Hydrotreatment and Hydrocracking of Oil Fractions (Studies in Surface Science and Catalysis)

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Author: G.F. Froment | B. Delmon | P. Grange

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studies in Surface Science and Catalysis 106 HYDROTREATMENT AND HYDROCRACKING OF OIL FRACTIONS

Technologisch Instituut

Studies in Surface Science and Catalysis Advisory Editors: B. Delmon and J.T. Yates Vol.106

HYDROTREATMENT AND HYDROCRACKING OF OIL FRACTIONS

Proceedings of the 1st International Symposium/6th European Workshop, Oostende, Belgium, February 17-19,1997 Editors G.R Froment Universiteit Gent, Gent, Belgium B. Delmon Universite Catholique de Louvain, Louvain-La-Neuve, Belgium P. Grange Universite Catholique de Louvain, Louvain-La-Neuve, Belgium

1997 ELSEVIER Amsterdam — Lausanne — New York — Oxford — Shannon — Tokyo

ELSEVIER SCIENCE B.V. Sara Burgerhartstraat25 P.O. Box 2 1 U 0 0 0 AE Amsterdam, The Netherlands

ISBN 0-444-82556-8 © 1997 Elsevier Science B.V. All rights reserved. No part of this publication may be reproduced, stored in a retrieval system or transmitted in any form or by any means, electronic, mechanical, photocopying, recording or otherwise, without the prior written permission of the publisher, Elsevier Science B.V., Copyright & Permissions Department, RO. Box 521,1000 AM Amsterdam, The Netherlands. Special regulations for readers in the U.S.A. - This publication has been registered with the Copyright Clearance Center Inc. (CCC), 222 Rosewood Drive, Danvers, MA 01923. Information can be obtained from the CCC about conditions under which photocopies of parts of this publication may be made in the U.S.A. All other copyright questions, including photocopying outside of the U.S.A., should be referred to the copyright owner, Elsevier Science B.V, unless otherwise specified. No responsibility is assumed by the publisher for any injury and/or damage to persons or property as a matter of products liability, negligence or otherwise, or from any use or operation of any methods, products, instructions or ideas contained in the material herein. This book is printed on acid-free paper. Printed in The Netherlands

COHTBNTS Introduction

xi

SCS7 WO^TM tSCTVWBB Processes and catalysts for hydrocracking of heavy oil and residues F. Morel, S. Kressmann, V. Harle and S. Kasztelan

1

An improved process for the production of environmentally friendly diesel fuels J. Grootjans and C. Olivier

17

Hydroprocessing to produce reformulated gasolines - the ISAL*"" process G.J. Antes, B. Solari and R. Monque

27

Molecules, catalysts and reactors in the hydro-processing of oil fractions W.H.J. Stork

41

Simultaneous HDN/HDS of model compounds over Ni-Mo sulfide catalysts L. Zhang and U.S. Ozkan

69

Kinetics of the catalytic removal of the sulphur components from the light cycle oil of a catalytic cracking unit G.F. Froment, G.A. Depauwand V. Vanrysselbergtie

83

A review of catalytic hydrotreating processes for the upgrading of liquids produced by flash pyrolysis R. Maggi and B. Delmon

99

Dual-functional Ni-Mo sulfide catalysts on zeolite-alumina supports for hydrotreating and hydrocracking of heavy oils H. Shimada, S. Yoshitomi, T. Sato, N. Matsubayashi, M. Imamura, Y. Yoshimura and A. Nishijima

115

VI

ORAL COMMDI^CAlllOHS CATALYTIC ASPECTS Hydrodracking of Cio hydrocarbons over a sulfided NiMo/Y zeolite catalyst J.L Lemberton, A. Baudon, M. Guisnet, N. Marchal and S. Mignard

129

Novel hydrotreating catalysts based on synthetic clay minerals B. Leiiveld, W.C.A. Huyben, A J. van Dillen, J.W. Geus and D.C. Koningsberger

137

Influence of the location of the metal sulfide in NiMo/HY hydrocracking catalysts D. Cornet, M. El Qotbi and J. Leglise

147

Acidity induced by H2S adsorption on unpromoted and promoted sulfided catalysts C. Petit, F. Mauge andJ.-C. Lavalley

157

Organo metallic siloxanes as an active component of hydrotreatment catalysts I.M. Kolesnikov, A.V. Yablonsky, M.M. Sugungun, S.I. Kolesnikov and M. Y. Kilyanov

167

Alumina supported HDS catalysts prepared by impregnation with new heteropoli-compounds A. Griboval, P. Blanchard, E. Payen, M. Fournierand J.L. Dubois

181

Genesis, characterizations and HDS activity of Mo-P-alumina based hydrotreating catalysts prepared by a sol-gel method R. Iwamoto and J. Grimblot

195

Effects of ethylenediamine on the preparation of HDS catalysts : comparison between Ni-Mo and Co-Mo based solids P. Blanchard, E. Payen, J. Grimblot, O. Poulet and R. Loutaty

211

Creation of acidic sites by hydrogen spillover In model hydrocracking systems A.M. Stumbo, P. Grange and B. Delmon

225

Application of ASA supported noble metal catalysts in the deep hydrodesulphurisation of diesel fuel H.R. Reinhoudt, R. Troost, S. van Schalkwijk, A.D. van Langeveld, S.T. Sie, H. Schuiz, D. Chadwick, J. Cambra, V.H.J, de Beer, J.A.R. van Veen, J.L.G. Fierro and J. A. Moulijn

237

Reactor runaway in pyrolysis gasoline hydrogenation E. Goossens, R. Donker and F. van den Brink

245

vu Surface property of alumina-supported Mo carbide and its activity for HDN T. Miyao, K. Oshikawa, S. Omi and M. Nagai

255

The design of base metal catalysts for hydrotreating reactions; temperature programmed sulphidation of NiW/AlaOa catalysts and their activity in the hydrodesulphurisatlon of thiophene and dibenzothiophene H.R. Reinhoudt, A.D. van Langeveld, R. Mariscal, V.H.J, de Beer, J.A.R. van Veen, S.T. Sie and J.A. Moulijn

263

THEORY AND CATALYTIC DEACTIVATION Surface science models of CoMoS hydrodesulfurization catalysts A.M. de Jong, V.H.J, de Beer, J.A.R. van Veen andJ.W. Niemantsverdriet

273

Molecular mechanics modelling of the interactions between M0S2 layers and alumina or silica support Ph. Faye, E. Payen and D. Bougeard

281

In-situ FT-IR study of NO adsorbed on C0-M0/AI2O3 sulfided at high pressure (

11

300

15

400

33 4. FCC GASOLINE POST-TREATMENT Extensive Hydrotreating of the total worldwide gasoline production capacity of more than 2x10^ barrels per day would be expensive.^ However, inspection of the gasoline process components indicates that approximately 98% of the sulfur in the pool is present in the FCC naphtha (Figure 5). Further, an analysis of one typical sample of FCC gasoline indicated that most of the sulfur is in the heavier cuts of the naphtha (Figure 6). Figure 6 also indicates the nature of the sulfur molecules present in the various cuts. Similar results confirming these observations have been reported elsewhere."^

HeayFOCrsbphtha

(13?-22?Q

U^FOCI^phtha (C^13?Q ,^2%

Oi^^-^

Figure 5. Sulfur Contribution in Gasoline Pool Hydrotreating the entire feedstock to the FCC unit is an attractive choice in terms of the ultimate FCC gasoline product; the tangible benefits are higher overall gasoline yields, lower sulfur in the gasoline, and lower SOx emissions. However, capital investment for feed desulfurization is estimated to be at least $20 to 40 million greater than for posttreatment. Thus, an FCC gasoline posttreatment process becomes the most likely scheme for achieving sulfur reduction in the pool.

34

8000

C2SH CjSH

C3SH

C4SH

C5SH ^

CH3

^C^Hg

^C3H,Q^|

s

6000 B

^ 3

4000

C/3

2000

0

L

FBR 109 138 151 169 194 214 237 271 286 324 347 392

EP

TBP Cutpoint, °F Figure 6. Sulfur Distribution in FCC Gasoline The various hydrotreating reactions that take place in posttreatment processing are: •

Sulfur compounds + H2 -^ H2S + Paraffins (thiophene > benzothiophenes > alkyl substituted benzothiophenes)

•

Olefins + H2 ^ Paraffins

•

Aromatics + H2 ^ Naphthenes

•

Paraffins + H2 -^ Cracked products -^ C4"

•

Paraffins ^ Iso-paraffins

In addition to the beneficial desulfurization reactions, octane loss may be a potential problem because olefins in particular are hydrogenated (Figure 7). If significant cracking to C4- products takes place, the loss of gasoline yield becomes a major expense for posttreating. Taking these factors into account, octane barrel loss during posttreatment can be significant if the process and catalyst are not optimized.

35

u ILU

•

-3

IZI

Q

I

o §

-6 IZI

-9

• IZI

12 1/;

0

1

5

10

15

1

20

25

30

1 35

Absolute Reduction in Olefin Content (vol% feed Olefins - vol% product olefins) Figure 7. FCC Gasoline Hydrofinishing, Olefin Saturation Results in Octane Loss 5. CATALYTIC PERFORMANCE The conversion of pure compounds is instructive in identifying the catalytic traits of the ISAL^^ catalyst. An ISAL-type catalyst was used. Microreactor studies were run with the pure compounds N-octane (NCg); M-xylene (M-X); 1,3 dimethylcyclohexane (1,3 DMCH); and N-octene (C8=). No nitrogen or sulfur compounds were added. Conditions chosen were in the range expected for the process, with constant temperature of 320°C at 400 psig reactor pressure and 1 LHSV. Microreactor test results are on a weight basis, assuming 100 g of feed. The reactivities are those representing the initial state of the catalyst and as such are extremely high. In actual operation with nitrogen- and sulfur-containing feedstocks, activity will be reduced. Cracking and aromatics production are expected to be reduced to a greater extent.

36 Table 1 exemplifies the catalytic features for the ISAL-type catalyst. Paraffins and olefins are the most reactive, and olefins exhibits the highest rate of conversion. Although isomerization appears to be the most significant reaction for all molecules, the catalyst is able to achieve molecularweight reduction, dehydrocyclization, and recombination of hydrocarbon fi-agments. Table 1. Pure Component Reactivities over ISAL-Type Catalyst NC8

M-X

1,3DMCH

C8=

Unreacted

45.5

75.1

91.6

1.2

Isomerized

26.2

21.6

5.9

29.5

Cyclicized

5.7

2.8

-

14.4

Aromatized

4.1

-

0.2

7.6

Cracked (C7")

18.5

0.5

2.3

47.3

Feedstock Products:

Overall, the catalyst exhibits reactivities that are desirable for FCC naphtha posttreatment: strong hydrogenation capability, isomerization capability, and ability to reduce molecular weight without sacrificing aromatics. As a result, the ability to reduce sulfixr level in the FCC gasoline with minimum impact on RON and MON and minimized Cs^ loss is a real expectation with the optimized catalyst functionalities of the ISAL^^ catalyst. Feedstocks from refineries demonstrate these same aspects of the ISAL^"^ catalyst performance. Figure 8 demonstrates the processing of a heavy virgin naphtha, a Ce^ FCC naphtha, and a Cg^ FCC naphtha over the ISAL™ catalyst. The road octane (RON+MON/2) of the product is shown as a flinction of volumetric liquid yield. The losses of road octane observed with hydrotreating of the FCC gasoline in Figure 7 can be offset by employing the ISAL'r'^ catalyst, as demonstrated in Figure 8. Yield loss may be minimized, and in fact overall production of octane barrels may be increased despite this loss.

37 Table 2 relates experimental data comparing the results of hydrotreating a Cg^ FCC naphtha with those obtained by processing the same stream with the I SAL technology. Significant reduction in the loss of road octane is observed with the ISAL process. Other data has been reported elsewhere. Table 2. Data Comparing Hydrotreating and ISAL Technologies Feedstock:

Product:

Cg^ FCC Naphtha

Hydrotreating Technology

ISAL Technology

Sulflir, wppm: 975

Aromatics + H2 Paraffins —> Aromatics + H2 Naphthenes, Paraffins -h H2 —> Hydrocracked Products Feedstock + Catalyst —> Coked Catalyst + H2 (Catalyst Deactivation) Fig. 2. Main reactions in catalytic reforming.

43

CATALYST

CENTER PIPEPRODUCT

Fig. 3. Radial flow reactor [1].

The reactions overall are strongly endothermic, and hence a series of reactors is used with interstage heating; these reactors are of the radial flow (Fig. 3) rather than axial flow type to reduce the pressure drop. Several commercial processes have been developed along these lines, such as UOP's Platforming, IFP's reforming, Engelhard's Magnaforming, Exxon's Powerforming, Chevron's Rheniforming and Amoco's Ultraforming. Differences between these are discussed in [3]; one of the variations is the envisaged lifetime of the catalyst, ranging from a week to a month (low. pressure operation) in cyclic, fully regenerative processes (Fig. 4a) where an extra swing reactor allows taking out one reactor of the train for regeneration, to semi-regenerative processes (operating at higher pressure), with a life time of about 1 year (Fig. 4b). For thermodynamic reasons one wants to operate the process at high temperature and low hydrogen partial pressure; this also enhances the rate of the difficult dehydrocyclization reaction. High hydrogen pressures, which reduce catalyst deactivation by coking (Fig. 5 [4]), on the other hand, also increase the rate of hydrocracking. Thefirstreformers using (monometallic) platinum catalysts were therefore run at a hydrogen pressure of about 35 bar [3]. The advent of the more stable bi(multi)metallic (Pt/Re/(S)) catalysts has allowed a significant reduction in operating pressure, significantly contributing to an increased yield in the "semi-regenerative" ("SR") reformers, as shown in Fig. 6 [5]. Recent refinements here are the use of multiple catalyst systems (e.g. UOP's R72/R56 [6], or Criterion's Pt/Sn//Pt/Re system, PS7/PR9 [12]), where the front reactors in which the rapid endothermic naphthenes dehydrogenation occurs on average at low temperatures use a different catalyst (with high selectivity/lower stability) than the back-end reactors, in which the slow dehydrocyclization occurs at high temperatures putting higher demands on stability, leading to overall yield improvements, and the use of skewed Pt/Re catalysts, (high Re), intrinsically very stable but sensitive to sulphur, often in combination with

44 (a)

NET HYDROGEN TO REFINERY ,

RECYCLE GAS

& \

m P. p^ l_^ L,^

FLASH DRUM

TO STABILIZER

NET HYDROGEN TO REFINERY ,

RECYCLE GAS

Pk pL

Krarrn L ^ t^

0

& FLASH DRUM

-REGENERATION PIPING TO STABILIZER

Fig. 4. Schematic flow diagram of semi-regenerative and cyclic reforming processes [1].

45

T(°C)

-

@ = Pressure (bar)

/

10

St 1 RON = C

@)

/

/

^^@^^

5 To

1

1

0

1

.

1

3

.

1

5 Time (arbitrary unit)

Fig. 5. Typical influence of pressure on stability [4]. 100 HYDROGEN Q

_J UJ

90

> § 80

'^^_ v:

0^

'^--/\

"

•^.^ '^-.^

C5+ PLATFORMATE

CL

70

143

0—

C C^ • Co + C3 +Â 1 ^

Q

O a.

n—

285

""^

428

570

PRESSURE,psig

Fig. 6. Effect of pressure on yield structure [5],

a sulphur guard bed. Sulphur guard beds appear imperative in the operation of the new, zeolitic type of aromatization catalysts, Pt/K/L, of the Aromax process, that are very selective to lower aromatics [7], However, the basic issue in reforming is that the conditions that favour a higher liquid product and hydrogen yield also lead to a lower catalyst stability. It was therefore a major step forward when the importance of catalyst life was drastically reduced by UOP's introduction in 1971 of the continuously regenerated reformer, "CCR", shown schematically in Fig. 7 (later also IFP introduced a continuous process). Here the catalyst is continuously replaced and regenerated after only days of operation which allows a drastic reduction in operating pressure, currently down to some 3.5 bar, which favours the desired reactions and which suppresses the undesired hydrocracking reaction. A large overall yield increase as shown in (Fig. 6) is the net result, and clearly the majority of the new units are of the low pressure, CCR type (Fig. 8 [8]). Catalyst improvements for a CCR unit will be targeted towards e.g. higher intrinsic selectivities, better regenerability, and high mechanical strength rather than towards the better stability desirable for a SR unit, and hence will follow a different set of rules: these are generally Pt/Sn catalysts. The above clearly illustrates the impact that a change in reaction

46

FUEL GAS

REGENERATED CATALYST

SPENT CATALYST

LIGHT ENDS TO RECOVERY

NET GAS

CHARGE PLATFORM ATE

Fig. 7. UOP Continuous plarforming process [5].

^^H

6h 0)

o

5

4

/ ^ 6 0 ' s & 70s

U

I

2\-

50S&60S

70 s & 8 0 s 1

1 10

1

15

1 20

1 25

1

30

35

PRESS (bars) Fig. 8. Increasing trends in the severity of naphtha reforming processes (Arrow indicates increasing severity of operation) [8].

conditions and reactor concept has by itself, and also how it affects the direction for catalyst developments. 3. HYDROTREATMENT AND HYDROCRACKING OF OIL FRACTIONS 3.1. Introduction As is well known, the first step in the processing of crude oil is the separation by distillation into a number of boiling fractions. The naphtha, with boiling point between 80 and 200°C is the basis for gasoline, and an important chemical feedstock; kerosine (bp between 150 and 250°C) is

47 used for aviation fuel, and gasoil (bp between 250 and 370°C) is used a.o. for diesel (automotive gasoil). All these fractions already have the right boiling point range and "only" have to be brought to specification, in which hydrotreating processes such as hydrodesulphurization and hydrogenation often play a major role. The crude oil also contains a significant amount of components boiling above about 370°C, the residue in the atmospheric distillation, and these can be converted into the valuable transportation fuels by (hydro)cracking, either by processing the entire residue, or (usually) by processing the vacuum gasoil, a fraction with boiling point between about 370 and 520°C that is recovered by vacuumflashingthe atmospheric residue. We will discuss here the reactor aspects of hydrotreating and hydrocracking by going to progressively heavier feedstocks, and then in the next section inventorize the various reactor issues encountered, and the common reasons behind those choices. 3.2. Hydrotreatment of light gases Light gases are produced in the refinery in various processes, such as catalytic cracking, and are often for used further conversion into e.g. gasoline fractions by e.g. alkylation. In most of these applications it is important to minimize the amount of acetylenes and diolefins by selective hydrogenation, processes for which dedicated catalysts have been developed. Thus one may estimate that the removal of the C4, C5 dienes from the alkylation feedstock yearly would save several millions of dollars in terms of reduced acid consumption in the alkylation unit. For cost-effectiveness, it would be very attractive to achieve this hydrogenation applying catalytic distillation technology [9]. In this operation the catalyst is placed inside the (existing) distillation column, requiring a sufficiently high reaction rate under the distillation conditions (obviously temperature, reactant concentration, depth of conversion, hydrogen partialpressure and reaction inhibition and selectivity are all relevant). If appUcable, however, the technology has intrinsic advantages in terms of reaction control, heat of reaction removal, removal of coke precursors and low capital costs. Indeed this technology, well-known from e.g. MTBE production [10], has now found a number of applications in selective hydrogenation of light gases (Fig. 9). In the commercial selective hydrogenation of C4, it is shown in Fig. 10 [11], that three catalytic

>-Distillate Product Hydrogen

•

Hydrocarbon feed

•

> • Bottoms Product Fig. 9. CD Hydro reaction column (ref. [11]).

48

Oligomer free Reflux

>-Offgas Catalyst Bed

Hydrogen Hydrocarbon Feed

^

Distillate Product

Fig. 10. CD Hydro Desulphurization and anti-fouling mechanism [11]. I 1,3-Butadiene

• Butene-1

50

"r 60

70

80

90

Packing Height (*/o of total bed) Fig. 11. Concentration profiles in CD Hydro Catalyst Bed [11].

functions are performed successively: first methyl mercaptan absorbs strongly and reacts with butadiene to form heavy sulphides going to the bottom of the column, then butadiene absorbs and is selectively hydrogenated and finally double bond isomerization occurs (see Fig. 11). Further appHcations of such technology may be expected.

49 Diolefins —> (Ni/S or Pd catalysts, 50-150°C) —> Monoolefins Monoolefins —> (NiMo/S catalysts, 250°C) —> Saturates Desulphurization —>(NiMo/S catalysts, 330°C) Diolefins —> (200°C or higher) —> Rapid fouling Fig. 12. Chemistry of pyrolysis gasoline desulphurization (schematic).

3.3. Hydrotreating of naphtha With the advent of the bimetaUic reforming catalysts, combined with the higher proportion of cracked feedstocks, the demands on naphtha desulphurization have only increased. Improved catalysts are available (see below), and while the technology itself has not changed much, the cracked feedstocks have favoured operation at higher pressures. New catalytic guard materials are available to remove silicon and iron [12] and protect the HDS catalyst; with the cracked feedstocks the recombination of H2S with olefins to mercaptans can limit the maximum operation temperature [13]. The hydroprocessing of pyrolysis gasoline from ethylene crackers requires a special set-up, since these contain large amounts of olefins and diolefins, defining the following chemistry (Fig. 12). The diolefins are especially nasty since they easily give severe fouling in an HDS unit. Therefore the pyrolysis gasoline is first processed at low temperatures over a selective hydrogenation catalyst using multiple catalyst beds to convert the diolefins into mono olefins, with product recycle of hydrogenated gasoil to limit the temperature rise over the reactor, to dilute the reactants and to dissolve formed polymers. In the next reactor the olefins are hydrogenated in a similar fashion, andfinallythe product is desulphurized (Fig. 13 [14]). 3.4. Hydroprocessing of middle distillates The specifications for automotive gasoil have become much more stringent recently, a trend which may very well continue; the sulphur specification in particular is now set to 500 ppm in many places. Apart from this, the Swedish and the Califomian authorities have promoted the introduction of a special grade of diesel, low in aromatics and sulphur, which is now also available in the UK. The basic chemistry of desulphurization, and aromatics hydrogenation in gasoils is shown in Fig. 14.. As is common in processing of oil fractions, in the ever deeper desulphurization of gasoil the most refractory sulphur species remain, which have been identified as the alkyl substituted dibenzthiophenes (e.g. ref. [15]). As a consequence, in the example quoted in ref. [16], a doubling of the reactor volume/halving of the feed rate would be required to go from e.g. 0.1 %S to 0.05 %S. Alternatively the reactor temperature could be increased by some 20°C, but this certainly reduces catalyst life and may also negatively affect product properties such as aromatics level and/or colour stability. The development of ever more active HDS catalysts has helped significantly, but the activity differences between various generations of Co/Mo desulphurization catalysts are about 15-20% (and quite difficult to realize) (Fig. 15), which is clearly by itself not enough to reach the new specification at otherwise unchanged conditions. A new option though apparently not yet commercial is the use of a dedicated ("alkyl-substituted DBT conversion") catalyst in the bottom part of this reactor [17,18]. Thus, at present, with many existing "single-stage" units, the severest sulphur specifications can only be achieved at a severe penalty to the catalyst runlength, especially if the other product properties also have

50 FIRST STAGE

SECOND STAGE

THIRD STAGE

^

LIGHT GASOLINE AROMATICS CONCENTRATE

^ HEAVY GASOLINE

GAS OIL HYDROGEN PYROLYSIS GASOLINE/GAS OIL

Fig. 13. Pyrolysis gasoline hydrotreating unit [14].

Sulphur compounds —> CoMo/S or NiMo/S catalysts —> Desulphurization -range of reactivities -alkylsubstituted DBT's most refractory compounds -inhibition by H2S Aromatics hydrogenation -equilibrium limited (partial pressure H2, 7) -Ni/Mo/S catalysts —> moderate activity, moderately poisoned -noble metals catalysts —> high activity —> strong poisoning by (organic!) S, N compounds Coke formation —> Deactivates CoMo/S and NiMo/S catalysts —> high T, low pp H2 Fig. 14. Chemistry of desulphurization and aromatics hydrogenation in gasoils.

to be met. Only single-stage units that operate at higher hydrogen pressures (favouring reaction rates, product properties at higher operating temperature and catalyst stability) offer significant scope. This holds for desulphurization, but even stronger for aromatics hydrogenation, where thermodynamics limits the use of the higher reaction temperatures. In this situation a process line-up in which a second reactor, with a fresh supply of hydrogen without H2S or NH3, is added, can be attractive ("two-stage" Hne-up). Even for conventional desulphurization catalysts hydrogen sulphide is a catalyst poison, due to competitive absorption, and the normal H2S partial pressures cause a significant loss in catalyst activity [19]. For noble metal catalysts, the detrimental effect of H2S (and NH3) of course is much stronger. A second-stage reactor, however, might employ dedicated, rather H2S sensitive catalysts to further reduce the sulphur levels [17,18].

51 IMPROVEMENTS IN CATALYST Rel.Vol.Activity

TECHNOLOGY Diesel Sulphur S p e c ^ ^ w t

200

150

100

2000

I Sulphur Spec

QRec.Voi.Act

Fig. 15. Activity of generations of HDS catalysts.

As was indicated above, the larger incentive for the two-stage unit is in aromatics hydrogenation using a noble metal catalyst in the second stage. The temperature in this stage must be low so that sufficiently high aromatics conversions are possible by thermodynamics, which puts special demands on catalyst activity. The favoured noble metal catalysts for this are those systems that have a relatively high tolerance for sulphur and nitrogen compounds in the feed; in this quite some progress has been made [16]. Theflowscheme of the integrated two-stage Shell Middle Distillate Hydrogenation process, running since 1992 in Gothenburg, Sweden, is shown in Fig. 16 [16]. Here the second-stage catalyst, at low pressure, can cope with (organic) S and N levels in thefirst-stageeffluent of some 100 and 50 ppm, respectively, and, on top of the aromatics hydrogenation, also achieves a desirable hydrodecyclization. The line-up implemented in the co-current Synsat process described in ref. [20] is similar, while another realization for this is shown in Fig. 17 [19]. An alternative line-up is based on the use of a counter-current second-stage reactor. In the normal co-current reactor the reaction rate always is highest near the top of the reactor (particularly for reactions of order 2 such as HDS) generating high concentrations of H2S already at the top of the reactor, negatively influencing the activity of the remainder of the catalyst (Fig. 18 [21]). In the counter-current line-up, this H2S (and/or NH3) is quickly removed from the reactor without affecting the activity of the bottom part of the catalyst. Hence, the counter-current operation potentially has the largest advantage when the catalyst is very sensitive to gases generated at the top of the catalyst bed, for noble metal catalysts therefore. The counter-current reactor was originally developed by Lummus in their Aerosat technology and later implemented together with Criterion in the Synsat unit in the Scanraff refinery in Sweden [22]. The detailed line-up is shown in Fig. 19, with the second-stage rector with "pure" hydrogen in counter-current operation. It is reported in ref. [22] that the counter-current operation allows a reduction in catalyst volume of 30-35% for reaching the same aromatics conversion relative to conventional all co-current operation (10-15% for a HDS target). Thus the SMDH and the Synsat processes are two options to produce low aromatics diesel oil; the paper by Grootjans [23] addresses the manufacture of this in more detail. (Ref. [16] gives a good listing of catalysts claimed for this duty.)

52 Fresh Gas rR-1 I XI

r*4m\^ FEED

^

I

i ^ ^

'^^P

1 / Jiff

t ^t ^

N

*

^

^

^?

^

^fi>

^

/ ^^^Sx

M M M^ G

TL

G

TL

G

TL

G

G

1 i1 G

Fig. 25. Three level porosity concepts for counter current operation [35].

