gas phase catalysis by zeolites

www.elsevier.com/locate/micromeso

Microporous and Mesoporous Materials 82 (2005) 257–292

Gas phase catalysis by zeolites

Michael Stocker *

SINTEF Materials and Chemistry, Department of Hydrocarbon Process Chemistry, P.O. Box 124 Blindern, N-0314 Oslo, Norway

Received 25 April 2004; received in revised form 15 November 2004; accepted 20 November 2004

Available online 8 April 2005

Dedicated to Diplom-Chemiker Ulf Blindheim on the occasion of his 70th birthday

Abstract

This paper provides an overview about today�s use of zeolites and related microporous materials as catalysts within the fields ofrefining, petrochemistry and commodity chemicals. The content of this presentation is devoted to gas phase catalysis—with focus on

acid catalysis, hydrocarbon conversion and formation, oil and natural gas upgrading as well as catalytic probe reactions for the

characterisation of zeolites and related microporous materials. The review is primarily meant for beginners who intend to get

acquainted with this field. However, for more detailed information the interested reader is invited to consult the dedicated papers

cited throughout this overview.

� 2005 Elsevier Inc. All rights reserved.

Keywords: Zeolites; Microporous materials; Gas phase catalysis; Crude oil upgrading; Natural gas conversion

1. Introduction

Catalysis by zeolites—with focus on hydrocarbon

conversion and formation—covers nowadays a broad

range of processes related to the upgrading of crude oil

and natural gas. This includes, among others, fluid cata-

lytic cracking (FCC), hydrocracking, dewaxing, aliphate

alkylation, isomerisation, oligomerisation, transforma-

tion of aromatics, transalkylation, hydrodecyclisation

as well as the conversion of methanol to hydrocarbons.All these conversions are catalysed by zeolites or related

microporous materials, based both on the acid pro-

perties and shape-selective behaviour of this type of

materials.

The first part of this chapter deals with the under-

standing of the chemistry of acid catalysis using zeolites

or related microporous materials, including the for-

mation of acid sites, carbocation chemistry and their

1387-1811/$ - see front matter � 2005 Elsevier Inc. All rights reserved.doi:10.1016/j.micromeso.2005.01.039

* Tel.: +47 98 24 39 33; fax: +47 22 06 73 50.

E-mail address: [email protected]

reaction mechanisms, as well as the importance of the

shape-selectivity of the microporous materials. The term‘‘zeolite’’ is used for the microporous aluminosilicate

systems, however, SAPO type catalysts belong to the

family of microporous materials as well.

The second part of this chapter covers the discussion

of the present situation and the new developments re-

lated to the above mentioned petroleum refining and

natural gas conversion processes using microporous

materials, with focus on the new requirements due tothe introduction of new fuel specifications world-wide.

Finally, the third part of this chapter is dedicated to

the different probe reactions with respect to the charac-

terisation of zeolites and related microporous materials.

Since the micropores of zeolites and related compounds

have diameters in the range of molecular dimensions,

the shape-selective effect reveals unique possibilities in

the catalytic conversion of, for example, hydrocarbons.However, proper utilisation of this behaviour requires

that the conversion occurs at active sites connected

to the internal pore structure and not at the external

mailto:[email protected]

258 M. Stocker / Microporous and Mesoporous Materials 82 (2005) 257–292

surface of the zeolite crystals. A suitable possibility for

the investigation of relative activities of the internal

and external surface of microporous solids is the appli-

cation of appropriate probe molecules, either for the

investigation of their interaction with the internal and

external surface of the zeolite or as model compoundsduring catalytic conversions.

Fig. 1. Brønsted acid sites (‘‘bridging hydroxyl groups’’) in zeolites

[1,2].

Fig. 2. Formation of Lewis acid sites in zeolites (simplified version—

not taking into account the model of ‘‘true Lewis acid sites’’) [1].

2. Acid catalysis and shape selectivity of zeolites and

related microporous materials with respect to

hydrocarbon conversion and formation

2.1. Acid sites

Zeolites and related microporous molecular sieves

consist of a three-dimensional network of metal–oxygen

tetrahedra (in a few cases also octahedra) which provide

the periodically sized microporous structure, in which

the active sites are part of the structure. Acid sites result

from the imbalance of the metal and the oxygen formal

charge in the primary building unit. This can easily berecognised in the case of zeolites, which consist of a

three-dimensional network of Si–O tetrahedra. A lattice

comprising of only Si–O tetrahedra is neutral (the 4+

charge at the silicon is balanced by four oxygen atoms

with each 2� charge, however, belonging to two tetrahe-

dra). Replacing one Si4+ atom by Al3+ causes a formal

charge on the tetrahedron of 1�. This negative charge

is then balanced by a proton or metal cation formingan acid site. The bare, negatively charged tetrahedron

is then the corresponding base. Please keep in mind that

these acid and base properties are not just a function of

the chemical composition, since other factors, like the

framework density, the type of cation or the local strain

have an influence as well [1].

In AlPO4 type microporous materials the framework

structure consists of a strictly alternating Al–O–P se-quence (Al3+ and P5+, balanced by four oxygen atoms

with each 2� charge, however, belonging to two tetrahe-

dra), resulting in a completely neutral lattice as well, like

in the case of pure silica zeolites. Depending on the com-

binations of the metal cation in the lattice, frameworks

with positive or negative charges are in principal possi-

ble, however, so far only cation exchanged microporous

materials are known.Several industrial applications of zeolites are based

upon technology adapted from the acid silica/alumina

catalysts originally developed for the catalytic cracking

reaction. This means, that the activity requested is based

on the formation of Brønsted acid sites arising from the

creation of ‘‘bridging hydroxyl groups’’ within the pore

structure of the zeolites. These ‘‘bridging hydroxyl

groups’’ are usually formed either by ammonium orpolyvalent cation exchange followed by a calcination

step. The ‘‘bridging hydroxyl groups’’, which are pro-

tons associated with negatively charged framework oxy-

gens linked into alumina tetrahedra, are the Brønsted

acid sites, as demonstrated in Fig. 1 [2].The protons are quite mobile at higher temperatures,

and at 550 �C they are lost as water molecules followedby the formation of Lewis acid sites, as shown in Fig. 2

[2].

For zeolites, it can be stated that the concentration of

aluminum in the lattice is directly proportional to the

concentration of acid sites. However, for other micropo-

rous solids, corresponding correlations are not straight-forward [3].

2.2. Carbocations

In general, the nature of acid sites in zeolites is well

understood, however, there is much less consensus on

the reaction mechanisms for hydrocarbon conversion

or formation over microporous materials. It is generallyaccepted that the reaction mechanisms of hydrocarbon

conversion and formation on acid zeolites and related

catalysts involve the formation of carbocations. How-

ever, whether these carbocations act as transition states

or as intermediates is still under discussion, and is, in

addition, depending on the type of hydrocarbon. The

behaviour of carbocations and their reaction pathways

Fig. 3. Representation of alkylcarbenium (a) and alkylcarbonium ions (b–c). R represents either hydrogen or alkyl group [4]. Reproduced by

permission of Elsevier, Amsterdam.

M. Stocker / Microporous and Mesoporous Materials 82 (2005) 257–292 259

in zeolites and related microporous materials are

strongly depending on the shape-selective effect due to

the confinement of the reacting molecules in the micro-structure of the catalysts—offering very restricted space

[4].

Concerning the type of carbocations related to the

conversion or formation of hydrocarbons, one has to

distinguish between alkylcarbenium ions (containing a

tri-coordinated positively charged C-atom with three

substituents being either alkyl groups or hydrogens)

and alkylcarbonium ions (consisting of a penta-coordi-nated positively charged C-atom with the same type of

substituents). See also Fig. 3 [4,5].

The stability of alkylcarbenium ions depends on the

inductive effect of the substituents on the positively

charged C-atom, with the tertiary alkylcarbenium ions

as the most stable ones. However, this effect is less pro-

nounced for the alkylcarbonium ions [4].

In the following, the behaviour of acid sites and theimportance of carbocations in connection with catalytic

conversions using zeolites or related microporous solids

is demonstrated for the cases of aliphatic hydrocarbon

cracking (C–C bond scission) and for the alkylation of

isobutane with n-butene (C–C bond formation). For

more detailed reviews regarding the reaction mecha-

nisms of acid catalysed hydrocarbon conversions the

interested reader should consult one of the following ref-erences [1,4].

2.3. Mechanistic pathways for catalytic cracking of

aliphatic hydrocarbons on zeolites (C–C bond scission)

In general, catalytic cracking reactions of hydrocar-

bons using zeolites can be classified according to the fol-

lowing three main mechanistic pathways:

1. Classical cracking mechanism consisting of a hydride

transfer step to a carbenium ion followed by b-scission.

2. Non-classical Haag-Dessau (protolytic) cracking

mechanism proceeding via a carbonium ion transi-

tion state.

3. Oligomerisation cracking.

The classical cracking mechanism is based on the fact

that a carbenium ion abstracts a hydride from an alkane

forming another carbenium ion, which cracks by b-scis-sion (cleavage of the C–C bond located b to the trivalentpositively charged carbon atom), forming an alkene—

see also Fig. 4 [6].

The overall process is governed by the stability of the

carbenium ions in the different states of the reaction. In

addition, the reaction rate decreases in the sequence

tertiary > secondary > primary carbenium ions formed.

Furthermore, the activation energy usually increaseswith increasing energy level of the final state. Therefore,

the rate for reactions starting from a tertiary carbenium

ion and ending with a tertiary carbenium ion (type A in

Fig. 5) is faster than the reaction starting from and end-

ing with a secondary carbenium ion (type C in Fig. 5).

The different reaction pathways for the b-scission mech-anism are summarised in Fig. 5. Please note that these

rather simple assumptions for the b-scission mechanismcorrespond quite well with the cracking selectivity ob-

served [1,7,8].

Dehydrogenations are efficiently catalysed on the me-

tal sites of bi-functional catalysts, since unsaturated

compounds are much more strongly adsorbed on the

acid sites forming classical carbenium ions than the

saturated ones forming non-classical carbonium ions.

Therefore, classical cracking clearly dominates in

Fig. 4. Classical cracking mechanism for an alkane molecule [6].

Reproduced by permission of Elsevier, Amsterdam.

Fig. 6. Non-classical (protolytic) Haag-Dessau cracking mechanism

for an alkane molecule [6]. Reproduced by permission of Elsevier,

Amsterdam.


bi-functional catalysts. However, classical cracking can

also take place on mono-functional acid catalysts, but

in this case the carbenium ions have to be formed in asterically demanding, bi-molecular hydride transfer [9].

Large pore zeolites like Y zeolite show usually a greater

tendency to crack the hydrocarbons according to the

classical cracking mechanism. However, the small and

medium pore zeolites like ZSM-5 favour the non-classi-

cal Haag-Dessau mechanism which allow mono-molecu-

lar reactions while restricting the bi-molecular (hydride

transfer) reactions due to steric limitations in the pores.The Haag-Dessau mechanism (see Fig. 6) is the key to

unravel the competing mechanisms of catalytic cracking,

including the classical cracking and oligomerisation

Fig. 5. b-Scission mechanism for secondary and tertiary alkylcarbenium

cracking. Understanding of non-classical cracking hashelped in the diagnosis of shape-selectivity and mass

transfer effects in zeolite-catalysed cracking [6].

From the work of Olah concerning the hydrocarbon

chemistry in superacids it was known that alkanes can

be protonated at low temperatures in liquid phase.

However, Haag and Dessau postulated their mechanism

in 1984 by demonstrating that even zeolites can proton-

ate alkanes to give carbonium ions—which are transi-tions states in cracking [10]. The carbonium ions

collapse to give the cracking products: to begin with alk-

anes (or hydrogen) and smaller carbenium ions, which

then release protons to form the final cracking product

ions [1,8]. Reproduced by permission of Wiley-VCH, Weinheim.

Fig. 7. Schematic reaction pathways for carbonium ion decay (protolytic cracking) of a protonated 3-methylpentane together with principle

transition states for dehydrogenation and cracking [1]. Reproduced by permission of Wiley-VCH, Weinheim.


consisting of alkenes [6]. Haag and Dessau�s suggestionis based on the carbonium ion decay of a protonated

3-methylpentane molecule, as shown in Fig. 7.Since the decay of the carbonium ion leads to break-

ing of the C–C or C–H bonds at the insertion point of

the proton, the reaction is called protolytic cracking as

well [1]. Furthermore, compared to the super acid chem-

istry in liquid phase, the formation of the carbonium

ions using zeolites is only significant at temperatures

higher than 450 �C, and they exist only in a transitionstate.In conclusion, Haag-Dessau cracking (also called

mono-molecular or protolytic cracking) dominates at

low conversions, high reaction temperatures, low reac-

tant pressures and with small and medium pore zeolites

having a low concentration of Brønsted acid sites. All

these conditions favour a low reactant concentration

in the pores and impede hydride transfer. The decay of

the carbonium ion into an alkane and a smaller carbe-nium ion is the main step in the Haag-Dessau cracking

mechanism [1].

Fig. 8. Simplified reaction network for the cracking of alkanes on zeolites

A simplified reaction network for the catalytic crack-

ing of alkanes using zeolites is shown in Fig. 8 [11].

At higher reactant partial pressure the classical crack-ing mechanism is gradually replaced by oligomerisation

cracking, where we observe substantial oligomerisation

preceding the cracking process. Experimental evidence

for such a route has been demonstrated by Werst et al.

[12] using labeling investigations, and revealing entire

scrambling of carbon-labeled olefinic cracking products.

The importance of this mechanism increases with higher

conversion and higher partial pressure as well as lowerreaction temperatures. However, the fundamental chem-

istry related to this cracking mechanism is basically the

same as observed for the classical cracking mechanism

[1].

