Aromatic Hydrogenation Catalysis: A Review

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Aromatic Hydrogenation Catalysis: AReviewAntonymuthu Stanislaus a & Barry H. Cooper aa Haldor Tops⊘e A/S Research Laboratories Lyngby, DenmarkPublished online: 23 Sep 2006.

To cite this article: Antonymuthu Stanislaus & Barry H. Cooper (1994): Aromatic HydrogenationCatalysis: A Review, Catalysis Reviews: Science and Engineering, 36:1, 75-123

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CATAL. REV.-SCI. ENG., 36(1), 75-123 (1994)

Aromatic Hydrogenation Catalysis: A Review

ANTONYMUTHU STANISLAUS and BARRY H. COOPER*

Haldor Tops@e AIS Research Laboratories Lyngby, Denmark

I. INTRODUCTION . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 76

11. TYPES OF AROMATIC COMPOUNDS IN PETROLEUM FRACTIONS . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 77

111. THERMODYNAMIC ASPECTS . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 78

IV. KINETICS AND REACTION MECHANISMS . . . . . . . . . . . . . . . . A. Kinetic Studies on Model Aromatic Compounds: Benzene

and Alkyl Benzenes . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . B. Hydrogenation of Di-, Tri-, and Polyaromatic Compounds . , . C. Kinetic Studies on Industrial Feedstocks . . . . . . . . . . . . . . . . . . .

83

83 90 97

V. CATALYSTS AND NATURE OF CATALYTIC SITES . . . . . . , . A. Catalysts . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . B. Nature of Catalytic Sites in Sulfide Catalysts . . . . . . . . . . . . . . . C. Nature of Catalytic Sites in Sulfur-Tolerant Noble

Metal Catalysts . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . INDUSTRIAL ASPECTS . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . SUMMARY AND CONCLUSIONS . . . . . . . . . . . . . . . . . . . . . . . . . .

102 102 103

108

110

114

VI.

VII.

*To whom correspondence should be addressed at Haldor Topscbe A h , Nymollevej 55, DK-2800, Lyngby, Denmark.

75

Copyright 0 1994 by Marcel Dekker, Inc.

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76 STANISLAUS AND COOPER

ACKNOWLEDGMENT . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 116

REFERENCES . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 116

Key Words: Aromatic hydrogenation; Aromatic reduction in diesel; Hydrotreating.

I. INTRODUCTION

High aromatic content in diesel fuel has been recognized both to lower the fuel quality and to contribute significantly to the formation of undesired emissions in exhaust gases [l, 21. Because of the health hazards associated with these emissions, environmental regulations governing the composition of diesel fuels are being tightened in both Europe and the United States, leading to limitations on aromatics [3, 41.

The California Air Resources Board (CARB) has passed legislative measures that limit the sulfur and aromatic content of motor vehicle diesel fuel to O.O5%(w) and IO%(v), respectively, effective October 1, 1993. A recent version of an administration-sponsored bill supports this ruling and would prohibit the sale of motor vehicle diesel fuel with >0.05 wt% sulfur and a capped aromatic standard defined as 35%(v) aromatics or a cetane index of more than 40 IS]. In Europe, Sweden has already tightened the aromatics specifications, requiring not more than 5%(v) aromatics for class 1 diesel fuel, and other countries in the region are considering measures to limit the aromatic content or stipulation of a minimum cetane index.

As a result of the stringent environmental regulations, processes for aromatic reduction in middle distillates have received considerable attention in recent years. Studies have shown that existing middle distillate hydrotreaters designed to reduce sulfur and nitrogen levels would lower the diesel aromatics only marginally [6-81.

At present, conventional hydrotreating technology is adapted for aromatic saturation and it has been recognized that aromatic hydrogenation (AH) is more difficult than hydrodesulfurization (HDS) and hydrodenitro- genation (HDN) under conditions that are normally used for hydrotreating. [n addition to this, there are thermodynamic equilibrium limitations on aromatic hydrogenation within the normal operating range of hydrorefining. A clear understanding of the effects of process variables, catalyst type, and the interaction of these variables on chemistry and thermodynamic equilibria of different types of aromatic compounds present in the feedstock is necessary for determination of the optimum operating strategies to handle the aromatics present in the diesel-blending streams.

Despite the importance of AH in the refining industry, the subject has not received much attention. Compared with the extensive literature on HDS [9-121, HDN [13-161, and hydrometallation (HDM) [17-191 pro-

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AROMATIC HYDROGENATION CATALYSIS: A REVIEW 77

cesses, the number of publications on aromatic hydrogenation is relatively small, although catalytic centers responsible for the hydrogenation in hydrotreating catalysts have received some attention in HDS and HDN literature. Particularly, comprehensive reviews on aromatic hydrogenation covering varous aspects of the AH processes are scarce.

The purpose of this paper is to review the work reported in the literature on both basic and industrial aspects of aromatic hydrogenation. Fundamental studies concerning the kinetic, mechanistic, thermodynamic, and catalytic chemistry aspects of aromatic hydrogenation reactions have been examined, and recent scientific literature on the nature of hydrogenation and hydrogenolysis sites is discussed. Particular emphasis is given to recent developments in catalysts and processing technologies for reduction of aromatics to acceptably low levels in diesel fuels in the light of stringent environmental regulations.

11. TYPES OF AROMATIC COMPOUNDS IN PETROLEUM FRACTIONS

Detailed analyses by several techniques [20-261 such as high-pressure liquid chromatography (HPLC), 13C nuclear magnetic resonance (I3CNMR), gas chromatography-mass spectrometry, and by ultraviolet (UV) and infrared (IR) techniques have shown that the aromatics found in petroleum and synthetic middle distillates mainly fall into four groups: (i) monoaromatics, (ii) diaromatics, (iii) triaromatics, and (iv) polycyclic aromatics. Among these, the polycyclic aromatics with four or more condensed ben-

TABLE 1 Aromatic Compounds in Kuwait Atmospheric

Gas Oil

Aromatic content Aromatic type (wt%)

1. Monoaromatics C3-C, alkyl benzenes 3.68 C, - C, benzothiophenes 14.24

Co-C4 naphthalenes 52.81 Co-C4 dibenzothiophenes 14.38 Co-C, fluorenes 4.12 C,-benzyls + dibenzofuran 4.07

Co-C4 phenanthrenes 6.20 C,-C, pyrenes/fluoranthrenes 0.5

Total 100.00

2. Bicycloaromatics

3. Triaromatics

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TABLE 2 Properties of Hydrotreated ad Unhydrotreated Gas Oils

(first-stage hydrotreating) [25, 271

LGO LCO

Aromatic Unhydro- Hydro- Unhydro- Hydro- type treated treated treated treated

Total aromatics 33.7 31.6 70.2 70.9 Mono- 17.7 22.4 11.2 38.5 Di- 11.5 5.7 49.5 26.9 T i + - 4.5 3.5 9.5 5.5 Polars 0.8 0 - -

~~ ~~

Note. LGO = light gas oil; LCO = light cycle oil.

Lene rings are largely present in many high-boiling petroleum fractions (BP > 35OoC), whereas the other three types are important components of middle distillates. Typical analysis data (26) on the type of aromatic compounds present in untreated gas oil fractions of Kuwait petroleum are presented in Table 1.

It is seen that diaromatics (mainly alkyl naphthalenes) constitute a major portion of the total content of aromatics in straight-run gas oil. However, in hydrotreated oils, monoaromatics (e.g., alkyl benzenes, ben- zocycloparaffins, and benzodicycloparaffins) are present in larger quantities than di- and triaromatics (Table 2).

Condensed multiring aromatic compounds are hydrogenated more eas- ily to the corresponding monoaromatics under mild hydrotreating conditions. This point is discussed in more detail in the section on kinetics and thermodynamics. Typical structures of some mono-, di-, and triaromatic compounds found in middle distillate fractions are shown in Table 3.

The amount and type of aromatics in middle distillates show large variations depending on the origin of the feedstock. Typical HPLC data on aromatic type distribution in gas oils from different sources are summarized in Table 4.

It is seen that the aromatic content in light cycle oils (LCO) from fluid catalytic cracking (FCC) units is very high (about 70%). Straight-run gas oils contain relatively low amounts of aromatics. The front-end LCO contains virtually all the monoaromatics, whereas the back-end LCO contains predominantly di- and tri + -aromatics.