("red death"). Several options exist to combat this phenomenon, one of which is a special absorption step to remove these species for subsequent disposal [34]. A clear current limitation of the counter-current reactor is the requirement of low pressure drop to avoid flooding of the reactor. To this end large shaped catalyst particles such as 5 mm Raschig rings could be envisaged, possibly also shell type catalysts. Sooner or later however the requirement of low pressure drop is conflicting with a high volumetric catalyst activity and a high catalyst utilization. Sie [35] therefore has proposed alternative solutions, based on the insight that in hydroprocessing the gas/liquid mass transfer is generally not rate limiting. In this way constructions are devised such as shown in Fig. 25 [35], in which the gas/liquid contacting is only periodically intensified. Such constructions, as yet still in the research phase, could allow higher gas rates and extension of the counter-current concept to areas such as hydrocracking, where one needs high gas rates, and also may wish to minimize the poisoning effect of the ammonia generated at the top of the catalyst bed by denitrogenation of the organic nitrogen compounds [21,35]. 3.6. Residue hydroprocessing Instead of recovering the vacuum distillate for conversion into transportation fuels, one can also process the entire residue of the atmospheric distillation, either to obtain a heavy low-sulphur fuel (desulphurization), or, again, to produce more transportation fuels (cracking, also called conversion in the oil industry). The atmospheric residue differs from the vacuum distillate in boiling point distribution (the vacuum residue, with components boiling above 520°C, is included), but even more importantly, in that it contains large amount of contaminant species, such as metals (in particular Ni and V) and asphaltenes (large aromatic compounds), which are very strong catalyst poisons. Furthermore, in the residues the concentration of sulphur and nitrogen containing molecules is higher than in the distillates, while on average these molecules have a lower reactivity. As a consequence of the basic chemical factors, outlined in Fig. 26, residue processing conditions have to be quite severe, high in temperature and in hydrogen pressure.

60 Sulphur compounds HDS —> CoMo/S or NiMo/S catalysts: low reactivity some diffusion effects Ni/V compounds HDM —> CoMo/S, NiMo/S or autocatalytic: low reactivity strong diffusion effects HDM: Ni and V deposition on thie catalyst, poisoning HDS activity. I interstitial deposition (would lead to) —> severe catalyst bed fouling residual compounds cracking —> mainly thermal, not catalytic residual compounds at high T —> severe catalyst coking and fouling hydrogenation of residues —> at a given hydrogenation level asphaltenes precipitate

oo

Fig. 26. Basic chemistry of hydroprocessing of residual oil fractions. PORE SIZE

WIDE

• ASPHALTENE PENETRATION HYDRO CONVERSION ACTIVITY METAL STORA6E

USE

1

TYPE

TOTAL VERY LOW VERY GOOD DEMETALLIZATION CATALYST A

MEDIUM

NARROW

MEDIUM

SHALLOW

MODERATELY HIGH

HIGH

FAIR

POOR

FRONT END TAIL END HYDROCONVERSION HYDROCONVERSION CATALYST CATALYST B

C

Fig. 27. ABC catalyst concept in residue HDS [37].

The reactor concept chosen strongly depends on the processing goal. The emphasis mostly has been on the desulphurization of residues to produce low-sulphur heavy fuel, or on the direct residue hydrocracking. For the desulphurization appHcation one desires a high degree of desulphurization, and therefore high catalyst activities and excellent staging. Fixed bed trickle bed reactors have been an obvious choice, and the technology has been developed by Chevron, Unocal, UOP, Exxon and Shell [36]. Catalyst deactivation was controlled by on the one hand high hydrogen pressures to reduce coke formation, and on the other hand by dedicated catalyst systems favouring demetallization in the front end, and highly stable desulphurization at the back end of the catalyst system (ABC concept. Fig. 27 [37]). In general the catalysts are contained in several large reactors; since the front end, demetallization, catalyst may deactivate more rapidly and the risk of interstitial deposition is to be minimized, sometimes small guard reactors are used that can be alternated, or put off-line; also the introduction of the feedstock at various points lower down the reactor with increasing runlength has been described [38], With the desire to increase cracking in thefixedbed units, by increasing temperature, also the demetallization rate increases

61 Fresh Catalyst Bin

Product to RDS Reactor

High Pressure Catalyst Vessel

OCR Reactor

Feed in

Spend Catalyst Bin Fig. 28. OCR catalyst replacement system [36].

and hence catalyst deactivation increases. The ultimate solutions, therefore, particularly for high metals feedstocks, are continuous HDM catalyst replacement systems, such as developed by Chevron (OCR, on line catalyst replacement, with a counter current operation, the feed moving upflow, Fig. 28 [36]) and by Shell HYCON (co-current downflow in the bunker flow reactors. Fig. 29a, b [37], processing even vacuum residue). They also have the advantage that the HDIVI catalyst leaving the system is fully loaded, contrary to the spent HDM catalyst from afixedbed reactor, where a metals profile will exist over the catalyst bed [36]. In residue hydroprocessing, more than in any other application, guard materials consisting of large, porous, low activity materials are used to catch Fe, scale and salt species before they plug the catalyst bed with its fine particles. Improved catalysts for the trickle bed units should have higher activity for HDM or HDS, higher stability (e.g. in terms of metals uptake capacity), or higher activity for cracking, e.g. by the use of zeolites (some approaches are discussed in refs. [12] and [39]). A recent example from our laboratory is given in Fig. 30 [12]. The main alternative is the use of expanded bed or ebullating bed reactors (see Fig. 31 [40]), which have a larger liquid hold-up, smaller catalyst concentration, with less risk of fouling, and with continuous catalyst replacement. A tight control of the inlet temperature however is critical [41]. The process has been developed in two versions, by ABB Lummus Crest ("LC-Fining") and by HRI ("H-Oil")- The catalyst bed is expanded by the liquid flow that is accomplished by pumping recycle liquid recovered near the top of the reactor to the plenum beneath the bed. These reactors are clearly intended and suited for the high-temperature cracking application rather than for deep desulphurization; important limitations are the lack of staging, including that of different catalysts, inherent to continuous stirred tank reactors. A further development of the concept has been implemented in the Texas City unit of Amoco, where three reactors

62 (a)

1

1

CTS

(CTS= CATALYST TRANSPORT

TO DISTILLATE SECTION

SYSTEM)

(b) FIXED BED REACTOR Fresh Catalyst

MOVING BED REACTOR (BUNKERFLOW)

Liquid > ' Gas

STATIONAR CATALYST BED

CONTINUOUS FLOW OF GAS AND LIQUID

Liquid Gas

CONTINUOUS FLOW OF GAS AND LIQUID

Liquid Gas Spend Catalyst

Fig. 29. Bunker flow system [37].

are used in series, improving staging and in principle allowing the use of optimized catalysts in each separate reactor (Fig. 32 [42]). Improved catalysts in the ebuUating bed units should allow higher conversions (cracldng) without leading to asphaltenesflocculation;some approaches are discussed in ref. [43]. Figure 33 [44] gives an impressive example of an improved catalyst that allows higher conversions without problems as to sediment formation; ref. [45] on the other

63 +50 (+90)

^4 •

^ +40 + ""^ (+72) "d) -30 b (-54) ?

CD

+20

^ (+36) ^^ +10 (+18) Base

RN -400^__^^^..*^—5—-rrW-

-L

^^Q -{XH L-OCHlH

G •D A

GO

H V

C JZ

c

GO

T

AC Fig. 32. Expanded bed reactors in series [42].

0.2

BFDSHDIMENT.WT'/o

0.15 h

STANDARD CATALYST

0.1

0.05 h

NEW CATALYST

( TEX 2710)

_L_

50

55

60

65

70

75

80

85

538 C + CONVERSION,V

o

u

0 5

7

9

11 13 15 17

M0O3, wt% (3% NiO) Pyridine

NiO, wt% (15% M0O3) Piperidine

-a-C5

Figure 2. Effect of loading on pyridine conversion, piperidine and C5 production rates at 360°C. When the effect of Mo loading on hydrogenation and hydrogenolysis rates is examined under the same conditions (Figure lb), we see that Mo does not show the same enhancement effect on piperidine production rate as Ni does. Instead, increasing Mo loading increases C5 formation and pyridine conversion rates and decreases piperidine formation rate. The increase in C5 and p5n:'idine conversion rates with increasing Mo loading becomes more pronounced at 360°C as shown in Fig. 2a. On the other hand, the Ni loading has a negligible effect altogether (Fig. 2b). Since at 360°C, piperidine production rate is controlled by thermodynamic equilibrium, this observation reiterates the conclusion that Ni has no effect on hydrogenolysis steps, as it is already seen at 320 °C through the C5 yield which essentially remains constant regardless of Ni loading. The trends seen in these figures are significant in providing the first clues about the role of Ni and Mo-associated sites. It appears that under conditions where the first hydrogenation step is kinetically controlled, the addition of Ni increases the p3n:-idine conversion rate. When the first step approaches thermodynamic equilibrium, the addition of Ni no longer has a promotional effect, suggesting an assignment of the hydrogenation function to Ni-associated sites. When the role of Mo is considered, we see that it has an enhancement effect on C5 production rates at all times by promoting the piperidine hydrogenolysis reaction. The effect of increased Mo loading on the pyridine conversion rate is mainly an indirect one and is facilitated through the

75 consumption of piperidine in the hydrogenolysis reaction, which drives the first step fiirther to the right. The effect of sulfur compounds (HgS and thiophene) on pyridine HDN is found to be strongly dependent on several factors. When hydrogenolysis of piperidine is the rate determining step, pyridine HDN is enhanced by gas phase thiophene and H2S over both Mo and Ni-Mo catalysts. This enhancement effect is most pronounced over the bimetallic catalyst. Figure 3 gives an example of the effect of sulfur compounds on pyridine conversion and C5 formation rates over the monometallic and bimetallic catalysts at 400°C. In the absence of sulfur compounds, significant levels of piperidine were observed in the product stream, making the C5 production rate much lower than the pyridine conversion rate over both Mo and Ni-Mo catalysts. Under these conditions, the pjrridine/piperidine equilibrium is established and the rate-determining step is the hydrogenolysis of piperidine. When thiophene or H2S is present, however, piperidine is never detected in the reactor effluent, making C5 formation rate equals to pyridine conversion rate. An implication of this observation is that the rate determining step is no longer piperidine hydrogenolysis, but pyridine hydrogenation. The enhancement effect of sulfur compounds is facilitated through an increase in the hydrogenolysis sites. The increased HDN activity is much more evident over the bimetallic catalyst due to the strong promotional effect of nickel on the hydrogenation step. It can be seen from Figure 3 that, the enhancement effect of thiophene on pyridine HDN over the Ni-Mo catalyst is more pronounced than that of H2S. Since HgS is the product from thiophene HDS, it seems that thiophene has a n additional enhancement role t h a n HgS and this additional role is linked to Ni promoter.

Mo

S HDN only

H HDN with thiophene

Ni-Mo

HDN with H2S

Figure 3. Effect of sulfur compounds on the pyridine conversion and C5 production rates (400°C, catalysts: 10%NiO/Y-Al2O3, 20%MOO3/Y-A12O3, 3%NiO15%Mo03/Y-Al203)

76 3.3. Temperature-prograinined reduction and temperature-programmed desorption studies TPD experiments over the sulfided Ni-Mo/y-AlgOg and Mo/y-AlgOg catalysts showed two HgS desorption peaks. For both catalysts, the first peak corresponded to the desorption of weakly adsorbed H2S fiôm the a l u m i n a support. The second H2S desorption peaks had the same on-set temperatures of 300°C and maximum temperatures of 480 and 490°C for the 20% MoO^y-Alfi^ and 3%NiO-15%Mo03/Y-Al203 catalysts, respectively. In addition, the shapes of the second HgS TPD peaks for the two catalysts were very similar. The TPD profile of the bare alumina support did not show the high temperature feature. Based on these TPD results, it is conceivable that the second HgS peak represents the desorption of H2S from the catalytic active sites associated with Mo atoms only. Prior to the H2 TPR of the sulfided catalysts, a degassing treatment at 500°C was performed to remove all the adsorbed H2S species so that any HgS that evolves during the TPR experiment is a result of the reduction reaction and is not due to desorption of H2S which is left on the surface from the sulfidation process. The TPR profiles for bimetallic Ni-Mo/Y-Al203 catalysts were quite different from those of the monometallic Ni/y-Al203, and Mo/y-Al203 catalysts. However, bimetallic Ni-Mo/y-Al203 catalysts with different Ni loadings showed very similar patterns. The detailed results are presented elsewhere [20]. There were two low temperatiu-e H2S peaks at ca. 160°C and 220°C, which could be due to the removal of sulfur from Mo and Ni centers, respectively. ^The ratio of H2S peak areas for Mo- and Ni-associated peaks was slightly larger than 1 for the 3%NiO-15%Mo03/y-Al203 catalyst and about 1 for the 5%NiO-15%Mo03/y-Al203 and 7%NiO-15%Mo03/y-Al203 catalysts. It is conceivable that these two peaks correspond to the formation of sulfur vacancies on the edge planes of Ni-Mo-S phase. The 1:1 ratio for the Mo- and Ni- associated S vacancies for catalysts with high Ni loadings is in agreement with the maximum Ni accessibility to the M0S2 plane that is concluded in the literature [1]. 3.4. Active sites and their catalytic functions Combining the results of our kinetic and characterization studies with some findings in the literature, we propose two major types of active sites promoting HDN of nitrogen heterocycles. Type I: these are hydrogenation sites consisting of sulfur vacancies associated with Mo (type la sites) or Ni in Ni-Mo-S phase (type lb sites). T5rpe II: these are hydrogenolysis sites consisting of Bronsted acid centers associated with Mo atoms only. According to Yang and Satterfield [91, the adsorption and dissociation of an H2S molecule can convert a sulfur vacancy to a Bronsted acid site and a sulfliydryl group (SH), but the adsorption is readily reversible if HgS is removed fiôm the reaction system. The results from our pyridine HDN studies can be explained in terms of these active site assignments. The catalytic job distribution of Ni and Mo associated centers in pyridine HDN can be summarized as Ni-associated sulfur vacancies in the Ni-Mo-S phase (type lb sites) being responsible for hydrogenation steps, whereas the primary function of Mo being to promote C-N bond hydrogenolysis reactions through Bronsted acid sites. For Ni-Mo

77 catalysts, Mo-associated S vacancies are not important for hydrogenation reactions due to their much lower intrinsic activity compared to that of Niassociated ones. The role of sulfur compounds in the pyridine HDN catalytic scheme is envisioned to be multi-faceted: 1) HgS in the gas phase helps maintain a certain content of Bronsted acid sites (type II sites). Thiophene promotes hydrogenolysis of piperidine indirectly via HgS formed during HDS reaction. 2) Thiophene in the gas phase also helps the pyridine hydrogenation step over the bimetallic Ni-Mo catalysts by keeping Ni active sites in an effective form for pyridine hydrogenation reaction, probably by converting double S vacancies to single vacancies. 3) Thiophene in the gas phase does not enhance the hydrogenation step over the mono-metallic Mo catalyst, but inhibits it by reducing the number of available hydrogenation sites. 3.4. HDN of indole in the presence of H^S, benzothiophene, and o-ethylaniline The reaction network of indole HDN based on the proposals from the literature [6, 21-29] and this work is depicted in Figure 4, including the acronyms used for the compounds discussed. There is general consensus that the reaction network in indole HDN starts out with hydrogenation of the heterocyclic ring in a reversible step which leads to indoline formation, dictated by thermodynamic equilibrium under most conditions. Several of the previous reports on indole HDN suggest that o-ethylaniline is the exclusive intermediate toward the formation of hydrocarbons [21-27] following indoline formation. One of the more recent studies on indole HDN which was conducted over a sulfided NiMoP/Y-Al203 catalyst [28] proposed the denitrogenation pattern for indole to be analogous to that found for quinoline [2, 3], i.e., the hydrocarbon products being formed predominantly by complete hydrogenation of both the heterocyclic and the benzene rings prior to the cleavage of C-N bond. There are also suggestions that both routes could be playing an important role in the overall network [6,29]. Indole

Indoline

OEA

r OHI

I OECHA " N

EB

t JLXV.

ECH Figure 4. Indole HDN reaction network.

78 A key question regarding the reaction network of indole HDN is the role of two N-containing intermediates, o-ethylcyclohexylamine (OECHA) and octahydroindole (OHI), since they were not detected in many of the studies reported. Another question that has not been addressed very much is the effect of sulfur compounds. This study was designed to determine the important intermediates and to differentiate among the individual steps involved in the indole reaction network in the presence and the absence of sulfur compounds. Benzothiophene was chosen as the model sulfur-compound in this study. Ethylbenzene (EB) was the only major product in the HDS of benzothiophene (BT) over the Ni-Mo catalyst. The hydrocarbon products from BT HDS were about 99% of EB and 1% of ethylcyclohexane (ECH) throughout the temperature range from 200 to 400°C at 100 psig. In contrast to benzothiophene HDS, ECH was always a major product for indole HDN over Ni-Mo catalyst, implying that the hydrogenation of benzene ring is important in the coiu^se of indole HDN. By selecting reaction parameters which favor hydrogenation reactions, OHI and both cis- and trans-OECHA were observed in our reaction experiments. Since the hydrogenation of EB to ECH does not seem to be important under our reaction conditions, as evidenced by the BT HDS results, the hydrogenation of benzene ring should occur prior to the C-N bond cleavage. This, in turn, suggests that OECHA is the intermediate in ECH formation. According to the reaction network proposed, there are two possible routes for OECHA formation, i.e., the hydrogenation of OEA and the hydrogenolysis of OHI. It has been reported in the literature that, in a mixture with other nitrogen-containing compounds, aniline-tjrpe molecules are least reactive [3, 28, 30, 31]. To determine the reactivity of OEA in an indole HDN reaction mixture, we performed a set of experiments keeping the total concentration of nitrogen compounds constant, but replacing one half of indole with OEA. The results from these experiments are summarized in Table 1. The percentage of OEA in the product stream resulting from a co-feed of indole and OEA, although somewhat higher than that resulting from a OEA-free feed, is less than 10 % of the OEA concentration which was in the feed. This result clearly shows that OEA remains highly reactive even in the presence of indole/indoline. It should also be noted that the ECH/EB ratio decreases when half of indole in the feed is replaced by the OEA intermediate. If the conversion of indoline is through the OEA route only, same ECH/EB ratio would be expected in this case. The lower ECH/EB ratio from indole + OEA HDN suggests that both OEA and OHI routes are important for the conversion of indoline. Table 1 HDN of indole and indole+OEA mixture over the Ni-Mo catalyst at 320°C and 100 psig

Feed indole alone indole + OEA

indole mol% inlet outlet inlet outlet

0.046 0.023 0.023 0.017

OEA mol% 0 0.0016 0.023 0.0022

ECH/EB 2.03 1.72

79 It has been reported in the hterature that hydrogenation reactions are enhanced by increasing H2 pressure whereas hydrogenolysis reactions are not very sensitive to Hg pressure [32]. Based on this observation, one would expect that increasing H2 pressure will increase the importance of OHI route in indole HDN reactions and will result in larger OHI and OECHA presence in the reactor effluent if hydrogenolysis reaction rates are relatively lower. The experiments we performed at 1000 psig for indole HDN in the absence of sulfur compounds over the Ni-Mo catalyst showed that it was indeed the case. Figure 4 compares the selectivity of some major products from indole HDN at 100 and 1000 psig H2 pressure. At low pressure, only trace amounts of OECHA and OHI were present, whereas at high pressure, OECHA had the highest selectivity among N-containing species and the relative amount of OHI was also increased. Also at high pressure, OEA, an intermediate from indoline hydrogenolysis, had the lowest selectivity among all N compounds, whereas it had the highest selectivity at low pressure. When hydrocarbon product selectivities are compared, the most pronounced difference is that ethylbenzene selectivity dropped while ECH selectivity increased when H2 pressure increased was from 100 to 1000 psig. This result suggests that due to a much stronger hydrogenation function of the catalyst at high pressure, the route via OEA in indole HDN becomes less important with increasing H2 pressure in the absence of sulfur compounds.

H

0^2 r

Cr 10

20

30

Selectivity %

0

10

20

30

40

50

Selectivity %

Figure 5. Comparison of indole HDN product selectivities at 100 and 1000 psig. (3%NiO-15%Mo03 catalyst, 320°C, feed: 0.046% indole in hydrogen)

80 Two routes are proposed for the ECH formation from OECHA HDN as shown in the network. The route via p-ehmination (OECHA -> ECHE ) and hydrogenation (ECHE -^ ECH) has been well estabhshed [4-6, 33]. The route without ECHE as an intermediate could be envisioned as a hydrogenolysis reaction or, more accurately, as a Hofmann degradation reaction [6, 33, 34] in which the -NHg group of OECHA is replaced by a -SH group whicli goes through hydrogenolysis very easily and rapidly. The results obtained from indole HDN by varying the HgS-to-Hg ratio in the feed from 0 to 1.63 xlO'^ u n d e r ICKX) psig over the bimetallic catalyst are presented in Figure 6. As shown in the figure, with increasing H2S-to-H2 ratio, the indole conversion and ECH production increased steadily up to a H2S/H2 ratio of 0.3x10'^. The ECHE and EB production rates, on the other hand, showed very Uttle change with HgS/Hg ratio. OECHA production rate decreased rapidly with increasing HgS concentration (Figure 6B) suggesting a direct correspondence between the ECH and OECHA production rates. These results suggest the presence of a second OECHA -» ECH route, which does not go through the ECHE intermediate.

0.4 0.8 1.2 1.6 0 0.4 0.8 1.2 l.( Mole Ratio, H2S/H2 xlOO Mole Ratio, H2S/H2 xlOO Figure 6. Effect of H2S-to-H2 ratio on indole HDN over the 3%NiO-15%]V[o03 catalyst (320°C, 1000 psig , feed: 0.046% indole in 300 cm^(STP)/min hydrogen). As shown in Figure 6B, there was a sharp decrease of indoline, OHI, and OECHA production rates with increasing HgS concentration, especially at lower H2S/H2 ratios. Since C-N bond cleavage is involved in the further reaction of each of these species, this observation indicates that the higher activity in indole HDN with increasing H2S/H2 ratios is directly linked to the enhancement effect of HgS on C-N bond cleavage (hydrogenolysis) reactions.

81 The appearance of significant amounts of these intermediate species at h i g h hydrogen pressures in the absence of sulfur compounds (Figure 5) seems to suggest that, at higher pressures, the hydrogenation function of the catalyst is relatively strong and more HgS is needed to balance the hydrogenation a n d hydrogenolysis functions of the Ni-Mo sulfide HDN catalysts. The necessity of replenishing the sulfur sites with HgS is especially relevant to industrial practice since high hydrogen partial pressures tend to deplete the sulfur on the catalyst more readily, leading to a greater loss of hydrogenolysis activity. 4. SUMMARY/CONCLUSIONS Combining the results of our kinetic and characterization studies with some findings in literature, a catalytic job distribution of Ni and Mo centers in HDN reactions is proposed. It is envisioned that, in the HDN of nitrogen heterocycles, Ni-associated sulfur vacancies in the Ni-Mo-S phase a r e responsible for hydrogenation steps, whereas the primary function of Mo is to promote C-N bond hydrogenolysis reactions through Bronsted acid sites, which are generated fi*om the adsorption and dissociation of H2S on Moassociated sulfur vacancies. The conversion of Mo-associated S vacancies to Bronsted acid sites by HgS is a reversible process strongly depending on temperature and HgS-to-Hg ratios. The reaction network of indole HDN is a complex one, consisting of multiple steps of hydrogenation, hydrogenolysis, P-elimination, Hofmann degradation, and dehydrogenation. The reactivity can not be characterized in terms of a single rate-determining step, but is mainly controlled by the relative strengths of the hydrogenation and hydrogenolysis functions of the catalyst. As for the effect of sulfur compounds on the HDN activity, a "universal" effect which accounts for all conditions can not be defined. This effect varies significantly depending on whether the hydrogenation of the heterocyclic-ring is kinetically or thermodynamically controlled and which steps and which sites are dominant in the overall catalytic scheme. At h i g h temperatures and high Hg pressures, more HgS is needed to maintain the proper balance between the hydrogenation and the hydrogenolysis functions of the catalyst. Acknowledgment The financial support provided for this work by the National Science Foundation through the Grant HRD-9023778 is gratefully acknowledged. REIFERENCES 1. 2.

Tops0e, H., Clausen, B.S., and Massoth, F.E., Hydrotreating Catalysis, Springer-Verlag, Berling,1996. C.N. Satterfield, Heterogeneous Catalysis in Industrial Practice, 2nd ed., McGraw-Hill, 1991, p383.

82 3. 4. 5. 6. 7. 8. 9. 10. 11. 12. 13. 14. 15. 16. 17. 18. 19. 20. 21 22. 23. 24. 25. 26. 27. 28. 29. 30. 31. 32. 33. 34.

G. Perot, Catal. Today, 10 (1991) 447. M.J. Girgis and B.C. Gates, Ind. Eng. Chem. Res., 30 (1991) 2021. T.C. Ho, Catal. Rev.-Sci. Eng., 30 (1988) 117. R. Prins, in Knozinger, Ertl and Weitkamp (eds.). Encyclopedia of Catalysis, to be published. C.N. Satterfield, M. Modell, and J.A. Wilkens, Ind. Eng. Chem. Process Des. Develop., 19 (1980) 154. C.N. Satterfield and S. Gultekin, Ind. Eng. Chem. Process Des. Develop., 20 (1981) 62. S.H. Yang and C.N. Satterfield, J. Catal., 81 (1983) 168. S.H. Yang and C.N. Satterfield, Ind. Eng. Chem. Process Des. Develop. 23 (1984)20. H. Tops0e and B.S. Clausen, Catal. Rev.-Sci. Eng. 26 (1984) 395. U.S. Ozkan, S. Ni, L. Zhang, and E. Moctezuma, Energy Fuels, 8 (1994) 249. U.S. Ozkan, L. Zhang, S. Ni and E. Moctezuma, J. Catal., 148 (1994) 181. U.S. Ozkan, L. Zhang, S. Ni and E. Moctezuma, Energy Fuels, 8 (1994) 830. L. Zhang and U.S. Ozkan, in J.W. Hightower et al. eds., 11th Int. Cong. Catal - 40th Anniversary, 1996, 1223. U.S. Ozkan, Y. Cai, M.W. Kumthekar, and L. Zhang, J. Catal., 142 (1993) 182. J.G. Weissman and J.C. Edwards, Appl. Catal. AiGeneral 142 (1996) 289. H. G. Mcllvried, Ind. Eng. Chem. Process Des. Develop., 10 (1971) 125. C.N. Satterfield and J.F. Cocchetto, AIChE J., 21 (1975) 1107. L. Zhang and U.S. Ozkan, to be summited to J.Catal. L.D. RoUmann, J. Catal. 46 (1977) 243. E.W. Stern, J. Catal., 57 (1979) 390. E.O. Odebunmi and D.F. OUis, J. Catal., 80 (1983) 76. J.-L. Olive, S. Biyoko, C. Moulinas and P. Geneste, Appl. Catal., 19 (1985) 165. P. Zeuthen, P. Stolze and U.B. Pedersen, Bull. Soc. Chim. Belg., 96 (1987) 985. J. Shabtai, G. Que, K. Balusami, N.K. Nag and F.E. Massoth, J. Catal., 113 (1988) 206. F.E. Massoth, K. Balusami and J. Shabtai, J. Catal., 122 (1990) 256. M. Callant, P. Grange, K.A. Holder, and B. Delmon, Bull. Soc. Chim. Belg., 104(1995)245. M.V. Bhnde, Ph.D. Dissertation, University of Delawere, Newark, 1979. S. Kasztelan, T. des Courieres, and M. Breysse, Catal. Today, 10 (1991) 433. S.-J. Liaw, A. Raje, K.V.R. Chary, and B.H. Appl. Catal. A:General 123 (1995)251. S.R. Shih, J.R. Katzer, H. Kwart, and A.B. Stiles, Am. Chem. Soc. Div. Petrol. Chem. Prepr., 22 (1977) 919. N. Nelson and R.B. Levy, J. Catal. 58 (1979) 485. J.L. Portefaix, M. Cattenot, M. Guerriche, J. ThivoUe-Cazat and M. Breysse, Catal. Today, 10 (1991) 473.

® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors

83

Kinetics of the catalytic removal of the sulphur components from the light cycle oil of a catalytic cracking unit G.F. Froment, G.A. Depauw and V. Vanrysselberghe Laboratorium voor Petrochemische Techniek, Rijksuniversiteit Gent, Krijgslaan 281, B-9000 Gent, Belgium ABSTRACT First order kinetics for the hydrodesulphurization of 40 sulphur components of the benzothiophene and dibenzothiophene families were determined using experimental data obtained with a light cycle oil. Substituents next to the sulphur atom reduce the reactivity with respect to the nonsubstituted component. Hougen Watson rate equations for substituted dibenzothiophenes were related to those for the head of the family, dibenzothiophene, through global multiplication factors f„ s^ny ^"^ frsDUT- These factors are the products of the electronic and steric effects of the substituents on the adsorption equilibrium constants and of the electronic effects on the rate coefficients. They were evaluated for a number of substituted dibenzothiophenes. 1. INTRODUCTION Light cycle oil (LCO) contains various aromatic sulphur components such as benzothiophenes (BTs), dibenzothiophenes (DBTs) and naphthothiophenes [1]. These components have to be removed for both technical and environmental reasons. Hence the sulphur content of diesel is limited to 0.05wt% from October 1996 onwards. Detailed kinetic equations for the sulphur components are required for the modeling of the trickle bed reactors used for hydrodesulphurization (HDS). LCO is an ideal feedstock for the determination of the kinetic data since all interesting refractory sulphur components are present in it and since the content of aliphatic sulphur components as well as the length of the alkyl side chains are limited, thus facilitating the quantification. A kinetic modeling of the hydrodesulphurization reactions based upon structural contributions was developed to limit the number of parameters [2]. The rates for reactions involving substituted components were related to those of a non substituted reference component in terms of the effect of the substituents on the rate coefficients and the adsorption equilibrium constants. Based upon a number of assumptions, multiplying factors were introduced for the electronic and steric effect of the substituents on the adsorption and for the electronic effect on the reaction rate. In that way the number of parameters for the HDS of a set of mono-, di- and trimethyl substituted DBTs was reduced from 1133 to 93. According to the rate equations derived for the HDS of DBT [3] and assuming that the rate expressions

84 for the HDS of the methyl substituted DBTs are of an identical form, this number can be further reduced to 35. A limited set of experiments with complex mixtures, such as LCO, can be used to obtain the global multiplication factors for various components. These global factors are the products of the effects of the substituents on both the rate coefficients and the adsorption equilibrium constants. The determination of the structural contributions as defined by Froment et al. [2] requires additional experimental data obtained from model components such as 4-methyldibenzothiophene (4-MeDBT), 4,6-dimethyldibenzothiophene (4,6-DiMeDBT) and one of the trimethyldibenzothiophenes. 2. EXPERIMENTAL SET-UP The liquid hydrocarbon was fed into the reactor with a high-pressure pump. The hydrogen, the hydrogen sulphide and the nitrogen feed were controlled and metered with a set of electronic mass flow controllers. Hydrogen sulphide and nitrogen were used in the pretreatment of the catalyst. The gases and the liquid feed were preheated and mixed before entering the reactor. The effluent section was also heated to avoid condensation. The reaction was carried out in a multiphase Robinson-Mahoney reactor. The temperature was measured by means of thermocouples and controlled by a PID temperature controller. The pressure was controlled by a back pressure regulator. The effluent of the reactor consisted of gas and liquid phases at high pressure and high temperature. Both phases were separated by means of a cyclone. The liquid was collected in the liquid holder. The cyclone and the liquid holder were kept at the same pressure and temperature as in the reactor, to avoid changes in composition of both phases. The gas phase was cooled, so as to condense heavy fractions, and was then scrubbed by means of a sodium hydroxide solution to remove hydrogen sulphide before venting. The liquid product was cooled and flashed under ambient conditions. The light gases, dissolved in the liquid phase, were partially desorbed and collected in a gas burette. Product samples were taken off-line for GC analysis. 3. ANALYSIS The GC-AED system (Hewlett-Packard 5921 A) was used for the quantification of sulphur components and the determination of the total sulphur content of the LCO mixture. The atomic emission detector (AED) is element specific and has a high dynamic range for C, H, N and S. The total sulphur content of the LCO (1.28wt%) was determined with the AED using hydrogen linearization. Detailed identification and quantification results of the sulphur components in LCO are presented in a separate paper by Depauw and Froment [1]. A quadrupole mass spectrometer, part of the Hewlett-Packard 5989A system was used for quantification as well. Electron impact ionisation mass spectra (m/z 40-400) were obtained at 70eV at a rate of 1.6 scans/second. The quadrupole temperature was lOO^'C, the ion source temperature 250''C and the transfer line temperature 250''C. The GC separation was performed with a Hewlett-Packard 5890 Series 11 instrument on a 50 m x 0.2 mm fused silica capillary column coated with a 0.5 ^m film of cross linked 100% dimethylsiloxane (HPPONA). Helium was used as carrier gas (0.645 ml/min at 35''C). The column was

85 temperature programmed from 35°C(5 min) to SCCCIS min) at a rate of 2.5°C/min and further to 200°C(5 min) at a rate of 2.0°C/min and finally to 250X at a rate of TC/min. The injector temperature was 250°C. One microliter of undiluted sample was injected at a split ratio of 63:1. The conversion of a molecule can be calculated by comparing the feed and effluent ion chromatograms. The effluent chromatograms are scaled using fluorene as an internal standard. This molecule is present in the LCO feed and is not produced nor hydrogenated or evaporated significantly under reaction conditions. It does not co-elute with other components with the same mass. The most important fragment of fiuorene with a m/z ratio of 166 was used for the scaling. An ion chromatogram (m/z =166) of an effluent is multiplied with a scaling factor in order to get the same surface for the fluorene peak as was obtained for this peak in the feed analysis. All other ion chromatograms of the effluent are multiplied with the same factor. The surfaces of the peaks in the effluent and the feed ion chromatograms are then linearly proportional in an identical way to the molar flows at the reactor exit and entrance respectively. The conversions of all components in the feed are derived from the relative decrease of their fragments. In case of co-elution of several molecules an ion fragment which is typical for the molecule considered has to be used for the calculations of the conversion. 4. EXPERIMENTAL PROGRAM Experiments were performed with a light cycle oil (LCO) of a catalytic cracking unit containing 1.28wt% sulphur and with a boiling range of 225-370''C. Experiments were carried out at temperatures between 240 and 320°C. The liquid pumping rate was varied between 15 and 27 ml/hr. The total pressure was 80 bar, and the hydrogen to hydrocarbon ratio 344 Nl/1. The molar hydrogen to methane ratio was 6.4. The number of experiments amounted to 25. The catalyst used was the commercial HDS catalyst AKZO Ketjenfine 742. It was crushed to a size between 710 and 800jitm to avoid diffusional limitations and 2.53 gcat were diluted with nonporous inert alumina. The absence of diffusional limitations was calculated using the Weisz-Prater criterion. The reaction mechanism and the intrinsic kinetic equations for the hydrodesulphurization of dibenzothiophene were derived by Vanrysselberghe and Froment [3]. In order to determine the kinetics for the hydrogenation of naphthalene into tetralin, experiments were carried out using a solution of 2wt% dibenzothiophene and l-5wt% naphthalene in a paraffinic mixture (Cjo-CiJ. The temperature was varied between 240 and 300''C, the molar hydrogen to hydrocarbon ratio between 1.10 and 1.36. The molar feed flow rate of dibenzothiophene was varied between 1.23 10'^ and 3.94 10'^ kmol/hr, the naphthalene molar feed flow rate between 1.71 10'^ and 1.42 lO'** kmol/hr. The total pressure was 80 bar.

86 5. RESULTS AND DISCUSSION 5.1. First order kinetics In order to compare the kinetic data obtained in this work with literature data and to provide quantitative results on the reactivity, first order rate coefficients were determined from the experimental conversions. The conversion X; of component i is defined as: pj. in

X.=

T;

^

out

(1)

'

The conversion Xj of component i is directly obtained from the experimental data. The continuity equation for this component in the perfectly mixed reactor is given by: F/"-r.W=Fr

(2)

Substitution of (2) into (1) results in: r.W

with

Fi'"=V'"Ci'"

(3)

Using the first order approach, the rate of removal is assumed to have the form: r=k.C. I I

(4)

The ratio of the concentration of i at reactor conditions and in the inlet CJC-^ can be written as (l-Xi)pV"'/p"'V°"^ The rate coefficient is finally written as a function of measured quantities: k.=—L-l-^ '

1-Xj W

p

(5)

The activation energy and preexponential factor are determined from the Arrhenius relationship: krAiexp(-JfL)

(6)

87 Results for the BT-family The results for 27 sulphur components of the BT-family are presented in Table 1. The conversions used for the determination of these properties were obtained with the GC/MS system, which shows less peak overlapping than the GC-AED analysis. The Arrhenius plots for a set of substituted benzothiophenes are shown in Figure 1. The correlation coefficients R^, the ratio of the regression sum of squares to the total sum of squares, varied from 0.986 for 7-EtBT to 0.999 for 2567-TeMeBT. Table 1 Preexponential factors, activation energies and first order rate coefficients at T=280°C for various members of the BT-family Sulphur component

A, [mVkg,,yh]

E,i [kJ/kmol]

k, [mVkg,,,/h] atT=280°C

BT 54-6-MeBT 3+4-MeBT 7-MeBT 27-DiMeBT 23-DiMeBT 2-EtBT 7-EtBT 24-DiMeBT 56-DiMeBT 45-DiMeBT 34-DiMeBT 35-h36-DiMeBT 257-TriMeBT 357-TriMeBT 267-TriMeBT 356-TriMeBT 234-TriMeBT 235+236-TriMeBT 7-PrBT 2357-TeMeBT 2367-TeMeBT 2567-TeMeBT

6.537 10^^' 4.253 10' 8.687 10' 2.025 10' 2.717 10^' 5.511 lO''* 2.173 10' 2.394 10^« 4.412 10'" 7.423 10' 9.943 10' 6.025 10*^ 3.717 10^2 5.102 10'" 1.212 10^" 4.371 10'^ 1.528 10" 9.660 10'' 1.816 10'2 8.112 10'^ 2.026 10'^ 5.629 10*2 5.297 10'"

121100 107900 85900 113300 187000 179100 88200 125300 172400 112200 115900 87170 142100 178400 123500 147500 130100 210000 150200 171100 142600 159700 134400

2.587 10' 3.109 10' 8.737 10-2 5.059 10-2 6.591 10-^ 9.990 10-' 1.016 10' 4.645 10-2 2.852 10-2 1.885 10' 1.087 10' 3.345 10-2 1.921 10' 8.197 10-' 3.180 10-2 6.515 10-' 8.258 10-2 1.363 10-2 1.382 10-2 5.840 10-2 7.185 10-' 4.906 10-' 1.147 10-2

88 1/T(*1000)(K'^) 1.78 1.83 H

1.

1.93 h

• 27-DMBT 1 B 7-MBT A 2367-TMBT xBT X 356-TMBT • 2567-TMBT + 7-EBT - 235+236-TMBT - 56-DMBT o 45-DMBT D 357.TMBT 1 A 257-TMBT

Figure 1

Arrhenius plots for a set of components of the BT-family.

Components with substituents in position 2 show an important decline in reactivity with respect to benzothiophene. This is less pronounced for substituents in position 7 and position 3. The decline in reactivity due to substitution in position 2 and 7 is caused by the steric hindrance of these substituents on the vertical adsorption through the sulphur atom. A substituent in position 3 causes steric hindrance on the adsorption through the double bond of the thiophenic ring. Substituents in the positions 4, 5 and 6 hardly affect the reactivity of the component. The first order rate coefficients of the methyl substituted benzothiophenes can be related to the rate coefficient of benzothiophene as follows: 1. AAAJAAA^ Kj-r2l7 131415 16 Kjj

(7)

where f4=f5=f6. The power 6j equals one if the substituted BT contains a methyl substituent in position j , otherwise 5j is zero. The product f2'f7Y3'f4'fJ*f6' 's a global multiplication factor if the removal of the benzothiophenes occurs mainly on one type of catalytic site. The parameters fj were obtained by minimization of the objective function:

89

A 2:[ln(iL)-ln(^)]^ >=1

'^BT

^^^

BT

where n=23. The parameter estimates for T=240X, T=260°C, T=280°C and T=300X are given in Table 2. The parameters f, could not be determined at T=320°C because benzothiophene was completely converted at this temperature for the flow rates that could be applied in the equipment. Table 2 Values of the fj parameters of the rate coefficient prediction model for four temperatures. Parameter

T=240X

T=260°C

T=280°C

T=300°C

f2

0.0479

0.0882

0.125

0.183

h

0.334

0.295

0.298

0.318

f3

0.395

0.429

0.481

0.609

f4 = f5 = f6

1.078

0.871

0.886

0.807

The calculated versus experimental first order rate coefficients for the substituted BTs are shown in Figure 2 for all temperatures. The correlation coefficient R^ is 0.973. From the parameter values of Table 2 it is seen that a methyl substituent in position 2 reduces the reactivity at T=240''C with 95% with respect to BT. For methyl groups in positions 7 and 3 the reduction at T=240°C is respectively 67% and 61 %. The effect predicted by the model of the 4, 5 and 6 methyl substituents is a small increase at T=240°C and a reduction in reactivity at the other temperatures. The l, parameter values increase with temperature for substituents on the thiophene moiety and tend to decrease for substituents on the benzene moiety. Results for the DBT-family These results are given in Table 3. Houalla et al. [4] determined first order rate coefficients for DBT, 4-MeDBT and 4,6-DiMeDBT at 300''C and 102 bar on a C0M0/AI2O3 catalyst. The ratio between the rate coefficient of the substituted DBT and DBT obtained by Houalla et al. [4] and in the present work is given in Table 4. The ratios for 4-MeDBT and 4,6-DiMeDBT are higher than those found by Houalla et al. [4] The reactivity obtained for 4-MeDBT with respect to 4,6-DiMeDBT is almost 3 in this work and 1.35 according to the results of Houalla et al. [4]

90

Calculated ki [m^ /kg cat hr] 1

0.001

0.003

0.01

0.03

0.1

0.3

Experimental ki [m^ /kgcathr] Figure 2 Calculated versus experimental first order rate coefficients at all temperatures for components of the BT-family. Table 3 Preexponential factors, activation energies and first order rate coefficients at T=320°C for various members of the DBT-family. Sulphur component

A; [mVkg^Jh]

E,i [kJ/kmol]

ki [m'/kg^/h] atT=320°C

DBT 1-MeDBT 2+3-MeDBT 4-MeDBT 46-DiMeDBT 24-DiMeDBT 13-DiMeDBT 23-DiMeDBT 4-PrDBT 4Et6MeDBT 146-TriMeDBT 346-TriMeDBT

3.974 10' 1.717 10' 7.579 la' 1.181 10" 5.161 10' 3.371 10^ 5.421 10' 4.714 10" 5.220 10' 5.499 10' 2.364 10' 9.252 10"

64700 57800 91400 76850 54100 59400 96300 72500 48700 63900 59300 82700

1.064 10-' 1.892 10' 1.035 10' 2.115 10-' 8.152 10-" 3.457 10-' 2.253 10-' 2.566 10' 3.060 10' 1.174 10-' 1.821 10-' 4.733 10'

91 Table 4 Comparison of the reactivities found by Houalla et al. [4] and the present work at T = 3 0 0 X for 4-MeDBT and 4,6-DiMeDBT with respect to DBT llaetal. [4] '^-MeDBT'l^DBT

Present work

0.090

0.35

0.067

0.12

5.2. Hougen Watson kinetics Introduction Aromatic sulphur components are converted on a sites by hydrogenolysis and on r sites by hydrogenation. Several authors [3,5,6] found for the HDS of benzothiophene and dibenzothiophene that the surface reaction between adsorbed species is the rate determining step on both sites. The resulting Hougen Watson rate equation is generalised here for all aromatic sulphur components. The rate equation for the disappearance of an aromatic sulphur component i in a complex mixture becomes: K^\o^Haî

^H,

r.=' DEN„(T,C„...,C„)

^rîr^Hrî

^H,

DEN^(T,C„...,C„)

(9)

where the first term relates to the hydrogenolysis and the second to hydrogenation. The adsorption characteristics on both catalytic sites differ, as reflected by the denominators DEN„ and DEN^. These contain the concentrations of all adsorbing species and their temperature dependence. The functions represented by DEN„ and DEN^ are identical for all rate equations. Hydrogenation reactions of aromatics occur only on the r sites. The surface reaction was observed to be the rate determining step for biphenyl [3,6] and naphthalene hydrogenation. Since the hydrogenation reactions are irreversible their rate can be written: _

k..K.^K„^C. C^

^^^^

' DEN^(T,Cp...,C„) The rate equations for the hydrogenation of sulphur components and aromatics contain the same denominator DEN^. Determination of the denominators DEN,, and DEN, In complex mixtures the denominators DEN,, and DEN^ cannot be calculated a priori, since not all adsorbing species and their corresponding adsorption equilibrium constants are

92 known. Relating the rates of substituted sulphur components in complex mixtures to those of the unsubstituted heads of the families, requires the knowledge of both denominators DEN„ and DEN^ for each LCO experiment. For the complex mixture LCO these can be calculated considering model components for which kj^Kj^KH^ and kj^Kj^Hr ^^^ known, since these products are invariant of the mixture composition. The product k,)BT„K[)BT^KH„ was determined for the hydrogenolysis of dibenzothiophene [3]. The products k^mTT^Dm^Hr^ kupH^BPHr^HT and kN^Nr^Hr were determined for the hydrogenation of dibenzothiophene [3], biphenyl [3] and naphthalene in the present work. These products can be substituted in the rate equations of the reactions of dibenzothiophene and naphthalene in the LCO mixture. Now the denominators of the rate equations of these model components can be calculated. These denominators DEN„ and DEN^ are identical for the rate equations of all the sulphur compounds in the LCO. The expressions for the conversion of DBT, the conversion of DBT into biphenyl and cyclohexylbenzene, the conversion of naphthalene and the conversion of naphthalene into tetralin in a LCO in the completely mixed reactor can be written:

" " ^ " " C ^ V ^ DEN„(T,C„...,C„) _

W

DEN/T,C„...,C„)

^ H , J- ^ D B T o ^ D B T o ^ H o ^ D B T _ * ^ B P H T ^ B 1 » H T ' ^ H r ^ B P H -I

'"'" C ^ V^ DEN„(T,C„...,C„) " DEN,(T,C„...,C„)^ " " ' c £ v ^ DEN,(T,C, . . - ^ % NT

"

c^

NT

C.) DEN,(T,C

V^'DEN^(T,C„...,C„)^

V„„ =2.44 10'» exp[. -122.8 10^ ] kmol/kgjh

ko„,

=3.36 10-" exp[ilM_!2!] m'/kmol

KJ

mVkmol

=2.87 10'* exp[Zl^^4^] k'1101/kg.a/h

KJ

k„p„^ =3.41 10" exp[-^^Jiy^] kmol/kgjh

KJ

(13)

Cj'

(15)

with

K„„T„ =7.57 10'

(12)

(14)

HT

C ; V-^DEN/T,C„...,C„)^

KH„

(11)

93 K„,

=1.40 10-" e x p [ i i | i L l ^ ]

K^^^ =2.50 10-' exp[

76.8 10^

mVkmol

]

m'/kmol

K „ _ =4.97 10-" exp[ 37.9 10^ ]

m'/kmol

-185.9 10^

k^,

=1.57 10" exp[-

K^,

=5.06 10-' e x p [ ^ 4 ^ i ^ ] R,a.T

] kmol/kgjh mVkmol

The values of the unknowns, DEN„ and DEN^ were estimated for each LCO experiment by means of regression. They are shown in Figures 3 and 4 at T=320°C for various liquid compositions expressed in terms of a molar averaged conversion defined as follows: X

1

=

E î y*

(16)

Eyi with Xj the conversions of a set of selected components (BT, DBT, Naphtho[2,/-Z7]thiophene, 4-MeDBT, 4,6-DiMeDBT, Naphthalene and Phenanthrene) and yj the corresponding mole fractions in the LCO feed. The adsorption is weaker at higher temperatures on both the r and a sites and the denominators decrease. For a given temperature, the coverage of the r sites increases and that of the a sites decreases with molar averaged conversion. The reacting species in the LCO mixture adsorb to a higher extent on the a sites than the reaction products. 220

200

27

30

33

Molar averaged conversion [%]

Figure 3: Denominator corresponding to the c7-sites at T=320°C as a function of the molar averaged conversion.

27

30

33

Molar averaged conversion [%]

Figure 4: Denominator corresponding to the T-sites at T=320°C as a Function of the molar averaged conversion.

94 Determination of multiplication factors In a second step the numerical values of both denominators can be used in the rate equations of substituted aromatic sulphur components. The rate equations for these substituted sulphur components, sDBT e.g., can be related to that of the unsubstituted head of the family, DBT [2]: f

]c

K

K

f

fsDBT-^sDBT ^",1 pEN (T,C„...,C„)

\c

K

K

DEN,(T,C„...,C„) ^

The global multiplication factors f„ ^DBT and f^soBT are the products of the electronic and steric influences of the substituents on the adsorption equilibrium constant and of the electronic effects on the rate coefficient. These factors depend on the temperature. The conversion of a sDBT in a completely mixed reactor can then be written: _ ^

^sDBT

Hj |. *asDBT^DBTo ^ D B T o

Ho

^ ° ' ' ^ ~ C ^ V^ DEN„(T,C„...,C„)

TSDBT^DBTT ^ D B T T ' ^ H T - .

/1 0\

DEN/T,C„...,C„)^

The global multiplication factors f^ SDBT and frsDBT are the unknowns in this equation. As mentioned ^mja^mTcf^Ha and kj)j,T^t)j,TrKHr were already derived by Vanrysselberghe and Froment [3]. The denominators were determined for the compositions reached in each LCO experiment as explained above. The multiplication factors at a given temperature are obtained by minimization of the objective function:

/ ^

(^SDBT"^.S[)BT)

^

^

where X',DBT are the experimental conversions. In Table 6 the global multiplication factors t,i)BT and t,[)BT at T=320X are given for 4-MeDBT, for 4,6-DiMeDBT and for 2 and 3MeDBT which were not separated by GC-MS. 6. CONCLUSIONS The reactivity of various sulphur components of the benzothiophene and dibenzothiophene family has been demonstrated using first order rate coefficients. Substituents next to the sulphur atom reduce the reactivity with respect to the unsubstituted component. Substituents in other positions can either increase or decrease the reactivity. The electronic and steric influences of the substituents on the rate coefficient and the adsorption equilibrium

95 Table 6 Global multiplication factors f„,i,BT and t,^,^ at T=320°C for 2+3-MeDBT, 4-MeDBT and 4,6-DiMeDBT. Component

K sDHT

I i sDBT

2+3-MeDBT

1.3

1.1

4-MeDBT

0.15

3.3

4,6-DiMeDBT

0.037

1.8

constant in the more refined approach of Froment et al. [2] were combined into global multiplication factors f„ .DBT and f^ ^DBT i^ the Hougen Watson rate equations. The denominators corresponding to the a and r sites in the LCO can be obtained by using the Hougen Watson rate equations of the hydrodesulphurization of dibenzothiophene and the hydrogenation of naphthalene into tetralin. The a sites were found to be more occupied than the T sites. The global multiplication factors can be determined in a next step using the numerical values for the denominators. The next level of refinement involves the determination of the structural contributions as defined by Froment et al. [2]. This approach requires experimental data with 4-MeDBT, 4,6-DiMeDBT and one of the trimethyldibenzothiophenes as model components. The contributions kELo''''''(m;0;0), k^J'''''(n\;0;0), Ki,L+sTr''''''(m;0;0) and the product KsTa^^'^(4;0;0)KELo^"'^(m;0;0) can be obtained from experiments with 4-MeDBT. Combining these results with the global multiplication factors obtained with the LCO experiments, KsTa^*^^(4;0;0) can be calculated by dividing the global multiplication factor fs^DBx for 4-MeDBT by that for the 1, 2 or 3-MeDBT. An identical approach has to be followed to determine the structural contributions for the di- and trimethylDBTs. ACKNOWLEDGEMENT This work was funded by the European Commission under the Joule program contract no. JOU2-0121. V. Vanrysselberghe and G.A. Depauw are also grateful for a contribution from the Center of Excellence Grant awarded to the Laboratorium voor Petrochemische Techniek by the Belgian Ministry of Science. We wish to thank R. Le Gall for his cooperation. REFERENCES L G.A. Depauw and G.F. Froment, Journal of Chromatography, to be published. 2. G.F. Froment, G.A. Depauw and V. Vanrysselberghe, Ind. Eng. Chem. Res., 33 (1994) 2975.

96 3. V. Vanrysselberghe and G.F. Froment, Ind. Eng. Chem. Res., to be published. 4. M. Houalla, D.H. Broderick, A.V. Sapre, N.K. Nag, V.H.J. De Beer, B.C. Gates, and H. Kwart, J. of Catal., 61 (1980) 523. 5. I.A. Van Parys and G.F. Froment, Ind. Eng. Chem. Prod. Res. Dev., 25 (1986) 431. 6. R. Edvinsson and S. Irandoust, Ind. Eng. Chem. Res., 32 (1993) 391. NOMENCLATURE Ai

c> Eai pin pout

kE.."»T(m;0;0)

W,™^(m;0;0)

K,J«'T(m;0;0) KE,,.sT.™^(ni;0;0)

K.sT„""(4;0;0)

•^gas

T yin yout

w XfiPH ^CHB ^DBT ^sDBT

preexponential factor liquid concentration of component i inlet liquid concentration of component i activation energy inlet molar liquid flow rate of component i outlet molar liquid flow rate of component i electronic effect of one methyl group on the rate coefficient for the hydrogenolysis of sDBT electronic effect of one methyl group on the rate coefficient for the hydrogenation of sDBT first order rate coefficient of component i rate coefficient of component i on s sites electronic effect of one methyl group on the adsorption equilibrium constant of sDBT on the a sites electronic and steric effect of one methyl group on the adsorption equilibrium constant of sDBT on the r sites steric effect of a methyl group in position 4 on the adsorption equilibrium constant of sDBT on the a sites adsorption equilibrium constant of component i on s sites gas constant total rate of disappearance of component i absolute temperature total inlet volumetric liquid flow rate total outlet volumetric liquid flow rate total catalyst mass conversion of dibenzothiophene into biphenyl conversion of dibenzothiophene into cyclohexylbenzene conversion of dibenzothiophene conversion of substituted dibenzothiophene conversion of component i conversion of naphthalene conversion of naphthalene into tetralin calculated conversion of substituted dibenzothiophene

m-Vkg,,yh kmol/mL^ kj/kmol kmol/h kmol/h

mVkgJh kmol/(kg,,ih)

mL^/kmol kJ/kmol/K kmol/(kg,,ih) K mVh mVh kgcat

97 GREEK SYMBOLS p p^ a r

liquid density at reactor conditions inlet liquid density hydrogenolysis site hydrogenation site

SUBSCRIPTS BPH CHB DBT H H2 a r

biphenyl cyclohexylbenzene dibenzothiophene atomic hydrogen molecular hydrogen with respect to the hydrogenolysis function with respect to the hydrogenation function

SUPERSCRIPTS ^ in out

calculated inlet conditions outlet conditions

kg/m^ kg/m^

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® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P, Grange, editors

99

A review of catalytic hydrotreating processes for the upgrading of Uqmds producedfayflash pyrolysis R. Maggi and B. Delmon Unit6 de Catalyse et Chimie des Mat^riaux Divises, University Catholique de Louvain, Place Croix du Sud 2/17,1348 Louvain-la-Neuve, Belgium ABSTRACT Liquids produced by flash pjnrolysis of biomass and solid wastes are intended to be used in direct combustion but they contain a high quantity of oxygenated molecules which causes unwanted characteristics such as thermal instability, tendency to polymerise, corrosion and low heating value. These properties can be improved by partial or total elimination of oxygen atoms by catalytic hydrotreating with production of hydrocarbons and water. This paper reviews the development of this process from the first empirical tests using real oils and industrial sulphided cobalt-molybdenum and nickel-molybdenum supported on alumina catalysts to the development of a specially dedicated catfidjrtic system. Special attention is given to the catalytic aspects: the reaction schemes, the identification of the different catalytic functions and the control of the reaction. Finally, we discuss the utilisation of activated carbon as a new support. 1. WHY UPGRADE PYROLYSIS OILS? All biomass available as energy feedstock consists of chemically complex ligno-cellulosic materials. Their thermochemical and/or catalytic conversion produces gases, liquids and charcoal in various proportions. The liquids offer advantages in transport, storage, handling, retrofitting and flexibility of use. In addition they have a much Idgher energy density than the starting biomass. Fundamental studies such as those carried out by Shafizaded (1) indicate that high liquid yields from biomass can be obtained by pyrolysis, which is the thermal decomposition of the ligno-cellulosic matter either in the complete absence of oxidising agent, or with a limited supply in order to avoid gasification. In the last 10 years, liquid jdelds as high as 85% have been obtained by fast or flash pyrolysis (2-6) which involves extremely short residence times and extremely fast heat exchanges followed by rapid quenching. The valuable intermediary products are thus preserved before further repolymerisation.