2.4. Mechanistic pathway for the alkylation of isobutane

with n-butene on zeolites (C–C bond formation)

Zeolites and related microporous materials are also

used as catalysts for the formation of carbon–carbon

[11]. Reproduced by permission of Imperial College Press, London.


bonds, like the very demanding alkylation of isobutane

with n-butene. The reaction pathway of the isobutane/

n-butene alkylation is quite complicated since several

competing reactions can take place besides the regular

alkylation, like self-alkylation, destructive alkylation,

multiple alkylation as well as oligomerisation and crack-ing. Due to the lower reaction temperature for the ali-

phate alkylation (thermodynamically favoured), the

desorption step is often difficult in these reactions, since

the reaction product is often more strongly adsorbed

than the reactants. However, the obtained product mix-

ture is an excellent blending component for gasoline,

and the reaction is industrially carried out applying

either HF or sulfuric acid as acid catalysts. There is anintense search looking for attractive alternatives for

those acids, with large pore zeolites (among other solid

Fig. 9. Mechanism of aliphate alkylation [4]. Repro

catalysts) as promising candidates, however, so far these

systems suffer from an unsuitable catalyst lifetime

[1,4,13–18].

The mechanism of the aliphate alkylation can be de-

scribed as follows: The reaction is initiated by the addi-

tion of a proton to the n-butene, forming the secondarybutyl-(2) cation, which abstracts a hydride ion from

isobutane forming tertiary butyl cations. These tertiary

butyl cations interact with n-butene forming isooctyl

cations (preferentially the high octane number repre-

senting trimethylpentanes). The isooctyl cations capture

hydride ions from isobutane forming isooctanes and

tertiary butyl cations, which then continue the reaction

cycle (see Fig. 9).The lifetime of the large pore zeolites (FAU, BEA

and EMT) is determined by the relative rates of hydride

duced by permission of Elsevier, Amsterdam.


transfer and n-butene addition. The latter is usually con-

trolled by using back-mixed reactors operating at

high conversions and low alkene concentrations. The

hydride transfer rate is depending on the stability of

the carbenium ion and the space limitations given by

the microporous framework of the zeolite. The slowerthe hydride transfer the more multiple alkylation takes

place, forming C12 and C16 units, which block the active

sites as well as the micropores and deactivate the cata-

lyst [1].

Typical side reactions are classical cracking and oli-

gomerisations over weak Brønsted acid sites, producing

larger olefins, which participate in alkylation and lead to

the formation of larger alkylate molecules, which againcontribute to increased deactivation of the catalyst [1].

The coke is formed through a combination of hydride

transfers, inter- and intra-molecular alkylation reactions

and oligomerisations leading to heavy, unsaturated cyc-

lic and acyclic compounds [4].

2.5. Shape selectivity of acid zeolites and related

microporous materials

In 1960, Weisz and Frilette [19] introduced the

expression ‘‘Shape-Selective Catalysis’’ by demonstrat-

ing that Ca A zeolite dehydrated 1-butanol at 260 �Cbut not isobutanol. This observation showed that the

conversion took place inside the microporous structure

of the Ca A zeolite (0.5 nm pore diameter), not available

for the branched isobutanol due to its large kineticdiameter. This size exclusion model has since then been

used to remove linear hydrocarbons from mixtures con-

taining both branched and linear hydrocarbons.

The mechanisms of molecular shape-selective cataly-

sis can be described and summarised as follows:

Reactant selectivity describes the phenomenon of

microporous catalysts acting as molecular sieves and

excluding bulky molecules from entering the intra-crys-talline void-structure while allowing smaller molecules

to enter. The critical exclusion limit can be varied over

a wide range of different zeolites and related micropo-

rous solids [4].

Product selectivity refers to discrete diffusivities of the

reaction products formed with respect to the micropo-

rous pore architecture and crystal size of the catalyst

particles. Sterically less hindered product moleculesmay easily leave the microporous framework, whereas

bulky product molecules may stay much longer in the

cavities of the zeolites. The term molecular traffic control

has been coined by Derouane and Gabelica in 1980

describing qualitatively the transport of molecules

with different shape and/or size in the microporous

framework of zeolites with two discrete sets of pores

[4,20].Restricted transition state-type selectivity occurs when

the spatial configuration around a transition state or a

reaction intermediate located in the intra-crystalline

volume is such that only certain configurations are pos-

sible. This means the formation of reaction intermedi-

ates and/or transition states is sterically limited due to

the shape and size of the microporous lattice allowing

the access of the species formed to interact with the ac-tive sites. This type of selectivity was first proposed by

Csicsery [21] and is usually connected to the suppres-

sion of undesired side reactions like coke formation.

Whereas the product selectivity depends on the crystal

size of the catalyst the restricted transition state-type

selectivity is not depending on the relative rates of dif-

fusion and reaction, hence both selectivities can easily

be distinguished by changing the crystal size of thecatalyst [22].

The different types of shape-selectivities of zeolites

and related microporous materials are summarised in

Fig. 10.

Zones et al. [23] introduced the term inverse shape-

selectivity for those cases where the restricted transition

state-type selectivity arises from a positive discrimina-

tion of specific transition states. An example is the skel-etal multiple branching of n-hexane in large pore zeolites

or related microporous solids, where the highest selectiv-

ity for multiple branched isohexanes was registered for

microporous solids with well defined pore diameters

and with an optimum interaction with the desirable

isomers [4]. See also Fig. 11.

The cage or window effect, observed in connection

with the hydrocracking of long n-alkanes, represents acertain case of molecular shape-selectivity in zeolites.

This effect is in operation when the diffusivities and/or

reactivities do not change monotonically within a

homologous series of compounds, due to the fact that

certain cracking products, which fit the cage dimensions

and are trapped in the cage of the zeolites, were not ob-

served as products [4]. Gorring reported for the first

time this effect in connection with the hydrocrackingof hexadecane using erionite [24]. Only small amounts

of C7–C9 alkanes were observed (although representing

the central cracking products of the probe molecule

and detected in large amounts when cracking hexade-

cane without using a shape-selective catalyst), since they

fit excellently within the cage of erionite [4].

Selective reactions at the pore mouth of zeolites have

been observed by Martens et al., for example the long-chain n-alkane isomerisation over Pt/H ZSM-22 (TON

structure) revealing large amounts of mono-branched

isomers although these isomers cannot desorb from

the narrow channels of this mono-dimensional zeolite

[25–28]. This observation has been termed pore mouth

catalysis since the product pattern is explained by

involving only the acid sites at the entrance of the small

pore zeolites such as Theta-1, ZSM-22 and erionite. Inaddition, the second branching of n-alkanes over Pt/H

ZSM-22 was shown to occur at approximately the

Fig. 11. Inverse shape selectivity observed for the skeletal multiple branching of n-hexane in large pore zeolites. AFI microporous solids represent the

optimum pore diameter [4]. Reproduced by permission of Elsevier, Amsterdam.

Fig. 10. Schematic representation of the three types of shape-selectivity. Concerning the restricted transition state-type selectivity: the lower

transition state molecule is easier to accommodate in the cavities than the upper one [1]. Reproduced by permission of Wiley-VCH, Weinheim.


distance between the pore openings at the surface of thezeolite crystallites [22]. This phenomenon has been

termed key-lock catalysis [26].

A schematic presentation of the latter type shape-selective effects in zeolites and related microporous sol-

ids is given in Fig. 12.

Fig. 12. Shape-selective environments in different zeolite structure types: (a) large molecules have access to interrupted cavities and channel

intersections for pore mouth catalysis; (b) molecules are plugged into the pore aperture; (c) molecules are converted in multiple pore mouths

according to key-lock catalysis; (d) molecules are converted in the intra-crystalline shape-selective environment [4]. Reproduced by permission of

Elsevier, Amsterdam.


3. Hydrocarbon conversion and formation, crude oil and

natural gas upgrading

3.1. World market of zeolites and related microporous

materials with focus on their application as catalysts

Zeolites and related microporous solids are impor-

tant components of adsorbents and catalysts used inconnection with the upgrading of crude oil and natu-

ral gas, among others due to the fact that they are

easily separated from the educts and products and

by representing a clean technology compared to their

predecessors.

The world market of zeolites and related micropo-

rous solids is still in a period of strong development.

Currently about 1.6 millions of tons are used per year,of which about 1.3 millions of tons refer to synthetic

zeolites and about 0.3 millions of tons to natural zeo-

lites, the latter mainly applied as adsorbent and ion ex-

changer [29,30].

Concerning the application of synthetic zeolites and

related microporous materials, the focus in terms of

amounts is definitely on detergent builders (1.05 millions

of tons per year), followed by catalysis (0.15 millions oftons per year) and finally adsorption (0.1 millions of

tons per year). A-type zeolites are by far the most com-

monly applied detergent builders. Furthermore, A-type

zeolites are mainly used with respect to the application

related to adsorption, separation and purification, which

covers, among others, insulating windows, purification

of olefins, natural gas as well as industrial gas, desicca-

tion of alcohols, separation of paraffins and xylenesand, finally, production of oxygen and hydrogen. X-type

zeolites are applied as adsorbents for the elimination of

trace amounts of polar impurities, whereas highly sili-

ceous mordenite and ZSM-5 are used for desiccation

of acid gases and the elimination of volatile organic

compounds [29,30].

Finally, almost all the zeolites and related micropo-

rous solids, which are used as catalysts, are applied in

the upgrading of oil and natural gas, that means oil refin-ing and petrochemicals. Within oil refining, the main

applications are fluidised catalytic cracking (FCC),

hydrocracking, C5/C6 isomerisation and dewaxing,

whereas in petrochemicals, the principal applications

are related to the different transformations of aroma-

tics (alkylation, transalkylation, isomerisation, . . .). TheY-type zeolite present in the FCC catalysts accounts

for almost 95% of the total world consumption of zeo-lites used within catalysis [30]. Table 1 shows an overview

of the zeolites and related microporous materials used as

catalysts in different modified forms on an industrial or

pre-industrial scale in connection with the corresponding

processes [30–33].

3.2. Crude oil and natural gas upgrading—present

scenario and the future

Handling all aspects of crude oil and natural gas

upgrading and their impact on corresponding catalyst

development requires an analysis of the current situation

as well as an evaluation of the major driving forces in

order to meet the future requests from an economic,

technological and environmental point of view.

Modern oil refineries use crude oil of various originsand they have to meet different market demands

Table 1

Overview of zeolites and related microporous materials used as catalysts in different modified forms on an industrial or pre-industrial scale in

connection with the corresponding processes [30,33]

Zeolite/microporous material Process or application technology

LTA (A-type zeolites) Detergent builder, separation, desiccation

FAU (X- and Y-type zeolites) Catalytic cracking, hydrocracking, separation, purification and desiccation, aromat alkylation

BEA (Beta zeolite) FCC additive, cumene and ethylbenzene production

MOR (Mordenite) Hydrocracking, hydroisomerisation, dewaxing, NOx reduction, adsorption, cumene synthesis,

transalkylation of aromatics

MWW (MCM-22) Ethylbenzene and cumene production

MFI (ZSM-5) Dewaxing, hydrocracking, ethylbenzene (Mobil-Badger) and styrene production,

xylene isomerisation, methanol to gasoline (MTG), benzene alkylation, adsorption,

catalytic aromatisation, FCC additive, toluene disproportionation

ERI (Erionite) Selectoforming, hydrocracking

LTL (KL-type zeolites) Catalytic aromatisation

CHA (SAPO-34) Methanol to olefins (MTO)

FER (Ferrierite) n-Butene skeletal isomerisation

TON (Theta-1, ZSM-22) Long-chain paraffin isomerisation

AEL (SAPO-11) Long-chain paraffin isomerisation


depending on which country or part of the world they

are serving. Consequently, the process facilities at differ-

ent refineries can vary, however many processes can be

implemented. The market demand is mainly related to

gasoline, diesel, kerosene and fuel oils, which have to

meet specifications which are requested by national gov-

ernments or the European Commission [34]. As an

example, the most advanced category of the car indus-try�s World-Wide Fuel Charter (WWFC) concerninggasoline specifications is summarised in Table 2 [35].

The quality of gasoline is usually defined by the mo-

tor octane number (MON, built up with isoparaffins and

ethers and reduced by the presence of alkenes) and the

research octane number (RON, obtained through the

presence of aromatics and ethers). Diesel is mainly char-

acterised by the cetane number, which is linked to thepresence of alkanes. The cetane number is decreased

by the presence of higher aromatics [34].

World-wide are about 40 million barrels (one barrel

corresponds to 159 l) crude oil refined very day in the

refineries. The basic processes for refining crude oil are

still the same but the tendency is in the direction of more

complex process technology. To begin with crude oil is

divided into various fractions by atmospheric distilla-

Table 2

Most advanced gasoline specifications proposed in the World-Wide

Fuel Charter (WWFC) [35]

Parameter

Low octane gasoline [(MON + RON)/2] 86.8

High octane gasoline [(MON + RON)/2] 93.0

Oxygen (wt.% max.) 2.7

Benzene (vol.%) 1.0

Aromatics (vol.%) 35.0

Olefins (vol.%) 10.0

Sulfur (ppm) 5–10

tion. During this procedure the main fractions of oil

products are obtained, which cover

Liquefied petroleum gas (LPG)
C1–C4 cut
Naphtha
C5 to about 180 �C Middle distillates 130–300 �C Diesel/gas oil 150–370 �C Lube base oils/atm. residue higher than 370 �C
The residue of the atmospheric distillation can be

used as feed for a vacuum distillation, leading to thefractions termed vacuum gas oil (VGO, 370–540 �C)and vacuum residue (higher than 540 �C).All these fractions are the primary oil products,

however, many of them have to be upgraded before

they can meet the requested specifications and be used

as commercial products [36]. In the case of large

amounts of gasoline to be produced, a FCC unit will

be installed (feed stocks: VGO, residues). Furthermore,polymerization of the light alkenes obtained or their

alkylation with isobutane is implemented in the refin-

ery. Isomerisation of C5/C6 n-paraffins may be installed,

if there is a need for an increase in the octane number

(getting branched paraffins). If there is a stronger need

for diesel oil a hydrocracker unit may be added, and if

the residue is too viscous to be handled, a visbreaker

unit may be requested. The refineries apply a lot of cat-alytic units in order to upgrade, to convert or to purify

their product streams, only visbreaking and coking are

thermal processes. Only a few catalytic units use liquid

catalysts, like the isobutane/n-butene alkylation (HF or

H2SO4), the conversion of mercaptans into disulfides

(Co-phthalocyanines) or alkene dimerisation (Ziegler–

Natta type catalysts). All other catalysts are solids,

and the processes applying zeolites or related micropo-rous materials will be presented and discussed in the

following sub-chapters [34]. However, a presentation

Fig. 13. Thermodynamic equilibrium for hexane isomerisation [36].