111. THERMODYNAMIC ASPECTS

The hydrogenation of aromatic compounds is reversible, and at typical hydrotreating conditions, complete conversion is not possible owing to equi-

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TABLE 3 Typical Structures of Some Aromatic Compounds Present in

Petroleum Fractions AROMATIC COMPOUND TYPE

~

1. MONOAROMATICS

a) Alkyl Benzenes

b) Benrocycloparaf f Ins

c) Benzodicycloparaf f ins

2. DIAROMATICS

a) Naphthalenes and

b) Biphenyl

alkyl naphthalenes

c) lndener

d) Naphthocycloparafflnr

3. TRIAROMATICS

a) Anthracenes

b) Phenanthrenes

c) Fluorener

TYPICAL STRUCTURES

librium limitations. The hydrogenation of an aromatic species, A, is given by:

A + nH, AH (1) where AH is the hydrogenated product (a naphthene). It can be shown [28] that the equilibrium concentration of the aromatic species can be approx- imated by:

where YA and YAH are the mole fractions of the aromatic and naphthene

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TABLE 4

Aromatic Type Distribution in Different Gas Oils [22]

HPLC analysis

Saturates Total aromatics Mono- Di- Trif- Polars

(wt%) LGO 74.0 26.0 16.8 8.8 0.4 0.0

LGO/LCO LCO (70%/30%) FE-LCO BE-LCO

29.0 59.4 25.2 58.4 70.2 40.6 74.8 41.6 11.2 12.6 29.0 0.1 49.5 24.3 45.8 32.0 9.5 3.7 0.0 9.5 0.8 0 0.0 0.8

Note. LGO: straight run light gas oil; LCO: light cycle oil from FCC unit; FE-LCO: front-end LCO (IBP-290°C); BE-LCO: back-end LCO (290°C + ).

FIG. 1. Equilibrium constants for the hydrogenation of the benzene series P71.

species, respectively; K, is the equilibrium constant, and PH2 the partial pressure of hydrogen. In the derivation of this equation it is assumed that liquid activity coefficients and fugacities for A and AH are equal, and that the hydrogen activity coefficient and the ratio of fugacity to total pressure at hydrotreating conditions are both unity.

It is seen that high pressures favor low equilibrium concentrations of aromatics (high conversions). This is particularly so for reactions where the number of moles of hydrogen, n , required for complete saturation is high. Aromatic hydrogenation reactions are highly exothermic, with heats of reaction typically in the range 63-71 kJ/mole H2 [29,30]. K , decreases with

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-2.5

-5

Y

a 0 -

-7.5

-1 0

1000 T'K

FIG. 2. Equilibrium constants for hydrogenation of phenanthrene [37].

increased temperature, and thereby the equilibrium aromatics concentration increases as the temperature is increased.

Experimental data for the calculation of equilibrium constants are sparse [31-341. Group contribution methods have been applied by a number of workers in order to be able to estimate equilibrium constants [35-371. As pointed out by Girgis and Gates [38], the uncertainties of these methods can result in more than an order of magnitude error in calculated K,.

Calculated equilibrium constants indicate that there is considerable variation from one family of aromatic hydrocarbons to another. In the hydrogenation of benzene homologues (Fig. l) , the value of the equilibrium constant decreases with an increase in both the number of side chains and the number of carbon atoms in each side chain [37,39]. The same is found for naphthalenes [28].

For aromatic hydrocarbons containing more than one ring, hydrogenation proceeds via successive steps, each of which is reversible. The equi-

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- 525 575 625 675 725 7 I U >- +

5

1

%25 575 625 675 725 775

525 575 625 675 725 775 TEMPERATURE, K

FIG. 3. Equilibrium concentrations for first- and last-ring hydrogenation of phenanthrene and naphthalene as function of hydrogen pressure.

librium constant is generally higher for the hydrogenation of the first ring (Fig. 2, Ref. 37), but since more moles of hydrogen are involved in the final ring hydrogenation (3 moles compared with 1 or 2 moles for hydrogenaton of the first ring in phenanthrene), hydrogenation of the first ring is usually less thermodynamically favored than hydrogenation of the final ring at typical hydrotreating conditions.

That this is so can be seen in Fig. 3, in which the data of Fig. 2 and Ref. 28 have been used to calculate the equilibrium aromatic concentration for the saturation of phenanthrene and naphthalene as a function of tem-

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perature at three levels of hydrogen pressure. The equilibrium aromatic concentration for first-ring hydrogenation is higher than that for final-ring hydrogenation at temperatures below 375°C (648 K) at 3 MPa, and at temperatures up to 435°C (708 K) at 10 MPa. These curves also demonstrate the strong effect of hydrogen pressure on the equilibrium aromatic concentration: At 350°C (623 K), the equilibrium concentration for the reaction naphthalene to tetralin decreases from 11.4% at 3 MPa to 1.2% at 10 MPa, while the equilibrium concentration decreases from 21.5% at 3 MPa to 4.5% at 10 MPa for the reaction phenanthrene to tetrahydrophenanthrene.

IV. KINETICS AND REACTION MECHANISMS

Several types of aromatic compounds have been used as representative model compounds in the study of A H kinetics. These include benzene, toluene and alkyl benzenes, naphthalene, biphenyl, phenanthrene, pyrene, fluorene, and fluoranthene [38, 40-511.

In most of the studies, particularly those involving large condensed ring aromatic compounds, the complex reaction network has been reported

The various kinetic studies reported in the literature for different model aromatic compounds and real industrial feedstocks are reviewed in this section. Earlier work carried out on the subject (before 1980) is only briefly discussed in this review. Readers may refer to Kiperman [52] and to Weiser and Landa [53] for earlier work on the subject.

[381.

A . Kinetic Studies on Model Aromatic Compounds: Benzene and Alkyl Benzenes

Hydrogenation kinetics of benzene have been studied extensively both on supported group VIII metal (e.g., Ni, Pd, Pt) and on metal sulfide (e.g., MoS,, WS2, Co-Mo-S/A1,03, Ni-Mo-S/A1203, and Ni-W-S/Al,O,) catalysts. A general reaction network for benzene hydrogenation over a sulfided Co- Mo/Al,03 catalyst is shown in Scheme 1 [44].

As is evident from Scheme 1, benzene hydrogenation to cyclohexane and formation of methyl cyclopentane are consecutive reactions. Some workers [54] argue that hydrogenation to cyclohexane and hydroisomerization to methyl cyclopentane start from a common intermediate product, cyclohexane, as shown in Scheme 2.

SCHEME 1. Reaction network for hydrogenation of benzene.

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SCHEME 2

SCHEME 3. Mechanism for the heterogeneous catalytic hydrogenation of benzene over group VIII metals.

While it is generally agreed in the hydrogenation of benzene over group VIII metal catalysts that the adsorbed state of benzene in active hydrogenation is associatively adsorbed, probably as a wcomplex, there is still considerable debate as to the nature of surface intermediates. Several workers [55-571 have proposed that benzene hydrogenation proceeds through cycloolefins .

Van Meerten et al. [58] found evidence of three forms of chemisorption assigned to benzene on Ni-Si catalysts. A stepwise reaction mechanism has been proposed recently by Van der Steen, Scholten et al. [57 ,59] (Scheme 3). The reaction intermediate, cyclohexene, has been observed to desorb before further hydrogenation. Desorption of the cyclohexadienes (C,H,) has never been observed. Apparently, these dienes are too strongly adsorbed to desorb before further hydrogenation.

Aben et al. [60] proposed the following Langmuir-type kinetic equation for benzene hydrogenation on supported Pt, Pd, and Ni catalysts in a flow system:

where b is the adsorption coefficient for molecular hydrogen, 0 is the fraction of surface covered, and ko is the rate constant. The form of the kinetic expression implies that hydrogen is molecular adsorbed and that the hydrogenation reaction is zero-order in benzene. The activation energy was found to be nearly the same (about 14 k 1 kcallmole) for the three catalysts. While many studies confirm that the reaction is zero-order in benzene [61],

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others report [62] that the reaction order increases from the low value 0.5 at 25°C to 2-3 at 200°C.

Recently, Chou and Vannice [40,41] examined the kinetics of benzene hydrogenation on supported Pd catalysts under a wide range of reaction conditions. At temperatures below 160"C, the apparent activation energies were found to be about 12.0 kcal/mole. The reaction order in hydrogen increased from 0.5 to nearly 4 as the temperature increased from 80°C to 300"C, whereas the reaction order for benzene increased from 0 to 0.8 in the same temperature range. The hydrogenation activity of the catalyst showed a temperature-dependent maximum near 220"C, as reported by others [52]. Such a maximum is apparently not associated with the thermodynamic equilibrium limitation, since the rate of the reverse reaction becomes considerable only at high temperatures. Chou and Vannice [40] attributed this behavior to the formation of hydrogen-deficient surface carbonaceous species by dehydrogenation of benzene, which inhibits the prin- cipal hydrogenation reaction.

They proposed a Langmuir-Hinshelwood-type rate expression in which benzene and hydrogen are assumed to be adsorbed on different sites and the reaction proceeds by stepwise addition of hydrogen to the benzene ring. Desorption of cyclohexane is assumed to be fast and irreversible. Rate terms for a concurrent equilibrated, dehydrogenation reaction sequence involving adsorbed benzene have also been included in the rate equation to account for changes in the number of active sites due to carbonaceous species formed under different reaction conditions. The complete rate expression is too complex to reproduce here. Readers may refer to the original paper [41] for further details.