100 These liqiiids, generally called oils, bio-oils, fast or flash p3n:olysis oils or biocrudes, rarely meet the standards required for fuels, but are nevertheless intended to be used as such in direct combustion in boilers, internal combustion engines or turbines. They have the aspect of a tar, they are viscous and not completely volatile and they do not mix with fossil fuels. In addition they are thermally imstable and tend to polymerise with time, temperature and light. These unwanted characteristics are related to the high oxygen content (up to 52%) present under the form of almost all oxygenated fimctions and as dissolved water (15-30%). Thus it is generally agreed that, to be used as fuels, bio-oils must be processed to remove oxygen. Two deoxygenation methods are currently proposed in the literature. One method proposes deoxygenation with simultaneous synthesis of gasoline-type compounds at atmospheric pressure through simultaneous dehydration-decarboxylation reactions over acidic zeolite catalysts and without reducing gases. This zeolite ZSM-5 is well-known for the production of gasoline from methanol. Its activity for the deoxygenation of other small oxygenated compoimds such as aldehydes and ketones has also been proved several times. At a typical temperature of 450 °C, oxygen is rejected as H2O, CO2 and CO (7). The maximal theoretical yield in hydrocarbons is 42% for flash p3rrolysis oils, but the literature indicates yields as low as 17-18% (8) because of the loss of carbon due to the high coke deposition (up to 15% of coke deposition on the catalyst and an extra 15% as suspended carbon) and because of the low conversion of the phenolic compounds (9). Moreover, the large molecules present in p3nrolysis vapours do not have access to the pores of the catalyst and, consequently, are not selectively converted. The other method proposed is hydrodeoxygenation (HDO) a t high temperature under hydrogen pressure in the presence of a catalyst (10). The reactions occurring are elimination of oxygen as water, elimination of nitrogen as ammonia, and hydrogenation-hydrocracking of large molecules. The reaction conditions and the catalysts (sulphided nickel-molybdenum or cobaltmolybdeniim supported on alumina) are similar to those used in the petrolexmi refining processes. The degree of deoxygenation can be easily modulated from simple stabilisation (elimination of more reactive functions such as carbonyl, olefins and carboxyles) to complete refining with a maximal theoretical hydrocarbon yield of 50%. This paper reviews the process of upgrading of bio-oils by catalytic hydrotreating from the first empirical tests to the recent development of new adequate catalytic systems. The reaction schemes, the potential inhibitors and the parameters enabling the control of the HDO reactions are discussed, as well as the catalyst deactivation by coke deposition. Finally, we discuss the utilisation of alternative neutral supports such as activated carbon. 2. THE EARLY YEARS The first studies concerning the upgrading by catalytic hydrotreating of vegetable oils were carried out in the early 80s' after the oil crisis. The main objective of these studies was to prove the feasibility of the hydrodeoxygenation process to produce hydrocarbonaceous fuels. Very little attention was paid to the setting up of the reaction conditions nor to the selection of adequate

101 catalysts, and most of the results were qualitative or obtained in inaccurate experimental conditions. The first quantitative and reproducible tests were performed by Elliott and Baker (11) in a flow reactor at 390°C and 13,5 MPa with hydroliquefaction oils. These oils contain more aromatics, polyaromatics and phenols than those produced by pyrolysis. This chemical composition is reflected by a low oxygen content (20%) compared to the high oxygen content of pyrolysis oils (52%). Different catalysts were tested: Ni, CoMo oxide, sulphided CoMo and NiMo, all supported on alumina. The best results were obtained with sulphided CoMo. The hydrogen consumption was 600 1 H2/1 fed oil and the yield in hydrocarbons was 75% of the fed oil. Later, on the basis of experiences with hydroliquefaction oils, the same authors hydrotreated oils obtained by p)rrolysis. However, the first experiences were relatively unsuccessful because of the extensive polymerisation of pyrolysis oils occurring at usual hydrotreating temperatures (12). The solution adopted by Elliott and Baker to avoid this thermal degradation was a pretreatment at lower temperature (270-280°C) aiming at eliminating more chemical functions such as aldehydes, ketones, carboxylic acids and esters (13). The yield in liquid hydrocarbons was 35% of the fed oil (dry basis), the maximal theoretical jdeld being 60%. This difference is due to the coke deposition leading to the quick deactivation of the catalyst. Elliott and Baker contributed greatly to the further development of hydrodeoxygenation process by setting up this two-step treatment, but very little attention was paid by these authors to the catalyst activation procedure and to keeping the sulphide state of the catalysts during the reaction. In fact, one important difference between pyrolysis oils and fossil fuels is that bio-oils do not contain any sulphur, thus H2S must be added during the process in order to preserve the sulphide state of the catalysts. Churin et al (14) hydrotreated bio-oils in a mechanically stirred batch reactor. The catalysts, industrial sulphided CoMo or NiMo supported on alumina, were directly suspended in the oil and CS2, which decomposes to H2S during the reaction, was added in order to maintain the sulphide state of the catalysts. This type of reactor does not allow a good contact between the liquid and the catalyst and, in addition, the concentration of the catalyst is low compared to fixed beds. In order to offset these problems, he used a hydrogen donor solvent (tetraline) which favours the hydrogen transfer, thus 40% yields in hydrocarbons were obtained. Later, it was demonstrated that tetraline could be replaced by a fraction of treated bio-oil or by a petroleum fraction such as diesel. Churin demonstrated that the hydrogen consumption begins at reaction temperatures of 200°C when a diluted bio-oil is treated. Other hydrodeoxygenation experiences were carried out in a batch reactor by Gagnon et al. (15). They treated bio-oils produced by vacuum pyrolysis which easily polymerise due to the high quantity of carbonyls. They developed a pretreatment at very low temperatures (60-100°C) using Ru/alumina catalysts. The stabilisation of the oil is explained by the hydrogenation of aldehydic functions in sugars contained in the oils. Gevert et al. (16) demonstrated that the alkalis contained in the bio-oils (ash) have a negative influence on the catalytic activity. They obtained good hydrocarbon yields (80%) when the oils were pre-treated by neutralisation.

102 Other groups also hydrotreated both Uquefaction and pyrolysis oils (17,18) in batch reactors but did not study the reaction parameters nor the catalysts or the process development. All these authors proved the feasibility of the deoxygenation by catalytic hydrotreating of oils produced by pjnrolysis. Important contributions such as the two-step treatment and the use of a solvent donor of hydrogen were carried out d u r i n g t h e s e early y e a r s . In p a r a l l e l , a n i n - d e p t h chemical characterisation of the oils was undertaken by different groups (19,20) in order to study the extremely complex composition of these oils containing hundreds of different molecules. Nevertheless, no optimisation of the reaction conditions nor of the catalytic system was done and, after this preliminary period, systematic studies leading to the optimisation and the scaling-up of the process became necessary. 3. THE MODEL COMPOUNDS APPROACH In 1989, Laurent (21) imdertook more fundamental studies concerning the hydrodeoxygenation of bio-oils in order to elucidate t h e main reaction pathways, the influence of the most important reaction parameters, the competition between different molecules and the possible inhibitors or poisons. This meant a complete kinetic and chemical study of the different reactions occurring during the upgrading process. For the whole study, he used industrial sulphided CoMo and NiMo supported on aliunina catalysts which had appeared to be the most adequate in the preliminary tests with real oils. Such a task would be extremely difficult, even impossible, with real oils because of the complexity of these liquids which are mixtures of hundreds of organic molecules: phenols, poly-substituted phenols, guaiacyls, aldehydes, ketones, carboxylic acids and esters, linear and aromatic ethers, sugars and others. In this case, model compounds representative of the real oil facilitate the analytical follow-up and the understanding of the different parameters. In addition, a lower concentration of the reactive chemical functions limits the polymerisation caused by the thermal reactions. It is very important that the tests carried out with model compound be as representative as possible of the real situation. This is why Laurent used model mixtures instead of isolated molecules, so as to establish the interactions and the competitions between molecules. Two different mixtures were defined, the first one contained the most reactive molecules: an aromatic ketone (4-methylacetophenone), a carboxylic ester (diethyl-decanedioate) and an aromatic ether (guaiacol). This "low temperature" solution mimicked the stabilisation step (< 300°C) prior to further hydrotreating. The second mixture, called "high temperature" solution, contained substituted phenols (4-methylphenol, 2ethylphenol) and dibenzofuran. It mimicked the total deoxygenation process requiring temperatures above 340°C. These molecules were selected on the basis of an in-depth chemical characterisation by fractionation into different chemical families followed by deep spectroscopic characterisation of each family (19).

103 4. FIRST STEP: REACTION SCHEMES The first part of the work performed by Laurent (21) in the frame of his PhD thesis was the elucidation of the hydrodeoxygenation reaction pathways of the four most abundant chemical functions in the bio-oils: phenol, carbonyl (ketone), carboxyl and methoxy. Special attention was given to the study of the catalytic activity and selectivity, as well as the possible competition between molecules. 4.1. Phenols The hydrodeoxygenation of phenols is a key reaction of the hydroprocessing of bio-oils when the production of highly refined hydrocarbon products such as transportation fuel is considered. Phenols and their derivatives represent an important portion of the organic part of these oils and, in addition, they are among the most difficult molecules to deoxygenate due to their high chemical stability. The hydrodeoxygenation process (HDO) has not been studied as much as hydrodenitrogenation (HDN) and hydrodesulfurisation (HDS), because oxygenated molecules are present in very low concentrations in petroleimi and because these molecules are not as harmful for the environment and for the catalysts as other heteroatom-containing molecules. Nevertheless, the HDO of phenolic compounds has been studied by several authors in the last years (2224). Two different catalytic functions of hydrotreating catalysts play a role in their conversion. One reaction scheme is the hydrogenation of the aromatic ring immediately followed by the deoxygenation of the intermediary cyclohexanol, which is generally not detected (25). The second pathway is the direct elimination of oxygen by hydrogenolysis of the Carom-0 bond. These two paths involve structurally different active sites (26,27).

Figure 1. Reaction scheme of 4-methylphenol (28) Laurent (28) studied particularly the hydrodeoxygenation of 4-methylphenol in solution with 2-ethylphenol and dibenzofuran in a batch reactor at 340°C

104 using sulphided CoMo and NiMo supported on alumina. CS2, which decomposes to H2S in the reaction conditions, was added in order to keep the sulphide state of the catalysts. 4-methylphenol is converted at a reasonable rate at 340°C, in typical runs its end conversion was 80-100%. 2-ethylphenol conversion rate was always 5 times lower than that of 4-methylphenol. Dibenzofiiran practically did not react. Figure 1 shows the hydrodeoxygenation reaction scheme for 4-methylphenol which is converted in toluene by hydrogenolysis and in methyl-cyclohexane by hydrogenation of the aromatic ring. 1 and 4-methyl-cyclohexene and 1,2cyclopentane are also produced by hydrogenation in minor amounts. In the expression of results, these minor products are grouped with methylcyclohexane. The intermediary 4-methyl-cyclohexanol has never been detected. Since these two path begin with strictly parallel reactions (fig. 2), a linear relation between toluene and methyl-cyclohexane is observed at moderate conversion rates. The deviation from the straight line at conversions higher than 50% is due to the hydrogenation of toluene to methylcyclohexane. This does not affect the measurement and expression of the selectivity methylcyclohexane/toluene, since Laurent performed it using the first conversion points.

T

0,0

0,1

0,2

0,3

'

I

0,4

0,5

CTOL

Co 4MP Figure 2. Relation between hydrogenation products and toluene (28) In the competition between molecules, the increase of 2-ethylphenol concentration causes a decrease of all reaction rates for both catalysts and results obtained by Laurent (28) indicate that this inhibition is more pronounced on the hydrogenolysis than on the hydrogenation. The influence of dibenzofuran was not considered by Laurent but, a priori y it has no specific influence in the conversion of 4-methylphenol since its concentration is almost constant and identical in all nins. 4.2. Carbonyl, carboxylic and guaiacyl grouiis The study of the hydrodeoxygenation of these groups is very important in the context of the upgrading of bio-oils, since they are the main cause of instability

105 and polymerisation because of their high chemical reactivity. The deoxygenation of these compounds leads to the stabilisation of the oils. This stabUisation could be the first step of a full refining process in order to avoid pol3rmerisation at the standard hydrotreating temperatures or, another interesting possibility, the reaction could be stopped at this stage leading to the production of a partially deoxygenated oil useful for the electricity production in turbines or diesel engines (10). However, the literature concerning the hydrodeoxygenation of these groups is very scarce. The reduction of carbonyls has been intensively studied in organic chemistry, as well as the transformation of ketones and aldehydes into alcohol over platinum group metal catalysts in very mild conditions (29). Maier et al studied the transformation of ketones in a methylene group over a metallic nickel catalysts (30). Weisser et al reported the hydrogenation of ketones over single metal sulphides (25). Concerning typical bimetallic hydrotreating catalysts, only Durand et al (31) worked on the hydrogenation of ketones over sulphided NiMo/alimiina. Laurent (32) studied the catalytic reaction schemes for these three groups using a model mixture containing 4-methyl-acetophenone, diethyldecanedioate and guaiacol. The tests were carried out in a batch reactor at different temperatures (260-300°C). The catalysts were industrial sulphided CoMo and NiMo supported on alumina. The sulphide state of the catalysts was maintained by addition of H2S. 4-Methyl-acetophenone: the conversion of the 4-methyl-acetophenone is very fast at 260°C, reaching 100% in less than 2 hours. The carbonyl is reduced to a CH2 with a very high selectivity. The only product observed is ethylmethylbenzene, the carbonyl group being hydrogenated to the alcohol which is quickly dehydrated under the reaction conditions. In addition, it is well known that yalumina catalyses dehydration reactions. The C=0 double bond hydrogenation is then the rate-limiting step (32). Figure 3 shows the reaction pathway.

,CHo

Figure 3. Hydrodeoxygenation reaction scheme for 4-methyl-acetophenone (32) Diethyl-decanedioate: according to the results obtained by Laurent (32), the hydrodeoxygenation of diethyl-decanedioate requires a temperature of aroimd 300°C over both NiMo and CoMo to be converted at a substantial rate. In fact, carboxyls are more refractory to deoxygenation t h a n carbonyls. Figure 4 shows

the reaction scheme leading to three linear alkanes: octane, nonane and decane. The intermediary products are Cs and C9 ethyl esters and their corresponding acids. It is reported in the literature (25) that carboxylic groups

106 are hydrogenated to CH3 groups with hemiacetal, aldehyde and alcohol as intermediary products which are quickly converted above 250^C on sulphide catalysts. These product have never been observed by Laurent. On the basis of these observations, Laurent proposes two mechanisms: one is the hydrogenation of the carboxylic group leading to a CH3 (reaction 1), the other is the rupture of the C-C bond leading to the complete decarboxylation (reaction 2). Octane is produced when reaction 2 occurs at both ends and decane si produced with reaction 1 at both ends. Nonane is produced by a combination of the two mechanisms. He defined the selectivity as the ratio between octane and decane. Laurent reports that a third reaction mechanism accounts for the production of carboxylic acids as intermediary. Using the corresponding pure acid (decanoic acid), he demonstrated that it is less reactive than the corresponding ester, giving slightly higher yields of decarboxylated products. Concerning the behaviour of the two catalysts, NiMo has a higher decarboxylating activity than CoMo, which could be due to the difference of acidity and the ensuing cracking activity between the two catalysts. Finally, he reported that the octane/decane selectivity slightly decreases with the conversion of the reactant.

^^ ( 1 ) 0 (2) O / H5C2-0fc^^(CH2)8 - C-0-C2H5^

O ( 1 ) ^ C10H22 H3C -(CH2)8 - C-O-C2H5 ^ ( > ^ ^ d ^ ^"20

(2) XH3C -(CH2)7 - C-O-C2H5 ^CT ( 2 ) ^ QH 18 Figure 4. Hydrodeoxygenation reaction schemes for carboxylic esters (32) Guaiacol: figure 5 presents the hydrodeoxygenation mechanism for guaiacol. The first step in the conversion of this molecule is the rupture of the C-CH3 bond leading to the formation of catechol which is then converted into phenol by elimination of one hydroxyl group. This phenol is subsequently transformed into benzene and cyclohexane. Under the reaction conditions used by Laurent, the conversion of guaiacol was limited to catechol and phenol (32). The total amoxmt of reaction products never accoimted for the converted guaiacol. Laurent reported 15-30% default in the molar balance at a conversion rate of 60%. Other authors reported 20% default at moderate conversion over CoMo at 250°C (33) and 10 to 50% conversion with the same catalysts at 300°C (34). These defaults in the molar balance can be related to the formation of heavy products or coke by analogy with the high tendency to form char during P)n:olysis of guaiacyls and hydroxyphenols (35).

107

Figure 5. Hydrodeoxygenation mechanism of guaiacol (32) Laurent (32) reported a decrease of the conversion rate of guaiacol over sxilphide catalysts. This decrease can be explained by the formation of coke and the ensuing blockage of active sites. But the production of catechol which adsorb strongly on y-alxxmina can also explain this deactivation (36). 5, SECOND STEP: ESIFLUENCE OF H2O, NHg AND HaS The refining or the stabilisation of highly oxygenated bio-oils require the total or partial elimination of oxygenated fimctions contained in molecules such as those above mentioned. The feasibility of this deoxygenation process has been proved with model compoimds at temperatures varying between 200-350°C over bimetallic sulphide catalysts. However, this reactivity can be different with real oils due to the presence of poisons or inhibitors: water which is dissolved in biooils (up to 30%), nitrogen which may be present in quantities as high as 3% in bio-oils and sulphur which is added during the hydrodeoxygenation reaction to keep the sulphide state of the catalysts (bio-oils contain extremely low quantities or even no sulphur). The influence of these three compounds, which are potential inhibitors or poisons but could also promote certain reactions, has been extensively studied in all hydrotreating reactions (24,37-41). Laurent (21,28,42) specifically studied their influence on hydrodeoxygenation reactions of model compounds representing bio-oils. Table 1 summarises the influence of ammonia, hydrogen sulphide and water on the different catalytic reactions involving the above mentioned model compoiinds. A scale going from — for a very strong inhibition to +++ for a very strong promoting effect has been used in the table in order to facilitate the comparison, 0 is used when there is no influence. As indicated in the table, water has no influence or very little inhibiting effect for all reactions except the hydrolysis of carboxylic esters which is promoted. Ammonia appears as a strong inhibitor of almost all reactions and this for both CoMo and NiMo. Surprisingly, it does not affect the hydrogenation of the ketonic group. Hydrogen sulphide has very little influence on the hydrogenation of ketones over sulphided CoMo, while it depresses the same reaction with NiMo. It has a promoting effect on the decarboxylation of carboxylic esters as well as on the hydrogenation of phenols. Demethylation of guaiacol is not affected.

108 Table 1 Inhibiting or promoting effects of NH3. H2S and H2O on the HDO reactions (21) Compound NH3 HgS HgO Reaction 4-Methyl-phenol hydrogenation

—

-0+

0

4-Methyl-phenol hydrogenolysis

—

~

0

4-Methyl-acetophenone hydrogenation

0

~

0

Carboxyl ester hydrolysis

—

0

++

Decarboxylation

—

+

0

Carboxyl ester hydrogenation

—

-

0

Guaiacol demethylation

~

0

0

These results suggest that both ammonia and hydrogen sulphide could be used for the control of the hydrodeoxygenation reactions: carbonyl groups could be selectively eliminated from complex feeds under a pressure of ammonia. On the other hand, carboxylic groups could be selectively eliminated by direct decarboxylation controlling the hydrogen sulphide pressure.

a CO

a o

0,05 0,10 0,15 0,20 H2S cone, (mole/1) Figure 6. Evolution of the hydrogenation (kj^cg • ) and hydrogenolysis (kôL ^^ rate constants of the CoMo catalyst as a function of H2S pseudo in concentration (28). Another important effect of H2S partial pressure is the control of the hydrodeoxygenation of phenols, since the hydrogenation/hydrogenolysis selectivity is strongly influenced, specially over CoMo. This control of the reaction could allow the control of the hydrogen consumption which is extremely important for the upgrading of real bio-oils. In fact, four molecules

109 of hydrogen are consumed via the hydrogenation pathway against only one via hydrogenolysis, the hydrogen consumption being multiplied by 2.3 in the H2S concentration range investigated (28). Figures 6 illustrates the influence of H2S partial pressure on the hydrogenation/hydrogenolysis selectivity over CoMo (28). Similar observations could be made for the conversion of carboxylic ester since t h e decarboxylation pathway consumes more hydrogen t h a n the hydrogenation one (42). & PROBLEM: DEACTIVATION BY COKE DEPOSITION Experiences performed in the early years with real bio-oils indicated a low stability of the catalytic system. Experiments could not r\in over a few days, nor even a few hours: the catalyst, embedded in coke, plugged the reactor. This could be explained by the still intense polymerisation occurring at the stabilisation temperatures (250-300°C), but also by the deactivation of the catalyst. It was t h u s necessary to understand this phenomenon. Laurent (21,28,32) calculated molar balances for each of the molecules he studied, and concluded t h a t these balances reached or were very close to 100% for the hydrodeoxygenation of the ketonic, carboxyl and hydroxyl groups.

100

150

200

250

300

Reaction time (min) Fig. 7. Evolution of the molar fraction of guaiacol (•), catechol (A), phenol (O) and their siim (H) as a function of the reaction time (CoMo, 280 °C), (32). The situation was completely different for guaiacol and catechol. As mentioned in section 4.2., up to 30% default in molar balance can be observed at a conversion rate of 60% of guaiacol. Figure 7 presents the molar balance for a typical run over CoMo, where a 20% default can be observed. The balance is still poorer with NiMo. This tendency to coke formation is tjîcal for phenols containing two or more oxygenated substitutes such as guaiacyls and hydroxyphenols. Klein et al. (33,34) gave valuable information concerning the

110 HDO reactions of anisole and guaiacol. They reported t h a t guaiacol reacts faster t h a n anisole but t h a t anisole is quantitatively converted, whereas for guaiacol it is not possible to close mass balances. Laurent and Centeno thus undertook a systematic study of this phenomenon of coke formation (43). They carried out experiences with guaiacol in the standard model mixture, guaiacol alone, hexadecane (solvent) alone, catechol, phenol and methylanisol over the traditional sulphided CoMo/y-alumina but also over y-alumina. Results obtained by these authors are presented in table 2. Two important conclusions can be drawn: 1) the coke deposition is very similar with guaiacol in mixture or alone, confirming that the other model molecules used by Laurent do not form coke and, 2) the mixture reacts over pure alumina with a poor rate constant but a high coke deposition, meaning t h a t the coke formation must be attributed to the aliunina support. Catechol has a reactivity similar to t h a t of guaiacol and also leads to a high coke formation. The catalysts used for phenol and methylanisol present a much lower coke content, indicating t h a t benzenic molecules containing only one oxygen have a lower propensity to this phenomenon. This coke would result from the condensation of guaiacol and catechol themselves. Table 2 Initial rate constant and quantities of coke deposited (43) Rate constant Reactant Catalyst (min-1. g.-l).103

Carbon content (%p) 1.8

Hexadecane (solvent) Guaiacol in mixture

C0M0/Y-AI2O3

9.4

8.9

Guaiacol in mixture

Y-AI2O3

2.9

10.3

Guaiacol

C0M0/Y-AI2O3

9.2

7.8 (85% conv.)

Catechol

C0M0/Y-AI2O3

1L7

5.5 (90% conv.)

Phenol

C0M0/Y-AI2O3

0

3.5 (0% conv.)

Methylanisol

C0M0/Y-AI2O3

19.7

2.8 (95% conv.)

C0M0/Y-AI2O3

7. THE FUTURE: A NEW CATALYTIC SYSTEM The conversion of guaiacol in experiences performed by L a u r e n t and Centeno can be attributed to the alumina support since, without active phase, it has a certain activity for the conversion of the reactant. Unfortunately, it also has a high activity for the formation of coke and heavy products. On this basis, neutral supports such as activated carbon or silica or even non supported catalysts could be a good alternative to avoid coke formation. Centeno et al.

Ill explored these possibilities (44). They prepared different catalysts: CoMo supported on activated carbon and on silica, and non supported. Table 3 presents results obtained with these catalysts compared to the traditional alumina.

Tables

Rate constants and phenol/catechol selectivity obtained with diflFerent cat. (44) Catalyst Rate constant Phenol/catechol (min'\ g. cat.'\cm^) (%) -^0 rAl203 0.35 12.6 1.30 COMO/Y-A1203 2.0 0.28 CoMo/Si 0.22 89.3 CoMo/C 0.39 8.0 CoMoS This table shows that the catalyst supported on activated carbon seems very promising because, even if its activity is very poor compared to alumina, it has a very high phenol/catechol selectivity. The other proposed catalysts do not show interesting possibilities since both activity and selectivity are low. Concerning the coke formation, the catalyst supported on carbon is also very interesting. Evidently, no values of carbon content are available for activated carbon-supported catalysts, but Centeno reported that the catalyst is not embedded by coke after reaction and remains active. In a recent congress (45), he presented the molar balance for the hydrodeoxygenation of guaiacol over CoMo/C compared to that obtained over CoMo/alumina (figure 8).

^

0,028 1

1

0,026
NM/HY2 > NM/HY7 > NM/HYio. The NM/Al catalyst presents high performances for overall reactions, again probably due to the enhancement of the acidity by unknown additives in the commercial catalyst. Neither of the NM/Ti nor NM/HM catalyst showed high performance for the upgrading of RC. The active sites of NM/Ti does not function effective to heavy feedstocks, although the superiority of a Ti02 supported Ni-Mo catalyst for HDS was reported in a recent paper [16]. Mordenite with strong acidity and smaller pore mouth (7 A) than Y-type faujasite (8 A) has often been used for hydroisomerization of light hydrocarbons [3]. Minja and Ternan [17] reported that the incorporation of mordenite into AI2O3 was effective for hydrodemetalization but not for hydrocracking of heavy oil. The strong acidity of the external surface of mordenite is presumably not appropriate for hydrocracking of polycyclic compounds. Table 6 illustrates the reaction results obtained for the hydrotreated feedstocks. Since the first-stage treatment was carried out in a semi-batch autoclave, light fractions produced during the hydrotreatment were removed with the hydrogen flow. This has resulted in the feedstocks for the second stage with relatively small portions of light fractions as shown in Table 2, while the asphalten and heteroatom contents are much reduced. Table 6 evidently displays that the HCK active sites function much more effectively in the reaction of H-VGO than in the reaction of VGO. Further, the M0/HY70 catalyst, HY (70wt%)-Al2O3 supported molybdenum sulfide catalyst without Ni promoter, exhibits higher HDS and HDN activities than the NM/Al and NM/HYQ catalysts. The most refractory nitrogen- or sulfide-containing

122 Table 6 Hydroprocessing of H-VGO and H-RC Catalyst NM/HYo NMmY2 NM/HY7 NM/HYio NM/Al Mo/HY7*^ H-VGO 254 138 277 70 59 59 Tio CC) 362 112 144 400 113 409 T50 ('C) 0.10 0.08 0.07 0.27 0.27 AH/C 4.5 13.5 5.7 3.7 H2 cons, (mg/g) 12.5 12.7 64 86 90 HDN (%) 87 91 91 94 81 74 98 98 79 HDS (%) H-RC 27 51 39 HDN (%) 0.11 0.05 0.08 AH/C 4.2 3.6 2.4 H2 cons. (mg/g)_ * 1: Mo (10 wt%)/HY(70%)-Al2O3 catalyst.

