Hydrotreatment of coal-derived naphtha. Properties of zeolite-supported Ru sulfide catalysts

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A PT PA LL E IY DSS CA I A: GENERAL

Applied Catalysis A: General 151 (1997) 423-435

ELSEVIER

Hydrotreatment of coal-derived naphtha. Properties of zeolite-supported Ru sulfide catalysts Shuh-Jeng Liaw, Rongguang Lin, Ajoy Raje, Burtron H. Davis * Center for Applied Energy Research, University of Kentucky, 3572 Iron Works Pike, Lexington, KY 40511, USA

Received 13 May 1996; revised 16 August 1996; accepted 21 August 1996

Abstract

A Ru (0.77 wt.%)/zeolite catalyst was prepared and tested for the hydrotreatment of Illinois No. 6 coal-derived naphtha. It was found that this catalyst exhibited significant hydrodenitrogenation (HDN) and hydrodesulfurization (HDS) activities. Furthermore, the nitrogen compounds convert as easily as sulfur compounds, in contrast to most of the hydrotreatment catalysts where the HDS reaction is much more rapid than HDN. A comparison on the basis of nitrogen converted per gram of active metal or on the cost of the catalyst indicates that for HDN, the Ru(0.77)/zeolite is superior to a Ru(0.77)/alumina, a commercial C o - M o / a l u m i n a or a N i - M o / a l u m i n a catalyst. Keywords: Y-zeolite; Ru-Y-zeolite; Naphtha, Hydrotreating; Naphtha, Coal-derived; Hydrodesulfurization; Hydrodenitrogenation; Ruthenium

1. Introduction Since the supply of crude petroleum is limited, attention is shifting to developing means to convert heavy fractions into transportation fuels. Furthermore, tomorrow's gasoline must have a much lower heteroatom content than is allowed even today. Many investigations of the hydrotreatment of coal-derived naphtha show that current commercial catalysts, based on molybdenum sulfide, do not have sufficient activity to meet today's refinery requirements determined by environmental standards [1-5]. To attain compliance in the future, a new hydrotreatment catalyst and/or process is necessary for hydrodesulfurization (HDS) and hydrodenitrogenation (HDN) of feedstocks. * Corresponding author. Phone: [email protected].

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Pecoraro and Chianelli [6] reported that unsupported Ru sulfide catalysts exhibited an activity for HDS of dibenzothiophene that was at least 10-times higher than that of Mo sulfide. In earlier studies, an unsupported Ru sulfide catalyst was found to be the most active catalyst among the transition metal sulfides for both HDS and HDN of coal-derived naphtha [7,8]. An alumina-supported Ru (5.1 wt.%) sulfide catalyst, which has the same phase as the unsupported Ru sulfide catalyst, RuS2, was found to be more active in total catalyst weight than commercial N i - M o and C o - M o catalysts for HDN and as active as for HDS of coal-derived naphtha [7]. However, from an industrial viewpoint, any new hydrotreating catalyst should not only present an improved activity over the commercial catalysts but the price must also be acceptable. A low loading of Ru metal, e.g. less than 1 wt.%, is needed if this metal is to be a component of a practical catalyst. Unfortunately, our study showed that a low loading of Ru (0.77 wt.%) on alumina was not as active as either the C o - M o or N i - M o commercial catalysts for hydrotreating of coal-derived naphtha at temperatures in the 250-400°C range [7]. A low loading concentration of Ru (less than 1 wt%) in a Y-zeolite is suitable for consideration for use in the refinery. Zeolite-supported ruthenium sulfide showed outstanding activity in the HDN of quinoline [9] and pyridine [10] as well as for the HDS of thiophene [11 ] and benzothiophene [9]. However, no attempts were made to study a low loading, e.g. less than 1 wt.%, of Ru/zeolite for the hydrotreatment of a real feedstock. In this study, we have carried out hydrotreatment of an Illinois No. 6 coal-derived naphtha using zeolite-supported ruthenium (0.77 wt.%) sulfide catalyst. This low loading Ru(0.77)/zeolite catalyst is compared with commercial C o - M o and N i - M o catalysts on two bases: activity per weight of metal and price.