Earlier work by van Meerten and Coenen [63] on the kinetics of benzene hydrogenation using Ni-Si catalysts considered the relative rates in stepwise addition of hydrogen and concluded that the hydrogen addition steps all have the same rate.

Very few kinetic and mechanistic data for the hydrogenation of toluene are available in the literature. Lepage [37] reported the first extensive studies on the kinetics of toluene hydrogenation over sulfided Ni- W/A1,03 catalysts. At low pressures (0-0.5 MPa), the reaction rate was found to increase linearly as a function of hydrogen and toluene partial pressures, indicating that the reaction has an apparent order of 1 with respect to both hydrogen and toluene. The kinetic behavior of the reaction was represented by the following Langmuir-Hinshelwood-type rate equation:

where k = rate constant; bA and bsi are adsorption coefficients .for aromatic and sulfur compounds; and PA, PHz, and Psi represent the partial pressures of the aromatic hydrocarbon, hydrogen, and sulfur compounds including H2S.

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b? e: .-

At higher pressures (above 0.4 MPa), however, the kinetic data ap- peared to be consistent with the following rate equation:

.-

>

The form of this equation permits the adsorption of hydrogen on the catalyst to be considered independently of the other compounds. It also implies competitive adsorption between the aromatic compound (toluene) and sulfur compounds.

A strong inhibiting effect of H2S on toluene hydrogenation as predicted by the above rate equation was observed (Fig. 4).

The apparent energy of activation for toluene hydrogenation over sulfided Ni-W/AI2O3 catalyst calculated from the kinetic data was 92 kJ/ mole.

In another study from the same laboratory, Ahuja et al. [64] proposed the following rate equation for hydrogenation of toluene in the presence of thiophene and cyclohexene.

with

bHZS >> bTh > b, >> bT in which kT is the rate constant, b is the adsorption coefficient, andp is the partial pressure of the reagents and products. This rate law is consistent with a Langmuir-Hinshelwood mechanism without competitive adsorption between hydrogen and the aromatic hydrocarbon and inhibition by the reactants, products, and sulfur compounds. Some other studies with model aromatic compounds have , however, shown that competitive adsorption and inhibition by the product naphthenes are negligible [62].

Several studies on the hydrogenation of toluene on group VIII metals (Pt, Pd, and Ni) have been reported in the literature [65-681, but most of

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these were intended to probe the relation between electronic structure and activity of the metal catalysts.

The first extensive determination of kinetics of toluene hydrogenation on supported Pd catalysts over a wide range of conditions was reported very recently by Rahman and Vannice [42,43]. These authors also examined the effect of methyl substituents on the rate of hydrogenation of benzene rings by comparing the kinetic behavior of benzene, toluene, and xylenes.

The kinetic data showed that all reactions were first-order in hydrogen and zero-order in aromatic compound concentration, indicating a near saturation coverage of the active sites by aromatic species. Activation energies were near those of benzene (50.2, 49.3, 50.2, 53.1, and 57.2 kJ/mole for benzene, toluene, rn-xylene, o-xylene and p-xylene, respectively). The use of highly acidic oxide supports (e.g., Si02-A1203, Ti02) was found to markedly enhance the hydrogenation rate. Furthermore, contrary to the case of benzene, there was no deactivation or significant coverages of hydrogen deficient surface inhibitors.

The following Langmuir-Hinshelwood-type rate equation involving two sets of active sites (one on the Pd surface and the other composed by acid sites in the Pd-support interfacial region with spilled-over hydrogen) was proposed to describe the hydrogenation kinetics of all these aromatic hydrocarbons.

At high enough pressures of the aromatic compound this just becomes:

This rate expression indicates that hydrogen and aromatic hydrocarbons adsorb on different types of sites, and that the reaction occurs simultaneously on Pd surface and on acid sites in the interface region. The aromatic species are adsorbed strongly, and near saturation exists at high enough partial pressures of aromatics so that KAPA >> 1 and the reaction is zero-order in aromatic compound.

These authors also observed that the hydrogenation rate depends on the number and position of methyl substituents in the order:

Benzene > toluene > p-xylene 2 rn-xylene > a-xylene

which is in agreement with the results previously reported by others [52, 69,701. However, a completely reverse order of reactivity has been reported for these hydrocarbons on sulfide catalysts such as Ni-W-S/A1203 and Ni- Mo-S/A1203, as seen in Table 5. The reasons for the difference in the effect

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TABLE 5 Relative Rate Constants for Hydrogenation of Benzene and Alkyl

Benzenes Over Metal and Sulfide Catalysts

HYDROGENATION REACTION

-~ ~~

REFERENCE

IELATIVE HYDROGENATION RATE CONSTANT

METAL SULFIDE CATALYSTS METAL CATALYSTS -

‘tlslo,

1

0.3

0.08

0.01

70 -

I40

1

0.63

0.65

0.47

0.11

42

m / W

1

0.52

0.3

0.24

71

Mi-w-SJ W%

1

1.6

1.7

3.5

69.72 -

of alkyl substituents on hydrogenation rate for metal and metal sulfide catalysts are not well understood.

The bonding of a molecule on a surface depends on the local electron density of states on the adsorbing metal atoms. It is generally known that aromatic compounds such as benzene and toluene are bonded to the metal surface via n-bonds involving an electron transfer from the aromatic ring to the unoccupied d-metal orbitals [71, 73, 741. Since the n-electron cloud density in toluene is higher than that of benzene as a result of the inductive

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effect of the methyl group, toluene is generally expected to be more strongly adsorbed than benzene. In confirmation of this, Gallezot and co-workers [75] observed that the ratio of adsorption coefficients of toluene and benzene (&I&) on group VIII metals is always higher than unity and increases with the electron-deficient nature of the metals.

Based on these principles, Vannice and co-workers [42] argue that toluene has a greater stability on the catalyst surface than benzene and may hence undergo hydrogenation at a lower rate.

Pd supported on strongly acidic supports is electron deficient [76-781, and as a result toluene would be expected to be more strongly adsorbed than benzene on the metal. In addition, a stronger interaction between toluene and the acidic site as compared with benzene could reduce the hydrogenation rate of the former. In support of this argument, the ratio of toluene to benzene hydrogenation rates has been found to be higher in the absence of acidic supports than in the presence of one [42].

Similarly, further increase in electron density with the addition of another methyl group may explain the lower reactivity of xylenes. However, it should be noted that o-xylene has the lowest reactivity among the xylenes. This has primarily been attributed to a steric effect. Steric hindrance of neighboring methyl groups in o-xylene could have a more significant effect on the formation of n-bonded complexes.

For metal sulfide catalysts, as mentioned earlier, the reactivity pattern of benzene and alkyl benzenes is the opposite. In other words, addition of methyl groups to benzene nucleus enhances the reactivity of these molecules for hydrogenation. This behavior could be discussed on the basis of the work of Geneste and Moreau [79-811 on the hydrogenation of a series of substituted benzenes. Based on experimental results for hydrogenation and hydrogenolysis of a series of benzenes substituted by different electronega- tive groups such as NH2, NHC6Hs, OH, OC6H5, SH, SC6H5, C2H5, C6H5, and halogens over a sulfided Ni-Mo/A1203 catalyst at 340°C and 7 MPa H2 pressure, these authors proposed that the reactions of hydrogenation and hydrogenolysis are mostly influenced by the n-electron delocalization through resonance, hydrogenation being favored by highly electron-donating substituents. Hydrogenation is easier when the ring to be hydrogenated is less aromatic [79].

These authors argue [80, 811 that electron-donating substituents in an aromatic compound can cause localization of the n-electron at a preferred position, leading to a diminution in the strength of the n-absorption and consequently higher reaction rates. This explanation is similar to the one advanced earlier by Nag [82]. According to Nag, aromatic double bonds, due to stabilization by resonance, are normally difficult to hydrogenate. However, when a wcomplex is formed, these double bonds are weakened by transfer of electrons from the bonding highest occupied molecular orbital (HOMO) to the antibonding lowest unoccupied molecular orbital (LUMO) of the reaction via the metal atom on the catalyst surface. The stronger the

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Ir-complex bond, the weaker the double bonds become, and the more vulnerable they are to the attack by hydrogen. Therefore, any factor that strengthens the Ir-complex bond should augment hydrogenation.

This explanation, however, contradicts the arguments of Vannice and co-workers [42], who claim that electron-donating substituents increase the density of the Ir-electron cloud, which in turn increases the bond strength between the metal and the aromatic compound and leads to a lower reactivity.

It is possible that the mechanisms of adsorption on the two types of catalysts are entirely different so that the hydrogenation of the aromatic ring is enhanced in one type of catalyst (group VI metal sulfides) and suppressed on the other by the presence of electron-donating groups in the aromatic ring. Steric effects of the alkyl substituents should also be considered. The question still remains to be solved and obviously more research is needed on this subject.

