Fluid Catalytic Cracking V

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

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

vii

List of Contributors

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

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

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

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

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

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

viii

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

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

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

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

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

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

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

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

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

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

M. Nakamura Nippon Ketjen Tokyo JAPAN

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

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

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

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

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

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

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

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

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

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

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

Table of Contents List of Contributors Preface

1

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

1

zeolites: effects of aluminum removal and of basic metal oxide

3

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

2

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

3

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

4

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

5

87 catalysts in industrial units

The improvement of catalytic cracking process through the

6

107 utilization of new catalytic materials

7

NExCC-Novel short contact time catalytic cracking technology

111

8

Effect of vanadium on light olefins selectivity

133

9

Reduction of olefins in FCC gasoline

141

Gasoline sulfur removal: kinetics of S compounds in FCC

10

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

11

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

12

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

13

201 tolerant technologies

14

Active site accessibility of resid cracking catalysts

209

Catalyst evaluation for atmospheric residue cracking, the effect of

15

219 catalyst deactivation on selectivity

16

Optimum properties of RFCC catalysts

227

An experimental protocol to evaluate FCC stripper performance in

17

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

18

251 coke formation in fluid catalytic cracking

19

Bifunctionality in catalytic cracking catalysis

263

Catalytic cracking of alkylbenzenes. Y-zeolites with different

20

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

21

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

22

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

23

fresh and contaminant-affected catalyst structure on FCC

311

performance Keyword Index

333

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

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

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

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

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

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

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

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

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

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

H+ 0 I O-Si-O I 0

Si I 0 H

+

3H20

b,-

Si-OH

HO-SI H 0 I Si

+

AI(OH)3

G.T: Kerr, 1979

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

on this m e c h o n i s m

Figure 1. Schematic representation of formation of ultrastable Y materials

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

29Si and 27A1 MASNMR

and microcalorimetry

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

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

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

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

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

Isolated or clustered

Most abundant

TF

T-site

isolated

mildly dealuminated

TNFA

A1 in small cage and/or

isolated

mildly dealuminated

TNFA

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

severely dealuminated

PNFA

intermediate between octahedral and tetrahedral

unknown

severely dealuminated

ONFA

alumina species in s-c

clustered

always present

s-c = supercage

[from Ref. 61]

a

T = tetrahedral, 0 = octahedral

b

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

27A1in zeolites [48] has shown two

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

27A1MASNMR and 27A1CP

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

10

4X

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

5X

T ~, It

"

]\

2X

A

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

o.o33..

A

Calcine 538~ 2 hr Steam 600~ 3 hr

l ~ j ~

T Parent USY (LZ-Y82)

1X 9 =+0

~5o

A z~

~oo"5'o . . . .

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

~

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

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

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

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

12

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

13

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

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

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

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

15

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

16

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

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

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

Si/AI 2O 15 i

C

o

!

i

10 i

5 i

!

i

i

i

24.5

c (D

E :15 24.4 fl)

c

~D

24.3

24.2

9

N~ , # Al/ucell

!

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

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

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

carb. ion

1. clean framework, high-silica Y

4.9

13.7

8.5

2. conventional USY c

5.1

2.456

35.8

28.7

80

3. USY formed from AHF-dealum. material d

8.1

2.441

43.5

33.1

76

2.454

% carb. ion 62

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

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

substantially

lower

activity.

Gorte

and

co-workers

[66]

carried

out

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

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

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

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

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

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

22

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

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

% carb ion act.

total mol/(h'g) x 103

5b

0.30

7.28

69

32.5

22.4

0.196

5

0.40

9.48

70

31.7

22.2

0.194

5

0.50

12.6

73

33.7

24.6

0.215

5

0.60

15.7

74

34.9

25.8

0.225

5

0.70

18.8

75

35.9

26.9

0.235

5

0.80

21.1

76

35.3

26.8

0.234

5

1.00

26.9

80

36.0

28.8

0.251

12.4 c

1.00

13.0

71

17.4

12.4

0.227

12.4

2.30

30.1

78

17.5

13.6

0.249

molec/(min'A1F)

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

a 500oc,

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

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

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

,so0utane

30

/I

/

n"~t-

x

---~

E~

20

. ~ x:: 0

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

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

240 ,,.-.,

._=

E 200 160 O

E ::I.

120

9~

80

o


T(2Si,2AI) > T(4Si,0AI)>> T(1Si,3A1); see Figure lB. As reported elsewhere (11), the hydrolysis of the fluorosilicate salt forms protons and fluoride anions that cause the dealumination of the

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

Si/A1

ao (nm,)

LZY-82

2.8

2.4480

785

0.16

LZ-210 LZ-210 LZ-210

3.1 4.4 6.4

2.4581 2.4474 2.4377

607 625 629

2.44 1.17 0.01

LZ-10

3.0

2,4310

600

0.17

2.9 8.3

2.4530 2.4360

742 740 . . . . . .

4.41 0.17

USHY USHY-DA

SA(m2/_g)

%Na2Q

Table 2.

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

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

-

HY type

Si/A1

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

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

LZY-82

87.6 (1.5)

90.9 (4.6)

95.8 (14.1)

101.5 (37.2)

106.9 (37.6)

110.8 (5.0)

4.5

2.8

LZ-210

--

8 9.0 (3.5)

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

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

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

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

3.4

3.1

4.4

4.4

13.1

6.4

101.5 (16.9)

107.5 (76.7)

110.6 (6.4)

24.0

3.0

101.3 (37.3) 101.3 (1.9.0)

t06.2 (32.2) 106.8

111.9 (9.7) 111.1

4.2

2.9

11.0

8.3

(64.0)

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

LZ-210 LZ-210

LZ-10

USHY USHY-DA

--

90.3 (7.6)

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

46

A

D

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

_._ _L L_..~

i~

G o

-io

.......

-ibo ..... ~ - ~ o

Chemical

Shift

-

-ioo

[ppm]

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

47

A

B

120

80

40

0

.-40

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

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

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

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

50

NH8 uptake (wn~/g)

~1000

.~,

,

2500

,, , . ,

"u_,.,

u~,J

__,

. .......

:.

,,,.

-

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

-+-

-4-

L,Z-2D-,54

- . e - L,Z-10

A

I.Z..210";4,4

2OOO

1600 I(XX)

0

0.1

0.2.

0.8

0,4

0.8

~0

0.7

1400

B

1200 1000 8OO

0OO

0G ~

o

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

i .....

oa

: . . . . . .

I:. . . . .

eL=

+

USHY

__+ :

Jt . . . . . .

.

,

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

~a cx4 P (torr)

.....

s~.__.,.

o~a

_

_

,tr .

~e

.

.

.

.

.

.

c~7

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

51

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

2.8

208

122

1020

680

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

3.1 4.4 6.4

140 147 191

220 211 94

1893 1919 799

1130 1316 525

5. LZ-10

3.0

190

35

373

188

6. USHY 7. USHY-DA

2.9 8.3

150 i 86

40 67

479 608

161 420

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

Si AI

200 11 . .

. . --

36

,

12 24 . . . .

---

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

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

:

LZ-'210"e.1

-+"

LZ-'210-~4

-'~

LZ-'210-~4

- =

LZ-IO

A |

O 0

_

.....

