New catalytic applications of zeolites for petrochemicals

H. Chon, S.I. Woo and S.-E. Park (Editors) Recent Advances and New Horizons in Zeolite Science and Technology Studies in Surface Science and Catalysis, Vol. 102 9 1996 Elsevier Science B.V All rights reserved.

323

NEW CATALYTIC APPLICATIONS OF ZEOLITES FOR P E T R O C H E M I C A I ~

C T O CONNOR, E VAN STEEN AND M E DRY

CATALYSIS RESEARCH UNIT, DEPARTMENT OF CHEMICAL ENGINEERING, UNIVERSITY OF CAPE TOWN, RONDEBOSCH, 7700, SOUTH AFRICA

1.

INTRODUCTION

The oil refining industry is presently facing a number of important challenges. Among these are the environmental laws relating to both the quality of the fuels it produces as well as the operation of its refineries. It is also being required to produce more gasoline and, at the same time, more products, such as light alkenes, to serve as feedstocks for the petrochemical industry. The emphasis of the present paper is on recent applications of zeolites for the production of alkenes and aromatics and their conversion to certain petrochemicals. Hence the focus will be on recent developments in the use of zeolites, firstly, in catalytic cracking, which is a key process in producing the classical building blocks of the petrochemical industry, viz. light alkenes and aromatics, and, secondly, on the conversion of these into higher-value products, via: 9 Alkylation of aromatics; 9

Aromatization ofalkanes/alkenes;

9

Skeletal isomerization ofn-butene;

9

Oligomerization ofalkenes;

9

Isomerization of long-chain alkanes.

Figure 1 shows a simplified flowsheet emphasizing the processes to be discussed in this paper. Developments in most of these processes have recently been extensively reviewed [1-7].

324

C=/C, ~

Crude

B'IX

LSR Naphtha V ''~ I Kerosine~ L Diesel -•Reforming] :Lube Oil-----1 C,= GO LVGOI Resid

iso-Cs

BTX

[;l

"I(c" c") Disti"atet

Alkylation ~Lubricating Oils] CCO HVG( HVGO

;~HyorocracKing~

Gases I

Tops I NapthaI Jet I Diesel ! Lubes I

Figure 1. Flowsheet emphasizing processes reviewed in this paper. One of the most important properties of zeolites is their ability to carry out shape selective reactions [5]. These can be classified as, firstly, product shape selective reactions in which the only products formed are those which can diffuse out of the pores of the zeolite, secondly, reactant shape selective reactions which occur when some of the molecules in a reactant mixture are too large to diffuse through the catalyst pores, and, thirdly, restricted transition-state selective reactions in which the only reactions which occur are those in which space exists in the pores or cavities to allow the formation of the activated transition state complex~ In some cases where the zeolite is three dimensional the size of the channel intersections will also be a determining factor. This unique catalytic property is related to the pore size of the zeolite and has led to the synthesis of zeolites with a very wide range of pore

sizes. Over the past fif~ years there have been si~ificant developments in the synthesis of zeolites of different pore sizes [6]. Many of these developments have resulted in new applications of zeolite catalysts for petrochemical processes. This is especially so, for example, in the case of catalytic cracking which ideally requires the pore sizes to be tailored to cater for reactions involving a wide range of molecular sizes thus leading to the synthesis of catalysts with macro-, meso- and micropores. Table 1 shows the pore size, dimensionality and structure type of some of the zeolites and molecular sieves which will be discussed in this paper.

325

Table 1 Structural details of some zeolites and molecular sieves Zeolite

Pore Dimension Dimensionality Structure Ring Size [Angstrom] Type [T atoms] ZSM-5 5.3 x 5.6 and 5.1 x 5.5 3-dim MFI 10-membered US-u 7.4 x 7.4 3-dim FAU 12-membered Mordenite 6.5 x 7.0 and 2.6 x 5.7 2-dim MOR ~ 8- and 12-membered Beta 6.5 x 5.6 and 7.5 x 5.7 3-dim BEA 12-membered ZSM-22 5.5 x 4.4 uni-dim TON 10-membered Theta-1 Femerite 4.2 x 5.4 and 3.5 x 4.8 2-dim FER ~10- and 8-membered Rho 3.6 x 3.6 . RHO 8-membered MCM-41 SAPO-11 KL

40 - 65 6.3 x 3.9 7.1 x 7.1

uni-dim uni-dim uni-dim

M41S AEL LTL

depends on synthesis lO-membered 12-membered

* 2 unconnected 3-dim. systems # Intersecting system

Zeolite catalysts are mainly used to promote acid catalysed reactions. The number and strength of acid sites can be controlled by methods such as ion exchange, de/realumination and isomorphous substitution of tetrahedral atoms. Their high thermal stability implies that they are able to be regenerated by the burning off of carbonaceous deposits although in-situ steaming can lead to dealu_mination of the zeolite which may result in a loss of catalytic activity. Finally they are able to act as hosts or carriers of guest atoms which can include transition metal ions or atoms, metal oxide clusters, complexes or chelates, etc. These unique properties enable zeolites to have a wide variety of applications as catalysts.

2.

CATALYTIC CRACKING

2.1

Introduction

The catalytic cracking of oil in modem refineries is mainly aimed at the production of gasoline and diesel and is also one of the main sources of light alkenes and aromatics. The worldwide crude oil processing capacity is about 600M tonne crude per annum which consumes about 300k tonne catalyst per annum There are some 450 catalytic cracking plants worldwide and this capacity is growing at a rate of about 8 000 m3/day [6]. Fhfidized catalytic cracking (FCC) constitutes by far the largest use of zeolite catalysts (some US$ 600M) and, in the US, for example, it is estimated that catalytic cracking accounts for 42% of the estimated US refining catalyst market in 1991 [8]. The impact of zeolites is shown in Figure 2 which shows the changes that have occurred in typical FCC yields as the catalyst has changed over the last fifty years [3].

326

Figure 2. Impact of catalyst type on yields

Developments in the manufacture of more efficient catalysts for FCC have been reviewed by various authors [1,2,3,4,6,9,10,11,46]. Generally the objective of most FCC operations is to maximize the production and the octane number of the gasoline t~action and to minimize the formation of by-products such as LPG, fuel gas, heavy fuel oil and coke. Since the focus of the present paper is on petrochemicals, however, only those developments in the design and formulation of FCC catalysts which are aimed at increasing the production of light alkenes, the key building blocks of the petrochemical industry, will be discussed. In the FCC process, which uses the solid catalyst, ultra-stable zeolite Y (USY), in a fluidized bed reactor, long chain alkanes, polycycloalkanes and alkylaromatics are cracked into alkanes, alkenes and cycloalkanes in the C 1-C 17 range. At the same time aromatics in the C6C12 range are formed by oligomerization, dehydrocyclization and the cracicing of polyalkylaromatics and oligomers. Table 2 shows a typical example of FCC yields.

Table 2 Typical yields from an FCC Unit Cut Yield [%] Gas (H2,C1,C~,C2=) 2-4 C3- LPG (65 - 75 % olefinic) 4-6 C4 - LPG (65- 75 % olefinic) 6-11 Gasoline 40-50 LCO 15-20 Bottoms 8-15 Coke 5-6

327

2.2

Role of the catalyst matrix

Modem refineries are required to process more and more heavy feedstocks with a constant increase in the amount of atmospheric resid fed to the FCCU. The consequent increase in the amount of Conradson carbon in the feed reduces the ability to achieve conversion by increasing regenerator temperature and reducing the cat/oil ratio. This problem can be addressed by a proper selection of catalyst in which the zeolite:matrix ratio is carefully tailored to achieve maximum bottoms cracking while not accompanying this with excessive dry gas or coke production. At the same time the matrix can be carefully controlled with respect to factors such as alumina or silica-alumina type, surface area, acidity, and pore size distribution. This will also require that the zeolite part is carefully designed with respect to the nature, content and distribution of extra-framework ahaninham along the zeolite crystals, as well as Si/AI ratio, rare earth content, and unit cell size. The FCC particle consists basically of about 35 wt % zeolite Y [12] incorporated into a matrix usually consisting of silica, which acts as a "glue", and ahtmina, which serves to crack the large molecules. The matrix often includes clay to provide the desired density. One of the roles of the matrix is to increase the resistance of the catalyst to metal poisoning. This will be discussed in Section 2.4. Increasing the matrix content can also increase the hydrothermal stability of the catalyst. Some of the properties which must be developed in the matrix are [13]: Macropores (>IO00A)of low activity to assist in the cracking of the large asphaltenes and to allow deposition/passivation of Ni and V in particular; Mesopores (30-1000A), with higher activity than in the macropores, which are able to cleave side chains from aromatics and napthenic rings and thus produce products boiling in the LCO (light cycle oil) range (220 - 360C); Small pores (<20A), with the highest activity, to crack straight chain alkanes.

Accessibility is a function of the size of the hydrocarbon molecule, the pore size distribution of the catalyst, the plugging of pore mouths as a result of coking, poisoning, contact time of catalyst and oil, etc. The macropores at the entrances to the channel strucu~es should have low activity and surface area but should also be able to act as a guide to transport molecules to the key functional sites in the smaller pores. Improving the accessibility of the feed molecules to the FCC catalyst is a primary area of current research. This improved access would result in better bottoms cracking, better resistance to poisons, improved stripping efficiency in the FCCU and reduced overcracking and secondary reactions. It has also been shown that the relative stripping efficiency can be almost doubled since the more open pore structure not only enhances diffusion into the catalyst particle but transfer out of the particle is faster as well. The improved diffusivity of molecules also results in reduced residence times of the reaction products in the pores and thus overcracking and secondary reactions are reduced. Figure 3 shows how the pore size distribution changes as the accessibility increases [14]. The catalyst with a very open structure and high accessa"oility allows the coke precursors to undergo cracking and thus coke formation is reduced. This implies that reactors with short residence times are more susceptible to coking since there is insufficient time available to bring the precursors into contact with the functional sites due to diffusional limitations.

328

Figure 3. Pore size distribution for catalysts of different accessibility. [ 14]

TONx 104

0

K [ r m "~]

i 10

i

1

20

30

0

i

i

i

i

10

15

20

25

AI/u.c.

AI/u.c.

a.

i

5

b.

Figure 4. Activity for cracking on steam (11) and SiCh([]) deahminated Y zeolites (a: nheptane; b" gas-off) [ 1]

329

Large gasoil molecules cannot penetrate deep into the zeolite pores and only an outershell of the cTystallites will be active for cracldng gasoiL Primary cracking is known to take place on the matrix [15]. Thus all the acid sites measurable by bulk techniques are not necessarily available for primary cracking. This is well illustrated in Figures 4 and 5 [1]. The steam dealuminated zeolite Y possesses mesopores which results in an increase in the ratio of external to internal surface area and thus to increased aceess~ility to acid sites. Thus although the steamed sample has less Bronsted acidity as measured by pyridine TPD [16] and thus shows a lower activity or TON for hexane cracldng, it has a higher activity. This implies that, although the steamed sample has fewer sites, the number of accessa"ole acid sites is greater on the zeolite with mesoporosity. Hence by increasing the ratio of external to internal surface area and generally by increasing the accessibility of molecules to acid sites in a controlled way it should be possible to increase the activity of a given zeolite Y [1]. A recent patent has described the treatment of an ultrastable zeolite Y to produce a catalyst with "primary" pores of about 50A and "secondary" pores of 100-600 A by steaming at about 600~ followed by washing in 0.SN nitric acid at 80~ for 3 h [17].

0.07 Steam

~-

0.05

o eq

i !i~i i ~

NNNI 0.03

E

o

"" I ~ ! ~ SiCI4

E ql

Q.

0.01

f --1

~10

~25

-0.01 ~d/U.r

Figure 5. Mesopore volume for Y zeolites dealuminated by different procedures [ 1].

The diffusion of molecules into the pores of zeolites is a function of their size (Figure 6) and in order to produce catalysts which are capable of selectively cracking the large molecules in crude oil it is necessary to make catalysts with pores large enough to accommodate such molecules. The opening to a pore of zeolite Y is about 7.5A whereas the average diameter of a heavy residue molecule is greater than 20A. The large pore 18 MR

330 molecular sieve VPI-5 theoretically represents a posm'ble candidate for cracking the large molecules. However, when desiring any FCC catalyst hydrothermal stability is of primary concern and it has to be stable when subjected to the high regeneration temperatures such as 720~ in the presence of steam VPI-5 has no acidity and when acidity is introduced by making the SAPO version of this catalyst, it is not thermally stable. Another large pore zeolite which could possa'bly satisfy the above requirement is MCM-41 which can be made with a pore size of between 20 - 100 ~ Recently a series ofMCM-41 zeolites has been prepared to investigate their application as FCC and hydrocracking catalysts [ 18] The Si/A1 ratio of these samples were 6.34, 6.09 and 5.65 and their mean pore sizes were 3.25nm~ 2.15 nm and 3.85 nm respectively. These pores are essentially cylindrical in nature. It was found that their steam stability as determined by meamuSng their surface area and pore volumes after 5 hrs steaming at 788~ was very low and they did not show any promise for resid FCC application. In a subsequent test the MCM-41 sample was used as a 10% additive to Y zeolite in a simple FCC formulation. This sample showed some promise in steam stability and micro - pore retention. In Table 3 some characteristics of the standard cracking catalyst and the catalyst incorporating the wide pore zeolite, MCM, are shown [ 19]. From these results it can be concluded that the catalyst is stabilized in the presence of the ultra-wide pore (UWP) zeolite. The surface area and micro-pore volume atter steaming indicate a bigger part of the Y zeolite is still intact thus leading to a higher activity in the MST test. However there is no indication of an enhanced bottoms conversion capacity [20]. The ultra-wide pore zeolites do not appear to result in an improved performance and, as mentioned above, it would also be necessary to improve the stability of these catalysts drastically in order to enable them to survive the severe regenerator conditions of the FCC unit.