B . Hydrogenaton of Di-, Tri-, and Polyaromatic Compounds

Kinetics of hydrogenation of aromatic hydrocarbons containing two or more rings are more complicated. Most kinetic studies on di- and multiring aromatic compounds reported in the literature deal mainly with the reaction pathways and reactivities rather than quantitative kinetic models that include adsorption mechanism of reactants, intermediates, and products [38].

Spare and Gates [44] proposed reaction pathways for hydrogenation of naphthalene, biphenyl, and 2-phenyl naphthalene over a sulfided Co- Moly-AI,O, catalyst at 325°C and 75 atmospheres pressure (atm) (Scheme 4). The reactions were approximately first-order in the aromatic reactant. The rate constants for various forward and reverse reactions are shown in the network. The rate data for benzene hydrogenation are also included in the network for the purpose of comparison.

The rate data for various reactions in the network (Scheme 4) show that the rate of hydrogenation of naphthalene to tetralin is considerably (about 21 times) faster than that of benzene hydrogenation as reported by others [53]. Presence of phenyl substituents in benzene or naphthalene has no significant effect on the reactivity for hydrogenation of these aromatic hydrocarbons, in contrast to the effect of alkyl substituents, which enhance hydrogenation. The enhancement in reactivity resulting from the electro- donating influence of the aryl substituents is probably compensated for by increased steric hindrance of adsorption.

Patzer et al. [83] studied the hydrogenation kinetics of 1-methyl naphthalene over a series of commercial coal liquefaction catalysts in the temperature range 343" to 482°C and found the reaction to obey first-order kinetics. These authors also noticed that the ratio of concentations of tetralin to naphthalene in the product remained almost constant at about 0.49 2 0.03, being unaffected by changes in catalyst composition, feed flow rate,

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BENZENE CYUOHMANE

W

SCHEME 4. Reaction networks for hydrogenation of benzene, biphenyl, naphthalene, and 2-phenylnaphthalene in the presence of sulfided COO-Mo0,ly- A1203 at 325°C and 75 atm [44].

SCHEME 5

and reaction temperature in the range 316"-399"C. They attributed this result to an equilibrium limitation and proposed the reaction network shown in Scheme 5.

Their interpretation of an equilibration of the products was disputed

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by Spare and Gates [44] in view of the lack of dependence on the ratio of tetralin to naphthalene on temperature. They pointed out that in the ex- periments reported by Patzer et al., the residence time (defined as reciprocal of hourly space velocity, LHSV-') was typically varied over a small range (0.3 to 0.6). Estimation of relative concentration profiles from kinetic parameters showed that both naphthalene and tetralin concentrations are almost linearly dependent on residence time (LHSV- l) in this range (Fig. 5 ) , and consequently the naphthalene-to-tetralin ratio is almost constant, giving the incorrect impression of an equilibrium distribution.

Spare and Gates [45] reported the first extensive determination of the kinetics of diphenyl hydrogenation catalyzed by sulfided Co-Mo/Al,O,, and proposed the following Langmuir-Hinshelwood form of equation to represent the rate data for the reaction network shown in Scheme 4.

where K ( i ) is the adsorption equilibrium constant for species (i). This rate expression accounts for the reversibility of the reaction and

inhibition by H,S and reactants (H, and biphenyl). At high hydrogen partial pressures, the reaction is nearly first-order in hydrogen concentration. The Langmuir dependence of rate on biphenyl concentration suggests that saturation of the catalytic sites with biphenyl is approached at higher concentrations. The effect of all the important process variables is predicted by this rate model and a good agreement between the observed and predicted rates is observed, as shown in Fig. 6.

Hydrogenation reaction pathways for condensed polynuclear aromatic

0 2 4 6 8 (LHSV)", hr .

0

FIG. 5 . Relative concentration profiles for catalytic hydrogenation of 1- methyl naphthalene at 70 atm and 316°C. The curves are predicted from the data of Patzer et al. [83].

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AROMATIC HYDROGENATION CATALYSIS: A REVIEW

I I 1 1 I

93

FIG. 6. Comparison of observed rates and the predictions of Eq. (9).

hydrocarbons such as fluorene, fluoranthene, and phenanthrene over sulfided Ni-W/A1203 catalysts are summarized in Figs. 7(a)-7(c). The results show that hydrogenation of these polynuclear aromatic compounds occurs ring by ring in a series fashion. The hydrogenation reactions were essentially found to be first-order in the aromatic hydrocarbon. No kinetic rate equations were proposed for the above reactions.

With polycondensed ring aromatic hydrocarbons, the first-ring hydrogenation has been observed to be most favored kinetically. The rates of hydrogenation of subsequent rings tend to become lower, and hydrogenation of the last ring proceeds with considerable difficulty compared with the initial hydrogenation steps.

The relative rate constants for hydrogenation of the first ring of a few condensed di- and triaromatics such as naphthalene, anthracene, and phenanthrene over different sulfide catalysts are presented in Table 6. The resonance energies of various aromatic compounds are also included in the table. A comparison of the resonance energies indicates the following order of aromaticity:

Benzene < naphthalene < anthracene < phenanthrene

The partial resonance energy or aromaticity of the different rings in fused multiring aromatic systems may be different. Thus, for example, the resonance energy per ring in naphthalene molecules is considerably lower than that of benzene (Table 6). The low aromatic character of one of the rings in the naphthalene molecule is experimentally shown by its ability to undergo addition reactions across 1,2 positions. The middle ring of anthracene is more reactive, and it has been reported that it behaves just like a double bond. In multiring aromatic compounds, the ring with the lowest

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STANISLAUS AND COOPER

U (a) Fluoranthene Hydrogenation Pathways (ref. 51)

Fluorene

41 K l

IUIF Iso-IUIF isomers

(b) Fluorene IIydrogenation Network (ref. SO)

Phenanthrene Dehydrophenanthrene

Wahydrophcnanthrene Tctrahydrophenanthrene

(C) Ilicnanthrene Ilydrogcnalion Network (rcf. 46. 47)

FIG. 7. Hydrogenation pathways for polynuclear aromatic hydrocarbons: (a) fluoranthene, (b) fluorene, (c) phenanthrene.

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TA

BL

E 6

d 5n

0

z

Rel

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ate

Con

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ydro

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of M

ono-

, Di-

, and

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atic

Hyd

roca

rbon

s O

ver

Sulfi

de C

atal

ysts

[79,

811

H

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Kea

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onst

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tal

z En

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Res

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59-75

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71-105

8

@@

+

H* s&

4

- 85-92

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STANISLAUS AND COOPER

1

0

-1

? 3 - 2 A M

3 -3

- 4

- 5

- 6

H experimental I -log k = calculated

/ &

0.99 1.00 1.01 1.02 1.03 1.04

Electron Density 0

FIG. 8. Relationships between .rr-electron density and aromatic hydrogenation. The rate constants are for hydrogenation of the most reactive ring, marked “a.”

aromaticity is hydrogenated first. The rate data presented in Table 6 indicate the following order of reactivity for one-ring hydrogenation:

Anthracene > naphthalene > phenanthrene > benzene

In a recent study, Neurock and co-workers [84] found good correlation between the rate constants for hydrogenation of a series of polycyclic aromatic hydrocarbons and the .Tr-electron density (Fig. 8). IT was determined at each potential site of hydrogenation, and the one with the largest value of IT was considered to be the kinetically significant site. The correlation of Fig. 8 is shown below.

where A is - 155.66 2 34.12 and B is 151.23 f 33.88. A possible explanation of the correlation between . ~ r and the reaction

rate constant is that there is a Coulombic attraction between the carbon site on the aromatic ring and the attacking reagent. The site where the highest IT can be found is the kinetically significant one.

log k = A + B(T) (10)

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C. Kinetic Studies on Industrial Feedstocks

Literature information on kinetic studies of aromatic hydrogenation in industrial feedstocks such as petroleum and synthetic middle distillates is relatively scarce, probably due to the complexity of the reactions. Some authors [37] have proposed an overall kinetic order of 1 relative to the concentration of aromatic compounds taken as a whole. The reversible reaction was, however, not taken into account in this approach. The conversion of aromatics to naphthenes is a reversible reaction and involves a chemical equilibrium which is temperature dependent. Several studies [85 - 871 have shown that the aromatic content of the product decreases with increasing reaction temperature , but then increases as the temperature is further raised; that is, the aromatic content passes through a minimum. Therefore, the rate of the reversible reaction and thermodynamic equilibria must be taken into account in deriving kinetic models for aromatics hydrogenation.

Wilson and Kriz [88, 891 used the following equilibrium reaction as a basis for developing a model for kinetics of hydrogenation of aromatics in middle distillates:

Aromatics + nH2 $ naphthenes

By assuming that the forward reaction is pseudo-first-order (since hydrogenation is carried out in large hydrogen excess at constant partial pressure) and the reverse reaction is first-order in naphthene, the rate expression is given by:

where: CA = fractional aromatics content (-); CN = fractional naphthene content (-); t = time, h; k, = forward rate constant, h-l MPa-l; k, = reverse rate constant, h-l; PH2 = hydrogen partial pressure, MPa; and n = reaction order with respect to hydrogen partial pressure (-).