of a modern refinery is not complete without taking

into account the need of the petrochemical industry

which requests synthesis gas, olefins and aromatics

(benzene, toluene and xylenes—BTX) as basic feed

stocks. Olefins and aromatics are usually manufactured

from naphtha via steam cracking [36], however, theincreasing demand of propene can only be satisfied by

using other technologies like deep catalytic cracking

(DCC) and dehydrogenation of propane in addition

to the steam cracker production. Furthermore, a new

route has been introduced by Weitkamp et al. convert-

ing surplus aromatics from pyrolysis gasoline into high

value steam cracker feed (ethane, propane and n-bu-

tane) through a hydrodecyclisation step using a bi-func-tional Pd/H-ZSM-5 catalyst [37]. This technology

combines a considerable decrease of the requested ben-

zene gasoline content with an improved utilisation of

surplus aromatics by increasing the yields and selectiv-

ities of ethene and propene in the steam cracking pro-

cess and represents a real break-through with respect

to high value basic products within refinery and

petrochemistry.The stagnation of crude oil reserves and the increase

of their prices have recently driven the attention towards

the production of fuels and chemicals from natural gas

via synthesis gas (syngas). This route is also known as

the ‘‘gas to liquids (GTL)’’-technology. Fuels produc-

tion from syngas (in former times obtained from coal)

has been reported by Fischer and Tropsch in 1923 for

the first time [38], using an alkali-promoted iron cata-lyst. Fuels manufactured via the Fischer–Tropsch route

reveal an excellent quality since they consist mainly of

linear paraffins and a-olefins and do not contain sulfurand aromatics. A Co-containing catalyst is applied for

the production of heavy paraffins via the Fischer–Trop-

sch route starting with natural gas, a technology devel-

oped by Shell and named the ‘‘Shell Middle Distillate

Synthesis (SMDS)’’ route [39,40]. Finally, diesel fuel(or gasoline) is produced by hydrocracking of the more

or less sulfur and nitrogen-free wax obtained through

the SMDS process using noble metal containing zeolites.

The more restricted fuel specifications currently intro-

duced in order to reduce the environmental impact of

hazardous emissions represent a driving force with

respect to an increased use of fuels prepared via the

Fischer–Tropsch route as a blending component of thegasoline and diesel pools in the future [39]. Besides

the SMDS technology an alternative has been presented

by SASOL/Chevron termed as the ‘‘Slurry-Phase-Distil-

late’’ process, again based on the Fischer–Tropsch route

producing wax (using a Co-containing catalyst) fol-

lowed by a hydrocracking step in order to get diesel or

gasoline [41]. Finally, besides the dehydrogenation of

propane and the DCC process other alternatives haverecently been introduced with respect to meet the

increasing demand of propene, like the ‘‘Methanol to

Propene (MTP)’’ process of Lurgi applying H-ZSM-5

based catalysts [42].

The following sub-chapters will focus on the current

technology and future developments related to different

processes dealing with upgrading of crude oil and natu-

ral gas and applying zeolites or related microporoussolids as catalysts.

3.3. Isomerisation of n-paraffins

3.3.1. C5/C6 isomerisation (light straight

run—LSR—naphtha)

Environmental restrictions have caused the phase-out

of lead additives and the elimination (or lowering) ofbenzene from gasoline, with the consequence of an in-

creased demand of isoparaffins in order to improve the

octane number of gasoline. Therefore, isomerisation of

light straight run naphtha, containing C5/C6 n-paraffins,

has advanced to be an important process in the oil

refinery. Skeletal isomerisation of n-paraffins is an

acid-catalysed and equilibrium limited reaction, which

is thermodynamically favoured at lower temperatures,see Fig. 13 [39].

Industrially, the C5/C6-isomerisation is performed

using a bi-functional catalyst (noble metal together with

an acidic carrier) and in the presence of hydrogen. The

main advantage of applying a bi-functional catalyst is

that stable operations are possible under a sufficiently

high hydrogen pressure. The mechanism of this reaction

is well accepted, and can be summarised as follows (seealso Fig. 14):

1. The n-paraffins are initially dehydrogenated on the

noble metal sites to give the corresponding n-alkenes.

2. The n-alkenes are then protonated at the acid sites,

resulting in carbenium ions.

Fig. 14. Mechanism of the n-paraffin isomerisation [36]. Reproduced by permission of Elsevier, Amsterdam.


3. The carbenium ions undergo a skeletal isomerisation

to form more stable branched carbenium ions, fol-

lowing the protonated cyclopropane (PCP) interme-

diate mechanism suggested by Brouwer [43].

4. Finally, the branched carbenium ions are hydroge-

nated at the noble metal sites and desorb as isoparaf-

fins [39].

Fig. 14 suggests that isomerisation and cracking may

occur as parallel reactions in the presence of hydrogen,

in addition to sequential cracking of pre-isomerised

compounds and post-isomerised cracking products.

Since cracking and isomerisation are both catalysed by

similar acid catalysts, it is not surprising that cracking

takes place besides isomerisation. Skeletal n-paraffin

isomerisation requires at least a chain of four carbon

atoms, whereas cracking requests a minimum of seven

atoms in the carbon chain. That means that pentanes

and hexanes easily can be isomerised but not easily

cracked. For paraffins higher than hexanes, cracking is

usually a competing reaction to skeletal isomerisation,

resulting in a lower selectivity for the isomerisation reac-

tion, in spite of the fact that isomerisation is carried out

at lower temperatures than cracking [44]. Weitkampdemonstrated the different results on isomerisation and

cracking of C6–C10 n-paraffins using a bi-functional Pt/

Ca zeolite Y. See Fig. 15 [45].

Commercial isomerisation catalysts contain both a

noble metal (Pt) based hydrogenation-dehydrogenation

function and an acid function. The acid function is pro-

vided by either a halogenated (Cl, F) alumina carrier, a

sulfated zirconia substrate or by a zeolite (usually mord-

Fig. 15. Comparison of the hydroisomerisation and hydrocracking of

C6–C10 n-paraffins over Pt/Ca zeolite Y [45]. Reproduced by permis-

sion of Elsevier, Amsterdam.


enite (MOR)). A zeolite omega (MAZ structure type)

based catalyst has been demonstrated to be superior to

the mordenite based system (higher activity and better

selectivity), however, no commercial use has been re-

ported so far. The higher yield achieved by using a Pt/

H-MAZ catalyst has been reported to be due to both

the unique structural properties of MAZ and the higher

acid strength of its Brønsted sites [46]. Besides mordenite(MOR) and zeolite omega (MAZ), Pt/Beta zeolites

(BEA) have been investigated as isomerisation catalysts

as well, however, the Pt/H-MOR is the only system

which has been commercialised so far [39].

The halogenated alumina and sulfated zirconia based

isomerisation catalysts are more strongly acidic than the

zeolite based catalysts, which means that they can

isomerise the LSR naphtha at temperatures below150 �C. Consequently, this favours the formation ofthe desired isoparaffins due to the thermodynamic con-

ditions. On the other hand, the tolerance level of the

non-zeolitic catalysts against water and sulfur is not

high, and this leads to a fast deactivation of the haloge-

nated alumina and sulfated zirconia based isomerisation

catalysts and to the request of a severe pre-treatment of

the naphtha feed. Finally, a continuous stream of halo-gen has to be added in order to keep the catalyst active,

leading to corrosion problems for the reactor system

[39]. The zeolite based isomerisation catalysts are also

lacking sulfur tolerance, however, to a much lesser ex-

tent [46].

The Pt/H-MOR catalysts are less acidic than the

Pt/halogen–alumina catalysts, and, consequently, they

have to be applied at higher reaction temperatures

(about 250 �C), which limits the formation of isoparaf-fins due to the thermodynamic conditions. However,

they are more robust than the non-zeolitic catalystsand can withstand low levels of impurities such as sulfur

and water in the feed [36,39].

The most known example of a zeolitic isomerisation

catalyst is the Pt/H-MOR catalyst, first developed for

the Shell Hysomer process. The process operates at

27–30 bar hydrogen pressure and a reaction temperature

of 250 �C. At this temperature, not all normal paraffinscan be converted to branched paraffins (see also Fig. 13).Therefore, it seems to be attractive to combine the

Hysomer isomerisation process with the ISOSIV iso/

normal paraffin separation process (using a Ca A zeolite

as selective adsorbent for the n-paraffins), commercia-

lised by Union Carbide (now UOP). The combined tech-

nology is known as ‘‘total isomerisation process (TIP)’’

and commercialised by UOP (see Fig. 16). The octane

gain in the TIP process is reported to be in the rangeof about 10 octane numbers [46]. UOP�s zeolitic isomeri-sation catalyst (Pt loaded mordenite) is commercialised

under the trade name HS-10.

The stability of the Hysomer catalyst is also demon-

strated by its long lifetime: catalyst charges have been

used for up to seven years in commercial operation, how-

ever, a catalyst deactivated by operational mishaps can in

most cases be regenerated by a simple coke burn-off [44].An important parameter controlling the isomerisa-

tion activity and selectivity is the framework Si/Al ratio

of the zeolite. A maximum of activity for n-pentane

isomerisation was observed for a Si/Al ratio of about

10, for which all framework aluminium atoms are iso-

lated, and therefore, supporting the strongest frame-

work Brønsted acid site possible within the mordenite

structure [47–49]. Increasing the lattice Si/Al ratio bydealumination is beneficial as well with respect to cata-

lyst deactivation by reducing the coking rate [50,51].

Dealumination by acid leaching decreases the number

of Brønsted acid sites and creates mesoporosity inside

the zeolite crystallites, which again leads to a decrease

of the diffusion limitations. The mesoporosity causes

shorter residence times and facilitates easier desorption

of the products, avoiding secondary reactions of theintermediates formed and improving the overall selectiv-

ity [46]. Besides the final lattice Si/Al ratio, the method

of dealumination applied has an influence on the cata-

lyst activity: steam dealumination leaves the lattice

aluminium removed in extra-framework positions

(EFAl), whereas acid leaching results in almost EFAl-

free samples. In addition, acid leaching forms alumin-

ium gradients along the zeolite crystallites, whereassteam treatment produces a more uniform aluminium

distribution. The most active Pt/H-MOR has been

Fig. 16. Paraffin Total Isomerisation Process (TIP) [36]. Reproduced by permission of Elsevier, Amsterdam.


obtained by dealuminating the zeolite through acid

leaching followed by mild steaming, leading to small

and controlled amounts of EFAl which create a syner-

getic effect on the Brønsted acid sites associated toframework aluminium (FAl), improving their overall

acid strength [39].

CEPSA and Sud-Chemie AG have commercialised

their isomerisation catalyst based on a strongly acidic

Pt/H-mordenite catalyst under the trade name HYSO-

PAR [46].

3.3.2. Isomerisation of long-chain n-paraffins

There is a strong interest to extend the isomerisation

reaction to n-paraffins containing carbon chains longer

than C6, mainly with respect to produce higher octane

multi-branched isomers. However, as mentioned earlier,

one has to take into account that the cracking tendency

of branched paraffins increases with the length of the

hydrocarbon chain and with the degree of branching.

As an example, Pt/H-MOR produces low yields of iso-mers in connection with n-heptane isomerisation due

to extensive cracking of the isoheptanes formed [11,39].

Furthermore, isomerisation of long-chain n-paraffins

has been used to improve the pour point, viscosity,

cloud point and freeze point of middle distillates and

lube oils. In this respect very good catalytic perfor-

mances have been observed for the n-heptane isomerisa-

tion using Pt supported on nano-crystalline Beta zeolite[52]. This experimental result is explained by a combina-

tion of Brønsted acid sites of lower acid strength than in

mordenite and a faster diffusion of the branched isomers

through the small crystallites (10–20 nm) of the nano-

crystalline Beta zeolite, leading to a decrease in the

cracking rate [11].

In addition, besides high isomerisation selectivities

combined with low cracking rates, the degree of branch-

ing should be minimized in order to keep a high quality

of the paraffinic product. In this respect, Pt/SAPO-11 (amedium pore sized microporous solid) has been demon-

strated to display high isomerisation selectivity and low

yields to multi-branched species in the n-octane isomeri-

sation [53]. The suppression of the formation of multi-

branched isomers on Pt/SAPO-11 catalysts has been

interpreted in terms of a transition state shape-selective

effect induced by the uni-dimensional pore structure of

SAPO-11 [39].However, in the domain of middle paraffin isomerisa-

tion (C7–C9 carbon chain length) there is still a need for

catalyst improvement in order to improve a thorough

isomerisation selectivity (two branches or more) while

minimizing the cracking rate [30].

Isomerisation of higher alkanes in the wax range is

going to play a much more important role in the future

since moderately branched alkanes formed by isomer-ised/cracked wax represent excellent components for

lube oils. The isomerisation and cracking of synthetic

wax produced through the Fischer–Tropsch route repre-

sent an excellent alternative for the manufacture of high-

quality fuels which are currently derived from petroleum

[44].

3.3.3. Isomerisation of n-butane

Isomerisation of n-butane can take place following

two different mechanisms, either via a direct isomerisa-

tion involving the formation of a highly unstable

primary carbenium ion (mono-molecular) or via a

dimerisation-cracking mechanism (bi-molecular). The

isomerisation of n-butane is even more thermodynami-


cally limited than the C5/C6 n-paraffin isomerisation.

When running the reaction at low temperatures, very

strong acid sites are requested. In fact, commercial n-

butane isomerisation processes make use of HCl/AlCl3(Phillips [54], Shell [55]) or Pt supported on chlorided

alumina as in the Butamer technology by UOP [56].Alternative solid acid catalysts—more beneficial with re-

spect to the environment—have been studied, among

others sulfated zirconia, heteropolyacids (in particular

12-tungstophosphoric acid) and zeolites. Sulfated zirco-

nia, however, deactivates very fast. Pt/H-mordenite is

active for n-butane isomerisation, however, this catalyst

requires higher reaction temperatures than sulfated zir-

conia or heteropolyacids in order to achieve reasonableconversions. In addition, the zeolite based catalyst forms

larger amounts of different by-products than the non-

zeolitic catalysts [11,39].