De, x 1017 [m21s] 100

~

e

10

I

"/"~/~-~ ~ 1,3-Diisopropy I - ~ X~J~--_j/ benzene benzene -~.v/ /,... /

2-El:hylnaphthalene .,,.

1-Methyl-

(~~~halene /

1-1RhylI-( ~ ~ i naphthalene 1, j "-4.~ " ~

0.1

LO

1,3,5Trimethylbenzene

et

0.01 1,3,5-Triethylbenzene 0.001 0.65

........1~~ (.

I

I

t

!

I

0.7

0.75

0.8

0.85

0.9

Critical Molecular Diameter [nm]

Figure 6. Diffusivity of molecules of various critical molecular diameters.

0.95

331

Table 3 Characteristics of an FCC standard catalyst incorporating unltra-wide pore (UWP) zeolite (MCM) [20] Standard Chemical Composition 35 Y-zeolite wt% d.b. UWP zeolite wt% d.b. Active matdx wt% d.b. 10 Physical analysis Surface area fresh m2.g1 237 Micropore volume ml.g 1 0.110 137 Surface area after Ni deactivation m2.g~ Micropore volume after Ni deactivation ml.g 1 0.053 MST activity at cat:oil = 3.5 63.3 Selectivity at 68 wt% conversion Gasoline wt% 43.7 LCO wt% 17.9 Bottoms wt% 14.1 Coke wt% 2.4 -

MCM 35 10

344 0.174 164 0.073 65 43.8 16.0 16.0 2.6

Another approach to selectively control the cracking process is to crack the heavier resid molecules on an active matrix and to crack the smaller fragments in the Y zeolite. The pre-cracking step can be carried out on a layer of active matrix coating, such as alumina [21]. There has however not been any report to date of a commercial application of such a concept. In order to improve the chances of success for a commercial application the pore size and acidity of the coating will have to be carefully adjusted to perform the pre-cracking. Currently some multi-porous sieves which may have FCC potential are: Boggsite, which is a natural mineral having a 12 membered ring on the outside; SSZ-26 / SSZ-33 / CIT-1, which are potentially very interesting but probably too expensive to prepare [61]; ZSM-50/MCM-22, the latter being an extension of the former which has 12MR pockets so that MCM-22 has interesting 10/12 MRs but only 10 MRs on the outside; 9

NU-87 (SSZ-37) which has 10 MR to the outside and 12 MR intersecting.

Although the concepts of a multi-porous FCC catalyst has much promise, the authors are not aware of any current commercial application. This is mainly due to the cost of these catalysts and the fact that this extra expense is not apparantely justified when the catalysts' performance is compared with the relatively less expensive option of a ZSM-5/Y mixture (cf.

2.5).

332 2.3

Metal contaminants

One of the roles of the matrix is to reduce the effect of metal contamination. Various matrices have been proposed to be suitable for this purpose. Magnesia-alumina has been found to be superior to most other inorganic oxides as a vanadium and nickel metal trap as well as increasing the resistance of the zeolite Y to steam deactivation [4]. Table 4 shows a comparison of the performance of a coated FCC zeolite and a standard FCC catalyst in a micro simulation test both after Ni impregnation and steaming and after cyclic deactivation with vanadium [19]. The coated catalyst shows a better steam stability after Ni impregnation and a much better stab'flit3, after cyclic deactivation with V in the feed. The higher conversion points to a better stability of the zeolite under deactivation conditions. The main advantage of the coated zeolite is the better protection against V attack. In another study it was shown that the zeolite retention, i_e. the zeolite activity after a particular treatment relative to the act'wity of the ~esh catalyst, doubled in the case of a catalyst with increased accessibility after the catalyst had been subjected to both V and Ni poisoning [22]. This is due to the fact that the functional sites which can neutralize these poisons are more accessible to the metal bearing molecules.

Table 4 Performance comparison after deactivation with Ni and V of an FCC catalyst and the same coated [19] NonCoated coated RE-Y REoY Ni impregnated 5h 788~ + 1000 ppm Ni Surface area m2.g-1 165 182 Micropore volume ml.g -1 0.048 0.054 Conversion wt% 77.5 77.2 Cat:oil = 3.5 wt/wt Selectivity 68 vwt% conversion Gasoline wt% 46.3 46.5 LCO wt% 17.9 18.1 Bottoms wt% 14.1 13.9 Coke wt% 3.1 3.0 Cyclic deactivation ~ 5000 ppmV Surface area m 2.g-~ 132 170 Micropore volume ml.g -1 0.035 0.051 Conversion wt% 71.8 75.4 Cat:oil = 3.5 Selectivity 68 wt% conversion Gasoline wt% 43.4 44.7 LCO wt% 17.3 17.6 Bottoms wt% 14.7 14.4 Coke wt% 5.9 5.2

333 2.4

Role of additives

2.4.1

ZSM-5 and Beta

Cracking catalysts using combinations of medium and large-pore zeolites in order to maximize the production of products to be used in reformulated gasoline has been reported [23]. In a study of the cracking of n-heptane over MCM-22, ZSM-5 and Beta it was shown that the yield of propene was greatest in the case of MCM-22 and the overall alkane/alkene ratio of the products lay between ZSM-5 (0.94) and Beta (1.17). The addition of ZSM-5 to the FCC catalyst is an important method to increase the amount of light alkenes without increasing coke or dry gas yield. There are now more than 50 commercial units worldwide using ZSM-5 as an additive [24]. The main reason for ZSM-5 being so widely used is that it is very easy for the refiner to add this catalyst when fight alkanes are needed and as soon as addition ceases alkene production stops shortly afterwards. Figure 7 shows that the addition of ZSM-5 is also able to increase both the KON and MON by increasing the iso/normal alkane and alkene ratios and the concentration of gasoline range aromatics [25,26]. At the same time it leads to an increased yield in propene and reduced gasoline yield and an increase in both iso-butene and n-butene (Figures 8 and 9) [25,27]. The addition of ZSM-5 also results in a decrease in the amounts of methyl-pentanes, hexanes and heptanes. The increase in the amount of C5s will result in an increase in the RVP values. Table 5 shows the effect of adding ZSM-5 and also of adding a mesoporous matrix to the KEY FCC catalyst. The effect of adding ZSM-5 is to enhance the cracking of C7 and higher alkenes. This will of course be accompanied by a slight increase in the amount of C5 and C6 alkenes which may crack further to increase light alkene yield [25].

95

82.5

94

, ,,

0.8

9

93.5

0.7 81

~' 92.5

, - ~' ~'&

,,..,

6.

0.65

81 ~

91 90.5

8

0.6

4. ,

~

~

15

80

ZSM-6 Additive ( w t % o n blend)

Figure 7. Effect of ZSM-5 additive on RON and MON [25,26].

~ O'

0.45 5~

10~

15 0.4

ZSM-6 Additive [ w t % o n blend]

Figure 8. Effect of ZSM-5 additive on yields of propane and propene.

334

55

3

==

45

o 0

5

10

15

!

0

5

ZSM-5 Additive [wt% on blend]

10

15

ZSM-5 Additive [wt% in blend]

Figure 9. Effect of ZSM-5 additive on yield of gasoline.

Figure 10. Effect of ZSM-5 additive on yield ofbutene.

Table 5 Different options for the production of light alkenes in FCC [25] Catalyst options Base cat Base cat + Low RE(*) + low RE 3 wt% ZSM-5 mesopore act Reactor temperature, ~ 525 525 525 C2= wt% 2.6 2.6 2.4 LPG wt% 16.4 18.8 15.9 Gasoline wt% 45.0 42.7 46.4

Base cat + high reaction 540 3.3 20.4 44.2

Deep cat cracking 600? 10.5 40.2 22.7

6.3 2.1 5.5 3.7

19.0 6.2 8.3 2.1

Desired Products: C3= wt% iC4= wt% nC4= wt% iC4 wt%

4.2 1.8 4.5 3.5

5.2 2.0 5.0 3.8

Undesired products: C~ wt%

1.4

1.7

1.2

1.8

3.3

1.0 19.0

1.1 21.3

0.6 18.4

1.0 23.6

1.1 50.7

nC4 wt% Total 'Vtet gas" wt% (*) ADZ-50containing catalyst

4.1 2.0 4.5 3.5

335 When ZSM-5 is incorporated as a cracldng catalyst the adsorption of alkyl aromatics, in which the benzene ring leads the alkyl chain, is favoured [28] leading to relatively high yields of benzene. With USY on the other hand the ratio of dealkylation to side chain cracking is an order of magnitude lower than with ZSM-5. It is estimated that with 20% ZSM-5 in the catalyst inventory and using a combination of high ten~erature and, since H-transfer reactions are intrinsically slower than catalytic cracking reactions, short cracking contact time (0.1 - 0.5 sec.) isobutene yield can be 5.4 voL % and combined propene and butene yield can be 35 vol. % [32]. ZSM-5 is particularly selective to making propene relative to C4s with 50 - 60% of the yield shifting to propene. The addition of P to ZSM-5 has greatly enhanced its activity such that only 5-10% ZSM-5 is in the additional catalysts compared to 25% in the original additive catalyst without P. It has been claimed that large pore zeolites such as Beta and ZSM-20 can be incorporated into a cracking catalyst with unusual selectivity for producing compounds boiling in the gasoline range which contribute to high octane. These components are low molecular weight alkenes produced by Beta and aromatics produced by ZSM-20 [33]. The Beta or ZSM-20 in the final catalyst are in the H + form It has also been claimed that when the Beta contains a small amount of gallium or zinc the aromatic content of the gasoline increases [34]. It is worth noting that Beta has a relatively low H transfer activity and this would be conducive to the production of more light alkenes but Beta also has lower gasoil cracking activity due to its problems with accessibility arising from its pore size and structural defects [35]. When Beta is used as an FCC additive in the same way as ZSM-5 much higher amounts of Beta zeolite than ZSM-5 need to be added to see sensible results and, since it is an expensive catalyst, it is not presently economically viable to use it extensively as an additive. 2.4.2

Rare Earth Elements

Cracking catalysts are thermally stable in air at 760~ even when loaded with 3-4% V. In the presence of steam however this stability is greatly reduced and dealumination occurs. Corma et al. [29], however, have shown, by examining the butane/butene ratio in the products obtained during the cracking of a vacuum gas oil at 480~ that, with steam dealuminated zeolites, a sharp decrease in the ratio of H transfer to cracking is observed when the number of A1 atoms per unit cell ~lls below 10. At these conditions the adsorption of alkenes decreases dramatically more that in the case of alkanes implying that the rate of H-transfer reactions will inevitably decrease relative to the rate of cracking. Stability to steaming can be increased however by incorporation of rare earths into the FCC catalyst. The RE impedes the dealumination of the zeolite structure and therefore increases acid site density. Concomitantly a reduction in the RE levels will result in a decrease in the unit cell size and in the acid site density [ 1,30,31] This is accompanied by an increase in the strength of individual sites which will favour the formation of LPG which requires strong sites. A reduced acid site density will also result in a reduced extent of hydrogen transfer which is accompanied by greater isomerization and a reduction in the amount of aromatics formed [30]. This has been explained by the fact that, while cracking is a unimolecular reaction needing one active site, hydrogen transfer, being a bimolecular reaction, needs two close active sites [ 1]. However it should be noted that the higher molecular weight precursors of light alkenes are preferentially saturated and the resulting gasoline range alkanes do not readily crack to light alkenes. A greater concentration in alkenes, which will increase MON, has a greater impact than the effect of reduced aromatics which would decrease the octane number. Generally the RE content in FCC catalysts has been decreasing in recent years in order to allow for an increase in alkene contents in the products by as much as 15% [3]. Low rareearth, high matrix activities also result in high isobutene yields which is also a favourable product if MTBE is being synthesized downstream Engelhard has achieved these effects on

336 their FCC Isoplus series of catalysts in which USY has a with unit cell size <24.29 A and extremely low Na (<0.1%) and rare earth levels. These are commercially proven catalysts which give excellent yields of isobutene and other light alkenes at much lower cost than Beta based catalysts. Currently about 6 units worlwide are using these catalysts in order to obtain high alkene yield [50]. In conclusion it is hnportant to appreciate in the case of catalytic cracldng that modifying reaction conditions can often result in a more si,~nificant increase in alkene production than can be achieved by catalyst modifications. This requires operating the FCC at high severity by increasing the cat/oil ratio which leads to an increase in the yields of C3s and C4s. Raising the reactor temperature, decreasing the feed temperature, and decreasing the regenerator temperature all lead to an increase in alkene production. Overcracking and secondary reactions can also be enhanced by increasing residence times although if the reaction time is too long there exists the possibility for more secondary H-transfer reactions. Obviously excessive overcracking (e.g. conversions > 75%) will lead to decreased gasoline yield and enhanced dry gas yield (Table 5). Lower partial pressures in the FCCU will also increase the formation of C3 and C4 alkenes.

3.