9

,

,,|

~,,

LI

I(XX)

i

i

|

|

. . . . . . . .

|

RtX)O

1500

250

-

t.~'412

"-I-- 081"~

200

-11- ~ ' O A

100

B 0 0

200

400

~10

,~nmon~ ~

800

1000

1200

1400

(uric/q)

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

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

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

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

T(C) 200 300 400 500

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

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

.

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

55

~--~

1600

~

~.

, -,.,,

,

,

"

w-

'-

'"

"'~

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

--'""

'~'1

1500 1400 WAVENUMBERS (1/cm)

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

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

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

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

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

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

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

59

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

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

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

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

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

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

29Si

62

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

2.454 2.431 2.426 2.425

0.27 0.23 0.22 0.17

785 668 649 493

697 580 569 428

88 88 80 65

.

0.16 --,-

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

63

q

A

.

B

A

D I

0

'

-'

'~'

"'--

9

"

-50

'

"

"'

"i"

'

"' . . . .

~'

-100

I

-150

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

64

100

50

0

-50

-100

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

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

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

4.HY-815 .

.

.

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

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

.

.

.

.

.

.

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

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

(72)

(28)

4.5

2.8

45*

2.8 2.8

108

---

2.8

....

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

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

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

2.8 45 ---

1020 319 222 174

340 155 109 98

680 164 113 76

208 158 156 157

122 33 20 15

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

67

1400i

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

1200 1000 LZY'S2 8O0

" + - HY-]'80 HY-TU

0OO

-,0- HY-818

400 200 0

0,1

0,2

0,8 P (tort}

0.4

(15

0.6

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

260

Q (kJ/mol) ....

_.

,,

,.

,

. ,

,

~,

,,

- :" LZY-82

, , . ,

~,

,,

~,

~, : _

,

,

,

.....

.,

-4-- HY-Te0

.....

,,

,.,

,

.

._

,. . . . . .

~::_~

- ' ~ HY-TU

-:

--=

. . . . . . .

:

.

~ . . . .

-.B- HY"818

[

1200

1400

200

100

0 t,

0

2O0

4OO ~nla

~X) 800 IOO0 uptak8 (llnol/o)

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

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

69

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

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

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

71

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

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

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

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

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

1 L L L L

2 H L L L

3 L L H M

4 H L H M

5 L L M H

6 H L M H

7 L H L L

8 H H L L

9 L H H H

10 H H H H

11 L H M H

12 H H M H

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

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

74

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

3.0

at 60%wt conversion

xA liB

2.5

J

El HYD-B O .ca O

eC

2.0

o HYD-C

J

1.5

J

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

1.0

Y

0.5

0.0 0.0

0.5

1.0

1.5

2.0

2.5

3.0

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

Gasoline factor, Catalyst: Base, T=970~

1.6 xA 1.4

at 60%wt conversion

'

LmB

/

O HYD-B

1.2

eC O

o HYD-C

1.0

O

0.8 O ~

jf

0.6 0.4 0.2 0.0

f/

J 0.0

0.2

0.4

0.6

0.8

1.0

1.2

1.4

1.6

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

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

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

=

13 X

11 Q

9

f

f

J

7

1~

1~

180

2~

2~

3~

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

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

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

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

77

HA

HA

LA

HA

HA

LA

LA

LA

HA

HA

LA

LA

HRE

HRE

HRE

HRE

LRE

HRE

HRE

LRE

LRE

LRE

LRE

LRE

MZ/M

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

HSA

HSA

HSA

LSA

HSA

MSA

LSA

HSA

HSA

LSA

H Z / M LZ/M MSA

LSA

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

HA

HA

LA

LA

HA

LA

HA

LA

HRE

HRE

HRE

HRE

LRE

LRE

LRE

LRE

HSA

HSA

HSA

M SA

HSA

HSA

HSA

M SA

MZ/M

HZ/M

MZ/M

HZ/M

MZ/M

HZ/M

HZ/M

HZ/M

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

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

[] LA catalysts9 HA catalysts-

[]

\

L)

3

A v

0

2

4

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

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

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

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

HA

LA

HA

LA

HA

LA

HA

LA

HRE

HRE

HRE

HRE

LRE

LRE

LRE

LRE

HSA

HSA

HSA

MSA

HSA

HSA

HSA

MSA

MZ/M

MZ/M

HZ/M

HZ/M

MZ/M

HZ/M

HZ/M

HZ/M

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

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

_

!

39.5 ~ [] LA catalysts 9 HA catalysts 39.0

~_~..~-~

/ i

-~.

_

38.5 o~

S

38.0 37.5

o

jJ

37.0

r,r

36.5 -

36.0 35.5

j

Z

i..a

35.0 1~

160

2~

220

2~

2~

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

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

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

LA LRE

HA LRE

HA HRE

HA LRE

LA LRE

HA HRE

LA HRE

LA HRE

HSA HZ/M

HSA MZ/M

HSA MZ/M

HSA HZ/M

MSA HZ/M

HSA HZ/M

MSA HZ/M

HSA MZ/M

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

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

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

0.70 0.66

l rn LA catalysts]

0.62 9

l " HA catalysts]

0.58 ._o

\

0.54 0.50

\

O

0.46 w

0.42

[]

[]

0.38 0.34 0.30 0

2

4

6

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

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

83

HA HRE HSA MZ/M

HA HRE HSA HZ/M

HA LRE HSA MZ/M

LA HRE HSA MZ/M

HA LRE HSA HZ/M

LA HRE M SA HZ/M

LA LRE HSA

LA LRE M SA HZ/M

HZ/M

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

22 21

I

9 HM catalysts [] L M catalysts

20 19

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

.

.

.

.

.

.

.

18

17

9 C)

16 15 14 13 12 4

26

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

84

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

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

REFERENCES 1. 2. 3. 4.

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

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

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

87

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

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

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

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

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

Catalyst 2 PC1

Surface Area (m2/g)

FC1

Fresh

PC2

.

Zeolite

176

119

123

167

106

Matrix

83

41

39

60

43

Ni (ppm)

n.a.

611

344

n.a.

212

V (ppm)

n.a.

2505

1497

n.a.

991

Fe (ppm)

n.a.

4600

5400

n.a.

3714

Na (%)

0.28

0.19

0.26

0.27

0.23

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

9.4

2.01

1.89

8.1

2.03

2.4580

2.4245

2.4270

2.4510

2.4268

Combustion

--

Partial

Full

Partial

Catalyst inventory (ton)

"-

180

260

180

Daily fresh catalyst makeup (%inventory)*

__

1.5

1.8

1.7

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

__

0.23

0.21

0.054

""

36,100

34,900

38,000

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

665

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

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

wt%

Ni days

Metals (ppmw) Ni V

ZSA (m2g1)

K (MAT)* UCS (nm)

PC1 A B C D

19 29 28 24

130 85 60 31

950 625 437 229

3372 2875 2212 1612

53 120 140 152

1.03 1.89 2.28 2.53

2.4220 2.4238 2.4245 2.4275

FC1 A B C D E

20 24 21 18 17

143 80 59 50 38

657 368 269 230 175

2372 1595 1235 1060 880

42 128 149 154 161

0.84 1.99 2.51 2.38 2.76

2.4243 2.4257 2.4269 2.4277 2.4286

PC2 A B C D

34 27 28 11

173 97 82 33

313 175 148 59

1523 920 775 303

80 130 138 146

1.36 2.32 2.34 2.65

2.4240 2.4259 2.4282 2.4301

*see section 2.5

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

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

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

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

(i)