On integration, and with suitable substitution, the following expression is derived:

where CAo is the initial aromatics concentration, C,, is the equilibrium aromatics concentration, and kR is the hydrogenation rate constant = kF (1 + CAe/CNe), where kf x Pg2.

At low temperatures [89, 901, equilibrium effects were found to be negligible, and simple pseudo-first-order kinetics were used to obtain values of hydrogenation rate constant (kF) from:

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A kinetic model based on a simple first-order reversible reaction was developed by Yui and Sanford [91, 921 for aromatics hydrogenation in middle distillates.

where rn = reaction order with respect to aromatics content. Assuming that m and IZ are equal to 1, integrating the equation and substituting l/LHSV for t , the following relationship is obtained:

kr

kr kfPH2 + kr

- - kfPH2 + k* kfPH, + kr - CA -

LHSV In CAO -

At equilibrium, the fractional aromatic content becomes:

which when substituted into the integrated rate equation gives: CA - CAe - CACI - CAe CAe LHSV ( k r ) In -

which is similar to the model proposed by Wilson and Kriz [88-901. The model was used to calculate kinetic parameters from published

catalytic aromatic hydrogenation data from a variety of sources [85, 87, 88, 931. Aromatic contents at various operating conditions were calculated using these kinetic parameters, and good agreement between the observed and calculated results was found (Fig. 9).

In a recent study on kinetics of aromatics hydrogenation of five bitumen derived gas oils, Yui and Sanford [94] used a similar method but made the assumption that the power term for PH2 was equal to 0.5 when the aromatic content was defined as aromatic carbon content. The reasoning was that, stoichiometrically, 0.5 mole of hydrogen is consumed in the conversion of 1 atomic weight of aromatic carbon, irrespective of the aromatic species. Furthermore, they introduced a power term for LHSV to account for deviations from plug flow experienced in bench-scale trickle-bed reactors. The model gave a reasonable fit to the experimental data, and it was found that the degree of hydrogenation at equilibrium was lower for the heavier feedstocks. However, the power term for LHSV was found to vary from feed to feed (ranging from 0.28 to 1.0), which is not to be expected if the power term is an expression of deviations from plug flow, since flow conditions and levels of conversion were similar for each feed. It is possible that the variation in the power term is an artefact stemming from the use of a model based on a single, reversible reaction to describe what is in fact a series of consecutive reactions. It is also a simplification to assume that the power

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o T . . . . . i s 573 613 653 693

0- 573 613 653 693

8 100 1 1 Cokrr HGO 8.8 MPo

573 613 653 693

TEMPERATURE, (K)

FIG. 9. Observed and calculated percent aromatics hydrogenation at various operating conditions. The solid lines indicate equilibrium control and dotted lines indicate kinetic control.

term for hydrogen pressure in the kinetic equation is equal to that given by stoichiometry.

Magnabosco [95] proposed a model that takes the pressure dependency in the kinetic equation into account, and that can be used for consecutive reactions. For each reaction step:

S1 + 01H2 S 2

the kinetic equation has the form:

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where T = ULHSV, x = the power term for hydrogen partial pressure in the forward reaction, E = the activation energy for the forward reaction, and K, is the equilibrium constant. (In Magnabosco’s article, the exponen- tial expression in the second term on the right-hand side of the equation is given as - ( E - AH)/RT, where AH is the heat of reaction, and ( E - A H ) is equal to the activation energy of the reverse reaction. However, based on a definition of K , = K1/(K2 x Pg;), the correct form of the kinetic equation is as given above).

Using values of K , obtained from the literature as well as tests on jet fuel and light gas oil for which detailed analyses were made of feed and products, Magnabosco determined the kinetic parameters for single, double, and triple aromatic ring hydrogenation by numerical integration of the kinetic equation. The reaction schemes considered were:

$-R 0-R $$-R f $0-R e 0 0 - R

4 0 - R 2 $00-R 0 0 0 - R 5 M 4 - R

The power term for hydrogen pressure. x, was found to be 1.8 for hydrogenation of monoaromatics, 1.5 for diaromatics hydrogenation (both steps), but only between 0.077 and 0.346 for triaromatic hydrogenation. The effective rate constant for hydrogenation, kf x PS2, was about 6 times higher for diaromatic species than for monoaromatics and twice as high for tetralins than benzene homologues. Surprisingly, values of kf x PL2 for the three-ring aromatics hydrogenation reactions were all at about the same level except for hydrogenation of $$$-R to $O$-R, which was at about one fifth the level of the other reactions.

Hannerup [96] developed rate equations for aromatic hydrogenation considering the equilibrium reactions between monoaromatics (MA) and diaromatics (DA):

$O$-R

where AH stands for hydrogenated aromatics. The rate equations used had the form [97]:

- - - -klDAPkl + k2MA dDA

dt

- - - -K3MAPh2 + klDAP&, - dMA

dt dAH - = - k4AH + k3MAP&,

dt

k,MA + k4AH

No values are given for the kinetic parameters, but several figures are shown

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that demonstrate the results of the model fit. Figure 10 shows how the concentrations of DA and MA change through a long reactor. The DA + MA reaction is much faster than the MA + AH reaction. Therefore, in the first part of the reactor, DAs disappear rapidly whereas the MAS increase, and the reduction of total aromatics is therefore small. As the equilibrium ratio between DA and MA is approached (the straight line from O.O), the fast DA + MA reaction cannot proceed until MAS are coverted to AHs at a low rate.

The other straight line in the diagram represents the equilibrium ratio between MA and AH. The equilibrium concentrations are where the two lines cross each other.

Figure 11 shows how the MAS and MA + DAs vary in the product as a function of the reactor temperature for one fixed LHSV. This pattern is very similar to what was found in the tests.

Cooper et a1 [202] report finding zero-order kinetics using a sulfur- tolerant noble metal catalyst.

None of the kinetic expressions discussed above include adsorption equilibrium constants for reactants, products, and inhibitors such as H,S. Consequently, the adsorption mechanisms and the competition between the reactants and products including H,S and their effect on the reaction rate are not represented in these kinetic models.

In general, the use of the Langmuir-Hinshelwood-Hougen- Watson (LHHW)-type of rate equation for representing the hydrogenation kinetics of industrial feedstocks is complicated, and there are too many coefficients that are difficult to determine. Therefore, simple power law models have been used by most researchers to fit kinetic data and obtain kinetic parameters.

0.3 0.28

3 0.24

0 0.2 -.. 0.16

(r

g 0.12

8 0.08

Q)

0.04

0 0 0.04 0.08 0.12 0.16 0.2 0.24 0.28

Mole Fraction of DA

FIG. 10. Model prediction of the hydrogenation of a feed containing 0.22 mole fraction diaromatics and 0.13 rnonoaromatics. Each square represents an equal distance along the hydrogenation reactor [96].

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0.4 1

300 320 340 360 380 400 Temp'C

A MAe t DAe o MAe oMA +MA+DA

FIG. 11. Model prediction of distribution of mono- and diaromatics during hydrogenation as function of reaction temperature 1961.

V. CATALYSTS AND NATURE OF CATALYTIC SITES

A. Catalysts

Aromatic hydrogenation in industrial feedstocks may be carried out over supported metal or metal sulfide catalysts depending on the sulfur and nitrogen levels in the feedstock. For hydrorefining of feedstocks that contain appreciable concentrations of sulfur and nitrogen, sulfided Ni-Mo, Ni-W, or Co-Mo on y-A2-03 catalysts are generally used, whereas supported noble metal catalysts have been used for sulfur- and nitrogen-free feedstocks. Catalysts containing noble metals on Y-zeolites have been reported to be more sulfur tolerant than those on other supports [98,99], and such catalysts have increasingly been used in recent years for hydrogenation of aromatics in light and middle distillates [loo, 1011 containing appreciable amounts of sulfur. Within the series of Co- or Ni-promoted group VI metal (Mo or W) sulfides supported on y-Alz03, the ranking Ni-W > Ni-Mo > Co-Mo > Co-W has been found [lo21 for hydrogenation, and Ni-W and Ni-Mo/Al,O, catalysts are widely used to reduce sulfur, nitrogen, and aromatics levels in petroleum fractions by hydrotreating. It would be useful to discuss the recent understanding of the nature of catalytic centers involved in aromatic hydrogenation and the interaction between the competitive HDS and HDN reactions that occurs simultaneously with AH.