AI~YLATION

3.1

Introduction

In refining processes alkylation of isobutane with propene or butene is important in order to obtain alkylate which has a high octane number and a low vapour pressure. This process is not, however, directly relevant to the focus of attention of this paper and will therefore not be dealt with in any detail It has been well reviewed recently [36]. It is, however, worth noting that recent attempts to develop a zeolite as an alternate to the currently used hydrofluoric or sulphuric acid do not appear to have been successful and it is now assumed that superacid catalysts are the most likely heterogeneous alternatives. For the petrochemical industry the alkylation of aromatics is an important route to the production of alkylaromatic such as ethylbenzene, xylenes, cume, C10 "C18 alkylbenzenes, alkylphenol, alkyl-napthalenes and alkyl-biphenyls which are used in many different processes. Alkylation is an acid catalyzed reaction which traditionally employs aluminium chloride based catalysts or hydrofluoric or sulphuric acid These processes thus require the use of highly corrosiveresistant materials of construction, have a high catalyst consumption and are associated with environmental problems. Consequently there is considerable incentive to replace these catalysts with solid acids such as zeolites. The most important alkylated aromatic compound is ethylbenzene 99 % of which is used after dehydrogenation for styrene production. Several technologies using zeolites are nowadays available for the production of e t h y l b ~ e . The Mobil-Badger process [37] is a vapour phase process at 400 - 450~ 2 - 3 MPa using H-ZSM-5 as a catalyst. The high molar ratio of benzene to ethylene of 5 - 20 ensures an essentially complete ethylene conversion and maximal ethylbenzene selecthhty. The main by-product is diethylbenzene which can be converted into ethylbenzene by disproportionation with benzene. The catalyst life-time is several weeks and it can be regenerated using air. This process is technically proven using a diluted ethylene feed (17.6 voL-% ) using similar ethene:benzene ratios to the pure ethene feed [38,39]. Thus the mass flow through the reactor is slightly increased. It is necessary to eliminate propene from this feedstock because propene alkylates benzene readily and additional distillation facilities are then required. Although the catalyst is much less sensitive to sulphur, the latter was removed from the feedstock. It has been chimed that the liquid phase ethylation of benzene using MCM-22 results in a lower yield ofpoly-substituted ethylbenzenes

337 which would reduce the recycle of these compounds [40]. Processes utilizing catalytic / reaction distillation have been described in literature [41,42]. The formation of ethylbenzene by side-chain alkylation of toluene with methanol using a basic catalyst like alkali-cation exchanged X and Y zeolites has been shown to be feasible [43,44]. Partial oxidation of xylenes yields hnportant monomers for the polymer industry such as phthtalic acid from o-xylene, isophthalic acid from m-xylene and therephthalic acid from pxylene. Of these terephthalic acid is mostly used in industry as a co-monomer like in the production of PET. The demand for p-xylene is hence higher than for the other two xylene isomers. The desired product distribution can be achieved by using product shape selective catalysts like zeolites. Xylenes can be produced either by toluene disproportionation or toluene alkylation using methanol. Toluene disproportionation is an attractive route because it does not require an additional alkylating agent and both the product xylenes and the byproduct benzene are valuable chemicals. In conventional toluene disproportionation processes, the p-xylene content in the fraction of xylenes is ca. 24 %, which corresponds to an equilibrium distribution of the xylene isomers [45]. The Mobil Selective Toluene Disproportionation Process (MSTDP) using large H-ZSM-5 crystals [46,47,48] at 455 470~ 2 - 4 MPa, H 2 to toluene ratio of 3 moFmol, yields a p-xylene content in the fraction of xylenes of 82 - 90 % at 30 % toluene conversion. The use of large crystals is necessary, because aU three xylene isomers are formed inside the pores but due to the orders of magnitude greater diffusivity ofp-xylene relative to the two other isomers the desired product will leave the zeolite crystals and the other two isomers will be converted to p-xylene inside the zeolite's crystal [46, 49]. The high temperature is required because of the high activation energy of the disproportionation reaction and the presence of a noble metal in the catalyst and hydrogen in the feed increases the life-time of the catalyst which operates for up to one year before regeneration becomes necessary [1]. Alkylation of toluene with ethene using shape selective catalysts can yield pethyltoluene which upon dehydrogenation yields p-methylstyrene [51]. Using a modified ZSM-5 catalyst p-ethyltoluene can be produced very selectively (97 %) at high toluene conversion [46]. The polymer from this starting material may possess more interesting properties than polystyrene. Cumene formed by the alkylation of benzene with propene is the major source for the co-production of phenol and acetone. To a minor extent it is also used as the source for amethylstyrene. Originally it was produced using a solid phosphoric acid catalyst at 200 260~ 3 - 4 MPa. In this process 95 wt.-% eumene selectivity could be obtained and the main by-products were diisopropylbenzene and polyaromatic compounds. The reduction in the amount of by-products formed is the main incentive to replace this catalyst and liquid phosporic acid also causes corrosion problems in the downstream apparatus. Cumene can be produced using zeolite Y at ca. 200~ [52]. The once-through selectivity of this process is significantly lower (70-90%) but the by-products viz., polyisopropylbenzenes, can be transalkylated so that an overall process selectivity of 99 % can be obtained. Dow [53] have developed a cumene production unit on the basis of deahlminated Mordenite with aSi/A1 > 40 and preferably-160 [54], and which operates between 130 and 200~ At higher temperatures n-propylbenzene, which cannot be recycled, is formed and at lower temperatures the formation of diisopropylbenzene is favoured [55]. Interesting alternative catalysts for cumene production in the liquid phase are zeolites Beta and ZSM-12 which have shown better stability and higher selectivity in comparison with Y-zeolites [56,57]. Linear alkyl benzene sulfonates, which are important as detergents, are formed by the alkylation of benzene with linear C10 -Cls alkenes and are produced either using anhydrous HF, H2SO, or AICI3 as a catalyst [52]. The corrosive and hazardous nature of these catalysts has led to efforts to replace them with solid acids. Normally, in the alkylation of benzene with

338 linear ct-alkenes a mixture of alkylbenzenes is obtained with the phenylgroup attached to different carbon numbers except for the 1-phenyl isomer which would require a primary carbenium ion as a transition state. This indicates that the double bond isomerisation is much faster than the alkylation step. Ideally, 2-phenylalkenes should be formed because they can be easily converted in the ambient environment. Higher selectivities for this isomer can be obtained using shape-selective catalysts but their major drawback is their relatively rapid deactivation [58,59,60]. Recently, there has been much interest in the shape-selective alkylation of polyaromaties and biphenylic compounds, especially 4,4'-diisopropylbiphenyl ang 2,6dialkylnaphthalene, since upon oxidation they yield monomers for high quality plastics which have interesting applications such as in LCDs [47,62]. For these high technology polymers the narrowest isomer of the dialkylaromatics and dialkylbiphenyls seems to be the one most applied. Thus shape selective acid catalysts can prevent the formation of unwanted isomers and thereby improve the economics of the production of these compounds. Mordenite seems to be the most ideal catalyst for both the selective formation of the dialkylbiphenyl and the dialkylnaphthalene [63,64]. This is supported by molecular modelling studies (Figure 11) which have shown that, in the isopropylation of naphthalene, the 2,6-isomer is the favoured dialkylisomer due to the differences in diffusivity [65].

Figure 11. Model structure of 2,6-DIPN in HM and L.

339 3.2

Mechanism of alkylation

The mechanism of alkylation has recently been extensively reviewed by Venuto [40]. Briefly, alkylation involves an electrophilic addition of a carbenhtm ion which is generated by the alkylating agent. Different alkylating agents, such as alcohols, alkenes, alkyl halides and aromatics, can be used. If alkylaromatics themselves are the alkylating agent yielding the corresponding dialkylaromatic and the aromatic compound it is called disproportionation. Although the formation of the arenium ion from the relatively stable alkylaromatic is less favoured than the formation of the carbenium ions from other alkylating agents like alkenes and alcohols at high reaction temperatures both mechanism might occur. This was also concluded by Mirth and Lercher [67] who showed with IR that in the methylation of toluene at 200~ methanol replaces adsorbed toluene forming methoxonium ions and alkylation takes place by the interaction of this ion with toluene. At temperatures higher than 300~ however, the concentration of these ions becomes very small and therefore the other reaction pathway might prevail [68]. Alkylation reactions can be observed in all processes involving zeolites where aromatics and an alkylating agent are present. Even m-xylene isomerisation, which is classically visualized as a 1,2 methyl shitt [69], proceeds over faujasites partially by a number oftransalkylation steps. This has been shown by using a mixture of hexadeuterated m-xylene (C6H4(CD3)2) and normal m-xylene (C6H4(CH3)2) [70]. The relative importance of the 1,2 methyl shift versus the transalkylation differs over various zeolites [71]. During the xylene isomerisation at 200~ over HY at least 20 % of the reaction occurs via the bimolecular transalkylation. The intermediate complex in the bimolecular reaction needs to be accelerated in a micro-cavity and therefore the importance of the bimolecular reaction is three times less in the isomerization of xylene over H-Mordenite and does not occur over Beta at 2000C. Of interest is the reported absence of the bimolecular reaction over amorphous silica-alumina at 400~ Para/ortho-substitution is strongly favoured in the alkylation of alkylaromatics because of the electron-releasing effect. From a statistical point of view the ratio of ortho to para during alkylation of alkylaromatics should be 2, but, based on resonance considerations, the attack in the para-position is slightly favoured. Although in homogeneous catalyzed systems a high selectivity towards the para-isomer is expected, at slightly elevated temperatures the isomerization yielding the meta-isomer is fast and this decreases the production of the desired para-isomer. Shape-selective zeolites possess the ability to deliver high yields of the paraisomer. The isomerization forming the recta-isomer must, however, be suppressed. It has been shown in the case of toluene alkylation with methanol over H-ZSM-8 between 400 450~ that the shape selective formation of p-xylene is slower than the subsequent isomerization yielding m-xylene [72]. Of interest from a mechanistic point of view, is the observation that toluene disproportionation over Beta at 350 - 400~ yields a p-xylene concentration higher than the equilibrium value thus indicating the primary formation of this isomer in the large pores of this apparently non-shape selective catalyst [73]. The disproportionation of C9 aromatics, however, yielded an excess of o-xylene [73,74] whose concentration approaches the equilibrium value at higher temperatures. This was explained [74] with a biphenyl carbenktm ion intermediate which yields o-xylene as primary product. In the ethylation of toluene the ortho-isomer is the primary product formed over amorphous silica-alumina, whereas the shape-selective zeolite H-ZSM-5 yields the p-isomer with pethyltoluene being essentially absent [75]. Methylation of toluene over wide-pore zeolites like HY also exhibits a product distribution, which, although not thermodynamically limited, cannot be explained in terms of geometrical and diffusion effects [76]. Quantum mechanical calculations seem to indicate that the orbital interaction in the para-position is larger than in the ortho-position [77,78]. This, however, seems to contradict the observations made with amorphous silica-alumina catalysts [75].

340 Shape selective catalysis has been described earlier in this paper [79]. Due to the large difference in the diffusivites of para-, recta- and ortho-xylene in medium pore sized zeolites, the large observed selectivity of the para-isomer has therefore been explained in terms of product shape selectivity. It has been postulated that this product shape selectivity depends on the length of the intracrystalline diffusion pathway and on the tortuosity of the channel system [80, 81, 82]. Hence the selectivity should be governed by the type of the zeolite structure, type of modification and crystal size. If the whole reaction takes place on the internal, shape selective surface of the catalyst, then the observed decrease of the paraselectivity in the toluene methylation with increasing conversion and temperature [83, 84] and decrease in para-xylene selectivity with decreasing crystal size in the toluene disproportionation [85, 86] can be explained. Zeolite crystals possess both an internal, shape selective and an external, non-shape selective surface. Therefore, it was postulated that paraisomers are formed inside the pores and these are converted in a secondary reaction on the external surface of the zeolite [87,88,89,90]. In order to distinguish between the influence of the external and internal surface and to monitor the influence of the diffusion pathway on the selectivity it would be useful to perform the alkylation over zeolites with crystal sizes which differ by at least an order of magnitude and to inertize the external surface of these crystals. The formation of di- and tri-alkyl aromatics and biphenyls l~om their aromatic precursor is a consecutive reaction, in which first the mono-substituted aromatic compound is formed which subsequently reacts to form the di-alkyl compund. This was observed in the shape selective ethylation of biphenyl [91] isopropylation of naphthalene [92] over HMordenite and the isopropylation ofnapthalene over HY [63]. Therefore, if a dialkyl-isomer is the desired product, the reaction conditions (reaction time/residence time, partial pressures and temperature) have to be optimiTed to obtain the the maximum yield of this desired product. With shape selective catalysts the formation of poly-substituted aromatic compounds can be suppressed because they are not able to diffuse out ofthe channels of the zeolites [63].

3.3

Effect of pore size

The alkylation of aromatics over zeolites offers the poss~ility for shape-selective catalysis if the di~sivity of the products in the zeolite pores differ greatly or if certain reaction pathways are blocked due to the geometrical constraint which are put on the transition state of the reaction occuring in the zeolite pore or cavity. The major drawback of the use of zeolites is the possibility of pore blockage which leads to catalyst deactivation. For one-dimensional zeolites in particular, pore blockage will lead immediately to a strong deactivation. This was shown for the transalkylation of C 7 and C 9 alkylaromatics at 400~ in which mordenite showed a high initial conversion and a strong deactivation [73], whereas HZSM-5 was stable for the conversion of the toluene but not active for the conversion of C 9 aromatics due to reactant shape selectivity. Beta showed some deactivation but achieved a steady state conversion which for C 9 aromatics was higher than that obtained with H-ZSM-5. A slight deactivation of Beta was also observed during the isopropylation of toluene and cumene [93]. The zeolites have to be selected according to their pore and cavity dimensions to obtain the desired result. Bellussi et al. [57] studied the propylation of benzene in the liquid phase over H-ZSM-5, H-Beta (Si/AI = 14), HUSY (Si/A1 = 3) and the classical phosphorous impregnated kieselguhr catalyst at 150~ 3 MPa, benzene/propene ratio 7.4 and a space time of 0.06 rain. They observed after one hour on stream, that H-ZSM-5 was hardly active and HY was at that time less active than H-Beta.