40 50 60 HY activity (%)

52 0.21 4.9

51 0.11 3.7


Fig. 1. H2Consumption during the reaction as a function of HYD activity a) O VGO A CL-VGO • SHALE (b) O RC ^ CL • SAND (a) 0 A over NM/Al) (

-

123 compounds are alkyl-substituted dibenzothiophenes or aery dines with steric hindrance [18, 19]. To break the steric hindrance, hydrogenation of the aromatic rings or removal of the alkyl groups is needed [20]. The above results suggest that the dual-functionality with high HYD activities are more suitable for deep HDS. The remarkable reactivity improvements of RC after the first stage treatment are evidenced by the larger AH/C values over the NM/HY7 and NM/HYio catalysts than the other catalysts. Also, HDN and AH/C during the reaction of H-RC increase with the increase of the zeolite content in the support. The H2 consumption in the reaction of RC does not result in the nitrogen removal or the increase in the H/C ratios of the product, while H2 in the reaction of H-RC is consumed to upgrade the liquid product properties. The HCK active sites of the zeolite thus function effectively in the reaction of H-RC. This is in good agreement with the fact that HCK is catalyzed prior to the complete hydrogenation of the aromatic rings over the acid sites [21] and metal sulfides [22 ]. The poor activities of the zeolite supported catalysts for the heavy feedstocks were often attributed to the pore diffusional limitation [3]. However, the molecular sizes of H-RC are not very different from those of RC, since the pre-hydrotreatment does not heavily crack the molecules. The zeolite-supported catalysts gave larger amounts of H2 consumption in the reactions of RC than in the reactions of H-RC in spite of much less total upgrading reactions. This indicates that the HCK of RC is catalyzed in the micropores, while other reactions take place on the external surface of zeolite which is poisoned during the reaction of RC. When the feedstock properties of RC and H-RC are compared, significant differences are observed in the sulfur and asphaltene contents. It is very unlikely that sulfur-containing compounds in RC poison the HCK active sites. The reactivity improvements of the RC by hydrotreatment is presumably due to the reduction of the asphaltene content; the polar nitrogen-containing compounds which are strongly adsorbed on the HCK active sites have been removed in the first-stage hydrotreatment. In the case of VGO, partial hydrogenation of the polyaromatic rings evidenced by the increase of the H/C ratio during the hydrotreatment probably results in a high degree of HCK. These results refer to the importance of hydrotreating reactions prior to the use of solid-acid catalysts. To further discuss the relationship between the feedstock properties and reactivity for heavy feedstocks, the hydroprocessing reactions of synfuels are discussed in the following session. 3.3 Hydrocracking and hydrotreatment of synfuels Fig. 1 shows the relationship between the HYD activity of the catalysts and hydrogen consumption during the reaction which is assumed to be an index for total upgrading reactions. For SHALE, CL-VGO and CL, H2 consumption increases with increasing HYD activity, whereas opposite relationships are observed for RC and VGO. An intermediate trend is observed for SAND. Comparison of the feedstock properties in Table 2 indicates that the reactivity differences among the feedstocks are very likely attributed to the nitrogen concentrations in the feedstock. In the reaction of VGO and RC with low nitrogen concentrations, the acid sites of zeolite catalyze HCK reactions with H2 consumption. In contrast, HYD is the major upgrading reaction for SHALE, CL-VGO and CL with high nitrogen concentrations. Figs. 2-5 show the HDS, HDN, AH/C and HDA activities over the catalysts as a function of HYD activity. It is evident that HDS (Fig. 2) and AH/C (Fig. 4) increase with increasing catalytic HYD activity except for the reaction of VGO over NM/HYQ. These reactions are

124

40 50 60 "^ 40 50 60 HY activity (%) HY activity (%) Fig. 2. HDS during the reaction as a function of HYD activity (a) O VGO • SHALE (b) O RC ^ CL • SAND ( ® A a overNM/Al)


40

50 60 HY activity (%)

Fig. 3. HDN during the reaction as a function of HYDactivity (a) O VGO A CL-VGO • SHALE (b) O RC ^ CL • SAND

40


40


Fig. 4. AH/C of feedstocks during the reaction as afiinctionof HYD activity (a) O VGO ^ CL-VGO • SHALE (b) O RC ^ CL • SAND

125

40 50 60 HY activity (%) Fig. 5. HDA during the reaction as a function of HYD activity O RC A CL • SAND

1

H f &

1

ô ^.' A

0-

95

53

43

46

0.048

2.4

1.24

NiMoA^-17

>95

>95

50

83

0.056

0.55

1.75

NiMo/Al203

-

-

-

-

0.075

sp

HY-5 + NiMo/Al203

r

1

0

20

. L —

1

1 ^^Sfa

40 60 80 Cracking Yield

100

Figure 3. Hydrocracking of n-heptane over NiMo catalysts containing zeolite Y-5. A second mixed catalyst (A") containing equal masses (0.17 g) of NiMo/Y-5 and NiMo/Al203 was elaborated with the hope of improving the hydrogenating capacity of catalyst NiMoA^-5 while keeping the same amount of acid sites. Accordingly, benzene conversion was higher with catalyst A" than with A. However the conversion of heptane was nearly the same on both catalysts (Figure 3a). Furthermore, they exhibited comparable yields of heptane isomers (Figure 3b). Similarly, two mechanical mixtures were realized with catalysts containing zeolite Y-17. The mixture (B') containing 0.17 g of the neat HY-17 and twice as much NiMo/Al203 converted heptane at a lower rate than 0.17 g of NiMoAf-17 (Figure 4a). Internal NiMo

154 appeared once more preferable, the rate over NiMoAf-17 being 5 times that of the mixture B'. The product distributions in tables 2 and 3 show that the reaction proceeds through similar reaction mechanisms. The selectivity for isomers observed with B' was consistently low, but not significantly under that of NiMoA^-17 (Figure 4b). Finally, mixture B" was made by admixing 0.17 g of NiMoA^-17 with a two-fold excess of NiMo/Al203. As for NiMoA^-5, the conversion of heptane over NiMoA^-17 was not improved by adding external NiMo (Figure 4a). However, a slight improvement of the selectivity was observed (Figure 4b). b - Selectivity for heptane isomers

a - Conversion of n-heptane 20

100 80 o

60 h

15 NiMoA^-17 + NiMo/Al203

NiMoA^-17 + NiMo/AhCb NiMoA'-n

NiMoA'-n

c o U

o 40 hHY-17 + NiMo/Al203

HY-17 + NiMo/Al203|

ao 20 0 240

260

280

300

Temperature (°C)

320

340

20

40

60

80

100

Cracking Yield

Figure 4. Hydrocracking of n-heptane over NiMo catalysts containing zeolite Y-17. 3.4. Distance between hydrogenating and acidic sites The conversion of heptane over mechanical mixtures A' and B' of the HY zeolites and the NiMo/Al203 proves some cooperation between the acidic and hydrogenation centers, since the level of activity was much higher than on either component. The NiMo zeolites are therefore bifunctional catalysts, and the hydroconversion proceeds through heptene intermediates. Good hydrocracking catalysts should possess a high hydro-dehydrogenation activity, together with a short distance between the metallic and the acidic sites. The first condition is clearly not fiilfiUed with the NiMoY catalysts where cracking prevails over isomerization. Thus, the steady concentration CQ of linear heptenes at the acid sites is likely well under the limit C^q at equilibrium with n-heptane. Therefore, the desorption of isoheptenes (step d) is rather slow. As mixtures A' and B' exhibited lower activities and selectivities than the NiMoY, it may be inferred that the migration of heptene toward the inside of the zeolite is rate limiting. In this case, the longer intersite distance results in CQ values lower than in the corresponding NiMoY. Some quantitative confirmation of the transport limitation in those mixtures may be sought fi"om the Weisz's criterion for polystep reaction systems (19), which is

155

dt Ceq D where dN/dt is the rate per unit volume, R the radius of the zeolite particles, and D the effective difRisivity of linear heptenes. If diffusion is slow, O will be much greater than 1. O was evaluated for the conversion of heptane at 320 °C over the mixed catalyst A' containing support HY-5. Measurements by light scattering afforded a surface-averaged radius R = 1.15 ^m. Diffusivity D was estimated at 10"^ m^ s'^ from NMR measurements on NaX (20). An equilibrium concentration Cgq = 1.7 10"^ mol m"^ was assessed from thermodynamic data (21). With the observed value dN/dt = 0.65 mol s"^ m"^, we deduced O = 0.5. This not clearly in favor of diffusional limitation. Ruckenstein for instance reported a value O = 0.23 for the conversion of methylcyclopentane at 470'^C and 1 atm over a p zeolite (22) and concluded to the absence of limitation. The same synergism between Pt/Al203 and various zeolites was reported for light alkane isomerization (23). However, with our A' and B' mixtures, some limitation is obvious from the rate and selectivity data. The low O value may be due to a poor estimate of D. Moreover C^q in the above formula really stands for the actual concentration CQ outside the zeolite grains; this may be well under Cgq since dehydrogenation over NiMo/Al203 is not fast enough. Therefore, in the NiMo-HY system, the intersite distance appears as the main factor governing the activity. In both NiMo zeolites, the NiMo sulfide seems rather well dispersed and the two kind of sites cooperate efficiently. There, the alkene concentration CQ is higher than with mixtures A' and B', although substantially below the equilibrium value. The internal Co is hardly enhanced when NiMo/Ai203 is admixed with the NiMoY zeolites because then the distance between the acidic and external NiMo is too high and transport limitation occurs. Such transport limitation however is less severe with catalysts containing HY-17 simply because they are less active. Accordingly, hydroisomerization on mixture B" is slightly improved over that of NiMoA^-17. The above conclusions were deduced from the heptane test performed at low temperature, and will not necessarily hold under more severe conditions. For instance, the homogeneous distribution of the NiMo inside zeolite Y appeared favorable for the conversion of heptane. A different behavior was noted when methylcyclopentane reacted over a Pt mordenite (24). There, because of coking, a distribution of Pt at the periphery of the grains was preferred. Besides, external NiMo proved useful in industrial hydrocracking catalyst (6, 25) probably because it improves the diffusivity of large molecules and the resistance to coking. REFERENCES 1. H.L. Coonradt and W.E. Garwood, Ind. Eng. Chem. Prod. Res. Dev., 3 (1964) 38. 2. J.F. Le Page, J. Cosyns, P. Courty, E. Freund, J.P. Franck, Y. Jacquin, B. Juguin, C. Marcilly, G. Martino, J. Miquel, R. Montamal, A. Sugier and H. van Landeghem, Applied Heterogeneous Catalysis, Technip, Paris ,1987. 3. J. Weitkamp, Hydrocracking and Hydrotreating, J.W. Ward and G.A. Quader (eds.), ACS Symp. Ser., Vol. 20, p. 1, Amer. Chem. Soc, Washington DC, 1975. 4. J.P. Franck, M. El Malki and R. Montamal, Rev. Inst. Fr. Petr., 36 (1981) 211.

156 5. P. G. Smimiotis and E. Ruckenstein, Catal. Lett., 17 (1993) 341. 6. P. Dufresne, A. Quesada and S. Mignard, Catalysis in petroleum refining 1989, Kuwait, 1989, D.L. Trimm, S. Akashah, M. Absi-Halabi and A. Bishara (eds.). Stud. Surf. Sci. Catal, Vol. 53, p. 301, Elsevier, Amsterdam, 1990. 7. M. Guisnet, J.L. Lemberton, C. Thomazeau and S. Mignard, Proc. 9^ Int. Zeol. Conf, Montreal, 1992, R. von Ballmoos, J.B. Higgins and M.M.J. Treacy (eds.). Vol. 2, p. 413, Butterworth/Heinemann, Boston/London, 1993. 8. J. Leglise, M. El Qotbi and D. Comet, Coll. Czech. Chem. Commun., 57 (1992) 882. 9. W.J.J. Welters, O.H. van der Waerden, V.H.J. de Beer and R.A. van Santen, Ind. Eng. Chem. Res., 34 (1995) 1166. 10. J. Leglise, M. El Qotbi, J.M. Goupil and D. Comet, Catal. Lett., 10 (1991) 103. 11. J. Leglise, J.M. Manoli, C. Potvin, G. Djega-Mariadassou and D. Comet, J. Catal., 152 (1995) 275. 12. W.J.J. Welters, O.H. van der Waerden, W. Zandgergen, V.H.J, de Beer and R.A. van Santen, Ind. Eng. Chem. Res., 34 (1995) 1156. 13. R. Beaumont and D. Barthomeuf, J. Catal., 26 (1972) 218. 14. J. Weitkamp, Ind. Eng. Chem. Prod. Res. Dev., 21 (1982) 550. 15. G.E. Giannetto, G.R. Perot and M. R Guisnet, Ind. Eng. Chem. Prod. Res. Dev., 25 (1986)481. 16. J.A. Martens, P.A. Jacobs and J. Weitkamp, Appl. Catal., 20 (1986) 239. 17. G.F. Froment, Catal. today, 1 (1987) 455. 18. A Lopez Agudo, A. Asensio and A. Corma, J. Catal., 69 (1981) 274. 19. P.B. Weisz, Adv. Catal., 13 (1962) 137. 20. M.F.M. Fost, Introduction to zeolite science and practice, H. van Bekkum, EM. Flanigen and J.C. Jansen (eds.). Stud. Surf. Sci. Catal., Vol. 58, Ch. 11, p. 391, Elsevier, Amsterdam, 1991. 21. RA. Alberty and C. A. Gehrig, J. Phys. Chem. Ref Data, 14 (1985) 803. 22. P. G. Smimiotis and E. Ruckenstein, Ind. Eng. Chem. Res., 33 (1994) 493. 23. G. Perot, P. Hilaireau and M. Guisnet, Proc. 6^ Int. Zeol. Conf, Reno, 1983, D. Olson and A. Bisio (eds.), p. 427, Butterworth, Guildford, 1984. 24. B.A Lemer, B.T. CaviU and W. M. H. Sachtler, Appl. Catal., 21 (1994) 23. 25. A. Nishijima, S. Yoshitomi, H. Shimada, Y. Yoshimura, T. Sato and N. Matsubayashi, Proc. 9* Int. Cong. Catal., Calgary, 1988, MJ. Phillips and M. Teman (eds.). Vol. 1, p. 174, Chemical Institute Canada, Ottawa, 1988.


157

ACIDITY INDUCED BY H2S ADSORPTION ON UNPROMOTED AND PROMOTED SULHDED CATALYSTS Cyrille Petit, Frangoise Mauge* and Jean - Claude Lavalley Laboratoire SPECTROCAT - URA CNRS 414 - ISMRA - Universite de Caen 6, Bd du Marshal Juin -14050 CAEN cedex (France)

The effect of H2S adsorption on the nature and the number of acid sites of alumina supported sulfided catalysts was characterized by IR spectroscopy. For monitoring the variation of the surface properties, dimethyl-2,6-pyridine, a probe molecule more convenient to detect Bronsted acidity than pyridine, was adsorbed at room temperature (r.t). Before H2S introduction, a few number of Bronsted acid sites was detected, particularly on the promoted catalysts. Whatever the catalysts, H2S adsorption at r.t led to the creation of Br0nsted acid sites at the expense of Lewis acid sites. The concentration of the created Bronsted acid sites depended on the presence and on the nature of the promoter. The highest number of Bronsted acid sites was detected on CoMo/AI2O3. 1 - Introduction The activity of hydrotreating catalysts was generally related to the number of coordinatively unsaturated sites (1). In particular, IR spectroscopy of probe molecules like CO allowed us to differentiate unpromoted and promoted coordinatively unsaturated sites and good correlations between absorption bands characteristic of promoted sites and HDS activity were previously obtained (2-4). Nevertheless, studies of the catalytic properties of sulfided catalysts indicate the importance of Bronsted acidity (5). Moreover, the concentration of sulfur vacancies can be modified under H2S which is present during the reaction since the formation of Bronsted acid SH groups is then expected. The direct detection of SH groups is difficult by IR spectroscopy and the adsorption of a probe molecule able to be protonated by such groups is a way to assess their presence. Pyridine is generally used and it was shown that at temperatures typical of the reaction, pyridinium species were formed on M0/AI2O3 and C0M0/AI2O3 (6,7). Previous studies pointed out that dimethyl-2,6-pyridine (DMP) was more sensitive to the presence of Bronsted acidity than pyridine (8,9). For example, we showed that no Bronsted acid sites were detected by pyridine adsorption on Ti02-Zr02 mixed oxides whereas, in the same conditions, the

158 formation of protonated species was observed by adsorption of dimethyl-2,6pyridine (10). This can be related to the higher basicity of this probe molecule (pyridine, pKa = 5.2 - DMP, pKa = 6.7). Moreover, it has been proposed that the presence of methyl groups adjacent to the nitrogen atom is unfavorable for coordination of DMP. The aim of the present work is to determine if dimethyl-2,6-pyridine is a suitable probe molecule to characterize the weak Bronsted acidity of sulfided catalysts and to establish if the presence of H2S increases the Bronsted acidity of these catalysts. 2. Experimental The catalysts were either prepared in our laboratory, by pore filling impregnation of a y-alumina : M0/AI2O3 (9.3% Mo), C0/AI2O3 (3.1% Co), NiW/Al203 (2.5% Ni - 16.6% W) or provided by Procatalyse CoMo/AI2O3 (HR 306) and NiMo/Al203 (HR 346). The samples were pressed into self-supporting wafers (disc = 2 cm^). All the studied samples were sulfided in situ in the IR cell. Firstly, the catalysts were evacuated at increasing temperature from 298 to 473 K (ramp : 0.15 K.s-^). Then, 80 torrs of a mixture of H2 + H2S (15% vol.) was introduced. The temperature was raised from 473 to 673 K (ramp : 0.15 K.s"l) with a plateau of one hour at 673 K. Subsequently, the catalysts were evacuated at 673 K for 10 min. Two other sulfidations for 2 hours and 15 hours were performed at 673 K. The ultimate sulfidation was followed by an evacuation at 673 K for one hour to 3 10"^ torr before cooling at room temperature under evacuation. Dimethyl-2,6-pyridine (3 torrs) was contacted with the catalyst and then evacuated at room temperature (r.t.) for 15 minutes (P = 210"^ torr). Calibrated small doses of H2S were successively introduced and finally a pressure of 15 torrs of H2S was contacted with the disc. This was followed by an evacuation at r.t. for 15 minutes (P= 8.10-^ torr). Spectra were recorded on a Fourier transform infrared spectrometer Nicolet 60 SX. For sulfided catalysts, the spectra reported below correspond to a mass of 10 mg whereas for alumina, they correspond to a mass of 8 mg which is almost the mass of the support present in a disc of 10 mg of a sulfided catalyst. 3. Results 3.1. Effect of H2S adsorption on various metal oxides Previous infrared studies on the adsorption of dimethyl-2,6-pyridine (DMP) reported that the wavenumbers of the bands of this probe molecule depended on the adsorption site, in particular, those corresponding to the vga and V8b vibrations occurring in the 1660-1580 cmr^ range (8-12). On pure alumina or

159 ^ 1580 A

b s o r b a n c e

1650

1600

1550

Wavenumbers (cm"^)

B

1602 ; 1613, /\ !

0.04

1

1654

A

J

•

1

1

'

t

"

^ 'A

/^'

\ /

1

by /aT.'
M0/AI2O3 > C0/AI2O3 > AI2O3 Negative bands observed at 1613,1602 and 1580 cm'^ correspond to a decrease of the number of Lewis acid sites and DMP weakly adsorbed. This decrease is smaller on Mo/AI2O3 and C0/AI2O3 than on alumina and maximum for CoMo/AI2O3. The effect of introduction of small doses of H2S on the C0M0/AI2O3 catalyst is presented in figure 4A. It shows a continuous creation of Bronsted acid sites at the expense of the bands characterizing Lewis acid sites and DMP species in weak interaction with the surface. In order to follow more quantitatively this effect, we studied the intensity variation, for each dose introduced, from spectra resulting from the subtraction of those obtained after and before H2S introduction, as previously described. The area of the bands corresponding to the creation of

162

1650 1600 1550 Wavenumbers (cm"^)

A1203

•

C0/AI2O3

M0/AI2O3

•

C0M0/AI2O3

H2S introduced (nmoles)

Figure 4 : A - Effect of the introduction of small doses of H2S on the spectra of DMP adsorbed on a C0M0/AI2O3 catalyst. B - Variation of the area of the bands characterizing the creation of Brjlnsted acid sites versus the amount of H2S introduced for the various components of a CoMo/ AI2O3 catalyst. Bronsted acid sites is measured between 1680 to 1615 cm'^ and its variation with the amount of H2S introduced is reported in figure 4B. For all the catalysts, the variation of the number of Bronsted acid sites versus the number of micromoles of H2S introduced is close; it quickly reaches saturation. Nevertheless, it reaches saturation for higher H2S doses on Mo and C0M0/AI2O3 (~10 jimoles) than on alumina and Co/AI2O3 (2 jimoles). This is in agreement with the higher amount of Bronsted acid sites created on M0/AI2O3 and C0M0/AI2O3. We remark, for these two catalysts, that the number of Bronsted sites created is similar. Nevertheless, figure 4B shows that, after evacuation of H2S at r.t, the number of protonated species decreases on Mo/AI2O3 whereas it does not change after evacuation on C0M0/AI2O3. This could indicate that the created Bronsted acid sites are stronger on the promoted catalyst than on Mo/AI2O3. 3.3 Effect of H2S adsorption on various sulfided catalysts This study was extended to the sulfided NiMo/Al203 and NiW/Al203 catalysts (figure 5). DMP adsorption leads to the same bands as on sulfided C0M0/AI2O3. The intensity of the massif of bands at 1613, 1602 and 1580 cm-^ is higher on NiMo/Al203 than on CoMo/AI2O3 whereas on NiW/AI2O3, it is weaker. By contrast, the band near 1650 cm"^ which characterizes Bronsted acidity, presents the same intensity on the three catalysts before H2S introduction at r.t.

163

Ai

1625 1646/\1613 1580 1602 r

1650

1600

1550

1650

1600

1550

Wavenumbers (cmO

Wavenumbers (cm-^

Figure 5 : Effect of the introduction of H2S on the spectra of DMP adsorbed on various sulfided catalysts: ( ) Before H2S adsorption; ( ) After introduction of 15 torrs of H2S followed by an evacuation at r.t.

Figure 6 : Effect of H2S or CH3SH on the spectra of DMP adsorbed on M0/AI2O3 : a - DMP species present after evacuation at r.t. ; b - after H2S adsorption followed by an evacuation at r.t ; c : after CH3SH adsorption followed by an evacuation at r.t.

The effect of H2S adsorption is qualitatively the same. Nevertheless, figure 5 indicates that the highest amount of Bronsted acid sites is created on CoMo/AI2O3 whereas the number of sites present on NiMo/Al203 and NiW/AI2O3 is similar. 3.4 Effect of CH3SH adsorption on sulfided M0/AI2O3 The previous experiments clearly show that H2S addition creates Bronsted acid sites on sulfided catalysts. However, these experiments do not allow us to specify if H2S adsorption occurs dissociatively or not. Indeed, the intensity of the v(SH) band due to adsorbed H2S is so weak that no information can be drawn concerning the H2S adsorption mode. The v(SH) band of CH3SH seems more intense, in particular when it is coordinatively chemisorbed. On pure alumina, O. Saur et al. (13) showed that the first doses of CH3SH adsorbed dissociatively since the v(SH) band was not detected, while for higher amounts of CH3SH introduced, its adsorption became undissociative since the v(SH) band was then detected at 2560 cm'^. Therefore, instead of adsorbing H2S, we introduced CH3SH in order to study the v(SH) band. Adsorption of CH3SH at r.t. leads to the same results as for H2S concerning the DMP adsorption (figure 6) : the intensity of bands at 1613, 1602 and 1580 cm-^ decreases while that of the band near 1650 cm-^ increases. The number of Br0nsted acid sites created by H2S and CH3SH introduction is close. We note, from

164 the intensity of the v(CH3) bands, that CH3SH adsorption resists to evacuation at r.t. whereas no band around 2560 cm"^ can be detected. This suggests that CH3SH adsorbs dissociatively on this sulfided catalyst. This result could be extended to H2S adsorption, at least for the first doses introduced. 3.5 Effect of H2 adsorption on sulfided Mo and C0M0/AI2O3 Following the same procedure as that used for H2S adsorption, hydrogen has been introduced at r.t on sulfided M0/AI2O3 and C0M0/AI2O3. Contrarily to H2S, H2 introduction does not induce any modification of DMP adsorption : no new species are formed and the number of coordinated species and weakly adsorbed species stays constant This absence of interaction of hydrogen at r.t with the sulfided catalyst surface validates the method used. Indeed, it points out that DMP species interact sufficiently strongly with alumina or the sulfided phases to resist to the introduction of a neutral gas whereas the adsorption is sufficiently weak to be sensitive to H2S introduction.

4. Discussion Sulfidation of alumina at 673 K, followed by an evacuation at the same temperature, does not strongly modify its acidic properties since DMP leads to similar species when adsorbed on sulfided or pure alumina (10). This is in agreement with Ziolek et al. (14) who reported close catalytic properties of alumina before and after sulfidation as well as a very low sulfur content Nevertheless, H2S adsorption at r.t induces some modifications on the alumina acidic properties since it slightly increases the number of Bronsted acid sites and decreases that of Lewis acid sites. On sulfided catalysts, DMP adsorption at r.t provides evidence for the presence of some Bronsted acid sites. These sites are maximum on sulfided C0M0/AI2O3. This shows that such sites are formed on the sulfided phase and it indicates the presence of residual SH groups. Further adsorption of H2S at r.t followed by an evacuation at r.t. leads to the formation of supplementary Bransted acid sites. The number of the sites created varies in the following order : C0M0/AI2O3 > M0/AI2O3 > C0/AI2O3 > AI2O3. These results show that, although H2S adsorption also induces the creation of Bronsted acidity on alumina, the presence of M0S2 slabs favors the creation of Br0nsted acid sites particularly when they are promoted by cobalt atoms. Therefore, such sites are involved in the apparition of acidic SH groups. Introduction of small doses of H2S reveals that the same niunber of sites is created on M0/AI2O3 and C0M0/AI2O3. However on C0M0/AI2O3, the protonated species are more stable. This could indicate that the SH sites present on C0M0/AI2O3 are stronger than on M0/AI2O3. By analogy with results obtained with CH3SH, a dissociative adsorption of H2S on sulfided M0/AI2O3 is suggested. It can occur on couples of coordinatively unsaturated (cus) Mo and S sites. We note that the number

165 of SH sites present after evacuation varies in accordance with the catalytic activity measured in thiophene hydrodesulfurization of these catalysts (15). The introduction of small doses of H2S provides also evidence for a simultaneous decrease of the number of Lewis acid sites and weakly adsorbed species. Taking into account the results obtained on silica, we can proposed that H2S displaces the weakly linked species without creation of Brensted acidity, while the poisoning of cus sites present on alumina and on the sulfided phases leads to the creation of Bronsted acid sites. Comparison with results obtained from H2S adsorption on NiMo/Al203 and NiW/Al203 shows that the number of Bronsted sites created is maximum on C0M0/AI2O3. This result can be related to a study of Van Gestel et al. (16) on the sensitivity to the H2S amount of NiMo/Al203 and C0M0/AI2O3 in HDS of thiophene. At high H2S levels, these authors provide evidence for a higher efficiency of sulfur sites created on CoMo than on NiMo catalysts. In our conditions, hydrogen introduction does not induce any acidity. This does not mean that hydrogen does not generate acidity in conditions closer to those of the reaction. 5. Conclusion This study shows the presence of Bronsted acid sites on the support and more specifically on the sulfided phases, promoted or not. Adsorption of H2S at r.t. increases the number of such sites at the expense of Lewis acid sites. Our results suggest that H2S adsorbed dissociatively, likely on cus Mo and S couples. The number of Bronsted acid sites created by H2S adsorption depends on the catalyst. For the Co-Mo catalysts, their variation is correlated to their activity for HDS of thiophene : C0M0/AI2O3 > M0/AI2O3 > C0/AI2O3 > AI2O3. Finally, on NiMo/Al203 and NiW/Al203, H2S adsorption induces a number of Bronsted acid sites lower than that created on CoMo/AI2O3. Acknoledgements The authors thaiJc J. Van Gestel for stimulating suggestions. REFERENCES 1 - R. R. Chianelli, Catal. Rev.-Sci. Eng., 26 (1984) 361. 2 - F. Mauge, A. Vallet, J. Bachelier, J.C. Duchet and J.C. Lavalley, J. Catal., in press. 3 - F. Mauge, J.C. Duchet, J.C. Lavalley, S. Houssenbay, E. Payen, J. Grimblot and S. Kasztelan, Catal. Today, 10 (1991) 561. 4 - J.A. De Los Reyes J.A., M. Vrinat, M. Breysse, F. Mauge and J.C. Lavalley, Catal. Lett., 13 (1992) 213.