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B. Nature of Catalytic Sites in Sulfide Catalysts

The nature of active phases of unpromoted and promoted molybdenum and tungsten sulfide hydrotreating catalysts has been widely studied and reported in several reviews [ 103- 1091. For unpromoted molybdenum or tungsten sulfile catalysts, it has been proposed that the coordinatively un- saturated (CUS) sites (or exposed Mo ions with sulfur vacancies at the edges and corners of MoS, (or WS2) structures are active in hydrogenation and hydrogenolysis reactions. Basal planes are inactive in adsorption of molecules and are probably unimportant in hydrotreating reactions. For Co- or Ni-promoted catalysts, several different structural models such as monolayer model [110, 1111, intercalation model [112,113], contact synergy model [114, 1151, Co-Mo-S phase model [116-1191, and catalytic Co site model [120, 1211 have been proposed to explain the role of the promoter and its location in the catalyst. It is not the aim of this section to discuss the various models in detail. However, the relevant points of two important models, namely the contact synergy model and the Co-Mo-S phase model are briefly discussed to gain a better understanding of the nature of catalytically active centers and the interaction between various reactions that occur simultaneously during the hydrotreating of industrial feedstocks.

In the contact synergy model proposed by Delmon [104, 114, 1151, it is assumed that MoS2 and Co9S8 exist as separate crystallites in contact with each other, and that the role of the promoter (Cogs8) is to activate and provide hydrogen atoms to MoS2. These spilled-over hydrogen atoms would create reduced centers on the MoS, surface, which would be the catalytically active sites.

In the Co-Mo-S (or Ni-Mo-S) phase model proposed by Topsge and co-workers [ 116- 118,123,1241, the promoter atoms (Co or Ni) are located at the edges of MoSJike structures in the plane of Mo cations. The relative amount of Co atoms present as Co-Mo-S phase was shown to correlate linearly with HDS activity [122, 1251. Some possible locations and coordination of the Co atoms decorating the edges of MoS, slabs are illustrated in Fig. 12. The evidence of the Co-Mo-S phase has subsequently been carefully examined by various techniques, and at present the location of the promoter ion as described by Topsge and co-workers is widely accepted.

Although the catalytically active sites involved in hydrotreating reactions are generally viewed as sulfur vacancies (or anionic vacancies), it is still a matter of debate whether both hydrogenation and hydrogenolysis occur on the same type of site or on different types of sites.

Various kinetic studies with model compounds have indicated the existence of two distinct types of catalytic sites, one responsible for hydrogenation and the other one responsible for hydrogenolysis of the heteroatom [81, 127-1291. Several authors have reported that H2S significantly enhances HDN [130- 1341 but moderately inhibits hydrogenation [132, 135- 1381. Nitrogen bases have also been reported to inhibit the HDS, HDN,

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0 s in top layer

I - .

S in bottom layer

O M 0 . c o

FIG. 12. Co-Mo-S model indicating some possible locations of Co at the different MoS, edge planes. The singly bonded sulfur atoms, coordinated to Co,, Co, and Co;, are located in the Mo plane [103].

and AH reactions differently [81,135, 139- 1421. Different inhibiting effects of H2S on hydrogenation and hydrodeoxygenation (HDO) of phenols have also been noticed recently [143]. On the other hand, no inhibiting effect was observed on HDS of thiophene [81] and HDN of 2,4-dimethyl pyridine in the presence of aromatic compounds [142].

The existence of two types of catalytic sites has been proposed to explain these observations [Sl, 126, 127, 130, 143, 1441. Based on a systematic study of hydrogenation of aromatic compounds and hydrogenolysis of S, N-, and 0-containing model compounds, Geneste and co-workers [79-811 showed that hydrogenation reactions were mainly affected by the aromatic properties of the molecules and not by the nature of the hetero- atoms. By contrast, hydrogenolysis reactions were found to depend on the nature of the heteroatom. It was suggested that hydrogenation and hydrogenolysis reactions could proceed via different adsorption mechanisms: hydrogenation through horizontal n-adsorption and hydrogenolysis through vertical adsorption by the heteroatom [79-811. These authors also proposed a dual-site mechanism involving Mo or W atoms at different oxidation levels, the higher oxidation state for the hydrogenation site and the lower one for the hydrogenolysis site [79, 811. Furthermore, based on theoretical calculations and experimental results on adsorption and reactivity of model S, 0, and N aromatic molecules, these authors demonstrated that hydrogenation is favored by strongly electron-donating substituents, and hydrogenolysis by slightly electron-donating (or electron-withdrawing)

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HYDROGENATION

substituents (Scheme 6). Concerning the nature of these two sites, these authors argue that Mo or W atoms with 3 sulfur vacancies (higher oxidation state) at the corners are primarily responsible for hydrogenation through n-adsorption, and that the hydrogenolysis site could be an edge site with two sulfur vacancies [81] .

However, in apparent contradiction with the above conclusion, Topsoe and co-workers [103, 1451 argue that for a given catalyst, HDS, HDN, and AH reactions can occur on the same sites. Their kinetic model assumes that the reactants, intermediates, and products chemisorb competitively on the same sites, and that under typical reaction conditions, a majority of the surface sites are covered by the atomic species H-*, S-*, and N-*, and not molecular species. Accordingly, one may visualize the HDS, HDN, and AH reactions as taking place in vacancies in a mixed surface-sulfide-hydride and nitride phase.

Based on kinetic data on simultaneous HDS, HDN, and A H reactions over a series of different, sulfided hydrotreating catalysts such as Mo/A1,03, Ni-Mo/A1,03, Ni-Mo-P/Al,O,, and Co-Mo/Al,03, they suggested that the major effect of the Co and Ni promoter atoms lies in the lowering of equilibrium constants for S-* and N-* formation. As a result, the promoted catalysts will have a larger fraction of free sites under identical conditions.

In Co- or Ni-promoted hydrotreating catalysts (e.g., Co-Mo/Al,O,, Ni-Mo/A1203), surface sites related to both promoted MoS, edge sites (Co- Mo-S) and unpromoted MoS, edges will be present. Both of these may catalyze hydrotreating reactions, and in general the total activity ktot can be expressed as [123]:

HYDROGENOLYSIS

where kCoedge and kMoedge are the specific activities for promoted and unpromoted edge sites, respectively. @Coedge is the coverage of the MoS, edges by promoter atoms, and &dge is the total concentration of MoS, edge atoms in the catalyst. In general, kcoedge is larger than kMoedge, but the extent

CATALYST

e--WITHDRAWING-SITE e’ DONATI NG-SITE I SCHEME 6. Dual-site mechanism proposed for hydroprocessing of C6H5-

X over sulfided NiMo HR 346 catalyst at 340°C and 7 MPa H,.

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depends on the reaction. Thus, unless OCoedge is very small, the catalysis is expected to be dominated by the promoted sites.

Some authors [14, 130, 132, 144, 1461 have suggested that hydrogenation sites on the catalyst surface can be transformed into hydrogenolysis centers by the adsorption of H2S (Scheme 7). This means that one type of sulfur or anion vacancy present on sulfided catalysts is enough to explain the results. The distribution sites I and I1 would depend essentially on the sulfidation state of the catalyst and on the partial pressure of H2S. Con- sequently, the effect of H2S on the relative rates of different reactions would be expected to depend on the H2S partial pressure during the reaction [14, 103, 147, 1481.

Enhancement of HDN reaction rate at low H2S partial pressures and inhibition at higher H2S partial pressures have been predicted by the model proposed by Topsae and co-workers I1031 based on one type of site for different reactions (Fig. 13), which is in good agreement with the experimental results reported in the literature by various workers [128, 130, 132, 147, 149, 1501.

Interconversion of hydrogenation and hydrogenolysis sites on hydrotreating catalysts is now widely accepted. For kinetic modeling of hydro-

SCHEME 7

FIG. 13. HDN conversions vs. the H,S/H, ratio [103]. Symbols represent observed data by various workers: 0, Satterfield et al.; 0, Massoth et al.; A, Topscbe et al. The solid line represents the data predicted by the kinetic model.

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treating reactions, one kind of catalytic site could therefore be assumed, as suggested by Tops@e et al. [103].

Recently, Vivier et al. [144] have suggested that two categories of site I , namely promoted and nonpromoted, are possible. After H2S adsorption both would lead to sites 11, but only promoted sites I are assumed to be really active in hydrogenation reactions.

Nitrogen compounds as well as H2S have been reported to inhibit the hydrogenation of aromatic compounds strongly (136, 137, 142, 143). These compounds are probably more strongly adsorbed on the hydrogenation centers than on other centers that can catalyze the essentially hydrogenolysis reactions. Competition between different nitrogen, sulfur, and aromatic compounds present in the feed toward adsorption on hydrogenation sites is possible, and the preference for adsorption in such a competitive environment depends largely on the values of adsorption strengths of different compounds. The following sequence has been proposed recently by Vivier et al. [144] for the adsorption strength on hydrogenation centers:

R-C6H5 < NH, < aniline, H2S << decahydroquinoline

< 1,2,3,4-tetrahydroquinoline = quinoline

This is basically consistent with the previous findings by La Vopa and Satterfield [ 1401, except for decahydroquinoline.