The present n-paraffin isomerisation capacity world

wide is roughly split equally between n-butane isomeri-

sation, LSR naphtha isomerisation using Pt/H-MOR

and LSR naphtha isomerisation over Pt/Cl/alumina

[44,57].

3.4. Skeletal isomerisation of light n-alkenes

C4 and C5 alkene skeletal isomerisation has been re-

garded as a suitable alternative for increased production

of isobutene and isopentenes. These isoalkenes are

mainly obtained from FCC units or steam crackers

and are used for the production of methyl tert-butyl

Fig. 17. Reaction mechanism for the acid catalysed skeletal isomerisation o

London.

and tert-amyl methyl ethers (MTBE and TAME), which

represent excellent fuel oxygenates with good octane

blending properties, in spite of the fact that MTBE

has been questioned as fuel additive for environmental

reasons [58]. In addition, isobutene is also an important

reactant for the petrochemical industry [39].C4 and C5 alkene skeletal isomerisation is an acid

catalysed reaction requesting strong acid sites. The reac-

tion mechanism can be described by the initial double

bond cis–trans isomerisation taking place at the acid

sites before skeletal isomerisation. The protonation of

the double bond leads to the formation of a secondary

carbenium ion, which then rearranges into a protonated

cyclopropane (PCP) structure, and finally ending upwith an isoalkene via the formation of an unstable pri-

mary carbenium ion in the case of n-butene, as shown

in Fig. 17 [11].

For thermodynamic reasons, low isomerisation tem-

peratures should be applied, since the equilibrium con-

centration of branched olefins decreases with increasing

temperature. However, at low temperatures the selectiv-

ity to isoalkenes decreases due to the competing olefinoligomerisation reactions. The extent of these side reac-

tions can be lowered by applying higher reaction temper-

atures and low alkene partial pressure, however, at

higher temperatures other non-desired reactions take

place, like cracking, hydrogen transfer and coking, lead-

ing to the catalyst deactivation [59].

In former times different solid acid catalysts have been

applied as light alkene isomerisation catalysts, like metal

f n-butenes [11]. Reproduced by permission of Imperial College Press,


halides and supported phosphoric acid. Later, environ-

mentally more friendly zeolite catalysts have been intro-

duced successfully, with medium-pore size zeolites as the

most promising systems since the competing (bi-molecu-

lar) oligomerisations can be lowered within the limited

space available in the smaller channels of these micropo-rous solids. Among the medium pore sized zeolites those

having a one-dimensional channel structure (like ZSM-

22 and Theta-1, both having the TON structure, as well

as ZSM-23) turned out to be the best catalysts. They gave

the highest yield of isobutene by working at relatively

low temperatures (below 400 �C) and with a low olefinpartial pressure. The selectivity to isobutene has been

shown to improve be decreasing the density of theBrønsted acid sites (increasing Si/Al ratio) and by

decreasing their acid strength by replacing lattice alumin-

ium by other trivalent cations, like gallium and iron. This

behaviour is explained by suppression of bi-molecular

reactions leading to side-products as the acid sites be-

come more and more isolated, and can be regarded as

indirect evidence that isobutene is mainly formed via a

mono-molecular mechanism, as shown in Fig. 17 [11,39].Finally, a ferrierite based catalyst has been intro-

duced by Shell, giving high yields of isobutene at

350 �C and long catalyst lifetimes [60]. This behaviourhas been attributed to the particular structure of ferrie-

rite possessing intersecting 10- and 8-membered ring

channels, which induce the selective formation of tri-

methyl pentene dimers and their cracking into C4 frag-

ments, including isobutene [61].In conclusion, quite active and selective alkene isom-

erisation catalysts can be prepared from microporous

solids by combining the proper pore architecture with

the presence of isolated and/or mild Brønsted acid sites

[39].

3.5. Aliphate alkylation

Aliphate alkylation in connection with oil refinery re-

fers to a technology dealing with the partial conversion

Fig. 18. Conversion of a liquid isobutane/1-butene mixture on a CeY zeolite.

ratio = 11, pressure: 31 bar) [63]. Reproduced by permission of Wiley-VCH,

of the C4 cut (isobutane and n-butenes) into the so-

called alkylate. This alkylate represents a very high qual-

ity and valuable component for the refinery gasoline

pool due to the high octane numbers of the formed iso-

octanes, preferentially represented by the trimethyl-

pentanes. Current alkylation technology applies eitherhydrofluoric (UOP and Phillips) or sulfuric acid (Stracto

and Kellogg). These traditional processes suffer from

several safety risks and drawbacks, like the high toxicity,

volatility and corrosiveness of hydrofluoric acid and the

high catalyst consumption of sulfuric acid, which re-

quires a regeneration plant close to the alkylation facil-

ity. Therefore, replacement of the existing alkylation

processes by new technology based on non-toxic, non-corrosive and environmentally friendly solid acid cata-

lysts is one of the most important research challenges

in the field of heterogeneous catalysis. Significant re-

search during the last two decades resulted in a variety

of different solid acids capable to produce an alkylate

with the same quality features as the conventionally

manufactured product. However, none of the alternative

solid catalysts have been commercially applied so far,mainly due to the short catalyst lifetime caused by the

decline of their hydride transfer activity [14,39].

The mechanism of aliphate alkylation has been de-

scribed in Section 2.4 for the case of isobutane/n-butene

alkylation. Due to typical side reactions like oligomeri-

sation and classical cracking the question arises as to

how long the formed product can be considered as an

alkylate. Arbitrarily, Weitkamp et al. placed this limitat a content of alkanes in the C8 product fraction of

90 mol.% [62]. In the example shown in Fig. 18, the alkyl-

ation stage then ends after about half an hour. This illus-

trates that the time-on-stream (TOS) behaviour of the

solid catalysts is still unsatisfactory—due to the decline

of the hydride transfer activity. In addition, the ratio be-

tween the trimethylpentanes (TMP) and the dimethylhex-

anes (DMH) formed can be taken as a measure of thealkylation/oligomerisation ratio for a certain solid acid

catalyst [11]. However, future research within this sub-

Composition of the C8 fraction (temperature 80 �C, isobutane/1-buteneWeinheim.


field of acid catalysis should concentrate on the reasons

for this loss of hydride transfer activity and measures to

extend the solid catalyst lifetime considerably [14].

In the following, the most recent advances related to

the use of solid acid catalysts for the aliphate alkylation

will be summarised, with focus on the zeolites. An excel-lent review covering this sub-field of heterogeneous

catalysis has recently been written by Weitkamp and

Traa, including the early studies with respect to the

search for solid alkylation catalysts [14].

Large-pore zeolites, like FAU, EMT and BEA, have

been in the focus concerning their application as ali-

phate alkylation catalysts. Extensive studies of the

behaviour of rare earth exchanged faujasites (RE-Y)were performed by Weitkamp, where he combined a

sampling system downstream of the reactor with high-

resolution GC in order to obtain detailed results on

the olefin conversion and product yields as a function

of time on stream [63,64]. With Ce–Y as catalyst he ob-

tained 100% butene conversion and a high quality alkyl-

ate during the first half hour on stream, after that the

activity decreased rapidly and the selectivity changedfrom the alkylate to the oligomerisate (C8 and C12 ole-

fins) as the zeolite became deactivated [62].

Corma et al. found a maximum initial conversion of

2-butene for USY zeolites with different unit cell sizes

(which means different framework compositions) cover-

ing a0 values between 2.435 and 2.450 nm [15]. The

TMP/DMH ratio continuously increased with the unit

cell parameter and the hydride transfer activity washigher for the mildly dealuminated samples [39]. In addi-

tion, the extra-framework aluminium (EFAl) formed

during the steam dealumination influenced the catalytic

performance of USY zeolites in the aliphate alkylation

as well [65].

Cardona et al. investigated the aliphate alkylation

using an USY zeolite at 50 �C, concluding with theobservation that their reaction pathway is in line withthe mechanism discussed in Section 2.4 [66]. Gardos

et al. applied rare earth exchanged Y zeolites in the ali-

phate alkylation using a batch-type autoclave and reac-

tion temperatures between 50 and 100 �C, arriving at theconclusion that the composition of the alkylate is

entirely controlled by the kinetics [67,68].

EMT, the hexagonal faujasite has been investigated

as aliphate alkylation catalyst by Stocker et al., usinga stirred-tank reactor and a reaction temperature of

80 �C [69–71]. La-H-EMT with a La3+ exchange degreeof 40% was observed to reveal the best alkylation perfor-

mance, followed by H-EMT and H-FAU and Ce Y zeo-

lite. The better alkylation performance of EMT was

ascribed to a higher strength of the Brønsted acid sites

and slightly larger cages in EMT as compared to the

FAU zeolite [39].Zeolite Beta (BEA), another large-pore zeolite with a

three-dimensional pore structure, has been used as cata-

lyst for the aliphate alkylation [72–76]. As for Y-type

zeolites, zeolite Beta deactivated rapidly as well [72].

Corma et al. demonstrated in their more in-depth stud-

ies that the performance of Beta as aliphate alkylation

catalyst depends on the synthesis recipe, the crystallite

size, the chemical composition, the post-preparationtreatments, the nature and amount of the extra-frame-

work Al and the density and strength of the Brønsted

acid sites [73,74]. As an example, it was concluded from

these studies that H-Beta zeolite prepared from tetraeth-

ylorthosilicate (TEOS) was more active than that syn-

thesised from amorphous silica, and that samples with

crystal sizes of 0.35 lm were more active than those of0.1 lm [39,73,74]. Finally, Kiricsi et al. [75] and Flegoet al. [76] studied La H-Beta zeolites and the nature of

the carbonaceous deposits formed after adsorption and

conversion of isobutene/1-butene mixtures.

Twelve-membered ring zeolites other than FAU,

EMT and BEA have only found limited attention as ali-

phate alkylation catalysts, as for example, ZSM-4 (zeo-

lite Omega), ZSM-20 (inter-growth between FAU and

EMT), ZSM-3, ZSM-18 and mordenite [14].Medium pore zeolites, like ZSM-5 and ZSM-11, have

also been investigated as aliphate alkylation catalysts,

however, these zeolites were found to be active for the

alkylation only at temperatures higher than 100 �C,which is not of interest from a thermodynamic point

of view [77,78]. In addition, these medium pore zeolites

produced less trimethylpentanes, indicating serious pore

restrictions for the formation of the desired alkylate [39].MCM-22 has been investigated, however, its behaviour

as aliphate alkylation catalyst was found to be in be-

tween those of 10– and 12-MR systems [72,79].

Non-zeolitic systems, like sulfated metal oxides, het-

eropoly acids, supported Lewis and Brønsted acids

and resins have been studied as aliphate alkylation cat-

alysts as well, however, a presentation of these systems

would be outside the scope of this overview. Finally, thischapter should not be finished before the current process

developments at the pilot plant stage are summarised

concerning the use of alternative solid aliphate alkyl-

ation catalysts, however, as far as this is known, no zeo-

lite based systems are among these catalysts [14]:

1. Trifluoromethanesulfonic acid on a porous carrier

(Haldor Topsøe A/S).2. Antimonypentafluoride on acid-washed silica (Chem-

ical Research & Licensing Co., Chevron Corp.).

3. Proprietary catalysts (UOP and Catalytica Inc, Neste

Oy, Conoco Inc.).

3.6. Catalytic reforming

Besides catalytic cracking, catalytic reforming is one

of the most important processes within a modern

Fig. 19. Schematic diagram of a typical FCC unit [39]. Reproduced by

permission of Wiley-VCH, Weinheim.


refinery, where low-octane naphtha components (nor-

mal- and cycloalkanes) are converted into high-octane

isoalkanes and aromatics suitable for gasoline produc-

tion. The catalytic reforming process consists of a num-

ber of hydrogenation and dehydrogenation reactions, in

addition to isomerisation, cyclisation and cracking withrespect to the formation of isoalkanes and aromatics.

The process is run at 425–525 �C with hydrogen pres-sures in the range of 0.5–3.0 MPa [39].

Conventional reforming catalysts are based on Pt/Cl

and Pt–Sn supported on alumina, however, the focus

here will be concentrated on zeolite based reforming cat-

alysts. The non-acidic Linde Type L zeolite (LTL, with

potassium as counter ions) as support for Pt and Bahas been investigated by several groups [80–82]. The

presence of highly dispersed Pt clusters inside the zeolite

channels and the shape-selective effects imposed by the

mono-directional framework structure of the KL-zeolite

is responsible for the very good aromatisation perfor-

mance of this zeolite, which has a pore diameter of

0.71 nm. One of the main drawbacks of this zeolite is

the high sensitivity towards sulfur poisoning, responsi-ble for the fast catalyst deactivation. The preparation

method and especially the Pt incorporation have a

strong influence of the catalyst performance with respect

to activity and stability [39].

Other zeolite based reforming catalysts have been

investigated as well, like Pt/ZSM-12, Pt/Beta and sul-

fided Pt/Cs-Beta [83,84]. Finally, large-pore boro-silicate

based zeolites (B-Beta, B-SSZ-33, B-SSZ-24 and B-SSZ-31) have been patented by Chevron as reforming cata-

lyst [85].

3.7. Catalytic cracking

The catalytic cracking unit is the most important con-

version facility in a modern refinery. This process con-

sists of the scission of the hydrocarbon C–C bondspresent in the feedstock (usually vacuum gas oils or res-

idues) in order to obtain gasoline, light alkenes or other

low molecular hydrocarbons. A number of different

FCC catalysts exist and catalyst changes in the world-

wide about 350 refinery FCC units are made often,

depending on the feedstock type and quality available

[86]. This process, which produces about 30% of the

total gasoline pool either directly or indirectly, is veryflexible with respect to different combinations of process

design and catalysts. This flexibility allows the refiners

to process a large variety of feedstocks and to adapt

the product pattern to the changing market demands

with respect to local fuel specifications and environmen-

tal legislation [39].