166 5678910 11 1213 14 15 16 -

F.E. Massoth and G. Muralidhar in Fourth International Conference on Chemistry and Uses of Molybdenum (H.F. Barry and P.C.H. Mitchell, Eds.) p.343. Climax Molybdenum Co., Ann Arbor, MI, 1982. N.Y. T0ps0e, H. Topsee and F.E. Masoth, J. Catal., 119 (1989) 252. N.Y. T0ps0e and H. T0ps0e, J. Catal., 139 (1993) 641. P.A. Jacobs and C.F. Heylen, J. Catal., 34 (1974) 267. E.R.A. Matulewicz, F.P.J.M. Kerkhof, L.A. Mouljin and H.J. Reistma, J. Colloid Interface Chem., T7 (1980) 110. C. Lahousse, A. Aboulayt, F. Maug^ , J. Bachelier and J.C. Lavalley, J. Mol. Cat, 84 (1993) 283. A. Corma, C. Rodellas and V. Fomest, J. Catal., 88 (1984) 374. S. Jolly, J. Saussey, J.C. Lavalley, N. Zanier, E. Benazzi and J.F. Joly, Ber. Bunsenges. Phys. Chem., 97 (1993) 313. O. Saur, T. Chevreau, J. Lamotte, J. Travert and J.C. Lavalley, J. Chem. Soc. Farad. Trans. 1, Tl (1981) 427. M. Ziolek, J. Kujawa, O. Saur and J.C. Lavalley, J. Mol. Catal., 97 (1995) 49. J. Bachelier, M.J. Tilliette, M. Cornac, J.C. Duchet, J.C. Lavalley and D. Cornet, Bull. Soc. Chem. Belg., 93 (1984) 743. J. Van Gestel, L. Finot, J. LegHse and J.C. Duchet, Bull. Soc. Chim. Belg. 4-5 (1995) 189.

® 1997 Elsevier Science B. V. All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F, Froment, B. Delmon and P. Grange, editors

ORGANO METALLIC SILOXANES AS AN ACTIVE COMPONENTS OF HYDROTREATING CATALYSTS. Kolesnikov, I. M., Yablonsky, A.V., Sugungun, M. M., Kolesnikov, S. L, Kilyanov, M.Y State Gubkin Academy of Oil and Gas 117296 Moscow, Leninsky prospect 65, RUSSIA.

ABSTRACT Structural and contents optimization of hydrotreating catalysts and their subsequent activation with organo metallic siloxane based on the theory of catalysis by polyhedra was discussed.

INTRODUCTION Industrially, for the removal of sulfur and or sulfur compounds from oil fractions under an increased hydrogen pressure and temperature a process known generally as hydrotreatment, various catalysts are used which may contain a mixture of Al-Mo-0 or Al-Ni-Mo-0, and even catalysts containing zeolite additives and many others [1,4 ]. Metal sulfides of Ni-W, Cu-W, Co-Mo, Ni-Mo and others are used for hyrodesulfurization of sulfurous feed [ 5,6 ]. The activity of the oxides catalysts is determined by the ratio of metal oxides in them, type of the metal and operating regimes of the hydrotreatment process in general. The optimal content of the catalysts can be determined by the theory of catalysis by polyhedra put forward by one of us in the 1960s [7,9 ]. Active centers of the hydrotreating catalysts based on one of the postulates of the theory were considered to be the ensembles of the following tetrahedras: [ NiO4.MoO4.AlO4 ] , [ M0S4AIO4 ], [ NiS4A104], [ WO4.M0O4.AIO4 ] and others. Tetrahedras of the types : [ C0O6.M0O4.AIO6 ], [ Ni06.Mo06.A106], [ M0S6AIO6 ] are less active [7,8]. Catalysts are prepared by impregnating Co, Mo and Ni salts to y- AI2O3, that allows a sequential shaping of the polyhedral structures. Uncontrolled interaction of the impregnated salts with y- AI2O3 lattice will results to an alternate and chaotic distribution of both the active and non-active polyhedras at the surface of the carrier and makes the activity regulation much more tedious. Thermodynamic method for optimization of hydrotreating catalysts and the application of organo metallic siloxane as both a catalyst and activator are discussed in this paper.

167

168 EXPERIMENTAL Synthesis of organo metallic siloxanes: Synthesis of the organo metallic siloxanes is carried under laboratory conditions using the following technique:Solutions of metal chlorides ( C0CI2, M0CI5, NiCb or FeCh ) are prepared and dissolved in a pure absolute acetone. Powdered aluminophenylsiloxane (APS ) is dissolved in a pure ethyl spirit. Both solutions are then mixed. The reaction mixture is heated for 1 hour and the solvents (acetone and ethyl spirit) were distilled out. A mixture of aluminophenylsiloxane (APS ) and chloride compounds at the following molar ratios were synthesized. APS :CoCl2= 1.0: 0.15 APS : MoCl = 1.0 : 0.3 ; 1.0 : 0.5 ; 1.0 : 1.0 APS : C0CI2: WCl6 = 1.0 : 0.003 : 0.04 Infra-red spectra within thefrequencyinterval of 400-3600 cm"^ was carried out. Molecular weight of each of the sample compound was determined. Ferro-Molybdenum catalysts were synthesized by crafting FeCls and M0CI5 salts to the surface of the Y-AI2O3 carrier. Surface areas were determined on "Sorptomatic". Activities of the organo metallic siloxanes were determined in the homogenous thiophene desulfiirization from toluene. Catalysts activities in hydrothermal decomposition of sulfiir compounds in micro flow reactors in the stream of helium gas were determined. Chromatographic analysis of the products was carried out off-line. THERMODYNAMICS OF POLYHEDRA TRANSFORMATION Activity of the catalysts (types: C0O.M0O3.AI2O3, NiO.MoGB.AbOB ) is defined by the ability of ensemble of theirs tetrahedras to redistribute electrons as in the scheme: electron

[iMe^^04] + [2Me^^04] M*>/v'»W^>**vl^'

15

v * ^ j » 'y>;vY^^.» f^Âj^^^^mtti-Â

-10

10

(b)

6(ppm)

15

-10

10

Figure 7. ^^P-RMN Spectra of: (a) H3PM012O40 +l,5Co(N03)2, (b) C03/2PM012O40

In Figure 8 is reported the NMR spectrum of the H7PM0 which shows a main line at - 12.3 ppm, characteristic of the existence in solution of the P-form of the four-electron reduced heteropolyacid [14].

5(ppm)

- 2 - 4 - 6 - 8

-10

-12

-14

Figure 8. '^P-RMN spectra of H7PM0

190 It is well known that the reduction of 12-molybdophosphoric acid results in the transformation of the a-form (Keggin structure) into the P-form because the reduced P4-form (i.e. four-electron reduced)) is more stable than the a-form (a4) [14]. This result confirms the presence of a four-electron reduced molybdophosphoric acid and indicates that this compound remains reduced when it is dissolved in water. A line is also observed at - 5.32 ppm which correspond to the a2-form i.e. reduced by two electron [14] (with a molar ratio a4/ P4 of about .2) It should be noticed that the solutions which contain a reduced heteropolycompound are impregnated right after the dissolution of the polyanion in water. The spectrum of the C07/2PM0 (not reported here) is similar to the spectrum of H7PM0 The only difference results in the shift (A5 « 18 ppm) and a broadening of the lines which are due to the presence of cobalt. However, the shift is greater than it would be expected if we only consider the magnetic susceptibility. The broadening of the line, which is in relation with the relaxation time, is ascribed to the strong interaction of the Co with the heteropolystructure. So it can be deduced that the reduced heteropoly compounds are stable in solution. These results also show that a strong interaction exists between the Co^^ ion and the heteromolybdate entities. 3.3 Alumina supported heteropolyanions The alumina supported heteropolycompounds were characterized by DRS and ^^P MASNMR. The spectra are respectively reported in Figure 9 and Figure 10. The UV spectrum of alumina supported C03/2PM0 (Fig 9a) shows a main band at 550 nm representative of a cobalt aquocomplex [16] whereas the dried alumina supported H7PM0 (Fig 9b) exhibits the band characteristic of the p4 form at 685 nm [17], This shows that upon impregnation the nature of the heteropolycompound is not modified. A band is also observed at 344 nm which is not identified with the available data. Upon calcination in N2 of the C07/2PM0 a broad band is observed between 500 and 700 nm whereas the bands usually observed for cobalt supported on alumina are observed after calcination in air [18]. The ^^P NMR measurement of bulk H7PM0 (Fig 10a) shows a line at - 12.41 ppm, which corresponds to the p4-form.The presence of cobalt in the bulk Co7/2PMo° (Fig 10b) induces a shift (5 = -10.25 ppm) and a broadening of this line but the variation is lower than the one observed for the solution, so this suggests a lower interaction of the cobalt with the heteropolyanion in the bulk compound. After impregnation on the alumina and drying in air, the line of Co7/2PMo° (Fig 10c) is observed at - 8.25 ppm. The calcination in air (Fig lOd) does not influence the results (5 = - 8.21ppm). The presence of a single line in the ^^P NMR spectra as well as its position are indicative of the preservation of the heteropolyanion structure after the impregnation on alumina and calcination, but the greater shift suggests a stronger interaction HPA-Co in the catalyst than in bulk compound.

191 ^nf (a)

M^

700

300 400

(d)

1-^ X(nm) 800 900

Figure 9. UV spectra of supported catalysts: (a) C03/2PM0 /AI2O3 dried in air, (b) H7PM0/AI2O3 dried in N2, (c) Co7/2PMo°/Al203 dried and calcined in air, (d) Co7/2PMo°/Al203 dried and calcined in N2 -12.14

100

5 (ppin)

Figure 10. ^^P MAS-NMR Spectra (7kHz)of: (a) H7PM0, (b) Co7/2PMo°, (c) Co7/2PMo°/Al203 dried in air, (d) Co7/2PMo°/Al203 dried and calcined in air

192 3.4. Activity It should be mentionned that the HDS conversions are identical if the impregnation is performed directly on the extrudates or on grinded alumina (size of the particle .5 mm). This means that the impregnating solution penetrates inside the pore of the carrier and that we have no mass transfer limitation during the impregnation of the extrudates. The conversions in HDS of thiophene are reported in table 1. It shows that the catalysts prepared with unreduced heteropolyanions are effective for hydrodesulphurization (N"" 3 to 9) and that the values obtained with phosphomolybdate compounds are similar to those obtained with silicomolybdate ones. However the promoting factor is lower than the one currently observed for solids prepared by a classical impregnation with ammonium heptamolybdate (see N° 3; 4 and 7, 8, 9). This is due to the low Co content of these solids. Table 1 Activity in HDS of thiophene (% conversion of thiophene) N° Compound Thermic Treatment 1 CoMoPl drying and calcination: N2 2 C0M0P2 drying and calcination: N2 3 drying : air H3PM0 4 drying: air C03/2PM0 5 drying and calcination: air 6 drying: air Ni3/2PMo 7 drying: air H^SiMo 8 Co2SiMo drying and calcination: air 9 Co3/2SiMo drying and calcination: air 10 C05/2PM0 drying: N2 11 drying : N2 C07/2PM0° 12 drying and calcination: N2 13 drying and calcination: air 14 drying and calcination: N2 C07/2PM0* 15 Co7/2SiMo drying and calcination: N2 16 Co4SiMo drying and calcination: N2 17 H3C0PM0 drying: N2 18 drying and calcination: N2 19 H3C02PM0 drying: N2 20 drying and calcination: N2 21 drying and calcination: N2 H4C62SiMo

Conversion (%) 27 22 4 11 15 21 4.9 19 14 20 19 32 30 29 29.5 30 12 12.5 18 18.5 21.5

Higher conversions are observed for solids with higher Co loadings (see N° 10 to 16) which can be obtained through a reduction process of the heteropolyanion. Moreover it should be pointed out that calcination in air or N2is favourable to the activity (N° 11, 12, 13). The results presented lines 10 to 21, show that the substitution of the H^ ion is needed to increase the Co loading and consequently the thiophene HDS conversion. However the calcination allows an increase of conversion only for catalysts in which the K" ions are

193 exchanged (lines 17-21 compared to lines 10-16). Further experiments are now in progress to explain these differences. Up to now, from the results obtained in this work, it appears that the most active catalyst is the C07/2PM0, the conversion of which is higher than the one observed for the reference CoMoPl catalyst. It is even more active than the CoMoP2 one which has the same stoichiometry as the C07/2PM0 catalyst but its Co/Co+Mo ratio is lower than .28, the optimum value currently admitted. This could be correlated to the aforementionned strong interaction between the promotor and the heteropolyanion identified in the impregnating solution as well as on the oxidic precursor. This interaction could decrease the fraction of Co atoms involved in the formation of the well known surface "C0AI2O4" species [19], which is not available for the decoration of the M0S2 crystallites.

4. CONCLUSION The main findings of this work can be summarized as follows: i) The reduction allows us to increase the Co/Mo atomic ratio of the bulk phosphomolybdate and the silicomolybdate saUs of the Keggin structure, ii) The nature of these heteropolyanions is not modified after solubilization in water and a strong interaction between Co^^ ion and this heteropolyanion in solution has been evidenced, iii) Phosphomolybdate anion is preserved after impregnation and drying or calcination, iv) The increase of the Co/Mo ratio improves the thiophene HDS conversion, the value of which is higher than the one observed with a catalyst prepared by conventional dry impregnation with AHM, although with a lower Co/Mo atomic ratio. In conclusion, this work has shown that Keggin heteropolyanions containing promotor and base metals are convenient precursors for the preparation of HDS catalysts.

REFERENCES 1. J.A.R. van Veen, P.A.J.M. Hendriks, R.R. Andrea, E.J.G.M. Romers, and A.E. Wilson, J. Phys. Chem., 94 (1990), 5282. 2. W.C. Cheng, N.P. Luthra, J. Catal., 109 (1988), 163 3. J.A.R. van Veen, E. Gerkema, A.M. van der Kraan, A. Knoester, J. Chem. Soc. Chem. Commun., 1684(1987). 4. R.J.J. Jansen, H.M. Van Veldhuizen, M.A. Schwegler and H. Van Bekkum, Rec. Trav. Chim., Pays-Bas, 113, 115 (1994). 5. A.M. Maitra, N.W. Cant and D.L. Trimm, Appl. Catal, 48 (1989), 187. 6. Y. Okamoto, T. Gomi, Y. mori, T. Imanaka, S. Teraniski, React. Kinet. Catal. Lett., 22, 3-4 (1983), 417. 7. C. Sanchez, J. Livage, J.P. Launay, M. Fournier, Y. Jeannin, J. am. Chem. Soc, 104 (1982), 3194.

194 8. C. RocchiccioIi-DetchefF, M. Foumier, R. Franck, R. Thouvenot, Inorg. Chem., 22 (1983), 207. 9. C. RocchiccioIi-DetchefF, M. Amirouche, M. Fournier, J. catal., (48) 138 (1992), 445. 10. M. Foumier, C. Rocchiccioli-Detcheff, L.P. Kazansky, Chem. Phys. Lett., 123 (1994), 294. 11. C. RocchiccioIi-DetchefF and R. Thouvenot, J. Chem. Res., Synop., 46 (1977), miniprint 549 (1976) 12. L.P. Kazansky, Contribution From the institute oF Physical Chemistry, Moscow 117071, USSR, (1979), 70. 13. R.I. Maksimovskaya, V.M. Bondareva, Russian Journal oF inorganic Chemistry, 39, 8 (1994), 1238. 14. D.Z. Herranz, Contribution a I'etude des heteropolyanions molybdo-tungsto phosphoriques etude par RMN de ^*P, These, Paris, 1981. 15. R. Massart, R. Contant, J.M. Fruchart, J.P. Ciabrini and M. Fournier, Inorg. Chem., 16 (1977), 2916. 16. L.G. Roberts, F.H. Field, J. Am. Soc, 72 (1950), 4232. 17. R. Massart, Ann. Chim., t.4. (1969), 365. 18. M. Lo Jacono, A. Cimino, G.C.A. Schuit, Gaz. Chim. Ita., 103 (1973), 1281. 19. H. Topsoe, B.S. Clausen, Appl. Catal. 25 (1986), 273.

© 79P7 Elsevier Science B. V, All rights reserved. Hydrotreatment and hydrocracking of oil fractions G.F. Froment, B. Delmon and P. Grange, editors

195

Genesis, Characterizations and HDS Activity of Mo-P-Alumina Based Hydrotreating Catalysts Prepared by a Sol-Gel Method R. Iwamoto ^^ and J. Grimblot« «Laboratoire de Catalyse Heterogene et Homogene, URA CNRS D402, Universite des Sciences et Technologies de Lille, 59655 Villeneuve D*Ascq Cedex, France ^ Central Research Laboratories, Idemitsu Kosan Co., Ltd., 1280 Kami-izumi Sodegaura, Chiba, Japan ABSTRACT Mo oxide - P oxide - Aluminum catalysts with a wide range of P loading (014 wt%) were prepared by a sol-gel method to elucidate the role of phosphorous on the textural, structural and catalytic properties of Mo based catalysts. Two different Mo loadings (expected ~20 and ~30wt%Mo) and two kinds of P precursors (phosphoric acid, phosphorus pentoxide) were examined. The structural properties of dried and calcined forms were studied by means of various characterization techniques. Specific surface area (S.S.A.) of catalysts were decreased proportional to the P loading in every series. Especially, the S.S.A. in the series of P2O5 precursor decreased drastically above 7.7wt%P loading. XRD measurements revealed that excess loading of Mo and P within the alumina framework provokes aggregation of bulk M0O3 (above 6.8wt%P in the series of 30wt%Mo for H3PO4 precursor and above 5.5wt%P in the series of P2O5 precursor). From IR measurements, it was found that P and Mo atoms interact with equivalent sites of alumina. From NMR measurements, predominant formation of Mo-P heteropoly complex were observed in the drying step. P interacted strongly not only with alumina framework but also with P itself. P2O5 prefers to polymerize by calcination. It was also found that Mo enhanced the interaction of P with alumina through the formation of P-Mo heteropoly complex. Water extraction tests revealed that Mo and P interacts strongly with the alumina framework. The HDS activity was not promoted by P while excess P decreased HDS activity with the formation of bulk M0O3. 1. INTRODUCTION The active phase of hydrotreating catalysts generally consists of Mo sulfide deposited on y-alumina which was produced by calcination of alumina hydroxides precursors. The Mo precursor is usually introduced to alumina by conventional dry or wet impregnation methods. However, only up to 10-12 wt% Mo can be dispersed by these methods. In previous works, new preparation methods of welldispersed Mo precursor based on a sol-gel method were proposed [1][2]. In this sol-

196 gel method, alumina is obtained by hydrolysis of aluminium tert-butylate or aluminium sec-butylate. Mo is incorporated homogeneously with the alumina precursor during the support preparation. This advanced method can give at least 30 wt% of weU dispersed Mo and higher HDS activity than conventional catalysts. Indeed, the physico-chemical properties of resulting soUds depend on the reactions involved in the sol-gel process (hydrolysis, condensation through alcooxolation, oxolation or olation steps), on the nature of the metal or the associated alkoxide and finally on the reaction condition (temperature, the ratio between the solvent and alkoxide). Furthermore, the sol-gel preparation method is very convenient not only for obtaining active catalysts but also for investigating what happens on the surface of catalysts because of their unique high S.S.A.. To achieve higher activity with the sol-gel catalyst, it is useful to investigate the effect of promoter and additives such as Co, Ni and P on the Mobased sol-gel catalysts. The role of P on the HDS activity for Mo based hydrotreating catalysts has been studied by many researchers, while the precise effect has not been well understood yet [3-5]. Eijsbouts et al. reported that P had no effect on the HDS activity for M0O3/AI2O3 [3]. On the other hand, Lewis et al. and Kim et al. reported a positive effect for HDS reaction in the region of low P loading [4] [5]. In this work, we wished to elucidate the role of P on the Mo-P-alumina sol-gel catalysts which contain high loading of Mo and a wide range of P amounts. Their main structural and textural properties will be compared as well as their performance in thiophene HDS. 2. EXPERIMENTAL 2.1 Catalyst preparation Mo-P-Alumina catalysts were prepared on the basis of a sol-gel method according to the procedures in Figure 1. Alumina was prepared by the hydrolysis of aluminium sec-butylate (ASB) dissolved in 2-butanol (2BN) and 1,3-butanediol (13BD). Mo and P were incorporated with the alumina precursor during the gel preparation. Mo was added to the aluminium alkoxide before hydrolysis as a dispersion of ammonium heptamolybdate (AHM) in 13BD. P was introduced by different ways depending on the nature of precursor. P2O5 precursor was introduced in ASB solution after dissolving in 2BN (Route A). On the other hand, 99% of ortho-phosphoric acid (H3PO4 precursor) was dissolved in 13BD simultaneously with AHM (Route B). The catalysts obtained at each stage are noted to MPDOC-Y)H, MPC(X-Y)P where the MPD, MPC means dried and calcined sample, X,Y means expected loadings in wt% of Mo and P respectively. H, P refer to the nature of P precursor such as H3PO4 or P2O5 respectively. The ratio of H2O/ASB was usually kept at 10. It was noted however as * if H2O/ASB was increased to 100. 2.2 C a t a l y s t s c h a r a c t e r i z a t i o n The chemical compositions were provided by "Service Central d'analyse du CNRS" (Vernaison, France). The obtained powders were characterized by BET

197 Preparation of Catalyst

Preparation of support ASB in 2-butanol at 85*C with stirring for 10 min. molar ratio butanol/ASB: 3

/

I

P205

\

Addition of 1,3-butanedioi

Addition of (NH4)6Mo7024 and 1,3-butanediol molar ratio butanediol/ASB : 2

molar ratio butanediol/ASB: 2

^4 Addition of H2O molar ratio H2O/ASB: 10

Stirring at 8S'C for 1h

Holding at room temperature for 1h without stirring

Drying under reduced pressure at 40«C for 1h and 60*C for 1h T I

MPD(X-Y)H or P*

Drying at 100*C overnight

|

Calcining at 500*C for 3h heating rate 40*C/min.

/

MPC(O-O) Support AI203

\

MPC(X-Y)H or P*

Catalyst Mo03-P205-Al203

|

Figure 1. Procedure for preparation of Mo-P-Al sol-gel catalysts specific surface area (QUANTASORB Jr., Quantachrome; pretreated at 200°C for 30 min.), X-ray powder diffraction(XRD) (Siemens D5000 Diffractometer equipped with a goniometer, a monochromator and a Cu X-ray tube). Infrared Spectroscopy (FTIR, NicoUetSlO Spectrometer, sample was pelleted with KBr), 27A1-NMR (BRUKER ASX400; resonance frequency 104.26MHz, recycling time 3 sec, pulse length 1 îsec, spinning frequency 15kHz and reference A1(H20)6^"^) and îP-NMR (BRUKER ASXIOO; resonance frequency 40.53MHz, recycling time 40 sec, pulse length 2 jisec, spinning frequency 7kHz and reference H3PO4). 2.3 Catalytic activity (HDS) Hydrodesulfurization of thiophene was carried out at atmospheric pressure in a flow type reactor packed with 0.2g of catalyst. The catalyst was sulfided at 400*^0 for 2h with a H2/H2S (90/10) mixture gas at flow rate of 50 ml/min. After cooling down to 300°C, thiophene purified by vacuum distillation was introduced in the reactor at constant pressure (50 torr) with a flow of dried hydrogen (lOml/min.). The reaction products were analyzed by gas chromatography.

198 3. RESULTS AND DISCUSSION 3.1 Specific surface area and chemical composition Table 1 shows chemical composition and specific surface area per gram of calcined catalyst (S.S.A.) of all prepared samples. Amounts of Mo and P were almost those expected except for MPC(30-14)P*. It is suggested that excess water enhanced dissolution of AHM and prevents incorporation with alumina. As already reported in a previous report [1][2], the bare alumina MPC(O-O) and Mo oxide alumina such as MPC(20-0), MPC(30-0) have higher S.S.A. compared with conventional ones [6]. With introducing P, the S.S.A decreased proportional to the amount of P loading in every series. Especially, the S.S.A. decreased drastically above 7.7wt%P loading in the series of P2O5 precursor. It is presumed that P and cracked alcohol residues may block the porosity of sample because large amount of carbon (5.9wt%) was found on the MPC(30-13)P. However the corresponding MPC(30-14)P* which was prepared with increasing H2O/ABS ratio showed less carbon residue but stiU low S.S.A. (13m2/g). Table 1. Chemical composition and S.S.A. of prepared catalysts Catalysts Mo (wt%) P (wt%) Carbon(wt%) MPC(O-O) 0 0.5 0 MPC(20-0)H 0.2 17.5 0 MPC(20-1)H 0.3 17.9 1.6 MPC(20-2)H 17.3 2.2 0.2 MPC(20-3)H 17.2 0.2 3.1 MPC(20-4)H 17.9 4.4 0.2 MPC(20-7)H 6.6 0.3 16.8 MPC(20-11)H 16.4 11.3 0.3 MPC(30-0)H 26.0 0 MPC(30-1)H 0.1 25.9 1.1 MPC(30-2)H 2.2 0.1 26.5 MPC(30-5)H 0.1 4.6 25.3 MPC(30-7)H 25.8 6.8 MPC(30-11)H 0.1 25.3 11.1 MPC(30-1)P 28.3 1.7 0.3 MPC(30-3)P 0.6 27.7 2.7 0.2 MPC(30-6)P 5.5 25.5 0.2 MPC(30-8)P 26.7 7.7 5.9 MPC(30-13)P 12.7 25.3 MPC(30-14)P* 16.5 13.7 0.3 MPC(0-11)H 0 0.3 10.6 MPC(0-1Q)P 0 -_ (10)

S.S.A.(m2/g) 503 586 570 560 559 525 505 428 523 505 443 411 405 263 508 490 439 238 4 13 474 486

3.2 X-ray powder diffraction (XRD) (a)H3P04 precursor Figure 2 shows the XRD patterns of sol-gel catalysts obtained from the H3PO4 precursor. The bare alumina MPC(O-O) can be identified as poorly

199 crystalline y-Al203. For P oxide-alumina MPC(0-11)H and Mo oxide-alumina MPC(30-0), no peak corresponding to P or Mo oxo-compounds can be detected. It is suggested therefore that Mo and P exist as a weU dispersed species. It was also observed that incorporation of Mo and/or P with alumina prevents the formation of structured y-Al203. They seem to be present in an amorphous matrix. For Mo-PAl, the weU dispersed state of Mo oxide is kept up to MPC(20-11)H in the series of 20wt%Mo (not shown here) and up to at least MPC(30-5)H in the series of 30wt%Mo. However, bulk M0O3 can be identified above MPC(30-7)H. This result means that high loading of Mo and P within the alumina framework provokes aggregation of bulk Mo oxide.