341 The choice of the right pore dimensions is very important in the alkylation of polyaromatics and biphenyls. The activity of the medium sized pores of ZSM-5 for the methylation of naphthalene is low in comparison to H-Mordenite and HY. ZSM-5, however, showed high selectivity for 2-methylnaphtalene, whereas Mordenite and did not show shape selective methylation [88,94]. In 1-methylnaphthalene isomerisation at 300~ the selectivity of 2-methylnaphthalene correlates with the Spaciousness Index (SI) [95]. Good isomerization catalysts posses SI between 3 and 20 (HL, H-Beta, HM, EU-1 and ZSM-12). Catalysts with higher value for SI display a low activity and catalysts with a lower SI value showed activity for the disproportionation reaction. Over medium pore-size zeolites, i.e. H-ZSM-5 and HZSM-11, the methylation of 2-methylnaphthalene at 300 - 550~ produced the narrower dimethyl-isomers [96]. The selectivity towards 2,6-dimethylnaphthalene increased with decreasing temperature. Lee et aL [64] described the use of highly siliceous Mordenite for the alkylation of biphenyls to produce selectively 4,4'-diisopropylbiphenyl. Sugi and Toba [47] studied the liquid phase isopropylation of biphenyl over various zeolites at 240~ They observed, that the conversion increased in the order H-ZSM-5 H-Beta > HY> amorphous silica-ahtmina >KE-Y,~HF. Although the production of the 2-phenyl isomer is desirable in the production of LAB, it is much more important to avoid the formation of diphenyl-isomers and branched phenyl isomers. The production of these compounds can be suppressed by using HY as a catalyst [98,60]. The pore size of the zeolites can be modified by introducing cations in the pores or by the formation of coke inside the zeolite crystals, which decreases the diffus'lvity of the products, but also of the reactants, in the zeolite channels. Chen et al observed an increase in p-xylene selectivity with modified and coked H-ZSM-5 [85].

3.4

Effect of crystal size and external surface

As mentioned above, it is di~cu]t to separate the influence of the crystal size and of the external surface on the shape-selectivity of zeolites rigorously. Increasing the crystal size of the zeolite means at the same time reducing the influence of the external surface. For the industrial production it is desirable to operate with small crystal sizes.

342 Using small zeolite crystals (< 0.5~tm) in the methylation of toluene, Chen et al [85] observed at 500~ an equih'bfium mixture of the xylenes (i.e. 23% p-xylene). Increasing the crystal size to 31xm enhanced the para-selectivity to 46%. A further increase in the paraselectivity up to 97% could be obtained by modifying the catalyst with phosphoric acid ending with P-loading of 8.5%. The enhanced para-selectivity was explained by the increase in the diffusional pathway by pore plugging which would favour the outward diffusion of paraxylene. Also for the ethylation of toluene, modification with phosphorous or metal oxides of ZSM-5 was required to obtain high para-selectivities [75]. Kaeding et al [99] stated that the modification with P effectively blocks the external surface. Paparatto et aL [ 100] observed, in the ethylation of toluene and in the isomerisation of m-xylene at low contact times over H-ZSM-5, an excess pf p-ethyltoluene whereas amorphous silica-ahamina yielded an excess of the o-isomer. This indicates that the ethylation of toluene over H-ZSM-5 takes place inside the micro-pores. At higher contact times the product composition reaches the thermodynamic equih'brium distn'bution. With zeolite samples with large primary particles the o-isomer was always absent, whereas small crystals yielded the equilibrium distribution at high contact times. They explained their observations by the primary formation inside the pores of the p-isomer, which can isomerize on the external surface, The conm'bution of the external surface to the total active surface per gram of catalyst becomes larger if the size of the catalysts is reduced. The increase in crystal size also reduced the observed conversion. External acid sites can be eliminated by building an inert iso-structural silica shell around the zeolite by continuing the synthesis in an Al-free synthesis gel. This increases the crystal size and the effective diffusional pathway and is an effective method to enhance the para-selectivity in toluene alkylation but reduces the conversion over the catalyst [ 101]. The modification of zeolites by Chemical Vapour Deposition (CVD) does not only eliminate the external acid sites but also causes pore mouth narrowing. I-h'bino et al. [102] showed that the rate of adsorption of xylenes is decreased by CVD-treatment of H-ZSM-5 with tetramethoxysilane and they ascn"oed their enhanced para-selectivity in the methylation of toluene to pore mouth narrowing. Wang and Ay [103] showed that larger crystals needed less silica on their surface to obtain high para selectivity in toluene ethylation and therefore they regard the role of the active sites on the external surface to be very important. Matsuda et aL [104] studied the disproportionation of 2-methylnaphthalene over HZSM-5 and this zeolite post-treated with (NH4)2SiF 6 to eliminate the external acid sites. They observed that the bulkier isomers were formed over H-ZSM-5, whereas over the treated zeolite only the isomers 2,6 and 2,7-dimethylnaphthalene ws observed. This was explained by a disproportionation reaction in the pores and an isomerization reaction over the external surface. The same was observed in the isopropylation of biphenyl over Mordenite, where modification of the external surface with tn'butyl phosphonate increase the 4,4'diisopropylbiphenyl content in the fraction of dialkylbiphenyls and reduced the deactivation [105]. This was also ascn"oed to the activity of the external surface. Another method to eliminate the external acid sites is the selective poisoning technique in which a stronger base is added to the feed which is too large to enter the pores of the zeolite. It was observed that injection of small amounts of [3-naphthoquinoline during the toluene ethylation over H-ZSM-5 (crystal size 21am) increased the selectivity towards p-ethyltoluene but at the same time decreased the ethene conversion [106]. Regular injections of the base molecule are necessary because of the reversa'ble adsorption and the decomposition of the base under reaction conditions.

343 3.5

Effect of silica to alumina ratio and dealumination

The effect of silica to alumina ratios in the alkylation of aromatics is difficult to study separately because upon changing the Si/A1 ratio in the synthesis gel both the ratio in the zeolite and the crystal size are changed [97], as well as the crystallinity and morphology. If the Si/Al ratio is changed in a post-treatment step by dealumination, which can be done either by acid washing or by steaming, extra framework aluminium species are formed. These species can block pores and thereby modify the diffush~es of the reactants and products in the pores and can form a complex with remaining framework aluminium which may result in a modified acidity of the catalyst. Vinek and Lercher [107] synthesized ZSM-5 with Si/A1 ratios between 20 and 240, but because the pyridine TPD yielded a lower Si/Al ratio the existence of extra-framework aluminium species which were ascrl'bed to be weak acid sites. They obtained a linear correlation between the specific rate of toluene disproportionation and xylene isomerization and the number of strong Bronsted sites, which indicates that the reaction rate is primarily a function of the concentration of acid sites. This was also concluded by Nayak and Riekert [89] and observed for the ethylbenzene disproporfionation [108,109,110]. For toluene dispropordonation a linear relation~ip between the rate constant, assuming the rate to be first order with respect to toluene, and the Si/A1 ratio was obtained [ 111]. This indicates, that the turn-over-number (TON) remains constant and independent of the silica to alumina ratio. Sastre et al. [112] studied the isomerization of m-xylene over Ot~etite and observed monotonical increase in m-xylene conversion upon exchange of the K+-cations. This was ascn'bed to the increase of the concentration of the protons and the increase in accessa'bility of the pores, which resulted in a higher selectivity for the isomerisation reaction at the expense of the disproportionation reaction. Only a slight increase in the p-xylene in the fraction of orthoand para-xylene was observed. Over Beta a maximum activity for the xylene isomersation was observed and this was explained by either a possible existence of a synergistic effect between extra-framework aluminium and the framework Bronsted acid sites or a concentration effect [113]. The alkylation of toluene with methanol is also catalyzed by both strong and weak acid sites [107]. The ideal alkylation catalyst should have a high concentration of weak acid sites and a low concentration of strong Bronsted sites in order to minimize the side-reactions, viz. disproportionation. On the other hand it was observed that for the alkylation of benzene with linear alkenes over zeolite Y the rate increased linearly with the number of t~amework aluminium atoms which means that the turnover number remains constant [98]. However, the turnover number increased with increasing degree of ion-exchange which causes an increases in acid strength. The selectivity towards the desired 2-phenyl-alkane increases with increasing degree of ion-exchange showing that alkylation is a demanding reaction. For the alkylation of polyaromatics and biphenyls Mordenite with high silica to alumina ratios seems to be the preferred catalyst. Lee at al. [64] observed that dealumination of Mordenite by acid washing with 6 N HNO 3 modified the pore structure of Mordenite resulting in an increase in the total pore volume and especially an increase in the volume of pores with a diameter between 20 and 1000 A. In the isopropylation ofbiphenyl an increase in the yield of diisopropylbiphenyl was obtained which might be ascribed to either the enhanced diffusion of the reactants and products via the newly created meso-pores or the decrease in the rate of deactivation during the alkylation reaction.

344 The effect of dealumination of Mordenite by acid washing, leaching with EDTA and steaming has been studied systematically [114]. The selectivity to 2,6-diisopropylnaphthalene in the alkylation ofnapthalene was enhanced by the removal of external sites by leaching with EDTA. On the other hand after deep bed calcination the catalyst with a high external acidity showed a high conversion and a high selectivity. Stezming followed by mild acid washing to remove the extra-framework a~minhnn showed the lowest external activity and the highest selectivity for the formation of 2,6-diisopropylnaphthalene. Coke formation during the alkylation of biphenyl over Mordenite is reduced by using Mordenite with a high silica to aluminium ratio [ 110, 61], but also the nature of the coke is different. Mordenite with a high Si/A1 ratio produces a volatile coke (Td==,vtion= 200 - 340~ which are mainly biphenyl derivates whereas mordenite with a low Si/AI ratio yields hard coke which is burnt off at ca. 500~ The content of 4,4'-diisopropylbiphenyl in the fraction of encapsulated diisopropyldiphenyl isomers in the highly siliceous mordenite is over 80% which indicates the effectiveness of the pore system of Mordenite to produce selectively the desired isomer 4,4'- diisopropylbiphenyl.

4.

AROMATIZATION OF ALKANES/ALKENES

4.1

Introduction

Besides being a key high octane component of gasoline light aromatics are important raw materials for the production of a wide variety of petrochemicals. Benzene ranks third in volume and together with ethylene and propylene accounts for about 75% of the world's petrochemical production. At present catalytic reforming of hydrofined naphtha is the main source of BTX (benzene, toluene and xylene). The standard Pt/Re/A1203/CI catalyst is not very effective for converting C 6 alkanes to benzene, the yield being typically only about 10% as against 60% for methycyclopentane (MCP) and 90% for cyclohexane [58]. When the phasing out of octane-boosting lead from gasoline was started there was considerable interest in the production of additional BTX. To this end several zeolite based processes were developed, e.g., BP/UOP's Cyclar process using refinery C3/C 4 gases as feed, Chevron's Aromax process using C6 to C 8 alkanes as feed. The benzene could also be sold into the growing petrochemical market. However, since the allowable benzene content of gasoline is being lowered to below 1% the interest in these processes waned. To lower the benzene content in reformate it can be alkylated to toluene and xylenes, or it can be extracted and sold into the petrochemical market. The latter option could further depress the need in the short term for new sources of benzene. Only in instances where there is a shortage of aromatics but an ample supply of C 3 to C 8 alkanes and alkenes (as in a Fischer Tropsch complex) may processes such as Cyclar or Aromax be of interest. Nevertheless there is continuing research interest in aromatization using zeolite based catalysts such as Cra/HZSM-5 and Pt/KL.

4.2

Acidic Catalysts

The catalyst of choice remains acidic Ga-HZSM-5. The BP/UOP Cyclar process [115] used this catalyst in the 1000 bpd plant at Grangemouth, Scotland, which operated for about two years and was shut down in December 1991. With butane as feed a typical product spectrum was 65% BTX, 5% hydrogen and 30% fuel gas. UOP's continuous catalyst

345 regeneration process was used. IFffs Aroforming process also uses Ga-HZSM-5 in isothermal tubular reactors which operates on dual cycles [132]. Mitsubishi's Z-Forming process was tested in a 200 bpd unit which was commissioned in September 1991. The success of the HZSM-5 catalyst is no doubt linked to the low coke forming tendency of this particular zeolite. Other acidic zeolites such as HY are initially active but deactivate very rapidly due to coke deposition. Bradley and Kydd [116] investigated the performance of several pillar interlayered clay minerals and although the Ga pillared montmoriUonite was found to be the most effective it had a much lower activity than Cra/HZSM- 5. As is well know gallium addition greatly improves the performance of HZSM-5, eg, HZSM-5 at 550~ has a BTX selectivity of only about 12% while the addition of 5% Ga increases the BTX selectivity up to 70% [117]. The conversion of alkanes or alcohols to aromatics over HZSM-5 involves firstly the formation of alkenes which are then subsequently converted to aromatics, strong acid sites being involved in both steps [118,119]. Alkenes react much faster than alkanes [ 120] and this is in keeping with the deduction that the initial alkane dehydrogenation is a slow step in the overall process. The addition of Ga provides additional routes for dehydrogenation of alkanes, alkenes and naphthenes thus increasing both the overall reaction rate and also the selectivity to aromatics. Dehydrogenation via acid sites involves hydrogen transfer with the formation of low molecular mass alkane such as methane and ethane [120,121] which being inactive represent a loss of feedstock carbon. Dehydrogenation via Ga, however, produces hydrogen gas (which is a valuable by-product in refineries) and so results in a better feedstock carbon utilization. Addition of zinc to HZSM-5 has also been found to be very effective but Ga is preferred because of its higher stability [122]. ZnO is slowly lost through volafflisation at the high operating temperatures. More recently other metals active in dehydrogenation have been investigated as co-catalysts with HZSM-5. Ibm et al [123] claim that when feeding n-pentane to a Ni HZSM-5 catalyst the aromatic yield was equivalent to that obtained with Ga or Zn. They report an aromatic selectivity of 64% with Ni as against 71% for Zn and 66% for Ga. It should be noted, however, that when feeding propane the aromatic selectivity was only 25% which is a poor result. Ono et al [124] found that their Ag-HZSM-5 catalyst produced less methane and ethane than G-a- or Zn-HZSM-5 and concluded that Ag enhances C-H bond cleavage whereas Ga or Zn enhances both C-H and C-C cleavage. With butane and isobutane at 500~ the Ag catalyst gave a higher aromatic selectivity, namely 50 to 60% as against 30% for Ga. It should be noted, however, that here again the reported Ga results appear to be poor. With butene and methanol as feeds aromatic selectivities of 85 and 73% were obtained respectively with the Ag-ZSM5 catalyst. Shpiro et al [125] investigated the effect of adding both Pt and Ga to HZSM-5. They report that Pt promoted Ga reduction and its migration, resttlting in a more stable catalyst with a higher aromatic selectivity.