The history of catalytic cracking started in the 1920

when Eugene Houdry (the father of catalytic cracking)used an acid treated natural clay as catalyst to convert

hydrocarbons into lower molecular weight products. A

significant change in this business occurred in 1942 with

the introduction of the FCC technology. A schematic

diagram of a typical FCC unit is shown in Fig. 19 [39].

The FCC process can briefly be summarised as fol-

lows: The pre-heated feedstock is contacted with the

hot catalyst coming from the regenerator at the bottomof the riser reactor, where most of the cracking reactions

take place at temperatures around 500 �C and contacttimes with the catalyst of about two to three seconds.

The cracking products are hydrocarbons, which are ex-

tracted from the catalyst pores in the stripper unit using

steam, and then passed to the regenerator to restore the

catalyst activity by burning off the coke formed during

the cracking reactions at temperatures of about 700 �C[39]. Part of the used catalyst is continuously replaced

by fresh catalyst, which results in a consumption of

about 10 tons catalyst per day for a medium sized

FCC unit. The term ‘‘fluidised’’ refers to the catalyst

particles in the range of 60–90 lm consisting of porousmicro-, meso- and macrospheres, which are fluidised,

e.g., intimately admixed in a stream of vaporised hydro-

carbon feedstock and steam. Since the cracking reac-tions are primarily endothermic, heat balance with the

exothermic regeneration reaction is required for the riser

to operate at appropriate cracking temperatures. Thus,

the continued operation of an FCC unit depends on

the heat balance of the riser reactor and regenerator [86].

With the introduction of zeolite (faujasite type) con-

taining cracking catalysts in 1962, replacing the amor-

phous silica–alumina, a tremendous change concerning

Fig. 20. Conceptual pore architecture design of a FCC catalyst [88].



FCC technology took place. Zeolite containing catalysts

are much more active, show higher gasoline yield perfor-

mances and produce less coke than the amorphous

silica–alumina based catalysts, resulting in higher feed-

stock conversions and severities as well as enhanced eco-

nomic benefits of the process [39,86]. The actual FCCtechnology involves the formulation of proprietary multi-

functional cracking catalysts, consisting of different

amorphous (catalytically active macroporous matrix,

clay type binder) and crystalline acid functions (repre-

sented by shape-selective (microporous) zeolites like

Y-type zeolite containing mesopores due to dealumina-

tion forming the ultra-stable Y zeolite—USY), and a

series of additives for metal passivation (mainly V andNi), sulfur removal, promoters for total combustion

and octane enhancing additives [11]. The two main com-

ponents of cracking catalysts are the zeolite Y and the

matrix. The matrix plays a critical role in the selective

cracking of the (high molecular) bottoms fractions when

residue containing feedstocks are processed. The main

functions of the matrix are to pre-crack large molecules

and adsorb Ni and V preferentially in order to protectthe zeolite Y of the catalyst particle. In an ideal situa-

tion, the pre-cracked large molecules from the matrix

macropores are further cracked in the mesopores of

the USY (to, i.e., gas oil fractions) before, finally, gaso-

line is formed in the micropores of Y-type zeolite (or

propene in the case of ZSM-5), see also Fig. 20 [87,88].

Concerning metal passivation, both vanadium and

nickel deposit on the cracking catalyst as their hostmolecules are converted to lighter products and coke.

Both are extremely deleterious when present in excess

of 3000 ppm on the FCC catalyst. Vanadium in the oxi-

dation state 5+ is converted to vanadic acid and reacts

with the zeolite framework by hydrolysing the zeolite

lattice structure, and, thus deactivating the zeolite part

Fig. 21. Trends in catalytic cracking catalyst performance [3

of the catalyst, whereas nickel forms a metal/metal oxide

site on the catalyst surface causing formation of coke

and hydrogen [86,87].

FCC catalysts are nowadays mainly produced byfour companies: Grace Davison, Akzo Chemicals,

CCIC and Engelhard Corporation. The high activity

and relatively low coke formation of zeolite containing

FCC catalysts enabled the reactor technology to ad-

vance from dense fluidised beds to short-contact-time

(SCT) risers with a corresponding improvement in the

performance [36], see also Fig. 21.

Concerning the mechanism of catalytic cracking, theacid sites of the zeolite component are regarded as the

catalytically active sites, and the mechanism has been

discussed already in Section 2.3. Most studies in flui-

dised catalytic cracking have focused on zeolite Y, as

this is still the dominant zeolite used in FCC. Besides

the acid properties of this zeolite, the unique pore archi-

tecture of Y zeolite is ideal for cracking gas oil compo-

nents into gasoline molecules. Moreover, it has beenobserved that the activity of the Y zeolite for gas oil

cracking has a maximum for a Si/Al ratio of 5–8, which

6] Reproduced by permission of Elsevier, Amsterdam.


corresponds to a unit cell size (UCS), a0, of 2.436–

2.440 nm. This clearly indicates that gas oil cracking

requires the presence of strong Brønsted acid sites.

Unfortunately, so far it has not been possible to prepare

Y zeolite with a framework Si/Al ratio above 4 by direct

synthesis. Therefore, highly dealuminated zeolites (lowa0 values) have to be prepared by dealumination of com-

mercially prepared Y zeolite samples with Si/Al ratios in

the range of 2.6. However, it was found that the catalytic

behaviour of dealuminated Y zeolites with given a0 val-

ues strongly depend on the method of dealumination ap-

plied [89]. In this respect, Y zeolites dealuminated by

steaming (USY) create a secondary porosity formed

during the partial destruction of the zeolite frameworkand forming mesopores which facilitate diffusion of lar-

ger molecules into the zeolitic channels. The obtained

USY type zeolites show, in addition, a much better

hydrothermal stability, which is a pre-requisite of the

application as FCC catalyst (cf. regeneration conditions

of the FCC catalysts) [39,86].

In certain cases the oil companies would like to in-

crease the amount of lighter components, like propene,n-butenes and isobutene, as they are important feed-

stocks for the petrochemical industry. ZSM-5 turned

out to be excellent in this respect, especially for the

enhancement of propene. This zeolite can significantly

improve the octane number of gasoline in catalytic

cracking. Addition of a few per cent ZSM-5 to a conven-

tional FCC catalyst gives an equivalent octane number

increase [90]. Due to the pore architecture, ZSM-5 in-creases the octane number of the gasoline by selectively

upgrading low octane gasoline components into lower

molecular weight compounds with a higher octane num-

ber [11].

The concept of using ZSM-5 as co-catalyst to modify

the performance of a generic FCC catalyst system can

significantly increase the product flexibility in the FCC

unit. An extension of this technology is the so-calledDeep Catalytic Cracking process (DCC), which involves

the use of ZSM-5 as the primary catalyst rather than the

co-catalyst in a FCC type of moving bed reactor system

in order to maximise lower olefins production (primarily

propene and a-olefins) with gasoline as a by-product[36,91]. SINOPEC has commercialised this technology

in China and a DCC plant is under construction in Thai-

land [92]. A modified DCC process has been offered bySINOPEC, termed as Catalytic Pyrolysis Process (CPP),

in which vacuum gas oils and atmospheric residues are

converted to ethene and propene [93].

Other zeolites, like Beta and MCM-22, have been

investigated with respect to catalytic cracking. They

have shown some use in modifying FCC reactions, how-

ever, none of them have managed to balance activity

and product selectivity as well as Y zeolite [94].Since only 60% of the worldwide demand of propene

can be produced using steam cracking technology, the

remaining amount must be manufactured either by

FCC/DCC or propane dehydrogenation. One future

challenge within catalytic cracking will focus on the en-

hanced production of light olefins by development of the

current technology however, with the continued growth

of residue processing in FCC units, further attentionmust be paid to the performance of catalysts for this

variant of FCC operation as well. Since both the zeolite

and the matrix play an important role in the optimal

performance of a residue catalytic cracking catalyst the

following properties have to be considered in detail with

respect to catalyst development [36]:

1. structure of the mesopores for the bottomsconversion,

2. unit cell size (UCS) for the coke make,

3. resistance of the zeolite structure to vanadium attack

and

4. defect structure with respect to the hydrothermal

stability.

3.8. Catalytic hydrocracking

Catalytic hydrocracking is an oil refinery process

which has been developed with respect to the conversion

of relatively heavy oil feedstocks (including residues)

into lighter transportation fuel products through C–C

bond scission. In contrast to catalytic cracking, hydro-

cracking is performed in the presence of hydrogen asco-feed at relatively high pressures (50–200 bar) and

at lower temperatures than catalytic cracking (300–

450 �C). However, catalytic hydrocracking accountsfor a more hydrogenated product than catalytic crack-

ing, due to the hydrogenation reactions taking place

under these conditions. In addition, the coke formation

rate and the gas yield are considerably lower in catalytic

hydrocracking compared to catalytic cracking [39,95].Catalytic hydrocracking is a very flexible technology,

allowing a wide range of feedstocks to be processed.

Whereas in the US the hydrocracking units are mainly

used to convert lighter feedstocks (straight run light

and heavy gas oils, coker gas oils, FCC cycle oils and

thermally cracked gas oils) into gasoline components,

the hydrocrackers outside the US produce a much wider

spectrum of products, like kerosene, jet fuel and dieselfuel (also called middle distillates) from vacuum gas oils

(VGO) processing (see also Table 3) [39,95].

Currently, the annual catalytic hydrocracking capa-

city world-wide amounts to about 200 million tons, dis-

tributed over around 120 catalytic hydrocracking units,

which mainly have been developed by UOP/Unocal

(Unicracking process), Chevron (Isocracking and Iso-

max technology), Shell and IFP [95,96]. Catalytic hydro-cracking processes can be performed according to three

main technologies: The single stage configuration, where

Table 3

Hydrocracking feedstocks and products [39]

Feedstock Product

Straight run gas oils (SRGO) LPG

Vacuum gas oils (VGO) Gasoline

FCC cycle oils Catalytic reforming feeds

Coker gas oils Jet fuels

Thermally cracked gas oils Diesel fuels

Deasphalted oils Heating oils

Straight run and cracked naphthas Olefin plant feedstocks

Lube oils

FCC feedstocks


the feed is processed in a single catalyst bed in one or

two reactors in series, recycling the un-converted feed.

Another configuration employs two catalysts either in

the same reactor (‘‘stacked beds’’) or in two reactors

in series. The first catalyst is the hydrotreating catalyst

(removal of sulfur and nitrogen), which partially hydro-

genates aromatics as well. Catalytic hydrocracking is

then performed on the second catalyst. The two stageconfiguration consists of two reactors, where the first

one contains the hydrotreating catalyst and the second

one the hydrocracking catalyst [39,97]. Finally, a more

modern and cost-effective process is the series-flow con-

figuration, with no product separation in between the

hydrotreating and hydrocracking steps, and, thus, re-

Fig. 22. Basic catalytic hydrocracking reactions [39]. Rep

quires very robust second stage catalysts such as those

based on zeolites [36].

In general, with increasing feedstock heaviness, the

amount of catalyst poisons (metals, aromatic coke pre-

cursors, sulfur and nitrogen) increases. Metals cause

irreversible deactivation of the first stage hydrotreatingcatalyst, while organic nitrogen compounds specifically

reduce the cracking activity of the acidic second stage

catalysts [95].

The main reactions occurring during catalytic hydro-

cracking are summarised in Fig. 22.

Catalysts used in the first stage for feedstock pre-

treatment are usually hydrotreating catalysts (Co/

Mo, Ni/Mo or Ni/W supported on alumina or silica–alumina). The real hydrocracking catalysts are bi-

functional systems, consisting of a hydrogenation/

dehydrogenation and an acidic cracking function. The

activity and selectivity of the hydrocracking catalyst de-

pends on the ratio between the hydrogenation/dehydro-

genation and acid functions as well as on the strength of

both. For example, Lemberton et al. reported the exis-

tence of an optimum hydrogenation/acid ratio in Ni/Mo/zeolite Y/alumina catalysts, resulting in a reduced

amount of coke formation with increasing intimacy of

mixing of the two functions at the submicron level

[98,99]. Amorphous silica–alumina (ASA) is still being

used as acidic carrier in some hydrocracking units [97].

roduced by permission of Wiley-VCH, Weinheim.


However, the incorporation of zeolites into the cata-

lyst formulation during the 1960s represents a real

break-through with respect to an enhanced catalytic per-

formance of hydrocracking catalysts [100]. The zeolite

based hydrocracking catalysts show a higher activity

(due to a higher acid strength), higher thermal andhydrothermal stability, better resistance to sulfur and

nitrogen containing catalyst poisons as well as lower

coking rate, thus prolonging the catalyst lifetime

[101,102].

The first generation of Y zeolites (either in their

hydrogen form or exchanged with rare-earth cations)

used in catalytic hydrocracking contained a high

amount of lattice aluminium, which resulted in a highunit cell size. However, the use of ultrastable Y zeolites

(USY) synthesised by steam dealumination made it pos-

sible to control not only the number and strength of acid

sites but also the amount of extra-framework aluminium

as well as the degree of mesoporosity or secondary

porosity. As the severity of the steaming increases, the

density of acid sites decreases along with the unit cell

size. Hydrocracking catalysts synthesised from USYzeolites with low unit cell sizes (a0 < 2.445 nm) produce

less gas but higher liquid yields and are more selective

towards middle distillates. Furthermore, acid-leached

dealuminated Y zeolites are more active and selective to-

wards middle distillates than parent USY. The presence

of secondary porosity in USY zeolites (due to steam

dealumination) as well as the macroporosity of the for-

Fig. 23. Creation of a secondary pore structure in Y zeolite as a result of dea

London.

mulated hydrocracking catalyst play an important role

with respect to processing heavy fractions and residues.

The mesoporosity or secondary porosity has been

shown to reduce the mass transfer limitations during

hydrocracking and, thus, suppress secondary cracking

(see also Fig. 23) [39,96,103].Another solution with respect to improvement of the

reactant accessibility and decrease of secondary cracking

reactions has relied on the preparation of small Y zeolite

crystallites (<0.5 lm). Furthermore, the preparation ofmesoporous Al-MCM-41 materials and the application

of their Ni/Mo derivatives with respect to hydrocracking

of vacuum gas oil revealed a good selectivity for middle

distillates and demonstrated a higher activity comparedwith Ni/Mo supported on amorphous silica–alumina

[104,105].

