figure 3 shows the change of butane conversion and selectivity for aromatics formation over three kinds of Zn loaded catalysts with process time at a higher reaction temperature (883K). The catalytic activity and the selectivity of Zn(Imp) decreased quickly with process time. On the other hand, the catalytic activity and selectivity of Zn-Sil and Zn(A-P) stayed at high levels for up to 10 h of process time. These results suggest that Zn loaded in the zeolite framework would be more stable than Zn loaded on the zeolite surface. 3.2. Ga loaded ZSM-5 catalysts It is well known that Ga(Imp) is also a good catalyst for the aromatization of lower paraffins. We found that the addition of Cu cations into the Ga supported HZSM-5 can improve selectivity for aromatics formation. Figure 4 shows the effect of Cu cation in Ga(Imp)Cu(Ex) on activity and selectivity. The conversion of butane decreased with an increasing exchange degree of copper cation. However, the selectivity for aromatics formation increased and then decreased through the maximum point with an increasing exchange degree of copper cation. The selectivity maximum value could be obtained in the region of 45-66 % of Cu ion-exchange degree. The reason why Cu ions can improve selectivity for aromatics formation will be discussed as follows. Figure 5 shows the effect of Cu cation in Cu(Ex) on activity and selectivity. The conversion of butane decreased with an increasing Cu cation. On the other hand, the selectivity for aromatics formation increased and attained the maximum with an increasing Cu cation exchange degree. These results suggest that Cu cation loaded on HZSM-5 slows catalytic activity for dehydrogenation. However, the maximum selectivity value can be obtained at around 90% of Cu ion exchange degree. Above this value for Cu ion-exchange degree, selectivity for aromatics formation decreased. These results suggest that the higher the degree of Cu ion exchange, the weaker the acidity of Cu(Ex). lO0 ,~, "; 80

--

-

Conversion -

U

U

..~

60

o

<

~

40 -

o 20

o

L)

0

I

20

I

I

I

I

I

4:0 60 80 100 120 Cu ion exchange / %

140

Figure 4. Effect of Cu ion exchange degree in Ga(Imp)Cu(Ex) catalyst at 823K and 11.4 g.h.mol 1.

452 100 ?~ 8 0 r ~ ~ l C O n v e r s i o n

~U

[

60

t~

<

40

o 20 o

m

0

I

I

20

0

I

I

I

40 60 80 100 Cu ion exchange /%

I

120

140

Figure 5. Effect of Cu ion exchange degree in Cu(Ex) catalyst at 823K and 11.4 g.h.mol "1. On Cu(Ex)Ga-Sil, Cu cation will be positioned very close to Ga. Therefore, if Cu cation shows any effect of direct interaction with Ga, Cu(Ex)Ga-Sil will have high selectivity for aromatics formation. Figure 6 shows the effect of Cu cation added to Ga-Sil (34) on activity and selectivity. The conversion of butane decreased with increase in the degree of Cu ion exchange. On the other hand, the selectivity for aromatics formation was improved slightly by the addition of Cu cation. These results suggest that the direct interaction between Cu cation and Ga would be not as large on catalytic activity for dehydrogenation. ~. 60

. '50 "~ 40 30

< ~.~.20

Conversion

0 .,..~

~10

o ~) 0

I

20

I

I

I

I

I

40 60 80 100 120 140 Cu ion exchange / %

Figure 6. Effect of Cu ion exchange degree in Cu(Ex)Ga-Sil catalyst at 823K and 11.4 g.h.mol ~.

453 The effect of addition of alkaline earth cations to Ga-Sil(26)is shown in Table 1. The addition of such kinds of divalent cations to Ga-Sil(26) did not improve the activity or the selectivity for aromatics formation. The introduction of alkaline earth cation could weaken the acidity of Ga-Sil. In the case of Ga-Sil, the addition of divalent cation including Cu cation would weaken the acidity too much for the aromatization. These results suggest that the addition of Cu cation could control the acidity of HZSM-5. Table 1. Product distribution in butane aromatization over alkaline earth cation exchanged Ga-Sil at 823K and 20 g-h-mol n of W/ F.* Cation none Conversion / % 92.5 Selectivity / C-% C1-C4 aliphatics 35.9 C5+ aliphatics 1.9 Aromatics 62.2 * Data at a process time of 1 h.

Be 51.5

Mg 58.9

Ca 63.4

Sr 63.8

Ba 51.6

42.5 9.3 48.2

46.6 8.2 45.2

47.3 8.1 44.6

49.3 8.4 42.3

47.8 9.9 42.2

In the first stage of this reaction, butane is converted to butenes and hydrogen, propylene and methane, and ethylene and ethane. In these primary products, methane and ethane are difficult to convert further. Therefore, for a high selectivity for aromatics formation, it is desirable for the primary reaction to be only the dehydrogenation of butane to produce butene and hydrogen. Figure 7 shows the effect of contact time (W/F) over Ga(Imp)Cu(Ex)(66 %) catalyst. At a short contact time, the main product was butenes. Therefore, on this catalyst selectivity for aromatics formation from butane is high. 70 60 50

O 40 ~ ,,,-i

"--~ 30 tJ

F

Butenes

m 20 -

Methane+Propylene

10 Ethane+Ethylene I

0 0

I

I

I

0.4 0.8 log(l+W/F)

1.2

Figure 7. Effect of W/F over Ga(Imp)Cu(Ex)(66%) at 823K.

454 4. CONCLUSION 1. Of the Zn loaded catalysts, Zn(Imp) showed the highest activity and selectivity to aromatics formation. 2. The Zn-Sil and Zn(A-P) catalysts in which Zn was introduced into the zeolite framework, showed a higher thermal stability than the Zn impregnated HZSM-5 catalyst. 3. Cu cation exchanged on HZSM-5 showed weak catalytic activity for dehydrogenation of hydrocarbons, and improved the catalytic selectivity of Ga loaded on zeolites for aromatics formation. 4. Cu cation exchanged on HZSM-5 could control the acid strength of HZSM-5, and this would depress the cracking of butane and C6-C9 olefins as the intermediates of this aromatization. Therefore, Ga(Imp)Cu(Ex)(66%) catalyst showed the highest selectivity for aromatics formation in this work. ACKNOWLEDGEMENT A part of this work was carried out using PEC funding subsidized by the Ministry of International Trading and Industries, Japan. REFERENCES 1. L.R. Aarund, The Oil and Gas J., July 19 (1976) 98. 2. Y. Ono, Catal. Rev,-Sci. Eng., 34 (1992) 179. 3. T. Mole, J. R. Anderson, G. Creen, Applied Catal., 17 (1985) 141. 4. H. Kitagawa, Y. Sendoda, Y. Ono, J. Catal., 101 (1986) 12. 5. G. Sirokman, Y. Sendoda, Y. Ono, Zeolites, 6 (1986) 299. 6. Ono, H. Kitagawa, Y. Sendoda, Sekiyu Gakkaishi, 30 (1987) 77. 7. T. Yashima, T. Sasaki, K. Takahashi, S. Watanabe, S. Namba, Sekiyu Gakkaishi, 31 (1988) 154. 8. T. Inui, F. Okazumi, J. Catal., 90 (1984) 366. 9. P.C. Doolan, P. Pujado, Hydrocarbon Process., 68 (9) (1989) 72.

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All rights reserved.

455

D E V E L O P M E N T OF LIGHT NAPHTHA AROMATIZATION PROCESS USING A CONVENTIONAL F I X E D B E D UNIT S. F u k a s e a, N. lgarashi a, K. Kato b, T. N o m u r a c and Y. lshibashi c

a Petroleum Laboratory, Japan Energy Corporation, 17-35 Niizo-minami 3 chome, Toda, Saitama, 335, Japan b Engineering Department, Petroleum Refining Division, Japan Energy Corporation, 10-1 Toranomon 2 Chome, Minato-ku, Tokyo, 105, Japan c Mizushima Oil Refinery, Japan Energy Corporation, 2- 1 Ushiodori, Kurashiki, Okayama, 712, Japan ABSTRACT A new process of light naphtha aromatization, LNA process, has been developed. The process converts light paraffins containing high concentration of C4-C6 paraffins to aromatics. The development of a new catalyst having a long term stability enabled us to use a conventional fixed bed unit. Based on the results of fundamental and scale-up studies, Japan Energy Corporation has operated 2,250 BPSD demonstration plant and confirmed the good stability of the catalyst. 1. INTRODUCTION Aromatics are mainly produced through the catalytic reforming of These days, light hydrocarbons have become alternative feedstocks production. Several processes have been developed for this reaction: former [2] and Aroformer [3]. The conditions of these processes technology of catalyst regeneration, such as continuous regeneration regeneration, due to rapid catalyst deactivation.

heavy naphtha. for aromatics Cyclar [1], Zrequire special or swing type

The economics of light hydrocarbon aromatization processes does depend on the initial investment cost, mainly construction cost, and the price difference between the feedstock and aromatics. Today's construction cost of refinery processes is becoming expensive. Due to the massive construction cost and no expected widening in the feedstock/BTX price difference, the payout years of a construction cost would be lengthy. One solution of this problem is to develop a new aromatization process using a conventional fixed bed, thus avoiding the need to construct CCR type or swing type reactor unit. Currently in many refineries, conventional "semiregenerated type" heavy naphtha reformers have been replaced by CCR reformers. A number of these units are currently unused and available for

456 another use of light naphtha aromatization. The objective of the development of LNA process, thus, is to develop a new catalyst having extended stability which enables us to use conventional fixed bed reactors, minimizing initial construction cost. Under these circumstances, Japan Energy Corporation has been conducted extensive research on the development of a new aromatization catalyst that exhibits high activity and excellent inhibition of coke formation. Based on this fundamental research, an LNA demonstration plant with a capacity of 2,250 BPSD has been operated in 1994. This paper describes the features of the LNA process and its performance. 2. FUNDAMENTAL AND SCALE-UP STUDIES 2.1. Experimental 2.1.1 Microflow Reactor The reaction was carried out in a stainless steel microflow reactor. In each run, a 2 g portion of catalyst was placed in the reactor and heated to 520 ~ under a nitrogen stream. The nitrogen stream was replaced by a light naphtha vapor fed by a micro- plunger pump. The reaction was carried out at 520 ~ under a pressure of 3 k g / c m 2 G with a WHSV of 0.7 h -1. The products were analyzed periodically by gas chromatography. The properties of the feedstock are shown in Table 1. Table 1 Properties and components of the Feedstock Density (g / cm 3) 0.6591 Sulfur (ppm) <0.1 Nitrogen (ppm) <0.3 H20 (ppm) 13 Components (wt%) n-C5 32.3 i-C5 + C5 naphthene 17.8 n-C6 15.1 i-C6 + C6 naphthene 24.9 n-C7 2.1 i-C7 + C7 naphthene 5.4 n-C8 0 i-C8 0.1 Benzene + Toluene 2.3

2.1.2. Small-scale pilot plant Studies using a small-scale pilot plant was performed. The stainless steel tubular reactor containing 170 g of catalyst was heated by an electric heater. The system pressure was set by a back pressure regulator. The catalyst bed was heated to 500 ~ under a nitrogen stream. The nitrogen stream was then replaced by a

457 light naphtha vapor fed by a plunger pump. Reaction was carried out under a pressure of 3 kg/cm2G, reaction temperature of 500 ~ and WHSV of 0.7 h -1. The p r o d u c t s were collected, w e i g h e d and a n a l y z e d p e r i o d i c a l l y by gas chromatography.

2.1.3. Catalyst Metallosilicates, such as gallo-aluminosilicate (Ga-Al-silicate), and zincoaluminosilicate (Zn-Al-silicate) were synthesized according to procedure described in the paper [4]. Hydrothermal crystallization was carried out in a stainless steel autoclave at 160 ~ for 20 h. Resulting crystals were separated from the solution through a centrifuging process and washed with water. The samples were contacted with a 1M NH4NO3 solution at 80 ~ for 2 h, and thereafter taken from the solution and washed with deionized water. The products were dried at 120 ~ for 12 h and then calcined in air at 540 ~ for 3.5 h. Secondary ammonium-ion-exchange was performed for the calcined sample. The ammonium-ion-exchanged zeolites were dried at 120 ~ overnight, and then calcined at 540 ~ for 3.5 h to obtain H § exchanged zeolites. Zinco-aluminosilicate (molar ratio Si/Al=30, Si/Zn=100) synthesized according to above method was designated as Zn-Al-silicate (A). Zinco-aluminosilicate synthesized without template (molar ratio Si/Al=22, Si/Zn=169) was designated as Zn-Al-silicate (B). Zn-Al-silicate (A) was furthur dealuminated by our proprietary technique of pressurized steaming and was named as Zn-Al-silicate (C). H-ZSM-5 was prepared according to the same procedure described above. Zinc-exchanged ZSM-5 (Zn/H-ZSM-5) was prepared by contacting H-ZSM-5 described above with 0.035M Zn(NO3)2 solution at 80 ~ for 2 h.

2.2. Results and discussion 2.2.1. Stability of Various Catalysts Prepared by Different Methods Aromatization of light naphtha was carried out in the microflow reactor, under the condition stated above. The catalysts examined were Zn-Al-silicate (A), Ga-Al-silicate (molar ratio Si/A1 = 25, Si/Ga = 35), H-ZSM-5 (Si/A1 = 25) and Zn/ZSM-5 (Si/A1 =25, Si/Zn = 150). All the catalysts were calcined at 540~ for 3.5h before use. Change in conversion is shown in Fig. 1. The conversion of light naphtha is defined here by the following equation and is calculated on a carbon number basis: Conversion = (Products- AR0)/(1 - AR0) (1) where, AR0 = fraction of aromatics in the feed naphtha Products = (H2 + (C1 to C4) + (C5 and C6 olefine) + aromatics) in product. All of the catalysts demonstrated 100% conversion at very initial stage of the reaction u n d e r the condition employed. They showed, however, great differences in the stability of the activity. While H-ZSM-5 showed rapid decline of conversion, on the other hand, Zn-Al-silicate (A), showed more stable activity. Accordingly, metallosilicate exhibited better stability.

458 1 O0

'

~'-""~z~

A80 o~ o~'

60

z,....

r o O

O .~. C

"'0

2 0 0

. . . .

i

0

....

H-ZSM-5

i

50

II

Ga-AI-silicate Zn/H-ZSM-5 Zn-Al-silicate

(A)

, , i

,

100

,

I , ,

,

150 (h)

200

Time on Stream F i g . 1 Stability of the Catalyst Prepared by Various Methods

2.2.2. Stability of Catalysts with Different Pore Volume and External Surface Area Zinco-aluminosilicates having different pore volume and external surface area were synthesized and their stability in aromatization was investigated. Table 2 shows some properties of catalysts. External surface area was measured by benzene-filled pore method described by Inomata et al. [5]. Average crystal size was measured geometrically by SEM images. Stability of those catalysts was examined under the same condition as stated above, and the results are shown in Fig. 2. Table 2 Properties of zinco-aluminosilicates Pore Volume (ml / g) Zn-A1-Silicate (A) Zn-A1-Silicate (B)

0.374 0.195

O0

1

Surface Area Total External (m2 / g) (m2 / g) 300 280

~

33 9

,-~----.~..,,~ 2~

~80 "60

2

P 0

>40

tO

O

C

20

Zn-AI-silicate Zn-AI-silicate

O

. . . .

0

i . ,

5 0

,0

I

100

, . . ,

I , ,

150

(B) | (A)

,,I

I

200

Time on Stream ( h ) F i g . 2 Stability of the Catalyst with Different Morphology

Crystal Size (~m) 0.3 3

459 Zn-Al-silicate (B), having smaller pore volume and external surface area, demonstrated greater aging rate. Coke deposited on Zn-Al-silicate (B) was found to be less than that on Zn-Al-silicate (A). These suggest that greater deactivation of Zn-Al-silicate (B) was due to pore mouth blocking. Therefore, morphology of the zeolite is a key factor in maintaining stability of the catalyst.

2.2.3. Stability of Catalysts with Different Acidity Stability of zinco-aluminosilicates having different acidity was examined. Temperature Programed Desorption (TPD) of ammonia was carried out to measure the acidity according to the method reported by Niwa et al. [6]. Comparison of TPD spectra for two zinco-aluminosilicates are shown in Fig. 3.

Zn-AI-silicate (A)

A

100 200 300 400 500 Temperature (C) Fig. 3 NH3 - TPD spectra

Their stability in aromatization was also examined in the same way stated above. Zn-Al-silicate (C), having a smaller intensity of the TPD peak showed far greater stability than Zn-Al-silicate (A) as is shown in Fig. 4. The analysis of the amount of coke deposited on the catalyst showed 28 wt% for Zn-Al-silicate (A) after the experiment of 170 h and 44 wt% for Zn-Al-silicate (C) after 2000 h of experiment, indicating lower coke-forming rate of the latter catalyst. This indicates the importance of acid-property control in order to prepare the catalyst with a long-term stability. Typical yield pattern of Zn-Al-silicate (C) is also shown in Table 3.

460 1 0 0 ~, - , ~,.," " C C CO ,..,,.,,~ ~80 v

c 60 O L_ O

>40 0

o

2 0

O

0

,

,

,

,

I

,

,

,

,

Zn-AI-silicate (C) Zn-Al-silicate (A) I

,

,

.

.

I

.

.

.

I

.

0

500 1000 1500 2000 Time on Stream (h) Fig. 4 Stability of the Catalyst with Different Acidity

Table 3 Typical Yield Product H2 C1 - C2 C3 - C5 Benzene Toluene Xylene C9 + A r o m a t i c s

(wt%) 1.8 25.8 21.3 12.9 22.9 11.9 3.4

Stability of the catalyst after regeneration The stability of the catalyst after regeneration was examined using the smallscale pilot plant. A catalyst having a relatively higher coke-forming rate was used for this experiment to carry out the experiment in a short period. Regeneration was carried out at 450 ~ under a diluted air pressure of 5 k g / c m 2 G for 96 h. The variation of the conversion is shown in Fig. 5.

2.2.4.

100

l

8O v

,..60

Regeneration

0 s._

>

9 40

C 0

2O 0

,

0

,

,

I

200

,

.

,

I

i

i

,

I

.

,

,

I

,

,

,

400 600 800 1000 Time on Stream (h) Fig. 5 Catalyst Stability after Regeneration

461 No substantial decline of activity was found after regeneration. No change in chemical or physical properties of the catalyst was also observed before and after the regeneration. The catalyst showed high stability towards regeneration. 3. PROCESS DESIGN The concept of LNA process is to utilize an existing fixed bed unit, which is typically unused catalytic reformer replaced by CCR, with minor modification. Based on the results of the fundamental and scale-up studies, a 2,250 BPSD demonstration plant was designed, which was originally used as a conventional platforming unit.

3.1. Reactor Design A kinetic reaction model was developed and used with a process simulator to conduct reactor design [7]. The model was also used to investigate the optimum catalyst loading pattern to achieve the highest conversion and aromatics yield in the demonstration plant which consists of three reactors with preheaters in series. This was because a large temperature decrease was expected through the adiabatic reactors of the plant due to the endothermic nature of the whole reactions. 3.2. Preheater furnace tube design Adiabatic and endothermic reaction requires preheating and inter-heating furnaces in the reaction section. Even though hydrogen is being produced in the reaction, there is no hydrogen in the feed preheater furnace. In contrast, in most refinery furnace operations, when hydrocarbon vapor is heated higher than 500 ~ the vaporized feed usually contains hydrogen or steam to prevent coking inside heater tubes. Thermodynamic considerations [8] indicate that our reaction condition also is within the coking region which could cause tube coking. We have, therefore, chosen shorter residence time in the heater tubes to avoid severe coking during furnace operations. 3.3. Catalyst regeneration Small-scale pilot studies have demonstrated that the extended stability of the catalyst is such that it is well suited for service in semi-regenerative type operation using conventional fixed bed reactors. Semi-regeneration is a remarkable feature of this process compared with other light paraffin aromatization processes. Attention is required in regenerative operation to avoid catalyst degradation, which is due to coke-burning, because zeolite structure may collapse in severe hydrothermal atmosphere at higher temperature. In a commercial adiabatic reactor there is some concern about temperature distributions inside the reactor, in comparison with the isothermal small-scale pilot plant. We investigated, therefore, temperature profile in the catalyst bed to scale-up the process. The following study was conducted to investigate the temperature distributions during regeneration of large packed column.

462 Heat transfer of packed bed has been the subject of numerous studies. For cylindrical packed columns, a solution for determining temperature distributions was given using Bessel functions. Here, it is important to find out exact effective thermal conductivity of bed because of flowing gas and relatively high temperatures. Radial temperature distributions are more important than that of axial direction because the latter can be measured and controlled during the operation. Results of investigation have shown that the maximum temperature drop through radial direction does not exceed 45 ~ thus supporting system's operability [9].

4. PROCESS DESCRU~ION OF THE DEMONSTRATION PLANT A simplified flow diagram of the LNA demonstration plant is shown in Fig. 6. The plant is designed to prove the technology at semi-commercial scale. Process facilities consist of two major sections : reaction and product recovery. The reaction section includes a preheater, interheaters, and reactors. The recovery section includes separator, atmospheric fractionator. The predominance of dehydrogenation and cracking reactions causes the overall sequence to be highly endothermic. The LNA process employs adiabatic reaction stages with interheater to achieve optimum conversion and selectivity to aromatic product. The fresh feed, which consists of C4, C5, and C6, is directly charged into the preheater without combining with a recycle stream of unconverted feed. This is because high conversion level can be attained during the operation and almost no unconverted feed is obtained. Hydrogen stream from the separator, however, can be recycled to investigate the effect of hydrogen to hydrocarbon ratio at the inlet of the reactor.

R(

Gas

Heaters

LPG

Reactors

~eparazor

Reformate Stabilizer

~ ) Feedstock Fig.6 Simplified Flow Diagram of LNA Process

463 Semi-regeneration is a remarkable feature of this process compared with other light paraffin aromatization processes. A temperature swing adsorption type dryer is installed to eliminate water produced by coke-burning during regeneration, in order to avoid catalyst degradation. It is designed to remove water from the reactor effluent gas, thus avoiding water accumulation in the recycle gas stream. The method of regeneration is similar to that of conventional catalytic reforming. By introducing diluted air into the catalyst bed, the coke on the catalyst is burned off. This is carried out in several stages. The first stages is carried out under relatively mild conditions. In later stages, the temperature and the inlet oxygen concentration are increased. The temperature of the catalyst bed and the oxygen concentration at the outlet are monitored. Based on the basic design stated above, a demonstration plant was set up to verify the LNA process in Japan Energy Corporation's Mizushima Oil Refinery in Japan. The plant has been operated to aromatize light naphtha at the rate between 1,500 ~ 2,250 BPSD in 1994. 5. PROCESS PERFORMANCE OF DEMONSTRATION PLANT After the start-up, the plant achieved long-term operation without any trouble. Similar reaction yield acquired in the small-scale pilot plant was observed for several sets of operating conditions. The stability of the catalyst in the demonstration plant is shown in Fig. 7.

100

'-'

'- . . . . . .

-'

"--

0

~,

~

o

~

80 v

60 40 20 0

,

0

,

,

I

200

,

,

,

I

,

,

,

I

,

,

,

J

I

i

I

400 600 800 1000 Time on Stream (h) Fig. 7 Stability of the Catalyst in the Demonstration Plant

464 6. CONCLUSION A new light naphtha aromatization process has been developed using a conventional fixed bed reactor. Fundamental study revealed the importance of preparation method, morphology, and acid property to increase the catalyst stability. Based on fundamental and scale-up studies, a demonstration plant was designed and operated. This operation confirmed the good stability of the catalyst. ACKNOWLEDGEMENT The demonstration plant work has been sponsored by Petroleum Energy Center in Japan, which is supported by Japanese Ministry of International Trade and Industry. The authors wish to thank Professor T. Inui of Kyoto University for his valuable suggestions. REFERENCES

1. C. D. Gosling, F. P. Wilcher, L. Sullivan and R. A. Mountford, Hydrocarbon Processing, 70 (12) (1991) 69. 2. S. Saito, K. Hirabayashi, S, Shibata, T. Kondo, k. Adachi and S. Inoue, Paper presented at 1992 NPRA annual meeting, New Orleans, March 22-24, 1992, AM-92-38 3. J. C. Barbier and A. Minkkinen, Paper presented at 1990 JPI Petroleum Refining Conference, Tokyo, October, 1990. 4. S. Fukase, H. Kumagai and T. Suzuka, Appl. Catal. A-General, 93 (1992) 35. 5. M. Inomata, M. Yamada, S. Okada, M. Niwa and Y. Murakami, J. Catal., 100 (1986) 264. 6. M. Niwa, M. Iwamoto and K. Segawa, Bull. Chem. Soc. Japan, 59 (1986) 3735. 7. K. Kato, S. Fukase, T. Amaya and Y. Sato, J. Japan Petrol. in press. 8. J. M. Harrison, J. F. Norton, R. T. Derricott and J. B. Mariot, Wekstoffe u. Korrosion, 30 (1979) 785. 9. K. Kato and S. Fukase, J. Japan Petrol. Inst. 37 (1994) 77.

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All fights reserved.

465

I M P R O V E M E N T IN THE PERFORMANCE OF NAPHTHA R E F O R M I N G CATALYSTS BY THE ADDITION OF PENTASIL ZEOLITE Jorge Norberto Beitramini a and Ronghui Fang b

aDepartment of Chemical Engineering - King Fahd University of Petroleum &Minerals, Dhahran, 312 61, SA UDI ARABIA bSwiss Federal Institute of Technology, CH-8092, Zurich, SWITZERLAND ABSTRACT Effective conversion of normal heptane to more valuable high octane number aromatics was studied over conventional bimetallic naphtha reforming catalysts physically mixed with pentasil zeolite. The addition of zeolite to the reforming catalyst resulted in an improvement of catalytic activity and stability. Aromatic selectivity was initially improved, even though not considerably, having demonstrated the right direction towards improving aromatic production. A reaction mechanism is discussed in light of the established catalytic behavior of the studied catalysts. 1. INTRODUCTION Mono and bimetallic Pt and Pt-Re naphtha reforming catalysts have long been used in the petroleum refinery industry. In spite of their good performance, refineries, driven by the demand for high octane, "lead free" fuel, have been looking for ways of improving reformer performance by either enhancing the octane number of the product or lengthening the catalyst life, i.e., decreasing coking rate. A potential way of meeting these goals would be an improvement of the catalytic properties by modification of the structure of the active sites of the catalyst. Several patents have been recently filed which claim the improvement in the production (yield) of the gasoline fraction while concurrently increasing its octane rating with the addition of zeolites to a conventional reforming catalyst. Zeolites, that have been investigated, include: mordenite, ZSM-type, L, X, Y, SAPOs and others [1,2]. However, no effort has been made to compare these zeolite-containing reforming catalysts to establish the best zeolite combination for this process, and the best arrangement for zeolite addition to the conventional reforming catalyst. On the other hand, there is the question of the close relationship between coke formation (catalyst stability) and aromatics cyclization (catalyst selectivity) over the acid sites present in the zeolites. On naphtha reforming, coke formation is a bifunctional reaction requiring the dehydrogenation capacity of the metallic function and the condensation capacity of the acidic function. Therefore, it is interesting to

466 study the changes in the acidity function due to zeolite addition in order to decrease its condensation capacity while keeping unaltered its aromatic selectivity.

Another aspect of catalyst deactivation is sulfur poisoning. Sulfur can cause poisoning by irreversible adsorption onto the active sites of the catalysts, although sulfur can be beneficial in suppressing the hych'ogenolysis activity of the metallic component [3]. This prevent carbon deposition and prolongs catalyst life. The behavior of zeolite-containing reforming catalysts in the presence of sulfur is therefore of importance for the stability of the process. With this concepts in mind, the catalytic behavior of n-heptane over Pt-Re/A120 3 and/or pentasil zeolite composite catalysts has been extensively studied under deactivation conditions. Specifically, the following points were examined with reference to n-heptane conversion, product distribution and catalyst stability: (i) Effect of the mode of addition of ZSM-5 to the Pt-Re/A120 3 catalyst bed. (ii) Effect of the amount of ZSM-5 added to the Pt-Re/A1203 catalyst. (iii) Effect of catalyst presulfidation. 3. E X P E R I M E N T A L A laboratory made Pt(0.3%)-Re(0.3%)/A1203-Cl(1.00%) prepared according to the techniques described elsewhere [4] was used in this study. Zeolite material used included ZSM-5 type laboratory made with Si/A1 ratio of 72.4 and average crystal size ~f 0.8 (~tm). All catalytic runs were performed in a fixed bed laboratory reactor. Operating conditions were as follows unless otherwise stated: pressure = 1 atm, temperature = 500 ~ WHSV = 8 h "l, hydrogen to hydrocarbon ratio = 8 mol/mol and time on stream = 4 hours. As normal heptane is a major component in naphtha, it was chosen as a typical feed for this set of experiments to give a realistic picture of the industrial reforming process. One gram of reforming catalyst and half gram of zeolite catalyst were used in each run. When ZSM-5 was put at the top or bottom of the reforming catalyst bed, a quartz wool plug was used to separate Pt-Re/A120 3 and ZSM-5 to facilitate a separate examination of the catalyst at the end of run. For runs involving intimately mixed zeolitic materials and reforming catalyst, the two were well mixed by shaking before loading into the reactor. Catalyst pretreatment and presulfidation were performed according to the procedure described by Bickle et. al. [5]. 4. RESULTS AND DISCUSSION

4.1. Effect of m o d e of zeolite a d d i t i o n to the n a p h t h a r e f o r m i n g c a t a l y s t The effect of method of zeolite addition to the catalyst bed on the reforming reactions was investigated in accordance with three different addition modes: 1. Mode T denotes the zeolite ZSM-5 added at the top of the reforming bed, where the feed is first contacted with ZSM-5.

467 2. Mode E refers to the zeolite ZSM-5 added at the end of the reforming bed, where the feed is first contacted with the Pt-Re/A1203 catalyst. 3. Mode M indicates that the zeolite ZSM-5 was intimately mixed with the conventional reforming catalyst. Figure 1 illustrates the extent of n-heptane conversion over those catalysts as a function of time-on-stream, working under deactivation conditions. The catalyst deactivation was modeled by the well-known Voorhies equation C= At -B [6], where C denotes percentage conversion, t denotes time-on-stream (min), and A and B are empirical parameters. Even though the Voorhies equation was originally used to correlate the coke content with time, the use of the equation for correlating activity and time was also confirmed by Magnoux et. al. [7].

w

I00

I

I

I

'

l

'

I

OPt-Re 13 Z S M - 5 V To P M~O Mix AEnd Z --_-d____e/Pr o_~_.___~n edi c t i

L

80 e'O

-9

60

L_

m"

o

^hAUl#

40

0 20 0

I00

200

500

400

T i m e - o n - s t r e a m (min)

Figure 1. Conversion of n-heptane as a function of time-on-stream over Pt-Re/A1203 and/or ZSM-5 with different modes of addition of ZSM-5.

It can be seen from Figure 1 that the Voorhies model correlates well with the performance of those catalysts. The values of the parameters obtained are summarized in Table 1. It is demonstrated that the composite catalysts possess higher activities than either Pt-Re/A1203 or ZSM-5, with the mixed catalyst (mode M) being the highest. These results may suggest different reaction paths for n-heptane over those catalysts.

468 Table 1 [nitial conversion and Voorhies deactivation coefficients in n-heptane transformation over Pt-Re/A120 3, ZSM-5 and the different addition modes.

COMPOSITE

Catalyst a

Pt-Re/AI203

Conversion A B Norm b

52.4 75.7 0.270 2.49

ZSM-5 65.8 76.7 0.124 6.35

CATALYSTS

Mode T

Mode M

92.9 99.8 0.095 12.62

99.4 112.7 0.081 4.89

Mode E 71.2 90.2 0.169 6.79

a Initial data taken at 4 minutes-on-stream. bNorm is defined as square root of the sum of squares of the residual. A and B empirical constants Over Pt-Re/A1203 n-heptane is first converted to heptenes on the metal surface [8]; the r e s u l t a n t olefin is then protonated over the alumina acid sites to generate a C7H15 + carbenium ion according to:

GH,, -0 G H . + H,

(1)

GHu

(2)

-/-H + --+ C z H

For hydrocarbon transformation on H-ZSM-5, it has been generally accepted t h a t the carbenium ions are the common intermediates, but the initiation of the reaction might vary. Possible initiation routes for the reaction include a direct protonation of the n-heptane molecule on Bronsted acid sites to produce a carbonium ion (protolytic route) [9], according to:

C7H16 -/-H § --+ G H/7

(3)

It has also been shown t h a t hydride abstraction on Lewis acid sites to generate carbenium ion may occur if non-framework aluminum species exist in the zeolite [10]. However, for the current experiments the contributions from Lewis acid sites of HZSM-5 seem negligible, since no non-framework aluminum species were identified by 27A1 NMR. On the other hand, the experimental results support the idea t h a t on the composite catalyst, the activation of the n-heptane molecule could be considered as a combination of the modes described above, represented by the following reaction: 4-

(i.n) -C, H2,+, +. -C7H16 ~ (i,n) -C, H2,+2 + n -C7H15

4-

(4)

469 The higher activities of composite catalysts compared to those of Pt-Re/A1203 are ascribed to the increase in acidity of the catalysts upon introduction of H-ZSM-5. It is well known that H-ZSM-5 has much stronger acidity than amorphous alumina, by the fact that the production of carbenium ions is energetically favored over that of carbonium ions [11]. The presence of metal components will facilitate the rate of isomerization and cracking of paraffins, since the paraffins can be dehydrogenated on the metal sites and the resultant alkenes are much easier cracked or isomerized [12]. It can be seen from the product distribution (Figure 2) that selectivity to cracking reaction dominate over aromatics and isomers selectivity between ZSM-5 containing catalysts and ZSM-5 free catalysts. Since Pt-Re/A1203 is not sulphided, a high level of cracked products is expected to be formed. They are mainly produced by the metal function via hydrogenolysis, and further crackates can be also generated by the acid function of the bifunctional catalyst. With the addition of ZSM-5, the major products are C 3 and C 4 species resulting from the acidic nature of the zeolite, which can be accounted for by the carbenium ion mechanism [13]. It is found that Mode E behaves similarly to the zeolite free Pt-Re/A1203. Both catalysts have a relatively high proportion of isomer products which could be formed over the metal surface via a bond-shift mechanism [8]. Isomers are formed by doublebond isomerization and skeletal isomerization reactions at both the acid sites of the alumina support and the metal sites. The later provides a dehydrogenationhych'ogenation function and the acid sites an isomerization function for the olefins to dehydrogenate from paraffins over the metal function, since it is known that olefin isomerization proceeds much quicker than the respective paraffin isomerization [8]. On the other hand, branched paraffins are less easily cracked than linear ones [10]. Therefore, once isomers are formed over conventional reforming catalysts, they are likely to be the final products. Evidently, the isomerization of paraffin requires the metal function in the bimetallic catalyst, and so does the paraffin aromatization. This can also explain the observed decrease in the isomers and aromatics production with time-on-line since it is well- known that coke preferentially deposits on a metal surface first [ 14]. Aromatic hydrocarbons can be formed either on metal sites via dehydrocyclization, or on acid sites via hydrogen transfer between olefins and naphthenes, or olefin oligomerization. The decrease in aromatic selectivity observed when zeolite is placed at the bottom of the catalyst bed (Mode E) can be explained by the fact that the residence time over the metal catalyst is considerably reduced (by 1/3) since the contribution of ZSM-5 alone to the aromatics is not substantial as can be seen from Figure 2. It is known that the hydrogen transfer reaction over acid zeolite will produce aromatics [15], but it must be realized that the conditions chosen here (high hydrogen partial pressure, low residence time) do not favor the formation of aromatic hydrocarbons. From the above discussion, it is conceivable that mode T would also produce less aromatics and isomers than Pt-Re/A1203 alone, since the configuration of zeolite at: the top of the catalyst bed would crack the feed into smaller hydrocarbon fragments.