As gallium plays a key role its effective distn'aution in the zeolite is important. Although it is generally assumed that Ga 3+ is the active form, migration occurs more readily in the reduced Ga "~ state. The addition of Pt promotes Ga reduction by hydrogen spill over [125]. Hamid et al [126] found that the Ga, prepared by ion exchange, was, as one would expect, concentrated on the outer skin of the zeolite particles but with reduction/oxidation cycles the Ga migrated into the interior. They speculated that Ga + migrated as Ga20 vapour. In the regeneration cycle the Ga + is oxidised to the more active Ga 3+ state resulting in an improved performance. Further studies showed that after several reduction/oxidation cycles the performance reached a plateau [ 127]. It was deduced from pyridine infra red studies that the reduction/oxidation cycles resulted in a decreased H + concentration, due to exchange by Ga ions, and an increase in the Lewis acidity due to better Ga dispersion. In another study [117] it was found that H 2 pre-reduction markedly increased the aromatic selectivity of

346 physically mixed Ga20 3 / HZSM-5 but it decreased the aromatic selectivity of samples prepared by incipient wetness impregnation or by ion exchange. It appears therefore that H 2 reductions only improves matters when the Ga is poorly distn'buted in the initial state of the catalyst. If Cra dism~aution is important it could be reasoned that HGa silicate (MFI) would be a good catalysts since the Ga here is atomically dispersed by being incorporated in the ~amework. It has in fact again been reported recently that this zeolite is more effective than Ga/HZSM-5 [128]. Choudhary et al [129] found that the aromatic selectivity of Hgallosilicate increased with the degree of I-I* exchange while it decreased with mcreasing calcination temperature or increasing steam content during calcination. The latter two effects would be due to sintering (ie lower dispersion) of the extra-~amework Ga. Lukvanov, Gnep and Guinet have modelled the kinetics of propene [119] and of propane [120] aromatization over both HZSM-5 and Ga-HZSM-5 obtaining good fits with the experimental results. Propane is converted to propene along two main routes, protolytic craclcing of C-H bonds and dehydrogenation at the Ga sites. Protolytic crack~g of C-C bonds produce methane and ethane. The acid site reactions result in a CH4/H 2 ratio of 2.6 while the Ga sites give a 0.26 ratio which is in line with the observation that Ga HZSM-5 produces more H 2 than H-ZSM5. The propene then oligomerizes to higher alkenes (acid reaction). The oligomers are converted to dienes with both acid and Ga sites contn"outing and the dienes are converted to cyclic alkenes (acid sites). Cyclic alkenes are then converted to cyclic dialkenes and then to aromatics (H transfer at acid sites and de-hydrogenation at Ga sites). It was estimated that the Ga sites contn'bute about 90% to the diene formation and about 50% to the formation of aromatics. The formation of aromatics via H-transfer should result in the production of alkanes but the majority of the latter are again converted to alkenes. The only stable alkanes to emerge are the nonaromatizable methane and ethane. The product spectrum when feeding hexene or octene is very similar to that when feeding propene which is expected i~ as was found, the primary reaction is the craclcing of these higher alkenes to propene and butenes [117]. In general the percentage conversion of the feed and the aromatic selectivity follow the same trend. Likewise C2H4 selectivity follows the BTX selectivity [117,129]. When considering the breakdown of the aromatics it appears that at low temperature (350 ~ xylenes are the dominant aromatics, the benzene being low. As the temperature is raised to 550~ the benzene increases, toluene remains fairly constant and the xylenes decrease [117].

4.3

Platinum on Neutral Zeolites

Although platinum alone or on a variety of neutral supports selectively converts nhexane to benzene most of these catalysts deactivate rapidly due to coke formation. With the neutral zeolite KL as a support, however, much longer on-stream times are feast~ole and within a few years of Bernard's original publication [130] the Aromax process had been developed by Chevron [131]. Table 6 compares the aromatic selectivity obtained with Pt-Ba KL and Pt Re Sn / Al203 - C1 reforming catalysts [58]. Associated with the much higher aromatic selectivity is a lower amount of light gas production. Since the neutrality of the support was an important aspect the influence of doping Pt KL with the alkali series Li to Cs has been investigated by various workers. Hicks and coworkers [133] exchanged BaKL individually with Li to Cs and then added Pt by incipient wetness impregnation. They reported that the activity for aromatic formation increased markedly ~om Li to Cs but that the selectivity only increased slightly. Earlier studies [134]

347 had reported that both the conversion and selectivity increased markedly as Pt/KL was promoted with Li to Cs. Clearly the more basic the catalyst the better the performance. From this point of view it is interesting that promotion with various halogen compounds enhanced performance [135,136] despite the electronegativity of the halogens themselves. Tatsumi et al [ 137] investigated the effects of added KF, KC1, KBr and KI on the performance of Pt/KL. They found that KF and KC1 gave the highest benzene selectivities but that KBr and KI were actually inferior to the unpromoted Pt/KL. Table 6 Alkane Aromatization over Pt on Neutral and on Acidic Supports [58] % Aromatic Selectivity Alkane Feed Pt-BaKL Pt Re Sn/Ai203CI C6 87 25 07 82 45 Cs 80 60

The high selectivity of Pt/KL has been ascribed to the presence of very small Pt particles [138,139] and thus that sintering of these particles is one of the causes of deactivation [140]. It has been shown previously that Pt on zeolites KL, HL, HZSM-5 and silicalite were dispersed by treating with C12 in nitrogen or HC1 in air at 350~ [141]. With standard Pt / A120 3 reforming catalysts the practice of redistn'bution of the Pt (after air regeneration) by treatment with chlorine is well known. In the light of the foregoing it appears probable that the positive effect of halogen pro-treatment [135,136,137] is largely due to its resulting in finely dispersed Pt clusters. Iglesia and Baumgartner [ 142] have pointed out that selective terminal adsorption and dehydrocyclization of hexane to benzene are intrinsic properties of any clean Pt particles and that the role of KL zeolite is that the size of the channels inh~it the formation of coke in these channels thus keeping the Pt clusters there clean. This is in keeping with the opinion expressed previously by Tamm et al [131]. If it were a matter of pore size then one could, however, expect neutral silicalite also to be a satisl~ctory support. It is well known that operating metal catalysts in a hydrogen atmosphere inh~its coke fouling. Hicks et al [143] found that with 0.6% Pt/KBaL the catalyst deactivated due to coke fouling at hydrogen partial pressures below 6 atmospheres. The conversion of heptane increased with increasing hydrogen pressure up to 6 atmospheres. Pt/AI20 3 reforming catalysts commonly also contain Re which improves the catalysts' resistance to coke deposition [131] and this raises the question whether the effect of adding Ke to Pt/KL has been investigated. Pt/KL has been shown to be very sensitive to sulphur poisoning [131,144] and the effect has been ascribed to sulphur accelerating Pt sintering and subsequent blocldng of the zeolite channels rather than normal surface poisoning [144]. The sensitivity to sulphur obviously requires very thorough desulphurization of this feed and from this aspect a feedstock derived from the normal Co or Fe based Fischer Tropseh catalytic process could present an advantage. Fischer Tropsch products, however, contain other non-paraffmic substances such as alkenes, alcohols and carbonyls. The effect of these and other contaminants on the aromatization of hexane over Pt/KL is currently being investigated in the authors' research group. Zeolite supports other than KL have also been investigated. Pt/K Beta because of its higher acidity yielded more isomerized and cracked products than Pt/KL [145]. Ion exchanging with Cs reduced its acidity and improved the aromatic selectivity and Ba improved

348 the dispersion of the Pt which also increased its aromatic selectivity but despite these improvements the Pt Beta catalyst was still inferior to that of Pt/KL. It was, however, less sensitive to sulphur than Pt KL. Ruckenstein et al [146] studied the performance of composite catalysts, consisting of Pt/Ba-K1 with either Pt/beta or Pt/USY. Feeding mixtures of n-hexane, methylcyclopentane and methylcyclohexane the composite catalysts gave higher C7+ aromatics than expected from theindividual catalysts and feeds. An interesting observation was that for all the various individual catalysts (including Pt/Ba-KL) and for the composites n-hexane gave a lower benzene sdecdvity than did methylcyclopentane. This is contrary to the results of others [131]. In normal Pt/AI203-CI naphtha reforming the reaction network is complex because both metal and acidic sites to varying degrees catalyze ring closure, isomerization, dehydrogenation and cracking. Pt apparently is mainly responsible for hydrogenation / dehydrogenation with the acid sites accounting mainly for isomerization [147]. The conversion of n-hexane to benzene apparently goes via methylcyclopentane (MCP) and this is supported by the observation that at low conversions the major product when feeding nhexane is MCP [ 147]. With neutral Pt/KL the reaction network is simpler because the Pt sites mainly account for all the products. A commonly assumed reaction sequence is depicted in Figure 12. 1-6 Ring closure of chemisorbed n-hexane yields cyclohexane while 1-5 closure yields methylcyclopemane. Ring opening of the latter accounts for the two isoalkanes (2MP and 3MP).

Benzene

m

Cyclohexane Hexenes

~~

2MP

T IT MGP

n-Hexane

, P e n t a n e + OH 4

l

Butane

+ OH4

3MP

Figure 12. Reaction sequence for n-hexane conversion over Pt/KL.

It is of interest to compare the observed product concentrations with those predicted by thermodynamics. Table 7 lists several relevant equilibrium ratios. From the values ofthe cyclohexane / n-hexane and the benzene / cyclohexane ratios one would expect that the amoum of cyclohexane emerging from the reactor would be low which indeed it is. At about 40% n-hexane conversion at 450~ a typical exit molar ratio of benzene to cyclohexane is about 300: I, which although much lower than predicted is in line with the known fact that cyclohexane is very rapidly dehydrogenated over Pt at high temperatures. The exit benzene:MCP ratio is about 7 (at 40% n-hexane conversion at 450~ which is also lower than the predicted value. This nevertheless indicates that I-5 ring closure occurs at a reasonable rate compared to I-6 ring closure (the latter being followed by rapid benzene formation). At 350~ and at conversions below 3% the MPC concentration in fact exceeds that of benzene by a factor of 3, which again shows that I-5 ring closure is fairly rapid.

349 Table 7 Equilibrium Ratios of varius mixtures at 1 atm H2 Ratio 600K 700K 800K Cylohexaneln-Hexane 0.017 0.062 0.17 Benzene/Cylcohexane 34 1.8X10 4 2x106 MCPIn-Hexane 0.10 0.56 2.1 Benzene/MCP 5.9 2x103 1.6x105 3MPI2MP 0.51 0.54 0.56 MCP = Methylcyciopentane; MP = methylpentane

The observed 3MP/2MP ratio is about 0.8 which is not very different from the predicted value of about 0.5. The latter ratio does not change much with increasing hexane conversion when, as expected, the benzene level increases. This shows that the near equilibrium state between MCP, 2MP and 3MP persists at different conversion levels indicating that ring opening and closing occurs fairly rapidly. The main cracked products are methane and pentanes which is in line with the known hydrogenolysis activity of Pt. The main alkenes in the product mixtures are trans and cis 2-hexene in that order. While the reaction pathway depicted in figure z is supported by the fact that feeding n-hexane, MCP, 3MP or 2MP individually over Pt KL all result in high and similar aromatic selectivities [131], the actual mechanL~m on a molecular level is still a matter of dispute [139,142].

5.