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

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

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

,A

o~" 1 0 0

,t,,

m

._z, ._ 80

~ .__

40

-5 E

20

0

I

0

0

50

100

150

200

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

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

>

1000

0

50

100

150

200

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

A = Ao / (kd T +1)

(2)

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

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

3.5

3.02.5 K

2.0 1.5-

A

1.0

II

0.5 ]

2OO

100

30

ZSA(m2/g) 3.5 3.0

II

2.5

K

2.0 1.5

A

1.0

II

0.5 I

0

100

i

200

300

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

95 70 60

J

5O ~" < 40 o~ N 30

i

20

i

10 0

0

r

i

i

50

100

150

200

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

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

(3)

1.2

-

~

0.8

9

0.4 0.0

-

-0.4 0

I

I

I

50

100

150

200

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

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

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

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

O E~ 221~ components. In a suitable setting the NExCC TM can also serve as a source of increased propylene production. While gasoline production is reduced by 40%, compensation comes from increased propylene, ether, and alkylate production. Multiple reactors can be operated in parallel with different feeds, with optimum catalyst and optimum process conditions. The construction of the cyclone offers a short and exactly controllable reaction time as the catalyst enters the cyclone simultaneously from each point of the riser top. In the NExCC TM reactor, the catalyst inventory is located in the diplegs. The inventory is much smaller compared with the FCC-units of equal capacity, which means a rapid response to market needs by enabling a fast catalyst change.

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

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

VGO 908

wt-% ppm

0.1

1.4 1100

Aromatics

wt-%

35

50

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

~ ~ ~ ~ ~

160 385 446 495 577

151 314 379 438 559

Density

kg/m 3

Sulphur Nitrogen

.

.

.

.

.

.

.

.

.

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

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

base

+2.2

-2.2

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

base base base base base base base base base

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

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

.

.

.

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4.2. The effect of oil partial pressure

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

0.10 0.60

0.17 1.07

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

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

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

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

-

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.

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

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

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estimated 3.2 kg.cat/kg air

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

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15 20 25 Particle size, microns

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i. . . . . . .

30

i

35

. . . . . . . . . . .

40

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

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

Fig. 7. Multi-entry cyclone test device.

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

123

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

Fig. 8. Calculated particle concentration profiles.

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

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

125

Fig. 9. Construction of the pulse reactor.

126

Fig. 10. Construction of the pulse regenerator.

127

Fig. 11. Construction of the TPO-regenerator.

128

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

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

10

100

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13. _.i

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(.9

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P,

.~_

I

~ 2o P,

o

r

.

b

0 ~0

4~

500

5~

600

6~

Reaction Temperature, ~

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

129

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

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

130

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

7. KINETIC MODELING

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

kl GAS OIL

"-GASOLINE PARAFFINS

k3

GASOLINE OLEFINS GASOLINE NAPHTHENES

COKE

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

GASOLINE AROMATICS

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

25Aromatics

Paraffins

20

(

]

~

15

~

A

a

[

Olefins / ~- 10

5 ~


643 K b.p. 494-643 K C5 - b.p. 494 K

H2, C1-C4 carbon deposit, CxH•

170

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

(5)

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

coke : e-C~~

(6)

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

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

d(@c~ d cc

= _kd. (~

conv

(Cc)) mcc

(7)

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

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

(8)

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

(9)

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

(10)

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

kdA +1

e(kd,4.Cc)

@conv,4 (Co) :

(l 1)

kdA + The deactivation c o n s t a n t kd, 4 represents the sum of the contributions of site coverage and pore blockage to the total deactivation, whereas kdArepresents the decrease in activity due to pore blockage relative to the deactivation by site coverage. (Note that whenkda5 x l 0 4 m/s, where no limitation to interfacial mass transport exists. Otherwise, for lower values (kg80A share ca. 88% of the total pore volume, but the pores in A-2 with diameter of were determined (figure 3) and correlated with surface area, generating a new model showing how conversion, gasoline and coke yield change with deactivation severity for catalyst E (tables 4 and 5). MAT results from the original catalyst E (old batch) were also used in the model calculations. 3

......

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

0

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4

6 Activity

Figure 2 - Coke versus activity for catalyst E

223

I

lDr,,,Oas+LPO+Coke I

......., Io

oiino I

[Drye~s+LpO I I Coke I

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

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

Old 20 153.1 0.045 70.2 51.0 850.0 12.9 4.7

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

224

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

100

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,..

t

._. 45

60

.~ 35

0 .~

m O

I1)

~ 4o

(5

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15 100 100

140 180 Surface Area (m2/g)

220

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

140

180

220

Surface Area (m2/g)

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

225

45

40 1

!

Ic/~ I /

30

//

~2o 8 lO 0

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

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

140

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180

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220

Surface Area (rn2/g)

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

15

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

0

,

10

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

Coke (w%)

20

30

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

3. Conclusions:

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

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

227

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

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

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

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

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

0.922 2.8 89.7 0.48 1.7 2.7 2000

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

~ 256 341 373 412 438 458 481 508 544 567

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

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

230

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

.

.

.

.

.

.

.

.

.

.

.

.

.

.

.

.

.

BET area m2/g

.

.

.

.

151 144 155 164 138 193 142 115 171 .

.

.

.

Matrix area

.

.

.

RE wt%

m2/g

Zeolite area m2/g

Unit Cell Size /k

30 35 36 40 75 86 46 73 92

122 109 119 124 63 108 97 42 79

3.31 1.64 2.37 2.34 0.79 0.59 1.62 2.12 2.07

24.31 24.28 24.26 24.31 24.31 24.31 24.31 24.33 24.29

.

.

.

.

.

.

.

.

.

.

.

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

231 0.5 A

~0.4 i

i

0.3

o 0.2 m ~0.i o

0

25

50

75

i00

IV~trixSurfac~Area (m2/g)

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

232

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

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Zeolite Surface Area (m2/g)

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

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

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A

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& Matrix

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

233

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

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Zeolite Surface Area (m2/g)

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

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

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

.-.18 dO

"-'16 ~3

"~14 9 0 ,-a 12

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Mesopore & M a t r i x Surface Areas

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

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

234

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

15

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

i

:

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(m~/g)

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

15

i ---13 do

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

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

235

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

25,

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i 9

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

50

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125

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(m 2 / g )

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

P. N20 .._..

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50 Surface

75 Area

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(m 2/g)

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

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

236 0.4

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w

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

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

0.4 dO

~0.3

"~0.2

O0.i

0.0

0

25 50 Matrix Surface Area

75 (m2/g)

i00

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

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

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

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

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

238 21.

22.