The above sequence indicates that inhibition of aromatics hydrogenation by quinoline and the products of its hydrogenation is stronger than that of ammonia, H2S, and alkyl anilines; and that ammonia is a stronger inhibitor of hydrogenation than hydrogen sulfide.

Thus, hydrogenation of aromatics in real feeds under industrial hydrotreating conditions would be inhibited by various organic sulfur and nitrogen compounds present in the feed as well as by H2S and NH, produced by HDS and HDN reactions. These inhibiting effects coupled with thermodynamic equilibrium limitations (under the normal operating range of hydrotreating) make aromatic reduction in industrial feeds (e.g. , diesel) more difficult.

Modeling of the inhibition by nitrogen compounds was attempted by several workers using Langmuir-Hinshelwood-type rate equations [151, 1421, and the adsorption parameters estimated from such models were found to be consistent with the above sequence.

In the case of the inhibition by sulfur compounds, especially H,S, modeling of the inhibition is complicated since H2S adsorption can modify the catalyst surface leading to interconversion of catalytic sites, as mentioned earlier.

In general , quantitative information on ,adsorption parameters char- acterizing inhibition by different types of sulfur, nitrogen, and oxygen compounds is lacking.

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C. Nature of Catalytic Sites in Sulfur-Tolerant Noble Metal Catalysts

Adsorption and hydrogenation of aromatic hydrocarbons on group VIII metals have been studied extensively and the subject has been reviewed by Moyes and Wells [152] and by Garnett [153]. It is now generally agreed that aromatic hydrogenation on these metals involves a .rr-complex adsorption mechanism.

Group VIII metal catalysts used in hydrogenation reactions are known to deactivate rapidly by adsorption of sulfur-containing species that are present in industrial feedstocks [154, 1551. Sulfur is generally considered to be a poison for hydrogenation catalysts, and extensive research has been undertaken to develop sulfur-tolerant metal catalysts. Discussions in this section are mainly focused on noble metal (e.g., Pt and Pd) on Y-zeolite catalysts that have received increasing attention in recent years as important industrial catalysts for aromatics hydrogenation [loo, 101, 197, 2011.

An important and interesting property of these catalysts is their high sulfur tolerance [98, 156-1611. It is reported that the metals are present in the form of highly dispersed small particles or clusters within the zeolite cavities. The zeolite modifies the electronic state of the metal particles (Pt or Pd), which in turn changes the strength of the Pt-S bond.

The resistance to sulfur poisoning is suggested to come from the small bonding energy of the electron acceptor sulfur atoms with the electron- deficient Pt or Pd clusters [156, 1621. Hence, the higher the electron deficiency, the better the sulfur resistance. This has been substantiated by investigations of the electronic structure of Pt and Pd clusters in Y-zeolites [98, 162, 157, 1631. The electron-deficient nature of Pt and Pd aggregates in zeolites has been firmly established by several techniques such as IR, electron spin resonance (ESR), x-ray photoelectron spectroscopy (XPS), extended x-ray absorption fine structure (EXAFS) , and Fourier-transform infrared (FTIR) [162-1701. The zeolite protons are proposed to be responsible for the formation of electron-deficient metal particles, as follows [168]:

n[Pd2+ + 202,;;,,] + nH2+ Pdf + 2nH-OWall

The electron deficiency of the metal should have an adverse effect on the resistance to poisoning by electron-donor molecules. This was indeed shown to be the case for poisoning of Pt in Y-zeolite with ammonia [162, 171- 1731.

The activity, selectivity, and resistance to poisoning of these catalysts depend to a large extent on the location, size/dispersion, and structure of the metal particles in zeolites, which in turn are strictly controlled by the preparation and pretreatment conditions.

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The preparation of noble metal/zeolite catalysts normally consists of three main steps: (i) metal loading, (ii) calcination, and (iii) reduction. The noble metals are usually introduced by ion exchange or impregnation by pore saturation using the tetramine salt [e.g., Pt(NH,),(NO,),] solutions, The main purpose of the calcination and reduction steps is to decompose the ammine ligands attached to the metal, to remove the physisorbed water, and to reduce the metal ions to metal atoms.

Several studies have shown that variations in catalyst calcination and reduction procedures have a major effect on the dispersion and location of the metals as well as on the activity and selectivity of the Pt and Pd/Y- zeolite catalysts [156, 172, 174-1841.

A schematic description of the processes that can occur during calcination and reduction of Pd(NH,),2+ ions in NaY zeolites is shown in Fig. 14, which includes the following: (1) destruction of ligands L, (2) migration of complexed or bare metal ions from supercages to sodalite cages and the reverse process, (3) reduction of ions in various locations and different degrees of complexation, (4) reoxidation of metal atoms in sodalite cages, and (5) migration of these atoms to supercages, where they might be attached to existing clusters, forming grape-shaped particles.

The metal particle size after reduction is dictated by the location and coordination of the metal ions in the calcination step. Accordingly, the conditions of calcination and reduction steps should be carefully optimized to retain well-dispersed metal particles in the large supercages of Y-zeolite. Metal dispersion will be low and metals will be located in the inaccessible

Supercages Sodalite Cages

FIG. 14. Fundamental processes in metal/NaY synthesis. C: calcination in 0,; R: reduction by H,; IM: metal ion migration; AM: atom migration; M(L): metal ligand [157].

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60- s z 0 50- I- s 3 40- I- a u) 13 30- F s 20-

10- a

sodalite cages or on the external surface of zeolites if calcination is carried out above or below the optimal temperature.

Zeolite-supported noble metal catalysts that might tolerate several hundred parts per million sulfur could be prepared by careful optimization of the preparation procedures. The performance of such catalysts is discussed in the next section, dealing with industrial aspects of the aromatic hydrogenation process.

VI. INDUSTRIAL ASPECTS

In view of stringent environmental regulations with regard to aromatics level [as low as 5%(v)] in diesel fuel, aromatics reduction has become a key upgrading parameter in processing middle distillates. Intensive research efforts have been made in recent years to develop catalysts and processes for producing low-aromatic diesel fuel. Several workers [87, 88, 185- 1881 have attempted to optimize the process conditions or maximizing aromatic hydrogenation using existing hydrotreating catalysts, that is, sulfides of Co- Mo, Ni-Mo, and Ni-W supported on y-alumina. Because of their poor activity, such catalysts require high temperatures leading to an equilibrium shift favoring the reverse reaction (Fig. 15). Very high hydrogen partial pressures are required at such high temperatures for shifting the thermodynamic equilibrium in favor of the forward reaction. In general, severe operating conditions such as high temperatures, low space velocities, and high pressures were found to be necessary to achieve acceptable aromatic reduction.

Diesel blending components are generally in the boiling range 220"-

70

O i I I I I I 580 600 620 640 660 680

TEMPERATURE, K

FIG. 15. Aromatic saturation as a function of reactor temperature and pressure on a Middle East heavy gas oil: m, 4.5 MPa; +, 6.5 MPa; *, 12.5 MPa [197].

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350°C with plus or minus some front or back end. They contain mainly paraffins, aromatics, naphthenes, and aromatics with some sulfur- and nitrogen-containing compounds. The carbon number of the component hydrocarbon structures is generally in the C12-C25 range. The aromatic stream may contain several types of mono-, di-, and condensed polyaromatic structures, and their hydrogenation reactivities are considerably different. First-ring hydrogenation of condensed polyaromatic hydrocarbons is most favored kinetically and occurs at relatively low temperatures and pressures.

For example, during hydrotreating of gas oil over a commercial Ni- Mo/A1,0, catalyst, Matarrese et al. [87] noticed that polyaromatics are hydrogenated to a considerable extent at 320"C, a temperature at which monoaromatics only start reacting. An increase of temperature and pressure did not have any significant effect on polyaromatic hydrogenation, whereas monoaromatic conversion increased sharply up to 360°C and then decreased due to thermodynamic equilibrium limitation which favors the reverse reaction at higher temperatures. It should be remembered that polyaromatics are saturated ring by ring in successive steps.

Table 7 summarizes the operating severities reported by various research groups for aromatic reduction in diesel and gas oils in single-stage reactors. The data presented in this table clearly show that achieving low levels of aromatics by single-stage hydrotreating over sulfided Ni-Mo- or Ni-W-type catalysts requires very low space velocities and very high hydrogen partial pressures. Addition of parallel reactors requiring additional capital for the new equipment would be needed to allow lowering of the space velocity without decreasing throughput.

At low temperatures and high space velocities, the amount of monoaromatics in the product has been found to be higher than that originally present in the feed. This is not unexpected since every mole of triaromatic compound that is saturated would add a mole to the diaromatics, and every mole of diaromatic compound hydrogenated would add a mole to the monoaromatic category, which is hydrogenated at a slower rate than the corresponding diaromatic compound.