SKEI~ETAL ISOMERIZATION OF 1-BUTENE

Isobutene is an important petrochemical starting material and best known for its use in the production of MTBE which is added to fuel as an octane-enhancer. It is also used as a monomer for the production of butyl rubber. Furthermore isobutene can be converted into isoprene which is an important monomer for elastomers by the modified Prins reaction with formaldehyde over, for example, H-ZSM-5 at 175 - 4000C [148]. Partial oxidation of isobutene yields methacrolein/methacrylic acid which upon esterification yields alkylacrylates, which are used e.g. for the production of polymers (plexiglass) and in water-soluble paint. Presently, the need for isobutene is covered by its production in the FCC-unit. However, with a strongly increasing demand for this raw material, especially for the m~nufacture of MTBE, alternative routes for the formation of isobutene need to be explored such as the acid catalyzed skeletal isomerization of linear butenes. Thermodynamically the skeletal isomerization of alkenes is favoured at low temperatures and the reciprocal temperature increases with increasing carbon number. The equih~rium concentration of isobutene in the fraction of butenes decreases from ca. 50 % at 200~ to 37% at 500~ [149]. Thus, the conversion of n-butenes into isobutene at these temperatures will be limited by thermodynamic constraints. The skeletal isomerization of the alkenes with more than 4 carbon atoms is a relatively facile reaction step, which is carried out at ca. 290~ over H-Ferrierite [150] or at 340~ over ZSM-5 [151]. This reaction proceeds via the skeletal rearrangement of a carbenium ion yielding a secondary carbenium ion. The singular reaction mechanism indicates that side product formation can be minimized. Even the skeletal isomerization of C 5- and C6-alkanes over Pt-Mordenite, which is thought to proceed

350 via a dehydrogenation step is a relatively facile process [ 151]. This is nowadays an important process for increasing octane numbers [ 153]. Contrary to the isomerization of longer chain alkenes, the formation of iso-butene from n-butene over acid catalysts is a difficult reaction, which proceeds e ~ e r via a mechanism involving oligomerisation, skeletal isomerisation of the oligomers and subsequent cracking or via an energetically unfavourable primary carbenium ion mechanism [ 154]. The first proposed mechanism imnlies the unavoidable formation of C5+-oligomers and C3.-cracked species as by-products m this process. Fluorinated alumina seems to be a promising catalyst for the skeletal isomerization of linear butenes [155,156], but the need to add fluorine to the feed stream together with the associated corrosion and environmental problems might prevent its industrial application [ 157]. A promising alternative to fluorinated alumina are zeolites. A number of zeolites have been studied for their activity and selectivity for n-butene isomerization [156-161]. The conversion of n-butenes over H-ZSM-5 can be observed at 377~ [157], but the iso-butene selectivity for this catalyst is rather low (14 %) and especially the selectivity for C 1-C3 products is quite high. At 500~ high conversions are obtained but the yield is then limited due to thermodynamic constraints [158,160]. The high activity and low selectivity of ZSM-5 has been ascribed to its strong acidity [158]. The acidity of zeolites can be reduced by the incorporation of boron in the zeolite framework [16_2,163] and therefore B-substituted ZSM-5, ZSM-11 and Beta were tested [158,164]. A13§ free boron zeolites are inactive, but these zeolites with low levels of A13+ ions which can be obtained by adding A120 3 binder to the A13+ free boron zeolite have weak acidity and are moderately active at 500 - 600~ and isobutene selectivities of up to 50 % have been reported. At these conditions the observed activity and selectivity of B/A1-ZSM-5, B/A1-ZSM-11 and B/A1-Beta were similar and therefore it was concluded that the pore structure did not play a decisive role in the conversion of n-butene into isobutene [164]. However, A1 which migrates into the pores not only modifies the acidity but also modifies the effective pore diameter. The importance of pore size has been frequently emphasized [150-161]. Figure 13 shows the performance of H-Ferrierite, H-Mordenite and SAPO-11 on the skeletal isomerization of n-butene. The widepore zeolite, mordenite, shows a low conversion and a low selectivity towards isobutene. The selectivity to isobutene obtained with the medium pore size SAPO-11 is higher but, due to the low conversion, a lower isobutene yield is obtained. With H-Ferrierite, a higher selectivity (> 80%) is obtained yielding a composition which is close to thermod~(namic equih'bfium The ~lectivity of zeolite Theta-1, which has narrower pores (5.5 x 4.4 A) than ZSM-5 (5.6 x 5.3 A and 5.5 x 5.1 A) for isobutene at 377 - 379~ was reported to be three times higher but also the conversion over Theta-1 at these conditions was significantly lower. Ferrierite (4.2 x 5.4 A and 3.5 x 4.8 A) showed at these temperatures over 90 % selectivity to isobutene [159]. A comparison of the activity and selectivity of SAPO-5 (ca. 8 A), SAPO-11 (6.7 x 4 A) and SAPO-34 (4.3 A) at 400~ showed that the selectivity to isobutene increased with decreasing pore size [159]. It was further noticed that the conversion and catalyst deactivation decreased with decreasing pore size. The large pore zeolite Mordenite itself is not selective for butene isomerization but Mg-Mordenite has been found to be more selective than H-Mordenite [150]. Zeolites with pores smaller than 4.2 A like Erionite are not useful because of the diffusional constraint of the product isobutene [150, 160]. Studies have shown that Ferderite is an attractive catalyst for n-butene isomerization, because it is both active and selective at relatively low temperatures of 350 - 400~ [150]. Initially small differences in butene conversion and isobutene selectivity for Ferriefite with

351 different Si/AI ratios (Si/A1 = 9 - 43) were observed but after some time on stream the differences were negligible indicating an influence of the acidity on the initial performance of the catalyst but not on its steady-state performance. The selectivity of this zeolite has been explained in terms of shape selectivity and substantiated with computer simulations [160] because the pore structure of the Ferrierite strongly inh~its the diffusion of the intermediate trimethylpentenes and therefore increases the probability of cracking which yields iso-butene. Most laboratory studies has been performed with a highly diluted feed at atmospheric pressure [158,157,159,163,164] because at a butene partial pressure of 1 bar much lower isobutene selectivities [161] and shorter catalyst life-times for medium and large pore zeolites were reported [150]. This is consistent with the postulated tmi-molecular isomerization reaction and multi-molecular oligomerization reaction [158] and the observed first order with respect to the partial pressure of n-butene for the formation of iso-butene and an order larger than one for the formation of by-products assuming low coverage for boron substituted zeolites [164]. For industrial operation, however, it is desirable to operate at higher partial pressure of n-butenes. Ferrierite can operate at higher butene partial pressures with a moderate catalyst life time [ 150,160]. With time on stream the activity of the zeolites for butene conversion decreases and the selectivity for isobutene increases [150,157,159,160,165]. Ferrierite initially produces cracking and oligome"nsation products [165] but with time on stream the rate of formation of these by-products decreases whereas the rate of formation of isobutene first increases before showing a slow decrease. The still active and selective catalyst can contain up to 8 - 10 wt.-% coke which is aromatic in character (H/C = 1) and which reduces the accessible pore volume dramatically [160, 165]. In the case of Theta-1 it was observed by using lower calcination temperatures (325~ instead of 500~ that the isobutene yield increased si~ificantly from 4.6 mol-% to 25.5 tool-% [157]. This has been ascribed to residual template inside the pores [161]. Figure 14 shows the effect of temperature and space velocity on n-butene isomerization over H-Ferrierite. It has been observed, that the isobutene selectivity increases with increasing space velocity and with increasing temperature [150,157]. The usual explanation for an observed increase in selectivity with increasing space velocity is the reduction of secondary conversion of the compound and if the space velocity is high enough its primary formation. Primary formation of isobutene is consistent with the observed different reaction orders for the isomerization and by-product formation [164]. The enhanced selectivity for the coked catalysts can then be explained with a lower diffusivit~ of the reactants into the zeolite and thus a lower butene concentration in the pores, which would favour the reaction with the lower reaction order, i.e. the skeletal isomerization. Assuming the primary formation of isobutene the observed increase in selectivity with temperature would indicate a higher activation energy for its formation in comparison to the oligomerisation. This is mechanistically consistent, because the oligomerisation will proceed via secondary carbenium ions whereas the direct skeletal isomerization of butenes involves an energetically unfavoured primary carbenium ion which might be stabilized by the zeolite structure.

352

40 ) 35 3O

~ F e r r i e r i t e

25 -o 20 SAPO-11

15

"---t

10

H-Mordenite 0

I

I

t

6

12

18

24

Tim 9 on Stream [m in]

Figure 13. Yield ofiso-butene from the skeletal isomerization ofn-butene.

100 90

9 9 X

80

E

70

~

100

400"C, 14 hr-1 425"C. 14 hr-1 400"C, l B h r - 1 425"c, 7 hr-1 Export. (400~, 14hr- 1) qlxxt. (425"C, 14hr-1)

90

7O

~~ _-; so

F ~ y. ( .;2-3~,,. 7"~ - i"/

60

0

>=

8

50

o

40

~

4 o l~ /

1~400"C' 18 hr'l

30

m

30 ~-

IX425~:.7",-~

20

20

10

10

O/ 24

48

7"2

Time on Steam [hr]

96

120

0

F''~'~""r-'

)

I

i

t

24

48

72

96

Time on Steam [hr]

Figure 14. Effects of temperature and space velocity pm n-butene isomerization over Ferrierite.

120

353 6.

ALKENE OLIGOMERIZATION

The oligomerization of light alkenes into dimers, trimers, tetramers and higher oligomers represents an important reaction for the production of aromatic-free higher alkenes. Although the use of ZSM-5, with Si/Al ratios of approximately 30 - 40, as an oligomefization catalyst was patented in the early 1980s and has been extensively reviewed [166,167] currently only the Mossgas Refinery in South Africa is using this technology [168]. The process is able to produce, after hydrogenation, mainly low branched alkanes and scarcely any aromatics. The high selectivity to alkenes in the diesel mode of operation is attn~outed to restricted transition shape selectivity which both favours alkene formation and inh~its the formation of typical cyclic coke precursors. These factors together with a high reaction pressure (typically 5MPa) and moderate reaction temperature (200 - 220~ are conducive to the formation of diesel fractions. The typical feed composition is 81.7% alkenes, 15% alkanes, 1.5% aromatics and 1.8% oxygenates and the typical liquid fuel yields, based on alkenes, of 97% and, when operated in distillate mode, yields 78% distillate and 19% gasoline. The product diesel has a high cetane number (about 53). The gasoline has an RON of between 81 and 85 and a MON between 74 and 75. The oxygenates may cause premature catalyst deactivation possa'bly due to stronger and irreversa'ble adsorption on the acid sites [169]. Apart from ZSM-5, there have been reports of the oligomerization activity of other zeolites. As expected the extent of chain branching, usually undesirable in most applications, increases as the pore size increases due to the shape selective nature of the reaction in the zeolite. Oligomerization of higher alkenes represents an important route to the formation of synthetic lube oils [ 180]. After hydrogenation such oils have excellent properties due to their low volat'flity for their viscosity, high thermal and oxidation stability, very low pour point and exceptional low-temperature performance [170]. Such oils however are usually expensive due to the relatively high cost of the olefinic feed. Synthetic lubricant base stocks can be prepared in good yields by oligomefizing long-chain alkenes using catalysts containing large pore zeolites with high Si/Al ratios. Internal alkenes are less reactive than the corresponding alphaalkenes and conversions decrease as the chain-length of the feed alkene increases. In general, however, the zeolites thus far reported are not as good as clay catalysts or the curently used boron trifluoride or aluminittm chloride catalysts. 7.

ISOMERIZATION OF LONG-CHAIN ALKANES

The need for lubes and middle distillate fuels with greater performance, safety, and environmental advantages is increasing. This need has focused attention on highly paraffmic feedstocks due to their high oxidation stability, low volatility for a given viscosity, and high viscosity index (>130). Because highly paraffmic feeds tend to have high wax contents, however, the production of lubes and fuels from these feeds has been limited due to the large loss upon wax removal. An alternative approach is to change the molecular structure of the wax by isomerization, such that low pour point, high performance products can be prepared with a high yield. Since isomerization preserves paraffmicity rather than lowering it, the quality of the feedstocks are maximized. The usefulness of wax isomerization will depend greatly upon its feed flexl"oility, i.e. its ability to produce high yields of dewaxed oils from feeds which vary extensively in boiling range and in chemical composition, particularly wax content [ 171]. Recently anumber of patents have appeared descn~oing the use of zeolites for this isomerization process [ 172]. The catalysts used for dewaxing are usually bifimctional in nature with Pt being the hydrogenation-dehydrogenation component and a large pore typically 12MR zeolite provides the acidic component [1]. ZSM-5 is the catalyst used in Mobil's Distillate Dewaxing (MDDW)

354 or Lube Dewaxing (MLDW) processes. In this process the straight chain, waxy normal or slightly branched alkanes are able to emter the pores where they are selectively cracked and the light products are removed by distillation. Currently more than 70% of the catalytic dewaxing units in operation are based on the Mobil zeolite catalyst and process technology [7].

Table 8 Isomerization of n-octane over Pt catalysts at 1000 psig, 2.8 WHSV, 16 H2/HC, and 30% conv.[177] Catalyst Temperature i-C8sel. 2M-C7/3M-Cr [*c] (wt.%) SiO2-AI2Oz HY ZSM-5 (80 SIO2/AI203) ZSM-5 (650 SIO2/AI203) Na-Beta SAPO- 11

C3+Cs/C4 molar ratio

i-CJn-C4

DM-C6sel. (wt. %)

371 257 260

96.4 96.8 56.6

0.67 0.71 1.54

0.95 0.64 2.1

0.96 3.5 1.1

8.5 12 1.8

343

58.4

0.88

1.2

0.98

5.6

367 331

74.3 94.8

0.70 1.07

0.68 1.0

1.7 0.92

10 2.3

302826242220-

= 9

=. . . . . . . . . . .

Pt-SAPO-11

a ........Pt-ZSM-5 - - - o - - R-Silica Alumina

""l,

.~ 1 8 16 .==, o 14=E 1 2 -

/

'1

:~.

1086

...

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

o,

,! 2 0 0

2

4

6

8

10

12

Carbon Number

Figure 15. Molar distn"oution of cracked product l~om hexadecane at 1000 psig, 3.1 WHSV, 30 HJI-IC and 94% conversion [177].