Besides Y zeolite, other large pore zeolites or related

microporous and mesoporous materials, such as Omega,

L zeolite, Beta, VPI-5, mordenite, UTD-1, MCM-48

and SBA-15, have been investigated as components for

hydrocracking catalysts, however, these systems have

not yet led to commercial application [39,96,106].Recent developments within hydrocracking catalysis

have been concentrated on the improvements in the

amorphous silica–alumina and Y zeolite base materials

and the control of catalyst pore architecture. Although

zeolite based hydrocracking catalysts offer the best pros-

pects for heavy feeds conversion, their activity is sup-

pressed due to the restricted access of the heavier

lumination [103]. Reproduced by permission of Imperial College Press,


molecules to the zeolite pores. Therefore, the main

focus concerning further research will be concentrated

on finding a zeolitic catalyst component dedicated to

the production of middle distillates, associated with the

activity, stability and proper pore architecture of the

final hydrocracking catalyst.

3.9. Catalytic dewaxing

Heavier oils, like gas oils, diesel oils or lube oils, con-

tain often wax (long-chain normal and/or slightly

branched C18 and higher paraffins), which crystallise at

low temperatures (below 100 �C). This affects their vis-cosity deleteriously, as quantified by pour point determi-nation [107]. These compounds are therefore often

removed, either by physical processes such as extraction

in solvent dewaxing, or nowadays by catalytic dewaxing

using shape selective catalysts. During catalytic dewax-

ing (carried out at about 400 �C), normal and slightlybranched long-chain paraffins are removed by selective

cracking to lighter products including gas. The basic

process resembles selectoforming, where C5–C9 normalparaffins are cracked to LPG. The selective removal of

the long-chain normal paraffins in gas oil dewaxing

using ZSM-5 as catalyst is illustrated in Fig. 24, where

the sharp peaks of the normal paraffins in the gas chro-

matograph of the feed are absent in the product after

catalytic dewaxing [108].

Fig. 24. Catalytic dewaxing of gas oil using shape selective ZSM-5 as catalys

chromatograph after catalytic dewaxing, RT means retention time). Reprod

Catalytic dewaxing has been connected to the med-

ium pore zeolites and related microporous solids, includ-

ing ZSM-5, ZSM-11, ZSM-23 and SAPO-11 [109]. BP

introduced the first generation dewaxing catalysts,

which was based on mordenite in order to remove nor-

mal alkanes from lube oils. However, Mobil�s discoveryof ZSM-5 forced the development of the next generation

of dewaxing catalysts, applicable to both gas oils

(MDDW process) and lube oils (MLDW process). The

new generation of ExxonMobil catalysts (MSDW: Mo-

bil Selective DeWaxing) selectively converts linear long-

chain paraffins to their corresponding isoparaffins

instead of cracking to lower molecules. Studies in this

direction involved zeolite Beta and MCM-22 as catalystcomponents. Chevron commercialised their Isodewax-

ing process, which is based on isomerisation rather than

cracking as well, using SAPO-11 as dewaxing catalyst

[110]. Finally, AKZO-FINA, Criterion/Zeolyst/Lyon-

dell and others offer catalytic dewaxing technology and

catalysts as well [36,95,109].

In conclusion, the most recent developments within

catalytic dewaxing have been concentrated on dewaxingby means of shape selective isomerisation rather than

cracking. Examples of catalysts which exhibit shape

selective isomerisation properties include SAPO-11

and zeolite Beta. The SAPO-11 catalyst has the advan-

tages of both low cracking activity and good isom-

erisation activity while suppressing the formation of

t [108] (top: gas chromatograph before catalytic dewaxing, bottom: gas

uced by permission of Elsevier, Amsterdam.


multi-branched isomers which, if formed, would easily

crack to undesired lighter products [95].

3.10. Transformation of aromatics/petrochemicals derived

from aromatics

3.10.1. Alkylation of aromatics

Aromatics alkylation represents world wide a large

scale industrial process for the production of intermedi-

ates, fine chemicals and petrochemicals. In principle, the

chemical reaction consists in the replacement of a hydro-

gen atom of an aromatic compound by an alkyl group.

An acid catalyst is required when the replaced hydrogen

is on the aromatic ring (electrophilic substitution), how-ever, base catalysts or radical conditions are required if

the hydrogen on the side chain of an aromatic com-

pound is substituted. Important products manufactured

by aromatics alkylation include ethylbenzene (further

converted to styrene for polymer production), cumene

(isopropylbenzene, an intermediate for the production

of phenol and acetone), alkylnaphthalenes (precursors

to advanced polymers) and alkylbenzene sulfonates(detergent builders). Acid catalysts applied for aromat-

ics alkylation are usually Brønsted or Friedel-Crafts

acids and cover mineral acids, metal halides, cation ex-

change resins, acidic oxides and zeolites. Especially the

last group of catalysts is well suited for the specific pro-

duction of dedicated alkylated aromatic compounds,

including single isomers of those compounds, due to

their discrete pore architecture and Brønsted acidity[111,112].

Ethylbenzene (EB) is mainly used as a precursor for

styrene monomer. About 90% of the world wide ethyl-

benzene production is based on alkylation of benzene

with ethylene, with AlCl3 as the primary Friedel-Crafts

catalyst. Due to the corrosive nature of this catalyst,

alternatively, solid acid catalysts have been developed,

like ZSM-5 for the Mobil-Badger vapour phase forma-tion of ethylbenzene, commercialised in the 1970s and

operating at 380–450 �C and 20–30 bar pressure. Highyields of ethylbenzene (more than 99%) can be achieved

and the catalyst deactivation is slow (bi-molecular hy-

dride transfer is largely suppressed due to steric hin-

drance), leading to long lifetimes applying suitable

regeneration procedures. About 35 plants are operating

world-wide using this technology, with an annualproduction capacity of nearly eight million tons. Alter-

natively, liquid phase alkylation of benzene with

ethylene, using zeolitic catalyst systems have been com-

mercialised, like ExxonMobil�s EBMAX technology

applying MCM-22 as catalyst [111,112].

Cumene (isopropylbenzene) is manufactured by

alkylation of benzene with propylene, mainly still using

the solid phosphoric acid (SPA) technology (UOP). Zeo-lites have been investigated extensively as alternative

solid catalysts for the cumene production, including

ZSM-5, mordenite, Beta and ZSM-12. However, so far

only a few technologies have been commercialised using

highly dealuminated mordenite (Dow Chemicals), Beta

(EniChem) or a proprietary catalyst system (ExxonMo-

bil) [111,112].

Gas phase ethylation of toluene with boric or phos-phoric acid modified ZSM-5 yields para-ethyltoluene

with up to 100% isomeric purity, which is an important

precursor for manufacturing para-methylstyrene (ad-

vanced polymer production). ExxonMobil and Deltech

Corp. have commercialised this technology [113].

Cymene (isopropyltoluene) is manufactured commer-

cially by alkylation of toluene with propene using solid

phosphoric acid catalysts (Sumitomo). Cymene is animportant intermediate in the production of meta-cresol,

and zeolitic materials have been tried as catalysts as well.

Flockhart et al. have applied Y zeolite to perform this

reaction [114].

Side-chain alkylation of toluene with methanol has

been performed using Cs-exchanged X zeolite as base

catalyst in order to produce ethylbenzene and, subse-

quently, styrene [115]. This technology is close to becommercialised [93].

Alkylation of naphthalene with propene over zeolite

catalysts yields mainly 2,6-diisopropylnaphthalene

[116], whereas the alkylation of naphthalene with meth-

anol can either yield preferentially 1-methylnaphthalene

(ZSM-12, 12-MR system, kinetic control) or 2-methyl-

naphthalene (dealuminated mordenite or ZSM-5) [111].

3.10.2. Isomerisation and transalkylation

of alkylaromatics

The transfer of alkyl groups between aromatic mole-

cules, also termed transalkylation, and the intra-mole-

cular isomerisation are commercially applied in large

scale. Both reactions are acid catalysed processes, and

most of the catalysts used are solid systems, with zeolites

as the most prominent representatives [117].The pyrolysis gasoline from naphtha crackers and the

naphtha reformate consist mainly of xylenes and ethyl-

benzene with respect to their C8 aromatics fractions.

They are isolated from these streams by distillation

and solvent extraction. Xylenes are used on a large scale

industrially, and they are precursors for a number of

important petrochemicals. Especially para-xylene has

an enormous market potential (terephthalic acid pro-duction for polyester formation), with an annual in-

crease of about 7% [93,118].

When C8 aromatics fractions are processed with re-

spect to xylene isomerisation, the remaining ethylben-

zene must be converted to xylenes. Catalysts used for

the xylene and ethylbenzene isomerisation contain al-

ways platinum, in former times mainly on chlorinated

aluminas or steamed silica–aluminas (Octafining processdeveloped by Atlantic Refining Corp.), nowadays pref-

erentially on mordenite. The process is conducted in a


fixed-bed reactor at temperatures between 370 and

430 �C and a pressure range of 7–15 bar. New genera-tions of zeolite based catalysts have been introduced

by UOP (Isomar process) and IFP (Oparis), claiming

even higher yields of para-xylene. Metal modified

H-ZSM-5 catalysts are most effective when the xyleneisomerisation is accompanied by ethylbenzene dealkyla-

tion to benzene and ethene under similar process condi-

tions (ExxonMobil) [118].

Disproportionation is a special case of transalkyla-

tion and is in place when alkyl groups are transferred

between identical molecules. Several processes are

commercialised dealing with toluene disproportionation

using mordenite or ZSM-5 as catalysts and arriving atthermodynamic mixtures of xylenes (ExxonMobil,

UOP and TotalFinaElf). Selective toluene dispropor-

tionation processes are on the market as well, using

ZSM-5 catalysts and arriving at para-xylene rich prod-

uct mixtures with selectivities of more than 80% (Eni-

Chem, ExxonMobil and UOP) [117,118].

The disproportionation of ethylbenzene to benzene

and diethylbenzenes has been studied extensively[119,120], but has not been commercialised until re-

cently. With large pore zeolites (Y zeolite), the reaction

occurs via a hydride transfer chain reaction through

diphenylethanes as intermediates. Medium pore zeolites

(ZSM-5) cannot accommodate this bulky intermediate,

and ethylbenzene disproportionation proceeds via a

dealkylation–realkylation path. This technology is com-

mercialised using a modified ZSM-5 catalyst [117]. TheCatalysis Commission of the International Zeolite

Association (IZA) has recommended the disproportion-

ation of ethylbenzene using LaNaY zeolite as a stan-

dard reaction for acidity characterisation of acid

zeolites [121].

3.10.3. Conversion of surplus aromatics

In connection with the introduction of the EuropeanAuto Oil Programme, the aromatics content of gasoline

has to be reduced currently from 43 to 35 vol.% in 2005.

This number is also expressed by the most advanced gas-

oline specifications proposed in the World-Wide Fuel

Charter (WWFC, see Table 2) [35]. As a consequence,

there will be a surplus of aromatics. The main source

of aromatics is the so-called pyrolysis gasoline, a by-

product rich in aromatics from the manufacture of eth-ene and propene by steam-cracking of straight run

naphtha (light hydrocarbons). Due to the predicted

world-wide increasing demand of ethene and propene,

the surplus of pyrolysis gasoline will increase further

[122]. A novel catalytic process for hydrogenative ring

opening of aromatics has been introduced, which allows

the conversion of pyrolysis gasoline from naphtha

steam-crackers into a high quality synthetic steamcracker feed composed of C2–C4 n-alkanes [37,123].

There are two process variants, the direct conversion

of aromatics with hydrogen on bi-functional zeolites

and the two-stage process comprising a conventional

ring hydrogenation to cycloalkanes followed by ring

opening of these compounds using acidic zeolites. A

large part of this research has been carried out in the

laboratories of J. Weitkamp at the University of Stutt-gart (Germany) [37,123–127].

The main advantage of the direct route is the fact that

one single reactor is sufficient to perform the desired cat-

alytic conversion, which is carried out using shape selec-

tive bi-functional zeolites, like Pd/H-ZSM-5 (Si/Al ratio

of about 20). For example, ring hydrogenation (6 MPa

hydrogen pressure) of toluene to methylcyclohexane

and skeletal isomerisation of the cycloalkane into ethyl-cyclopentane and dimethylcyclopentanes occurs at tem-

peratures up to about 250 �C, however, ring openingwith respect to the desired formation of C2þ n-alkanes

(73% of the total synthetic steam cracker feed) requires

temperature of up to 400 �C [124]. C2þ n-alkanes refersin this context to ethane, propane and n-butane.

Advantages of the two stage route are an easier re-

moval of the heat generated in the two exothermic stepsand the possibility to optimise the reaction conditions of

ring hydrogenation (performed by conventional technol-

ogy using metal catalysts) and ring opening (carried out

by novel shape selective zeolites in their acid form as cat-

alysts). Applying H-ZSM-5 with a Si/Al ratio of 20,

yields of C2þ n-alkanes, comparable to those obtained

during the direct conversion of toluene on Pd/H-ZSM-

5, are achieved when processing methylcyclohexane at400 �C as second step of the two stage route. ZSM-5and ZSM-11 turned out to be the best suited catalysts

for this technology, whereas large pore zeolites (like

Beta or Y zeolite) deactivated rapidly. Zeolites with

too narrow pores, like ZSM-35 or ZSM-22, did not

show high degrees of ring opening capacities, probably

due to the hindered diffusion of the hydrocarbons in-

volved [124].

3.11. Aromatisation of light alkanes

Aromatics (BTX) are precursor compounds for the

petrochemical industry of large importance. Since the

liquified petroleum gas (LPG) fractions in the refinery

are supposed to increase due to more severe operations

of the FCC units, there is a need to convert the lowvalue LPG into aromatics.

The aromatisation of light alkanes is easier to per-

form as the size of the alkane increases, taking into

the thermodynamic relations. For example, propane

and higher alkanes can be aromatised at temperatures

lower than 500 �C, whereas temperatures of up to575 �C are requested for the aromatisation of ethane.In addition, the proportion of benzene in the BTX frac-tion tends to decrease with increasing size of the alkane

[11,39,128].