470 They are less reactive, especially over conventional reforming catalysts which possess less acidity. I

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Product Yield (mol/lO0 mol feed) Figure 2. Product selectivity of n-heptane transformation over Pt-Re/A1203 and/or ZSM-5 with different modes of ZSM-5 addition. The intimate mixture of ZSM-5 with Pt-Re/A1203 (Mode M) presents a different picture. With those composite catalysts, fewer isomers are produced while more crackates are generated. This mode gives the highest activity, which can be ascribed to the closeness of the metal and acid sites. When the two types of acid sites are close enough, the metal dehydrogenation reaction have a sinergistic effect. The metal sites produce olefins while the strong acid sites from H-ZSM-5 crack them into smaller fragments, since olefins are easily cracked. On the other hand, it can also be speculated that the restricted spatial volume of ZSM-5 will inhibit the formation of certain bulkier intermediates (transition state selectivity), and will therefore promote the breaking of C-C bonds and or C-H bonds to generate smaller hydrocarbons. Therefore, it can be concluded that by generating olefin intermediates, the dehydrogenation capacity of the metals in the presence of ZSM-5 enhance the rate of cracking of normal paraffins. The extent is greater with catalyst not exhibiting transition-state selectivity toward the cracking reaction. A similar result has also

471 been reported by Riley and Anthony [16], who studied n-heptane cracking over metalZSM-5 catalysts. The observed improvement in stability for ZSM-5 containing catalysts (Figure 1 and Table 1), is mainly due to the unique structure and novel configuration of ZSM-5. The well-known transition-state shape selectivity restricts the formation of aromatic hydrocarbons with carbon number higher than 10 [17], decreasing the rate of formation of heavier aromatics that are believed to be the precursors of coke, that mainly cause catalyst deactivation by occupying or blocking the way to active sites. In conclusion, for the different modes of zeolite addition, it appears that the use of a completely mixed Pt-Re/A1203 with ZSM-5 is the best combination. It gives higher activity while maintaining comparable aromatic selectivity. 4.2 Effect of t h e a m o u n t of z e o l i t e a d d e d to t h e Pt-Re/A1203 c a t a l y s t The initial and final product breakdowns of n-heptane transformation over PtRe/A1203-ZSM-5 composite catalysts which differ in ZSM-5 amounts, are summarized in Table 2. It is very interesting that with an addition of a small amount of ZSM-5 to the reforming catalyst, the activity and product distribution can be changed dramatically, manifesting the strong interaction of ZSM-5 with the feed and possible change on the way of the n-heptane reforming reaction. The predominant C 3 and C 4 yields are indicative of strong acid centers with ZSM-5 as discussed in the previous section. Table 2 The initial and final conversion and catalytic selectivity in n-heptane transformation over Pt-Re/A1203 + ZSM-5 with different amount of ZSM-5.

ZSM-5 (wt%) Conversion

0

Pt-Re/AI203 + ZSM-5 10

5

I

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,Selectivity (mol/1 O0 mol feed) I

c1 +c2 C3 + C4 Cracking Isomers Arom. Liquids RON I: initial

F

I

F

12.3 5.5 16.4 11.8 14.9 8.3 142.8 1 0 8 . 4 35.6 16.1 167.0 1 2 5 . 0 22.0 6.6 2.3 1.3 13.6 0.9 16.3 2.5 91.6 95.1 26.8 44.7 37.0 6.7 101.1 17.8 value at 4 min, F: final value at

20.8 13.7 143.1 96.0 173.1 115.8 2.4 1.5 15.0 1.9 26.6 51.0 100 15.3 4 hours.

I

24.4 148.4 181.4 2.1 12.8 23.6 98.1

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17.2 84.2 107.5 1.2 1.2 56.2 12.3

39.0 150.4 196.4 2.2 10.2 19.6 95.9

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34.0 95.1 135.4 0.7 1.7 47.3 15.7

Figure 3 shows a systematic decrease in i-C4/C 4 ratio with an increase in the C1 + C2/ C 3 + C 4, C3/C 4 and C2/C 5 ratios when the amount of ZSM-5 added to the PtRe/A1203 catalyst increases, suggesting that the activation of the n-heptane molecule

472 follows a direct protonation on a strong Bronsted acid site that produces a pentacoordinated carbonium ion according to reaction 3. The carbonium ion C7H17 + in turn undergoes simple cleavage to generate smaller alkanes (C3H 8, C4H10, C2H 6, C5H12, CH4). It can also be seen from Figure 3 that the initial i-C4/C 4 ratio is greater than the final one and the intial C 1+ C2/C 3 + C4, C3/C 4 and C2/C 5 ratios are smaller than the final ones. In fact, it can be easily claimed that the metal components are preferentially deactivated as the reaction proceeds, as observed with the decrease in i-C4/C4 ratio that reveals a decrease contribution from the classical carbenium ion mechanism (reaction 2). ' I ' I ' 1.6 I ' I ' '

0.8

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ZS M - 5 Content (wt%) Figure 3. Light hydrocarbons selectivity ratio of n-heptane conversion over PtRe/A1203 and/or ZSM-5 as a function of ZSM-5 additions. (O initial, 9 final) On the other hand, CH 4 can mainly occurs via hydrogenolysis reaction catalyzed by metals. Present results shows that CH 4 is reduced to the increased portion of zeolite in the composite catalysts. It was also found that deactivation on ZSM-5 containing catalysts is lower than on Pt-Re/A1203 alone. The results from Table 2 also suggest that the addition of ZSM-5 improves the stability of the conventional reforming catalysts. On the basis of the high RON, it would appear that the addition of different amounts of ZSM-5 to the reforming catalyst is the best choice. This is not true since also has the lowest liquid yield. The selection of the best catalyst combination should be considered as a balance between liquid yield and RON. To low a liquid yield results in most of the feed being wastefully cracked to less useful products. 4.3 Effect of c a t a l y s t p r e s u l f i d a t i o n on the c o m p o s i t e c a t a l y s t The composite catalysts described above demonstrated better activity and stability than conventional reforming catalysts. However, the liquid and aromatic yields were

473 not increased to any extent due to cracking reactions promoted by the addition of ZSM-5. On the other hand, methane and ethane were also produced by the metal functions, especially the Re metal which is well-known for its high hydrogenolytic capacity [1]. To reduce the hydrogenolysis capacity of those catalysts, sulfur is widely employed for its selectivity poisoning effect in reducing hydrogenolysis [3]. It would then be interesting to study the effect of sulfur on the performance of the composite catalysts. For instance to answer questions like "Will the catalytic activity and aromatic selectivity decrease upon suppression of the metal function by sulfur? and, Are there any potential links between coke formation and sulfur poisoning on composite reforming catalysts?", would be pertinent to understanding reforming reactions catalyzed by sulfided composite catalysts. This set of experiments were designated to answer the above questions, since there is no literature available regarding the effect of sulfur on zeolite-containing reforming catalysts. , I00

,

,

I

'loPt-Re I

~

I(D Pt - Re - S

\9

JE3 P t -

Re+ZSM-5

S 80 0

N 9

60

tD C

o

4o

r 20 ,,

I00

Time-on-

200

stream

.

300

400

(min)

Figure 4. Effect of time-on-stream on the catalytic activities of n-heptane over un- and presulfided Pt-Re/A1203 and Pt-Re/A1203+ZSM-5 catalysts. The catalytic activities and selectivities of presulfided and unpresulfided Pt-Re/A1203 and Pt-Re/A1203-ZSM-5 catalysts are depicted in figures 4 and 5. It is clearly demonstrated that presulfidation will decrease the initial activity of both zeolite-free and -containing catalysts, which is expected due to the selective poisoning action of sulfur [3]. However, sulfided catalysts show no deactivation over time, especially for the ZSM-5 containing catalyst, which actually demonstrated to produce higher activity and aromatic selectivity after 4 hours-on-stream than the corresponding unsulfided catalyst. The result concerning sulfur effects on performance of naphtha reforming catalysts is in line with results of previous authors [18]. For the composite Pt-Re/A1203-ZSM-5 catalyst, the increased activity and aromatic selectivity can be ascribed to the suppression of hydzogenolysis by sulfur adsorption on the rhenium surface reducing the ensemble size of platinum [3]. As a result, the

474 contribution of acid reactions will increase since the routes of metal reactions to produce methane and ethane by hydrogenolysis, aromatic by dehych'ocyclization and isomerization via double-bond shift are inhibited. More important, the reaction which converts olefins to paraffins on metal sites by hydrogenation is also suppressed. As a consequence, relatively large portions of olefins are available for olefin oligomerization which can be catalyzed by the acidic ZSM-5 zeolite. Therefore, more aromatics will be produced as shown in figure 5. On the other hand, sulfur has no negative effect on the performance of H-ZSM-5 as demonstrated previously [19]. In conclusion, the fact that the aromatic selectivity over the ZSM-5 containing reforming catalysts is considerable increased after catalyst presulfidation implies that presulfided zeolite containing conventional naphtha reforming catalysts may offer the potential to boost the octane number in the product pools of the naphtha reforming process. 20o~ , , , , , , , 150 I00

-

Crackets

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. t

.

. I

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lib

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I00 Time-

200 on-

stream

300

400

(min)

Figure 5. Effect of time-on line on product selectivities over Pt-Re/A1203 andJor PtRe/A1203+ZSM5 with and without addition of sulfur. CONCLUSIONS Catalytic conversion of n-heptane over Pt-Re/A1203 and H-ZSM-5 composite catalysts has been studied under deactivation conditions. Ample experimental evidence was given suggesting that both the carbonium ion route and the classical carbenium ion

475 route are effective in the present reaction system. It appears that the completely mixed composite catalyst is the best choice regarding high activity while maintaining substantially enhanced aromatic selectivity and catalyst stability. However still showed to have the lowest liquid yield due to dominance of cracking reactions resulting from the high acidity levels of ZSM-5 zeolite. A larger increase in aromatic selectivity and liquid yield was observed after presulfidation. This finding suggest that a potential naphtha reforming catalyst based on presulfided Pt-Re/A1203-ZSM-5 could be formulated to boost the octane number of reforming products in a commercial reforming unit. REFERENCES

1. P.K. Coughlin and R.J. Pellet, European Patent 0,242,616 (1987) 2. N.Y. Chen, W.E. Garwood and F.G. Dwyer, Shape Selective Catalysis in Industrial Applications, Marcel Dekker, N.Y., (1989). 3. J.H. Sinfelt, Bimetallic Catalysts: Discoveries, Concepts and Applications, John Wiley & Sons, N.Y., (1893). 4. J.N. Beltramini and D.L. Trimm, Appl. Catal., 32, 71 (1987). 5. G.M. Bickle, PhD Thesis, University of Queensland, (1989). 6. A. Voorhies, Ind. Eng. Chem. 37, 318, (1945). 7. P. Magnoux, P. Cartraud, S. Mignard and M. Guisnet, J. Catal., 106, 242, (1987). 8. B.C. Gates, J.R. Katzer and G.C.A. Schuit, Chemistry of Catalytic Processes, McGraw-Hill, N.Y., (1979). 9. A. Corma, J. Planelles, J. Sanchez-Marin and F. Tomas, J. Catal., 93, 30, (1985). 10.A. Corma, J. Planelles and F. Tomas, J. Catal, 94, 445, (1985). 11.A. Corma, Stud. Surf. Sci. Catal., Elsevier, Amsterdam, 49, 49, (1989). 12.B.W. Wojciechowski and A. Corma, Catalytic Cracking: Catalysts, Chemistry and Kinetics, Marcel Dekker, N. Y., (1986). 13.H. Pines, The Chemistry of Catalytic Hydrocarbon Conversions, Academic Press, N.Y., (1981). 14.C.L. Pieck, E.L. Jablonski, R.J. Verderone and J.M. Parera, Appl. Catal., 55, 1, (1989). 15.N.Y. Chen and W.O. Haag, Hydrogen Effect in Catalysis, Z. Paal and P.G. Menon, (Eds.), Marcel Dekker, N.Y., 695, (1988). 16.A. Riley and H. Anthony, private communication. 17.E.G. Derouane, Stud. Surf. Sci. Catal., Elsevier, Amsterdam, 5, 5, (1980). 18.J.N. Beltramini, Deactivation by Poisoning and Sintering, in Catalytic Naphtha Reforming- Science and Technology, G.J. Antos, A.M. Aitani and J.M. Parera, (Eds), p.313, Marcel Dekker, N.Y., 1995. 19.J.N. Beltramini and R. Fang, Proc. Int. Symposium on Zeolites and Microporous Crystals, Nagoya, (1993).

ACKNOWLEDGMENT Dr Jorge Norberto Beltramini wish to acknowledge the support of the King Fahd University of Petroleum and Minerals during the preparation of this work.

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Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All fights reserved.

477

Z E O L I T E CATALYSTS IN THE UPGRADING OF L O W - O C T A N E H Y D R O C A R B O N F E E D S T O C K S TO UNLEADED GASOLINES

V. G. Stepanov, K. G. lone and G. P. Snytnikova Scientific - Engineering Centre Novosibirsk, 630090, Russia

"Zeosit",

G. K. Boreskov Institute o f

Catalysis

1. ABSTRACT Processing different hydrocarbon raw materials into high-octane gasolines over zeolite-containing catalysts with pentasil structure has been studied in the absence of hydrogen in dependence on the technological parameters of the process. Correlations have been found between the hydrocarbon composition of the feed and the yield of gasoline and its composition. The possibility of increasing the octane numbers of light petroleum naphtha and gas condensate from 56-60 to 85-90 MON without hydrogen application has been shown. For low-octane number raw material, upgrading the catalyst IC30 without previous hydropurification of the feed results in a decrease of the total sulfur content in synthesised gasolines to 0.1 wt % and simultaneous improvement in their antidetonate parameters. The non-hydrogen transformation of light petrol fractions of oils and gas condensates over IC-30 catalysts to increase the octane number and to reduce the sulphur content was performed on a pilot and industrial scale over three years. 2. I N T R O D U C T I O N Nowadays because of deterioration of the ecological situation worldwide, there is a tendency towards reduction in the use of leaded gasoline followed by a complete cessation of its use. As a result, increasing interest is shown in development of new catalysts and processes based on them that allow, first the obtaining of gasoline with sufficiently high octane numbers, and second the involvement of unconventional hydrocarbon feedstocks, e.g. gas condensates, petroleum gas, gas gasoline, etc., into standard gasoline production. It is well known that zeolites with pintail structures are active in reactions of isomerization, cracking, aromatization, alkylation, etc., which makes possible their use as an active component of catalysts for a number of processes. Thus, ZSM-5 zeolites are used as catalysts for the transformation of lower alkanes into aromatics [1]; Ni/ZSM-5 zeolite is applied in the M-Forming process to increase octane numbers of reformates [2]; catalysts prepared on the basis of pentasil-type zeolites are employed in "zeoforming" - the process of unleaded high-octane gasoline obtaining from gas condensate and gas gasoline fractions [3-7]. Here we describe the results of systematic investigations of zeolite H-ZSM-5-type behavior during the processing of different hydrocarbon raw materials with gasoline boiling ranges that depend on reaction conditions.

478 3. EXPERIMENTAL PROCEDURE The catalyst containing H-ZSM-5 zeolite (SIO2/A1203=96, Na20<0.1 wt.%) granulated with 7-A1203 as a binder was investigated in a fixed-bed, down-flow, electrically heated reactor (10 ml, fraction 0.25-0.50 mm) at a temperature interval (TO of 300-460~ pressures (P) of 0.1-0.4 MPa and LHSV of 0.5-7.0 h~. The feed was delivered to the reactor by a metering liquid pump. Overpressure was maintained by N2 from buffer capacity at its consumption 21/h. Reaction products after reactor were passed through a cold-water condenser and then separated into gas and liquid (catalysate) phases. It was established, by independent experiments with quartz instead of catalyst, that in the reactor catalysis as described does not take place. The yield of catalysate was 90 wt % and that of the uncondensed C4-C6 fraction was about 10 wt %. Before the experiment the catalyst sample was activated in air flow in situ at 500~ for 2h, and then was blown by N2 at the same temperature for 0.5 h. During the experiments, i.e., a period of 4 h with hourly collection of probes for gas chromatographic analysis changes in catalyst stability and selectivity were not observed. Light gasoline fractions of oils and gas condensates from different deposits, gas petroleum, reformates, and reforming rafinates as well as model hydrocarbon mixtures were used as feedstocks. 4. RESULTS AND DISCUSSION The hydrocarbon composition of gasoline produced from different raw materials under the same process conditions may essentially differ individually as well as in groups. The main differences are observed in the composition of the paraffin-naphthene fraction of gasoline, and aromatics are represented basically by the toluene-xylene fraction. As a result of the transformation of hydrocarbon feedstocks over H-ZSM-5/AI203 catalyst, hydrocarbon and fraction composition of liquid products widen which leads to an increase of the end boiling point of the catalysate. Simultaneously, the formation of C1 - C4 hydrocarbons and of a small amount of H2 (not more than 3-4 wt % in the gas phase) occurs. The composition of the gaseous products obtained (P>5 MPa) differs only slightly. 4.1 The influence of the reaction temperature

With elevation of temperature, gas formation increases (Figure l a). The content of methane and ethane in the gas phase is increased, the portion of butanes is reduced and the concentration of propane, the main gaseous product, passes through its maximum. Increasing Tr from 340 to 450~ in process under atmospheric pressure leads to the growth of the C2-C4 olefin content from 5 to 20 wt % in the gas phase. When P>0.5 MPa the process temperature actually does not influence the concentration of the gas phase. It does not exceed 3-5 wt %. With an increase of the Tr obtained, the end boiling point ofcatalysates also rises to 220280~ depending on the conditions of the process; the portion of distilled off fraction >195~ is equal to 2-8 vol %. In this case, the content ofparaffines and naphthenes in the gasoline fraction decreases while that of aromatic hydrocarbons increases. It has been shown by the independent experiments that for the model hydrocarbon mixture (isooctane : n-octane : cyclohexane = 1 : 1 : 1 wt), the degree of conversion of isoparaffines owing to shape selectivity is much smaller than that of n-paraffines and

479 a ~d

.

b

c

80" 3

d

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60

-v

v

v"

00 O

40 O

d I

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350

450

Temperature, ~

I

9

I

I

~

-I

2 4 Pressure, MPa

I

I

I

I

=---

2 4 LHSV, h 1

Figure 1. Influence of process temperature, pressure and LHSV on the yield and hydrocarbon composition of the gas phase during the conversion of: (a)gas condensate (45-125~ P=IMPa, LHSV = 2 hl, (b) oil fraction (85-180~ at T=360~ LHSV = 2 h"~ and (r gas condensate (60-155~ at T=380~ P =lMpa. (1- gas yield; 2- content of methane+ethane; 3-propane; 4- butanes; and 5- C2-C4-olefins in C1-C4-fraction).

100

-

d

80-

r~

60-

0

40-

O .F-I .t.a

1

3

5

20~ !

300

350

400

Temperature,

450

~

Figure 2. Model hydrocarbon mixture conversion vs. temperature at P=I MPa, LHSV = 2 h"1. (1- yieM of gasoline fraction; 2- content of Cs+ n-paraffines; 3isoparaffines; 4- naphthenes; and 5- aromatics). naphthenes; along with the rise of Tr, the content of naphthenes and n-paraffines in catalysate falls, the amount of aromatics increases and the isoparaffine passes through its maximum (Figure 2). Increase in T, lead to substantial changes in the composition of the n-paraffines fraction; the portion of n-C7+ is essentially reduced and that of n-C4-C5 is increased. The dependence of the individual composition of the aromatic fraction of gasoline on Tr is of a more complicated nature, that depends on the composition of the initial raw materials. In general, with T, increasing from 360~ to 450~ the amount of benzene and toluene in the aromatic fraction increases while that of the C9+ aromatics is reduced (Figure 3).

480

6O

C8

d .~ 40 ~

~0 20Benzene I

I

I

I

300

350

400

450

Temperature, ~

Figure 3. Dependence of distribution of aromatics on reaction temperature at P = 1 MPa and LHSV = 2 h~. At the transformation of sulfur- and high-sulfur containing hydrocarbon fractions, the desulphurization of gasoline takes place and sulfur evolves in the gas phase in H2S. With increases of the process temperature, the degree of desulphurization increases; then it stabilizes and can reach 95-96 wt %. Liquid products contain about 300-600 PPM total sulfur. With increases of the process temperature, the yield of the gasoline fraction falls and the content of high-octane components is increased while the quantity of total sulfur and lowoctane components is reduced. 4.2 The influence of LHSV With an increase of the LHSV of reaction products, decreasing gas yield and increasing liquid hydrocarbon yields are observed, caused by the diminution of the stock conversion degree. In gaseous products, with increases in the LHSV, the content of Ci-C2 paraffines decreases while the concentration of C2-C4 olefines increases slightly (Figure lc). In gasoline obtained with an increase of the LHSV from 0.5 to 7.0 h"1 , the content of initial stock components, paraffines and naphthenes, increases practically linearly while the amount of aromatic hydrocarbons decreases (Figure 4). On the whole, with increasing the LHSV, increases the yield of gasoline fraction, but the amount of high-octane hydrocarbons in it diminishes and that of the low-octane components rises. 4.3 Composition and properties of the gasoline obtained

When performing the process, the reactions of C-C bond cleavage, isomerization, hydrogen transfer, alkylation of hydrocarbon stock components and intermediates taking place on the active surface of the zeolite result in the transformation of low-octane hydrocarbons (nalkanes, monomethylalkanes and naphthenes) into high-octane components (isoparaftines and arenes). Strongly branched stack paraffines, as a result of shape selectivity of the catalyst, in all practicality, do not undergo the conversion, which preserves the high-octane feed components. The gasoline obtained corresponds to standard motor fuels (Table 1).

481

80 1 3

60

,~ 4o 0 2

4

6

LHSV, h t Figure 4. Gasoline yield and composition vs. space velocity at the conversion of gas condensate (33 - 155~ at Tr = 420~ P = 1 Mpa. (1- yield of gasoline; 2content ofn-paraffmes; 3- iso+~cloparatNaes; and 4- arenes in C5+-fraction. Table 1 Composition and octane numbers of feedstocks and gasoline obtained N

1

2 3

Index

Groupcomposition,wt% C3-C4 Normal alkanes Cyclo + iso-alkanes Arenes Total S content, wt % Fractional composition, ~ Initial boiling point 10 vol %

4

.......................................... F.~..d..s.tp..e.~......~.......a...n...d...g...m....o.!.i.n..e....(.G..) "......................................... Gas Gasoline Gas Condensate Pet.Naphtha F G F G F G F G

50 90 End boiling point MON

2.2 33.1 64.1 0.6 0

5.1 5.3 32.8 56.8 0

30.8 61.5 7.7 0

3.2 6.7 34.3 54.8 O.

.1

.1

.1

1

33 63 82 105 109 68

35 58 112 169 193 86

44 63 94 137 149 66

35 56 94 157 181 86

1.8 31.5 51.1 15.6 1.3

7.0 11.2 38.2 43.6 0.06

36 56 89 109 134 56

31 45 98 150 185 80

32.7 44.2 23.1 0.05

2.3 16.8 48.3 32.6 0.02

85

35 67 115 163 196 78

108

128 159 185 62

The middle-and wide-pore zeolites were investigated in comparison with ZSM-5. As a result of these investigations the technology of the new catalyst IC-30 was developed for low octane number hydrocarbon mixture upgrading. The technology for obtaining high-octane, unleaded gasoline was elaborated on a pilot-scale using real feedstocks: 9 From the low-octane gasoline fraction of gas condensates at the Novo-Urengoy gascondensate plant and Luginetsk deposit; 9 From compressates of oil gas, at the Nizhne-Vartovsk gas plant. (This installation has been in operation for 3 years.); 9 From sulfur-containing gas condensates, at the Orenburg gas-refining plant.

482 5. CONCLUSION Hydrocarbons of different nature having low octane numbers can be converted into gasoline with the properties of motor fuels using middle-and wide-pore-type zeolites. The yield and composition of gasoline obtained are determined by the composition of the initial feed as well as the process conditions. Depending on process conditions, increases in octane numbers from 56 MON to 85-86 MON and higher are possible. The successful industrial application of this technology and type of catalyst has been in progress in the northern Siberia for three years. REFERENCES 1. Y.Ono, Transformation of lower alkanes into aromatic hydrocarbons over ZSM-5 zeolites, Catal. Rev. Sci. Eng. 34(3) (1992) 179-226. 2. Y.Chen, W.E.Garwood and R.H.Heck, Ind. Eng. Chem. Res., 26 (1987) 706 3. K.G.Ione, V.G.Stepanov et al. Patent of Russia Federation No.1325892. Method of producing gasoline fractions. 18.03.1993, appl. 03.10.1984. 4. V.G.Stepanov, K.G.Ione et al. Patent of Russia Federation No.1141704. Method of producing motor fuels from gas condensate. 18.03.1993, appl. 17.06.1983. 5. G.P.Snytnikova, M.N.Radchenko, K.G.Ione, V.G.Stepanov. The production of high-octane gasoline fractions. Gas Industry (Russia) No.4 (1988) 54-55. 6. V.G. Stepanov, A.J. Getinger, G.P. Snytnikova, V.L. Nebykov, K.G. Ione. Catalyic upgrading of gas gasoline of Nijnevartovsk plant on zeolite catalyst. Nettepererabotka and nettehimia, No.12, (1988) 3-6. 7. V.G. Stepanov, G.P. Snytnikova, L.G. Agabalian, K.G. Ione. Autogasolines from fractions of gas condensate. Gas industry (Russia), No. 1 (1989) 54-57.

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All rights reserved.

483

CATALYSTS FOR C6 ALKANE CYCLIZATION N. Ph. Toktabaeva, G. D. Zakumbaeva, and L. V. Gorbacheva

D. V. Sokolsky Institute of Organic Catalysis and Electrochemistry of National Academy of Sciences, 142, Kunaev str., Almaty, 480100, Republic of Kazakhstan ABSTRACT Non-oxidizing dehydrogenation of n-hexane was investigated on high dispersed promoted Pt/A1203 catalysts. The influence of chemical composition and catalyst structure on the direction of reaction and arenes yield was found. 1. INTRODUCTION The basic sources of petrochemical synthesis are benzene and its homologues. The production of these compounds from petroleum is profitable. In 1996, the world requirements for benzene will grow up to 24-26 million tons per year. Non-oxidizing dehydrogenation of alkanes is a subject of intensive investigation. So, the selection and increase of the assortment of highly effective catalysts for the synthesis of olefins and aromatic hydrocarbons from alkanes are very important for development of this branch of industry. There are three main catalysts for non-oxidized dehydrogenation: 1. Zeolite-containing catalysts [ 1,2]. 2. Metal oxide systems and heteropolyacids [3]. 3. Multi-component supported metal catalysts, including industrial reforming catalysts. It is known [ 1,2] that zeolite-containing promoted catalysts are used widely in petroleum processing. They meet the strict standards of industry and ecology. The activities of these catalysts are usually related to their acid-base properties and the presence of structural OHgroups. The conversion of n-hexane on ZnO/H-ZSM-5 catalysts was studied in the range of 2-32 h-1 volume velocity in a quartz microreactor at 500~ and 1 atm [1]. The selectivity of the lowest olefin formation has been shown to be slightly higher than that of paraffins. The lowest olefins and paraffins are formed in the process ofhexane cracking. The selectivity of benzene, toluene, xylene methane, ethane, and propane formation rises with increasing hexane conversion in proportion to the contact time, as the selectivity of the olefin formation decreases. At the same time, the selectivity of olefin formation decreases. This fact shows that nhexane is converted into aromatic hydrocarbons through the intermediate olefin form on the ZnO/H-ZSM-5. The processes of dehydrogenation of n-hexane into hexene and conversion of oligomers into aromatic hydrocarbons occur in the aromatization of n-hexane over ZnO/HZSM-5. The zinc oxide is involved in the n-hexane activation as well as H-ZSM-5.

484 The mechanism of n-hexane conversion over ZnO/H-ZSM-5 catalyst is described by the scheme: n-hexane

ZnO H-ZSM-5 H-ZSM-5 -~ hexene ~ lowest ~ " olefins q

oligomers

ZnO

; aromatic hydrocarbons

The aromatization of n-hexane over ZnO/H-ZSM-5 is a bifunctional reaction and its rate depends on the content of ZnO [ 1]. Platinum bifunctional catalyst on zeolite 13(zeolite with high content of SiO2) has been investigated in the process of n-hexane, methylcyclo-pentane and cyclohexane reforming and compared with that supported on A1203 [2]. For reaction of skeleton isomerization to obtain products with higher octane number than the initial hydrocarbons, the zeolite has been found more active than A1203. Products with more carbon atoms in their molecules than the initial compounds are formed by bimolecular alkylation reaction, which takes place in the zeolite pores. The formation of aromatic compounds is possible only in the case of methylcyclohexane and cyclohexane under specific conditions. The properties of Pt supported on zeolite depend on the relatively high acidity of zeolite 13and the Optimum sizes and structure of pores which promote drawing together different molecules and their interaction. Unlike this, Pt/A1203 catalyst promotes the processes of dehydrogenation/hydrogenation and does not increase the rate of skeleton isomerization and ring enlargement reaction because of the lower acidity of Pt/A1203 catalyst and its larger pores [2]. It is known that conversion of (C7-C10) fractions of petroleum into aromatic hydrocarbons is carried out on industrial Pt/A1203 reforming catalysts. To convert (C2-C6) light petroleum fractions over this catalyst is difficult. The yield of benzene from methane or natural gas can be increased by promoting the catalysts by metals, such as Ir and Re [4] or Ni and Re [5]. The reaction of aromatization of light petroleum fractions over Pt/AI203 is poorly known. Therefore, the investigation of this reaction over complex multi-component catalysts containing platinum is a subject of a great interest. 2. EXPERIMENTAL The non-oxidizing dehydrogenation of n-hexane was taken as a model reaction. The PtRe/AI203 and Pt-Re-Bi/AI203 catalysts with different content of metals have been prepared on a base of highly dispersed colloidal sol of Pt (d=10 A). The procedure includes the impregnation of spherically shaped alumina (S=220 m2/g) with a solution of mixed monodispersed platinum sol, ammonium perrhenate and bismuth oxychloride in the amount calculated for moisture capacity impregnation. Catalysts were dried at 100~ for 1 hour. The ratio of the metals in catalysts was varied within range: Pt = 0.15-0.35; Re =0.35-0.55; Bi = 0.025-0.1 mass percent. Electromicroscopic analysis of the samples showed the presence of finely dispersed platinum particles on the catalyst surface (davorago=10-12 A); large particles (d=20-35 ,/~) were observed very rarely. These catalysts did not give micro-diffraction images.

485

80

YIELD

vol%

A

o__

6O

40

_

20

_

t P-~

I

0

2

i

-

I

4

-~x

I

I

6

I

i

8

I

I~'Y..,6

10

I

I

12 V,h-1

Figure 1. Aromatic hydrocarbons yield vs. velocity of n-hexane feed. The reaction of non-oxidizing dehydrogenation of n-hexane was carried out in a flow quartz reactor with a stationary layer of catalyst at atmospheric pressure in a stream of high purity helium. The optimum conditions of the reaction have been found by variation of volume velocity from 2.4 to 12 h~ and temperature in the range 500-700~ Reaction products were analyzed by chromatography (Chrom-5), chromatomasspectroscopy (MX 1331) and IRspectroscopy (Specord) methods. 3. D I S C U S S I O N

Conversion of n-hexane over Pt-Re/A1203 catalysts depends on the temperature, the velocity of alkane feed, and the catalyst composition. Pt-Re/Al203 catalysts are widely used in industrial reforming. The effect of Pt/Re ratio on the n-hexane aromatization was investigated. Figure 1 shows the influence of velocity of alkane feed on the aromatic hydrocarbons yield over 0.35%Pt-0.35%Re/AI203 catalyst. Maximum yield of benzene was observed at 2.8-4 h1 velocity of n-hexane feed. The yield of toluene in this range is 20-23%. Total yield of aromatic hydrocarbons and benzene sharply declined at 5.5 h"1 velocity of n-hexane feed. Optimum volume velocity of hexane feed is in the range 3-4 h~. The n-hexane conversion under these conditions reached 87-95%. Table 1 presents the dependence of n-hexane conversion on reaction temperature and platinum content in the catalyst. Hexane conversion increased when the reaction temperature was increased from 550~ to 650~ Hexane conversion was also found to depend on catalyst composition at 550~ Increasing Pt content in the catalyst from 0.15 to 0.35% lead to an increase in n-hexane conversion from 10 to 35%. Nearly complete conversion of n-hexane was observed on the catalysts with higher Pt content.