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

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

239

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

a

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

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

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

24I

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

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

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

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

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

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

3

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Stripping Time (Minutes)

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

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

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

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

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

chlorofo rm-d

SOFT

SOFT

COKE-

II

COKE

- I

PPM

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

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

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

wc

(,r

WC

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

250

200

150

100 PPH

50

L_~ 0.86* 0.88 0.90 single pulse excitation

0

-50

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

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

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

//"

4.5

3.5 ~_ 3 _~_

O

2.5 1.5 1

0.5 0

j 0

2

4

6

8

12

10

14

16

Stripping Time (Minutes)

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

Heavy decant oil (evaporative loss)

(--

Soft coke I

--~ kl

Light products

Soft coke II

--~ kl'

Light products

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

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

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

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

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

REFERENCES

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

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

251

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

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

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

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

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

2. EXPERIMENTAL

2.1 Feedstock and Product Preparation The liquid products and cokes investigated were prepared from a low and high sulfur vacuum gas oils (VGO) using a microactivity test (MAT) reactor (ASTM D3907, D5154). The low sulfur VGO was doped with 8% w/w toluene and benzene, and the high sulfur VGO one with 8% w/w toluene, but only 4% w/w benzene. The carbon in the methyl position of the toluene was labelled as opposed to all carbons in the case of the benzene (purity >99% for both the labelled compounds). Therefore, when the 0.8 g of feed used in the MAT experiments was doped with 8% toluene, the 13C enrichment equated to approximately 8 mg and, similarly for benzene, doping gave enrichments of either 30 or 60 mg excess lac. The reactor contained 4.08 g of steamed zeolite making the catalyst to feed (0.8 g) ratio equal to 5.0. The feed was delivered in 18 seconds to the reactor at 510~ before the system was flushed with nitrogen for 14 minutes. The liquid products were trapped in the collection bulb cooled with a beaker of iced water. 2.2 Analysis Compound specific ~13C measurements of the liquid products, generated from the MAT experiments, were carried out on a VG Isochrom II | GC-IRMS instrument as described by McRae et al [ 13]. A gas chromatograph is used to separate the individual components of the sample that then pass into a copper oxide combustion furnace. The resultant CO2, produced for each separated species, is continuously analysed by a single inlet, triple-collector mass spectrometer. Differences in the isotopic composition of carbon-containing substances are expressed in the conventional ~5-notation giving the permil (%0) deviation of the isotope ratio of the sample (sa) relative to that of a standard (st), i.e.

253

J

613 Csa -" (13 C/12 C)sa -- 1 x 10 3 (13 C/12 C)st

(i)

The standard commonly used is Peedee belemnite (PDB), whose 8~3C value defines 0 %0 on the 8-scale. In this initial study isotopic compositions have also been expressed on an atomic percent basis (13C %) which represents the excess of 13C above the natural abundance:

i 13Csal 1 13c%-[ / 13Csa/ ooo + R~t~~-j

+1

1.11

(ii)

Where R~t is the 13C/12C isotopic ratio of PDB (0.011237) and 1.11% is the natural abundance of 13C. Helium was employed as the carrier gas and a temperature program of 40~ (5 mins) to 320~ (5 mins) at 5~ rain -~produced the best possible separation of sample peaks. For convenience compounds were grouped, according to the classes listed in Table 1. However, it was not possible to determine isotopic values for components eluting from the column within two minutes of the injection period. The liquid products were also analysed by normal GC-MS (Varian 3400 GC linked to a Finnigan MAT TSQ70 triple quadruple MS, ionising energy 70eV; ion source temperature 150~ transfer line temperature 300~ to identify the major constituents present. Table 1 Approximate boiling point distributions and compositions of the liquid products. Group Boiling Point Product Distribution Distribution (~ 1 80 - 130 C6-C 8 alkanes, cyclo-alkanes, toluene 2 130 - 180 C9-C10 alkanes, C2-C 4 alkyl benzenes 3 180 - 220 Cll-C~2 alkanes, C4- Cs alkyl benzenes, naphthalene 4 220- 270 C13-C14 alkanes, C~-C2 substituted naphthalenes 5 270- 340 C15-C21 alkanes, substituted naphthalenes 6 270- 340 C~5-C21 alkanes,C2-C4 substituted naphthalenes 7 340 3-membered polycyclics i.e. anthracene, phenanthrene 8 340- 360 Alkyl substituted 3-ring polycyclics 9 340- 360 Alkyl substituted 3-ring polycyclics 10 360 -450 Alkyl substituted 3-4 ring polycyclics

254 Equation (iii) was used to calculate the mass of 13C incorporation into the liquid product groups. The mass of each group was determined from normal GC analysis of the liquid products doped with an internal standard (triacontane, C30H62). A Carlo Erba 4130 instrument was used, equipped with an FID detector and an SGE 25m fused silica capillary column coated with BP-1. The same temperature program used for the GC-IRMS analyses was employed.

13C Mass =

Mass Standard / Respo---~se~ S-i----andard,]x Response Area Group x 13C O~

(iii)

Unfortunately, due to excessive evaporative losses, solution state 13C NMR was not carried out on the liquid products detailed above. Therefore, a second series of samples were generated to obtain ~3C NMR data; the low sulfur feed was doped with 4 % ~3C toluene and 8 % ~3C benzene, respectively. The spectra for the whole liquid products were obtained in deuterochloroform (50% solution) with the free induction decays being collected in 32K data sets over a spectral width of 16 kHz, using a 40 ~ pulse, a pulse delay of 2.5 s and a 0.524 s acquisition time. A known weight of TKS (tetrakistrimethylsilane) standard was added to a predetermined weight of liquid product. In order to quantify the extent of 13C incorporation into the major peaks of the labelled samples, the ~3C NMR spectra of the corresponding unlabelled samples were also obtained. These quantitative experiments were run overnight with a small amount of 0.1M solution of chromium acetylacetonate [Cr(AcAc)3] in deuterochloroform added. All the spectra were processed using an exponential multiplication with a 10 Hz line broadening factor. Sealed-tube combustion of the coked catalysts and ~3C labelled liquid products was conducted using copper oxide in a sealed evacuated tube at a temperature of 850~ The tube was then broken under vacuum, combustion products collected and directed through a series of vacuum lines. Water was removed from the combustion products by a dry ice/acetone slush trap. The remaining carbon dioxide was collected for stable isotope analysis by mass spectrometry. Equation (iv) was used to determine the mass of ~3C present in the coke. To calculate the enrichment of '3C, the mass of ~3C present in the cokes derived from feeds doped with ordinary toluene and benzene was subtracted from the value obtained in equation (iv). ]3C Mass = (Carbon Content Coked Catalyst x Mass Catalyst Used)x ~3C %

(iv)

3. RESULTS AND DISCUSSION

3.1 Liquid Products Figures 1 to 4 present the distributions of the 13C labels, expressed as the excess atomic percent (Equation (ii)), for the different product groups (Table 1). The general similarity between the profiles, implies that the sulfur content of the feed did not markedly affect the product distribution.

255

Figure 1. Distribution of 13C label based on ~3C % values (Equation (ii)) from the low sulfur feed doped with 8 % ~3C toluene.

Figure 2. Distribution of ~3C label based on ~3C % values (Equation (ii)) from the low sulfur feed doped with 8 % 13C benzene. Table 2 contains the 8~3C data on both a permil and an excess atomic percent 13C basis. To illustrate the portion of the starting enrichment being incorporated into each group, the mass of ~3C is expressed as a percentage of the total mass of ~3C added to the feeds. These results are summarised in Table 3. The accumulative errors for the balance are estimated to be at

256 least approximately + 10%. These are due to inherent errors associated with the integrated peaks in normal GC analyses, coupled with the fact that in GC-IRMS, the smaller peaks are below the threshold level for detection. Also, due the transfer times of over 30s from the GC though the combustion furnace to the mass spectrometer, the peaks are considerably broader than in normal GC analysis.

Figure 3. Distribution of ~3C label based on ~3C % values (Equation (ii)) from the high sulfur feed doped with 8 % ~3C toluene.

Figure 4. Distribution of ~3C label based on 13C % values (Equation (ii)) from the high sulfur feed doped with 4 % ~3C benzene.