40 2theta/«

Figure 2. XRD patterns of Mo-P-Al sol-gel catalysts prepared from the H3PO4 precursor. (a)MPC(O-O), (b)MPC(0-ll)H, (c)MPC(30-0), (d)MPC(30-5)H, (e)MPC(30-7)H, (f)MPC(30-ll)H

Figure 3. XRD patterns of Mo-P-Al sol-gel catalysts prepared from the P2O5 precursor (a)MPC(30-0), (b)MPC(30.3)P, (c)MPC(30-6)P, (d)MPC(30-13)P, (e)MPC(30-14)P*

200 (b)P205 precursor Figure 3 shows the X-ray powder dififraction patterns of sol-gel catalysts obtained from the P2O5 precursor. The addition of P2O5 precursor showed similar effect as the H3PO4 precursor. However, the intensity of bulk M0O3 in MPC(30-6)P which contains 5.5wt%P was almost the same as that observed in MPC(30-7)H which contains 6.8wt%P. It was concluded therefore that the P2O5 precursor enhanced the formation of bulk M0O3 compared with H3PO4 precursor at the same loading of Mo. 3.3 Infrared spectroscopy (a)H3P04 precursor The assignment of IR bands in Mo-P-Alumina based catalysts have been already reported by many researchers [7-13]. Figure 4 shows IR spectra of dried and calcined catalysts obtained from the H3PO4 precursor. For all dried catalysts, a broad band at ~750 cm^ which is assigned to Al-0 stretching was observed. Furthermore, many small bands and shoulders were observed (i.e. at 1458, 1370, 1135 and 1055 cm-i etc.), though it is sometimes difficult to identify. These bands could be assigned to residual alcoholate incorporated in the alumina framework or supported metal complexes, because they are well corresponding to IR spectra of the solvents (2BN,13BD). This fact indicates that the hydrolysis reaction of ASB does not proceed completely in this preparation condition. Bands at 1070 cm-i for MPD(O-O) and MPD(0-11)H is considered as sol-gel boehmite [7]. P containing catalysts such as MPD(30-11)H, MPD(0-11)H have a broad band at ^1100 cm-^ This band can be decomposed into three bands at 1115, 1080 and 1055 cm-i which is assigned to stretching vibration of P=Ot, P-0 and P-O-Mo of heteropoly acid [8] [9]. This result assumes that P-Mo heteropoly compound was formed during the gel precipitation. Specific bands at ~1404, 900 and 845 cm-i which can be assigned to AHM were observed in MPD(30-11)H. It is considered that P prevents the incorporation of Mo within the alumina framework even at the drying step since the intensity of these bands are weU correlated to the P content. For the calcined samples, all the spectra are rather broad. With increasing loading of P, a large broad band appeared again at about 1100 cm^ which can be decomposed into two bands at 1125 and 1090 cm-i. They are assigned to the P=Ot and P-0 respectively [9][10]. In MPC(30-11)H, bands at 1000, 880 and 823 cm-i which are assigned to bulk M0O3 were also detected. This result is in well agreement with the results of XRD. It is suggested that the main part of M0O3 derived from decomposition of bulk AHM with calcination. In MPC(O-O), characteristic three bands at 1640, 1503 and 1425 cm-i which might be assigned to physically or coordinately adsorbed H2O were observed. However, the two of three bands were disappeared with the introduction of Mo and/or P [11]. This result indicates that Mo and P are interacting with equivalent sites of AI2O3. This is a reason why a part of Mo cannot interact with AI2O3 when P content increases.

201

MPD(O-O)

MPC(O-O)

1500

1000

WAVENUMBER(cm-l)

1500

1000

500

WAVENUMBER(cm-l)

Figure 4. IR spectra of Mo-P-Al sol-gel catalysts prepared from the H3PO4 precursor. (A: after drjdng 100°C , B: after calcination at 500°C). (b)P206 precursor Figure 5 shows IR spectra of dried and calcined catalysts obtained from the P2O5 precursor. In MPD(0-10)P, a characteristic band which might be assigned to monomeric P species observed at 1000 cm-i. In MPC(30-13)P, intensity of bands for the residual alcoholate was much more higher than those in MPD(30-11)H. It is suggested that P2O5 prevents the hydrolysis of ASB and eventually, it remains more alcoholate in the final compound. It is also assumed that part of these bands are attributed to organic P complex formed with the alcohol solvent, since the same IR spectra was obtained from dried P2O5 after dissolving it in 2BN. From the literature, the formula of these complexes are P0(0H)2(0but), P0(0H)(0but)2 or P0(0but)3 [12]. These complexes are considered to be formed by the following reactions. P2O5 + 2but-0H+ H2O -> 2PO(OH)2(Obut)

(1)

P2O5 + 4but-0H ^ 2PO(OH)(Obut)2 + H2O

(2)

P2O5 + 6but-0H ^ 2PO(Obut)3 + 3H2O

(3)

202

For calcined samples, MPC(0-10)P showed broad bands between 1000 and 1330 cm-i which are assigned to highly polymerized P oxo-compounds [13][14]. For MPC(30-13)P, the intensity of bands at 1090 and 1125 cm-i which are assigned to P-0 and P=Ot vibration decreased comparing with those in MPC(30-11)H. On the contrary, the intensity of band at 1200 cm-i which is assigned to polymeric P oxocompounds increased. It is suggested that MPC(30-13)P contains also more polymerized P oxo-species than MPC(30-11)H. If the ratio of H2O/ASB increases from 10 to 100, the bands at 1330 cm-i increased significantly. This means that the excess of water during the gel preparation provokes the aggregation of P. It is assumed that P has less interaction with alumina in the drying stage because a large part of P is involved in complexes with the alcohol solvent. In such a case, P prefers to polymerize than to interact with the alumina framework. The bands for bulk M0O3 were also observed at 1000, 880, and 823 cm-i in MPC(30-13)P and MPC(30-14)P*. From the IR measurements, it was found that the P precursor affects significantly on the physicochemical properties of resulting catalysts.

MPC(30-0)

MPD(30-0)

lij

o z < CO

2000

1500 1000 WAVENUMBER(cm-l)

500

2000

1500

1000

500

WAVENUMBER (cm-1)

Figure 5. IR spectra of Mo-P-Al sol-gel catalysts prepared from the P2O5 precursor. A: after drying 100*^0 , B: after calcination at 500°C.

203 3.4 27A1.NMR (a)H3P04 precursor Top peak value of ^Âl-NMR spectra are listed in Table 2. The assignment of 2'Âl-NMR spectra in this region have already been reported by many researchers [15-22]. MPD(O-O) has a single broad signal at 7.2 ppm. This signal is assigned to octahedral alumina [15]. For all the Mo and P containing catalysts, tailing of spectra between 0 and -30 ppm or even presence of a shoulder at -5 ppm were observed depending on the content of Mo or P. This tailing should correspond to octahedral surface aluminium sites shell in which P or Mo are located in a second coordination [16]. These signals seem to be characteristic for sol-gel catalysts since surface informations are emphasized by the extremely large S.S.A.. In addition, P containing catalysts showed another weak signal at ~41 ppm which is assigned to AIPO4. This result indicates that P interacts strongly with alumina framework even in the drying step. It was also revealed that the degree of interaction between P and alumina increased in the presence of Mo because the intensity of AIPO4 in MPD(30-11)H was more stronger than that in MPD(0-11)H. Table 2. Results of ^Âl -NMR obtained from the Mo-P-Al sol-gel catalysts Catalysts MPD(0-0) MPD(0-11)H MPD(20-0) MPD(30-0) MPD(20-11)H MPD(30-11)H MPD(30-13)P MPD(30-14)P* MPD(0-10)P

Before calcination 7.2 40.2 7.1 6.5 6.2 41.2 6.0 41.2 5.9

61.0

. 1.1 13.7 13.6 -0.4 5.2 33.0

(ppm)

-5.0 -5.0 -5.0

After calcination (ppm) 65.5 33.0 6.9 -12.0 65.5 39.0 6.6 5.2 62.2 30.0 -13.6 54.7 27.3 5.3 -13.6 55.0 37.8 4.8 -13.6 36.7 6.6 36.7 37.9 55.0

27.0 13 -2.9 -14.1 26.8 -2.9 -12.4 26.8 6.3

On the calcined bare alumina MPC(O-O), another new signal of tetrahedral aluminium site was observed at ~65 ppm [15]. Furthermore, a broad shoulder appeared at 33 ppm which might be attributed to 5-fold coordinated aluminium sites [17]. This signal is characteristic for the sol-gel alumina since it possesses a highly disordered and poorly crystalline structure as shown by XRD. The signal at ^30 ppm is also observed in the Mo loaded catalysts such as MPC(20-0) or MPC(30-0). The intensity of this signal increased with the increasing amount of Mo. This could be assigned to 5-fold coordinated aluminium sites since introduction of Mo prevents the crystallization of alumina and leads to much more distortion as already shown by XRD. However, another explanation as being due to the presence of a surface tetrahedral Al(OMo)4 cannot be neglected.

204 Furthermore, MPC(30-0) gave a weak shoulder spectra at -13.6 ppm which is assigned to Al2(Mo04)3 [18]. This compound is supposed to be derived from a following equation. 3 MoOaCbulk) + 2 Al203(surface) ^ Al2(Mo04)3

(4)

The formation of Al2(Mo04)3 is more apparent in MPC(30-11)H, because the high P loading favors the formation of bulk M0O3. However, MPC(0-11)H which contains only P also showed the shoulder signal at ~-13 ppm. This signal can be assigned to A1(0P)6 in this case [19]. Hence, the signal of ~-13, 14 ppm might be considered as multiple states of octahedral surface alumina in which terminal OH are exchanged by Mo or P. The spectra for AIPO4 was observed at 37 to 39 ppm for all the P containing catalysts. Though the AIPO4 already existed in the drying step, the main part of AIPO4 forms during calcination. The intensity of AIPO4 for MPC(20-11)H and MPC(30-11)H were much more stronger than that for MPC(0-11)H catalyst. This result indicates again that Mo provokes the formation of AIPO4. Concerning the chemical shift, the top peak values of 6, 5 and 4-fold aluminium sites tend to decrease with the increasing amount of Mo. This might be caused by the increase in distortion of the alumina framework or by the decreases of the aluminium density in the shell of each aluminium sites when the Mo loading increases [20]. (b)P205 precursor MPD(0-10)P had three signals at 61 ppm(weak), 33 ppm(weak) and 5.2 ppm(strong) which correspond to tetrahedral, 5-fold and octahedral aluminium sites respectively. MPD(30-13)P showed a large broad signal at -1.1 ppm and a small sharp signal at 13.7 ppm. The former signal could be assigned to less condensed octahedral aluminium sites [15]. The later signal might be assigned to (Al(OH)„(H20)6.n)(Mo04) or A1(0P)5 [19][21]. This result suggests that the P2O5 precursor prevents drastically the hydrolysis and condensation of the Al-alkoxide. Zaharescu et al. also reported that the rate of hydrolysis and condensation of Sialkoxide is strongly influenced by PO(OR)x complexes in the P-TEOS system [22]. P complexes might affect on the accessibility of the metal alkoxide to water molecules or to other alkoxides for condensation. In calcined catalysts, MPC(0-10)P showed also three signals at 55 ppm, 26.8 ppm(strong) and 6.3 ppm(strong) which correspond to tetrahedral, 5-fold and octahedral aluminium sites respectively similar to MPD(0-10)P. However, their respective populations were strongly modified. It is remarkable that no signal for the AIPO4 was observed in MPC(0-10)P even after calcination. MPC(30-13)P gave a large broad signal at 36.7 ppm and a smaU sharp signal at -14.1 ppm which are assigned to AIPO4 and Al2(Mo04)3 or A1(0P)6 respectively.

205 3.5 31P-NMR (a)H3P04 precursor Table 3 shows the top peak value of ^^P-NMR spectra. The assignment of ^T-NMR spectra in this region has been also reported by several researchers [2327]. For the dried MPD(0-11)H, a broad signal which could be decomposed into 2 signals at -11 and -21 ppm was obtained. They are assigned to monomeric and pol5mieric P oxo-species respectively [23]. On the other hand, the Mo and P containing catalysts such as MPD(20-11)H, MPD(30-5)H and MPD(30-11)H showed another signal at about -15 ppm. This signal could be assigned to a P-Mo heteropoly compound in agreement with the IR observation. All the calcined catalysts showed broad overlapping signals at about -18 and ~-24 ppm which corresponds to polymeric P oxo-species and AIPO4 respectively. Mo containing catalysts such as MPC(20-11)H and MPC(30-11)H gave less polymeric P than MPC(0-11)H. It is suggested that Mo is effective for dispersing P on the alumina through the formation of a P-Mo heteropoly complex. It was found that the top peak value of AIPO4 signal for MPC(30-5)H and MPC(30-11)H shifts 2 ppm to the lower value. Table 3. Results of 3iP-NMR obtained from the Mo-P-Al sol-gel catalysts. Catalysts MPD(0-11)H MPD(20-11)H MPD(30-5)H MPD(30-11)H MPD(30-13)P MPD(30-14)P* MPD(0-10)P

Befor calcination (ppm) -11 -21 -11 -15.9 -21 -11 -14.3 -11 -15.9 -21 -4 to-11 -14.6 -21 .4 to-11 -14.6 -21 -4 to-11 ^21__

After calcination (ppm) -18 -24 -24 -18 -26 -18 -26 -18 -19 -18 -19

-25 -23

b)P206 precursor MPD(0-10)P showed several overlapping signals between -4 and -11 ppm which might be assigned to multiple states of monomeric P such as P0(0H)(0But)2 and P0(0But)(0H)2 including their isomer structures. Zaharescu et al. reported that these complexes are not hydrolyzed by water [22]. MPD(3013)P showed another signal at -14.6 ppm which is assigned to a Mo-P heteropoly compounds. This result means that a part of the organic P can form complexes with Mo as weU as the H3PO4 precursor. Considering from the equilibrium studies by Jian et al. and Cheng et al. [24][25], it is supposed, that the P-Mo heteropoly compound was formed, for example, by the following reactions. 2H^ + 14PO(OH)2(OBut) + 5M07O24 6- ^

7H2P2M05O23 ^- + 14But-0H+ H2O

(5)

206 or M07024 6- + 4H20 ->7Mo04 2- + 8H+ and 6H+ + 2PO(OH)2(OBut) + 5M0O4 2-

(6) ->

H2P2M05O23 ^' + 2But-0H + 3H2O

(7)

After calcination, it was found that MPC(0-10)P showed only polymeric P at '--IQ ppm. No signal for the AIPO4 was obtained at -24 ppm in agreement with 2'7A1-NMR. It is considered that the P2O5 precursor tends to polymerize rather than to interact with the alumina framework because the alcoholate P complexes prevent the interaction with alumina. On the other hand, MPC(30-13)P showed two overlapping signals at -19 and -25 ppm which are attributed to the polymeric P oxo-species and the AIPO4 respectively. These data indicate that the P-Mo heteropoly compounds induce the formation of AIPO4. In the "P2Mo5023"structure, the two P atoms are located at top and bottom of the cluster respectively [25]. Therefore, it is thought to be easier for P to be in contact with the alumina framework. From the above investigation, it was found that the hydrolysis and condensation reaction of the Al-alkoxide are extremely prevented by the P2O5 precursor. Therefore, another MPC(30-14)P* was prepared with increasing the ratio of H2O/ASB=100 to accelerate the hydrolysis and condensation reactions. Figure 5 showed, however, that MPC(30-14)P* gave much more polymeric P oxospecies than MPC(30-13)P. It is considered that excess water shifted the equilibrium equation (5) and (7) to the left hand and consequently prevented the hydrolysis of alcoholate P precursor. Since increasing the H2O/ASB ratio did not improve the hydrolysis reaction, P2O5 might prevent the access of alkoxide molecules to each other. The amounts of Mo and P remaining after water extraction were also investigated in Table 4. In general, M0O3, H3PO4 and heteropoly compounds are easily extracted by water while monolayer molybdate and AIPO4 are hardly extracted [26][27]. The extent of extraction depends strongly on the degree of Table 4. Effect of water extraction on the atomic ratio Mo/Al and P/Al Before water extraction Catalysts MPC(20-0) MPC(20-11)H MPC(30-0) MPC(30-11)H

Mo/AI 0.13 0.18 0.23 0.36

After water extraction

P/Al

Mo/Al

0.37

0.14 0.14 0.25 0.21

0.50

P/Al

0.40 0.55

207

interaction between those compounds and support. It was found that the amount of Mo and P for MPC(20-0), MPC(20-11)H, MPC(30-0) catalysts after the water extraction gave ahnost same value as that of the initial catalysts. Therefore, all the Mo and P oxo-species in these catalysts have strong interaction with the alumina surface. On the other hand, the amount of Mo for MPC(30-11)H apparently decreases with the water extraction. Therefore, it is considered that a part of Mo cannot interact with alumina and leads to the formation of bulk M0O3. As the main conclusions from the characterizations, it appears that reactivity of P with Mo and alumina depends strongly on the nature of the P precursor and on the preparation conditions. The scheme of interaction between P and other component is shown in Figure 6. P0(0H)x(0R)3.x +ROH polymeric ^ P oxo-species +Mo

P-Mo heteropoly species

>

AIPO4

Figure 6. Schematic diagram for P transformation 3.6. Thiophene HDS activity (a)H3P04 precursor Figure 7 shows the thiophene HDS activity and selectivity for the sol-gel Mo-P-Al catalysts prepared from the H3PO4 precursor as a function of P content. It was found that no effect was detected in the series of 20wt%Mo. On the other hand, a negative effect was obtained above 4wt%P in the series of 30wt%Mo. The decrease in the HDS activity should be attributed to the formation of bulk MoOs since bulk MoOs possesses less activity than dispersed Mo. Selectivity of C4 products did not changed significantly whUe the selectivity of hydrogenated compound (butane) decreased slightly with the formation of bulk M0O3

208

/-

I 6.56^- r 1 0

•

t

•

5.5 - [ •

• Mo20wt% • Mo30wt%

\#

54.5 - [

•

43.5 3-

•

•

•

•

1 '^

H

4

6

\

1

10

15

8

P content (wt%) Figure 7. Thiophene HDS activity and selectivity products on Mo-P-Al sol-gel catalysts prepared from H3PO4 precursor. In B, • , • mean the series of 20wt%Mo and n , 0 mean the series of 30wt%Mo. b)P205 precursor Figure 8 shows the thiophene HDS activity and selectivity for the Mo-P-Al sol-gel catalysts prepared from the P2O5 precursor. The almost same trend was obtained as the H3PO4 precursor, though the activity started to decrease above -'4wt%P. This limit is lower than that found in H3PO4 series because the P2O5 precursor favors the formation of bulk MoOs and carbon in the preparation procedure (see Table 1). It is considered that only dispersion of Mo affects on the thiophene HDS activity in Mo-P-Alumina sol-gel catalysts.

209 7 6

12N7 > 12N7C5 ^ 12R4 ^ 12N7C9 as shown in Figures 1 and 7. 3.6. Surface Model of Mo Carbide Figure 8 depicts tiie active sites on tiie surface Mo carbide as evaluated from the differences in activity resititing from differing amounts of carbon. Half of the tetrahedral structure of Mo metal in a-MOjC is filled witii carbon atoms, while a third of the structure for T1-MO3C2 is

262 composed of carbon. Since a-MOjC shows greater activity than does T1-MO3C2 on AI2O3 it can be deduced that the presence of carbon depleted sites contributes to the activity of the catalyst. 4. CONCLUSIONS (1) The catalytic activities for the HDN of carbazole decreased in the following sequence.

12N7C7 > 12N7 > 12N7C5 ^ 12R4 ^ 12N7C9

Carburization enhanced the activity of the M0/AI2O3 catalyst for the HDN of carbazole, compared with the reducing and nitriding. The catalyst carburized at 700**C was 2.2 times more active than was the catalyst carburized at 500T in the HDN of carbazole. (2) The TPR analysis show that neither transformation of the Mo nitrides to the Mo carbides nor free-carbon deposition occurred below SOO^C for the dumina• vacancy supported catalysts, although above # carbon atom 900*'C large amount of free carbon is deposited on the surface of the Mo Figure 8. The structure of Mo carbide carbides catalysts. catalysts. (3) From elemental analysis, the 100N7C7 was composed two phases: Mo nitrides and Mo carbides. (4) For the unsupported Mo catalysts, the XRD analysis showed that MOjN was converted to a-Mo2C at 700*C, and above 900**C the a-MOjC crystal was transformed to ri-MOgCj crystal. REFERENCES 1. M. J. Ledoux, C. Pham-Huu, J. Guille, and H. Dunlop, J. Catal., 134 (1992) 383. 2. M J. Ledoux, C. Pham-Huu, H. Dunlop, and J. Guille, "Proceedings, 10th International Congress on Catalysis, Budapest, 1992" (L. Guczi, F. Solymosi, and P. Terenyi, Eds.), p. 955. Elsevier, 1993. 3. C. Pham-Huu, M. J. Ledoux, and J. Guille, J. Catal., 143 (1993) 249. 4. J.-G. Choi, J. R. Brenner, and L. T. Thompson, J. Catal., 154 (1995) 33. 5. J. C. Schlatter, S. T. Oyama, J. E. Metcalfe, III, and J. M. Lambert, Jr., Ind. Eng. Chem. Res. 27 (1988) 1648. 6. J. S. Lee, M. H. Yeom, K. Y. Park, I. Nam, J.S. Chung, Y. G. Kim, and S. H. Moon, J. Catal., 128 (1991) 126. 7. M. Nagai, T. Miyao, and S. Omi, "Hydrotreating Technology for Pollution Control" (M. L. Occdli and R. Chianelli, Eds.), Chap. 18, Marcel Dekker, New York, 1996. 8. M. Nagai, T. Masunaga, and N. Hanaoka, Energy and Fuels, 2 (1988) 645. 9. J. S. Lee, S. T. Oyama, and M. Boudart, J. Catal., 106 (1987) 125.


263

The design of base metal catalysts for hydrotreating reactions; Temperature programmed sulphidation of NiW/Al203 catalysts and their activity in the hydrodesulphurisation of thiophene and dibenzothiophene' H.R. Reinhoudf, A.D. van Langeveld', R. Mariscar, V.H.J. de Beer^ J. A.R. van Veen\ S.T. Sie' and J.A. Moulijn". ' Delft University of Technology, 2628 BL Delft, The Netherlands ^ Eindhoven University of Technology, 5600 MB Eindhoven, The Netherlands ABSTRACT NiW based hydrotreating catalysts have a good performance in hydrodesulphurisation reactions. It appeared that the their performance in the hydrodesulphurisation of dibenzothiophene and thiophene strongly depends on their sulphiding degree, which can be controlled by both the calcining and sulphiding temperature. By combining temperature programmed sulphiding with quasi in-situ XPS and activity measurements, it was concluded that the active phase for the HDS of DBT consist of either micro-crystalline sulphided M on a WO3 substrate, or of Ni^^ dissolved in the in NiW04. In contrast, the active phase for the HDS of thiophene seems to consist of Ni-promoted WS2 structures. 1. INTRODUCTION Hydrotreating processes play a central role in refineries, amongst others to upgrade transport fiiels. A continuous growth in the demand for transport fliels and changes in the relative importance of different crude oil fractions have put more pressure on the performance of conventional hydrotreating catalysts. More active catalysts with higher selectivity for specific hydrotreating reactions are needed to meet legislation for fuel quality in an economically attractive way. For example, in western Europe nowadays an important part of automotive transport is powered by Diesel engines, thus making dieselfiielproduction a growing segment in oil refining. However, the most important drawback of Diesel engines compared to emission controlled Otto engines is the emission of particulates which are suspected to cause serious health problems. Although the debate on the effects of the various diesel fuel components on particulate formation is still going on, aromatics and sulphur content (which are supposed to play a role in the particulate formation) are currently restricted by legislation. Apartfromthe fact that reaching low sulphur levels is a problem of its own, the presence of small amounts of sulphur also hampers the de-aromatisation of diesel fiiel. Therefore, deep hydrodesulphurisation (HDS) of dieselfiielis a key process in the upgrading of diesel fuel properties. The sulphur components which remain in diesel fiiel after conventional HDS are maidy dibenzothiophenes (DBT). It has been well established [1] that especially DBT's with alkyl substituents on the 4 and 6 position have a low reactivity over conventional hydrotreating

264 catalysts like C0M0/Y-AI2O3. This is due to the steric hindrance of the alkyl groups which block direct hydrogenolysis. It has been demonstrated that hydrogenation of one of the benzene rings of DBT lifts the steric hindrance [2], resulting in a higher reactivity in accord with molecular modelling. This observation implies the search for new catalysts with a high activity for reactions like isomeriation and hydrogenation which lift steric hindrance. Recent work on the development of catalysts for deep hydrodesulphurisation [3] revealed that NiW/y-Al203 catalysts are very promising for this application. Despite the importance of hydrotreating catalysts, the production of these catalysts is still mainly based on experience and empirical knowledge. Considering the increasing relevance for dedicated hydrotreating catalysts, the need for a thorough description and understanding of preparationpretreatment-activity relationships is clear. In the past a lot of work was focused on CoMo/y-Al203 and NiMo/y-Al203 catalysts. However, with the expected change to more specific hydrotreating reactions, and possibly also different reaction conditions, other catalysts than CoMo/y-Al203 and NiMo/y-Al203 might be more suitable. In the light of the promising activity in the deep HDS of gasoil, it was concluded to investigate the genesis of the active phase in NiW/y-Al203 catalysts in a detailed and systematic way. Scheffer et al. [4] have shown that especially NiW/y-Al203 is an interesting sytem, since it allows to steer the formation of different phases and morphologies by applying different pretreatment conditions. In this paper we will demonstrate the importance of pretreatment conditions on the activity of NiW/y-Al203 catalysts for different model reactions. The differently pretreated catalysts were characterised with Temperature Programmed Sulphiding (TPS) and quasi insitu X-ray Photoelectron Spectroscopy (XPS). The final aim of the work is the assessment of catalyst design rules for the NiW catalysts based on the knowledge of the correlation between pretreatment conditions, development of the active phase and the catalytic performance in different hydrotreating reactions.

2. EXPERIMENTAL 2.1 Catalyst Preparation The NiW/Al203 catalyst was prepared by pore volume co-impregnation of y-Al203 (Ketjen 0001.5E CK300, high purity, SBEf=190 m^.g'\ pore volume 0.6 ml/g). The aqueous solution contained (NH4)6.Wi2039.xH20.(Aldrich) and Ni(N03)2.6H20 (Aldrich), resulting in a catalyst with 1.2 wt% Ni and 15.2 wt% W, corresponding to 2.5 W/nm^ and 0.6 Ni/nm^. Table 1 Pretreatments of the NiW based catalysts investigated Tsulph [K]

613 673 823

1

923

1

Tcalc [K]

393

673

* •

• • *

823 * * * *

923 * t

1 1

265 Next, the catalyst precursor was dried overnight in air at 393 K, followed by calcining at various temperatures during 1 hour. For the activity measurements, temperature programmed sulphidation and the XPS analysis, the extrudates were ground and sieved for the 125 and 250 jam particle size fraction. The various catalyst investigated are collected in Table 1. 2.2 Catalyst characterisation Temperature programmed sulphiding was performed in an atmospheric plug flow reactor, more details on the equipment can be found in [5]. About 100 mg of the catalyst was diluted with a same amount of SiC. After purging at room temperature with Ar, the catalyst was exposed to the sulphiding mixture containing H2S, H2 and Ar (3, 25 and 72 vol %, respectively) at a totalflowrate of 33 [j,mol/s. After 30 minutes the temperature program, with a linear heating rate of 0.167 K/s, was started. Upon reaching the highest temperature of sulphiding, the sample was cooled in the sulphiding mixture. The signals of all TPS profiles have been normalised to the amount of catalyst. The sulphiding degree of the samples is based on a quantitative sulphiding of the nickeloxide into K13S2, and of the tungstenoxide into WS2. Clearly, the sulphiding degreefromthe TPS also includes chemisorbed S on the active phase of the catalyst, which is not taken into account m the reference point for complete sulphiding. For subsequent XPS analysis, the samples were purged with Ar at room temperature in order to remove residual traces of H2S. Then, the reactor was closed, disconnected from the TPS equipment and transferred into a glove box where the catalyst sample was transferred into the quasi insitu transfer facility for the XPS machine. XPS analysis was performed on a Perkin-Elmer PHI 5400 ESCA system equipped with a hemispherical analyser. Sample excitation was done by AlKa X-rays (1486.6 eV). The catalyst was pressed into an In foil attached to the sample holder under protective conditions in a glove box. Partial pressures of oxygen and water are lower than 0.5 . 10"^ mbar, typically. Transfer of the samples form the glove box into the XPS machine was performed by use of a commercially available transfer chamber. Peak shifts due to charging of the samples was corrected for by taking the Al 2p line of the AI2O3 at 74.2 eV as reference [6]. 2.2 Catalyst performance testing The batch autoclave reactor set-up used in the experiments for the DBT hydrodesulphurisation was described elsewhere [5]. About 200 mg of the catalysts was sulphided quasi insitu in an integrated reactor with 15% H2S in H2 with a flow of 40 [xmol/s at 1.2 MPa. The reactor was heated at 0.167 K/s up to the highest sulpiding temperature and kept isothermal for 1 h. Upon cooling down in the sulphiding mixture, the sulphided catalyst is transferred into the batch autoclave, where is submerged in the model feed, thus avoiding exposure to air. The model feed,which consisted of 2 g dibenzothiophene (Aldrich, 98%) in 100 g n-hexadecane (Aldrich, 99%+), was used to test the activity of the catalysts at 633 K and 100.0 MPa. No initial H2S was added to the reactor, the H2S/H2 ratio at 50% conversion was 2 mol%. Thiophene HDS was performed at atmospheric pressure in a flow reactor. Sulphidation of the catalyst in a mixture of H2S and H2 (50 vol% of both gases) at a total flow rate of 16 lamol/s. The catalyst was heated at 0.167 K/s up to 543 K, kept isothermal at this temperature during 30 minutes, followed by heating at 0.167 K/s up to the desired sulphiding temperature, i.e. 673 K, 823 K or 923 K, where it was kept isothermal during 2 hours. The

266 total flow was 39.5 famol/s, the thiophene content being about 6%. Thiophene conversions were determined at 623 K after 4 hours stabilising the catalysts. From the conversion the first order rate constant for the HDS of thiophene was evaluated. 3. RESULTS 3.1. Temperature programmed sulphidation In Figure la and lb the temperature programmed sulphidation profiles are shown for the catalysts calcined at 673 K and 823 K, respectively. For both catalysts the isothermal room temperature uptake is not shown in thefigures,however, the amount has been taken into account for the quantification of the sulphiding degree. Directly after the start of the heating program, a small amount of H2S is produced without accompanying H2 uptake, followed by a H2S uptake which starting at about 375 K. For the catalyst calcined at 673 K, this H2S uptake is accelerated above 490 K, while simultaneously a minor H2 consumption can be seen. Next, a sharp increase in the H2S concentration is observed which effectively results in a production maximum at about 615 K. Simultaneously, a similar H2 uptake occurs.