486 Table 1 Effect of reaction temperature on the n- hexane conversion over Pt-Re/AI203 catalysts with different content of metals (V=3.6 hq) Pt-Re content % mass

Reaction T~

n-hexane conversion %

Products' yield vol % benzene toluene

Benzene selectivity%

0.15+0.55

550 600 650

10 95 93

2 56 66

24 17

C1-C5 alkanes 4 1 -

0.20+0.50

550 600 650

35 82 96

3 47 74

29 17

2 -

9 57 77

0.35+0.35

550 600 650

34 98 100

18 70 82

14 16 15

2 1 -

53 71 82

20 59 71

Benzene yield (from 2 to 18%) in non-oxidizing dehydrogenation of n-hexane also depends on the platinum content at 550~ In addition, hydrogen and C1-C5 hydrocarbons (14%) were found among the reaction products. The benzene yield increased from 66 to 82% at 650~ when platinum content was increased. C1-C5 hydrocarbons were found among the products. The process of non-oxidizing dehydrogenation is accompanied by coke formation on the catalyst surface. Its amount was controlled by CO+CO2 formation during the heat treatment of the catalyst in an air stream. The influence of 0.025-0. l%mas, bismuth additions to the Pt-Re/A1203 catalysts was studied in n-hexane dehydrogenation. Bismuth has been taken as a promoter because of its influence on the energy of the interaction of hydrocarbons and hydrogen with the catalyst surface. Heavy metals, such as cadmium, tin, zinc, bismuth, and lead, when interacting with platinum catalysts in a hydrogen atmosphere, form Pt-Bi, Pt-Pb, Pt-Cd surface clusters [ 1,6]. Measurement of hydrogen heat adsorption indicated the decrease of binding energy of active center-Had s due to a decrease of platinum free energy in the cluster composition (on the surface of the catalysts). Weak surface adsorption of reaction products has to promote higher selectivity of alkane conversion. Probable mechanisms of non-oxidizing dehydrogenation of n-hexane are as follows: Z C6H14 Z C6H13 Z C6H12

+ + +

Z Z Z

~ ~, ,

Z C6H13 Z C6H12 Z C5H9

+ + +

Z H Z H Z CH3

ZH

+

Z CH3 Z CH3

+ +

ZH ZH

, ~

+ +

Hz(gas) CH4 (gas)

ZCH3

~

2Z 2Z 2Z

+

CzH6(gas),

where Z is the active center.

487 The reaction rate depends on the limiting slow stage of hydrogen breaking from alkane and its recombination into H2 (gas). This reaction is carried out easier on Pt-Bi centers than on Pt due to the low hydrogen binding energy in Pt-Bi clusters. Reaction of n-hexane aromatization is bifunctional and includes dehydrogenation on metal centers and alkanes cracking on the acid centers. All active centers take part in the aromatization process. Benzene formation can be described by following mechanism:

CH3-(CH2)4-CH3 -2Hads ; ~

~ads

"~ C H 2 - ( C H 2 ) 4 - C H 2 a d s

' ~(:~~~ads

+3H2

It is known [2] that 92.8% benzene is formed from cyclohexane over Pt/A1203 at 300~ Table 2 presents data on n-hexane conversion over Pt-Re-Bi/Al203 catalyst. Extreme change ofhexane conversion was observed with increasing Bi content from 0.025 to 0.1% in the catalyst. Optimum conversion of hexane under the same conditions was found on catalyst with Bi content of 0.05%. For example, n-hexane conversion was observed to change within the range 82%-93%-90% with Bi percent in the catalysts was increased in the range 0.025% - 0.05 - 0.1%mas. at 600~ Benzene yield changes from 61 to 68% on these catalysts (Table 2).

Table 2 Hexane conversion over Pt-Re-Bi/A1203 catalysts (V = 3.6 h -1) Pt-Re-Bi

Reaction

N-hexane

content

T ~

conversion

% mas.

0.2+0,5+0.025

0.2+0.5+0.05

0.2+0.5+0.1

%

550 600 650

34 82 100

550 600 700

25 93 100

550 600 650 700

37 90 96 97

Products_s

C 1-C5 alkanes 4 -

2 -

-

% ....

Benzene

benzene 2 61 45

toluene 10 19 16

polycyclic traces much

selectivity

6 68 52

11 21 13

traces much

24 73 52

% 6 74 45

2

-

-

5

60 58 50

18 17 16

traces much much

67 60 52

488 Comparison of data in Tables 1 and 2 shows that the addition of 0.05% Bi to Pt-Re catalysts leads to an increase of n-hexane conversion and benzene yield from 82 to 93%, and from 47 to 68%, respectively. Maximum toluene yield decreases under these conditions from 29 to 21%. C~-C5 hydrocarbons were not found in the products, but there were polycyclic aromatic compounds (naphthalene, anthracene, etc.). Their amounts were observed to increase with the increase in reaction temperature up to 650~ The appearance of these products has been confirmed by IR-spectroscopy and chromatomassspectroscopy methods. Thus, bismuth addition to Pt-Re/Al203 leads to a significant change of the reaction mechanism of non-oxidizing dehydrogenation of n-hexane. This reaction is accompanied by the formation of complex polycyclic aromatic hydrocarbons from benzene and toluene. 4. CONCLUSION Mono-dispersed catalysts show high activity and selectivity in the process of nonoxidative dehydrogenation of n-hexane. Bismuth addition in the content of Pt-Re/AI203 catalyst reduces the temperature of conversion of n-hexane and changes the mechanism of the reaction. REFERENCES 1. 2. 3. 4. 5. 6.

J. Kanai and N. Kawutu, J. Catal, 114 (1988) 284. P. G. Smimiotis and E. Ruckenstein, J. Catal., 140 (1993) 526. H. H. Kung and M. A. Claar, Pat. USA N69284, 07 C 5/09 (1988). Inventor sert. USSRN 1608180 A1 ,C07 c15/04, 2/00 (1988). Inventor sert. USSRN 1811153 A1, C07 c15/04, 2/00(1990). D. V. Sokolsky and G.D. Zakumbaeva. Adsorption and catalysis in liquid phase over VIII Group metals, 1973.

Catalysts in PetroleumRefining and PetrochemicalIndustries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All rights reserved.

489

H I G H Q U A L I T Y G A S O L I N E SYNTHESIS BY SELECTIVE O L I G O M E R I Z A T I O N OF L I G H T OLEFINS AND SUCCESSIVE H Y D R O G E N A T I O N

T. Inui and J.-B. Kim

Division of Energy and Hydrocarbon Chemistry, Graduate School of Engineering, Kyoto University, Sakyo-ku, Kyoto 606-01, Japan. ABSTRACT The conversion reactions of light olefins were conducted using a two-stage reactor connected in series under the following mild reaction conditions: atmospheric pressure at 290~ for the first-stage reactor, and at a temperature range of 100-250~ for the second-stage reactor. An olefin-rich gasoline produced by light olefin conversion on an MFI-type H-Fe-silicate packed in the first-stage reactor was successively converted to the isoparaffin-rich gasoline on various Pt-modified catalysts packed in the second-stage reactor. The platinum supported by an incipient impregnation method on a nonmetal silicate showed a lower dispersion and lower hydrogenation activity. Although the dispersion of Pt in Pt-ion-exchanged H-ZSM-5 (Pt/H-Al-silicate) was higher than that of Pt-modified nonmetal silicate, the selectivity to isoparaffins in the gasoline range was lower owing to the hydrogenolysis and isomerization of isoparaffins to n-paraffins. Since nonmetal silicates possess a very small number of strong acid sites, the very strong catalytic activities of Pt and/or Pt/zeolite were moderated, and the undesirable reactions like hydrogenolysis proceeded minimally. Consequently, MFI-type nonmetal silicate modified with a small amount of Pt by means of adsorption treatment (almost same procedure as the ion-exchange method) showed a higher dispersion of Pt, and was the most effective in the hydrogenation of olefins in the gasoline range even at very low temperatures of 100-150~ 1. I N T R O D U C T I O N One of the possible ways to effectively utilize light olefins, which are largely produced by thermal cracking of heavy oil, is by their successive conversion to valuable products like high octane-number gasoline and aromatics for transportation fuels and as a source for engineering plastics, respectively. In particular, nonaromatic high octane-number gasoline, for which as a clean fuel recently there has been a growing demand, can be produced by oligomefization of light olefins. On the other hand, because of concerns over global warming due to CO 2 accumulation [1], the conversion of CO 2 into valuable compounds such as methanol, light olefins and gasoline have become important. We have already found that the new active catalysts and reaction methods were effective in CO 2 hydrogenation [2,3]. Since light olefins can be produced from CO 2 and hydrogen, high octane-number gasoline synthesis from the light olefins also has special significance for the effective utilization of heavy oil and/or CO 2. ZSM-5 is known to be effective for the conversion of light olefins to higher molecular weight distillate fuels and/or gasoline [4-8]. Since H-ZSM-5 zeolite has a strong hydrogen shift function, olefins in the reactants are hydrogenated to corresponding paraffins by the hydrogen evolved accompanying aromatization, resulting in a decrease in selectivity to gasoline range

490 hydrocarbons. To modify the ZSM-5 acid sites too strong and to give the silicate crystal a new catalytic function of maintaining pore structure, many researchers have studied the isomorphous substitution of various kinds of transitional metal elements for the aluminum in ZSM-5 at the stage of gel formation before crystallization [9-17]. MFI-type H-Fe-silicate [12,18,19] and H-Co-silicate [12,18] could convert light olefins completely into a high octane-number gasoline fraction with an extraordinarily high space-time yield. For example, a 95.6% propylene feed was converted at a gas hourly space velocity (GHSV) of 4500 h1 to liquid hydrocarbon products with a space-time yield as high as 8.09 kg/l.h. The product mainly consisted of iso-mono-internal olefins and a smaller fraction of aromatics. However, the olefin-rich product has the potential to act as the precursor of photooxidants. The olefins in the gasoline range produced on the H-Fe-silicate mainly branched, and hence, these can be converted into isoparaffins by simple hydrogenation using noble-metal-modified catalysts. However, undesirable reaction like hydrogenolysis should be depressed, and the higher reaction activity corresponding to the high space-time yield of gasoline fraction is required for the second-stage catalyst. In order to produce the effective synthesis of isoparaftqns in the gasoline range from light olefins, in this study, a two-stage reactor connected in series was employed. H-Fe-silicate was adopted in the first stage for the oligomerization of light olefins. In the second stage, the hydrogenation of the olefins in the gasoline range produced from the first stage was studied on various Pt-modified MFI-type nonmetal silicates, and the effect of differences in the modification methods were investigated. The properties of Pt-modified H-ZSM-5 having higher acid-site density were compared with those of nonmetal silicate. Moreover, the effect of various reaction conditions, such as flow rate and hydrogen concentration on product distribution was also investigated. 2. EXPERIMENTAL DESIGN

2.1 Catalyst Preparation Fe-silicate, with a silicon-to-iron atomic ratio of 100, and nonmetal silicate were prepared by the rapid crystallization method [20]. For Fe-silicate synthesis, the aluminum sulfate used for ZSM-5 synthesis was replaced by ferric nitrate at the stage of gel formation. For nonmetal silicate synthesis, no metal ingredient was introduced except the contaminant in the sodium silicate solution used as a silicon source. Pt-modification of nonmetal silicate by adsorption treatment, almost the same as ion-exchange treatment, was carried out with an aqueous solution ofPt(NH3)4CI 2 at 98~ for 3 h. The resultant product was washed with distilled water and dried. The platinum ammonium complex was then thermally decomposed in an air stream of 100 ml/min by heating to 350~ at a constant heating rate of 3~ and holding at that temperature for 10 min. The thermally decomposed complex was then treated in a stream of 10% 1-12-90% N 2 (50 ml/min) by heating to 400~ at a constant heating rate of 3~ and holding at that temperature for 30 min. The charged amounts of Pt in the solutions were 0.1, 0.5 and 1.0 wt% of the nonmetal silicate crystals, and they are designated as Pt(0.1), Pt(0.5) and Pt(l.0)/non-metal silicate, respectively. 1.0 wt% Pt-modified nonmetal silicate was prepared by impregnation with an aqueous solution of Pt(NH3)4CI2, and designated as 1.0 wt% Pt(imp.)/nonmetal silicate. Thermal decomposition and hydrogen reduction were carried out under the conditions described previously. The calcined crystals were made into tablets and crushed to 10-20 mesh to provide the catalysts for the reaction.

491 2.2 Catalyst Characterization

In order to confirm the phase of synthesized crystals and the presence of crystallized Pt particles in the Pt-modified catalysts, powder X-ray diffraction (XRD) analysis was carried out on a Shimadzu XD-D 1 XRD unit with nickel-filtered CuKc~ radiation at a scanning speed of l~ The Pt contents of the catalysts were measured by atomic absorption spectrophotometry (Shimadzu AA-640-01). A transmission electron microscope (TEM) (Hitachi H-800) was operated at 200 kV. XPS studies were performed by using a PerkinElmer ULVAC-PHI Model 5500 with a monochromatic MgKc~ source. The catalyst samples reduced by hydrogen were introduced into a spectrometer chamber. The carbon Is line was used as an internal energy standard, being set to 284.6 eV. The CO uptakes were measured at room temperature by the CO pulse method. A weighed amount (ca 0.2 g) of the catalyst was reduced under the same conditions as described for catalyst preparation, and then the adsorbed hydrogen was removed by helium gas flowing through the catalyst bed at that temperature for 30 min. 2.3 Reaction Method

A two-stage reactor connected in series was used for the conversion of light olefins to gasoline under atmospheric pressure. Two 1.0 g (ca 1.5 ml) portions of the catalysts (H-Fesilicate in the first-stage reactor, and Pt-modified catalysts in the second-stage reactor) were packed in quartz, tubular reactors with 8-mm inner diameters, and then they were pretreated with a nitrogen flow at 500~ for 30 min to standardize the state of the catalyst surface by removing pre-adsorbed water and other gases. The reaction temperature of the first-stage reactor was 290~ which was the most appropriate temperature for the oligomerization of light olefins [ 18,19], while the second-stage reactor was 100-200~ Non-diluted light olefins (ethylene, propylene and 1-butene) were introduced at a GHSV of 1000-5000 hl. The same mole or one-third mole hydrogen of the light olefins feed was introduced into the second-stage reactor. The products were analyzed using three gas chromatographs equipped with integrators. Columns of MS-5A, VZ- 10 and silicon-OV- 101 were used to analyze hydrogen and the whole range of hydrocarbons produced. 3. RESULTS AND DISCUSSION 3.1 Comparison of a Two-Stage Reactor and a Single Reactor for the Synthesis of Isoparaffin-Rich Gasoline from Propylene

In Figure la, propylene conversion reaction was carried out by using a single reactor in the presence of hydrogen of the same mole as the propylene feed; the catalyst was H-Fe-silicate modified with a very small amount (0.01 wt%) of Pt by adsorption treatment. On the 0.01 wt% Pt/H-Fe-silicate, the selectivity to gasoline range hydrocarbons was largely decreased because the hydrogenation of the propylene occurred prior to oligomerization, although the modified amount of Pt was smaller than the Pt(0.1)/nonmetal silicate (Figure lb). Only at 250~ was the selectivity to C5+ hydrocarbons above 60 wt%; however, the C5+ olefins were hardly hydrogenated.

492

a. Single reactor 0.01 wt% Pt/H-Fe-silicate, hydrogen/propylene = 1/1, total GHSV=2000 hl, * non-modified H-Fe-silicate.

b. Two-stage reactor 1st stage: H-Fe-silicate, temperature = 290~ GHSV = 1000 hl; 2nd stage: Pt-modified nonmetal silicate, temp. = 150~ H2 added at 1/3 mol of C3H6 fed.

Figure 1. Comparison of the performance a single reactor and a two-stage reactor for gasoline synthesis from propylene.

The iso-mono-internal olefins produced from propylene on the H-Fe-silicate were successively converted into isoparaffins and/or naphthenes on the various Pt-modified nonmetal silicates with the addition of hydrogen in the amount of 1/3 mol of propylene feed at 150~ (Figure lb). In the first stage, high selectivity to gasoline-range hydrocarbons was obtained, and the fraction of olefins was very large as shown in the top of Figure lb. Since aromatics formed less than 10 wt% in selectivity on the H-Fesilicate, the coke deposition caused by formation of fused-ring aromatics hardly proceeded. It had already been confirmed that the high activity for the oligomerization of H-Fe-silicate could be maintained for at least 100 h [12]. Although the aromatics increased slightly in second reactor, the total amount of aromatics finally produced was still below 10 wt% in selectivity. Therefore, the second-stage catalysts might be hardly deactivated by coke deposition. Time dependency of product distribution over 20 h and the change of colors of the catalysts after the reaction under the present reaction conditions were not observed. The olefin-rich gasoline produced on the H-Fe-silicate was effectively hydrogenated on the Pt(0.5) and Pt(l.0)/nonmetal silicate. The final product consisted mainly ofisoparaffins. Moreover, the high selectivity to gasoline-range hydrocarbons obtained on the H-Fe-silicate was also maintained after the olefins hydrogenated in the second stage. These results indicate that olefins are selectively hydrogenated on second-stage catalysts without hydrogenolysis and/or isomerization of branched isomers into corresponding straight-chain hydrocarbons. Only the catalyst modified with the smallest amount of Pt, i.e. Pt(0.1)/nonmetal silicate, was ineffective in the hydrogenation of olefins, especially in the gasoline range.

493

Figure 2. Synthesis of isoparaffins in the gasoline range from propylene using a two-stage reactor at a much higher flow rate. (propylene GHSV = 5000 h1 ; other reaction conditions as in Figure lb).

3.2 Effect of Flow Rate and Modification Method of Pt on Product Distribution

Figure 2 shows the results obtained at a much higher GHSV of 5000 h"l. The first-stage catalyst was same as in Figure 1, and the second-stage catalysts were Pt(0.5) and Pt(1.0)/nonmetal silicate, which were effective in hydrogenation of olefins produced from the first stage at the lower GHSV of 1000 h1 (Figure 1). In addition, the influence of the Ptmodification method was investigated by adopting the 1.0 wt% Pt(imp.)/nonmetal silicate as the second-stage catalyst. In the first stage, shown in the top of Figure 2, high selectivity to gasoline-range hydrocarbons was also obtained, although the light hydrocarbons increased slightly with the rise of GHSV from 1000 h1 to 5000 h~. The Pt(1.0)/nonmetal silicate was the most effective catalyst for the hydrogenation of gasoline-range olefins produced from propylene on H-Fesilicate even at the higher GHSV of 5000 hl, although the amount of residual olefins in the second-stage product was larger than that at the lower GHSV of 1000 h1. On the 1.0 wt% Pt(imp.)/nonmetal silicate, i.e. Pt-impregnated nonmetal silicate, the gasoline fraction decreased when compared with the first-stage product. However, on the Pt(0.5) and Pt(1.0)/nonmetal silicates, i.e. Pt-adsorption-treated nonmetal silicate, the total amount of C5+ hydrocarbons increased slightly. At the lower GHSV of 1000 hl, the ratio of paraffins to what was more than 90% in the range of C4.~7. It was confirmed by GC-MS analysis that the residual olefins with carbon numbers higher than 7 consisted mainly of cyclo-olefins. There was only a little difference in the ratio of paraffins between Pt(0.5) and Pt(1.0)/nonmetal silicate at the lower flow rate. However, it became more apparent by raising the flow rate that the hydrogenation activity of Pt(1.0)/nonmetal silicate was higher than that of Pt(0.5)/nonmetal silicate. On the other hand, straight-chain olefins slightly increased with an increase in flow rate at the first stage, and could

494 100

a ...... n...

.... i~..=.---n.. "~k

o~

""

"9

80

o

60 t,.~

m

40

ffl

~

20

< o

I

I

I

I

I

4

5

6

7

8

Carbon number

Figure 3. Effect of hydrogen concentration in the reaction gas on the ratio of paraffins produced in the range of C4--C8. (Pt(1..0)/nonmetal silicate adopted as the second-stage catalyst; (O) GHSV = 1000 h1 in the first stage; (e) GHSV = 5000 hi in the first stage; ([]) 1000 hi, same mol H 2 as C3H6 feed; (zx) 1000 hl, 1/3 mol H 2 as C3H 6 feed; (11) 5000 h~, same mol H2 as C3H 6 feed; (A) 5000 h1, 1/3 mol H 2 as C3H 6 feed). be more rapidly hydrogenated than the iso-olefins due to the difference in diffusivity in the zeolite pores [21-23]. Consequently, the ratio of straight-chain paraffins of second-stage product increased with increases in the flow rate, although the total amount of straight-chain paraffins was below 10 wt% in selectivity. Among those various Pt-modified nonmetal silicates, the Pt(1.0)/nonmetal silicate was the most effective catalyst for hydrogenation of olefins in the gasoline range. However, the final products contained significant amounts of Cs+ olefins. In addition, the olefins in the gasoline range were not sufficiently hydrogenated at the higher flow rate, although lower olefins were almost completely hydrogenated. The added amount of hydrogen (1/3 mol of the propylene feed) might be insufficient for the effective hydrogenation of all the olefins produced from the first stage. Therefore, the same mole hydrogen as the propylene feed was added the second reactor, and the results were compared with those obtained with the addition of 1/3 mol hydrogen, (Figure 3). With the addition into the second reactor of 1/3 mol hydrogen of the propylene feed, considerable C8+ olefins remained in the final products even at the lower flow rate (GHSV=1000 hl). By raising the amount of hydrogen added, however, it became possible to hydrogenate the higher olefins more effectively. Actually, at the lower flow rate (GHSV=1000 h'l), the ratio of paraffins in C 8 aliphatic hydrocarbons rose about 20% by introducing the same mole hydrogen as the propylene feed, compared with that at the addition of 1/3 mole hydrogen. Even at the higher flow rate (GHSV=5000 hl), the ratio of paraffins in the whole range of hydrocarbons increased markedly. 3.3 Comparison with Pt-Modified H-ZSM-5 Figure 4 shows the results of propylene conversion on the H-Fe-silicate (first stage) and Pt(1.0)/nonmetal silicate (second stage), at a GHSV of 1000 hl with the addition of the same

495

(a) Pt(1.0)/nonmetal silicate (2nd stage) 1st stage: as shown in Figure lb. 2nd stage: same mol added H 2 as C3H6 fed.

(b) 0.5 wt% Pt/H-ZSM-5 (2nd stage) Reaction conditions are the same as in Figure 4a.

Figure 4. Comparison of the product distributions on the Pt(1.0)/nonmetal silicate catalyst with 0.5 wt% Pt/H-ZSM-5 (Si/AI=100) catalyst in the second stage. mole hydrogen as the propylene feed. The resutls were compared with those obtained by adopting the 0.5 wt% Pt/H-ZSM-5 as the second-stage catalyst. On the Pt(1.0)/nonmetal silicate (Figure 4a), the high selectivity to isoparaffins in the gasoline range was maintained in the temperature range of 100-200~ However, on the 0.5 wt% Pt/H-ZSM-5 (Figure 4b), the selectivity to normal paraffins and light paraffins increased markedly with a rise in the reaction temperature. The isomerization from isoparaffins to normal paraffins and the increase of light paraffins by hydrogenolysis of gasoline range hydrocarbons could be due to the excessive acidity and bifunctional catalytic activity of Pt/H-ZSM-5. The isoparaffins in the final product consisted of mono-methyl paraffins mainly, and a smaller fraction of dimethyl paraffins. Isoparaffins which possess more than four alkyl groups were not detected. On the other hand, the selectivity to naphthenes, which were composed mainly of cyclopentane and cyclohexane substituted by less than three alkyl groups, was about 15 wt% totally and increased with an increase in the carbon number. There was a small amount of benzene, toluene, ethyl benzene, xylenes, and some higher aromatics having two ethyl groups and/or a propyl group which were detected. The selectivity of the total aromatics was only about 4-~8 wt%. Even by means of GC-MS analysis, indane, indene, naphthalene, and their derivatives were hardly detected. It was difficult to identify all the higher hydrocarbons; however, aromatics were differentiated from other hydrocarbons, and aromatics with carbon number 10 were usually the highest molecular-weight aromatic hydrocarbons. Moreover, the highest molecular-weight hydrocarbons in the whole products were usually not higher than carbon number 12. The finally produced gasoline fraction including C 4 hydrocarbons indicated a calculated research octane number (RON) of about 80, although selectivity to aromatics was very low.

496 Table 1. Characterization of Pt particles in nonmetal silicate Catalyst a

Pt content

XPS data for Pt 4f7/2

CO uptake

FWHMb(eV) ([.tl/~ e,nt,)

Particle size of Ptc

(rim)

obs. (wt%)

Bindingenergy (eV)

0.5(ads.)

0.35

70.9

2.61

36

8.4

1.0(ads.)

0.53

70.8

2.50

48

9.5

0.5(imp.)

0.56

71.3

2.14

19

25

1.0~imp.)

1.1

71.4

2.07

34

28

a 0.5(ads.): Pt(0.5)/non-metal silicate, 1.0(ads.): Pt(1.0)/non-metal silicate, 0.5(imp.): 0.5 wt% silicate, 1.0(imp.): 1.0 wt% Pt(imp.)/nonmetal silicate. ~t(imp.)/nonmetal Full width at half maximum. c Calculated from CO uptake. 3.4 Characterization of Pt Particles in the Nonmetal Silicate

From the TEM observations, a large number of Pt particles were shown of the Pt(1.0)/nonmetal silicate, with particle sizes (ca 5-15 nm) that were significantly smaller than those of the 1.0 wt% Pt(imp.)/nonmetal silicate (ca 20--50 nm). Moreover, XRD peaks due to crystallized Pt particles were observed only from the Pt-impregnated nonmetal silicates. The higher dispersion of Pt particles could be considered to be the reason for the higher hydrogenation activity of the Pt(1.0)/nonmetal silicate, although its experimental Pt content (ca 0.53 wt%) was less than that of the 1.0 wt% Pt(imp.)/nonmetal silicate (ca. 1.1 wt%) as shown in Table 1. The particle sizes of Pt calculated from CO uptake also agreed with TEM observation. Table 1 gives the binding energy (BE) and full width at half maximum (FWHM) of Pt 4f7/2 peaks observed by XPS measurement. The BE of the Pt 4f7/2 peaks could correspond to those for Pt metal (Pt~ The FWHM values of the Pt 4f7/2 peaks for the Pt-adsorption-treated nonmetal silicates were larger than those of the Pt-impregnated nonmetal silicates. The broad Pt 4f7/2 peaks for the Pt-adsorption-treated nonmetal silicates indicates that there might have been more than a single Pt peak contributing to each doublet component, that is, metallic Pt, cationic Pt and/or Pt particles having an strong interaction with the zeolite surface might have existed in the Pt-adsorption-treated nonmetal silicate [24,25]. On the other hand, the particle sizes of the adsorption-treated Pt in the nonmetal silicate was larger, compared with the other Pt-adsorption-treated acidic zeolites [24-26]. The number of acid sites in the nonmetal silicate, which was only due to the contaminated aluminum in the water glass used as silica source, was very small. Consequently, the amount and size of metallic Pt particles concentrated at the external surface of the nonmetal silicate was larger than that of the zeolites with a higher strong acid-site density. The highly dispersed Pt particles in the acidic zeolites could be more effective in various catalytic activities [26-28]. However, in the present studies, Pt-adsorptiontreated nonmetal silicate having a medium-sized distribution of Pt particles was suited to the hydrogenation of olefins in the gasoline range without a decrease in the yield of gasoline fraction by hydrogenolysis.

497

Figure 5. Extension to ethylene and 1-butene conversion to isoparaffins in the gasoline range on H-Fe-silicate catalyst in the first stage and successively on Pt(1.0)/nonmetal silicate catalyst in the second stage. (Reaction conditions are the same as given in Figure 4). 3.5 Extension to Ethylene and 1-Butene Conversion

Figure 5 shows the results of ethylene and 1-butene conversion. The conversion of ethylene was lower than that of propylene under the same reaction conditions Figure 5a. Under those reaction conditions, the unreacted ethylene was easily hydrogenated in the second stage, and the gasoline fraction was less than that of the propylene feed. However, it had already been confirmed [ 19] that the same conversion level as propylene could be obtained on higher metal containing H-Fe-silicate and/or at a higher reaction temperature of about 310~ although the selectivity to aromatics slightly increased. At the 1-butene fed, as shown in Figure 5b, a higher yield of gasoline-range hydrocarbons was obtained. When a mixture of ethylene, propylene and 1-butene was introduced as the reactant, the product distribution hardly changed except for a slight increase in the ethylene conversion, compared with the results of individual reactions. 4. CONCLUSIONS The olefin-rich gasoline synthesized from light olefins on the MFI-type H-Fe-silicate was effectively hydrogenated into isoparaffins in the gasoline range on the MFI-type nonmetal silicate modified with a small amount of Pt in the second-stage reactor at 100-200~ The second-stage temperature of about 100~ could be maintained by the large amount of exothermic hydrogenation reaction heat without any extra heating. The Pt-adsorption-treated nonmetal silicate was more effective in the hydrogenation of olefins in the gasoline range than the Pt-impregnated nonmetal silicate. The dispersion of Pt particles in the Pt-adsorptiontreated nonmetal silicate was lower than that in the other acidic zeolites because of the very small number of acid sites. However, the overly strong catalytic activities of Pt and/or Pt/zeolite were moderated by using nonmetal silicate possessing Pt particles of medium sizes (5-15 nm), and then undesirable reactions like hydrogenolysis and excess aromatization rarely occured. As a result, selectivity to the isoparaffins in gasoline range markedly increased and the high space time yield of gasoline fraction was constantly maintained.

498 REFERENCES

1. J. Hansen, D. Johnson, A. Lacis, S. Lebedeff, P. Lee, D. Rind and G. Russell, Science (Washington), 213 (1981) 957. 2. T. Inui, T. Takeguch, A. Kohama and K. Tanida, Energy Convers. Mgmt., 33 (1992) 513. 3. T. Inui, K. Kitagawa, T. T akeguch, T. Hagiwara and Y. Makino, Appl. Catal., 94 (1993) 31. 4. C.D. Chang and A.J. Silvestri, J. Catal., 47 (1977) 249. 5. S.A. Tabak, F.J. Krambeck and W.E. Garwood, AIChE J., 32 (1986) 1526. 6. R.J. Quann, L.A. Green, S.A~Tabak and F.J. Krambeck, Ind. Eng. Chem. Res., 27 (1988) 565. 7. J.M. Baker, S. Bessell and D. Seddon, Appl. Catal., 45 (1988) LI. 8. S. Schwarz, M. Kojima and C.T. O'Connor, Appl. Catal., 56 (1989) 263. 9. T. Inui, O. Yamase, K. Fukuda, A. Itoh, J. Tarumoto, N. Morinaga, T. Hagiwara and Y. Takegami, In Proceedings of the 8th International Congress on Catalysis, West Berlin, 1984, Vol. m, Verlag-Chemie, Berlin (1984) 569. 10.L.A. Vostrikova, V.K. Ermolaev and K.G. Ione, React. Kinet. Catal. Lett., 26 (1984) 259. 11. V.N. Romannikov, L.S. Chumachenko, V.M. Mastikhin and K.G. Ione, J. Catal., 94 (1985) 508. 12. T. Inui, J. Tarumoto, F. Okazumi and H. Matsuda, Chem. Express, 1 (1986) 49. 13. R. Szostak and T.L. Thomas, J. Catal., 100 (1986) 555. 14.G. Perego, G. Bellussi, C. Camo, M. Taramasso, F. Buonomo and A. Esposite, In Y. Murakami, A. fijima and J.W. Ward (eds), New Develop. in Zeolite Science and Technology (Studies in Surface Science and Catalysis, Vol. 28), Elsevier, Amsterdam (1986) 129. 15.G. Coudurier and J.C. Vedrine, in Y. Murakami, A. fijima and J.W. Ward (eds), New Developments in Zeolite Science and Technology (Studies in Surface Science and Catalysis, Vol. 28), Elsevier, Amsterdam (1986) 643. 16. W.F. H01derich, in Y. Murakami, A. fijima and J.W. Ward (eds), New Developments in Zeolite Science and Technology (Studies in Surface Science and Catalysis, Vol. 28), Elsevier, Amsterdam (1986) 827. 17.R.B. Borade, A.B. Halgeri and T.S.R. Prasada Rao, in Y. Murakami, A. Iijima and J.W. Ward (eds), New Developments in Zeolite Science and Technology (Studies in Surface Science and Catalysis, 28), Elsevier, Amsterdam (1986) 851. 18. T. Inui, React. Kinet. Catal. Lett. 35 (1987) 227. 19.T. Inui, F. Okazumi, J. Tarumoto, O. Yamase, H. Matsuda, H. Nagata, N. Daito and A. Miyamoto, J. Jpn. Petrol. Inst., 30 (1987) 249. 20. T. Inui, in M.L. Occelli and H.E. Robinson (eds), Zeolite Synthesis (ACS Symposium Series, Vol. 398) (1989) 479. 21.N.Y. Chen and W.E. Garwood, J. Catal. 52 (1978) 453. 22. V.Y. FriUette, W.O. Haag and R.M. Lago, J. Catal. 67 (1981) 218. 23 R.M. Dessau, J. Catal. 89 (1984) 520. 24.K. Foger and J.R. Anderson, J. Catal. 54 (1978) 318. 25. S. Fukase, H. Kumagai and T. Suzuka, Appl. Catal. 93 (1992) 35. 26. T. Inui and F. Okazumi, J. Catal. 90 (1984) 366. 27. C.W.R. Engelen, J.P. Wolthuizen and J.H.C. van Hooff, Appl. Catal. 19 (1985) 153. 28.T. Inui and A. Matsuoka, Preprints, Division of Petrol. Chem. ACS Annual Meeting, New York, (August 25-30 1991) 705.

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All fights reserved.

499

H Y D R O G E N A T I O N OF A R O M A T I C C O M P O U N D S R E L A T E D TO FUELS O V E R A HYDROGEN STORAGE ALLOY t

S. Nakagawa a, T. Ono a, S. Murata a, M. Nomura a'* , and T. Sakai b

aDepartment of Applied Chemistry, Faculty of Engineering, Osaka University, 2-1 Yamadaoka, Suita, Osaka 565, Japan bOsaka National Research Institute, 1-8-31 Midoriga-oka, Ikeda, Osaka 563, Japan ABSTRACT Hydrogenation of several aromatic compounds was conducted using the activated hydrogen storage alloy MmNi3.5Co0.7AI0.sH 4 at a relatively high than temperature (160-240~ under a nitrogen atmosphere. Selective reduction was observed with several aromatic compounds such as biphenyl and 2-phenyipyridine. The details and mechanistic aspects of this reaction were discussed by referring to the results of the reduction reaction with the alloy activated by deuterium. 1. I N T R O D U C T I O N Hydrogenation is one of the key technologies in the refinery in the petroleum industry. Many solid catalysts have been developed to attain an economical process for refining fuels. In order to get a better understanding of the catalysis of these solid catalysts, extensive studies have been conducted concerning the characterization of catalysts. The authors now consider it to be very important to understand the chemistry of hydrogenation at the molecular level for the development of more efficient catalysts for hydrogenation. Hydrogen storage alloys are expected to act as catalyst or reagent of hydrogenation [1] because these alloys can absorb hydrogen at high pressure and release it rapidly at low pressure [2,3]. In addition, the structure of these alloys has been clarified by many researchers so that these results are available at the molecular level [4-8]. The authors investigated the hydrogenation reaction of aromatic hydrocarbons over the hydrogen storage alloy MmNi3.5Co0.7A10. 8 (Mm: La, 30; Ce, 52; Pr, 5., Nd, 13 wt%) (1), and found that biphenyl (2a) was reduced efficiently and selectively to either phenylcyclohexane (3a) or dicyclohexyl (4a) by this system [9]. This is the first example, to our knowledge, of the application of such an alloy for hydrogenation of aromatic hydrocarbons. In this paper, we would like to report on the details and mechanistic aspects of the hydrogenation of compounds related to fuels such as biphenyl, 2-phenylpyridine, and quinoline. t

This work was supported by Grant-in-Aid for Scientific Resaerch No. 06750877 from the Ministry of Education, Sience and Culture, Japan. To whom correspondence should be sent.