257

Table 2 Isotopic ratios (813C, Equation (i)) and the concentrations of excess 13C (Equation. (ii)) for the liquid products Feed Low sulfur doped with High sulfur doped with Group 8% 13C C7H8 8% 13C C6H6 8% 13C C7H8 4% 13C C6H6 813C 13C% 813C 13C~ 813C 13C~ 813C 13C% 1 2 3 4 5 6 7 8 9

2490 3110 149 29.8 42.6 44.3 17.1 -3.2 -0.1

2.66 3.30 0.16 0.03 0.05 0.05 0.02 r.-

(Pretreot) Reoqlor I F

~ P i ......... ~:

(Crocking)

,s

Reoctor Lr'I ..

>

-

$omNing loop

[

..... ooo~

:2

g o1_! tD ga 9s ~ 50 s')

_

~a

tD u J..

o

0 200

300

400

500

Temperature in reoctor [ , *C

n-El~tcne Hydrogen

-6

.5 .4

-3

I0~0 10 tO

-~ ~0

.~ )0

Otefin Concenlrahon (relo)ive)o butane )

Fig. 1. Weisz demonstrates bifunctionality in butane cracking

3. RADICAL CATALYSIS IN ISOBUTANE CONVERSION In a seminal study on the conversion of isobutane over solid acids G. B. McVicker et. al. proposed that some solid acids catalyzed radical-like chemistry (2). This was very provocative since radical chemistry was generally believed to be thermal and not catalytic chemistry. McVicker showed that thermal decomposition of isobutane results in equilmolar production of methane and propylene and also butenes and hydrogen. He then proceeded to demonstrate that amorphous solid acids such as alumina, silica-alumina, and halogenated aluminas produced the same products in the same proportions but at lower temperatures indicating that these materials were indeed catalysts for radical-like chemistry (Table 1). The halogenated aluminas were most active in promoting the radical-like catalysis. This demonstrated that not only did radical catalysts exist but that it is conceivable to formulate fairly high activity versions of such catalysts. McVicker went on to postulate that the radical mechanism involved electron acceptor sites on the solid acid catalyst which resulted in the production of radical cations. The radical cations subsequently decompose into methane and propylene or into butenes and hydrogen.

266 Table 1 Various amorphous aluminas act as radical catalysts in isobutane conversion CATALYST

NONE

AL203

SIO2-AL203

0.9% CLAL203

TE,MPERA'D,YRE. K

0.9% FAL303

923

873

873

873

823

1.6 ! .8 2.5 2. ! 9 3.7

1.6 1.4 4.2 2. ! 3 4. I

1.5 ! .7 4.9 2.12 5.5

8. I 7.4 7.6 2.25 6.5

7.3 7.0 9.0 2.22 14.2

-

0.1 0.2

0.1 TR

0.6 0.3

RADICAL PRODUCTS, MOL% METHANE PROPYLENE BUTENES H/C RATIO I-I2 (ESTIMATE) (~ARBONIUM ION PRODUCTS, MOL% PROPANE N-BUTANE PENTANES I_SOBUTANE, MOL% MINOR, MOL%

TR

-

94.0

92.6

91.4

76.0

74.7

0.1

0.2

0.2

0.8

I.!

REFERENCE2

The choice of isobutane as a molecular probe for radical catalysis was inspired. Isobutane is not very susceptible to conversion by carbonium ion chemistry since it cannot undergo beta scission. Therefore it is well suited to probe other catalytic sites which may be present near the carbonium ion sites. 4. C A R B O N I U M ION C A T A L Y S I S IN I S O B U T A N E C O N V E R S I O N

McVicker also studied the conversion of isobutane over faujasite (2). He used a sample of Union Carbide's LZ-Y82 (ultrastable Y, USY) for this purpose. Thermal treatment used in the production of LZ-Y82 results in some of the zeolite framework alumina being driven from the structure to extra framework positions as an amorphous component (zeolite nonframework alumina). Over this zeolite he noted the appearance of methane and propylene but also large quantities of saturated products (see Table 2). He proposed that the saturated products observed with faujasite arose from carbonium ion assisted chemistry. Implicit in McVicker's conclusion is that two catalytic chemistries are operative in the conversion of isobutane over LZ-Y82. One is radical-like in character and the other is carbonium (carbenium) ion chemistry. He further proposed that radical catalysis is prerequisite to carbonium ion chemistry. On the amorphous solid acid catalysts he found very little production of saturates indicating little or no carbonium ion assisted chemistry for the conversion of isobutane over these materials

267 Table 2 Faujasite produces secondary carbonium ion products from isobutane cracking 1,5% F-AL203

FAUJASITE

CATALYST

CONDITIONS TEMPERATURE, K CATALYST, GM _RADICAL PRODUCTS. MOL% .. METHANE PROPYLENE BUTENES H/C RATIO H2 (ESTIMATE)

823 1

773 1

773 1

773 0.3

723 1

9.7 8.2 9.7 2.23 17.0

! .g 2.0 2.8 2.19 4.7

2.5 0.9 ! .5 2.52 0

1.0 0.7 0.8 2.48 0

0.4 0.2 0.7 2.48 0

1.3 0.4 TR

0.3 0.3 TR

I 1.6 7.9 3.9

2.5 2.2 i .0

2.7 4.0 ! .7

68.5

92.4

69.5

91.2

89.8

CARBONIUM ION PRODUCTS, MQL% PROPANE N-BUTANE PENTANES !SOBUTANE. MOIL.*/, REFERENCE 2

It is instructive to examine the faujasite data of Table 2 more closely. At 773 ~ K yield data are provided for two different catalyst charges, 1.0 and 0.3 grams. Isobutane conversions in these experiments were 30.5 and 8.8 mol.% respectively. The magnitude of the various yield changes with changing conversion can provide information about initial and secondary products (3). For instance a reduction in conversion will reduce secondary products more than primary products. Note the large reduction in the saturates (carbonium ion products) as conversion is decreased by a factor of ca. 3. This indicates that the carbonium ion products are secondary products of isobutane conversion. They are not formed directly by some protolytic catalysis but certainly by some further chemistry involving reactive intermediates. Note also that methane changes more nearly in proportion to conversion. This suggests that methane is a primary product of isobutane conversion. If methane is a primary product of conversion then its co-product would be propylene. Methane and propylene have been found to be radical-like products of isobutane conversion over amorphous solid acids. They are likely formed as initial products of radical catalysis by the zeolite nonframework alumina. As mentioned earlier zeolite nonframework alumina is an amorphous alumina. It would be expected to have high Lewis acidity. Methane is a stable initiation product. Propylene is a reactive initiation product which can pick up a proton from a Bronsted site to become a carbenium ion. From this perspective it is more accurate to term the Bronsted acid catalyzed chemistry carbenium rather than carbonium ion chemistry. Propylene can convert to other molecular weight olefins by disproportionation. All olefins can convert to saturates by hydride transfer. Hydride transfer is a form of carbenium ion chemistry (4). As reactive intermediates, olefins are both being formed (radical catalysis) and destroyed (carbenium ion catalysis). Thus they should change

268 less with conversion than either stable primary products (methane) or stable secondary products (saturates). The data of Table 2 show this to be true. 5. RADICAL CATALYSIS BY LEWIS ACID SITES Another study reports on the conversion of isobutane over various amorphous solid acid catalysts (5). This study reported a strong correlation between a rate constant for isobutane conversion and the Lewis acidity of various amorphous solid acid catalysts. The rate constant for isobutane conversion was found to increase in direct proportion to the concentration of Lewis acid sites (Figure 2). The simplest explanation is that the Lewis acid sites catalyze the radical-like conversion of isobutane. An alternative explanation is the Lewis acid sites simply enhance the acidity of the Bronsted sites by charge withdrawal (3, p. 69). However if this were the case one should see evidence in increased production of saturates. This is not seen.