H,S

H2 T — I — I — 1 — I — I — r

300

500

700

900

300

1 — I — \ — I — r

500

700

900

-^T[K] Figure 1. Temperature programmed sulphidation profiles of the catalyst calcined at 673 K (left) and 823 K (right). Signals are normalised in micromol of H2S and H2 per 100 mg of catalyst. Note that a negative deflection of the upper profile corresponds to a H2S consumption, whereas for H2 (lower profile) a positive deflection corresponds to a H2 consumption. Quantification of the integrated sulphur uptake at 615 K, yields a sulphidation degree of the active phase of 63 %. The H2S production peaking at 615 K is followed by a second uptake with a maximum consumption at about 700 K, followed by a slow progressive sulphiding of the sample up to 923 K. At the highest temperature (923 K) the sulphiding degree is 87 %. For the catalyst calcined at 823, the H2S and H2 profiles exhibit two major uptakes. However, the amounts of H2S and H2 consumed are much less than in case of the catalyst calcined at 673 K. At 615 K, the total sulphur uptake corresponds to a sulphiding degree of 32 %, while at 923 K the total sulphiding degree is 64 %. Quite remarkably, the H2S production and H2 consumption peaking at 615 K are virtually absent.

267 3.2. Quasi in-situ XPS of sulphided catalysts The XPS spectra of the Ni 2p and the W 4f emission line regions of the catalyst calcined at 823 K in the various stages of sulphidation are collected in Figure 2. Note, that the reference spectra of the oxidic precursors are not shown. The quantitative data, that is, the peak position of the oxidic and sulphidic contribution of the Ni 2p3/2 and W 4fia emission lines and sulphiding degree of both elements upon the various sulphiding steps are collected in Table 2. Upon isothermal sulphiding of the catalyst at 293 K, no significant shift of the emission lines of Ni and W could be observed. However, the Ni 2p3/2 line broadens by about 10% at half height. The peak position of the oxidic M 2p3/2 was nearly constant at 856.6-856.7 eV, for the sulphidic contribution, the peak position was found at 853.6-853.8 eV. At 540 K, 39 % of the Ni was sulphided. At 613 K, the amount of sulphided Ni was increased up to 49 - 58 %, dependant on the time of isothermal sulphiding. After sulphiding at 823 K the amount of sulphidic Ni was about 70 %, and almost independent of the time of isothermal sulphiding, leaving 29 % of the Ni in the oxidic state at the highest temperature of sulphiding. The peak position of the oxidic W 4fâ line was found at 34.4 - 34.8 eV, while that of the sulphidic contribution was found at 31.6 - 32 eV. Sulphiding of the W (9 %) could only be observed at 540 K. After sulphidation at 613 K, the relative amount of sulphidic W increased up to 21 % after 2 hours isothermal sulphiding. Pronounced sulphiding of the W up to 75 % only occurred at 823 K. Note, that about 25 % of the W is not sulphided at the most severe sulphiding conditions. Table 2 Quantification of the XPS spectra shown in Figure 2. 1

A sulph

(K)

Ni 2p3/2 (eV)

tsulph

(min)

W4f7/2

(eV)

1

oxid

sulph

% Wsulph 1

oxid

sulph

0

856.7

-

0

-

-

30

856.2

853.2

3

35.7

-

1

856.5

853.8

39

35.4

-

1 613

5

1 856.4

853.6

49

35.5

32.2

1 613 1 613

30

856.6

853.6

49

35.3

31.8

120

1 856.6

853.6

58

35.4

31.9

60

856.6

853.9

70

35.7

120

856.9

853.9

71

35.6

-

1 298

1 ^"^^ 1 ^23 1 823

% Nisuiph

1

0 0 0 9

1 1 1 1

32.2

11 21 74

1 1 1

32.2

75

1

268 1

1

1

1

1

1

'^' ^ P

1

\

1 1—

854.0 eV

—1—I—I—I—r

W4f

3/2: s

5/:

1/2

32.0 eV 7/2

. y ^ ^ (0 c (D

a

J

1

1

880

1

1

L

I

860

'

'

J

-^ BE [eV]

40

1

1

L

30

Figure 2. Quasi in-situ XPS spectra of the Ni 2p (left) and the W 4f lines of the catalyst calcined at 823 K in its various stages of sulphidation, a) 30 min. at 298 K, b) 1 min. at 540 K, c), d) and e) 5, 30 and 120 min, respectively at 613 K, f) and g) 60 and 120 min. at 773 K, respectively. 3.2. Catalyst activity in the HDS of thiophene and dibenzothiophene Figure 3 gives thefirstorder reaction rate constant for the thiophene hydrodesulphurisation over the various catalysts. Clearly, a progressively increasing activity is observed at higher sulphidation temperatures, both for the catalyst sulphided at 623 K and 823 K. Note, that under mild sulphiding conditions the catalysts calcined at 673 K seems to perform better in the thiophene HDS. However, this difference levels off at a sulphiding temperature of 823 K

269

673 5.7

r

2.8

613

5.3

3.4

673

823

613

923

673

823

923

Tsulf [ K ]

Figure 3. The reaction rate constants for thiophene HDS of the various catalysts calcined at 673 K and 823 K, followed by sulphiding at 613 K, 673 K, 823 K and 923 K. Thefirstorder reaction rate constant for the hydrodesulphurisation of dibenzothiophene over the various catalysts is shown in Figure 4. For both catalysts calcined at 673 K and 823 K the same trend, i.e., a decreasing DBT HDS activity is observed at increasing sulphidation temperatures. The highest activity for the HDS of DBT is found upon 'low temperature sulphiding' of the catalyst calcined at 823 K. Note, that at the highest sulphiding temperature both catalysts have the same performance in the conversion of dibenzothiophene. Quite remarkably, the catalyst calcined at 673 K already reaches the low level of activity upon sulphidation at 823 K, whereas the catalyst calcined at 823 K still has an enhanced activity for the HDS of dibenzothiophene. 12 £.

'

673

"

12 ] 1 0.5

r 8f

8 4

6.1

c(D

5.5

•

3.0

3.3

^

^

D > F > B > E > A , the pseudofirst order rate constant of 4,6-DMDBT (also 4-MDBT, DBT) in gas oil C > D > F > B > E = A (fig. 6b - fig. 6d). The similarity of these rankings shows that the best catalyst is most efficient for removal of refractory sulfur species. The ranking obtained with 4,6-DMDBT in decane (fig.3) is similar to that of Total Sulfur Content in gas oil ; but, definitely, NiMo catalysts exhibited a much higher activity for 4,6-DMDBT in decane than for gas oil. The inhibitors present in gas oil (aromatics, nitrogen compounds, ...) and produced by the desulfurization (H2S) retarded more severely the HDS reaction over NiMo catalysts than over CoMo catalysts. The most active C catalyst desulfurized essentially through the isomerization and cracking reactions, both acid catalyzed routes, giving a number of products. Among them, mono and poly alkylated DBTs (not desulfurized

337 products), mono and polyalkylated biphenyls and hydrogenated derivatives (desulfurized products)(2), cracked mono-ring products (benzene, toluene, xylenes, cyclohexane, methylcyclohexanes ...)^^^ were detected. k(min-1 g-1)

4,6-DMDBT in dec.

0.025

k.a (min-1 g-1) Tots in gas oil

0.1

0.08 0.06

0.015

n H

4,6-DMDBT in decane Tots in gas oil

fig.3 : HDS activity

0.04 ^^^ 9^^ ^'' (pseudo-1.65 0 02 oi'der) and for 4,6DMDBT (first 0 order) in decane A D r n f f ^^^^ constants NiMo B had a high activity through its high hydrogenation route, giving a higher ratio of (HYD/DirectDesuif.)route products at 270^ = 12 compared to 4 for CoMo D and F catalysts. Hence, the activity difference between catalysts D and F is ascribed to their different number of active sites. For spent C catalyst, the former ratio is about 1.5. At higher temperature, the DirectDesuif. route is much favored, hence this ratio decreased to 0.3 at 360°C over CoMo and NiMo catalysts. 3.3. Aromatic and H2S Inhibitors in HDS of 4,6-DMDBT in decane * Naphthalene, tetralin and isobutylbenzene (0-40%w in decane) were used as aromatic model inhibitors present in the real feed. Such aromatics inhibited the HDS in the order of Naphthalene > Tetralin > Isobutylbenzene as shown in figure 4. The HDS products showed that the hydrogenation route was more affected than the direct desulfurization route on NiMo catalyst. Nevertheless, the extent of inhibition by aromatic partners appeared similar regardless of catalysts. Another inhibitor in gas oil may deactivate NiMo catalyst more compared to CoMo catalysts. According to a series of HDS tests of 4,6-DMDBT in decane+naphthalene over some Ni-Co-Mo catalysts, it was found that the rate constant at the hydrogenolysis step (the C-S bond breaking of the tetra-hydrogenated 4,6DMDBT, see fig. 7) is proportional to the %Co on the catalyst, for a total metal loading %Co+%Ni=3%w. The acidic catalyst C is very strongly inhibited by naphthalene which eliminated its major route, that of the cracking, giving then the classical 4,6DMDBT HDS products, dimethylbiphenyl and methyl-phenyl-methyl-

338 cyclohexane. This important inhibition is explained by the high aromatic electron density which leads to a strong adsorption on acidic sites. - F+Napht.

1 X, Ratio of Activity A

0.9

m

iBuBz series

— X —

• B+Napht. • D+Napht. " F+Tetral.

A- —•

B+Tetral.

X- —•

D+Tetral. F+iBuBz

10

20

30

40

----A---

B+iBuBz

X - - - D+iBuBz %Arom. in decane Fig.4 Ratio of Activity for HDS of 4,6-DMDBT in decane + aromatics

* Figure 5 shows 4,6-DMDBT HDS inhibition by H2S which was produced by dimethyldisulfide (D.M.D.S., 0-1.5%wS) in decane. At 270°C, H2S is a strong inhibitor as aromatics. Particularly, NiMo catalyst B is more inhibited by H2S than CoMo, explaining partly the lower activity of NiMo in the gas oil. The HDS products distribution showed the hydrogenation route is more affected by H2S than the direct desulfurization route on NiMo. An opposite trend was found over CoMo. According to a series of HDS tests of 4,6-DMDBT in decane+D.M.D.S. over some Ni"Co-Mo catalysts, it was confirmed that the hydrogenation rate constant of 4,6-DMDBT (first step in its HDS hydrogenation route) is proportional to the %Co on the catalyst, for a total metal loading %Co+%Ni = 3%w. Ni has a low promoting efficiency in presence of H2S. On the other hand, the hydrogenolysis step rate constant (second step in its HDS hydrogenation route, see fig. 7) is also proportional to the %Co on the catalyst, for the same total metal loading %Co+%Ni = 3%w. At least, the direct desulfurization route is more inhibited by H2S over CoMo catalysts than over NiMo catalysts. Except for C catalyst, H2S behaved as an inhibitor like aromatics : the direct desulfurization route on CoMo catalysts and the hydrogenation route on NiMo catalysts are the most affected. On C catalyst, whereas H2S is a weak inhibitor, the aromatics are very strong inhibitors.

339 Relat. Activ, ^ -D A

m

n D

r 1

0

I

I

0.5

1

_ %w S in dec. T

1.5

Fig.5 Ratio of Activity for HDS of 4,6-DMDBT in decane + H2S 3.4. HDS of L.C.O. compared with G.O. Figs. 6a-6d show the comparison between the HDS tests of L.C.O. and gas oil over the 6 catalysts under the same conditions. k in LCO

kinGOOM Cond.: 340^; - 2.4 MPa H2 - 0-60 min.

A

B

C

D

E

F

Fig.6a Total Sulfur Content in L.C.O. and G.O. (1.65 order)

0.018 k 4,6-DMDBT 0.016 (1/min.) 0.014 0.012 0.01 0.008 0.006 0.004 0.002 0 Hg^4,6-DMDBT in L.C.O. and ' (first order)

340

k 4-MDBT

kDBT (1/min.)

B

C

i

B

Fig.6c 4-MDBT in L.C.O. and G.O. Fig.6d DBT in L.C.O. and G.O. (first order) (first order) In L.C.O., all the catalysts had about the same activity for the total sulfur content, except catalyst D showed a definitely higher activity. For the identified sulfur species, it appeared that the ratio kin L.co. / kin G.O. is decreasing from DBT («0.5) > 4-MDBT («0.25) > 4,6-DMDBT («0.68). It means that the high aromatic content of L.C.O. affected more the hydrogenation route (main route for 4,6DMDBT) than the direct desulfurization route (main route for DBT). To distinguish the catalysts, NiMo catalysts showed a slightly higher activity than CoMo catalysts for HDS of DBT ; on the contrary, for 4,6-DMDBT, CoMo are superior to NiMo. It confirms that the direct desulfurization route is more inhibited by aromatics on CoMo than on NiMo ; on the contrary, the hydrogenation route is more inhibited by aromatics on NiMo than on CoMo. Acidic catalyst C had a comparable activity with the other ones, showing that isomerization & cracking routes are probably become minor routes in L.C.O. desulfurization because of large aromatic content in L.C.O.

4, D I S C U S S I O N The present study compared the catalytic activities of 6 catalysts for HDS of a gas oil and aim to clarify the high activity of CoMo on acidic supports. One can examine the HDS reactivity of 4,6-DMDBT in decane with inhibitors such as aromatic compounds and H2S to evaluate the performances and schemes for the available catalysts. The present comparisons of the catalysts clarified several points of discussion. The catalysts achieved the deep desulfurization by the desulfurization of 4,6-DMDBT tiirough the hydrogenation and acidic routes, which are both inhibited severely by aromatics partners, similarly to basic compounds. Both routes can be evaluated by standard tests using hydrogenation of naphthalene and dealkylation of isopropylbenzene

341

Hydrodesulfurization Reaction Scheme for 4^6-DiMethylDiBeiizoThiophene

rt^^S— r^^H

UJr^^

TT tion

ca

& Direct Desulfurization

Cracking & Isomerization

VA CH3

CH3

4,6.DMDBT

CH3 CH3 H-4,6DMDBT

Direct Desulfurization

CH3

CH3

C Product

Hydrogenolysis

CH3

CH3

B Product

CH3

CH3

A Product

J

Cracking •> Benzene, Toluene, Xylenes, Cyclohexane, Methylcyclohexanes Fig. 7: HDS reaction scheme for 4,6-dimethyldibenzothiophene

342 respectively. Very acidic zeolite-containing catalyst allows the isomerization and cracking of sterically hindering methyl groups, such as for 4,6-DMDBT in decane. However such catalysts suffer the decrease of these properties with the time on stream by coking and are inhibited by aromatic partners. So, the importance of acidic routes is lower for HDS of real feed stocks than for that of 4,6-DMDBT in decane. Silica-alumina supported and zeolite containing acidic catalysts exhibited a very higher HDS activity for 4,6-DMDBT in decane compared to the other catalysts, but their activity is leveled off for HDS of high aromatic content L.C.O. real feed stock. Inhibition by H2S is important for the studies on model molecule at 270°C, but this effect may be lower at higher temperature, under practical conditions, owing to the H2S adsorption constant decreasing with the temperature. Interestingly, acidity of the catalyst appears to enhance the hydrogenation activity of CoMo which accelerates the desulfurization. Lower coking acidity obtained with silica-alumina support provides a high hydrogenation which is less inhibited by aromatics, explaining the high activity for HDS of G.O. and the highest activity for HDS of L.C.O. Strong adsorption of both aromatics and H2S inhibitors is involved on alumina-supported catalysts. The same adsorption phenomena could explain the similarity of qualitative inhibition effects ; indeed, aromatics and H2S affected severely the direct desulfurization route on CoMo and the hydrogenation route on NiMo. Deep desulfurization (ex. 4,6-DMDBT desulfurization) going through mainly hydrogenation route, CoMo suffer much less inhibitions by aromatics partners and H2S than NiMo, being superior in the practical desulfurization where the inhibitors are always present. The inhibition of H2S appears less on the acidic catalysts. Hence, the catalyst which is active and selective for the hydrogenation of 4,6-DMDBT in presence of inhibitors can be a target of better performance. The support for CoMo catalysts can be thus explored in terms of controlled acidity and better dispersion of CoMo.

Acknowledgments:

We gratefully acknowledge ELF-ANTAR FRANCE for supporting this work, Japan Energy Co. and Haldor Topsoe A/S for supplying experimental and reference commercial catalysts. 1. R. Gerdil and E. Lucken, J. Am. Chem. Soc. 87 (1965) 213 2. T. Isoda, S. Nagao, X. Ma, K. Sakanishi, I. Mochida Japanese Petroleum Institute bi-annual Conference, October 1994 3. D. Yitzhaki, M.V. Landau, D. Berger, M. Herskowitz Applied Catalysis A: General 122 (1995) 99


343

HYDROTREATING OF COMPOUNDS AND MIXTURES OF COMPOUNDS HAVING MERCAPTO AND HYDROXYL GROUPS T.-R. Viljava and A. O. I. Krause Helsinki University of Technology, Department of Chemical Technology Kemistintie 1, FIN-02150 Espoo, Finland ABSTRACT Simultaneous hydrodesulfurization (HDS) and hydrodeoxygenation (HDO) of mercapto and hydroxyl group containing benzenes was studied using a commercial presulfided C0M0/7AI2O3 catalyst under hydrotreating conditions (150-280 °C, 7 MPa). Mercaptobenzene, phenol and 4-mercaptophenol were used as model compounds, and CS2 was used as precursor for HjS. The HDS rate of a mercapto group in the presence of a hydroxyl substituent in the para position was higher than that for the molecule containing only a mercapto group. When the hydroxyl group was present as phenol, the HDS rate of the mercapto group was about 30% lower than that for mercaptobenzene without an oxygen-containing additive. The decrease in the HDS rate was independent of the initial molar ratio of sulfur and oxygen within the ratios studied (5:1-1:1). The HDO rate of a hydroxyl group was suppressed by the mercapto group present either in the same or in a separate molecule. HDO reactions did not start until HDS conversion was almost complete. CS2 also decreased the HDO rate of phenol. When compared to the reactions of phenol alone, the rate of the hydrogenolysis route to benzene was decreased in the presence of a sulfur additive more than the hydrogenolysishydrogenation route to cyclohexane. 1. INTRODUCTION Hydrodesulfurization (HDS) is of great importance in the oil-refining industry. In the long term, it is likely that biomass will be used as an alternative raw material for liquid fuels and chemicals as such or in mixtures with traditional feedstocks. A new kind of hydroprocessing is needed to treat the feeds containing considerable amounts of both oxygen and sulfur. Reactions taking place during hydroprocessing have mostly been studied using model compounds [1-3]. Benzothiophenes and dibenzothiophenes represent the least reactive organic sulfur compounds in fossil fuels, and thus their reactions have been most intensively investigated. Phenol and naphthol derivatives and heterocyclic oxygen compounds have generally been used as oxygen-containing model compounds [1]. Interactions between different heteroatoms during hydrotreating have typically been studied with mixtures of model compounds. HDS reactions have turned out to be slightly inhibited by the oxygen-containing compounds [1,3].Only a few studies deal with the hydrotreating of

344 compounds containing both sulfur and oxygen in the same molecule [4-10]. In the cases where the model compound contained both oxygen and sulfur in different substituents of a benzene ring, the HDS reactions were promoted considerably in the presence of the oxygencontaining substituent compared to the corresponding oxygen-free compound. This phenomenon has been explained by the increased electron density of the sulfur atom in the presence of an oxygen-containing substituent [8-10]. Organosulfur compounds seem to have only a weak effect on hydrodeoxygenation (HDO) [1]. However, the role of sulfur, especially the role of H2S [3], in HDO is not quite clear so far. Sulfur is, to some extent, needed in HDO to maintain the sulfidation of the catalyst. On the other hand, competitive adsorption of the sulfur compound and HjS formed from the sulftir compound may have an inhibiting effect in HDO [1,11]. In addition, sulfur-containing compounds may alter the selectivity of HDO [11,12]. So far, the hydrotreating of binary mixtures of sulfur and oxygen compounds has only been studied with a thiophenic compound, or H2S as a sulfur-containing reactant. No reports have been presented concerning interactions in hydrotreating of mixtures of sulfur and oxygen compounds with similar structures, e.g. mercaptobenzene and phenol. In order to get deeper on understanding, we have carried out a series of experiments with compounds and binary mixtures of compounds containing a hydroxyl group and a mercapto group either in the same or in a separate molecule. The effect of H2S, formed from CS2 precursor, on the HDO of a phenolic hydroxyl has also been investigated. The kinetic parameters of the HDS and HDO reactions of the substituents are compared in this paper. 2. EXPERIMENTAL 2.1. Catalyst The catalyst was a commercial hydrodesulfurization catalyst, Ketjenfine 742-1.3Q, which contained 4.4 wt-% of CoO and 15 wt-% of M0O3 on 7-AI2O3. The catalyst was crushed and sieved to a fraction of 0.75-1.0 mm, and presulfided off site with CS2/n-hexane at 280 °C under hydrogen. 2.2. Hydrotreating procedure Hydrotreating of mercaptobenzene (Merck, >98%), 4-mercaptophenol (Aldrich, >90%), phenol (Carlo Erba, >99.5%) and carbon disulfide (Merck, 99.99%) was studied using a 50 ml stainless steel autoclave at 150-280 °C and at a total pressure of 7.2-8.4 MPa. The substrate or the mixture of substrates in m-xylene (Merck, >99%) was added to the preheated reactor via a high pressure feed vessel. Decane (Fluka, >99.5%) was used as a tracer during the reactions. The amount of solvent, tracer and catalyst in the experiments was constant, being 15 ml, 200 /xl and 0.5 g, respectively. CS2 was used as a precursor for H2S. 4-10 runs were carried out with each model compound or a mixture of model compounds. Four to six samples of 100-200 mg were withdrawn from the reactor liquid phase and analyzed by gas chromatography (HP 5890 A, flame ionization detector, capillary column DB-1). The accuracy of the analysis method was within + 5% for sulfur-containing aromatics, and ±2% for phenol and hydrocarbons. A more detailed description of the hydrotreating procedure is presented in our previous paper [10].

345 2.3. Phase equilibrium in the reactor To compensate for the lack of quantitative gas phase analysis, the gas phase composition in equilibrium with the analyzed liquid phase was estimated as described in detail earlier [13]. 2.4. Conversions and kinetic parameters The total conversions of HDS and HDO were defined as conversion of sulfur or oxygen in the reactant to HjS and H2O. The reactions were assumed to be first-order with respect to the sulfur or oxygencontaining reactant and the concentration of hydrogen in the reaction mixture was assumed to be constant. The following rate equation was used for the HDS and HDO reactions: rate = r^m^^Cj,

(1)

where k' is the pseudo reaction rate constant, mât the amount of catalyst and C^ the concentration of the reactant. The parameters for the reactions of the known intermediates were fitted separately. All the reaction data available for one reactant or a mixture of reactants were processed simultaneously. The MODEST model estimation program [14] was used for calculations. 3. RESULTS Examples of the composition of the hydrotreated product at 225 °C are given in Table 1. On the basis of the product compositions, the simplified reaction networks for mercaptobenzene, phenol and 4-mercaptophenol can be presented as shown in Figure 1. Direct hydrogenolysis of the aromatic carbon-sulfiir bond was the main reaction of the mercapto group. No ring hydrogenated products were detected for mercaptobenzene, and no compounds containing sulfur but not oxygen were found in the reaction product of 4mercaptophenol. The HDO of phenol proceeded via both hydrogenolysis of the carbonoxygen bond to form benzene and the hydrogenation-hydrogenolysis route to form cyclohexane and cyclohexene [10]. HDS of mercaptobenzene was much easier than HDO of the oxygen compound of similar structure, phenol, under the same experimental conditions. The HDS conversions were almost complete in the reaction times in which the HDO conversions were still below 10%. The HDS conversion the disubstituted model compound, 4-mercaptophenol, was clearly higher than that of the corresponding oxygen-free compound, mercaptobenzene under the same reaction conditions (see Figure 2a). On the other hand, the conversion of HDO of 4mercaptophenol was much lower than that of phenol (see Figure 2b). In studies with binary mixtures of phenol and a sulfur-containing compound, mercaptobenzene or CS2, the same reaction products for HDS and HDO were detected as with mercaptobenzene and phenol alone. However, the HDO started at higher temperatures and the HDO conversion was much lower than that for phenol alone under the same reaction conditions. The conversions as a function of reaction time at 250 °C for pure phenol, 4mercaptophenol and phenol in the presence of CS2 arepresented in Figure 2b. It was not possible to determine the HDO conversion in the experiments with mixtures of phenol and

346 |SH

(a)

0^ —

+ H2,-H2S

0 ^ ^

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Fluid Catalytic Cracking V (Studies in Surface Science and Catalysis)

Photochemistry on Solid Surfaces (Studies in Surface Science and Catalysis)

Fluid Catalytic Cracking V (Studies in Surface Science and Catalysis)

Hydrocracking science and technology

Past and Present in DeNOx Catalysis: From Molecular Modelling to Chemical Engineering, Volume 171 (Studies in Surface Science and Catalysis) (Studies in Surface Science and Catalysis)

Surface and nanomolecular catalysis

Surface and Nanomolecular Catalysis

Catalysis and Automotive Pollution Control (Studies in Surface Science and Catalysis)

Catalysis and Automotive Pollution Control III (Studies in Surface Science and Catalysis, Volume 96)

Catalysis by Metals and Alloys (Studies in Surface Science and Catalysis)

Catalysis by Acids and Bases (Studies in Surface Science and Catalysis)

New Frontiers in Catalysis, Parts A-C (Studies in Surface Science and Catalysis)

New Developments in Selective Oxidation by Heterogeneous Catalysis: Proceedings (Studies in Surface Science and Catalysis)

Fluid Catalytic Cracking VII:, Volume 166: Materials, Methods and Process Innovations (Studies in Surface Science and Catalysis) (Studies in Surface Science and Catalysis)

Advanced Zeolite Science and Applications, Volume 85 (Studies in Surface Science and Catalysis)

Studies in Surface Science and Catalysis 110. 3rd World Congress on Oxidation Catalysis

12th International Congress on Catalysis (Studies in Surface Science and Catalysis)

Catalysis by Zeolites: International Symposium Proceedings (Studies in surface science and catalysis)

Laboratory Studies of Heterogeneous Catalytic Processes (Studies in Surface Science and Catalysis)

Equilibria and Dynamics of Gas Adsorption on Heterogeneous Solid Surfaces (Studies in Surface Science and Catalysis)

Reaction Kinetics and the Development of Catalytic Processes (Studies in Surface Science and Catalysis)

Reaction Kinetics and the Development of Catalytic Processes (Studies in Surface Science and Catalysis)

Spillover and Mobility of Species on Solid Surfaces ,Volume 138 (Studies in Surface Science and Catalysis)

Equilibria and Dynamics of Gas Adsorption on Heterogeneous Solid Surfaces (Studies in Surface Science and Catalysis)

Hydrotreatment and Hydrocracking of Oil Fractions (Studies in Surface Science and Catalysis)