500 2. EXPERIMENTAL DESIGN

2.1 Samples, Reagents, and Apparatus The hydrogen storage alloy (1) [10], 4-methoxybiphenyl (2b) [11] and 2-phenylpyridine, (7) [ 12] were prepared by the method reported previously. Other reagents were commercially available and were purified by recrystallization or distillation prior to use. 1H-, 2H- and 13CNMR spectra were recorded on a JEOL JNM-GSX-400 spectrometer as a CDCI 3 (for 1 H and 13C_NMR) or CHCI 3 (for 2H-NMR) solution. MS spectra were obtained with a JEOL JMSDX-303 spectrometer. 2.2 Activation of the Alloy The alloy, MmNi3.5Co0.7AI0. 8 (1), (3.5 g), was put into a 50 ml autoclave (HastelloyX), and hydrogen was introduced up to 5 MPa. The autoclave was put into an electric furnace preheated to 200~ and held at this temperature for 1 h. After being cooled to room temperature, hydrogen (5 MPa) was introduced again. This process was repeated three times. During this activation process, about 18 mmol. of hydrogen were introduced to 3.5 g of (1) to produce MmNi3.5Co0.7AI0.sH4 (I-H4). 2.3 General Procedure for Hydrogenation with the Alloy A substrate (5-10 mmol.) was added to the autoclave containing 1-H 4 (3.5 g)at-78~ under a nitrogen stream. Then, the apparatus was heated to 120-180~ under nitrogen pressure (0.5 MPa) for 3 h with shaking. After the end of the reaction, products were extracted with benzene and analyzed with GC and GC-MS. In the case where nitrogen-containing compounds were used as a substrate, the products were isolated and purified by silica gel chromatography (with benzene-ethyl acetate as the eluant) and were identified with NMR and MS analysis.

2.4 Semi-empirical MO Calculations All MO calculations were carried out on a Titan 750V workstation (KubotaPacific Computer Co.) using the semiempirical molecular orbital calculation program MOPAC (version 5.0) [13]. For each calculation, the AM1 method was used. The Titan version of this program was purchased from Simulation Technology Inc. 3. RESULTS AND DISCUSSION

3.1 Hydrogenation of Biphenyls (2a-d) with the Activated Hydrogenated Alloy Scheme 1 summarizes the various reactions investigated. Using the activated, hydrogenated alloy, I-H4, we tried to hydrogenate biphenyl (2a), one of the typical aromatic hydrocarbons. The results are summarized in Table 1. At 120~ the conversion of 2a and the yield of hydrogenated products were very low. As the reaction temperature was raised to 160180~ the yield of phenylcyclohexane (3a) reached a maximum value of 41-42%. In this case, 3a was observed as the only detectable hydrogenated product. The effect of the ratio of 2a (1-10 mmol., 0.154 to 1.54 g) to 1-H 4 (3.50 g) on product distribution was investigated (Table 1). When a small amount of 2a was used, dicyclohexyl (5a) became the major product. With an increase of the ratio of 2a to 1, the yield of 3a

501

• 2

+ • 3

4

5

a" X=H, b: X=MeO, c" X=Me, d" X=Br increased. Consequently, with 5.5 mmol. of 2a, 3a was obtained almost exclusively. These results indicate that either 3a or 5a can be prepared selectively by changing the ratio of 2a to 1. Hydrogenation of 4-substituted biphenyls (2b-d) was also investigated (Table 2). In the case of 4-methoxybiphenyl (2b), no reaction occured. With 4-methylbiphenyl (2c), conversion was lower than with 2a, and the yield of 3c was slightly higher than that of 4c. These results indicate that electron donating groups retard the reaction, that is, the hydrogen in 1 attacks the substrates nucleophilically. The reaction of 4-bromobiphenyl (2d) gave a 95% yield of 2a and further hydrogenation did not take place. This is probably due to the deterioration of 1 caused by the hydrobromic acid that evolves from 2d as the reaction proceeds. We also tried hydrogenation of naphthalene; however, the conversion of naphthalene was rather low (2-5%) even at an elevated temperature (200-240~

Table 1 Hydrogenation ofbiphenyl (2a) over the hydrogen storage alloy (l-H4) a Substrate

Temperature

(mmo~.)

(oc)

Conversion of 2a ..........................Y!e!.d...(..~ .......................... (%) b) 3a 5a 10 120 5 4 10 140 27 26 10 160 42 41 10 180 56 42 95 1 160 98 89 160 99 7 2 41 3 160 98 48 11 5 160 95 83 4 5.5 160 95 90 a The hydrogenation reaction of 2a was performed in the presence of I-H 4 in a 50 mL ~utoclave for 3 h. Determined by GLC analysis. =

502 Table 2 Hydrogenation of 2b-d with l-H4 .a Substrate

Conversion of 2

2b

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

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

3

4

21

15c

-

37 99 d

2c

2d

a The hydrogenation of 2b-d (5 mmol.) was carried out in the presence of 1-H4 at 160~ for 3h. b Determined by GLC analysis. c No detectable amount of products was observed. A mixture of trans- and cis-isomer was produced. d A 95% 2a yield was obtained.

o~

"~ = o E

100 ! '[ 801 60

~0.5 h -~lh ~ 3 h ~

,

E n o E

,L 40

~0.5 h + lh ~ 3h -e-6h

50

40 30

9

.__. "~ n"

60

20 0 ~-----0 1

2

3

N u m b e r of deuterium

4

>

20

"~ n"

10 0. 2

3 4 5 6 7 8 9 1 0 Number of deuterium

Figure 1. Distribution of deuterium in 2a recovered (left) and 3a (right). 3.2 Reaction

Mechanism

To obtain insight into the reaction mechanism, the reaction of 2a with deuterated alloy was carried out. Deuteration of the alloy 1 was undertaken by a method similar to that used for \ hydrogenation. During this activation process, about 15 mmol. of deuterium was introduced to 3.5 g of I to give MmNi3.5Co0.7AI0.8D3.5(1-D3.5). Deuteration of 2a with l-D3. 5 at 160~ for 3 h afforded 19% 3a yield, the resulting product and recovered substrate were submitted to GC-MS analysis and 2H-NMR measurement. GC-MS measurement indicated that deuterium was introduced not only into the 3a, but also into the recovered 2a (Figure 1).These results imply that dehydrogenation of partially hydrogenated products occurs. In the case of 2a, the relative amount of deuterated biphenyls was increased by lengthening the reaction time to 3 h. A further elongation of the reaction time to 6 h, however, did not change the distribution of deuterated biphenyls, although the conversion to 3a increased. In the case of 3a, the hexadeuterated product is the major product. The distribution of deuterium in the product was found to ve independent of the reaction time.

503

mil o-H

"1'".~!7.8 7....7.61....7!5"'I.~''"77:3.... 7.21'-~ Chemical shift (ppm) Figure 2.2H-NMR spectrum of recovered 2a from the reaction of 2a with l-D3. 5.

R

R

major R

6

R=Ph or cyclohexyl R minor

0

Scheme 1. Ph

"6 50" E o

48

cO

46

E ~0

44

0

*" "1"

Ph

Ph

42

Figure 3. Energy diagram of the first stage of the hydrogenation of 2a.

Ph

J

0

504

#'--"PhZ :7

'

'

'

I

. . . .

I

. . . .

I

'

'

2.5 2.0 1.5 Chemical shift (ppm) Figure 4.2H-NMR spectrum of the 3a produced by the reaction of 6 with l-D3. 5.

2H-NMR spectra for 2a and 3a produced in the reaction of 2a with I-D3. 5 are shown in Figure 2. The relative amounts of deuterium in o-, m-, and p-position of 2a were 21, 48, and 3 1 % , respectively. The amount of deuterium in o-position was less than that in the other two positions, suggesting that hydrogenation of 2a took place mainly at the 2,3- and 3,4-positions of the benzene ring (Scheme 1). To calculate a heat of formation of these hydrogenated products, semi-empirical molecular orbital calculation was carried out using the MOPAC program [ 13]. The results are illustrated in Figure 3. They also indicate that the formation of both 2,3- and 3,4-di-hydrogenated biphenyls is preferable to that of 1,2-di-hydrogenated biphenyl. Hydrogenation of 1-phenylcyclohexene (6) was carried out in order to get detailed information about the manner in which hydrogen is added into aromatic rings. A 2H-NMR spectrogram of the product is shown in Figure 4. From this spectrogram, it was found that the main product was (cis-l,2-di-deuterio-l-phenylcyclohexane, suggesting that the addition of hydrogen might take place in a syn-1,2-addition. These findings also agree well with the results reported by Imamoto et al [3].

3.3 Hydrogenation of Nitrogen-Containing Compounds Hydrogenation of 2-phenylpyridine (7) (5 mmol.) over activated 1-H4 (3.5 g) was also carried out in a 50ml autoclave at several temperatures for 3-6 h. An elevated temperature, such as 240~ was needed to obtain higher conversion rates. The reaction at 240~ for 6 h yielded 2-cyclo-hexylpyridine (8) (53 %) as a major product along with substantial 2phenylpiperidine (9) (23 %). It was reported that hydrogenation of 7 with H 2 over platinum or nickel catalysts mainly yields 9. These results indicate that the catalytic nature of 1 is different from that of nickel powder [ 14,15].

505

7

H

8

9

H 10

11

12

Hydrogenation of quinoline (10) (5 mmol.) was also carried out at 240~ for 3 h in the presence of 3.5 g of I-H 4. The reactivity of 10 was found to be higher than that of naphthalene or 7 and a 67% yield of 1,2,3,4-tetrahydroquinoline (11) was obtained along with 5,6,7,8-tetrahydroquinoline (12) (18%). Denitrogenation is believed to proceed by saturation of the nitrogen-containing ring followed by the fission of C-N bond. Thus, these results are very interesting because this reaction shows preferential hydrogenation of the nitrogencontaining ring. Therefore, the activated alloy can be used as a selective catalyst for this purpose: the formation of perhydroquinoline is harmful for the economical process of refining. 4. CONCLUSIONS The results obtained in this paper are summarized below. 1. Hydrogenation of biphenyl was found to proceed at 160~ over the hydrogen storage alloy, and either phenylcyclohexane or biphenyl could be selectively prepared by changing the ratio of the substrate to the alloy 2. The results of the hydrogenation of 4-substituted biphenyls suggested that hydrogen absorbed in the alloy attacks the aromatic rings nucleophilically. 3. The results of the reactions ofbiphenyl and 1-phenylcyclohexene with the deuterated alloy indicated that hydrogen addition at the 2,3- and 3,4-positions favourably occured in the first step. This reaction appeared to include dehydrogenation of the di-hydrogenated products and addition of hydrogen mainly proceeded in a ~yn-manner. 4. Hydrogenation of 2-phenylpyridine or quinoline over the hydrogen storage alloy could be carried out at an elevated temperature to give 2-cyclohexylpyridine or 1,2,3,4tetrahydroquinoline, respectively. In the case of 2-phenylpyridine, the product distribution was observed to be different from the hydrogenation over nickel or platinum catalysts.

506 REFERENCES

1. J. Barrault and D. Duprez, J. Less-Common Met., 89 (1983) 537. 2. K. Soga, H. Imamura, and S. lkeda, Chem. Lett. (1976)1387;NipponKagakuKaishi, 1977, 1299; 1978, 923; J. Phys. Chem., 81 (1977) 1762; K. Soga, Y. Sano, H. Imamura, M. Sato, and S. lkeda, Nippon Kagaku Kaishi (1978) 930. 3. T. Imamoto, T. Mita, and M. Yokoyama, J. Org. Chem., 52 (1987) 5695. 4. H.C. Siegmann, L. Schlapbach, and C. R. Brundle, Phys. Rev. Lett., 40 (1978) 972. 5. W. E. Wallace, R. F. Karllcek, Jr., and H. Imamura, J. Phys. Chem., 83 (1979) 1708. 6. E. D. Snijder, G. F. Versteeg, and W. P. van Swaij, AIChE J., 39 (1983) 1444. 7. J. J. Reilly and J.R. Johnson, J. Less-Common Met., 104 (1984) 175. 8. M. Miyamoto, K. Yamaji, and Y. Nakata, J. Less-Common Met., 89 (1983) 111. 9. S. I. Nakagawa, S. Murata, and M. Nomura, Chem. Lett. (1994) 431. 10. T. Sakai, T. Hazama, H. Miyamura, N. Kuriyama, A. Kato, and H. Ishikawa, J. Less Common Met., 172/174 (1991) 1175. 11. Y. Kiso, K. Yamamoto, K. Tamao, and M. Kumada, J. Am. Chem. Soc., 94 (1972) 4374. 12. J. C. W. Evans and C. F. H. Allen, Org. Synth. Coll., 11 (1966) 517. 13.M.J.S. Dewar, E.G. Zoebisch, E. F. Healy, and J. J. P. Stewart, J. Am. Chem. Soc., 107 (1985) 3902. 14. H. Adkins, L. F. Kuick, M. Farlow, and B. Wojcik, J. Am. Chem. Soc., 56 (1934) 2425. 15. F. W. Vierhapper and E. L. Eliel, J. Org. Chem., 40 (1975) 2729.

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All rights reserved.

A THEORETICAL STUDY OF ETHYLENE O R G A N O M E T A L L I C N I C K E L CATALYSTS

507

OLIGOMERIZATION

BY

L. Fan a, A. Krzywicki a, A. Somogyvari a and T. Ziegler b

aNovacor Research & Technology Corporation, 2928-16 Street, N.E. Calgary, Alberta, Canada T2E 7K7 bDepartment of Chemistry, The University of Calgary, Calgary, Alberta, Canada T2N 1N4 ABSTRACT The mechanism of ethylene oligomerization catalysed by (acac)NiH has been studied by Density Functional Theory. The energy barrier for ethylene insertion (chain propagation) is calculated to be 5.7 kcal/mol. Chain termination by 13-hydrogen elimination is endothermic by 44.7 kcal/mol. Three alternative termination reaction pathways involving 13-hydrogen transfer to monomer have also been investigated. The lowest energy pathway reflects a two-step reaction via an intermediate of nickel hydride. The energy barrier leading to the intermediate from a n-complex is 6.5 kcal/mol, and the barrier leading to the termination product from the intermediate is 8.5 kcal/mol. The [3-hydrogen transfer reaction is thus suggested to be a possible cause for chain termination. 1. I N T R O D U C T I O N Dimerization and oligomerization of ethylene to 1-butene; and higher a-olefins are processes of considerable industrial importance. A variety of catalysts; has been reported to be active in producing ct-olefins by oligomerizing ethylene. A number of nickel-based catalysts has been developed by Keim and co-workers[ 1], and by others[2]. All of those catalysts contain a bidentate chelating ligand, X-Y, where X, Y = O, S, N, P. The reaction mechanism of the oligomerization has not yet been well established. It is suggested[ 1] that the actual catalyst is a nickel hydride and the oligomerization processes follow a mechanism shown by Scheme 1.

(acac)NiH + C2H 4

~'~ (acac)NiC2H 5 p-

+ C2H4,~ + C2H4~.~ ,,--- (acac)Ni~, '4H9 f Chain growth

- C4H 8 temaination Scheme 1. Mechanism of the Catalytic Cycle. In the initial step, an ethylene approaches the coordinatively unsaturated nickel hydride to form a rt-complex, followed by a four-center transition state that leads to the insertion of the

508 ethylene into the Ni-H bond. The vacant coordination site is released when the insertion completes, and similar reactions continue until the oligomer chain is eliminated. In a previous study [3], we have determined the structures of some of the important intermediates involved in Scheme 1 by Density Functional Theory (DFT). We have shown that the nickel hydride, (acac)NiH, (acetyl acetonate was modelled by 1,3-propanedione) is very active in the presence of ethylene and leads to (acac)NiC2H 5 with exothermicity of 44.7 kcal/mol. On the other hand, the butene elimination process shown by Eq. (1) is energetically unfavorable with the (acac)NiC4H 9 ~

(acac)NiH + Call 8 - 44.7 kcal/mol

(1)

endothermicity of 44.7 kcal/mol, while the process of Eq. (2) is essentially thermoneutral. (acac)NiC4H 9 + C2H4 -~

(acac)NiC2H 5 + C4H8+ 0 kcal/mol

(2)

We have therefore proposed a modified mechanism of the catalytic cycle, Scheme 2, in which-the nickel hydride is bypassed and the catalyst is considered to be (acac)NiC2H 5. The optimized structure of (1,3-propanedionato)NiC2H 5 is shown as structure 1. The question remaining was the detailed mechanism of Eq. (2) and the reaction energy of the chain elimination processes compared to the chain growing reactions.

(acac)NiH + C2H4

~.- (acac )NiC2H5

l, ,

+ c:H4 ,--- (acac)NiC4H9-'-a---~" Chain growth

+ C2H4 and -C4H8

Scheme 2. Modified Mechanism. We shall discuss the chain termination reaction, Eq. (2), in the present study based on Density Functional Theory at a high level with non-local gradient corrections. A variety of reaction pathways and the associated transition state structures as well as the reaction barriers will be analyzed. 2. COMPUTATIONAL DETAILS Density Functional theory [4] (DFT) has been widely recognized as a powerful alternative computational method to traditional ab initio schemes, particularly in studies of transition metal complexes where large size of basis set and an explicit treatment of electron correlation are required. The local spin density approximation [5] (LDA) is the most frequently applied approach within the families of approximate DFT schemes. It has been used extensively in studies on solids and molecules. Most properties obtained by the LDA scheme are in better agreement with experiments [4a] than data estimated by ab initio calculations at the HartreeFock level. However, bond energies are usually overestimated by LDA. Thus, gradient or nonlocal corrections [6] have been introduced to rectify the shortcomings in the LDA. The non-

509 local corrections can be introduced as a perturbation or incorporated in a fully variational calculation. In the perturbative approach, the non-local energy functional is evaluated based on the LDA electronic density while in the variational approach the electronic density itself is determined by optimizing the gradient corrected energy. The variational procedure is computationally more demanding than the perturbative approach. We have shown in previous studies [7] that the density change induced by non-local corrections is minor and the two approaches lead to similar results for most of the molecular properties that have been studied. In the present investigation all calculations were carried out by the ADF program due to Baerends [8] et al. and the molecular geometries were optimized based on the LDAin the parameterization due to Vosko et al [9]. Single-point energy evaluations were then carried out with Becke's non-local exchange correction [6b] and Perdew's non-local correlation correction [6c]. The basis set [ 10] used for the 3s, 3p, 3d and 4s valence shells on nickel was of triple-~ quality and augmented by three 4p Slater-type-orbitals (STO). A double-~, basis set was applied for the 2s and 2p shells of oxygen and carbon as well as the 1s shell of hydrogen. An additional 3d STO was added to oxygen and carbon whereas hydrogen was given a single 2p STO. All inner shell orbitals were kept frozen in the variational calculations [8]. A set of auxiliary [11] s, p, d, f, and g type of STOs centered on each atoms was used to fit the electronic density. The numerical integrations were carried out according to the scheme [12] proposed by Boerrigter et al. 3. RESULTS AND DISCUSSION We have studied the energetics of the ethylene insertion reactions in a previous paper [3]. The energy change from the 7t-complex 2a to the direct product of insertion 2b was found to be 10.7 kcal/mol, indicating the chain growing reactions are thermodynamically favorable. However, the kinetic feature of the insertion, i.e. the transition state structure and the reaction energy barrier have not been discussed.

i-'-O 1.50

1. Catalyst

2a. H-complex

In the present study, the transition state structure for the ethylene insertion reaction 2a---~2b has been fully optimized by the standard algorithm of transition state optimization [ 13]. Most of the important geometric parameters are indicated in 2r The Ni-C(ethyl) bond is elongated from 1.98A in the reactant 2a to 2.07A in 2c, and a partial C-C bond of 1.97A is

510 formed in 2e. Thus, the structure 2r is a typical four-center transition state that is similar those found in ethylene polymerization by metallocene catalysts [14]. The transition state 2e is 5.7 kcal/mol and 16.4 kcal/mol higher in energy than the reactant 2a and the product 2b, respectively.

x. 2b. Insertion Product

2c. Transition State for Insertion

As mentioned earlier, chain termination via 13-hydrogen elimination is energetically demanding. We have proposed [3] an alternative approach to explain the chain termination process, i.e. by hydrogen transfer from the oligomer to monomer. Eq. (2) models a simplified process by replacing the oligomer chain with an ethyl group. The advantage of such a simplification is that the transition state should adopt symmetric structures if the reaction takes place by an elementary step.

+

H

Scheme 3. Ethylene Approaches on the Chelate Plane Scheme 3 illustrates a possible reaction pathway by which the incoming ethylene molecule attacks the catalyst through the chelate plane and the transition state is of C2v symmetry with R=H. Structure 3a has thus been optimized with a C2v constraint. It is clear from 3a that the strong 13-agostic interaction in I does not exist in 3a with a remarkably long Ni-H distance of 3.11A. The structure 3a is 26.8 kcal/mol higher in energy than the reactants C2H4 + 1. Scheme 3 is therefore unlikely a realistic route for chain termination.

511

3.11 95.1*

! I

3a. Transition State for Scheme 3

~

'

~

~

~

\ 1.49 ;~

3b. Transition State for Scheme 4

Scheme 4 shows another approach by which an ethylene attacks the catalyst from the top of the chelate plane. The transition state of C~, symmetry is given as 3b. The reaction barrier of Scheme 4 is 18.4 kcal/mol, which is substantially lower than that of Scheme 3 while still much higher than the insertion barrier of 5.7 kcal/mol.

+ i

Scheme 4. Ethylene Approaches on Top of the Chelate Plane Reducing the constraint from C2v to Cs, an alternative approach depicted by Scheme 5 leads to a transition state with the energy barrier of 17.3 kcal/mol, which has been reported previously [3].

Scheme 5. An Alternative Approach of Ethylene Attacking on Top of the Chelate Plane We have found another structure of Cs symmetry which is shown as 4a. Structure 4a is clearly a five-coordinated nickel hydride with the Ni-H distance of 1.43A. Two ethylene

512 molecules are coordinated to the nickel center by their rt-orbitals. A similar nickel hydride has been identified by experiments [15]. The energy of 4a is 1.8 kcal/mol lower than C2H4 + 1, and thus it should exist as an intermediate which supports the experimental evidence. The transition state structure has been determined by decreasing the C-H distance, which is 2.40 A in 4a, step by step while optimizing the other geometric parameters to obtain an initial structure and followed by non-constraint optimization. A well converged transition state structure is shown by 4b. The forming C-H bond in 4b is 1.75 A and the Ni-H distance is slightly elongated to 1.44 A. Further reduction of the C-C distance led to a x-complex 4e. Structure 4e is not a very stable n-complex since it is only 0.2 kcal/mol lower in energy than C2H4 + 1. The hydrogen transfer reaction is essentially a two step process as shown by Scheme 6. In the first step a loosely bonded r~-complex 4e is formed and evolved to an intermediate 4a through a transition state 4b. The hydrogen transfer is then completed in the second step by overcoming a related transition state that leads to another n-complex.

1.26

o~ t

I

r

1.441'~ ~ 397.4,~ 4 ~ ''v

t~.

4a. Intermediate

4b. Transition State for Scheme 6

4c. H-complex for Hydrogen Transfer The energy barrier for the first step is only 6.5 kcal/mol, and slightly higher for the second step with 8.3 kcal/mol. The energy profile of Scheme 6 is compared with that of the insertion reaction by Figure 1.

513

v

re-complex 4e transition state 4b

Q

intermediate 4a

transition state symmetric to 4b

n-complex symmetric to 4c

Scheme 6. Proposed Mechanism for Chain Termination 4. C O N C L U S I O N Compared to the other reaction pathways, Scheme 6 illustrates the most plausible mechanism for chain termination. The reaction barrier of 8.3 kcal/mol is higher than the insertion barrier of 5.7 kcal/mol. Keim and co-workers have successfully trapped the nickel hydride as evidence to support their catalytic mechanism in which the nickel hydride is considered as the active catalyst [15]. We have found in the present study that the nickel hydride actually exists as an intermediate of the chain termination process. The premise for Scheme 6 to be practically competitive to the ethylene insertion reactions is the formation of the n-complex 4c. Based on our calculations, 4e is only a shallow minimum with the stabilization energy of 0.2 kcal/mol. Higher ethylene concentration is thus expected to facilitate the formation of the 7t-complex and hence to increase the possibility of chain termination in order to generate dimers and trimers. 10.0 ~o

transition state 4b .........

5.0

terminationprofile

C2H4 o.o 0

,-gff, ---~--." 5.7 9 , intermediate4a ,,'..___~_, ! x-complex 2a , ",,

,

O

",

= -5.0 ua O > .,..4

4c

9

-~ -10.0

'; insertion profile

-15.0

2b

'I 9

-20.0.

insertion product 2c

,

Figure 1. Energy Profiles for Insertion and Termination

514 REFERENCES

1. W. Keim, Angew. Chem. Int. Ed. Engl., 29 (1990) 235. 2. (a) S. J. Brown and A. F. Masters, J. Organomet. Chem., 367 (1989) 371. (b) R. Abcywickrema, M. A. Bennett, K. J. Cavell, M. Kony, A. F. Masters, and A. G. Webb, J. Chem. Soc. Dalton (1993) 59. 3. L. Fan, A.Krzywicki, A. Somogyvari and T. Ziegler, Inorg. Chem., 33 (1994) 5287. 4. (a) T. Ziegler, Chem. Rev., 91 (1991) 651. (b) R. G. Parr and W. Yang, Density Functional Theory of Atoms and Molecules, Oxford University Press, New York, 1989. 5. J. P. Dahl and J. Avery (eds.), Local Density Approximation in Quantum Chemistry and Solid State Physics, Plenum, New York, 1984. 6. (a) C. D. Hu and D. C. Langreth, Phys. Rev., B33 (1986) 943. (b) A. D. Becke, Phys. Rev., A38 (1988) 3098. (c) J. P. Perdew, Phys. Rev., B33 (1986)8822. Also seethe erratum: Phys. Rev., B34 (1986) 7046. (d) L. C. Wilson and M. Levy, Phys. Rev., B41 (1990) 12930. 7. (a) L. Fan and T. Ziegler, J. Chem. Phys., 94 (1991) 6057. (b) L. Fan and T. Ziegler, J. Chem. Phys., 95 (1991) 7401. (c) L. Fan and T. Ziegler, J. Chem. Phys., 96 (1992) 9005. (d) L. Fan and T. Ziegler, J. Phys. Chem., 96 (1992) 6937. 8. E. J. Baerends, D. E. Ellis and P. Ros, Chem. Phys., 2 (1973) 41. 9. S.H. Vosko, L. Wilk and M. Nusair, Can. J. Phys., 58 (1980) 1200. 10. (a) G. J. Snijders, E. J. Baerends and P. Vernooijs, At. Nucl. Data Tables, 26 (1982) 483. (b) P. Vernooijs, G. J. Snijgers and E. J. Baerends, Slater Type Basis Functions for the Whole Periodic System (Internal Report), Free University of Amsterdam, Amsterdam, 1981. 11. J. Krijn and E. J. Baerends, Fit Functions in the HFS-Method (Internal Report), Free University, Amsterdam, 1984. 12. P. M. Boerrigter, G. te Velde and E. J. Baerends, Int. J. Quantum Chem., 33 (1988) 87. 13. (a) J. Baker, J. Comput. Chem., 7 (1986) 385. (b) L. Fan and T. Ziegler, J. Chem. Phys., 92 (1990) 3645. 14. (a) L. Fan, D. Harrison, L. Deng, T. Woo, D. Swerehone and T. Ziegler, Can. J. Chem., in press. (b) T. Woo, L. Fan and T. Ziegler, Organometallics, 13 (1994) 432. (c) T. Woo, L. Fan and T. Ziegler, Organometallics, 13 (1994) 2252. 15. U. Muller, K. Keim, C. Kruger and P. Betz, Angew. Chem. Int. Ed. Engl., 28 (1989) 1011.

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All rights reserved.

515

IFP-SABIC PROCESS FOR THE SELECTIVE ETHYLENE DIMERIZATION TO BUTENE-1 Fahad A. AI-Sherehy

Development Department, SABIC Industrial Complex for Research and Development, P.O. Box 42503, Riyadh 11551, Saudi Arabia ABSTRACT This paper outlines the various routes for the manufacturing of butene-1 used as a comonomer for the production of polyethylene (LLDPE, HDPE). Analysis of the characteristics of each route is provided, together with a comparison between the different processes. The preferred route for the manufacturing of butene-1 employing IFP-SABIC Alphabutol technology is highlighted. The advantages this technology offers over possible competing processes are identified. The first commercial plant at a wholly owned SABIC affiliate (Petrokemya) employing the IFP-SABIC technology has now been in operation since 1987. The main features of IFP-SABIC Alphabutol technology are that it is very selective to butene-1 production and offers a simple process sequence resulting in lower capital and operating cost. The technology over the years has improved considerably in dealing with polymer formation and deposition which is detrimental to the plant performance, changes incorporated in the original design together with accumulated experience has resulted in significantly less polymer deposition. 1. I N T R O D U C T I O N Butene-1 is the first member of the linear alpha olefins (LAO)family. It is a basic petrochemical and it can be converted to products such as polybutene-1 and butylene oxide. However, the main use of butene-1 is as a co-monomer with ethylene for the polyethylene production (LLDPE and HDPE)which accounts for approximately 80% of the butene-1 market

[1]. Butene-1 can be produced by a variety of methods, including the following [2]: 1. Refinery Operations (about 15% of the generated effluents from the fluid catalytic cracker in the refineries is butene-1) [ 1]. 2 Steam Cracking of C4 hydrocarbons 3 Butane dehydrogenation 4. Co-product from alpha olefin manufacturing 5 Ethylene dimerization 6 Butyl alcohol dehydration 7 Pyrolysis of butyl acetate and butyl chloride

516 Table 1 [ 1] Global Butene-1 Production Capacity, 1993 (thousand metric tons)

Source of Butene-1 Refinery and steam cracking Ethylene Oligomerization Ethylene dimerization TOTAL

World Production 300 190 105 595

Percentage 50 32 18 100

Only the first five methods are of direct interest to industry. About 50% ofbutene-1 is produced from refinery and steam cracking operations. The remainder is obtained as alpha olefin co-product. The United States and European countries are the major sources for butene1 in the world market. The breakdown of the 1993 global butene-1 production capacities is presented in Table 1. Selective dimerization or oligomerization of ethylene have been considered as economic routes for butene-1 production. However, due to the wide range of products associated with ethylene oligomerization and market limitation for some of these products, ethylene dimerization to butene-1 appears to be a more attractive option for the butene-1 production. About 18% of the butene-1 is produced by ethylene dimerization as shown in Table 1. The objective of this paper is to review the currently available processes for ethylene dimerization to butene-1, and to highlight the development of the IFP-SABIC process as the only commercially proven technology. 2. ETHYLENE DIMERIZATION PROCESSES Development in the catalytic dimerization of ethylene into butene-1 was pioneered in 1952 by the studies of Ziegler which were originally aimed at producing higher-chain polymers via the growth reaction of the organoalurninum compounds (multiple insertion of ethylene into the A1-C bonds). One particular batch gave the opposite result, namely the quantitative formation of butene-1 from ethylene [ 1]. This field became, later, the subject of interest for research in industry and academia. Normally, the efficiency of a dimerization process is determined by: (i) the selectivity to butene-1, (ii) the yield of butene-1 per unit weight of the catalyst, and (iii) the required process equipment [3]. So far, the only commercial process is the IFP-SABIC Alphabutol process, which is an indication of its technical advantages and economic potential. There are other processes which have not as yet reached commercial stage. These processes are offered by Phillips, MIT and Dow. The main features and characteristics of these dimerization processes are as follows:

Phillips Process The Phillips dimerization process [4,5] catalytically converts ethylene to butene-1 utilizing a nickel based catalyst system consisting of ethyl aluminium dichloride and bis(tri-nbutyl phosphine) nickel dichloride prepared in dry n-pentane. The process consists of three steps, a reaction step and two quench steps. In the reactor section, ethylene is fed to the reactor where it comes in contact with a mixture of diluent butenes and the two catalyst

517 components which are circulated through the reactor by an external pump. The circulating mixture is passed through a cooler before it enters the reactor to remove the heat of reaction. Fresh catalyst components are continuously pumped into the reaction system where the Al/Ni molar ratio is controlled at a value in the range of 0.7 to 1.0 during the initial start-up period and at a higher value (3-5) during the operating period. Typical reaction conditions are 48 ~ and 13.7 atm. and average residence time of the reactants in the circulating loop being 30 minutes. The liquid product from the reaction section containing unreacted ethylene, product butenes and catalyst is sent to the catalyst quench section. The catalyst is deactivated by contacting with 2 wt. percent acetic acid and separated from the butene product in an extractor. The catalyst-free butene stream of the extractor effluent proceeds to a neutralization vessel where it is contacted with dilute caustic soda solution. The butene stream leaving the neutralization vessel is filtered, distilled and recovered as product.

Massachusetts Institute of Technology (MIT) Process A conceptual process for ethylene dimerization in the presence of tantalum or niobium based catalysts has been developed by MIT researchers [6-8]. The technology is based on a metal hydride-based homogeneous catalyst that selectively dimerizes ethylene to butene-1. The particular catalyst is neopentylidene complex of tantalum or niobium. The preparation of the homogeneous catalyst is rather a complex process; the tantalum complex is prepared by reacting tri-neopentyl tantalum dichloride, Ta(CH2CMe3)3Cl2 and neopentyl lithium LiCH2CMe3 in octane solvent to yield thermally stable neopentylidene tantalum catalyst in quantitative yield. Fresh and recycled ethylene plus octane solvent are fed into the dimerization reactor operating at 100 atm. and 80~ The dimerization takes place in a homogeneous liquid phase and proceeds rapidly at the rate of one mole of butene-1 formed per min. per mole of the catalyst. The heat of reaction is removed by excess ethylene leaving the reactor as an overhead stream. Ethylene is cooled and recycled to the reactor along with fresh ethylene feed. The reaction is conducted under an oxygen-free, anhydrous environment to prevent deactivation of the catalyst. The reaction is maintained at 80~ in order to minimize the potential isomerization of butene-1 to butene-2. Ethylene conversion is about 20 percent per pass and the assumed product yield from this process are 95 percent butene-1 and 5 percent butene-2. Effluents leaving the dimerization reactor are sent to a liquid-vapor separator, where ethylene is separated from butenes mixture and recycled to the reactor. The separator bottoms proceeds to the solvent recovery column which produces a butenes overhead stream and bottoms solvent stream (containing the catalyst in solution). The solvent stream is recycled back to the reactor while the butenes are sent to an extractive distillation column. A high purity butene-1 (99.9%) is produced from the extractive distillation column.