0[

T.s

E.

;c

1.0

o u o N

t

.

0.5

|

) IO4 .x

,,

2 4 6 8. Concenlrolion of Lewis ocid Qroq)s,mn~lm 2

Fig. 2. Lewis acidity drives isobutane conversion on amorphous solid acid catalysts

6. INITIATION/PROPAGATION IN BIFUNCTIONAL ISOBUTANE CRACKING McVicker further reports on isobutane partial pressure effects (2). He observes that the partial pressure dependency over the amorphous solid acids is first order for methane and

269 olefin formation while the appearance of saturates over faujasite is second order. This may be regarded as evidence that initiation chemistry is monomolecular while the propagation chemistry is bimolecular. Initiation by demethylation or dehydrogenation would be expected to be monomolecular. On the other hand the secondary conversion of light olefins has been found to be by dimerization/recracking, a bimolecular reaction (Figure 3)(3). Also the conversion of olefins into paraffins by hydride transfer would involve bimolecular reactions. This is further evidence that initiation of isobutane conversion involves radical-like monomolecular chemistry while propagation is provided by bimolecular carbenium ion chemistry.

0.7-~

I UN~MOt.ECULAqo ~'xJN

_ o.s-

~vt

'~'. 0.5 -

E ~ 0.4 ~,

0.3-

/ X__. /

0.2 -

CRAC,,,,o

0

o.I-

" ,,0 9. ~ ~

0 -- 0 ~ " 9

/

\(~

" ~

~'--~------O

0,----0

I.

I _

I ,

.I

2

3

4

5

I, 6.

I

"I

7

8

J

9

CARBON NUMBER

Fig. 3. Light olefins convert by bimolecular cracking

In fact McVicker goes on to propose kinetics which incorporates both initiation and propagation catalysis in the conversion of isobutane (2). In his formulation the initiation catalysis provided by the electron acceptor sites is found to be multiplicative to the propagation catalysis provided by the faujasite in facilitating total isobutane conversion activity. Thus not only do solid acid catalysts provide catalysis for both radical-like and carbenium ion assisted chemistries, but the chemistries interact synergistically to provide the overall conversion of isobutane. All the above information is assimilated in an understanding in which initiation is associated with the formation of olefins by dehydrogenation of isobutane over Lewis acid

270 electron accepting sites. These olelfins, once formed, diffuse to Bronsted acid sites, pick protons and become carbenium ions. Conversion then proceeds by beta scission. In this schema there are distinct sites involved in the initiation and propagation chemistries. No conversion will occur if the initiation of olefins does not happen. No propagation will occur without the presence of initiated olefins. Conversion involves the formation of olefinic intermediates. The appearance of saturates is due to secondary hydride transfer reactions catalyzed by Bronsted sites. 7. HISTORICAL EVIDENCE FOR BIFUNCTIONAL CAT CRACKING CATALYSIS Bifunctional catalysis is well understood. As noted above, Weisz was prominent in the development of such understanding for metal/acid bifunctional catalysts (1). A key feature of bifunctional catalysis is a synergistic interaction between the two types of sites in providing activity. A catalyst activity synergy between amorphous silica-alumina and faujasite in composite cat cracking catalysts has long been noted (7). The present authors believe that the interaction between radical initiation and carbenium ion propagation is the explanation for this synergy. More recent evidence is provided by the use of amorphous alumina as an activity enhancement in zeolitic cracking catalysts. By itself alumina has been found to provide very low cracking activity (4, p. 46). However when combined with faujasite in a cracking catalyst it much enhances the activity of the composite catalyst (Table 3). For example the inclusion of 20 wt.% alumina in a catalyst containing 20 wt.% USY increased the catalyst activity 78%. This improvement is much more than would be expected from the additive contribution of a very low activity ingredient. It is suggestive of a synergistic contribution as may arise from bifunctional catalysis. The authors contend that the amorphous alumina has very high Lewis acidity but very little Bronsted acidity. Thus it provides excellent initiation but little ability to propagate the conversion. Faujasite, on the other hand, has a much larger ratio of Bronsted to Lewis acidity. Thus it provides much higher propagation to initiation capability. Combining both the amorphous alumina and the faujasite in a common catalyst provides a better balance of initiation and propagation than either ingredient alone; hence improved activity. Table 3 Addition of matrix alumina much increases activity o f zeolite cracking catalyst CATALYSTCOMPOSITION.WT% ULTRASTABLEZEOLITE AL203 SI02-AL203 GEL

20 0 80

20 20 60

50 1

64 1.78

pERFORMANCE 430~ CONVERSION, VOL.%

SECONDORDERACTIVH'Y REFERENCE 8

271 8. ZEOLITE N O N F R A M E W O R K ALUMINA AS A RADICAL CATALYST Zeolite nonframework alumina is derived from zeolite framework alumina which has been driven from the framework by exposure to severe hydrothermal conditions such as exist in a cat cracker regenerator. The academic community has long been aware of a synergistic interaction between zeolite framework and nonframework alumina (9). Prominent explanations for the enhanced activity associated with zeolite nonframework alumina involve chemical interactions between the framework and nonframework aluminas which enhance the activity of the framework sites. The present authors propose that the nonframework alumina is an amorphous alumina that has very high Lewis acidity and thus very high initiation activity but low propagation activity because of its low Bronsted acidity. Its high initiation activity complements very nicely the very Bronsted acidity of the zeolite framework sites. The two activities are multiplicative in providing enhanced bifunctional zeolite activity. 9. GASOLINE COMPOSITION AS EVIDENCE FOR INITIATION Further evidence for initiation is found in gasoline composition data obtained on amorphous and zeolitic cracking catalysts (10). These data (Table 4) show that the amorphous catalysts, which have a greater ratio of Lewis/Bronsted acidity than the zeolitic catalysts provide much higher naphtha olefins and less of the paraffin and aromatic products of secondary hydrogen transfer. Hydrogen (hydride) transfer was noted earlier in this paper to be catalyzed by Bronsted acid sites. Early explanations for the lower naphtha olefins with the zeolitic catalysts focused on the high hydrogen transfer activity of these catalysts.