The mechanism of the light alkane aromatisation in-

volves the initial dehydrogenation of paraffins to form

lower olefins and hydrogen, followed by olefin oligomer-

isation and cyclisation reactions to produce aromatics,

in which hydrogen transfer plays an important role

[36], see also Fig. 25.Several zeolite based industrial processes have been

introduced with respect to the aromatisation of light

alkanes. The most known process is BP/UOP�s CY-CLAR process using C3/C4 refinery gases as feedstock

and Ga/H-ZSM-5 as catalyst (1–5 wt.% Ga, Si/Al ratio

between 15 and 30). However, other technologies are

known, including Mobil�s M2 Forming process (usingH-ZSM-5), Chevron�s AROMAX process applyingC6–C8 alkanes as feed and Pt/L zeolite as catalyst as well

as Selectoforming over erionite [128]. An advantage for

the CYCLAR process is the fact that significant

amounts of hydrogen are formed, which is needed in

the refinery and which makes this process economically

very attractive [11].

The preferred catalyst for alkane aromatisation is Ga/

H-ZSM-5. The incorporation of Ga can be performedeither by direct synthesis, impregnation or ion exchange.

The success of the modified H-ZSM-5 catalyst is con-

nected to the limited coke formation by using this

zeolite. Ga modification improves this behaviour even

more [128]. Besides gallium, other metals such as Zn,

Pt, Ni and Ag have been used in combination with

ZSM-5 for the light alkane aromatisation, however,

the Ga- and Zn/H-ZSM-5 modified catalysts are the pre-ferred systems. Finally, Pt/KL zeolite has been applied

in the selective formation of benzene from hexane.

There is a strong interest in the aromatisation of eth-

ane since this compound is an important component of

Fig. 25. Simplified reaction mechanism for the aromatisation of propane

Imperial College Press, London.

both refinery and natural gases. However, from a ther-

modynamic point of view, it is very difficult to convert

ethane into aromatic hydrocarbons, since a highly

unstable primary carbenium ion will be formed in the

acid catalysed oligomerisation step, whereas for higher

alkanes a more stable secondary carbenium ion isformed. Anyway, several investigations have been re-

ported using Pt, Pd, Ga or Zn modified H-ZSM-5 zeo-

lites and temperatures of about 575 �C [39,128–130].

3.12. Natural gas upgrading/gas conversion technologies

The evolution of the known crude oil and natural gas

reserves world-wide indicates a dramatic increase in thelatter compared to a levelling off concerning the crude

oil. This trend is expected to continue, which will—in

addition to the price development with respect to the

crude oil based upgrading—most likely generate a grad-

ual shift towards the application of natural gas as a feed-

stock for the production of fuels and petrochemicals.

This situation has forced an enhanced global interest

in processes, which can convert natural gas into liquidsand higher added value products—without going via

methanol as intermediate. This route is known as the

‘‘gas to liquids (GTL)’’-technology, based on the

Fischer–Tropsch route. The interest to manufacture

fuels and petrochemicals from natural gas is driven by

the desire to apply this technology directly, for example

at remote natural gas field sites, in order to minimize

transportation costs and gas burning at the recoverysites [36].

The following sub-chapters will deal with the

‘‘GTL’’-technology, based on the Fischer–Tropsch

route as well as the methanol to hydrocarbon conver-

over modified H-ZSM-5 catalysts [11]. Reproduced by permission of


sions, where zeolites and related microporous materials

have been demonstrated to be superior catalysts.

3.12.1. ‘‘Gas to Liquids (GTL)’’/conversion of synthesis

gas to fuel

Fuels production directly from syngas (in formertimes obtained from coal) has been reported by Fischer

and Tropsch in 1923 for the first time [38], using an al-

kali-promoted iron catalyst. Fuels manufactured via the

Fischer–Tropsch route reveal an excellent quality since

they consist mainly of linear paraffins and a-olefinsand do not contain sulfur and aromatics. A Co-contain-

ing catalyst is applied for the production of heavy paraf-

fins via the Fischer–Tropsch route starting with naturalgas, a technology developed by Shell and named the

‘‘Shell Middle Distillate Synthesis (SMDS)’’ route

[39,40]. In addition, diesel fuel (or gasoline) is produced

by hydrocracking of the more or less sulfur and nitro-

gen-free wax obtained through the SMDS process using

noble metal containing zeolites. The more restricted fuel

specifications currently introduced in order to reduce the

environmental impact of hazardous emissions representa driving force with respect to an increased use of fuels

prepared via the Fischer–Tropsch route as a blending

component of the gasoline and diesel pools in the future

[39].

Besides the SMDS technology an alternative has been

presented by SASOL/Chevron termed as the ‘‘Slurry-

Phase-Distillate’’ process, again based on the Fischer–

Tropsch route producing wax (using a Co-containingcatalyst) followed by a hydrocracking step in order to

get diesel or gasoline [41].

The methanol to gasoline (MTG) plant in New

Zealand has been combined with a methane steam

reforming unit for production of synthesis gas and a

methanol plant to produce gasoline from natural gas.

The process economics can be improved considerably

by a clever combination and close integration of the dif-ferent steps. In the TIGAS process developed by Haldor

Topsøe AS for the manufacture of gasoline in a pilot

plant scale, the methanol synthesis and the MTG reac-

tions are integrated—without the separation of metha-

nol as an intermediate product. A multi-functional

catalyst has been developed, however, these process

technologies do not usually apply catalysts based on

zeolites or related microporous materials [36]. Finally,ExxonMobil has introduced the so-called ‘‘Advanced

Gas Conversion for the 21st Century’’ (AGC-21) tech-

nology, again based on the Fischer–Tropsch route [131].

3.12.2. Methanol to hydrocarbons

Besides the direct route from synthesis gas to hydro-

carbons (GTL-technology, based on the Fischer–Trop-

sch route), a strong focus has been concentrated onthe indirect route via the production of methanol from

synthesis gas (mixture of carbon monoxide and hydro-

gen), which is made by steam reforming of natural gas

or gasification of coal, and the consecutive formation

of hydrocarbons. Of coarse, these technologies are alter-

natives to the chemical conversion of methane, either via

direct coupling, which is thermodynamically not favour-

able, or via oxidative coupling, a route not successfullyso far from an industrial point of view.

To begin with, the methanol–hydrocarbons technol-

ogy was primarily regarded as a powerful method to

convert coal into high-octane gasoline. This concept

has been expanded since, not only with respect to the

formation of other fuels, but also to chemicals in gen-

eral. Of coarse, light olefins are important components

for the petrochemical industry, and the demand ofhigh-quality gasoline is increasing as well. In fact, with

this new technology, one can make almost anything

out of coal or natural gas that can be made out of crude

oil.

Methanol is converted into an equilibrium mixture

of methanol, dimethylether and water, which can be

processed catalytically to either gasoline (methanol to

gasoline, MTG) or olefins (methanol to olefins, MTO),depending on the catalyst and/or the process operation

conditions (see Fig. 26).

Although methanol itself is a potential motor fuel or

can be blended with gasoline, it would require large

investments to overcome the technical problems con-

nected with it. The commercial MTG reaction runs at

temperatures around 400 �C at a methanol pressure ofseveral bars and uses a ZSM-5 catalyst. These are theoptimal conditions for converting the olefins that form

within the catalyst into paraffins and aromatics. How-

ever, at one point in the MTG reaction, the product

mixture consists of about 40% light olefins. The impor-

tance of light olefins as intermediates in the conversion

of methanol to gasoline was recognised early. Conse-

quently, a number of attempts were made to selectively

form light olefins from methanol, not only on medium-pore zeolites but also on small-pore zeolites, SAPO type

molecular sieves and over large-pore zeolites (however,

to a much lesser extent). If one interrupted the reaction

at the point of about 40% light olefin formation, one

could harvest these C2–C4 olefins. By adjusting the reac-

tion conditions (such as, for example, raising the tem-

perature to 500 �C) as well as the catalyst applied, onecan increase dramatically the olefin yield. This discoveryled to the development of the MTO process, which gen-

erates mostly propene and butenes, with high-octane

gasoline as a by-product. However, the catalyst can be

modified in such a way that even more ethene is pro-

duced [132].

1973 marked the beginning of the energy crisis, and

new interest in synfuels and other chemicals favoured

the continuation of the methanol–hydrocarbon research[133,134]. Already the MTG and MTO processes repre-

sent a sort of chemical factory, to be brought on stream

Fig. 26. Methanol to hydrocarbons reaction path [132]. Reproduced by permission of Elsevier, Amsterdam.


as the technological and/or economic demands arise.

One can go a step further and convert the olefins to

an entire spectrum of products, through another ZSM-

5 based process: Mobil�s olefin–gasoline and distillateprocess (MOGD), originally developed as a refinery pro-

cess, which works well coupled with the MTO process.In the MOGD reaction, ZSM-5 oligomerises light ole-

fins, from either refinery streams or MTO, into higher-

molecular-weight olefins that fall into the gasoline,

distillate and lubricant range (see also Fig. 27) [135].

In 1979 the New Zealand government selected the

MTG process over the Fischer–Tropsch (SASOL) pro-

cess for converting natural gas from their extensive

Maui field to gasoline. At that time, Mobil�s fixed-bedMTG process was unproven commercially, whereas

the SASOL technology was already commercialised

[135]. The New Zealand plant started to produce about

600000 ton per year gasoline from April 1986, supplying

one-third of the nation�s gasoline demand [133]. Thegasoline production part of the factory was later closed

down, due to the price available for gasoline vs. the price

Fig. 27. Gasoline and distillate production via methanol and Mobil�s ZSM-5

of methanol, however, the methanol production part is

still in operation.

In the scale-up to commercial operation of the MTG

plant two factors were of certain interest, catalyst deac-

tivation and heat production, respectively (the MTG

reaction is highly exothermic). Despite the relativelyhigh stability and selectivity of the ZSM-5 catalyst

(due to the high silica content and the catalytic effective-

ness based on the unique pore structure), deactivation

during the MTG operation is pronounced. The catalyst

is progressively coked and must be regenerated by calci-

nation in air. The heat production in the MTG process

is high, leading to an adiabatic temperature rise of some

650 �C [36,136,137].The MTO technology seems now to be ready for

commercial use. The MTO process of Mobil has been

demonstrated in the same experimental 4000 ton per

year plant at Wesseling (Germany) used to prove the

fluid-bed MTG process, applying ZSM-5 as catalyst

[133]. The fluid-bed technology provided all advantages

in terms of increased product yield, better quality and a

technology [132]. Reproduced by permission of Elsevier, Amsterdam.


very efficient heat recovery. Depending on the world

market situation (price, demand, etc.) the fluidised-bed

technology is available to produce liquid fuels via meth-

anol [138].

UOP has, in co-operation with Norsk Hydro, an-

nounced in 1996 their SAPO-34 based MTO processto be realised for construction of a 250000 ton per year

plant using a natural gas feedstock for production of

ethene. A 0.5 ton per year demonstration unit operated

by Norsk Hydro has verified the olefin yields and cata-

lyst performance. SAPO-34 is extremely selective to-

wards ethene and propene formation (about 80%) with

the flexibility of altering the ratio between the two ole-

fins by varying the reactor conditions. The MTO processcan be designed for an ethene to propene ratio between

0.75 and 1.5—at nearly complete methanol conversion.

The high selectivity to ethene gives SAPO-34 (which

has the chabazite (CHA) structure) a significant advan-

tage over other types of catalyst systems, like ZSM-5 or

SSZ-13 (synthetic aluminosilicate with CHA structure).

In addition, the SAPO-34 has a significantly better sta-

bility due to a lower rate of coke formation than theother catalytic systems with comparable acid site densi-

ties. The need to remove the high exothermic heat of the

MTO reaction as well as the need for frequent regener-

ation led to a fluidised-bed reactor and regenerator

design. UOP and Norsk Hydro have commercially

manufactured the MTO catalyst (MTO-100TM), based

on SAPO-34, which has shown the type of attrition

resistance and stability suitable to handle multipleregeneration steps and fluidised-bed conditions [139,

140].

The discovery of the MTG reaction happened by

accident. One group at Mobil was trying to convert

methanol to other oxygen-containing compounds over

a ZSM-5 catalyst. Instead, they received unwanted

hydrocarbons. Somewhat later, another Mobil group,

working independently, was trying to alkylate isobutanewith methanol over ZSM-5 and identified a mixture of

paraffins and aromatics boiling in the gasoline range—

all coming from methanol. Although the discovery of

MTG was accidental, it occurred due to a balanced ef-

fort in catalysis over many years. The MTO reaction

seems to benefit from this development, although inde-

pendent research has been performed since. The evolu-

tion of the methanol–hydrocarbons technology, fromits discovery until its realisation on a demonstration

and/or commercial scale, has been accompanied by

extensive research related to the basic question of the

mechanism of formation of the initial C–C bond. To ad-

dress this topic in detail would be out of scope for this

paper, however, the interested reader can follow this dis-

cussion in the reviews mentioned in the list of Refs.

[132–137].Finally, an alternative to produce propene from

methanol has been introduced by Lurgi, the so-called

‘‘Methanol-to-Propene (MTP)’’ process, applying an

H-ZSM-5 based catalyst from Sud-Chemie AG

[42,131].

4. Catalytic probe reactions with respect to the porearchitecture characterisation of zeolites and related

porous materials

4.1. General remarks

Different types of probe molecules have been used to

investigate various properties of zeolites and related

microporous and mesoporous materials, for example,the pore architecture with respect to internal channel

dimensions, possible network connectivities, contribu-

tions of the external surface to the overall behaviour

of the materials, accessibility of pockets and caves in

connection with the structure of those compounds. Fur-

thermore, probe molecules are used to evaluate the sur-

face hydrophobicity and hydrophilicity as well as to

determine the acidity and basicity of porous systems.These type of investigations are done either by phys-

ico-chemical studies of the interaction of probe mole-

cules with the surface of the porous material (for

example by adsorption of the probe molecule and mon-

itoring the interaction by applying spectroscopic, calori-

metric, volumetric, gravimetric or other methods) or by

using probe molecules as model substrates in catalytic

reactions. Since the pores of zeolites and related micro-porous materials have dimensions comparable to those

of actual probe molecules, the phenomenon of shape-

selectivity in catalytic reactions can be observed. How-

ever, in order to take advantage of this, the catalytic

reaction must take place at active sites on the internal

surface and not on the external one. Therefore, the

probe molecules applied are usually designed with re-

spect to evaluate the influence of both the internal andexternal surface activities on the properties under inves-

tigation, like for example, a mixture of linear and

branched/bulky molecules. Branched and bulky probe

molecules are usually incapable to penetrate the micro-

pores whereas linear molecules can.