Dow Process [9,10]. A mixture of ethylene and butene-1 is prepared by the dimerization of ethylene in the presence of organic aluminium compound A1R3in a boiling solvent reaction zone. High purity ethylene is fed into the dimerization reactor operating at 27 atm. The dimerization takes place

518 in a homogeneous liquid phase of A1Et3 and tetradecane solvent. Ethylene flows through a gaseous diffuser to disperse the ethylene gas for better contact. Ethylene to AIEt3 weight ratio is in the range of 4,000 to 8,000. A prepared solution of 0.4 wt percent A1Et3 in tetradecane is added to the reactor and maintained at a specific level. A conventional heating device is used to heat the liquid-gas mixture to 277 ~ At the upper end of the vertical dimerization reactor, a conventional contact device such as mesh packing is used. The reactor effluent proceeds to a cooler where dissolved ethylene is separated from the butenes stream. A reflux drum is provided for the condensation of solvent vapor and the liquid is recycled back to the reactor. The mixture of ethylene and butene-1 proceeds from the reflux drum into the outlet line. After 5 hrs of reaction time, ethylene conversion reached 25.7 percent and the product distribution was mainly butene-1 at 95.5 percent selectivity and small amounts of hexenes and other oligomers but without any polymer formation. The unit ratio for the grams ofbutene-1 produced per gram of triethylaluminum was about 159. 3. I F P - SABIC ALPHABUTOL TECHNOLOGY IFP developed a process for the selective dimerization of ethylene to butene-1 over a homogeneous titanium based catalytic system. The world largest operating butene-1 plant using this technology has been on stream since 1987 at a wholly owned SABIC affiliate (Petrokemya) in Jubail, Saudi Arabia with a name plate capacity of 50,000 metric tons/year. Since 1987, extensive process modification, contributed by SABIC and IFP, has enabled the smooth running of this first and world's largest plant. As a result of this collaboration, the two parties now jointly own this technology, referred to as IFP-SABIC Alphabutol technology for butene-1 production. Today, sixteen butene-1 plants using this technology have been licensed throughout the world with five of them already gone on-stream.

Process Chemistry The IFP-SABIC Alphabutol process utilizes a proprietary homogeneous titanium based catalyst which demonstrates high dimerization activity coupled with excellent selectivity to butene-1 at moderate pressures and temperatures. This performance is influenced by the catalyst composition and reaction parameters. The catalytic ethylene dimerization to butene-1 is widely regarded as a degenerate ethylene polymerization reaction and therefore the formation of higher molecular weight byproducts (oligomer and/or polymer) is expected [1]. However, in IFP-SABIC process, the judicious choice of the titanium based catalyst [ 11-13] (Ti(OR)4 compound activated by an alkyl aluminum A1R3) and the reaction conditions [14] (20-30 atm. pressure and 50-60 ~ temperature) lead to the selective generation ofbutene-1 (93 % wt.) at a conversion of(8085%). Small amounts of by-products such as hexenes, cis/trans butene-2 and butane are formed. Typical analysis of the butene-I produced by IFP-SABIC Alphabutol process is presented in Table 2. Catalyst 2 CH2=CH2 Ethylene

~

CH3CH2CH=CH2 + Butene-1

By-products ..........

(1)

519 TABLE 2. Typical Analysis oflFP-SABIC Alphabutol Butene-1 [ 15]

Composition

Concentration Limits

Butene- 1 Butenes and butanes Ethane Ethylene C60lefins Dienes, Acetylenics CO, CO2, Oz, 1-120

99.5 wt % min 0.3 wt % max 0.05 wt % max 0.15 wt % max 50 wt ppm max 5 wt ppm max 5 wt ppm max

A special feature of the IFP-SABIC Alphabutol technology is the inhibition of the catalyst toward polymer formation during the production of butene-1. Generally, polymer formation results with the use of Ziegler-type catalysts based on titanium, these catalysts are known for their ability to polymerize ethylene to high molecular weight materials. In the IFPSABIC process such polymerization reaction is inhibited by adding a modifying agent to the catalyst formula.

Process Description A simplified representation of the process scheme is shown in Fig. 1. Three main sections are involved in this process, Reaction Section, Catalyst Removal Section, and Distillation Section, and are described as follows:

Reaction Section The reactor is operating in liquid phase at bubble point conditions. Fresh and recycled ethylene are fed to the liquid phase of the reactor through a gas distributor. The homogeneous catalyst is continuously fed to the reactor section. The dimerization reaction is carried out at about (50-600~ and 20-30 atm.)[14] with a reaction time of about (4-6)hrs. The homogeneous catalytic reaction proceeds at an ethylene conversion of about 80-85 percent per pass with a selectivity to butene-1 approaching 93%. The exothermic heat of reaction is removed by means of external pump-around loop equipped with a cooler. The reactor effluent is withdrawn from the reactor as a liquid containing the catalyst. Catalyst Removal Section In the catalyst removal section, the active catalyst in the reactor effluent is deactivated by adding a catalyst deactivating agent. The catalyst is then separated from the reactor effluent by means of vaporization where the liquid withdrawn from the reactor is vaporized and the residue contains the spent catalyst and a small amount of hydrocarbons. Distillation Section At the distillation section, catalyst-free hydrocarbon portion of the reactor effluent proceeds to the first column where unconverted ethylene is recovered as a distillate and recycled to the dimerization reactor at an adequate pressure. The bottoms from the first column are fed to the butene-1 purification column where co-monomer grade butene-1 (99.7%) is distilled overhead as a final product. The purification column bottoms are mainly oligomers of C6.

520 Ethylene Feed

ETHYLENE RECYCLE

BUTENE-1

.

REACTION SECTION

CATA LYS T

CATALYST REMOVAL SECTION

S P E NT CATA LYS T

.

.

.

.

.

DISTILLATION SECTION

C6

Figure 1. IFP-SABIC Alphabutol Process Scheme

Influence of Reaction Conditions influence of reaction conditions such as temperature, pressure, catalyst molar ratio (AI/Ti), and impurities on the overall ethylene conversion and Butene-1 yield are analyzed as follows: Temperature. The catalyst activity is quite sensitive to temperature changes. As temperature increases, the catalyst activity and the ethylene conversion increase. However, the selectivity to butene-1 production is adversary affected through the increase in by-product, mainly hexenes formation. Another undesirable effect of temperature increase is the extent of polymer formation. The optimum reaction temperature range is generally between 50 to 60~ [141. Pressure. Since, the reactor is operating at the bubble point of the liquid, the pressure is directly related to the conversion, because as the pressure in the system is decreased the ethylene concentration in the liquid phase will also decrease, hence, the conversion of ethylene increases. However, the selectivity to butene-1 production is reduced through the increase in by-products formation. On the other hand, the formation of higher olefins decreases with increasing the pressure. The optimum reaction pressure range was found between 20 to 30 atm. [14]. Catalyst Ratio (AI/Ti). The molar ratio of aluminum alkyl to titanium alkoxide is recognized as an important parameter in the dimerization of ethylene to butene-1. A molar ratio less than 10 favors dimerization while a ratio higher than 10 favors polymerization [ 1]. An optimal catalyst activity was found to exist at AI/Ti range of (2- 4) (subject to the reaction temperature), where it was also noticed that there was no remarkable increase in the polymer formation. Impurities. Impurities such as H20, CO, CO2,02 reduce the catalyst activity. However, if the impurities content in the ethylene feed stock increase, this can be balanced by increasing the catalyst rate. Operational Efficiency Enhancement. Polymer formation characteristics of this process were known to exist at the pilot plant scale. The first commercial plant was scaled up from a

521 pilot scale by a considerable factor. Mitigation design aspects dealing with the problem of polymer formation were not fully built into this plant. Following plant start-up a number of studies and plant observation programs were carried out. These resulted in making some physical modification to the plant such as elimination of dead legs in the reaction section where polymerization could be enhanced. At the same time, operational control aspects were modified such that corrections to the operating conditions could be made before the reaction become unstable. Furthermore, efficient means of heat removal from the reactor were implemented i.e. chilled water cooling loop. The cumulative operational knowledge ofPetrokemya (SABIC affiliate) was a significant factor in increasing the over all plant operational efficiency, in that downtime for polymer cleaning was reduced. As a result of all these steps, the plant's capacity has recently been improved to about 8% over its design capacity. 4. C O M P A R I S O N

OF DIMERIZATION

PROCESSES

IFP-SABIC Alphabutol process shows a high butene-1 selectivity associated with minimum formation of by-products such as cis/trans butenes-2, n-butane and higher olefins. The Phillips process is characterized by the formation of cis/trans butenes-2 due to the isomerization activity of the nickel based catalyst used in the process. Low ethylene conversion and high butene-1 selectivity are obtained in the processes assigned to MIT and Dow. MIT process uses a neopentylidene tantalum complex at high operating pressure, while Dow uses a triethyl-aluminum catalyst at high displacement temperatures. Table 3 presents the operating conditions, ethylene conversion, butene-1 selectivity and other features of the above mentioned dimerization processes TABLE 3. Comparison of ethylene dimerization processes. ,

T

Process Assignee

IFPSABIC Phillips

MIT

Dow

Catalyst System

Operating Conditions

Temperature (~ Titanium50-60 based Nickel48 based

Tantalum based

Triethylaluminum '* mixture ofbutenes

80

277

Ethylene Conversion

Butene-1 Selectivity

Pressure (atm) 20-30

(%) 80-85

(%)

13.7

85-90

50-85*

100

20

95

27

25.7

93

95.5

Remarks

- Low isomerization and polymerization activity - High formation of cis/trans-butene-2 - Product superfractionation is needed. - Catalyst preparation is a complex method. - Catalyst recovery is required. - Low conversion.

522 5. CONCLUSION 1. Production of butene-1 can be achieved by a number of routes and processes; mainly ethylene oligomerization and dimerization. Process employing oligomerization tend to produce a range of products in addition to butene-1. If butene-1 is the main product of interest, then, the ethylene oligomerization processes are less competitive as compared to the ethylene dimerization route. 2. We have reviewed the processes for dimerization of ethylene to butene-1. IFP-SABIC Alphabutol process as yet remains the only commercially proven process. A plant based on this technology has been in operation since 1987. This plant has achieved targeted capacity of 50,000 metric tons/year and has met design requirements. 3. The distinguishing features of the IFP-SABIC Alphabutol technology are: - The process configuration is simple; involving few steps hence it offers lower capital cost as compared to other technologies. - This process employs once through catalyst addition, this catalyst is relatively inexpensive. Hence that operating cost is expected to be lower. - This technology has been proven over a range of plant capacities, hence scale up problems for a new plant are not envisaged. - The butene-1 quality is superior and ideal for production of polyethylene, because of very few by-products and efficient means of separation. - Further improvement of the performance of this process is continuing within SABIC R&D. ACKNOWLEDGMENTS I would like to thank SABIC R&D management for their support and their permission to publish this paper. Special thanks to the referees who made very valuable comments and suggestions about the content of this paper. REFERENCES:

1. A.W. AI-Sa'doun, Appl. Catal. A: General; 105 (1993) 1-40. 2. G. P. Belov, T.F. Dzhabiev, and F. S. D'Yachkovsky in "Mechanisms of Hydrocarbon Reactions", Elsevier, Amsterdam, pp. 507-516, 1975. 3. A.M. AI-Jarallah, J.A. Anabtawi, M.A.B. Siddiqui, A. M. Aitani and A.W. AI-Sa'doun, Catal. Today; 14 (1992) 1-122. 4. C. Carter, (Phillips Petroleum Co.), Surface Conditioning in Olefin Dimerization Reactors, US Patent No. 4,538,018, 1985. 5. C. Carter, (Phillips Petroleum Co.), Olefin Dimerization in a Loop Reactor, US Patent No. 4,242,531, 1980. 6. R. Schrock and J. Fellmann, J. Am. Chem. Soc., 100(1978) 3359-3370;. Chem. Abs. 89, 129635 (1978). 7. G. Parshall, Oligomerization of Olefins, Homogeneous Catalysis, McGraw Hill, New York, pp 56-63, 1980.

523 8. Y. Eidus, S. Minachev, P. Lapidus, A. Avetisyan, and I. Isakov, Butylenes, USSR Patent No. 23 5,016 (1969). 9. D.M. Maschmeyer, A. E. Flower, S. A. Sims, and G. E. White, Process for Making a Mixture of Ethylene and Butene-1, US Patent No. 4 484 016 (1984). 10. D.M. Maschmeyer, Mixtures of Butene-1 and Ethylene, Japan Patent No. 61 122 230 (1986). 11. N. LeQuan, D. Cruypelinck, D. Commereuc, Y. Chauvin, and G. Leger, Butene-1 by Ethylene Dimerization, European Patent No. 135,441 (1985). 12. D. Commereuc, J. Gaillard, and G. Leger, Butene-1 by Dimerization of Ethylene, French Patent No. 2,546,488 (1984). 13. Y. Chauvin, D. Commereuc, and Y. Glaize, Pure Butene-1 from the Crude Product of Ethylene Dimerization, European Patent No. 200,654 (1986). 14. N. LeQuan, D. Cruypelinck, D. Commereuc, Y. Chauvin, and G. Leger, US patent No. 4 532370(1985). 15. IFP Industrial Department, Alphabutol Process for Butene-1 Manufacturing, IFP Publications Paris, 1988.

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Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All fights reserved.

525

C O B A L T C O N T A I N I N G ZSM5 ZEOLITES PREPARATION, C H A R A C T E R I Z A T I O N AND S T R U C T U R E SIMULATION

A. Jentys, A. Lugstein, O. E! Dusouqui* and H. Vinek Institut f~r Physikalische Chemie, TU Wien, Getreidemarkt 9, A-I060 Wien, Austria M. Englisch and J. A. Lercher University of Twente, P.O. Box 217, 7500 AE Enschede, The Netherlands ABSTRACT The incorporation of Co into ZSM5 by direct synthesis, ion exchange and impregnation methods is described. The influence of the preparation method on the density of sites and the stoichiometry of the ion exchange reactions was investigated. On samples prepared by solid state ion exchange, two new Lewis acid sites were formed, while the incorporation of Co by liquid phase ion exchange, synthesis and impregnation resulted in the creation of only one new Lewis acid site. A complete exchange could only be achieved with a stoichiometry of Co+2/AI+3 = 1 by solid state ion exchange. Liquid phase ion exchange resulted in a maximum exchange of about 22 % with a ratio of Co+2/A1+3 = 0.5. 1. I N T R O D U C T I O N Bifunctional catalysts are extending the utilization ofzeolites from typical applications based on their acid properties (e.g., cracking, alkylation or isomerisation), to reactions such as hydrocracking, HDS and HDN, where a hydrogenation function is indispensable (1). The transition metal cations incorporated can be used for various roles such as keeping the acid sites clean from coke and / or to establish the olefin / paraffin equilibrium. Additionally, transition metal containing zeolites might be applied in hydrogenation reactions, when a high sulfur concentration of the reactants limits the use of traditional noble metal catalysts. Ni-Mo or NAt present, industrial hydrotreating processes are based on alumina supported Co-Mo or Ni-W sulfide catalysts (2,3). We expect that transition metal containing molecular sieves will provide attractive alternatives, as the size of the sulfide particles is controlled by the pores and the interaction between the metal and the zeolite lattice will stabilize the cluster. Moreover, the shape of the zeolite pores will constrain the environment available for the reactants and the acid / base properties of the zeolite can be fine tuned to the specific needs of the reaction. In this communication, we discuss the structure of Co containing ZSM5 zeolites prepared by direct synthesis, ion exchange and impregnation methods and compare it with results obtained from an atomistic simulation.

*Present Address: Departement of Chemistry, University of Kuwait, Kuwait.

526 2. EXPERIMENTAL Cobalt containing ZSM5 zeolite was synthesized hydrothermally in a stirred autoclave in 48 h at 743 K according to refs. (4, 5) using COC12.6H20 as cobalt source. After synthesis the sample was calcined to remove the template. Ion exchange was carried out by a liquid phase (6) and by a solid state ion exchange reaction (7, 8). For the ion exchange in liquid phase, 1.0 g ofHZSM5 (Si/Al=26) was heated to 353 K in 20 ml of a 0.2 molar COC12solution for 48 h. The ion exchange procedure was repeated up to five times. The solid state reaction was carried out by grinding HZSM5 (Si/AI=26) together with COC12.6H20 or Co(NO3)2.6H20 and heating the mixture in He atmosphere at 773K for 6 - 14 hours. The amount of Co was selected in order to achieve the desired Co2+/AI3§ ratio in the final material. After all ion exchange reactions the materials were washed repeatedly with water until no anions could be detected in the solution and dried subsequently at 400 K. Impregnation of the zeolite was carried out using COC12.6H20 and Co(NO3)2.6H20 applying the incipient wetness technique. Ir spectra were measured after activating the samples at 773 K in vacuum. The bands of the lattice vibrations between 2090 and 1740 cm~ were used to normalize the intensities of all bands (9). X-ray absorption spectra were measured at liquid nitrogen temperature after drying the samples in He at 373 K. The weight of the samples was selected to achieve an absorption of less than lax =2.0 for the activated catalyst, in order to optimize the signal to noise ratio (10). The energies of all absorption edges were aligned to that of bulk Co. The analysis was carried out using standard analysis procedures as described, e.g., in ref. (10). This included a polynomial baseline approximation, an isolation of the contributions of the coordination shells by Fourier transformation and a determination of the structural parameters of the first coordination shell under the assumption of single scattering and plane waves. The phaseshift and amplitude functions were obtained from experimental data of Co reference components. XANES provides information about the density of vacant states near the Fermi level of the absorber atom (10). By comparing the features of the XANES with those of reference compounds with known oxidation state and structure, changes in the XANES region can be interpreted as changes in the oxidation state or as structural information. The latter can be only of qualitative kind, as XANES is usually the result of the superposition of a multiple scattering process with a final state electron excitation effect. The concentration of acid sites was determined by temperature programmed desorption (t.p.d.) of NH3. The degree of ion exchange was calculated from the difference in concentration of the strong Bronsted acid sites present before and after ion exchange. The structure of the zeolites after synthesis or postsynthetic modification was verified by XRD. The stability and geometry of (CoO)x particles within the zeolite was modeled using atomistic simulation techniques. The techniques for calculating the lattice energy used in this work have been described extensively (11,12) and therefore, will be only briefly discussed here. Long-range Coulombic potentials, defined between ions, were summed to infinity using the Ewald technique (13). Short-range interactions, which were also defined between ions, were

527 parameterized into a Buckingham potential form. Thus, the interaction between two ions separated by a distance rij was calculated by: E(rij)_ qi-qj + A.eC-r~/~)

r,j

C

r

where qi is the charge on ion i, A, P and C are parameters for describing the interaction between the atoms, which were obtained from quantum mechanical cluster calculations and from crystallographic data. The anions were treated as polarizable by virtue of the shell model of Dick and Overhauser (14) and the directional properties of the covalent bonding were modeled using harmonic three body terms around the tetrahedral angle of the silicon atoms. The formation of clusters within the lattice was determined using periodic boundary conditions. For the simulations a purely siliceous material was used. 3. RESULTS The chemical composition of the samples determined by X-ray fluorescence analysis is reported in the Table. The samples are denoted by the Co/AI ratio and the Co source chosen for the synthesis. Sample HZSM5 CoZSM5 CoZSM5 CoZSM5 CoZSM5 CoZSM5 CoZSM5 CoZSM5 CoZSM5 CoZSM5

XRF Results [mol/mol] Co/AI Co/C1 Preparation Method Co [%] Si/A1 0.0 26.6 0.28 1.05 Solid State Ion Exchange COC12 1.0 26.3 0.61 1.67 COC12 2.2 25.9 1.22 1.89 COC12 4.5 24.6 0.30 Co(NO3)/ 1.0 29.3 0.68 Co(NO3)2 2.3 28.3 1.36 Co(NO3)2 4.5 26.2 0.11 1.32 Liquid Phase Ion Exchange COC12 0.4 28.3 Co(NO3)2 6.4 Impregnation 0.21 Direct Synthesis COC12 0.8 25.0 Co Source

(0.3C1) (0.7C1) (1.4C1) (0.3N) (0.7N) (1.4N) (LE) (Imp) (Syn)

The XANES of the samples investigated are shown in Fig. 1. The XANES of the samples prepared by ion exchange with COC12showed almost identical features compared to that of COC12, while samples prepared by using Co(NO3)2 as Co source showed a XANES similar to that of Co304. In the XANES of the synthesized CoZSM5 sample two Co phases were observed. One of them could be clearly identified as CoO from the XANES. The analysis of the EXAFS revealed, that only in the case of the samples prepared with COC12, the chemical environment of the synthesis was still preserved, i.e. in both samples a CoC1 coordination was observed. In the samples prepared with Co(NO3)/oxidic species were formed. Similar to CoO and Co304, Co-O and Co-Co coordinations at short distances were observed in these samples (15).

528

2000

Co(NO:}):, CoCI2

r..-.i

CoZSM5 (1.4 CI)

~~500 Iz i,.,.,.i

CoZSM5 (LE)

LLI ..,..

CoZSM5 (1.4N)

o Ii LL

t.ul000 O O

CoZSM5 (IMP) CoZSM5 (Syn)

E

CoO

500

C%O4

I

I

I

I

I

7700

7750

7800

7850

7900

7950

ENERGY [eV]

Figure 1. XANES of the CoZSM5 Samples The NH3 t.p.d, of the samples prepared by solid state ion exchange using COC12and Co(NO3)2 are compiled in Fig. 2 and Fig. 3, respectively. The NH3 t.p.d, of the samples prepared with liquid phase ion exchange, impregnation and synthesis are shown in Fig. 4. Figures 2a, 3a and 4a show the t.p.d, of NH3 and Figures. 2b, 3b and 4b the changes in the t.p.d, alter subtraction of the starting material. The concentration of the strong Bronsted acid sites vs. the CoZ+/AI3+ ratio are shown in Fig. 5.

After ion exchange, two new desorption states were observed for the samples prepared by the solid state reaction, while only one was observed for the samples prepared by liquid phase ion exchange and impregnation. The t.p.d, profile of the synthesized CoZSM5 can not be directly compared to those of the other samples, as this material was entirely generated during the preparation, while the others were postsynthetic preparations, all of them starting from the same HZSM5 zeolite. The concentration of acid sites present alter the solid state ion exchange are compiled in Fig. 6 for the samples prepared with COC12 and in Fig. 7 for the samples prepared with Co(NO3)2.

529 a

a

=.-

HZSM5

~Zw I-___-

CoZS~15(0.3CI) CoZSM5(0.7CI)

HZSM5

~

CoZSM5(0.3N) CoZSM5(0.7N)

-J <1:

CoZSM5(1.4CI)

CoZSM5(1.4N)

Z-

l ~=j

~' f ~ ~ \

/ ~

~

CoZSM5(1.4CI)

400

500

~

CoZSM5(1.4N) CoZSM5(0.7N)

CoZSM5(0.3CI)

CoZSM5(0.3N)

|

300

~

CoZSM5(0.7CI)

600 700 800 900 TEMPERATURE [K]

|

1000 1100

Figure 2. N H 3 t.p.d, of the samples prepared by solid state ion exchange with COC12

-

300

400

500

600 700 800 900 TEMPERATURE [K]

1000 1100

Figure 3. N H 3 t.p.d, of the samples prepared by solid state ion exchange with Co(N03)2

a

Solid S t a t e Ion E x c h a n g e

,..~ 9 I

,..,

HZSM5

80

r _~ z I~ z

CoZSM5 (Imp) CoZSM5 (LE)

=~n~

CoCI2 9 C~

'. ~

&. ""

9

'X~ ~N ~ "........

Liquid P h a s e Ion E x c h a n g e

' ~

4,

CoCl 2

CoZSM5 (Syn) ;.

-

Z (9 u~ -

~

~

CoZSM5 (Syn)

~

CoZSM5 (LE)

9

2O CoZSM5 (Imp) 300

I

I

400

500

I

I

I

I

600 700 800 900 TEMPERATURE [K]

/

1000 1100

Figure 4. N H 3 t.p.d, of the samples prepared by liquid phase ion exchange with COC12, impregnation and by synthesis

0

0.00

0.25

0.50

0.75

1.00

Aml wl

1.25

A ,,.

1.50

Co=*/AP* Ratio

Figure 5. Stoichiometry of the ion exchange reactions

The ir spectrum of the activated CoZSM5 (1.4C1) sample is compared with that of the starting material in Fig. 8. After ion exchange, the intensity of the band at 3610 cm1, assigned to strong Bronsted acid sites, decreased and the formation of a new band at 3680 cm1 was observed. This band can be assigned to the stretching vibration of a Co-OH group. Assuming similar molar extinction coefficients for Co-OH and Si-OH-A1 hydroxyl stretching vibrations, it can be seen that the density of OH groups alter the ion exchange is much lower than that in the starting material.

530 100'

\\

\N

80

\

80

o'~

A

,oo %

............a

60-

\

....~...-A

60 o.o-~176

,...~~ NN

40

40 ooo ~

~

20

20-

/

4

I

I

I

I

I

0.0

0.2

0.4

0.6

0.8

1.0

0~,. 1.2

0.0

1.4

I

I

I

I

I

0.2

0.4

0.6

0.8

1.0

Co2+/AI3+ Ratio I

'~

ZI

1.2

1.4

Co2+IAI3+ Ratio

9 Bf~nstedAcid Sites

9 Lewis Sites

9 Co Sites I

Figure 6. Concentration of sites in

Figure 7. Concentration of sites in

the CoC12exchanged samples

the Co(NO3)2exchanged samples

U.I_ 0 Z < 1330 133 <

HZSM5

CoZSM5 (1.4Cl) i

i

i

i

3800 3700 3600 3500 3400 3300 3200 3100 3000

WAVENUMBER [cm-1] Figure 8. Ir spectra of the activated HZSM5 and CoZSM5 (1.4el) samples

531

9 O

-3

C00

Zeolite

C00

Gas Phase

5" -4

8 cr-

0-5 N

W

"5 LU~- -6

-7

-8

i

i

i

i

i

i

i

i

i

2

4

6

8

10

12

14

16

18

20

rico

Figure 9. Energy of (CoO)x clusters in ZSM5 and in the gasphase a function of the cluster size

The energies of (CoO)x clusters inside the straight channels of ZSM5 and in the gasphase are shown in Fig. 9.The energy of(CoO)x clusters with a size between 10 and 12 Co atoms was lower inside the pores of ZSM5 than that of the clusters in the gasphase. An attractive interaction between the (CoO)x cluster and the zeolite, which is increasing with the cluster size, leads to an increased stability for larger clusters. However, this interaction is overcompensated by sterical limitations of the cluster caused by the zeolite. Clusters consisting of 10 to 12 Co atoms show the highest stability, smaller clusters can increase their stability by forming larger units by maximizing the cluster-host interaction, while larger clusters are already limited in their size by the structural properties of the lattice. 4. DISCUSSION In order to examine the stoichiometry of the ion exchange reaction, the solid state ion exchange was carried out at various C02+/A13+ratios. Formally two protons (from the SiOHA1 group) can be exchanged with one Co 2§ cation. Therefore, 100% exchange should be achieved at a molar ratio of C02+/A13+ = 0.5 (stoichiometric ratio). However, only for samples with a C02+/A13+ ratio higher than 1, the complete removal of the strong Bronsted acid sites could be observed. On samples with Co2+/Al 3+ ( 1, strong Bronsted acid sites were still present. This indicates that during the solid state ion exchange one Co 2§ cation was exchanged with only one proton. This ratio was obtained for all samples prepared by solid state ion exchange, independent of the ratio and the Co source used. Moreover, this results indicate that achieving an equimolar C02+/AI3+ ratio is preferred over the exchange of one Co 2+ cation with two protons, even at low Co 2§ concentrations.

532 In contrast to the solid state ion exchange, the ion exchange carried out in liquid phase did not lead to a complete removal of the strong Bronsted acid sites. A maximum exchange degree of 22 % was observed, even aider repeating the ion exchange up to five times. By extrapolating the decrease of the number of strong Bronsted acid sites and the corresponding Co concentration of the sample to the Co2+/Al3+ ratio which would be observed in the case of a complete exchange, as shown in Fig. 5., a Co2+/AI3+ ratio of 0.5 was obtained in the ion exchange reaction carried out in liquid phase. This indicates, that in this reaction one Co 2+ cation was replaced by two protons. The incorporation of Co into ZSM5 caused the formation of new acid sites. The samples prepared by solid state ion exchange revealed two new sites that were observed during t.p.d, of NH3. One was at a lower (475 K) and the other at a higher temperature (700 K) than that of NH3 desorbing from the strong Bronsted acid sites (610 K). The number of the former sites was indirectly proportional to the concentration of the remaining Bronsted acid sites. The concentration of the latter sites increased up to a Co2+/Al3+ of about 1.2 and remained constant at higher Co concentrations. Both sites revealed Lewis acid character in the ir spectra of adsorbed pyridine. The XANES of the samples prepared by using COC12showed, that the Co cations are in an environment similar to that found in CoC12. This is further supported by the Co-Cl coordination found in the analysis of the EXAFS and by the presence of Cl" in the material even aider careful and repeated washing of the samples. For the samples prepared by solid state ion exchange with Co(NO3)2, XANES and EXAFS indicated that the Co cations are present as an oxidic phase. Since an equimolar ratio between A1 and Co was obtained for these samples, only one of the charges of the Co 2+cation is balanced by the negative charge of the lattice and (CoCl) +, species as shown in the XAS and the chemical analysis, (CoOH) + species, as observed in the it- spectra, and Co-oxide species, verified by XAS, are present in these samples. In the CoZSM5 samples prepared by direct synthesis, XANES indicated that Co is present in two different states The analysis of the EXAFS revealed that the Co atoms were surrounded by O atoms and that. Co neighbors were present at short distances. If the Co atoms were substituted entirely into the lattice, Co-Co coordinations could not be observed. The impregnation of the zeolite led to Co clusters which were also present in an oxidic environment. The size and location of the clusters were not specific, but the maximum metal loading, obtained by this preparation method was higher compared to the two techniques mentioned above. The simulation of the structure of (CoO)x clusters in the pores of the zeolite clearly showed a preference for the formation of clusters with a size between 10 and 12 Co atoms. Larger clusters were sterically limited by the zeolite lattice. 5. CONCLUSIONS Ion exchange methods and direct synthesis techniques can be used to incorporate Co ions into ZSM5 zeolites at well defined positions. The solid state ion exchange led to the formation of two new (Lewis acid) sites. One was affiliated with the presence of Co cations equimolar exchanged with the protons of SiOHAI groups. The other site, which revealed a very strong Lewis acid character, was associated with the presence of (CoCl) + and (CoOH) + species. Only by solid state ion exchange methods, a complete exchange could be achieved. By of liquid phase ion exchange only a partial cation exchange (22 %)was possible, even aider

533 multiple ion exchange treatments. The stoichiometry of the ion exchange reaction was found to be Co+2:AI+3 = 1:1 for the solid state ion exchange and Co+2:AI+3= 1:2 for the liquid phase ion exchange. The direct synthesis of CoZSM5 results not only in a lattice substitution of Si with Co, but also in the formation of small Co-oxide clusters. The preparation of CoZSM5 by impregnation led to the material with the highest Co concentration, but also the lowest dispersion and the least defined clusters. The results of the simulation showed a preference for the formation of CoO clusters consisting of 10 to 12 atoms. Smaller clusters will have the tendency to combine to form larger clusters, while clusters above this size are structurally restricted by the zeolite lattice. 6. ACKNOWLEDGEMENTS The work was supported by the "Fonds zur FOrderung der Wissenscha~lichen Forschung" under project FWF P9167. We are grateful to P. Wiede and K. Mereiter for the XRD and XRF measurements. X-ray absorption spectra were collected at the National Synchrotron Light Source (Beamline X23A2), Brookhaven National Laboratory, which is supported by the U.S. Department of Energy, Division of Materials Sciences and Division of Chemical Sciences. REFERENCES 1

J. A. Martens, P. A. Jacobs and J. Weitkamp, Appl. Catal. 20 (1986) 239

2

B. Delmon, Catal. Lett. 22 (1993) 1

3

C. T. Douwes and M. Hart, Erdl und Kohle 21 (1968) 202

4

J. A. Rossin, C. Saldarriaga and M. E. Davis, Zeolites 7 (1987) 295

5

T. Inui and D. Medhanavyn, Appl. Catal. 18 (1985) 311

6

H. S. Sherry, in Ion Exchange (ed. J. Marinsky), Chap. 3 (1969).

7

H. G. Karge, H. K. Beyer and G. Bartdy, Catalysis Today 3 (1988) 41

8

A. V. Kucherov and A. A. Slinkin, Zeolites 7 (1987) 43

9

A. Jentys and J.A. Lercher, Stud. Surf. Sci. Catal. 46 (1989) 585

10

D. C.

11

EXAFS, SEXAFS and XANES, Chemical Analysis Vol. 92 (Wiley, New York, 1988). 'Modelling of Structure and Reactivity in Zeolites' C.R.A. Catlow ed. (Academic Press

12

London, 1992). 'Computer Simulations of Solids', C. R. A. Catlow and W. Mackrodt ed., Lecture

13

Notes in Physics 166, (Springer Verlag, Berlin 1982). M. P. Tosi, Solid State Phys. 16 (1964) 517

14

B. G. Dick and A.W. Overhauser, Phys. Rev. 112 (1958) 90

15

A. Jentys, A. Lugstein, H. Vinek, M. Englisch and J. A. Lercher to be published.

Koningsberger and R. Prins, eds., Principles, Applications, Techniques of

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Catalysts in PetroleumRefining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All fights reserved.