Table 4 Cat cracked gasolinecomposition reflectsradicalinitiationchemistry FEED CATALYST

CALIFORNIAVIRGIN GAS OIL

CALIFORNIA COKER GAS OIL

GACHSARAN GAS OIL

DURAB~ 5

DURABEAD1

DUP.ABEAD5

DURABEAD1_

.DURABEAD5

D U R A B ~ 1_

21.0 19.3 14.6 45.0

8.7 10.4 43.7 37.3

21.8 13.4 19.0 45.9

12.0 9.5 42.8 35.8

31.9 14.3 16.3 37.4

21.2 15.7 30.2 33.1

GASOLINE, VOL.% PARAFFINS CYCLOPARAFINS OLEFINS AROMATICS

"

DURABEAD 5 IS EARLYGENERATIONZEOLITE(REHX) DURABEAD 1 IS AMORPHOUS SILICA-ALUMINA REFERENCE 10

It is, however, possible to set forth expectations for naphtha olefinicity based on kinetic analyses. One begins by writing out the differential equation for the production and destruction

272 of naphtha olefins. This is accomplished in equation I below. dCN=/dT = klkp(1 -X) z -

k H C N --

-kxCN=

(1)

In this equation CN= is the yield of naphtha olefins, T is the contact time, kl is the initiation rate constraint, kp is the propagation rate constant, X is the fractional conversion of the feed, kH is the hydrogen transfer rate constant, and kx is the rate constant for the recracking of naphtha olefins to light olefins. Thus the equation depicts the formation of naphtha olefins by the second order conversion of feed molecules. Both initiation and propagation are involved in the conversion of feed molecules into naphtha olefins. The equation also depicts the destruction of naphtha olefins by hydrogen transfer and recracking. Both of these destruction steps involve only propagation or Bronsted sites. The maximum in naphtha olefins may be determined by setting the derivative in equation I equal to zero. The maximum in naphtha olefins is then determined by equation 2. CN=Max = klkp(1 -XMax) 2/(kH+kx)

(2)

Note that in this equation k p , kH, and kx are all related solely to propagation or Bronsted acidity in the numerator and denominator of the equation. Thus the Bronsted acidity effectively cancels out and the maximum in naphtha olefins is expected to be a function of initiation or Lewis acidity only. This kinetically determined expectation is in direct opposition to the opinion that naphtha olefins are set by the hydrogen transfer activity of the catalyst. This is very important for catalyst design. One increases naphtha olefins by increasing the initiation activity of the cracking catalyst not by minimizing Bronsted acidity. 10. ZEOLITE N O N F R A M E W O R K ALUMNA AS SUPERIOR RADICAL CATALYST Some convincing corroboration of bifunctionality in cat cracking catalysis is provided by data generated by Lercher (11). In this publication Lercher provides MAT data on a series of commercial cat cracking catalysts. The MAT test is defined by ASTM method 3907 and uses a standard vacuum gas oil (VGO) as feed as opposed to model compounds. Complete yield structures were provided at constant conditions. From this data it is possible to calculate rate constants for the hydrogen transfer which converts isobutene to isobutane (the ratio of isobutane to isobutene at the constant conditions of the MAT test). Summarized data are provided in Table 5. These data provide a six fold variation in hydrogen transfer activity. These hydrogen transfer rate constants were taken to be a measure of the carbenium ion activity of the subject catalysts. In addition second order rate constants for overall conversion at 421~ F were calculated. These rate constants were calculated by excluding the aromatic content of the feed (23.7 wt.%) as unconvertible. The rate constants for overall conversion covered a range of nearly five fold in activity. Thus by these two measurements we have quantified indications of a rate constant for initiation (radical activity) multiplicative to a rate constant for propagation (carbenium. ion activity) in overall conversion and a rate constant for

273

carbenium ion activity as measured by isobutene hydrogen transfer. Simple division of the former by the latter should provide a measure of the initiation (radical) activity. Table 5 Deconvolution of carbenium ion and radical activities f,ATAI,X~

UTE

UTA

UTO

trrc

UTS

UTR

5.68 1.68 1.98 71.5

3.07 2.40 1.22 63.6

3.39 2.23 1.38 61.1

4.76 2.03 1.79 68.6

6.75 0.88 2.97 72.5

4.66 1.79 1.66 64.2

3.42 14.9 4.35

1.28 5.02 3.92

1.52 4.01 2.64

2.35 8.93 3.81

7.67 18.8 2.45

2.61 5.32 2.04

YIELDS.WT.% ISOBUTANE ISOBUTENE COKE 421 F CONV. RATE CONSTANTS IC.4 HYDROGEN TRANSFER 421 F CONV. RADICAL ACWIVITY REFERENCE 11

The results of this computation are provided in Figure 4. This figure is a plot of radical activity versus zeolite nonfrarnework alumina. Notice the strong correlation. This suggests that zeolite nonframework alumina is a very strong radical catalyst. Zeolite nonframework alumina is known to have predominantly Lewis acidity (9). Thus our exercise sought to compute a measure of initiation catalysis and the result identified a prominent Lewis acid material as the source of such catalysis.

I

4.5

I

3.5

/

> v3 o < J 2.5 < _O 2 es

/

-

v

1.5

1 0.5

0

0.5

1

t.5

2

2.5

3

3.5

4

ZEOLITE NONFRAMEWORK ALUMINA, WT.%

Fig. 4. R a d i c a l activity v e r s u s z e o l i t e n o n f r a m e w o r k alumina

4.5

274 Linear regression of the radical catalysis rate constants against catalyst properties showed that matrix alumina may also play a role in the radical catalysis measured with the feed used and at the conditions studied in Lercher's report (11). The regressed equation relating radical contributions to catalyst properties are shown on Figure 4. The coefficient is much smaller for matrix relative to the zeolite, nonfrarmework alumina. However there is much more matrix than zeolite nonfrarmework alumina in these catalysts. In one catalyst (UTR) the matrix provides more than a third of the total initiation activity. The very high activity of the zeolite nonframework alumina may be due a combination of very high intrinsic activity and its location in close proximity to the propagating zeolite framework sites. How well the bifunctional model fits Lercher's MAT data is shown in Figure 5. His figure presents a plot of actual versus predicted 421~ F conversion. The authors believe that the excellent fit is strong corroboration for radical initiation and carbenium ion propagation in cat cracking catalysis.

75 i g w

Y

70

y = 1.00Z 0 u w w IIL w

Al

65

7

u o w Fa w w a.

50

55

60

65

70

75

ACTUAL COKE FREE CONVERSION,WT.'/,

Fig. 5. Actual versus predicted 421 F coke free conversion

11. COKE SELECTIVITY IN BIFUNCTIONAL CAT CRACKING CATALYSIS

Coke selectivity is the ratio of coke yield to the second order conversion activity. It is a very important parameter because most cat crackers operate at a coke burning limitation. At this limitation improved (lower) coke selectivity can provide conversion credits in the reactor. The MAT data of Lercher (11) provide some interesting insights on coke production. In Figure 6 we plot the MAT coke yield from the six catalysts of this study versus the square root of the carbenium ion activity of these catalysts. Recall that isobutene hydrogen transfer was

275 used as a measure of carbenium ion activity. Note the strong correlation. This suggests that most of the coke made from the feed used in this study is associated with carbenium ion activity provided by the Bronsted acid sites. Likely much of this coke is due olefin oligermerization/condensation chemistry. The square root relation likely results from a Voorhies type coke make formulation in which the rate of coke formation is inversely related to the amount of coke on the catalyst. 3.5

2.5

v i.08:

z~

.
0 equal to the concentration in the bulk phase times the ratio between the saturation concentration inside and outside the sphere. 3. the initial (t = 0) concentration within the spheres is identically zero. 4. The concentration and flux are equal at the interface between the shell and interior. Reaction rate, per se, was not considered, but was instead represented by a Freundlich adsorption isotherm: 1

C,d, = re'C"

Two separate isotherm cases were examined:

328

1) Linear Isotherm Here, n = 1, and a continuous reaction with no time dependency is represented. There is no deactivation and a continuous coke gradient results. 2) Shrinking Core Here, where n >> 1 (50), a very fast reaction rate relative to diffusion is represented. Deactivation is rapid and a "black and white" coke gradient is observed. Practice will lie somewhere between these two models. Assumptions in this model include: 9 9 9 9 9