Thus conversion of the monoaromatics is the key step in the production of low-aromatic diesel fuel by hydrogenation. Since the reaction is primarily controlled by kinetics at lower temperatures, it would be possible to achieve high conversion by using highly active catalysts that could promote aromatic hydrogenation reactions at relatively low temperatures and pressures.

Extensive research has been undertaken by various catalyst manufac- turing companies in recent years to develop more active hydrogenation catalysts. Information available on current research indicates that highly active hydrogenation catalysts, utilizing supported group VIII metals such as Pt, Pd, Ru, and Ni, may significantly improve the efficiency of these processes [39, 189-1961. Aromatic saturation reactions may be carried out at thermodynamically favorable conditions. However, the stability of these

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L

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supported metal catalysts in processing sulfur-containing feedstocks could be a problem. Thus, for example, a silica-supported nickel catalyst was found to exhibit an initial activity much superior to that of conventional hydrotreating catalysts, but its tolerance to sulfur poisoning was low and inadequate for commercial use [39, 1901. This problem might be alleviated through the use of sulfur-tolerant noble metal catalysts. The performance of such catalysts has been reported in patent literature. Supported RuS, catalysts have been reported to be highly active for hydrogenation in an H,S environment [191, 192, 194-1961. However, for catalysts containing Pt, Pd, and Ni, the sulfur and nitrogen in the feedstocks should be reduced to certain minimum levels by hydrotreating prior to processing over these noble metal catalysts. As a result, two-stage processing technologies have been developed and their performance in producing diesel fuels at the specified low-sulfur [loo, 101, 1971 and low-aromatics limits has been dem- ons trated.

A two-stage process [loll has been developed by IFP for upgrading LCO by aromatics hydrogenation. In the first stage, deep hydrodesulfurization reduces the sulfur and nitrogen contents to much less than 50 wppm and 1 wppm, respectively. In the second stage, deep aromatic hydrogenation takes place at a relatively low hydrogen partial pressure with a very active catalyst that tolerates the small amounts of residual sulfur.

A similar two-stage process has been reported by Shell [201]. In the first stage, an Ni-Mo/Al,O, catalyst (C 424) is used for deep HDS and HDN, after which the product is stripped to remove H2S and NH3, and is processed over a noble metal catalyst (e.g. , Pt on amorphous silica: C 614) for deep aromatic hydrogenation. The second-stage catalyst can tolerate organic sulfur levels of about 5-10 wppm.

Criterion catalyst company has introduced a new catalyst technology jointly with Lummus Crest Inc. for aromatics conversion at low pressure [198]. The process known as SynSat (short for synergetic saturation) is claimed to be an economically attractive route to aromatic reduction in diesel fuels. The SynSat reactor design utilizes both concurrent and coun- tercurrent liquidlvapor contacting in separate catalyst beds. The design is geared to optimize both hydrogen partial pressure and reactor temperature profiles. The reactors can either be added to existing hydrotreaters or built as new units. The performance data [198], however, indicate that high hydrogen partial pressures and very low space velocities (e.g., 0.3 h-l) would be required to reduce the aromatics from 29 wt% by about 60% (to 12 wt%) in a feed blend containing 67% virgin gas oil and 33% coker gas oil.

Criterion has also developed a two-stage process [loo] using conventional hydrotreating catalysts in the first stage to reduce the sulfur and nitrogen levels to less than 10 wppm and using a noble metallzeolite catalyst for deep aromatic hydrogenation (>95%) in the second stage at relatively low temperatures (260-315°C) and pressures (4.2 MPa), and high space

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velocities (LHSV 3-4 h-l) . At higher sulfur (1200 wppm) and nitrogen (200 wppm) levels in the feedstock, hydrogen partial pressure as high as 10.5 MPa was found to be necessary to achieve 85% aromatic hydrogenation in this process.

Recently, Haldor Topsoe has developed new technologies for moderate and deep aromatic hydrogenation in diesel fuel. Using a conventional Ni-Mo catalyst in the first stage and a sulfur-tolerant noble metal catalyst (TK 908) in the second stage, 85% aromatic saturation was achieved on a feed containing 1.7 wt% sulfur, 225 wppm nitrogen, and 36 wt% aromatics [197,199,202] at low H2 pressures (5.0 MPa) and low temperatures (315°C). The relative reactor volume in this process was 50% lower as compared with single-stage operation using a conventional Ni-Mo catalyst despite lower operating pressure and temperature. On another feedstock, stable operation was demonstrated in a 6000-h test producing diesel with max 5% aromatics. For moderate levels (25-50%) of aromatics saturation, a two- stage process with NiMo catalyst in the first stage and Ni-W in the second stage has been recommended.

VII. SUMMARY AND CONCLUSIONS

New environmental regulations governing diesel fuel specifications limit the aromatic content to low levels, and as a result processes and catalysts for aromatic hydrogenation in middle distillates have received renewed interest. In this paper we have reviewed the literature regarding the state of aromatics hydrogenation (AH) processes and catalysts. Fundamental studies concerning the kinetic, thermodynamic, mechanistic, and catalytic chemistry aspects of aromatic hydrogenation reactions have been examined critically. Recent scientific literature on the nature of active hydrogenation sites in hydrotreating catalysts has been reviewed with particular emphasis on the interaction between AH, HDS, and HDN reactions. The factors that influence the reactivity of aromatic compounds have also been addressed. Special attention has been paid to new developments in catalysts (including sulfur-tolerant noble metal catalysts) and processing technologies for reduction of aromatics to acceptable low levels in the diesel blending streams in the light of stringent environmental regulations.

Unlike HDS and HDN, aromatic hydrogenation is more difficult to achieve over conventional hydrotreating catalysts under conditions that are normally used for lowering sulfur and nitrogen levels, and the reactions are controlled by thermodynamic equilibrium. Conversion of di- and triaromatics to monoaromatics occurs even at low severity, but the monoaromatics remain nearly unchanged. Hydrogenation of the monoaromatics is the key step in the production of low-aromatic diesel fuel. With conventional hydrotreating catalysts this can only be achieved at high hydrogen pressures and low space velocities, generally requiring new and expensive equipment.

Aromatic saturation reactions may be carried out under both kineti-

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cally and thermodynamically favorable conditions using highly active noble metal catalysts. However, the sulfur and nitrogen levels of the feedstock should be reduced to certain minimum limits by hydrotreating prior to processing over these catalysts. As a result, two-stage processes for com- bined deep HDS and AH have been developed. Production of diesel fuels with less than 5%(v) aromatics and less than 10 wppm sulfur is possible in such two-stage processes.

Reaction kinetic studies with model aromatic compounds have pro- vided valuable insight into the adsorption and reaction mechanisms. In industrial feeds, however, several types of aromatics are present, whose AH reactivities differ considerably.

The polyaromatic hydrogenation occurs ring by ring in successive steps which are all in equilibrium. The kinetic models for aromatic hydrogenation in real industrial feeds should take the reversible consecutive reactions into consideration. The use of Langmuir-Hinshelwood-type rate equations for representing hydrogenation kinetics of industrial feeds is complicated. Sim- ple power law models have been used by most researchers to fit kinetic data.

As regards the nature of active hydrogenation sites, although our understanding of the nature of surface sites and their functions has improved remarkably in recent years, there is still considerable debate as to whether both hydrogenation and hydrogenolysis occur on the same type of site or on different types of site. No clear explanation is available in the literature of the opposite trends in the relative rates of hydrogenation of alkyl benzenes over metal (e.g., Pt, Pd, Ni) and metal sulfide (e.g., MoS,, Co-Mo-S, Ni-Mo-S, Ni-W-S) catalysts. It is possible that the adsorption mechanism of aromatic molecules on the two types of catalysts is entirely different so that the hydrogenation of the aromatic ring is enhanced in one type of catalyst and suppressed in the other by the presence of electron- donating groups in the aromatic ring. The question still remains open and obviously more research is needed on this subject.

A wealth of information available on sulfur-tolerant noble metal catalysts including the peformance of such catalysts in the hydrogenation of aromatics in industrial feedstocks has been discussed in this review. Since the resistance to sulfur poisoning is suggested to result from electron deficiency of the noble metals, the presence of electron-donor molecules such as NH3 or organic nitrogen compounds in the feed would have an adverse effect. In other words, the sulfur resistance property of the catalyst may be lost if appreciable amounts of nitrogen compounds are present in the feed. More research is needed for developing both sulfur- and nitrogen-resistant catalysts.

Finally, aromatic hydrogenation in middle distillates will grow in importance in the petroleum refining industry and the corresponding catalysts and processes will continue to improve steadily. A fundamental understanding of the nature of the catalytic sites, reaction mechanisms, and chemical nature of different types of aromatic compounds present in industrial feeds

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and their reactivities will facilitate more efficient use of the existing catalysts as well as provide a basis for developing improved catalysts and processes.

ACKNOWLEDGMENT

The authors would like to thank Henrik T o p s ~ e for helpful advice and discussions.

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