14

355 Catalysts containing a hydrogenation component and an intermediate-pore silicoaluminophosphate (SAPO) molecular sieve have recently been found to have a high selectivity for wax isomerization [ 173,174,175]. A new process for dewaxing high alkane lubricating oils, called Isodewaxing [176], is being commercialized by Chevron using Pt/SAPO-11 catalyst. Table 8 shows the hexadecane isomerization selectivities of a number of Pt -loaded catalysts [ 177,178]. SAPO-11 has a low selectivity to dimethyl isomers such that fewer branches are required to obtain a given degree of pour point reduction. Since increased branching reduces the wide-temperature range fluidity of an oil the oil made using SAPO-11 catalyst has a lower sensit'wity of viscosity to temperature. Figure 15 shows that SAPO-11 had a more even distn'bution over the carbon numbers than is commonly associated with intermediate-pore sieves such as ZSM-5 [177]. The cracked also contained fewer isomers with methyl branches separated by less than two carbons than, for example, silica-alumina. Secondary hydrocracking is low and the cracked by-product is all liquid and at the same time hydrocracking of the long chain alkenes is inh~ited. These properties of SAPO-11 for this application appear to be associated with its moderate acid activity and the one-dimensional nature of its pores. Much of the catalysis appears to occur at or near to the sieve external surface. [178]. In a separate study of the relative activities of USY, mordenite, ZSM-5, Beta and SAPO-11, the SAPO-11 was found to be the only catalyst capable of isomerizing normal alkanes in the presence of iso-alkanes without large yield losses due to unwanted cracking [179]. Pt-H mordenite and ferrierite have also been used for this reaction. Recently bifimctional forms of Beta have been found to give better isomerization selectivities relative to hydrocracking and this may represent a superior and economically attractive dewaxing process [3,66]. 8.

ACKNOWLEDGEMENTS

The authors wish to thank all their colleagues both from academic and industrial research groups who kindly contributed much of the source material used in this paper. Their kind assistance and ready response to a request for information is much appreciated.

REFERENCES

A. Corma, Zeolite Microporous Solids: Synthesis, Structure and Reactivity", E.G. Derouane et al (eds.), Kluwer Academie Publishers (1992) 373. W.O. Haag, in "Zeolites and Related Microporous Materials: State of the Art 1994", J. Weitkamp, H.G. Karge, H. Pfeifer and W. Holderich (eds.), Stud. Surf. Sci. and CataL 84, Elsevier, Am~erdam, (1994) 1375. '~atalysis by Microporous Materials", ILK. Beyer, H.G. Karge, I. Kiricsi and J.B. Nagy (eds.), J. Weitkamp, U. Weiss and S. Ernst, Stud. Surf Sci. Catal., 94, Elsevier, Am~erdam, (1995) 363. S.T. Sie, in "Advanced Zeolite Science and Applications", J.C. Jansen, M. Stocker, H.G. Karge and J. Wekkamp (eds.), Stud. Surf. Sci. Catal., 85, Elsevier, Am~erdam, (1994) 587. S.M. Csiscery, in '~atalysis by Microporous Materials", ILK Beyer, H.G. Karge, I. Kiricsi and J.B. Nagy (eds), Stud. Sure Sci. Catal., 94, Elsevier, Am~erdam, (1995) 1.

356 J.E. Naber, K.P. de Jong, W.H.J. Stork, H.P.C.E. Kuipers and M.F.M. Post, in "Zeolites and Related Microporous Materials: State of the Art 1994", J. Weitkamp, H.G. Karge, H. Pfeifer and W. Holderich (eds.), Stud. Surf. Sci. Catal., 84, Elsevier, Amsterdam~ (1994) 2197. I.E. Maxwell and W.H.J. Stork, in '~ntroduction to Zeolite Science and Practice", I-l, van Bekkum~ E. Flanigen and J.C. Jansen (eds.), Stud. Sure SoL Catak, 58, Elsevier, Amsredam, (1991) 571. .

A. Corma, Catalysis Letters, 22 (1993) 3. J. Dwyer and D.J. Rawlence, Catalysis Today, 18 (1993) 487-507.

9. 10.

A.A. Avidan, M. Edwards and I-I, Owen, Oil and Gas Journal, (Jan. 1990) 33.

11.

J. Biswas and I.E. Maxwell, AppL Catal, 63 (1990) 259.

12.

R. Rajagopulan and E.T. I-labib, Jr., Hydrocarbon Processing, (Sept. 1992) 43.

13.

C.J. Groenenboom, in "Zeolites as catalysts, sorbents and detergent builders", H.G. Karge and J. Weitkamp (Eds), Elsevier, Am~erdam, (1989) 99. E. Brevoord and J.R. Wilcox, Akzo Nobel Catalysts Symposimn, Drakensberg, South Africa, March 1995.

14. 15.

C. Thomas and D.S. Barmby, J. Catak, 12 (1968) 341.

16.

A. Corma, V. Fomes, ,4_ Martinez, F. Melo, O. Pallota, in "Innovations in Material Science", P. Grobet (ed.), Stud. SurE. Sci. Catak, 37, Elsevier, Am~erdam, (1988) 495.

17.

P.E. Dai, D.E. Sherwood, B.I-I, Bartley, US Patent, 5, 354,452 (1994).

18. 19.

P.J. Nat and E.T.C. Vogt, Akzo Nobel, WD 94/26847 (1994). H. Th. Rijnten, Akzo Nobel Catalysts Symposium, Drakensberg, South Africa, March 1995. B.A. Aufdembrink et ak, US Patent, 5, 258, 114 (1993). R.J. Nozemarck, Brit. Pat. 2 079 737 (1982). P. van de Gender and ICY. Yung, Akzo Nobel Catalysts Symposium, Drakensberg, South Africa, March 1995. J. Biswas and I.E. Maxwell, AppL Catat, 58 (1990) 1.

20. 21. 22. 23. 24. 25. 26.

A.A. Avidan, Akzo Catalysis Sy~osium, 1991. K.Y. Yung, P. O'Connor, N.M. van't Klooster and S.J. Yanik~ Akzo Catalysts Symposium, Drakensberg, South Africa, March 1995. F.G. Dwyer and T.F. Degnan, in "Fktid Catalytic Cracking: Science and Technology", J.S. Magee and M.M. Mitchell (eds.), Elsevier, Am~erdam, (1993) 499.

27.

F.N. Guerzoni and J. Abbot, AppL Catat A: General, 120 (1994) 55.

28. 29.

V. Choudhary and D. Akolekar, J. Catak, 117 (1989) 542. A. Corma, M. Faraldos, A. Martinez and ,4_ Mifsud, J. Catak, 122 (1990) 230.

30.

J. Dwyer, J. Dewing, I~ Karim, S. Holmes, A,F. Ojo, A.A. Garforth and D.J. Rawlence, in "Zeolite, Chemistry and Catalysis", P.A. Jacobs et al (eds.), Elsevier, Amgerdam (1991) 1. A. Corma, P.J. Miguel and A.V. Orchilles, J. Catal., 145 (1994) 171.

31.

357 32.

Oil and Gas Journal, Mar, 28 (1994) 41.

33. 34.

G.C. Edwards and A.W. Peters, EP Pat., 0 243 629 (1987). J.P. Gilson, EP Pat., 0 258 726 (1987).

35.

L. Bonetto, M.A. Camblor,/L Corma and J. Perez-Pariente, Appl. Catal., A, 82 (1992) 37. A. Corma and A. Martinez, Catal. Rev. SO_ Eng., 35(4) (1993) 483.

36. 37.

F.G. Dwyer, P.J. Lewis, and F.H. Schneider, Process Technology, ChelxL Eng., 83 (1976) 90.

38.

P.J. Lewis, and F.G. Dwyer, Oil and Gas Journal, 75 (1977) 55.

39.

Newsbdet~ Badger, 11/11/1991.

40. 41.

P.B. Venuto, Microporous Materials, 2 (1994) 297. Chem_ Week, 143(22) (1988) 9.

42.

Hydrocarbon Process, 66(6) (1987) 57.

43. 44. 45.

Itoh, H., Miyamoto A. and Miwakami Y., J. CataL, 64 (1980) 284 J. Engelhardt, J. Szanyi and J. Valyou, J. CataL, 107 (1987) 296 S. Han, D.S. Shihabi, ILP.L. Absil, Y.Y. Huang, S.M. Leiby, D.O. Marler and J.P. McWilliams, Oil & Gas J., Aug. 21 (1989) 83.

46. 47. 48.

N.Y. Chen and T.F. Degnan, Chem_ Eng. Progress, (Feb. 1988) 32.

49.

G. Mirth, J. Cejka and J.A. Lercher, J. Catal. 139 (1993) 24.

50.

US Patent, 5,243,121; D. Harris, Private Comm--~cation.

51

W.W. Kaeding, L.B. Young and A.G. Papras, Chemtecla, 12 (1982) 418.

52.

Kirk-Othmar Encyclopedia of Chemical Technology, 4th ed., John Wiley & Sons Inc., 1992. G.R. Meima, M.J.M. van der Aaist, M.S.U. Samson, J.M. Garces and J.G. Lee, Proceedings of 9th Int. Zeolite Conference, IL von Ballmoos et al (eds.) Montreal, (1992) 327. US Patent, 5,198,595. D.L. King, New Cumene Process, Catalytica Highlights, 1992. B.S. Rao, I. Bahkrishnan, V.1L Chttmnhala, A.1L Pradhan and P. Ratnasamy, in "Proc. 1st Tokyo Conf. on Adv. Catalyst Science Technology", VoL 1, S. Yosida et al. (eds.), Tokyo, 1990, 361.

53.

54. 55. 56.

Y. Sugi and M. Toba, Catal. Today, 19 (1994) 187. F. Gorra, L.L. Breckenridge. and 1LA. Sailor, Oil and Gas Journal, Oct. 12, (1992).

57.

G. BeUussi, G. Pazzuconi, C. Perego, G. Girotia and G. Terzoni, J. CataL, 159 (1995).

58.

A. Corma, Catalysis Letters, 22 (1993) 3.

59.

J.L. Goncalves de Almeida, M. Dufaux, Y.B. Taarit, C. Naccache, Appl. Catal. A: General, 114(1994) 141.

60.

S. Sivasanker, A. Thangaraj, 1LA. Abdulla and P. Ramasamy, Proceedings of 10th Int. Congress on Catalysis, 19-24 July 1992, Budapest, Hungary, (1992) 397. 1L Lobo and M.E. Davis, J. Amer. Chem_ Soc., 117 (1995) 3766.

61.

358 62.

J. Weitkamp, U. Weiss and S. Ernst, Stud. Surf. Sci_ Catal., 94 (1995) 363.

63.

S-J. Chu and Y-W Chen, Appl. CataL A: General, 123 (1995) 51.

64.

G.S. Lee, J.J. Maj, S.C. Rocke and J.M. Garces, Catalysis Letters, 2 (1989) 243.

65.

J.A. Horsley, J.D. Fellmann, E.G. Derouane and C.M. Freenmn, J. Catal., 147 (1994) 231.

66.

N.Y. Chen et al., US Pat. 4 851 109.

67.

G. Mirth and J.A. Lercher, J. Phys. Chem, 95 (1991) 3736.

68.

W.W. Kaeding and S.A, Butter, J. Catak, 61 (1980) 155.

69.

A.P. Lien and D.A. McCauley, J. Amer. Chem Sot., 74 (1952) 6246.

70.

H.G. Karge, Y. Wada, J. Weitkamp, S. Ernst, U. Girbach and H.I~ Beyer, Stud. Surf Sci. CataL, 19 (1984) 11. A. Corma and E. Sastre, Chemical Communications, (1991) 594.

71. 72.

74.

Y.S. Bhat, A.B. Halgeri and T.S.IL Prasado Rao, Ind. Eng. Chem. Res., 28 (1989) 890. J. Das, Y.S. Bhat, A.I. Bhardwaj. and A.B. Halgieri, Appl. CataL A: General 116 (1994) 71. I. Wang, T.-C. Tsai, S.-T. Huang, Ind. Eng. Chem. Res., 29 (1990) 2005.

75.

W.W. Kaeding, L.B. Young and C.-C. Chu, J. Catal., 89 (1984) 267.

76.

A. Corma, G. Sastre, F. Melo, P. Viruela, and C. Zicovich-Wilson, Proceedings of 9th Int. Zeolite Conference, 1L von Ballmoos et al (eds.) Montreal, (1992) 343.

77.

A. Corma, F. Llopis, F. Melo, P. Viruela and C. Zicovich-Wilson, J. Amer. Chem. Soc. 116 (1994) 134.

79. 80.

S. Csicsery, Zeolites, 4 (1984) 202. P. Weis~ Pure AppL Chem., 52 (1980) 2091. W. O. Haag, 1LM. Lago and P.B.Weisz, Faraday Discuss. Chem. Sot., 72 (1982) 317.

73.

81.

J.A. Wei, J. CataL, 76 (1982) 433. H. Vinek and J.A. Lercher, J. Mol. Catal., 64 (1991) 23.

82. 83. 84.

L.B. Young, S.A. Butter and W.W. Kaeding, J. CataL, 76 (1982) 418.

85.

N.Y. Chen, W.W. Kaeding and F.G. Dwyer, J. Am. Chem. Soc., 101 (1979) 6783.

86.

G.V. Bhaskar and D.D. Do, Ind. Eng. Chem_ Res., 29 (1990) 355.

87.

E.G. Derouane, Stud. Surf. Sci. CataL, 5 (1980) 541.

88.

D. Fraenkel, M. Chemiavsky and M. Levy, Proc. 8th Int. Congr. on Catalysis, Berlin, DECHEMA (ed.), Verlag Chemie, Weinheim, VoL 4, (1984) 545.

89.

V.S. Nayak and L. Riekert, Appl. CataL, 23 (1986) 403.