Concerning the selection of suitable probe molecules

one should note that the crystallographic pore diameters

do not represent suitable threshold molecular diameters,since molecules about 20% larger than this diameter can

be accommodated, especially at elevated temperatures.

Furthermore, the type of activity under investigation

(acid catalysed reaction, hydrogenation, oxidation etc.)

will influence the choice of the probe molecules as well.

In addition, framework aluminium tends to reduce the

pore volume and to broaden the pore size distribution.

Finally, pre-adsorption of polar molecules like ammoniaor water can reduce the apparent pore size and, there-

fore, certain attention should be paid with respect to


the hydration state of the zeolite or related porous mate-

rials [141,142].

Catalytic probe reactions for studying the pore archi-

tecture of zeolites and related porous materials have

much in common with adsorption tests with the same

objective. Of course, adsorption takes place during thesecatalytic tests as well. However, a shape-selective chem-

ical reaction occurs in addition. The basic understanding

of shape-selective catalysis is, therefore, an important

item with respect to the design and appropriate interpre-

tation of the results of catalytic test reactions.

This review deals with the topic of catalysis by zeo-

lites. Consequently, the focus within this chapter will

be concentrated on catalytic reactions using probemolecules. However, the interested reader will find excel-

lent reviews in the literature addressing the physico-

chemical studies of the interaction of probe molecules

with the surface of porous materials as well [141–145].

A number of catalytic test reactions for the pore

architecture characterisation of zeolites and related por-

ous materials have been introduced so far, and the

majority of the published data refer to Brønsted acidcatalysis. In the following sub-chapters these test reac-

tions will be presented and discussed, including recent

advances within catalytic test reactions using probe

molecules.

4.2. Constraint index (CI)

The constraint index (CI) introduced by researchersfrom the Mobil Oil Company more than 20 years ago,

has been the first technique with respect to the charac-

terisation of the relative pore width and shape-selective

properties of zeolites using a catalytic test reaction

[146]. This method is based on the competitive cracking

of an equimolar mixture of n-hexane and 3-methylpen-

tane on acid zeolites. Originally, the test reaction was

designed to distinguish between small-, medium- andlarge-pore zeolites, composed of 8-, 10- and 12-mem-

bered ring systems, respectively [147]. Based on the

shape-selective effect, the CI was defined as the ratio

of first-order rate constants (k) of the cracking of n-hex-

ane and 3-methyl-pentane:

Constraint index ðCIÞ¼ kðn-hexaneÞ=kð3-methylpentaneÞ

As long as the catalyst pores are sufficiently spacious,

branched alkanes are cracked at higher rates than their

linear isomers. In the literature, the following experi-

mental conditions for the determination of the con-

straint index have been given: reaction temperature

between 290 and 510 �C, LHSV between 0.1 and

1 h�1, 10 vol.% of each reactant in He as carrier gas,

catalyst mass of 1 g, fixed bed reactor at atmo-spheric pressure and an overall conversion of 10–60%

[141,146].

A higher constraint index arises from the preferential

cracking of n-hexane compared with the branched iso-

mer. The 3-methylpentane would be easier to crack in

the absence of steric hindrance. According to the Mobil

researchers, the constraint index can be used to classify

the molecular sieves into small-, medium- and large-porezeolites and related microporous materials [146]:

Small-pore (8-MR) systems
12 < CI
Medium-pore (10-MR) systems
1 < CI < 12
Large-pore (12-MR) systems
CI < 1
The attempt in the original paper by Frilette et al.

[146] was to provide a guideline for the determination

of small-, medium- and large-pore zeolites. However,

the discovery of new zeolite structure types has lead to

CI values which are misleading to incorrect conclusionsabout the pore size and structure. Among those exam-

ples are zeolites with 14-membered ring systems and

pore openings larger than 8 A as well as zeolites with

large internal cavities and pores composed of 8- or 9-

membered ring systems [147].

In conclusion, the introduction of the constraint

index represents the first example of probing the pore

architecture of zeolites and related microporous materi-als by applying a catalytic test reaction. In spite of a

number of disadvantages, the constraint index has been

extensively used, and a number of literature data are

available [141,146,147].

4.3. Modified or refined constraint index

Whereas the constraint index acts as a test reactionfor mono-functional acidic molecular sieves, completely

different reaction mechanisms apply concerning test

reactions dedicated to bi-functional zeolites and related

microporous materials. The search for an appropriate

test reaction covering the use of bi-functional molecular

sieves has been concentrated on the isomerisation and

hydrocracking of long-chain n-alkanes. The common

feature of these reactions is the fact that they are per-formed using hydrogen, which is activated by the noble

metal component of the catalyst. Therefore, the isomeri-

sation of n-decane at low conversions has been used to

define the modified or refined constraint index CI* for

the characterisation of bi-functional zeolites [148]:

Modified or refined constraint index ðCI�Þ¼ Yield2-methylnonane=Yield5-methylnonane

The modified or refined constraint index CI* is now well

approved concerning the pore width characterisation of

medium-pore zeolites, that means 10-membered ring

molecular sieve systems. However, the nature and exact

origin of the shape-selective effects on which the index is

based is not completely understood yet. Since 2-methyl-

Fig. 28. Modified or refined constraint index (CI*) for different zeolites

[141] (data taken from Refs. [149–151]). Reproduced by permission of

Wiley-VCH, Weinheim.

Fig. 29. Spaciousness index (SI) for different zeolites and related

microporous materials [141] (data taken from Ref. [155]). Reproduced

by permission of Wiley-VCH, Weinheim.


nonane is the kinetically preferred isomer in 10-mem-

bered ring zeolites, the amount of 2-methylnonane

formed from n-decane at low conversions increases rela-tive to the other methylnonanes with decreasing pore

width of the zeolite.

Fig. 28 summarises CI* values taken from the litera-

ture [149–151]. The CI* values for 10-membered ring

zeolites cover a broad range from about 3 to 15, which

represents a range where the CI* is quite suitable.

The only feature which the CI and CI* values have in

common is that their numerical values increase withdecreasing pore size of the zeolites under investigation.

On the other hand, the modified or refined constraint

index CI* is of little use concerning the pore width char-

acterisation of 12-membered ring zeolites or related

microporous materials. However, there is another index

based on a different test reaction which is complemen-

tary to the CI* value, the so-called spaciousness index

(SI) for the characterisation of 12-membered zeolites[141].

4.4. Spaciousness index

The mechanisms of hydrocracking and isomerisation

of cycloalkanes and n-alkanes are essentially identical.

However, different carbon number distributions have

been observed for the hydrocracking of n-alkanes andcycloalkanes, respectively, both of them consisting of

10 carbon atoms. Weitkamp et al. registered during

hydrocracking of C10 cycloalkanes the formation of iso-

butane and methylcyclopentane almost exclusively in

the absence of spatial constraints. The carbocation

intermediates which govern the selectivity of hydro-

cracking of C10 cycloalkanes seem to be perfectly suited

for investigating the space available—covering the entirerange of 12-membered ring zeolites and related micro-

porous materials [141,152,153].

The spaciousness index (SI) is defined as the yield

ratio of isobutane and n-butane in the hydrocracked

products of a C10 cycloalkane (butylcyclohexane or pen-

tylcyclopentane) [154,155]:

Spaciousness index ðSIÞ ¼ Yieldisobutane=Yieldn-butaneSI values of different zeolites and related microporous

materials are given in Fig. 29. There is no doubt that

the spaciousness index is quite suitable for evaluation

of 12-membered ring microporous materials, extending

a broad range of SI values from about 3 to 20. On the

other hand, 10-membered ring zeolites have all together

a spaciousness index of about 1, indicating that the SIsystem is not appropriate for the pore width evaluation

of medium-pore zeolites.

In conclusion, the spaciousness index is the method

of choice with respect to the pore width characterisation

of large-pore zeolites and related microporous materials

[141].

4.5. Disproportionation of ethylbenzene

As discussed in Section 3.10.2, the disproportionation

of ethylbenzene to benzene and diethylbenzenes has

been studied extensively [119,120]: With large pore zeo-

lites (like Y zeolite), the reaction occurs via a hydride

transfer chain reaction through diphenylethanes as

intermediates. However, medium pore zeolites (like

ZSM-5) cannot accommodate this bulky intermediate,and ethylbenzene disproportionation proceeds via a

dealkylation–realkylation path. To begin with, this reac-

tion has been proposed in order to receive information

about the number of strong Brønsted acid sites in zeo-

lites, however, later the reaction was found to be suit-

able for monitoring of the pore width as well [156].

Comparative experiments revealed that 12-membered

ring zeolites showed an induction period (followed byno or limited deactivation) whereas 10-membered ring

systems did not, however, considerable deactivation

was observed in the case of medium-pore zeolites

[157,158]. In conclusion, the disproportionation of eth-

ylbenzene allows a safe discrimination between large-

and medium-pore zeolites and related microporous

materials, however, so far a clear ranking of these

Fig. 30. Isomerisation and disproportionation of m-xylene using acidic catalysts: Main reaction pathways [141]. Reproduced by permission of Wiley-

VCH, Weinheim.


systems according to their effective pore width is not

possible [141]. Finally, the Catalysis Commission of

the International Zeolite Association (IZA) has recom-

mended the disproportionation of ethylbenzene usingLaNaY zeolite as a standard reaction for acidity charac-

terisation of acid zeolites [121].

4.6. Isomerisation and disproportionation of m-xylene

It is well known that m-xylene can perform both

isomerisation into o- and p-xylene as well as dispropor-

tionation into trimethylbenzene isomers and tolueneunder acidic conditions (see Fig. 30) [159,160].

Gnep et al. suggested for the first time the application

of m-xylene conversion for exploring the effective pore

width of zeolites more than two decades ago [161]. They

proposed several criteria with respect to the character-

isation of zeolites and related microporous materials in

terms of pore architecture:

1. the relative rates of the formation of o- and p-xylene,

2. the rate ratio between isomerisation and dispropor-

tionation and

3. the distribution of the trimethylbenzenes as products

of the disproportionation.

The first criterion is based on the observation that o-

and p-xylene are formed at about the same rates as longas no shape-selective effect is in operation. However,

with decreasing pore width, the formation of the para-

isomer is increasingly preferred in comparison to the

bulkier ortho-isomer (product shape selectivity).

The second criterion refers to the finding that isom-

erisation of m-xylene is more favoured compared to dis-

proportionation in the case of decreasing pore width,

since suppression of the disproportionation is due tothe restricted transition state shape selectivity rather

than to mass transfer limitations.

Finally, the isomer distribution of the trimethylbenz-

enes reflects the restricted transition state selectivity as

well, when discussing the hindered formation of 1,3,5-

trimethylbenzene in mordenite as compared to a largeramount of this isomer using zeolite Y as catalyst.

The m-xylene test reaction has found certain interest

with respect to probing the pore architecture of zeolites

and related microporous materials, however, the appli-

cation has been limited concerning the determination

of the pore width of those porous solids [141].

5. Conclusions and outlook

During the last 30 years the main focus on zeolites

and related microporous materials as industrial catalysts

has been concentrated on crude oil and natural gas

upgrading as well as petrochemical industry—mainly

based on the acid properties of those porous materials.

To some minor extent, zeolites have been used as redoxcatalysts as well. However, new demands and challenges

require the use of porous catalysts, besides the tradi-

tional applications, within new fields as well. The pro-

duction of fine and special chemicals, drugs and other

compounds will in the future to a much larger extent

be performed utilising the catalytic shape selective prop-

erties of porous materials. Furthermore, the changes in

order to adapt the demands and constraints imposedby the environmental legislation will, to a large extent,

be pursued by the application of porous materials

[162]. Porous materials in this context will not only

cover traditional zeolites and related microporous mate-

rials, but also include mesoporous materials, micro- and

mesoporous composite materials, metal organic open

framework materials, porous alumino silicates, organic-

inorganic hybrid materials and others [163].However, for a number of important applications,

the pore size of the microporous materials are too small


for processing bulkier molecules. As a consequence, new

porous materials with larger pores have to be prepared.

With respect to the improvement of the catalytic proper-

ties of porous materials and taking into account the pro-

cessing of bulkier molecules, the following challenges

have to be considered: Synthesis of ultra-large porezeolites or related microporous materials as well as del-

aminated systems. Furthermore, the preparation of

nano-crystalline zeolites or related microporous materi-

als seems to be an interesting alternative as well, allow-

ing large ratios of external to internal surfaces to be

achieved. Consequently, the reaction of bulky molecules

will take place at the external surface and/or at the pore

mouth of the zeolites. The preparation of micro- andmesoporous composite systems would complete the

application of nano-crystalline zeolites in this context

quite well. In addition, delaminated zeolites show a very

large external surface with excellent accessibilities to ac-

tive sites for bulky molecules. Finally, the synthesis of

new structures with ultra-large pores and three-dimen-

sional framework systems of connected pores would be

the most direct way to expand the possibilities of zeolitesand related microporous materials within catalysis. A

number of efforts in this direction are in progress [164].

Catalysis will have to play a major role in overcoming

the technical challenges we will face in the future—some

of them will require scientific and technological break-

throughs—and without doubt, porous materials will

have a strong impact in relation to this development

[162].

Acknowledgments

The author is indebted to the organisers of the 14th

International Zeolite Conference Pre-Conference School

for the invitation to present this contribution. Thanks

are due to Andreas C. Moller for his assistance in con-nection with the technical preparation of the figures.

Finally, the author gratefully acknowledges financial

support from SINTEF and funding from the European

Commission in connection with the TROCAT project

(contract no. G5RD-CT-2001-00520) and BIOCAT

project (contract no. ENK6-CT-2001-00510).

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gas phase catalysis by zeolites

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