535

ACID-BASE PROPERTIES OF SOME ZEOLITES AND THEIR ACTIVITY IN THE D E C O M P O S I T I O N OF N-HEXANE Susumu Tsuchiya

Department of Advanced Materials Science and Engineering, Faculty of Engineering, Yamaguchi University, Toldwadai, Ube, Japan 755 ABSTRACT Since zeolites are typical acid-base catalysts, their acid-base properties are of great importance in investigating the catalytic decomposition of hydrocarbons. Three methods --titration, temperatureprogrammed desorption, and characterization by test reaction -- are employed to measure acid-base properties. In this study, n-hexane was used as a model hydrocarbon and its decomposition over HY, HCeY, HSmY, and HCuY zeolites was investigated. Depending on the metal exchanged, n-hexane conversion and product distribution were observed to vary in the higher conversion region. The relation between product distribution and the acid-base properties of the zeolites are discussed. 1. INTRODUCTION Since zeolites are typical acid-base catalysts, their acid-base properties are of great importance in investigating the catalytic decomposition of hydrocarbons. Three methods-titration, temperature-programmed desorption (TPD), and characterization by test reactions -are now employed for the measurement of the acid-base properties of catalysts. We previously proposed a method to estimate the acid-base properties of catalysts by means of the reaction profiles of n-butenes isomerization [ 1]. In this paper, n-hexane was employed as a model hydrocarbon, and the decomposition of n-hexane over zeolites was investigated. Focus was centered on the relation between the product distributions of the reaction and the acid-base properties of zeolites. 2. METHOD 2.1. Materials and Procedure Some kinds of metal ion-exchanged zeolites were prepared by ion exchange [2]~ These zeolite samples were used for the isomerization of n-butenes and the decomposition of nhexane. The apparatus used for the n-butene, isomerization was a closed, circulating system with a U-shaped reactor, a conventional vacuum line and a gas-chromatograph. With the gas chromatograph, we analyzed the reaction mixtures at suitable times following the reaction. The volume of the closed circulating system corresponds to about 288 cm3at a reaction temperature of 373K. The apparatus used for the decomposition of n-hexane was a conventional flow-type reactor, where nitrogen was used as a carrier gas. The reaction mixtures were occasionally sampled with a syringe after the reaction for analysis by gas chromatograph.

536 2.2 Network Analysis

The n-butene isomerization reaction is often first-order with respect to the pressure of the reactants in most cases Since the total amount of substance of the butenes does not change with the extent of isomerization, xl + xr + xt = 1

(1)

where Xl, x~, and xt are the mole fractions of 1-butene, cis-2-butene, and trans-2-butene, respectively. Equations express the overall reaction rates accordingly. The mole fractions, xi (i = 1, c and t), of individual isomers during the interconversion reaction can be estimated by solving the following equations: dxl/dt = - (klr k~lX~ + ktlXt dxddt = kleXl - (k~t+k~Oxc + kt~xt dxt/dt = kttxt + ketxe - (ktt+kt~)xt

(2) (3) (4)

The rate constant of the formation reaction of j-isomer from i-isomer, kij, can be determined from the reaction rates and the selectivity ratios in the initial stage of the reaction. The mole fractions, xi, thus calculated are shown in a triangular graph, and the set of curves which correspond to the change of the mole fraction with time obtained is called the nbutenes reaction profile. Tsuchiya and Imamura [ 1] calculated a lot of n-butenes reaction profiles as a function of ( = k e t / k l t ) o n the basis of the Wei-Prater method [3]. In view of the acid-base catalysts, they classified these profiles into four types: (1) cis-convex, (2)rake, (3)transconvex and (4) tree. The reaction profiles for base catalysts are of the cis-convex type. Reaction profiles of the trans-convex type are obtained by acid catalysts. X (=kle/klt) and y

3. RESULTS AND DISCUSSION 3.1. Isomerization of n-Butenes Figures la to ld show the typical profiles ofisomerization of n-butenes over (a) HNaY (0%), (b) HNaY (7.6%), (c) HNaY (33%), and (d) HNaY (55%) at 473K, where the percentage exchanged is shown in parentheses. The profile for HNaY (0%), HY, is of the trans-convex type, and that for HNaY (7.6%) is of the rake type. The profiles for HNaY (33%) and HNaY (55%) are of the cis-convex type. Judging from Figures,. la to ld, the profile seemingly varied from the trans-convex to cis-convex type via the rake type by exchanging the protons in the zeolites by means of sodium ions. In HNaY (7.6%), sodium ions replaced the minor portion of the protons, and the transconvex type profile for HNaY (0%) was modified to the rake-type profile. In HNaY (33%) and HNaY (55%), on the other hand, sodium ions replaced the major portion of the protons, and the profiles became cis-convex in type. The type of profile produced is dependent on the degree of proton-sodium ion displacement. The acid-base property of zeolites can accordingly be characterized by these profiles. Figure 2 shows the reaction profile of n-butenes isomerization over HSmY (10%) at 273 K. The curves which start at both the left and fight sides of the triangular graph go straight downward in the wide range of conversion. This profile, therefore, is of the rake type.

537

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I 1oo

c 100

80

60 ,

40

20

0 t

c 100

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o

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- c-?-B tool Z

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oo

t

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Figure 1. Reaction profile of n-butenes isomerization over (a) HY (0%) at 373K, (b) HNaY (7.6%) at 473K, (c) HNaY (33%) at 473K, and (d) HNaY (55%) at 473K. Figures 3 and 4 show the reaction profiles of n-butenes isomerization over HCeY (10%) at 273K and over HCuY (10%) at 273K, respectively. These profiles are of the rake type, suggesting the existence of both acid and base sites on the surface of the catalysts.

3.2. Decomposition of n-Hexane The decomposition of n-hexane was investigated over HY, HCeY, HSmY and HCuY. Figure 5 shows typical product distributions ofn-hexane decomposition over these zeolites. The reaction temperature was 723K and the W/F was 0.85 g.h/mol. The conversions ofnhexane decomposition over HY, HCeY, HSmY and HCuY were 55%, 55%, 78% and 74%,

538 1-B

8,,~o 4 I

,fro

/I I IX~,~

0

c-2-B I00

80

60 40 c-2-B tool I

20

00 t-2-B

0

Figure 2. Reaction profile of n-butenes isomerization over HSmY at 273 K.

1-B

.or /

2~~

,~ /I

~o

/ / / / // X\~0~. "~,. I I

-'~/1111

/

/X

IX

§

0 c-2-B 100

100 80

-

60

40 c-'2-B mol Z

20

0 t-2-B

Figure 3. Reaction profile of n-butenes isomerization over HCeY at 273 K.

539

Figure 5. Product distribution of n-hexane decomposition over zeolites. (Reaction temperature: 723K. C l" methane, C2" ethylene, C2: ethane, C3" propylene, C3" propane, C4" butenes, C4" butanes, C5': pentenes, and C5" pentanes)

540

Figure 6. Product distribution of the reaction over HY, (Reaction temperature: 723 K. CI' methane, C2': ethylene, C2: ethane, C3': propylene, C3: propane, C4': butenes, C4: butanes, C5': pentenes, and C5: pentanes).

respectively. The product hydrocarbons observed were methane, ethane, ethylene, propane, propylene, butane, butenes, pentane and pentenes. Depending on the kinds of metal exchanged, the products distributions of the reaction were different. To obtain information on the initial state of the decomposition, the product distribution of the reaction at very low conversion were investigated. In these experiments, the amounts of catalysts used were small, and fresh catalyst samples for the same batch were used in individual runs.

Figure 6 shows the product distribution of the reaction over HY at 723K with the conversion of n-hexane in the lower conversion region. The distribution was also obtained on the basis of the statistical treatment of the experimental data. The formation of methane was observed. The relative amounts of ethylene and ethane formed were much smaller than those of propylene and propane, and they were seemingly not varied much by the conversion. The relative amounts of propylene formed were larger than those of propane, and decreased with increases in the conversion. The formation of propane, on the other hand, increased with increases in the conversion. The relative amounts ofbutenes and butanes formed were smaller than those ofpropylene and propane. The relative amounts ofbutenes was not varied much by conversion, and those of butanes, pentane and pentenes decreased with increases in the conversion. Figure 7 shows the product distribution of the reaction over HSmY at 723K with the conversion of n-hexane in the lower conversion region. The product distribution for HSmY resemble that for HY.

541

Figure 7. Product distribution of the reaction over HSmY.(Reaction temperature: 723 K.C1: methane, C2': ethylene, C2: ethane, C3': propylene, C3: propane, C4': butenes, C4: butanes, CS': pentenes, and C5: pentanes)

Figure 8. Product distribution of the reaction over HCuY. (Reaction temperature: 723 K. C 1: methane, C2': ethylene, C2: ethane, C3': propylene, C3: propane, C4': butenes, C4: butanes, C5': pentenes, and C5: pentanes.)

Figure 8 shows the product distribution of the reaction over HCuY at 723 K with the conversion of n-hexane in the lower conversion region. The product distribution for HCuY is seemingly a little bit different from that for HSmY.

542 The formation of methane was also observed. The relative amounts of ethylene and ethane formed were much smaller than those of propylene and propane. The relative amounts of propylene formed were larger than those of propane, and decreased with increasing the conversion. The formation of propane increased with increases in conversion. The relative amounts of butenes, butanes and pentenes formed were smaller than those ofpropylene and propane, and decreased with increases in the conversion. The relative amount of pentane formed did not vary much with change in the conversion. The product distribution of the reaction varied with change in the conversion. In the lower conversion region, propane and propylene were predominant. The product distributions obtained by decomposition over HSmY and over HCuY were not very different from that over HY. The results suggest that the different product distribution in the higher conversion region that were dependent on the kind of rare earth metal exchange employed was is due not to the initial reaction, but rather to the second or subsequent reactions. The oxides of alkali metals, alkali earth metals, and rare earth metals can be classified as base catalysts, and the surface basicity of Y-zeolites may be increased by rare earth ion exchange. The copper ions may also increase the basicity. The product distribution in the higher conversion region may accordingly be influenced by the increase in basicity. 4. CONCLUSIONS 1. The reaction profiles of n-butenes isomerization over HSmY, HCeY and HCuY were of the rake type, suggesting the existence of both acid and base sites on the surface of the zeolites. 2. The product distributions obtained by decomposition over HSmY and HCuY were not very different from that obtained over HY in the lower conversion region. 3. The product distributions obtained by decomposition over HSmY and HCuY were different from that obtained over HY in the higher conversion region. This difference is due to the second or subsequent reactions. Product distribution may be influenced by the increase of their surface basicity. ACKNOWLEDGEMENT A part of this work was carried out as a research project for the Japan Petroleum Institute commissioned by the Petroleum Energy Center with a subsidy from the Ministry of Trade and Industry, for the author's thanks are due. REFERENCES 1. Tsuchiya and H. Imamura, Shokubai (Catalyst) Tokyo, 25 (1983) 133; S. Tsuchiya, AcidBase Catalysis, eds.. K. Tanabe, H. Hattori, T. Yamaguchi and T. Tanaka, V. C Weinheim, (1989) p. 189. 2. Namba, Handbook for Experimental Method in Catalytic Research,. Y. Murakami, (ed.), Kodansha Scientific, Tokyo, 1986. 3. Wei and C. D. Prater, in Adv. Catal vol. 13.,. D. D. Eley, P. W. Selwood, P. B. Weisz, (eds.), Academic Press, New York, (1962) 203.

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All rights reserved.

543

R E D U C T I O N AND SULFIDATION PROPERTIES OF IRON SPECIES IN FETREATED Y-ZEOLITES FOR HYDROCRACKING CATALYSTS Kazuhiro Inamura and Ryuichiro lwamoto

Central Research Laboratories, Idemitsu Kosan Co., Ltd., 1280 Kami-izumi, Sodegaura, Chiba 299-02, Japan

ABSTRACT Temperature programmed reduction (TPR) and temperature programmed sulfiding (TPS) were used to characterize reduction and sulfiding properties of Fetreated Y-zeolites, which were prepared by treating NH4Y-zeolite with an aqueous ferric nitrate solution (Fe-treatment). It was demonstrated that three types of the Fe-species are present in the Fe-treated Y-zeolites: ion-exchanged type species, small Fe-oxide clusters, and Fe oxides without interaction with the zeolite framework (including aggregated ferric oxide), the proportion of which is dependent on the extent of the Fe-treatment. The small Fe-oxide clusters, which are probably situated inside the supercages through a coordination with the framework oxygen atoms, are responsible for high activity for toluene disproportionation. A realistic production control for the active Fe-treated Y-zeolite catalyst has been achieved for the first time by using the TPR and TPS techniques. 1. INTRODUCTION Recently Idemitsu Kosan Co., Ltd. has developed hydrocracking catalysts for hydrotreating of atmospheric residual oil in a fixed bed process (Figure 1). Especially, the catalysts containing Fe-treated Y-zeolites showed a high activity and selectivity to middle distillates. Typical activity data are shown in Figure 2. R-HYC5, which contains the Fe-treated Y-zeolite (FellY), exhibited the highest conversion of all, although R-HYC4, which contains the ultra-stable Y-zeolite (USY), already showed higher activity than a conventional HDS catalyst (alumina supported Co-Mo catalyst).

the

The most characteristic catalyst itself exhibits

feature of the Fe-treated Y-zeolite catalyst was that high activity and low coke deposition for toluene

544 disproportination especially in the presence of H2S [ 1], which is commonly known as a strong poison against the catalytic active species. This may be advantageous to the hydrocraking catalyst, because H2S always exists in a catalyst bed during the hydrodesulfurization (HDS) and hyrdocracking (HYC) reactions.

Figure 1. An application of the R-HYC catalyst for the existing RC-HDS process.

Figure 2. Hydrocracking activity of RHYC catalyst: Feed; KW-AR, LHSV = 0.3 h -1, PHz = 135 kg cm -2.

The objective of the present study is to examine the reduction and sulfidation properties of the Fe-treated Y-zeolite by using temperature-programmed reduction (TPR) and sulfiding (TPS), in order to quantitatively determine the active species and support the production control of the commercial Fe-treated Y-zeolite catalysts. An interpretation of the reduction and sulfidation mechanisms of several types of Fetreated Y-zeolites is presented.

2. EXPERIMENTAL

According to a conventional procedure [2], Fe2+-exchanged Y-zeolite was prepared as a reference compound by treating NH4Y [LZY-82 (UCC)] with an aqueous solution of FeSO4 (0.25 M, 333 K, 60 min) under a nitrogen atmosphere, and is denoted Fe-LZY. In addition, Fe-LZY(A) was prepared by treating LZY-82 with a aqueous solution of FeSO4 (1.0 M, 363 K, 120 min) in air. A series of Fe-treated Y-zeolites was prepared by treating LZY-82 with a 0.250 M Fe(NO3)3 solution at 293 K for 0.5 h, followed by heating to 323 K and keeping at

545 323 K for 2 h [1]. Samples were taken from the solution every 10 - 30 min intervals, washed with hot distilled water, and dried at 363 K for 3 h in static air. The resulting Fe-treated Y-zeolites are denoted Fe/LZY(a) - (h) depending on the sampling time, as shown in Figure 3B. Fe/LZY(i) was obtained from the residual slurry after cooling to room temperature overnight. Before the TPR and TPS measurements, the samples were calcined in situ at 650 K for 2 h in dry air. The instruments and the experimental procedures have been described elsewhere [3, 4].

I

8

B

1

i

c ~

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h

e

I

I

I

I

!

A 323 ~303 1.5

....

1 283

1.0 0

40

80

120

O

160

T r e a t m e n t Time / min

Figure 3. The change in physicochemical properties of Fe/LZY (a) - (h) as a function of treatment time: (A) pH (left) and temperature (right) of the zeolite suspension, (B) the amount of supported iron (as Fe203) on the zeolite.

3. RESULTS AND DISCUSSION 3.1. Fe-Exchanged Y-Zeolites The reduction of the Fe-exchanged Y-zeolites can be considered to proceed as Fe3+ --, Fe2+ in the a-TPR peak (low temperature) and Fe2+ ~ Fe in the [~-TPR peak (high temperature) as shown in Figure 4(a). This interpretation is consistent with the view that Fe2+-species are stabilized inside the sodalite cages and/or the hexagonal prisms of Y-zeolite [2]. The unique TPS patterns for the Fe-exchanged Y-zeolites can be also interpreted in the same manner. The detailed descriptions will be appeared in the following section.

In the case of Fe-LZY(A), the aggregated ferric oxide deposited on the zeolite surface can be easily distinguished from the ion-exchanged type species by evaluating the TPR profile as shown in Figure 4(b). The H2 consumption due to the reduction for the aggregated ferric oxide is estimated to be 66% of the total H2 consumption.

546

[Fe]dep= [ a ] - [ 19]12

~[a

[B]

~o.= [ B ]12

Fe . LZY (A )

J a -peak

.

.

.

.

19-pea i/k~

~

.

I

300

_ j___~

I

I

I

I

I

I

I

I

I I<

I

z~ I~"

500 700 900 1100 1300 1350 Reduction Temperature (K)

Figure 4. T P R patterns of Fe-exchanged Yzeolites. Possible contribution of the aggregated ferric oxide to the T P R profile of Fe-LZY(A) can be determined as [Feld~p = [a] - [13]/2, with [a] and [13] the amount of H2 .onso,

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Temperature /K

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300 500 700 900 1100 1300 1350

Figure 5. T P R patterns of a series of Fe-treated Y-zeolites [Fe/LZY (b) - (i)].

\

..... [I/I

(b)

'

i

,

600 900 1200 1350 Temnerature/K

TPS patterns (H2S) of LZY-82

and a series of [Fe/LZY (b) - (i)].

Fe-treated

Y-zeolites

547 3.2. Fe-Treated Y-Zeolites [Fe/LZY(b) - (i)]

The TPR and TPS (H2S) patterns for Fe/LZY(b) - (i) are shown in Figures 5 and 6, respectively. In the beginning of the Fe-treatment [Fe/LZY(b) and Fe/LZY(c)], the TPR patterns are identical to that of the Fe-exchanged Y-zeolite (Fe-LZY in Figure 4(a)). This is further confmned in the TPS patterns of the same samples. According to the previous report [3], peak II (a set of an H2S evolution and H2 consumption, the H2consumption side of the TPS patterns is not shown here) appeared at ca. 600 K and peak III (a steep H2S consumption) at 1310 K in Figure 6(b) and (c) are in fair agreement with that of Fe-LZY. The results of both TPR and TPS data suggest that the Fe-species of the Fe-treated Y-zeolite are introduced in the zeolite in the ionexchanged form at the first stage of the preparation. During the heat treatment to 323 K in solution, the Fe-speeies change their structure, as revealed by their easier reduction (the low-temperature shifts of the 13 peaks in TPR) and sulfidation properties (the low-temperature shifts of the peak II and the decrease in intensity of the peak III in TPS). A possible explanation for the change in the state of the Fe-speeies in the zeolite may be the migration of Fe ions from the ion-exchanged sites (in the sodalite cages and/or hexagonal prisms [2, 5, 6]) to the supereages, where the Fe-oxide species are considered to facilitate the reduction and the sulfidation because of a smaller size restriction. On the other hand, the invariance of the ~x peak position and [13]/[ct] ratio in TPR during the treatment indicates that Feoxide species are still stabilized through a coordination with the framework oxygen atoms. Accordingly, this can be attributed to the formation of small Fe-oxide clusters inside the supercages with coordination to the framework oxygen atoms. A prolonged Fe-treatment of the zeolite suspension causes substantial changes on both the TPR and TPS patterns of Fe/LZY(i). The most remarkable change can be seen in twin H2S-evolution peaks (with twin H2-consumption peaks) in the peak II region in TPS (Figure 6(i)). The appearance of the twin hydrogenative desorption peaks of the sulfur [4] suggests that at least two different types of Fe-speeies exist in Fe/LZY(i). It is conceivable that an oligomerization of the Fe-oxide clusters proceeds with the prolonged Fe-treatment and finally the small Fe oxide aggregates to bulk ferric oxide. The reduction of such species can be much easier and thus assigned to the r peak in TPR (Figure 5(i)). Taking into account the above considerations, it can be concluded that three types of the Fe-speeies are present in the Fe-treated Y-zeolites: ion-exchanged type species, small Fe oxide clusters inside the supereages, and Fe oxides without interaction to the zeolite, and classified by combining TPR and TPS as the stmunary in Table 1.

548 Table I TPR

TPS

Classified and Quantified Fe-Species by TPR and TPS [13]/[cq = 2: A or B 13/cxpeak ratio [[3]/[r > 2: +C 13peak peak II

cluster size of B A or B cluster size of B twin: +C

single:

A peak III A: ion-exchanged type species, B: Fe-oxide clusters inside the supercages, and C: Fe oxides without interaction to the zeolite (aggregated ferric oxide).

3.3. Production Control for the Active Fe-Treated Y-Zeolites

A possible Fe-species distribution estimated from the TPR and TPS results during the Fe-treatment is schematically presented together with results on the catalytic activity for toluene disproportionation in Figure 7. The activity of the Fe-treated Y-zeolites increases dramatically by applying the heat treatment [Fe/LZY(e), and Fe/LZY(f)]. This increase in activity coincides with the formation of the small Fe-oxide clusters at the expense of the ion-exchanged species. It can be presumed that toluene molecules cannot approach the ion-exchanged species even inside the sodalite cages as well as that inside the hexagonal prisms. The high activity can be accounted for by the formation of small Fe-oxide clusters inside the supercages.

Figure 7. (a) Catalytic activity for toluene disproportionation, and (b) possible Fespecies distribution estimated from TPR and TPS results: (A) ion-exchanged type species, (B) Fe-oxide clusters inside the supercages with strong interaction to the framework oxygen atoms, and (C) Fe oxide without interaction to the zeolite.

549 Hidaka and co-workers have proposed that the adsorption of H2S on such oxidic Fe-species accounts for the generation of the unique acidity required to catalyze the hydrocracking reaction as well as the toluene disproportionation [ 1 ] . From the standpoint of the production control, it is desirable that the Fe-treated Y-zeolite should be taken out at a time when the amount of the small Fe-oxide clusters inside the supercages reaches its maximum, and when the amount of the aggregated ferric oxide is still at a minimum. Although the exact structure of the such active Fe-species is still needed to examine, such quality control over Fe-species in the Fe-treated Y-zeolites has been applied to the development of commercial resid hydrocracking catalysts. Commercial applications using the R-HYC catalyst and resid HDS units have been realized in Idemitsu, Japan, and Valero, U.S.A. [7, 8]. Idemitsu has been developing new catalysts, which contain new Fe-treated Y-zeolites and which have fia'ther improved in a hydrocracking activity and in a selectivity to middle distillates.

REFERENCES

1. S. Hidaka, A. Iino, M. Gotoh, N. Ishikawa, T. Mibuchi, and K. Nita, Appl. Catal., 43 (1988) 57. 2. R.L. Garten, W.N. Delgass, and M. Boudart., J. Catal., 18 (1970) 90. 3. K. Inamura, T. Takyu, Y. Okamoto, K. Nagata, and T. Imanaka, J. Catal., 133 (1992) 498. 4. K. Inamura, R. Iwamoto, A. Iino, and T. Takyu, J. Catal., 142 (1993) 274. 5. W.N. Delgass, R.L. Garten, and M. Boudart, J. Phys. Chem., 73 (1969) 2970. 6. J.R. Pearce, W.J. Mortier, and J.B. Uytterhoeven, J. Chem. Soc. Faraday Trans. 1, 77 (1981) 937. 7. T. Yamamoto, H. Sue, and T. Ohno, "Resid Upgrading to Produce Transportation Fuels", Paper presented 13th World Petroleum Congress, Buenos Aires, 1991. 8. S. Uchiyama, and T. Ohno, Proc. 5th China-Japan Joint Seminar on Research and Technology for Petroleum Refming, (1994) 33.

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Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All fights reserved.

551

PREPARATION OF HIGHLY ACWIVE ZEO[ATE-BASED HYDRODESULFURIZATION CATALYSTS: ZEOIATF.,-SUPPORTED Rh CATALYSTS M. Sugioka, C. Tochiyama, F. Sado and N. Maesaki

Department of Applied Chemistry, Muroran Institute of Technology, 27-1 Mizumoto-cho, Muroran 050, Japan ABSTRACT It was revealed that Rh/USY showed the highest activity among Rh supported on various zeolites and its catalytic activity was higher than commercial CoMo/A1203 catalyst for the hydrodesulfurization of thiophene at 400~ The catalytic activity of Rh/USY decreased gradually with the reaction time. However, the catalyst deactivation of Rh/USY with reaction time was remarkably improved by the addition of small amounts of alkali metal salts. It is concluded that Rh/USY modified with alkali metal salts are potential highly active second generation hydrodesulfurization catalysts for petroleum feedstocks. 1. INTRODUCTION Hydrodesulfurization of petroleum feedstocks is one of the important processes in the petroleum industry to produce clean fuels. CoMo/AI203 catalyst has been widely used in hydrodesulfurization process of petroleum. However, recently, the development of highly active hydrodesulfurization catalysts with higher activity than commercial CoMo/AI203 hydrodesulfurization catalyst has been claimed in the petroleum industry to produce lower sulfur content fuels because of serious problems of air pollution on global scale by burning petroleum feedstocks. It has been reported that metal-zeolite catalysts have high possibility as new hydrodesulfurization catalysts for petroleum [1-9]. The authors have been investigating the catalytic desulfurization of organic sulfur compounds over zeolites [ 10-12] and also developing highly active zeolite- based hydrodesulfurization catalysts [ 13-19]. In the present work, the authors prepared various zeolite-supported Rh catalysts and examined their catalytic activities for the hydrodesulfurization of thiophene in order to develop highly active second generation hydrodesulfurization catalysts for petroleum feedstocks. 2. EXPEREVIENTAL Hydrodesulfurization of thiophene was carried out at 400~ under 1 atm by use of a conventional fixed bed flow reactor. Thiophene was introduced into the reactor by passing hydrogen through thiophene trap cooled at 0~ The reaction products were analysed by gas chromatograph(FID). Zeolite-supported Rh catalysts were prepared by impregnation method using RahC13 aqueous solution and the Rh loading was 0.5-5 wt%. Alkali metal-modified Rh/USY catalysts were prepared by addition of alkali metal salt aqueous solutions to Rh/USY catalyst. All

552 Rh/zeolite catalysts were calcined at 500~ for 4 hr in air and were reduced at 450~ for 1 hr prior to use. Presulfiding of Rh/USY catalyst was carried out at 400~ for 1 hr by using 5%HzS-H2 mixture. 3. RESULTS AND DISCUSSION 3.1 Activities of Rh/zeolite catalysts

The catalytic activities of Rh supported on various zeolites such as NaY, NaX, NaA, NaZSM-5, NaMordenite, HY, USY, HZSM-5, HMordenite, etc. for the hydrodesulfurization of thiophene were examined at 400~ Table 1 shows the catalytic activities of various Rh/zeolite catalysts in the hydrodesulfurization of thiophene. It was revealed that the activities of Rh/zeolite catalysts were markedly changed by the kind of zeolites. Rh supported on proton type zeolite(HZ) showed high catalytic activity but Rh supported on sodium zeolites(NaZ) except NaY showed low activity. Especially, Rh supported on HZ with large pore diameter such as USY and HY zeolites showed considerably high catalytic activity for the hydrodesulfurization of thiophene in comparison with that on HZ with small pore diameter like HZSM-5 and HMordenite. It was found that Rh/USY showed the highest initial activity and this activity was higher than commercial CoMo/A1203 catalyst as shown in Figure 1. The reaction products in the hydrodesulfurization of thiophene over Rh/USY catalyst were mainly hydrogen sulfide and C4 hydrocarbons and small amount of C1 -C3 hydrocarbons were also formed. Furthermore, the effect of presulfiding with 5%H2S-H2 mixture on the catalytic activity of Rh/USY was examined. It was revealed that the catalytic activity of Rh/USY was enhanced by the presulfiding treatment as shown in Figure 2. This indicates that Rh/USY catalyst is not poisoned by sulfur compounds and this catalyst has high sulfur-tolerant ability in the hydrodesulfurization of thiophene. On the other hand, the catalytic activities of various H-zeolites used as carriers in Rh/Hzeolite catalysts for the cracking of thiophene and cumene were also examined at 400~ by use of a pulse reactor under helium stream. It was ascertained that USY zeolite showed the Table 1. Catalytic activities of Rh/zeolite catalysts for the hydrodesulfurization at 400~ (W/F = 37.9g.hr/mol.; H2/Thiophene=30). Catalyst 5wt%Rh/NaY 5wt%Rh/NaZSM-5 5wt%Rh/NaMord. 5wt%Rh/NaX 5wt%Rh/NaA a) After 10 min.

Conversion(%) a) 79.5 34.3 23.5 22.3 17.5

Catalyst 5wt%Rh/USY 5wt%Rh/HY 5wt%Rh/HZ SM- 5 5wt%Rh/HMord.

CoMo/AI203

Conversion(%) a) 98.2 93.4 38.2 34.3 77.2

553

to0 --- 80 o

9 t.,,,I

t%

Rh/HY

60

~ 4o 0

~

5

20

w t % 5wt% Rh/HZSM-5

R/vtH-Mordenite

!

0 0

...... 1'

2' .......... 3' . . . . . .

4"

Time on Stream W/F =

37.9 g.hr/mol,

7

5. . . . . . . . .

(hr)

H 2 / Thiophene =30

Figure 1. Hydrodesulfurization of thiophene over Rh/H-zeolite catalysts at 400~

too ,~

80

o

60

" ~

Presulfided5wt% Rh/USY 2

* t,,ll

3

5wt%

Rh/USY

0

(.) 20

0

1

2 Time W/F =

3 4 on Stream

37.9 g.hr/mol,

5 (hi')

H 2 / Thiophene - 3 0

Figure 2. Effect of presulfiding on the activity of Rh/USY catalyst.

6

554

100 80

o

d_ C'oMo/AI203

60 " ~ , . ,

9 ~,,,I

3wt%

Rh/USY

5wt% Rh/USY

~. 40 o L)

"

i

I .

0

~

lwt%

Rh/USY

0.5wt Rh/USY

!

!

1

2

.,

I

j

3 4 T i m e on S t r e a m

W/F = 37.9 g.hr/mol,

5

6

(hr)

H 2 / Thiophene =30

Figure 3. Effect of amount of Rh loading on the activity and catalyst life of Rh/USY.

highest activity among H-zeolites for the cracking of both thiophene and cumene. This indicates that the strong BrOnsted acid sites of USY in Rh/USY catalyst play an important role for the hydrodesulfurization of thiophene. That is to say, it is assumed that the strong BrOnsted acid sites of USY in Rh/USY catalyst act as active site for the activation of thiophene, whereas Rh acts as active site for the activation of hydrogen in the hydrodesulfurization ofthiophene. In other words, Rh/USY catalyst behaves as bifunctional catalyst for the hydrodesulfurization of thiophene as well as the reduced MeY zeolite catalysts as described in our previous papers[ 13-15]. 3.2 Improvement of catalyst deactivation of Rh/USY

It was revealed that Rh/USY showed higher catalytic activity than commercial

CoMo/A1203 catalyst in the hydrodesulfurization ofthiophene. However, the catalytic activity of Rh/USY decreased gradually with the reaction time as shown in Figure 1. This may be due to the accumulation of carbonaceous deposit on Rh/USY catalyst surface. Thus, we tried to improve the catalyst deactivation of Rh/USY by various procedures. We attempted to change the dispersion of Rh on USY by changing Rh content in Rh/USY catalyst in order to enhance the hydrogenating ability for carbonaceous deposit on Rh/USY catalyst. Figure 3 shows the effect of Rh content on the catalytic activity and catalyst life of Rh/USY in the hydrodesulfurization of thiophene. It was found that the catalyst deactivation of Rh/USY was not improved by changing the Rh content in Rh/USY catalyst as shown in Figure 3. It is assumed that the strong BrOnsted acid sites are prerequisite for the activation of thiophene in the hydrodesulfurization ofthiophene on Rh/USY. However, strong Br6nsted

555

100 j0.Swt%

Na-5wt%

Rh/USY

80

CoMo/ o t,t)

o

O3

60 0.25wt%

Na-5wt%

Rh/USY

4o 20

0

I

I

I

1

2

3 Time

I

!

4 on Stream

W/F = 37.9 g.hr/mol,

5 (hr)

!

,,I

6

7

H 2 / Thiophene =30

Figure 4. Effect of amount ofNa loading(NaOH) on the activity and catalyst life of Rh/USY. sites also act as active sites for the formation of carbonaceous deposit which brings about catalyst deactivation. Therefore, it is necessary to control the strength and number of strong Br6nsted acid sites in Rh/USY in order to prepare highly active Rh/USY catalyst with long catalyst life. The modification of Rh/USY with alkali metal salts such as NaOH, NaNO3, Na2CO3, NaCI, etc. was, therefore, performed in order to control the strength and number of strong Br6nsted acid sites of Rh/USY catalyst. It was revealed that the catalyst deactivation of Rh/USY was remarkably improved by the addition of small amount of alkali metal salts. Modification with NaOH was the most effective and 0.5wt% addition of Na using NaOH was optimal amount for the improvement of the catalyst deactivation with reaction time as shown in Figure 4. It is evident that 0.5wt%Na-5wt%Rh/USY catalyst shows higher and more stable catalytic activity for the hydrodesulfurization of thiophene than 5wt% Rh/USY and CoMo/AI203 catalysts. Therefore, it can be concluded that there is a possibility of usage ofNa- Rh/USY as highly active second generation hydrodesulfurization catalyst for petroleum feedstocks. 3.3 Mechanism of hydrodesulfurization of thiophene on Rh/USY catalyst It was revealed that Rh/USY showed higher catalytic activity than commercial CoMo/Al203 catalyst in the hydrodesulfurization of thiophene. We also studied the mechanism of hydrodesulfurization of thiophene over RH/USY catalyst. As mentioned above, Rh/USY catalyst acts as bifunctional catalyst for the hydrodesulfurization of thiophene, in which both Br6nsted acid sites of USY and Rh in RH/USY catalyst act as active site.

556

30

A B Rh/Quartz+USY

20 o

-v,=(

9

C,r

9

;> o L) 10

0 lg Rh/Quartz (A) O.lg ....