90. 91.

D. Fraenkel, Ind. Eng. Chem. ges., 29 (1990) 1814. T. Xu, M. Matsumoto, T. Maeda, Y. Sugi, T. Matsuzaki, T. Hanaoka, Y. Kubota, J.H. Kim~Microporous Materials, 3 (1995) 593.

92.

J.-H. Kim~ T. Matsuzaki, I~ Takeuchi, T. Hanaoka, Y. Kubota, Y. Sugi, X. Tu and M. Matsumoto, Trans. Mat. ges. Soc. Jpn., 15A (1994), 153.

359 93.

P.A. Parikh, N. Subrahmanyan, Y.S. Bhat and A.B. Halgeri, J. MoL CataL, 88 (1994) 85.

94. 95.

D. Fraenkel, M. Chemiavsky, B. Ittah, M. Levy, J. Catal., 101 (1986) 273 J. Weitkamp, M. Neuber, Stud. Surf ScL CataL, 60 (1990) 291.

96.

European Patent, 0 280 055

97.

S-J. Chu and Y-W. Chen, Ind. Eng. Chem ges., 33 (1994) 3112.

98.

J.L. Almeida, M. Dufaux, Y.B. Taarit, C. Naccache, Appl. Catal. A:General, 114 (1994) 141. W.W. Kaeding, C. Chu, L.B. Young, B. Weinstein and S.A. Butter, J. CataL, 67 (1981) 159.

99. 100.

G. Paparatto, E. Moretti, G. Leofanti and F. Gatti, J. Catal., 105 (1987) 227.

101.

C.S. Lee, T.J. Park and W.Y. Lee, AppL CataL A: General, 96 (1993) 151.

102.

T. Hibino, M. Niwa, Y. Murakami~ J. CataL, 128 (1991) 551.

103.

I. Wang and C.-L. Ay, Proc. 9th Int. Congr. on CataL, Calgary, 1 (1988) 324.

104. 105.

T. Matsuda, K. Yogo, Y. Mogi and E. Kikuchi, Chem_ Lett., (1990) 1085. T. Matsuda, T. Urata, I~ Saito and E. Kikuchi, Prepr. 37th S3anp. on Petrochemistry, Japan Petroleum Institute, ( 1991) 80.

106.

I. Wang, C.-L. Ay and B.-J. Lee, Chen M.-H., Appl. Catal., 54 (1989) 2571. I-I. Vinek and J.A. Lercher, J. MoL CataL, 64 (1991) 23.

107. 108.

I-LG. Karge, L. Ladebeck, Z. Sarback and tC Hatada, Zeolites, 2 (1982) 94.

109.

H.G. Karge, Y. Wada, J. Weitkan~, S. Ernst, U. Girbach and H.I~ Beyer, Stud. Surf ScL Catal., 19 (1984) 11.

110.

N.P. Rhodes and 1L gudham, J. Chem_ Soc. Farady Trans., 89 (1993) 2551.

111. 112.

G.V. Bhaskar and D.D. Do, Ind. Eng. Chem ges., 29 (1990) 355.

113. 114. 115. 116. 117. 118.

E. Sastre, A. Corma, F. Fajula, F. Figueras and J. Perez-Pariente, J. CataL, 126 (1990) 457. I. Kiricsi, C. Flego. and G. BeUussi, Appl. CataL A" General, 126 (1995) 401. A.W. O'Donovan and C.T. O'Connor, 11th Int. Zeolite Conference (In Press). P.C. Doolan and ElL Pujado. Hydrocarbon Processing, 68 (1989) 72. S.M. Bradley and ILA. Kydd, J. CataL, 142 (1993) 448. 1L Nash, M.E. Dry and C.T. O'Connor, AppL CataL, (In Press). W.O. Haag, Acid Catalysis with Medium Pore Zeolites. Proc 6th Int. Zeolite Conference, Olson and Bisio (eds.), Butterworths, 1984.

119.

D.B. Lukyanov, N.S. Gnep and M.1L Cnfisnet, Ind. Eng. Chem_ ges., 33 (1994) 223.

120.

D.B. Lukyanov, N.S. Gnep and M.1L Cnfisnet, Ind. Eng. Chem ges., 34 (1995) 516.

121.

S.B.A. Hamid and E.G. Derouane, Stud. Surf Sci. CataL, 84 (1994) 2335.

122.

D. Seddon, Catalysis Today, 6 (1990) 351.

123.

S.K. lhm tC-H. Yi, Y.-K. parL Stud. Surf. Sci. Catal., 84 (1994) 1765.

124.

Y. Ono, IC Osako, G.-J. Kim and Y. Inoue, Stud. Surf. Sci. Catal., 84 (1994) 1773.

360 125. 126.

E.S. Shpiro, D.P. Shevchenko, O.P. Tkachcnko and ILV. Dmitfiev, Appl. Catal. A: GeneraL, 107 (1994) 147. S.B. Harold, E.G. Dcrouanr G. Demorticr, J. Riga and M.A. Yarmo, Appl. CataL A General, 108 (1994) 85.

127.

P. Mcriaudeau, S.B. Abdul I-lamid, C. Naccache, J. CataL, 139 (1993) 679.

128.

G. Giannctto, A. Montes, N.S. Creep, A. Florentino, P. Catraud and A, Cmisnet, J. CataL, 145 (1994) 86.

129.

V.1L Choudhary, A.K. Kinage, C. Sivadinarayana and S.D. Sansare, Catalysis Letters, 33 (1995) 401.

130.

J.R. Bernard, Proc 5th IntL ConfZeolites, Heydon, London, (1980)686.

131.

P.W. Tamm D.H. Mohr and C.1L Wilson, Stud. Surf. Sci. Catal., 38 (1987) 335.

132.

L. Mank, A. Minkldnen and 1L Shaddich, H.T.I. Quarterly, (1992) 69.

133. 134.

W. Hau, A.B. Kooh and ILF. Hicks, Catal. Letters, 18 (1993) 219. C. Besoukhanova, J. Guidot and D. Barthomettt~ J. Chem Soc. Far. Trans., 1.77 (1981) 1595. M. Su~moto, H. Katstmo, T. Hayasaka, N. lghikawa and K. I-~asawa, AppL CataL A: General 102 (1993) 167.

135. 136.

C. Dossi, 1L Psaro, K Bartsch, hi. Fusi, L. Sordelli, IL Ugu, M. Bdlatreccia, IL Zanoni and G. Vlaic, J. CataL, 145 (1994)377.

137.

T. Tatsumi, L. Dai and H. Sakashita, CataL Letters, 27 (1994) 289.

138.

D.J. Ostgard, L. Kustov, K.1L Pocppelmeier and W.M. Sachtler, J. CataL, 133 (1992) 342.

139.

E. Mielczanski, S.B. Hong, 1LJ. Davis and M.E. Davis, J. Catal., 134 (1992) 349. M. Su~moto, T. Murakawa, T. tgtrano and I4_ OhashL Appl. CataL A: General, 95 (1993) 257. F. Foger and H. Jaeger, AppL CataL, 56 (1989) 137. E. Iglcsia and J.E. Baumgartner, Proceedings 10th Int. Congress Catal, Budapest, (1993) 993. A.B. Kooh, W. Han and 1LF. Hicks, CataL Letters, 18 (1993) 209. G.B McVicker, J.L. Kao, J.J. Ziemiah, W.E. Gates, J.L. Robbins, M.M. Treacy, S.B. Rice, T.H, Vanderspurt, V.IL Cross and A.K. Ghosh, J. CataL, 139 (1993) 48.

140. 141. 142. 143. 144. 145.

J. Zheng, J. Dong, Q. Xu, Y. Liu and A. Yan, Appl. Catal., A: G~eral, 126 (1995) 141.

146.

P.G. Smimiotis and E. guckestein, Appl. Catal, A: General, 123 (1995) 59.

147.

J.a. Sinfelt in '~atalysis Science and Technology", J.IL Anderson and M. Boudard (eds.), Vol 1, Chapter 5, Springer-Verlag (1981).

148.

E. Dnmitriu, V. Hulea, C. Chelara and T. Hulea. in "Zeolites and Related Microporous Materials: State of the Art 1994", J. Weitkamp, H.G. Karge, H. Pfeifer, W. H61derich (eds.), Elsevier, Amsterdam, Stud. Sure Sci. Catal, 84 (1994) 1997.

149.

V.1L Choudhary, Chem Ind. Dev., 8 (1974) 32.

150.

D.H. Powers, B.D. Murray, B.H.C. Winquist, E.M. CaUendar and J.I-L Vamer, European Patent No. 0 523 838 A2 (1993)

361 151.

ILL Pellet, P.IC Cou~hlin, E.S. Sham~oum and J.A. Rabo, ACS Syrup. Ser. 368 (1988), Ch. 33.

152.

A. Corma, J. Frontela, J. Lazaro and M. Perez, American Chemical Society, New York City Meeting, August 25-30 (1991).

153.

J.A. Moulijn, ILA. Sheldon, H. van Bekkum and P.W.N.M. van Leeuwen, Catalysis, Elsevier, Amqterdam, (1993) 33.

154.

J. Szabo, J. Perrotey, G. Szabo, J.C. Duchet and D. Comet, J. Mol. Catak, 67 (1991) 79

155. 156.

V.IL Choudhary and L.I~ Doraiswamy, J. Catal., 23 (1971) 54. Z.X. Cheng and V. Ponce, Appt Catal. A: General 118 (1994) 127.

157.

S.A.I. Ban~ D.W. Walker and IL Tahir, European Patent No 247 802 (1987).

158.

C.-U O~/'oung, W.-Q. Xu, M. Simon and S.L. Suib in "Zeolites and Related Mieroporous Materials: State ofthe Art 1994", J. Weitkamp, tLG. Karge, H. Pfeifer and W. H61derich, (eds.), Elsevier Science, Amsterdam, Stud. Sure Sci. Catal., 84 (1994) 1671.

159.

S.M. Yang, D.I-I. Guo, J.S. Lin, G.T. Wang, in "Zeolites and gelated Microporous Materials: State of the Art 1994", J. Weitkamp, H.G. Karge, H. Pfeifer and W. Holderich, (eds), Elsevier, Amgerdam, Stud. Sure Sci. Catal, 84 (1994), 1677.

160.

H.H. Mooiweer, K.P. de Jong, B. Kraushaar-Czarnetzld, W.H.J. Stork and B.C.H. Krutzen, in "Zeolites and Related Microporous Materials: State of the Art 1994", J. Weitkamp, H.G. Karge, H. Pfeifer and W. H61derich (eds.), Elsevier, Amsterdam, Stud. Surf. SoL Catal., 84 (1994) 1677.

161.

A.C. Butler and C.P. Nicolaides, Catak Today, 18 (1993) 443.

162.

ICF.N.J.G. Scholle, A.P.M. Kentgens, W.S. Veemans, P.F. Frenken and G.P.M. van der Velden, J. Phys. Chem_, 88 (1984) 5.

163.

W.-Q. Xu, S.U Suib and C.-L. O~fotmg, J. Catal., 144 (1993) 285.

164.

D. Bianchi, M.W. Simon, S.S. Nam, W.-Q. Xu, S.L. Suib and C.-L. O~/'oung, J. Catak, 145 (1994) 551.

165.

M. Cmisnet, P. Andy P., N.S. Gnep, C. Travers and E. Benazzi, J. ChexIL Sot., Chem_ Commun., (1995) 1685. ILJ. Quarto and F.J. Krambeck in Chemical Reactions in Complex Mixtures, Proceedings of Mobil Symposium, F.J. Krambeck and A.V. Sapre(eds.), van Nostrand, New York, (1991) 143. C.T. O'Connor and M. Kojima, Catalysis Today, 6(3) (1989) 329. E. Kohler, F. Schmidt, H.J. Wemicke, M. de Pontes and H.L. Roberts, HTI Quarterly, Summer, (1995) 37.

166.

167. 168. 169.

S.T. Langford, J.C.Q. Fletcher and C.T. O'Connor, Proceedings of 9th Int. Zeolite Conference, IL von Ballmoos et al (eds.) Montreal, (1992) 467.

170.

ILL. Shublin, M.S. Baylerian, A.IL Mayle, Ind. Eng. Prod. ges. Dev., 19 (1980) 15.

171.

S.J. Miller, J. Xiao and J.M. gosenbaum, Science Technology in Catalysis 1994, Kodansha, (1995) 379.

172.

J. Lucien, E.P. Pat. 0 324 528 B 1 (1989).

362 173.

S.J. Miller, US Pat. 4 689 138 (1987).

174.

S.J. Miller, US Pat. 4 859 311, (1989). S.J. Miller, Appl. Catal. A., 119, 121-138, (1994). S.J. Miller, M.A. Shippey and G.M. Masada, NPRA Natl_, Fuels and Lubr. Mtg, (1992), Paper No FL-92-109. S.J. Miller, Microporous Materials, 2 (1994) 439. S.J. Miller, in "Zeolites and Related Microporous Materials: State of the Art 1994" J. Weitkamp, H.G. Karge, H. Pfeifer and W. Holderich(eds.), Elsevier, Am~erdarn~ Stud. Surf. Sci_ Catal., 84 (1994)2319. R.J. Taylor and R.H. Petty, Appl, Catal. A: General, 119 (1994), 121. P.S.E.Dai, J.R, Sanderson and J.F. Knifton, in "Zeolites and Related Microporous Materials: State of the Art 1994" J. Weitkmaxp, H.G. Karge, I-I. Pfeifer and W. Holderich (Eds), Elsevier, Am~erdam, Stud. Surf. Sci. Catal,, 84, (1994) 1701.

175. 176. 177. 178.

179. 180.