No interparticle mass transfer Only adsorption No reaction kinetics No desorption of reaction products No influence of flux of reaction products on intraparticle mass transfer

The general conclusions are represented in Figures 15 and 16 for the two isotherm cases. The utilization of the particle, quantified by the effectiveness factor Cp(t) Eff ectivenness

F a c t o r = c p (equilibrium)

is plotted against the Fourier Number Ff--w

D.t dp2

in which D t dp

= = =

Diffusion coefficient time Particle diameter

The skin was assumed to occupy 2% of the particle radius. Fourier numbers were calculated for various carbon number molecules by estimating the diffusivity and assuming a particle size of 60~ and a time of 3 sec. Effectiveness factors were calculated for a number of cases representing lower relative diffusivity in the shell versus the bulk. Comparing Figures 15 and 16 we see that there is not much effect of the isotherm assumption on the results. As expected, as contact time is lowered, the effect of a

329

less permeable shell is magnified. Obviously, the lower the permeability of the shell, the more dramatic the effect. For a molecule with 40 carbon number, modest reductions of 10 to 20% in catalyst utilization are calculated. For an 80 carbon number, the penalty calculated for the lowest shell permeability is around 40%. It should be realized that the diffusivities are simply estimates and the calculations only serve to illustrate the plausibility of the penalty of a low permeability on catalyst performance. However, the relative carbon number calculations suggest that the penalty will be most severe on bottoms cracking.

6. MASS TRANSFER AND HIGH ACCESSIBILITY CATALYSTS

If, in fact, a low permeability shell is responsible for loss in performance in highly contaminated FCC catalysts, two remedies are possible: A catalyst architecture that renders the shell more porous, or A catalyst architecture that significantly increases the bulk mass transfer capabilities.

Figure 15. Catalyst Particle Utilization Assuming Linear Isotherm

330

Figure 16. Catalyst Particle Utilization Assuming Shrinking Core Model

Figure 17. Higher Accessibility Catalyst Retains Advantage as an Fe is Added

331

It is somewhat difficult to conceive of how to achieve the first remedy. Possibly the surface composition could be engineered to resist glazing, but there is no evidence to date of this being the case. Certainly, if the permeability of the shell were low enough, increasing the mass transfer in the interior of the particle would never help. However, if the mass transfer rates in the two regions were not too disparate, the second remedy could be effective. This seems to be the probable case based on experimental data to date. This is illustrated in Figure 17 in which AAI data for catalysts that have been contaminated with Fe are plotted. It appears that catalysts with high accessibility structures maintain their advantage over less accessible structures contaminated with incremental Fe that either forms a shell or further defines an already present shell. From a modeling point of view, this can be visualized by starting with the low skin permeability curves of Figures 14 and 15 and observing what happens if the diffusivity ratio is increased. In the following table, some cases in which high accessibility catalysts were utilized to improve performance in high contaminant environments are summarized:

Case A1 A2 B1 B2

Fe (wt%) 0.8 0.8 ....

Delta Conversion LA to HA (wt%) 5 4 2.1 2.0

Delta Bottoms LA to HA (wt%)

-3.4 -2.5

In all of these cases, neither the zeolite character nor content was changed appreciably. Improved performance can be ascribed almost totally to improved accessibility.

7. CONCLUSIONS

It has been shown that shells exist in both fresh and contaminated FCC catalysts that have the potential to negatively affect performance. Based on microscopic examination, these shells appear to be quite dense with the probability of a significantly lower permeability. A mathematical model representing diffusion into a catalyst particle with such a shell indicates the plausibility of this argument. Depending on parameters assumed, catalyst utilization penalties of up to 40% were predicted for large petroleum molecules.

332

The solution to resisting this problem commercially appears to be utilization of a catalyst with a high initial accessibility. These catalysts appear to maintain their accessibility advantage over less accessibility catalysts as a contaminant shell is formed or better defined. 8. REFERENCES

1. P. O'Connor, P. Imhof and S.J. Yanik, Catalyst assembly technology in FCC. Part I: A review of the concept, history and developments. This volume, p. 299. 2. M.O. Coppens, Catalysis Today 53 (1999) 225. 3. M.C.J. Hodgson, C.K. Looi and S.J. Yanik, Akzo Nobel Catalysts Symposium 1998, Paper F-4. 4. G.W. Young and K. Rajagopalan, Ind. Eng. Chem. Process Dev., 24 (1985) 995. 5. P. O'Connor and F.W. van Houtert, Ketjen Catalyst Symposium 1986, Paper F-8.

333

Keyword Index

Accessibility, 71, 279, 299, 311 Accessibility measurement, 209 Acidity, 41, 263 Activity, 219, 263 Alumina, 41 Basic metal oxide, 3 Bifunctionality, 263 Butylenes, 111 13CNMR, 251 Catalyst, 71, 263 Catalyst activity, 209 Catalyst aging, 87 Catalyst assembly, 299 Catalyst deactivation, 187, 201, 219, 227 Catalyst preparation, 293, 299 Catalyst structure, 311 Catalytic coke, 239 Coke, 187, 239, 251 Coke analysis, 251 Coke formation, 167 Contact time, 153 Crystal size, 279 Deactivation, 87, 133, 311 Dealumination, 3 Entrainment coke, 239 Equilibrium catalysts, 201 Extraction of coke, 239 FCC contaminants, 133 Feedstock, 71 Fluid bed, 111 Fluid cracking catalyst preparation, 59 Gasoline composition, 153

Hard coke, 239, 251 High rare-earth on Y, 141 High silica alumina ratio, 107 History, 299 Hydrogen transfer, 133, 141, 153 Hydrothermal treatment, 59, 87, 187, 219 Industrial units, 87 Kinetic models, 187 Kinetics, 153, 167, 209, 251 Lab scale pilot plant, 167 Labelling, 251 Light olefins optimization, 111 Low-alumina Y (LAY) zeolites, 107 Mass transfer and diffusion, 311 MAT testing, 209 Mathematical model, 87 Mathematics, 187 Matrix, 263 Matrix properties, 71, 227 Mechanisms, 153 Metal oxides, 107 Metals, 201, 219, 227, 263, 263, 311 Micro-riser, 167 Modelling, 311 Models, 219 New test unit, 153 Nickel resistance, 201 Non-framework alumina, 3, 41, 59 Olefins cracking, 141 Olefins in gasoline, 141 Olefins optimization, 133 Organized mesoporous silica, 293

334 Particle binders, 293 Particle formation, 299 Pilot plant, 111 Pilot riser testing, 227 Pore size distribution, 299 Probe molecules, 279 Propylene, 111 Reactor, 111 Resid cracking, 227 Riser testing, 71 Short contact time, 71, 111, 167, 311 Silica, 293 Silica-alumina ratio, 3 Silica structures, 293 Sintering, 311 Soft coke, 239, 251 Solid state aluminum 27 NMR, 41, 59 Solid state silicon 29 NMR, 41, 59 Strippability, 209, 239 Structural design, 299

Sulfur in gasoline, 153 Sulfur species, 153 Surface area, 219, 279 TCC, 107 Total olefins management, 141 Tri-iso-propyl-benzene, 279 Ultrastable zeolites, 3 USY, 3 Vanadium, 133 Vanadium tolerance, 201 Zeolite aging, 59 Zeolite deactivation, 59 Zeolite properties, 227 Zeolite synthesis, 41 Zeolites, 3,263, 279 Zeolites, HY, 41, 59 ZSM-5, 141