0

I

1

30 60 90 T i m e o n S t r e a m (min)

I,

120

Figure 5. Hydrodesulfurization of thiophene over Rh/quartz(A), USY(B)and mechanically mixed (Rh/quartz(A) + USY(B)) catalysts at 400~ Furthermore, it was assumed the existence of spillover hydrogen in the hydrodesulfurization of thiophene over RhAJSY catalyst. Thus, we tried to confirm the existence of spillover hydrogen in the hydrodesulfurization of thiophene over RH/USY catalyst. The catalytic activity of Rh/SiO2(quartz) mixed mechanically with USY in the hydrodesulfurization ofthiophene was examined. It was found that the activity of mixed catalyst obtained experimentally was higher than that calculated theoretically as shown in Figure 5. This implies that there exists the spillover hydrogen on Rh/USY catalyst in the hydrodesulfurization of thiophene. Therefore, we Propose a possible mechanism for the hydrodesulfurization of thiophene over Rh/USY catalyst as shown below; In this mechanism, thiophene is adsorbed on the Br6nsted acid sites and hydrogen is activated on Rh to form spillover hydrogen. The spillover hydrogen formed on Rh attacks the reaction intermediate like S=C=CH-CH=CH2, which is formed by the decomposition of thiophene adsorbed on the strong Br6nsted acid sites of H- zeolite [ 16]. On the basis of the proposed mechanism, it can be possible to develop much more highly active zeolite-based hydrodesulfurization catalysts for petroleum feedstocks. 4. CONCLUSION It was revealed that Rh/USY showed higher catalytic activity than commercial CoMo/ A1203 in the hydrodesulfurization of thiophene. The catalyst deactivation of R h ~ S Y with

557

H~ Ti

Spillover Hydrogen

H

H

eta

Tl

+ Ca~C4 Hydrocarbon l

H

'

S\

1

0

H§ [S=C=CH-CII=CH 2 ] <

o

I lr..~nrm.~iii.~*.i~.r~ri~*.:r~:m.T~rim~iiiiiiirm..rij:.~i~'~iii~r~iii~i~i.ii~ii~i~iFii......p~iiiiii~r..!.i~i..iriiii.im..%...ir.T.rm.r.i.!i~.iiii....mrii...iiL..irrr.:iiii.i!i~iiii : ::: :: :::.:" :S::.$.::." : : : : : ~ : . | | ; | , l

$:: ".l l l;: : : : S : | l : : : .;:l:l;||".'l~.;t|zz~:~:::.': 7. ~ . : : ~ .:h'~r

( Rh )

.: le:'~ ..:.: :,, . ' : . : : . . , . . . : . ::..~. : : : : : : 7 K : :::.:: : : : . : : : : : :: :.: :: ::: :. :: :::: : : : : : : : : : : : : : : : : : ; : : : : : : : :: ~~:": :.':::: !: :: ! :" ~: :~: ::: :::

( USY )

Scheme 1. A Possible mechanism for the hydrodesulfurization of thiophene over Rh/USY catalyst. reaction time was remarkably improved by the addition of small amount of NaOH. Therefore, there is a possibility of use of Rh/USY modified with NaOH as highly active second generation hydrodesulfurization catalyst for petroleum feedstocks.

Acknowledgment A part of this work has been carried out as a research project of The Japan Petroleum Institute commissioned by the Petroleum Energy Center with the subsidy of the Ministry of International Trade and Industry. REFERENCES 1. 2. 3. 4. 5. 6. 7.

M. L. Vrinat, C. G. Gachet and L. de Mourgues, Stud. Surf. Sci. Catal., 5 (1980) 219. C. S. Brooks, Surf. Technol., 10 (1980) 397. K. E. Givens and J. G. Dillard, J. Catal., 86 (1984) 108. T. G. Harvey and T. W. Matheson, J. Catal., 101 (1986) 253. R. Cid, F. Orellana and A. L. Agudo, Appl. Catal., 32 (1987) 327. Y. Okamoto, A. Maezawa, H. Kane and T. lmanaka, J. Mol. Catal., 52 (1989) 337. S. Gobolos, M. Breysse, M. Cattenot, J. Decamp, M. Lacroix, J. L. Portefaix and M. L. Vrinat, Stud. Surf. Sci. Catal., 50 (1989) 243. 8. M. Laniecki and W. Zmierczak, Zeolites, 11 (1991) 18. 9. P. Kovacheva, N. Davidova and J. Novakova, Zeolites, 11 (1991) 54. 10. M. Sugioka and K. Aomura, Intern. Chem. Eng., 13 (1973) 755. 11. M. Sugioka and K. Aomura, Bull. Japan Petrol. Inst., 17 (1975) 51.

558 12. M. Sugioka, T. Kamanaka and K. Aomura, Prepri. Am. Chem. Soc., Div. Petrol. Chem., 24 (1979) 740. 13. M. Sugioka and K. Aomura, Prepri. Am. Chem. Soc., Div. Petrol. Chem., 25 (1980) 245. 14. M. Sugioka and K. Aomura, J. Japan Petrol. Inst., 26 (1983) 216. 15. M. Sugioka and K. Aomura, J. Japan Petrol. Inst., 26 (1983) 362. 16. M. Sugioka, J. Japan Petrol. Inst., 33 (1990) 280. 17. M. Sugioka, Y. Takase and K. Takahashi, Proc. of JECAT'91, p.224 (1991). 18. M. Sugioka, Zeoraito(Zeolite), 10 (1993) 121. 19. M. Sugioka, Erd61 & Kohie, Erdgas, Petrochemie (1995), in press.

Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All rights reserved.

559

HIGH-DISPERSED SUPPORTED CATALYSTS ON BASIS OF MONODISPERSE PT-SOLES IN PROCESSES REDUCTIVE OF TRANSFORMATION OF HYDROCARBONS N. A. Zakarina and A. G. Akkulov

Institute of Organic Catalysis and Electrochemistry, National Academy of Sciences, 142 Kunaev st., Almaty-480100, Repubfic of Kazakhstan ABSTRACT Hydrogen adsorption and catalytic activity of ml203,supported monodispersed Pt-soles (AP) and the same catalysts promoted with Re (APR) have been studied in gas-phase benzene hydrogenation. The tennoprogrammed desorption (TPD) of H2, adsorption-calorimetry, electron microscopy and X-ray photoelectron spectroscopy methods have been used. The aggregation of colloidal Pt-particles supported on 3t-A1203 increased with decreasing of the particle size from 3.2 to 0.6 nm at the redox treatment of catalysts. The correlation between strongly bound hydrogen species and catalytic activity of the AP- and APR-catalysts for investigated reaction has been found independent of the method of preparation. 1. INTRODUCTION AP- and APR-catalysts are widely used in reforming plants to produce aromatic hydrocarbons and high-octane gasoline [ 1,2]. It is known that the decrease of the activity and selectivity of these catalysts is observed with the decrease of dispersity of supported active metal. Therefore, the investigation of properties of metal particles at the different stages of catalyst preparation has great practical interest. The nature of the initial metal compound has influenced on the most important catalysts' properties (electronic state, dispersity, particle size distribution), which depend the chemical reactions carried out during catalyst preparation [36]. The use of Pt-hydrosoles with narrow distribution of particle size for catalyst synthesis gives the opportunity to regulate the dispersity of active metal and prepare catalytic systems with unique properties. 2. EXPERIMENTAL Pt-soles with different dispersity (dav.=0.6; 1.5; 2.1; 3.2 nm) have been obtained by variation of temperature and time of liquid phase reduction of water solution of HzPtC16 by sodium citrate in inert atmosphere [7]. AP catalysts have been prepared by impregnation of ~/-A1203 with water solution of H2PtC16 in air (industrial method (I)), in hydrogen atmosphere (method II) and by impregnation of Al/O3 with monodispersed Pt-soles (method III). Three methods of APR-catalyst preparation have been used: by impregnation ofT-A1203 with a mixture of water solutions of HzPtC16 and HReO4 in air (industrial method (I)); by treatment of APC(II) catalysts with Re ions in hydrogen atmosphere (method II); and by

560 impregnation of y-ml203 with mixture of Pt-soles and n R e O 4 water solution in air (method III). Platinum content in all catalysts was 0.36 wt% and rhenium was varied from 0.2 to 0.6 wt%. Alumina with different porous structures (predominant porous radius was 7.0 nm for A1-support and 6.0-10.0 nm for A-64) were used for catalyst preparation. All AP- and APRcatalysts were exposed to high temperature redox treatment (anneal at 500 ~ in dry air flow and then reduction of catalysts at 500 ~ in hydrogen flow). Electron microscopy method (EM-125 K, 120000 x magnification) was used to estimate metal particle sizes in soles and on supports. Samples were prepared by replic method. Distribution of metal particle sizes over the catalyst surfaces were determined by size statistic processing of 1500 -2000 particles. Measurement of H2 activated and inactivated adsorption, and H2 adsorption heats on AP- and APR-catalysts was carried out on the adsorption-calorimetric apparatus consisting of a vacuum part and microcalorimeter. The destination vacuum part was used for vacuumthermal purification of catalyst surface and measurement of H2 portions during adsorption or desorption. The DAK 1-1 microcalorimeter was used for registration of adsorption heats of H2 at 35~ After vacuum-thermal cleaning of the surface (P
3.1. Pt/AI203- Catalysts. High dispersed metal on the surface of A-64 support is obtained by high-temperature treatment of APC. Pt-crystallites with the most clear face are observed on the surface of the catalyst prepared by method II. About 80% of particle sizes are varied from 1.5 to 2.5 nm in the samples obtained by methods I and II. There were particles of mostly two different sizes on the surface of the catalyst prepared by method III (dav.Pt-sole = 1.5 nm). The results of electron microscopy investigation of APC are in good agreement with the calculation of dispersity

561

"~2 4 0

2.0

4.0

d initial, nm

Figure 1. Degree of aggregation of Pt particles on ]t-A1203 (support A-1) aider redox treatment APC(III) versus average particle size Pt in sole.

based on low-temperature (35~ chemisorption of hydrogen. The average sizes of crystallites in catalysts prepared by methods I, II, III aider redox treatment were 1.9, 1.7, and 3.8 nm, respectively. The variation of the average Pt-particle size of soleused for impregnation leads to the change of its crystaUite structure and electronic state of separate crystallite composing atoms. As the result of the extent of interaction of the crystallites with support was the change of their following behaviour on it. The most intensive aggregation was observed for Pt particles with 0.6-1.5 nm initial sizes (on the surface of alumina, it became day.=2.6-2.8 nm) atter thermal redox treatment of APC (III), supported on A-1. The changes of the larger Pt particles were insignificant (Fig. 1). For example, Pt-sole with 3.2 nm particles supported on alumina formed crystallites with 3.4 nm sizes. According to TPD data, the H: chemisorption on APC(I) and APC(II) at 500~ was characterized by a rather broad energy spectrum with one form of adsorbed H2, which is desorbed at 200~ (Fig.2). These catalysts were prepared from water solutions of HzPtC16. Another picture was observed in the case of APC(III) synthesized from Pt-soles with different dispersity. As it was shown by the TPD method, two hydrogen chemisorption species with desorption peaks at 150 and 400~ appeared in spectra. According to the prevailing mechanism of the ionic exchange, the more monodispersed active centers in APC(I and II) as compared with APC(III) are observed due to the anchoring of the [PtCI6]2 anions on the surface of alumina and formation of the outersphere >AI+[PtCI6]2"; >AI(OH)2+[PtCI6]2 associates. Formation of two different types of chemisorbed H: on APC(III) can be explained by the appearance of two different types of active centres, which are formed due to both mixed (ionic exchange and ligand replacement) mechanism of interaction of surface functional groups of A1203 with different Pt-particle dispersity and location of particles inside support's pores with different diameters [8]. These results are also confirmed by the microcalorimetric volume adsorbed on the APC(I and II) are greater than on the APC (III) and this hydrogen is more energetically homogeneous (Fig.3). In the case of APC(I) and APC(II), the more extended parts of qm Q curves with constant H2 adsorption heats and more gentle sloping were observed compared with APC(III). The presence of the extended part with constant heat (qH, = 90 kJ/mol) for APC(II) are proven by the homogeneous nature and size of about 50% of active centres of H2 adsorption. -

562 25,

~

.

.

.

.

.

.

.

20 "

2

4

167 1

zlO

b

60

300

480

-

2

8

16

Degree of coverage, mol H2" g pt-1(10-4)

Temperature, ~ C

Figure 2. Effect of Preparation Method: Thermodesorption of H2 over APE (support A-64). (1-APE(I); 2-APC(II); 3-APC(III, 0.6 nm Pt-sole); 4-APC(III, 1.5 nm Pt-sole)).

Figure 3. Effect of Preparation Method: Differential heats of H2 adsorption over APE (support A-64). (1-APE(I); 2-APC(II); 3-APC(III, 0.6 nm Ptsole); 4-APC(III, 1.5 nm Pt-sole).

Table 1. Properties of APE prepared by different methods (Tann.=500~ Tred.=500~ Preparation Method of APE I II III*

H/Pt (500~

H/Pt (35~

1.3 1.1 1.4

0.62 0.67 0.31

dav" (chem), nm 1.9 1.7 3.8

dav" (el.micr), nm 1.5 - 2.5 1.5 - 2.5 2.0 &5.0

S, m2/gPt

Econ.4ds/2 Pt, ev

npt/nAl

146.2 150.0 73.7

317.3 315.0 315.5

0.046 0.030 0.020

*for APC(III) obtained from 0.6 nm Pt-sole

According to X-ray photoelectron spectroscopy, the surface atoms of Pt in APC(II and III) were in a more reduced state than in APE(I) (Table 1). On the APC(III) prepared from Pt-sole with dav.=l.5 nm, the platinum is present in the forms ofPtn § and Pt ~ with prevailing ofPt ~ species. Conditions of the synthesis of APC(II) and APC(III) promote this phenomenon. Reduced state of Pt prevented strong interaction of Pt-atoms with support. It was found that destruction of surface structures depends on the strength of interaction of metal with support. The following tests of APE with different genesis in benzene hydrogenation showed that APC(III) prepared from Pt-soles with dav.=l.5-2.1 nm on y-AI2Os and average radius of pores 6.0-10.0 nm are more active than APC(I and II) (Table 2 and 3). The highest activity per m 2 is also observed on the same catalysts.

563 Table 2. Hydrogenation of benzene on APC(I & II). Method of preparation of

........................................Act!.v!t.y...of..~.C..at...1.5.0..~ ........................................

APC

mol C6Hlz/mol Pt-s

mol C6Hlz/m2"s(10"6)

I

0.48

16.9

II

0.74

23.9

Table 3. Hydrogenation of benzene on APC (III). Activity of APC The average size ofPt crystallites .........................m.o!..C6H!z./mo!..Pt.~s.......................................m o!..C6H.~..z./.m2?.s.(..1.0.~.).................... in sole, nm 150~ 180~ 150~ 180~ support A- 1 0.6 0.31 0.46 15.2 21.9 1.5 0.40 0.52 22.0 28.0 2.1 0.50 0.57 26.0 30.0 3.2 0.23 0.43 14.0 26.9 support A-64 0.6 0.81 1.06 56.4 73.8 1.5 1.02 1.19 80.1 93.4

3.2. P t - R e / A I 2 0 3 - catalysts.

No effect of Re on the dispersity of AP-catalysts has been shown by electron microscopy investigation of catalysts atter high temperature redox treatment [(excepting APRC(II)]. However, the density of particles on the surface of the catalyst was increased. The addition of Re to the catalyst independent of the method of preparation resulted in increase of Hads.(500~ (Fig.4) and decrease ofHad~(35~ (Fig.5). The differential heats of H2 adsorption were decreased under these conditions. The increase of activated extra additive adsorption of hydrogen occured mainly due to those forms of chemisorbed H2 that are desorbed at above 300~ This tendency is the most significant for APRC(III). The peak of desorption of this form of chemisorbed H2 was shitted in the more high-temperature region (Fig.4c). Higher thermostability of Pt-Re-catalysts was displayed in the forms of activated H2 adsorption. For example, the decrease ofHa~ amount (500~ with the increase of temperature treatment from 500 to 700~ for APC(I) was 6.4 times, and for APRC(I) with 0.36 wt% Re content, the increase was 1.4 times. It was supposed that the crystallisation of large platinum particles was hindered by the presence of Re in the catalyst. The further increase of activated adsorption and decrease of low-temperature H2 adsorption has been observed with increasing of Re content to 0.6 wt% in the catalyst. Results of the physic-chemical studies of APRC prepared by different methods are

564

20L

a ~

1. i

16 12

0

0

b

3

I

I

~2e

12 ~

a

i-""Tl~ i

. ~

i. "i

b

8

42

~

O"

I

"~. I

I

--l-"s~ I

i"

i

~L

-o

"~ E

40 30

C

3 I

20 10 0

z

o"

-

C

126 -

I

PI~~~~ 60

300

480

Temperature, o C

Figure 4. Effects of method of preparation and content of Re in catalyst: Thermodesorption of 1-12 over APRC (support A64). (a- APRC(I); b- APRC(II); r APRC(III)*; 1 - APC; 2 - 0.20 wt% Re; 3 - 0.36 wt% Re; 4 0.60 wt% Re. *-used 1.5 nm Pt-sole)

0

--

2

6

10

16

Degree of coverage, mol H2-& pt-l(ll) 4)

Figure 5. Effects of method of preparation and content of Re in catalyst: Differential heats of H2 adsorption over APRC (support A-64). (a- APRC(I); b- APRC(II); cAPRC(III)*; 1 - APE; 2 - 0.20 wt% Re; 3 - 0.36 wt% Re; 4 - 0.60 wt% Re. *-used 1.5 nm Pt-sole)

summarized in Table 4. The composition of active phase was the same as for catalysts used in industry. Influence of the dispersity of initial Pt-soles on the amount of activated-adsorbed H: and the ratio of different 1-I2 forms on APRC(III) depended on the Re concentration in the catalyst. For example, on APRC(III) with 0.2 wt% Re content, the total amount of Haas.(500~ decreased from 4.7x10 "4 to 30.9x10 "4 mol H2/g Pt with theincreasing of sole particle size from 0.6 to 3.2 nm. At the same time, the amount of Hads.(500~ increased from 41.6x10 -4 to 60.0x10 4 mol H2/g Pt for APRC with 0.5 wt% Re. It indicated the change of the electronic state of Pt and Re atoms during variation of these parameters. It has been proposed that it is possible to obtain optimal ratio of these metals to produce effective catalyst for the conversion of benzene into cyclohexane. The most active catalysts are APR with 0.36 wt% Pt - 0.36 wt % Re ratio, usually used in industry. These catalysts has been prepared by method III, using ?-AI:O3 with preferential pore radius 6.0 - 10.0 nm. Average dispersity of the Pt sole was 1.5 nm. The activity of platinum-rhenium catalysts depends on the methods of preparation and increase in the raw: I < II < III. The activity of catalysts was shown 1.61, 1.70, and 2.25 mol C6H~Jmol Pt's, respectively, at this range (Tables 5 & 6).

565 Table 4. Properties of APRC (0.36 wt %Pt-0.36 wt %Re), prepared by different methods (Tan.=500~ Tred.=500~ Prepar. H/Pt Method of (500~ APRC I 2.34 II 2.05 III* 2.76 *-used 1.5 nm Pt-sole

H/Pt dav S, (35oc) (chem.) m2/gPt

Econ. electron, eV Pt 4d5/2 Re 4f7/2 317.6 43.7 316.9 44.5 314.7 44.7

Relative atom. concentration Pt/AI Re/A1 0.019 0.011 0.007 0.003 0.030 0.020

........................................................................................................... a m

0.35 0.37 0.15

3.4 3.2 7.8

82.7 85.6 35.9

Effects of the dispersity of used sole on the activity of the APRC(III) and APC(III) have extreme character. The maximum of the activity of the catalyst with the rhenium content within 0.2 -0.36 wt % in tested reaction corresponded to sole with the average dispersity of 1.5 nm (Table 6). It might be suggested that the reason for the increase of catalytic activity of the best APRC(III) was the change of electronic and structural characteristics of the small Pt clusters, stabilizated by low valence rhenium ions. Thus, the XPS-investigations showed that the addition of Re into the catalyst leads to the essential change of the character of electron spectra. In the range of APR-catalysts with 0.36 wt % Re concentration, prepared by methods I II, and III the degree of reduction of Pt surface atoms increased. XPS investigation showed that in catalyst prepared by impregnation of alumina (predominant porous radius 6.0 - 10.0 nm) with Pt-sole ( day= 1.5 nm) and HReO4 solution, the surface Pt-atoms were mostly in zero valence state, and Re-atoms were in low-valence state. Probably, in this case, other conditions of catalyst genesis and optimal ratio of Pt and Re (1:1) concentration at given Pt-sole dispersion promoted the shift of the electron density from Re-atoms and lattice oxygen to Ptatoms and more complete reduction of Pt after high temperature redox treatment. It was supposed that the formation of Pt zero valence cluster was stabilised by lower valence Re ions. The probability of its existence was confirmed by character changing of Pt and Re concentration on the surface at change of their contents in catalyst. Table 5. Hydrogenation of benzene on APRC(I & II). Method of preparation of APRC

I

Content of Re in catalyst, wt%

0.20 0.36 0.60 0.20 0.36

Activity of APRC at 150~ .............................................................................................................................

mol C6Hlz/mol Pt.s

mol C6Hl//m2-s(10"6)

1.09 1.61 1.66 1.49 1.70

72.3 99.8 120.0 68.0 102.0

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

II

566 Table 6. Hydrogenation of benzene on APRC(III). The average size Activity of APRC(III) of crystallites in .........mo!..C6H!..z./.mol..Pt:.s............... .mo!...C6H.L2/.m...z.-..s...(1.0-~) ....... Pt-sole, nm 150~ 180~ 150~ 180~ 0.20 wt% Re (support A-1) 0.6 0.86 1.03 79.8 95.7 1.5 0.88 1.11 136.3 172.6 2.1 0.71 0.82 104.3 120.5 3.2 0.29 0.36 34.2 43.4 0.36 wt% Re (support A-64) 0.6 2.15 2.46 292.4 334.6 1.5 2.25 2.88 321.4 411.4

=

2.4

2.8 -

60-500~

9 Q -

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~ (/I

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-

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T~RH2, mol/g Pt (10"4) Figure 6. Activity of APC and APRC of different genesis vs. amount of desorbed H 2 curves. (APC: A- method I; !-I- method II,; O- method III; APRC: painted symbols). Comparison the data of activated H2 adsorption on AP- and APR-catalysts with activity of these catalysts in benzene hydrogenation showed correlation between amount of strongly bounded hydrogen forms (Enm) and catalytic activity (Fig.6). REFERENCES

1. M.A.Ryaschentseva and Ch.M. Minachev. Re and Its Compounds in the Heterogeneous Catalysts, Moskva, Nauka, 1983, p. 248

567 2. R.W.Joyner and E.S.Shpiro. Catal.Lett., .9, No.3-4, (1991) 239-244. 3. A.F.Flores, R.L.Burwall, and J.B.Butt., Chem. Soc. Faraday Trans. 88 (1992) 1191-1196 4. F.L.Marvin, V.M.LeRoy. J.Catal..35 (1974) 434-440. 5. B.B.Garkov, A.Z.Rubinov, S.V.Schapoval and J.D.Jakovleva. Zhurnal Fizicheslkoi chimii (j.Phis.Chim.), No.7 (1990) 1783-1788. 6. S.Engels, E.Hernold, No.3, (1992) 100-103.

H.Mayer,

H.Meinerg and H.Lausch. Chem.Tech. (DDR) 44,

7. K.Aika, L.L.Ban, I.Okura and J.Turkevich, J.Res.Inst.Catal., No.1(1976) 54 8. I.E.Smirnova, A.S.Beliy, M.D.Smolikov and V.K.Duplyakin, Kinetica i katalys (Kinetics and Catalysis-in Russian) .31 (1990) 686.

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Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All fights reserved.

569

INFRARED SPECTROSCOPY OF CO/E[2 COADSORPTION ON Ni/A1203 H Y D R O T R E A T I N G CATALYSTS: EVIDENCE F O R PERTURBED M E T A L SITES M. I. Zaki

Chemistry Department, Faculty of Science, Kuwait University, P.O. Box 5969-Safat, 13060Kuwait ABSTRACT In-Situ infrared spectroscopy was implemented to probe carbonyl species formed in adsorbed CO on 10 wt% Ni/A1203 catalyst at 160-300 K. The results characterize terminal (at vCO = 2060-2035 cm~) and bridging (<2000 cm~) carbonyls established on aggregates of Ni ~ metal sites. They also characterize additional terminal-CO exhibiting vCO frequency in uncommonly high range (2095-2075 cm-~), which is nevertheless sufficiently lower than the frequency range (2135-2120 cm~) at which vNi+C-O is observed. The high-frequency carbonyl has been considered manifesting perturbed CO/Ni ~ adsorptive interactions with electron-deficient metal sites. Causes of the metal electron drain are discussed on basis of metal/support interactions, considering the electronically "hard" nature of the alumina support. 1. I N T R O D U C T I O N Supported nickel catalysts have been the focus of attention of a number of infrared (IR) spectroscopic studies adopting CO as a surface probe molecule. This research has been motivated by the pronounced hydrotreating activity of the catalysts in processes of industrial importance [1,2]. Results published before 1978, have been critically assessed and compared to results then available on Ni ~ metal thin films and single crystal surfaces by Sheppard and Nguyen [3 ]. Accordingly, the following observations can be highlighted. First, irreversible CO adsorption at 300 K on completely reduced Ni ~ surfaces gives rise to two vCO absorption bands at 2045-2025 and 1950-1930 cm ~, due respectively to terminal and two-fold (132) bridging carbonyl species. Second, the terminal-CO bonding to the surface is slightly weakened and the vCO frequency thus shifted to a higher range (2065-2050 cm-~), as low site-density Ni ~ crystal planes become dominantly involved [4]. Third, on isolated or non-crystalline Ni ~ sites (i.e., metal sites next to oxygen atoms or oxide support), the terminal-CO bonding is rendered even weaker and a further high vCO frequency shift to 2100-2070 cm ~ occurred [5-7]; these carbonyls are referred to below as terminal-CO*. Fourth, vCO-absorption of the (132)bridged carbonyls suffers a slight high-frequency shift (from 1950-30 up to 1975-50 cm 1) upon adsorption on partially oxidized metal surfaces [8]. Whereas, the shift is to the low frequency range of 1911-1810 cm 1 when high site-density surfaces are involved [9], or following conversion into the three-fold (133) configuration [ 10,11 ].

570 Systematic IR studies of CO adsorption on supported nickel catalysts [3,12] have described the surface sites of the terminal-CO* as being electronically perturbed. The perturbation has been attributed [3,12] to oxygen-mediated electron-exchange interactions between Ni~ + couples established on incompletely reduced catalysts. This type of perturbation is referred to below as O-perturbation. Similar terminal-CO* species were also observed at 2095-2080 cm ~ by Primet and Sheppard [ 12], and Galuzska and Amenomiya [13], while examining CO and 1-12co-adsorption on supported Ni ~ These authors ascribed the perturbation in this case to a hydrogen-induced (H-induced) change in the vCO frequency on a particular site, rather than to a change of site. This type of perturbation is designated below as H-perturbation. Blackmond and Ko [ 14], investigating CO and CO/H2 adsorption on completely reduced Ni/SiO2, have also observed the emergence of a terminal- CO* type of vCO band at 2090-2080 cm ~. These authors have, however, assigned the band to multi-CO species bound to defect Ni ~ sites, i.e. Ni(CO)x and x>l. In the past decade numerous studies of interactions in reactive [ 15-18] and non-reactive [ 19-27] co-adsorption of CO and 1-12on Ni ~ single crystal surfaces have been performed, using a range of ultra-high vacuum (UHV) analytical techniques. Within this context, HREELS studies [ 19,20,26,27] have observed on Ni(100) surface a H-induced vCO high-frequency shift (up to 2100-2080 cm ~) for terminal-CO species. A similar H-perturbation to that encountered on the high-area catalysts was concluded [12-14]. Hence, a comparison between IR spectral features of CO adsorbate on surfaces of Ni ~ supported particles and self-supporting single crystals should help elucidating adsorption sites exposed on the catalysts. To justify such a comparison, genesis of catalysts containing large metal crystallises of extensive facets must be ensured. This experimental approach was pioneered by Pritchard et al [28] for Cu ~ Accordingly, the present investigation employed a heavily loaded Ni/A1203 catalyst (10 Ni% by weight) prepared by H-reduction at a higher-than-normal temperature (873 K). 2. EXPERIMENTAL IR spectra were taken from the "catalyst + adsorbed CO" over the vCO frequency range 2300-1700 cm ~, using a model 580B PERKIN-ELMER spectrophotometer equipped with a model 3500 P-E data station for spectra acquisition and manipulation. The spectra were signal ratioed and obtained with a slit programme yielding a maximum resolution of 5.3 cm "~ acquired at 1 point per cm ~ with data acquisition time of 1.6 s/cm "~. Spectra of the "adsorbed CO" were obtained by subtracting the "catalyst" background spectrum taken under identical pretreatment and spectroscopic conditions. The 1R-Cell capable of operation at 120-1400 K and equipped with CaF2 windows used in this study was that devised and described previously by Muha et al. [29]. The catalyst parent material is deposited by spraying onto a tungsten grid which is held rigidly by nickel clamps through which controlled electrical heating power may be conducted to the grid. In addition, the grid and, hence, the catalyst can be cooled using VN2. The catalyst temperature is measured by chromel/alumel thermocouple spot-welded to the top central region of the grid. The grid support is held in the center of the stainless steel cell body containing ports for gas delivery and for admission of the IR-beam.

571 The stainless steel gas/vacuum handling system used for this work facilitates a base pressure of 5x108 Torr (1 Torr = 133.3 Pa). It is equipped with a t-N2 cooled zeolite sorption pump, a 30 L/s ion pump, a BARATRON capacitance manometer (0.001-1000 Torr), and a model M100M DYCOR quadrupole mass spectrometer. The catalyst parent material consisted of nickel nitrate impregnated alumina. The support was DEGUSSA aluminium oxide C (104 mZ/g) and the precursor was ALPHA ultrapure Ni(NO3)2.6H20. The amounts required of these materials to obtain 10 wt% Ni/A1203 were added simultaneously into an appropriate volume (10 ml/g-support) of a liquid mixture of water and acetone (1:9 volume ratio), and the resulting suspension was agitated ultrasonically for 30 min. The slurry thus obtained was uniformly sprayed by a N2-pressurized atomizer, onto the entire exposed grid area (5.2 cm2). During spraying, the grid was electrically heated to 323-333 K to flash evaporate the liquid phase [30]. The net weight of the material sprayed onto the grid was 40.4 mg (= 7.8 mg/cm2). The catalyst (Ni/A1203) was prepared inside the cell by heating in vacuum at 473 K for 15 h, and reducing at 873 K with three successive exposures of H2 (using 10 Torr H2 for the first two exposures and 50 Torr for the last one) each followed by 10 min evacuation of the gas phase at the reduction temperature prior to cooling to 160 K under dynamic vacuum. A hydrogen covered catalyst (H/Ni/A1203) was obtained by, first, cooling to 160 K in the presence of H2 (g), and, second, outgassing at 160 K for 10 min. Carbon monoxide (99.99% pure) and hydrogen (99.995% pure) were used as obtained from MATHESON gas products. 3. RESULTS 3.1 Carbonyi spectra from CO/Ni/AI203 at 300 K

Spectrum (a), Fig. 1, shows that in presence of 40 Torr of CO gas phase the adsorption on Ni/A1203 at 300 K gives rise to two vCO absorption bands in the bridging-CO frequency region (<2000 cm ~) at 1935 and 1875 cm "~, and a band at 2085 cm ~ and a shoulder at 2045 cm ~ in the terminal-CO region (>2000 cml). Following outgassing at 300 K, spectrum (b), Fig. 1, was obtained. It indicates persistence of the bridging-CO bands but with slight highfrequency shiits to 1940 and 1980 cm ~, respectively. It also indicates disappearance of the terminal CO band and conversion of the shoulder into a weak, but stable, band at 2035 cm ~ The instability of the high-frequency terminal-CO may reflect weak zc-contribution to the known 6-zr bonding system of CO to metal sites (2). The zc-bonding involves electron-rich sites and results in vCO low-frequency shit, s. 3.2 Carbonyl spectra from CO/Ni/A1203 at <300 K Fig. 2 displays the spectra taken at 160 K from the catalyst following CO adsorption at 160 K and subsequent outgassing at different temperatures in the range 160-275 K. Spectrum (a) indicates that the application of the low temperature stabilizes and enhances the absorption of the high-frequency terminal-CO (now at 2090 cm"1) against outgassing. The shoulder signifying the low-frequency terminal-CO remains observed but at a slightly higher vCO frequency (2054 cm ~) than at 300 K. The low-frequency bridging-CO (at 1863 cm "~) is suppressed relative to the high-frequency one (at 1945 cm~), and a weak, but distinct, band emerged at 2123 cm 1.

572 CO/Ni/AI:,O3

CO/Ni/Ai~,O3

u)

Tad.. = 3 0 0 K T=.,,., = 3 0 0 K

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[

2100

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2000

1900

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1800

2200

.~

Figure 1. Room-temperature IR carbonyl spectra taken from CO/Ni]A1203 in presence (a) and absence (b)of the gas phase.

2100

1__ 2000

1900

1800

W a v e n u m b e r / c m -~

Figure 2. Low-temperature IR carbonyl spectra taken from CO/Ni/A1203 following CO adsorption at 160 K and subsequent desorption at the temperatures indicated.

Upon increasing the temperature, under dynamic vacuum, the following changes are observed in the spectra (b-d) compared in Fig. 2. At 180 K, the 2123- cm 1 band is eliminated; and at 225 K, the low-frequency bridging-CO band grows notably stronger and shi~s to 1848 cm 1. At 275 K, the terminal-CO band (at 2090 cm~) weakens markedly and shiPts down to 2072 cm 1, the bridging-CO band at 1848 cm~ also weakens, and a new bridging-CO band emerges at 1872 cm ~ (probably at the expense of the 1848 cm1 band). At 300 K, a spectrum (not shown) similar to spectrum (b) in Fig.1 was obtained. Thus, it revealed complete disappearance of the high-frequency terminal-CO band (at 2090-2072 cm'~), a further shifting of the low-frequency terminal-CO band from 2045 down to 2035 cm 1, and dominance of the two bridging-CO bands at 1940 and 1890 cm1. 3.3 Carbonyi spectra from CO/H/Ni/A1203 at 160-300 K

CO adsorption on H-covered Ni/A1203 at 160 K gave rise to a largely similar carbonyl spectrum (a, Fig. 3) to that exhibited by CO adsorption on the H-free catalyst at the same temperature (b, Fig. 3), except for two differences. These are (i) a considerable absorption intensification of the high-frequency terminal-CO band (now at 2095 cm'l), and (ii) a slight frequency mismatch between the high-frequency bridging-CO bands. At higher temperatures the spectral features monitored were very similar to those observed on the H-free catalyst (Fig. 2).

573 Lib

=o T ....

=

160 K

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160

K

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2200

2100

I

I

2000

1

I

I

1900

w a v e n u m b e r / c m "1

1800 ,

Figure 3. IR carbonyl spectra taken at 160 K from CO adsorbed on hydrogen-covered (a) and hydrogen-free (b) catalyst surfaces. 4. DISCUSSION 4.1 Nature of CO adsorbed on Ni/A1203 Table 1 summarizes vCO frequencies, assignments and characteristics of surface carbonyl species established following CO adsorption on the catalyst. The results indicate dominance of the two (132)bridging-CO species, in contrast to the low-frequency terminal-CO species, at 300 K and 10 "7 Tort. Presence of the gas phase is shown to be essential for stabilization of the high-frequency terminal-CO* species at 300 K (Fig. 1). However at _<275 K the terminal-CO* species are stabilized against outgassing (Fig. 2), and intensified markedly in presence of co-adsorbed hydrogen (Fig. 3). 4.2 CO-probed characteristics of the metal sites Bridging-CO. Bridging-CO species is multi-centered and a common feature of CO adsorption on various surfaces of metal single crystals [3,8,10]. Thus, formation of bridgingCO implies presence in the catalyst ofNi ~ metal particles. Indeed, a previous XRD analysis of the catalyst [31] determined a metal average crystallise size of 25+3 nm. The occurrence of (132) bridging-CO species of different stretching frequencies (Table 1) means that the metal particles expose differently-packed crystal facets. The frequency values and relative intensifies displayed (Table 1) may account for dominance of the low site density crystal facets Ni(100) and Ni(110). The high-frequency bridging-CO (at 1945 cm1) is stable to temperature elevation from 160 up to 275 K (Fig. 2). In contrast, the low-frequency band is shown to occur weakly at 5180 K (at 1863 cm]), and to grow stronger with a slight shift down to 1848 cm ], near 225 K. At 275-300 K, it shit, s up to 1890 cm~. In view of HREELS results [4], the high frequency (1945-35 cm ~) indicates carbonyls bridging metal sites on Ni(100)-like surfaces.

574 Table 1 oCO frequencies, assignments and characteristics Ni/A1203. u/cm ~ 1890-48

Assignment (132) bridging-CO

1952-35

(132) bridging-CO

2054-35

terminal-CO

2095-72

terminal-CO*

for surface carbonyls established on

Characteristics occurs irreversibly at 160-300 K, and grows more populous the higher the adsorption temperature. occurs irreversibly at 160-300 K, and grows more populous the lower the adsorption temperature. occurs reversibly at 160-300 K, assuming lower stretching frequencies the higher the adsorption temperature. occurs reversibly at 300 K, and irreversibly at <300 K. It grows stronger the lower the temperature, and much stronger on H-covered surfaces.

Terminal-CO. Room-temperature stable terminal-CO species exhibiting vCO at 20452020 cm -1 probe adsorption sites exposed on Ni o metal particles [3]. They are minority species as compared to the bridging species (Fig. 1). The corresponding IR absorption intensifies and shills to higher frequencies (up to 2054 cm "1) at <300 K (Fig. 2). The eventual band position remains unchanged by hydrogen co-adsorption (Fig. 3). Hence, these terminal-CO species probe, similarly to the bridging species, aggregates of metal sites. The population of the terminal-CO improves with the CO coverage as well as with the involvement of low sitedensity surfaces [3]. Accordingly, their detection at low-coverages (at 300 K) may provide a further support to the dominance of the low site-density surfaces ofNi(100).

Terminal-CO vibrating at 2123 cm1, which is stabilized only at <160 K, detects the presence of Ni § sites [3]. CO adsorption on metal ions is weak, due to the low rc/t~ bonding contribution. Thus, application of low temperature regimes is essential for irreversible adsorption of CO on metal cations [32]. The presence ofNi § sites indicates that the catalyst examined is incompletely reduced under the conditions applied. Earlier studies [3] have consistently found, that the complete reduction of Ni ions on alumina is difficult to accomplish. Incomplete reduction of a number of transition metal ions, including Ni 2§ seems to be the trend on alumina and alike electronically "hard" support materials [33]. According to Pearson [33], the electronic hardness of a metal oxide means the occurrence of weakly polarizable metal-oxygen bonds. Terminal-CO*. In accordance with earlier studies [3], the occurrence of terminal-CO* with such an exceptionally high vCO frequency (2095-2072 cm "l) discloses electron-deficient nature for the metal adsorbing sites. The electron-drain may be effected via direct interactions between the metal atoms and the oxide support [5,6], or by means of oxygen-mediated Ni~ § electron-exchange interactions. A model pioneered by Zener [34] expounds the physical justification for such interactions.

575 4.3 Surface Attributes of the Metal Site Perturbation.

Intensification of vCO absorption of the terminal-CO* species (at 2095-2072 cm 1) is observed following two specific events: (i) application of low temperature regimes (<300 K); see Figs. 1 and 2, and (ii) co-adsorption of hydrogen (see Fig. 3). The hydrogen impact is shown to maximize at 160 K (lowest temperature applied) and to minimize near 240 K. Online mass spectra indicated that the desorbing species while heating CO/H2/Ni/AI203 up to 240 K were merely CO and H2 molecules. These results attribute the formation of terminal-CO* species to two different, but relevant, reasons, depending on whether the surface is H-covered or H-free.

H-covered surfaces. TPD studies [15,18,20,22] of CO and Hz co-adsorption on Ni(100) concluded that CO+H interactions in the co-adsorbed layer lead to formation of a new surface entity exhibiting new low-temperature desorption states of CO and H2. Relevant LEED investigations [22,25,26] brought about consistent results, by characterizing formation of CO+H interaction species. HREELS studies [18,25,26] of a similar system proved the existence of a strong link between the CO+H interaction species and a terminal-CO stretching vibration at 2100 cm 1 i.e. close to that observed here at 2095 cm q for the terminal-CO* species (spectrum (a), Fig. 3). These features were observed [18] only for CO and 1-12coadsorption on low site-density surfaces of Ni ~ namely, Ni(100). In effort to define the link between the CO+H interactions in the co-adsorbed layer and the high- frequency vCO band (at 2100 cmq), Mitchel et al. [19] suggested that hydrogen adsorption decreases the surface Ni ~ d-band electron density which is then not available for NiCO bonding. In other words, the authors [ 19] suggest through-Ni CO+H interactions to cause a suppression of the metal n-back bonding and the consequent vCO high frequency shift. A different, though rather relevant, view point may be drawn from the metal cluster literature. Morton and Preston [35] observed by means of ESR spectra the compound HNi(CO)3 to exhibit smaller g-shifts than Co(CO)4 compound. This is despite the greater spinorbit coupling constant for the metal atom in the former compound. The authors [35] attributed this behaviour to a strong interaction of the o-donor orbital with dz, which is expected for HNi(CO)3 since the H(ls) lies 1.4 eV above the CO 5o. This interpretation was based essentially on the close similarity between estimated 3dz spin population for the two HNi(CO)3 and C0(CO)4 compounds. It is worth emphasizing, however, that through-Ni CO+H interactions on Ni/A1203 catalysts are by no means anticipated to advance H-Ni-CO bondings as strong as the bondings involved in HNi(CO)3molecules.

H-free surfaces. On hydrogen-free surfaces, the formation of terminal-CO* species reveals the involvement of electron deficient (perturbed) Ni ~ metal sites. These sites can be generated consequently to (i) strong interactions of atomically dispersed (isolated) metal atoms with the support, and/or (ii) through-oxygen electron-exchange interactions with adjacent Ni + impurity sites. Hence, the adsorption is weakened and the IR vCO frequency shifted upwards~ As shown by Pearson [33], both incomplete reduction of metal ions and atomic dispersion of metal atoms are quite likely on surfaces of the electronically hard alumina.

576 4.4 Catalytic prospects for perturbed metal sites

Ni/Al203 is a potential hydrotreating catalyst as compared to Ni/SiO2 [2]. The metal in the former catalyst is distinct by high dispersions and incomplete reduction [3,5,6]. Thus, electronically perturbed metal sites are generated, as is evidenced here by the formation of terminal-CO* species in CO/H/Ni/A1203. The structure of the perturbed sites facilitates coordination to Ni ~ atoms, whereas the electron availability is accompalished via d-d electron exchange interactions with adjacent Ni + sites. The electron-mobile environment thus generated would promote the necessary redox electron migration across the surface and, hence, the hydrotreating activity. Analogously, perturbed metal sites have been considered to form in HDS catalysts (Mo ~+ [36]) and Fischer-Tropsch catalysts (Rh8+ [37]), and to promote the catalytic activity. In those catalysts, however, neither the perturbation nor the perturbed sites were sufficiently characterized [36,37]. ACKNOWLEDGMENT Excellent facilities provided during a Fulbright Research Fellowship at the Surface Science Center of Pittsburgh University were essential for the accomplishment of the present work, as were the fruitful discussions with Prof. J.T. Yates and the assistance provided by Dr. T.H. Ballinger with the spectroscopic techniques. Thanks are due to the Fulbright Foundation

(USA). REFERENCES 1. H.H. Storch, H. Golumbic and R.B. Anderson, The Fischer Tropsch and Related Synthesis, J. Wiley & Sons, New York, 1951. 2. M.A. Vannice, Catalysis Rev. Sci. Eng., 14 (1976) 153. 3. N. Sheppard and T.T. Nguyen, in: R.J.H. Clark and R.E. Hester (eds.), Advances in Infrared and Raman Spectroscopy, Vol. 5, Heyden, Philadelphia, 1978, pp. 67-134. 4. A. Andersson, Solid State Commun., 21 (1977) 75. 5. J.T. Yates, Jr., and C.W. Garland, J. Phys. Chem., 65 (1961) 617. 6. C.W. Garland, R.C. Lord and P.E. Troiano, J. Phys. Chem., 69 (1965) 1195. 7. M. Primet, J.A. Dalmon and G.A. Martin, J. Catal., 46 (1977) 25. 8. S. Andersson, Solid State Commun., 24 (1977) 183. 9. J.C. Bertolini, G. Dalami-Imelik and J. Rousseau, Surf. Sci., 68 (1977) 539. 10. L. Surnev, Z. Xu and J.T. Yates, Jr., Surf. Sci., 201 (1988) 1. 11. J.B. Peri, J. Catal., 86 (1984) 84. 12. M. Primet and N. Sheppard, J. Catal., 41 (1976) 258. 13. J. Galuzska and Y. Amenomiya, in: S. Kaliaguine and A. Mathay (eds.), Catalysis on the Energy Scene, Elsevier, Amsterdam, 1984, p. 163. 14. D.G. Blackmond and E.I. Ko, J. Catal., 96 (1985) 210. 15. T.E. Madey, D.W. Goodman and R.D. Kelley, J. Vac. Sci. Technol., 16 (1979) 433. 16. D.W. Goodman, J.T. Yates, Jr., and T.E. Madey, Surf. Sci., 93 (1980) L135. 17. D.W. Goodman, R.D. Kelley, T.E. Madey and J.T. Yates, Jr., J. Catal., 63 (1980) 226.

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590

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591

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592 -Mr. V. D. Douplyakin Director Omrfu Dept. of Borzescov Institute Catalysis Omsk, Heftezavodskay str. 45, Russia

-Dr. Usman A. E1 Nafaty Lecturer King Fahd University of Petroleum & Minerals

(KFUPM) KFUPM 1709 Dhahran 31261, Saudi Arabia

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(KFUPM) Department of Chemical Engineering P.O.Box 1809 Dhahran 31261, Saudi Arabia -Mr. Suliman A. El Muhanna Ph.D. student King Fahd University of Petroleum & Minerals

(KFUPM) P.O.Box 798 Dhahran 31261, Saudi Arabia

~M)

KFUPM Box 83 Dhahran 31261, Saudi Arabia

(KFUPM) The Research Institute KFUPM, P.O.Box 432 Dhahran 31261, Saudi Arabia -Mr. Jamal A. Anabtawi King Fahd University of Petroleum & Minerals

~M) The Research Institute Saudi Arabia -Dr. Saeed M. A1 Zahrani Assistant Professor King Saud University Chemical Engineering Dept. P.O.Box 800 Riyadh 11421, Saudi Arabia -H. E. Mr. Hisham A1 Nazer Minister of Petroleum and Mineral Resources Ministry of Petroleum and Mineral Resources Saudi Arabia -Mr. Hideo Yamamoto Director General of PEC PEC, Riyadh Office, Saudi Arabia -Mr. Mohammed H. AI Hazmi Chemist Catalysis Group Research and Development (SABIC) P.O.Box 42503 Riyadh 11551, Saudi Arabia -Dr. Altaf Husain Advisor, Catalysis Group Development Dept. Sabic Research & Development (SABIC) P.O.Box 42503 Riyadh 11551, Saudi Arabia -H. E. Hisham AI Nazer Minister of Oil Saudi Arabia Ministry of Oil AI Riyadh, Saudi Arabia

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-Dr. M de Wind Area Technical Service Manager Akzo Nobel Chemicals BV P. O. Box 247 3800 AZ Amersfoort, The Netherlands -Mr. Mohamed S. AI Dhahiri ADNOC P.O.Box 898, U.A.E. -Mr. Khalaf R. AI Etabi Director, Industrial Div. ADNOC Industrial Division P.O.Box 898 Abu-Dhabi, U.A.E. -Mr. Ali M. A. A1 Jabri Head of Industrial Div. ADNOC Industrial Division A1 Riwas Refinery P.O.Box 11885, AI Riwas, U.A.E. -Mr. Mohamed R. M. A1 Sadiqi ADNOC Industrial Division P.O.Box 898, Abu-Dhabi, U.A.E. -Mr. Adel A. M. Khalifa Head of Industrial & Location Div. ADNOC Industrial Division Am AI Nar Refinery P.O.Box 3087, Abu-Dhabi, U.A.E. -Dr. Keith A. Holder Senior Chemist BP Oil M & S Technology Development Unit Chertsey Road, Sunbury-on-Thames Middlesex TW 16 7LN, U.K. -Mr. Andrew Holt Owner Catalyst Developments P.O.Box 507 Sheffield S10 3YT, U.K. -Dr. Simon J. W. Nutting Area Manager Chevron International Oil Company Technology Marketing 93 Wigmore Street, London W1H 9AA, U.K. -Mr. D. S. Nicholson Sales Manager, Catalyst Crosfield Catalysts P.O.Box 26 Warrington, Cheshire WA5 lAB, U.K.

594 -Mr. Kelvin A. Halliwell Account Manager Criterion Catalyst Co. Ltd. 1651 Parkway, The Solent Business Park Whitley, Fareham PO 15 7AH, U.K. -Mr. Shaun Rory Boardman Technical Support Engineer Criterion Catalyst Company 1650 Parkway Solent Businees Park, Whiteley Hampshire, P015 7AH, U.K. -Mr. Leonard Mumro Williamson V.P. International Sales Criterion Catalyst Company 1650 Parkway Solent Businees Park, Whiteley Hampshire, P015 7AH, O.K. -Mr. M. A. Murphy Technical Manager Engelhard Chancery House St. Nicholas way, U.K. -Mr. Layton M. Bamon Sales Manager Engelhard Industries Chancery House St. Nicholas Way Sutton, Surrey SMI 1JB, U.K. -Mr. F. W. Zemichael Ph.D.. Research Student University of Manchester Institute of Science and Technology (UMIST) Chemistry Dept., P.O.Box 88 Manchester M60 1QD, England, ILK. -Prof. Ronald Hughes Head of Department University of Salford Department of Chemical Engineering Safford M5 4WT, U.K. -Dr. Osama H. El Bayoumi Program Manager Science & Technology US Air Force "London" European Office of AeroSpace Research (EOARD) 223/231 Old Marylebone Rd. London, NWI 5TH, U.K. -Mr. Mike L. Sharp Vice President, Sales & Marketing Discovery Aluminas 16010 Barker's Point Lane Suite 250 Houston Texas 77079, U.S.A

-Mr. Philip G. Thiel Sales Manager, Hydro-processing Catalysts GRACE Davison & Co. P.O. Box 2117 Baltimore, Maryland 21203, U.S.A -Mr. Charles Joseph DiBona President American Petroleum Institute 1220 L. Street, NW Washington, DC 20005, U.S.A. -Dr. John C. Fetzer Senior Research Chemist Chevron Research and Technology Company P.O.Box 1627, Richmond, CA 94802-0627, U.S.A. -Mr. J. Kent Murray President Chevron Research and Technology Company 100 Chevron Way Richmond California 94802, U.S.A. -Mr. Dennis J. O'Rear

Team Leader - New Materials Chevron Research and Technology Company 100 Chevron Way Richmond CA 94802-0627, U.S.A. -Mr. Charles W. Olsen Technical Services Manager Crosfield Catalysts 4099 West 71st Street Chicago Illinois 60629, U.S.A. -Dr. Michael J. Pearson Director of Technology Discovery Aluminas 16010 Baker's Point Lane Suite 250 Houston, Texas 77079, U.S.A. -Dr. Mario L. Occelli Program Director, Zeolites & Clays Research Georgia Tech Research Institute 925 Dalney Street, Baker Bldg. Mail Code 0827 Atlanta Georgia 30332, U.S.A. -Mr. John Juracek Director of Hydroprocessing Catalyst Grace Davison Co. WR Grace & Co., P.O.Box 2117 Baltimore Maryland 21203, U.S.A. -Mr. Harald K. Lemke Senior Chemist Nalco/Exxon Energy Chemical, L.P. P.O.Box 87 Sugar Land,Texas 77487-0087, U.S.A.

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-Dr. Opinder K. Bhan Research Engineer Shell Oil Company Westhollow Technology Center M-2211 3333 Hwy. 6 South HOUSTON, TX 77082, U.S.A. -Dr. John W. Ward Senior Technical Director Tricat Industries 2854 Santiago Boulevard Orange CAL 92667, U.S.A. -Mr. Tayfun Altug Business Manager Union Carbide Corporation 39 Old Ridgebury Road Danbury CT 06817-0001, U.S.A. -Dr. Mahmoud R. Rifi Research Associate Union Carbide Corporation P. O. Box 670 Boundbrook, New Jersey 08805, U.S.A. -Mr. J. P. Yimoyines Vice President Union Carbide Corporation 39 Old Ridgebury Road Danbury CT 06817-0001, U.S.A. -Prof. C. Bruce Gates University of California Dept. of Chemical Engineering & Material Science Davis CA 95616, U.S.A. -Dr. Alfredo Morales Catalysis Refining Advisor INTEVEP S.A., Refining and Petrochemical Department Apartado Postal 76343, Caracas 1070A,, Venezuela -H. E. Dr. Erwin Arrieta Valera Minister of Energy and Mines Ministry of Energy and Mines Venezuela

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597 Author Index

Abotteen, S., 253 Absi-Halabi, M., 189, 243 Ahmed, M., 419 Aika, K., 397 Akkulov, A. G., 559 A1 -Sherehy, F. A., 515 A1-Mashan, M. M., 283 A1-Nasser, A., 171 A1-Qahtani, H., 437 A1-Zahrani, S. M., 383 Ali, S. A., 225 Alkemade, U., 303,339 Anabtawi, J. A., 225 Bannayan, M. A., 273 Bartholdy, J., 117 Beltramini, J. N., 465 Bhan, O. K., 135 Bhatacharya, S., 171 Billon, A., 209 Blidisel, I., 217 Callant, M., 199 CarrazAn, S. R. G., 1 Chaudhuri, S.R., 171 Chopra, S., 243 Cooper, B. H., 117 Dashti, H., 283 de Wind, M., 157 Delmon, B., 1, 199 Dorbon, M., 209 Dougan, T. J., 303 Duch, A., 91 Dufresne, P., 253 Ekiri, S., 447 E1-Dusouqui, O., 525 E1-Nafaty, U. A., 355 Englisch M., 525 Fan, L., 507 Fang, R., 465 Faramawy, S. M., 407 Fetzer, J. C., 263

Fujimoto, K., 235 Fujita, K., 157 Fukase, S., 455 Gates, B. C., 49 George, S.E., 135 Gorbacheva, L. B., 483 Grange, P., 199 Hamakawa, S., 375 Hasan, S. A., 407 Hayakawa, H., 293 Hayakawa, T., 375 Holder, K., 199 Hughes, R., 313,427 Hutchings, G., 313 Igarashi, N., 455 Inamura K., 543 Inui, T., 489 Ione, K.G., 477 Ishibashi, Y., 455 Iwamoto, R., 543 Jentys, A., 525 Kam, E. K. T., 283,427 Kaminisky, W., 91 Karasuda, T., 397 Kasztelan, S., 209 Kato, K., 447, 455 Khan, Z., 189 Kim, J. B.,489 Kitou, Y., 181 Komatsu, T., 447 Koon, C. L., 313 Korili, S., 1 Koyama, H., 147 Krzywicki, A., 507 Kumagai, M., 147 Lemke, H. K., 273 Lercher, J. A., 525 Lobban, L., 383 Lugstein, A., 525 Maesaki, N., 551

598 Mandani, F., 427 Mann, R., 355 McGhee, B., 313 Mekewi, M. A., 407 Mignard, S., 209 Miyauchi, Y., 157 Mizutani, Y., 181 Morales, A., 125 Muhammad, O. H. J., 365 Murata, S., 499 Nagai, E., 147 Nakagawa, S., 499 Nakamura, I., 235 Namba, S., 447 Nomura, M., 499 Nomura, T.,455 O' Connor, P., 323 Occelli, M. L., 27 Okamoto, Y., 77 Ono, T., 499 Paloumbis, S., 339 Qamra, A., 243 Rahman, A., 419 Ruiz, P., 1 Sadek, S. A., 407 Sado, F., 551 Sakai, T., 499 Sarrazin, P., 209 Sato, K., 375 Shibata, Y., 181 Siddiqui, M. A. B., 225 Snape, C. E., 313

Snytnikova, G. P., 477 Sobalik, Z., 1 Solari, R. B., 125 Somogyvari, A., 507 Sonnemans, J. W. M., 99 Stanislaus, A., 189, 243 Stepanov, V.G., 477 Stephenson, W. K., 273 Sugioka, M., 551 Takatsuka, T., 293 Takehira, K., 375 Tochiyama, C., 551 Toktabaeva, N. Ph., 483 Torii, H., 147 Trimm, D. L., 65 Tsuchiya, S., 535 Tsunoda, T., 375 Valeri, F., 253 Vicente Rodriguez, M. A., 1 Vinek, H., 525 Watari, R., 293 Yamamoto, Y., 181 Yamazaki, H., 181 Yanik, S. J., 323 Yashima, T., 447 Zaidi, S. M. J., 225 Zakarina, N. A., 559 Zaki, M. I., 569 Zakumbaeva, G. D., 483 Zamfirache, R., 217 Ziegler, T., 507

599

STUDIES IN SURFACE SCIENCE AND CATALYSIS Advisory Editors: B. Delmon, Universite Catholique de Louvain, Louvain-Ia-Neuve, Belgium J.T. Yates, University of Pittsburgh, Pittsburgh, PA, U.S.A.

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Preparation of Catalysts I.Scientific Bases forthe Preparation of Heterogeneous Catalysts. Proceedings of the First International Symposium, Brussels, October 14-17,1975 edited by B. Delmon, P.A. Jacobs and G. Poncelet The Control of the Reactivity of Solids. A Critical Survey of the Factors that Influence the Reactivity of Solids, with Special Emphasis on the Control of the Chemical Processes in Relation to Practical Applications by V.V. Boldyrev, M. Bulens and B. Delmon Preparation of Catalysts II. Scientific Bases forthe Preparation of Heterogeneous Catalysts. Proceedings of the Second International Symposium, Louvain-Ia-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 byJ. Bourdon Catalysis by Zeolites. Proceedings of an International Symposium, Ecully (Lyon), Septem ber 9-11, 1980 edited by B.lmelik, C. Naccache, V. 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 ed ited by T. Seiyama and K. Tanabe Catalysis by Supported Complexes by Vu.i. Vermakov, B.N. Kuznetsov and V.A. Zakharov Physics of Solid Surfaces. Proceedings of a Symposium, Bechyne, September 29-0ctober 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.lmelik, 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 ed ited by P.A. Jacobs, N.i. Jaeger, P. JiriJ and G. Schulz-Ekloff Adsorption on Metal Surfaces. An Integrated Approach edited by J. Benard Vibrations at Surfaces. Proceedings of the Third International Conference, Asilomar, CA, September 1-4,1982

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ed ited by C.R. Brundle and H. Morawitz Heterogeneous Catalytic Reactions Involving Molecular Oxygen by G.I. Golodets Preparation of Catalysts III. Scientific Bases for the Preparation of Heterogeneous Catalysts. Proceedings of the Third International Symposium, Louvain-Ia-Neuve, September 6-9,1982 edited by G. Poncelet, ~ Grange and P.A. Jacobs Spillover of Adsorbed Species. Proceedings of an International Symposium, Lyon-Villeurbanne, September 12-16,1983 edited by G.M. Pajonk, S.J. Teichner and J.E. Germain Structure and Reactivity of Modified Zeolites. Proceedings of an International Conference, Prague, July 9-13, 1984 edited by ~A. Jacobs, NJ. Jaeger, P. JiriJ, V.B. Kazansky and G. Schulz-Ekloff Catalysis on the Energy Scene. Proceedings of the 9th Canadian Symposium on Catalysis, Quebec, P.Q., September 3D-October 3,1984 edited by S. Kaliaguine and A. Mahay Catalysis by Acids and Bases. Proceedings of an International Symposium, Villeurbanne (Lyon), September 25-27,1984 edited by B. Imelik, C. Naccache, G. Coudurier, V. Ben Taarit and J.C. Vedrine Adsorption and Catalysis on Oxide Surfaces. Proceedings of a Symposium, Uxbridge, June 28-29, 1984 edited by M. Che and G.C. Bond Unsteady Processes in Catalytic Reactors by Vu.Sh. Matros Physics of Solid Surfaces 1984 edited byJ. Koukal Zeolites: Synthesis, Structure, Technology and Application. Proceedings of an International Symposium, Portoroz-Portorose, September 3-8, 1984 edited by B. Driaj, S. Hocevar and S. Pejovnik Catalytic Polymerization of Olefins. Proceedings of the International Symposium on Future Aspects of Olefin Polymerization, Tokyo, July 4-6,1985 edited by T. Keii and K. Soga Vibrations at Surfaces 1985. Proceedings of the Fourth International Conference, Bowness-on-Windermere, September 15-19,1985 edited by D.A. King, N.V. Richardson and S. Holloway Catalytic Hydrogenation ed ited by L. Cerveny New Developments in Zeolite Science and Technology. Proceedings of the 7th International Zeolite Conference, Tokyo, August 17-22, 1986 edited by V. Murakami, A.lijima and J.W. Ward Metal Clusters in Catalysis edited by B.C. Gates, L. Guczi and H. Knozinger Catalysis and Automotive Pollution Control. Proceedings of the First International Symposium, Brussels, September 8-11, 1986 edited by A. Crucq and A. Frennet Preparation of Catalysts IV. Scientific Bases for the Preparation of Heterogeneous Catalysts. Proceedings of the Fourth International Symposium, Louvain-Ia-Neuve, September 1-4,1986 edited by B. Delmon, ~ Grange, P.A. Jacobs and G. Poncelet Thin Metal Films and Gas Chemisorption edited by ~ Wissmann Synthesis of High-silica Aluminosilicate Zeolites edited by P.A. Jacobs and J.A. Martens Catalyst Deactivation 1987. Proceedings of the 4th International Symposium, Antwerp, September 29-0ctober 1,1987 edited by B. Delmon and G.E Froment

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Keynotes in Energy-Related Catalysis edited by S. Kaliaguine Methane Conversion. Proceedings of a Symposium on the Production of Fuels and Chemicals from Natural Gas, Auckland, April 27-30, 1987 edited by D.M. Bibby, C.D. Chang, R.F. Howe and S. Yurchak Innovation in Zeolite Materials Science. Proceedings of an International Symposium, Nieuwpoort, September 13-17,1987 edited by P.J. Grobet, W.J. Mortier, E.F. Vansant and G. Schulz-Ekloff Catalysis 1987. Proceedings 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 IUPAC Symposium (COPS I), Bad Soden a. Ts., April 26-29,1987 edited by K.K. Unger, J. Rouquerol, K.S.W. Sing and H. Kral Physics of Solid Surfaces 1987. Proceedings of the Fourth Symposium on Su rface 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 by Yu. Sh. Matros Successful Design of Catalysts. Future Requirements and Development. Proceedings of the Worldwide Catalysis Seminars, July, 1988, on the Occasion of the 30th Anniversary of the Catalysis Society of Japan edited byT.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 Surfaces. Proceedings of a European Conference, Trieste, September 13-16, 1988 edited by C. Morterra, A. Zecchina and G. Costa Zeolites: Facts, Figures, Future. Proceedings of the 8th International Zeolite Conference, Amsterdam, July 10-14,1989. PartsAand B ed ited by ~A. Jacobs and R.A. van Santen Hydrotreating Catalysts. Preparation, Characterization and Performance. Proceedings of the Annual International AIChE Meeting, Washington, DC, November 27-December 2, 1988 edited by M.L. Occelli and R.G. Anthony New Solid Acids and Bases. Their Catalytic Properties by K. Tanabe, M. Misono, Y. Ono and H. Hattori Recent Advances in Zeolite Science. Proceedings of the 1989 Meeting of the British Zeolite Association, Cambridge, April 17-19, 1989 edited by J. Klinowsky and P.J. Barrie Catalyst in Petroleum Refining 1989. Proceedings of the First International Conference on Catalysts in Petroleum Refining, Kuwait, March 5-8,1989 edited by D.L. Trimm, S. Akashah, M. Absi-Halabi and A. Bishara

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Future Opportunities in Catalytic and Separation Technology edited by M. Misono, V. Moro-oka and S. Kimura Volume 55 New Developments in Selective Oxidation. Proceedings of an International Symposium, Rimini, Italy, September 18-22,1989 edited by G. Centi and F. Trifiro Olefin Polymerization Catalysts. Proceedings of the International Symposium Volume 56 on Recent Developments in Olefin Polymerization Catalysts, Tokyo, October23-25,1989 edited by T. Keii and K. Soga Vol ume 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 ed ited by J.L.G. Fierro Vol ume 58 Introduction to Zeolite Science and Practice edited by H. van Bekkum, E.M. Flanigen and J.C. Jansen Volume 59 Heterogeneous Catalysis and Fine Chemicals II. Proceedings of the 2nd International Symposium, Poitiers, October 2-6,1990 edited by M. Guisnet, J. Barrault, C. Bouchoule, D. Duprez, G. Perot, R. Maurel and C. Montassier Volume 60 Chemistry of Microporous Crystals. Proceedings of the International Symposium on Chemistry of Microporous Crystals, Tokyo, June 26-29,1990 edited by T. Inui, S. Namba and T. Tatsumi Natural Gas Conversion. Proceedings of the Symposium on Natural Gas Volume 61 Conversion, Oslo, August 12-17,1990 edited by A. Holmen, K.-J. Jens and S. Kolboe Volume 62 Characterization of Porous Solids II. Proceedings of the IUPAC Symposium (COPS 11), Alicante, May 6-9,1990 edited by F. Rodriguez-Reinoso, J. Rouquerol, K.S.W. Sing and K.K. Unger Volume 63 Preparation of Catalysts V. Scientific Bases for the Preparation of Heterogeneous Catalysts. Proceedings of the Fifth International Symposium, Louvain-Ia-Neuve, September 3-6,1990 edited by G. Poncelet, P.A. Jacobs, P. Grange and B. Delmon Volume 64 New Trends in CO Activation ed ited by L. Guczi Catalysis and Adsorption by Zeolites. Proceedings of ZEOCAT 90, Leipzig, Volume 65 August20-23,1990 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, Balatonfured, September 10-14,1990 edited by LJ. 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, NJ. Jaeger, L. Kubelkova and B. Wichterlova

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Poisoning and Promotion in Catalysis based on Surface Science Concepts and Experiments by M. Kiskinova Catalysis and Automotive Pollution Control II. Proceedings of the 2nd International Symposium (CAPoC 2), Brussels, Belgium, September 10-13,1990 ed ited by A. Crucq New Developments in Selective Oxidation by Heterogeneous Catalysis. Proceedings of the 3rd European Workshop Meeting on New Developments in Selective Oxidation by Heterogeneous Catalysis, Louvain-Ia-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,August17-20, 1993 edited by T.lnui, K. Fujimoto, T. Uchijima and M. Masai Heterogeneous Catalysis and Fine Chemicals III. Proceedings of the 3rd International Symposium, Poitiers, April5 - 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, ~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 MicroporousCrystals. Proceedings of the International Symposium on Zeolites and Microporous Crystals, Nagoya, Japan, August 22-25,1993 ed ited by T. Hattori and T. Vashima 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

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Oscillating Heterogeneous Catalytic Systems by M.M. Slin'ko and NJ. Jaeger Characterization of Porous Solids III. Proceedings of the IUPAC Symposium (COPS III), Marseille, France, May9-12, 1993 edited by J.Rouquerol, F. Rodriguez-Reinoso, K.S.W. Sing and K.K. Unger Catalyst Deactivation 1994. Proceedings of the 6th International Symposium, Ostend, Belgium, October 3-5,1994 edited by B. Delmon and G.F. Froment Catalyst Design for Tailor-made Polyolefins. Proceedings of the International Symposium on Catalyst Design for Tailor-made Polyolefins, Kanazawa, Japan, March 10-12, 1994 edited by K. Soga and M. Terano Acid-Base Catalysis II. Proceedings of the International Symposium on Acid-Base Catalysis II, Sapporo, Japan, December 2-4, 1993 edited by H. Hattori, M. Misono and Y. Ono Preparation of Catalysts VI. Scientific Bases for the Preparation of Heterogeneous Catalysts. Proceedings of the Sixth International Symposium, Louvain-La-Neuve, September 5-8, 1994 edited by G. Poncelet, J. Martens, B. Delmon, ~A. Jacobs and P. Grange Science and Technology in Catalysis 1994. Proceedings of the Second Tokyo Conference on Advanced Catalytic Science and Technology, Tokyo, August 21-26,1994 edited by Y. Izumi, H. Arai and M.lwamoto Characterization and Chemical Modification of the Silica Surface by E.F. Vansant, ~ Van Der Voort and K.C. Vrancken Catalysis by Microporous Materials. Proceedings of ZEOCAT'95, Szombathely, Hungary, July 9-13, 1995 edited by H.K. Beyer, H.G.Karge, I. Kiricsi and J.B. Nagy Catalysis by Metals and Alloys by V. Ponec and G.C. Bond Catalysis and Automotive Pollution Control III. Proceedings of the Third International Symposium (CAPoC3), Brussels, Belgium, April 20-22, 1994 edited by A. Frennet and J.-M. Bastin Zeolites: A Refined Tool for Designing Catalytic Sites. Proceedings of the International Symposium, Quebec, Canada, October 15-20,1995 edited by L. Bonneviot and S. Kaliaguine Zeolite Science 1994: Recent Progress and Discussions. Supplementary Materials to the 10th International Zeolite Conference, Garmisch-Partenkirchen, Germany, July 17-22, 1994 edited by H.G. Karge and J. Weitkamp Adsorption on New and Modified Inorganic Sorbents edited by A. D"browski and V.A. Tertykh Catalysts in Petroleum Refining and Petrochemicals Industries 1995. Proceedings of the 2nd International Conference on Catalysts in Petroleum Refining and Petrochemical Industries, Kuwait, April 22-26, 1995 edited by M. Absi-Halabi, J. Beshara, H. Qabazard and A. Stanislaus