1.1. Catalyst preparation Zeolite Y was obtained from Davison Chemical. Prior to use, zeolite Y was further ion-exchanged in aqueous solution of NaC1 (4M) at room temperature for 24 h and dried at 450°C for 5 h. This NaY zeolite has Si/A1 = 2.7, BET surface area = 270 m2/g, and pore volume = 0.23 cm3/g. To obtain a 0.77 wt% of Ru on NaY zeolite catalyst (denoted as Ru(0.77)/zeolite), an aqueous solution of Ru(NH3)6C13 (0.005M, 150 ml) was slowly titrated into a continuously stirred NaY zeolite slurry (10 g of NaY in 400 ml of water). After adding all the Ru(NH3)rC13 solution to the slurry, a sample of the aqueous phase was taken periodically to follow the Ru(III) ion exchange into the NaY zeolite using Inductively Coupled Plasma Atomic Emission Spectroscopy to analyze for Ru content. The data in Fig. 1 show that a period of 24 h is sufficient to exchange essentially all of the Ru(III) ions into the zeolite. After the exchange step, the slurry was filtered and washed with HzO until no C1- was detected (by AgNO3); it was then dried in an oven (ll0°C, air) for 2 h. The dry solid was

S.-J. Liaw et al. /Applied Catalysis A: General 151 (1997) 423-435

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Fig. 1. The amount of Ru left versus exchange time.

then re-slurried into 250 ml of H20 and a 10 wt.% H2S/H 2 gas mixture was bubbled through this slurry at room temperature to convert the Ru(III) to ruthenium sulfide. The pre-sulfidation at room temperature was conducted to simulate the method of preparing the unsupported Ru sulfide which was prepared at room temperature and produced a higher surface area than a similar one prepared at a higher temperature [6]. After a period of 24 h, the color of the slurry changed from white to gray, indicating that at least some of the Ru had been converted to the sulfide. The gray slurry was then filtered, washed and dried in air. For calcination and further sulfidation of the catalyst, the solid was placed in a glass tube and heated in a rate of l ° C / m i n to 400°C under a flow of 10% HzS/H 2. After maintaining the sample at 400°C for 5 h, it was cooled to room temperature under a flow of 10% HzS/H 2. The 0.77-wt.% Ru/alumina catalyst was prepared by impregnation of a y-alumina (surface a r e a = 210 mZ/g, pore v o l u m e = 0 . 6 6 cm3/g) with (NH4)zRuC15 .H20 solution, dried in air at ll0°C, then calcined in air at 400°C. The catalyst was then sulfided in gas flow consisting of 10% HzS in H e at 600 psig by heating at a rate of l ° C / m i n to 375°C and then maintained at this temperature overnight. Details of the preparation of the Ru(0.77)/alumina catalyst can be seen elsewhere [7]. Two commercial catalysts, Ni(3.1%)-Mo(13.3%) on alumina (Akzo KF-840 1.3 Q) with a BET surface area of 138 m 2 / g and Co(2.7%)-Mo(11.1%) on alumina (America Cyanamid HDS-1142A) with a BET surface area of 288 m2/g, have also been used to provide data to compare with the results for naphtha hydrotreatment with a zeolite-supported Ru sulfide catalyst. 1.2. Activity test

A fixed-bed reactor, operated in a concurrent downflow mode, was used for these studies. A Brooks Mass flow controller, Model 5850 E, was used to deliver a constant flow of gas. A Milton Roy miniPump Solvent Delivery system was used for adding the naphtha at the reaction pressure. The reactor, assembled from 1 / 4 in. O.D. tubing, was placed in a 1 in. X 12 in. Lindberg

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Table 1 Characterization of Illinois No. 6 coal-derived naphtha Element

Amount

C H N S O

85.6 wt.% 13.2 wt.% 1420 ppm 818 ppm 1.24 wt.%

1200°C tube furnace held in a vertical position. Preliminary studies indicated that 1/4-in. O.D. should be used since mass transfer resistance was observed for the 1/2-in. reactor [12]. A hand-loaded back-pressure regulator, rated to 4000 psig, was used to regulate the reaction pressure. During a period of catalyst testing, each condition was maintained for 24 h. Three samples were taken during the final 6 h of each steady state period to obtain data for heterocompound conversion for each condition. Prior to analysis, a sample was washed three times with distilled water to remove dissolved H2S and NH 3.

1.3. Naphtha analyses The naphtha used as the feedstock for this study was produced during the liquefaction of a bituminous Illinois No. 6 coal at the Wilsonville, Alabama Advanced Integrated Two Stage Liquefaction Plant. The sample was collected during Run 261. The elemental analyses of the naphtha are shown in Table 1. Total carbon and hydrogen analyses were performed using a Leco CHN analyzer. Total nitrogen contents of the feed and product were determined using a Dohrmann DN-100 total nitrogen analyzer equipped with a chemiluminescence detector. Total sulfur content was determined using a Xertex C-300 microcoulometer. The feed and hydrotreated Illinois No. 6 naphtha samples were analyzed for individual nitrogen compounds using a Thermionic Specific Detector (TSD) coupled to a Varian 3700 gas chromatograph fitted with a KOH-treated Carbowax column. Identification of the nitrogen compounds was accomplished by comparison of retention time and doping with standard compounds. Details of the identification of N and S heteroatom containing compounds have been published elsewhere [13]. 2. Results and discussion

2.1. Catalyst characterization Three samples, NaY Zeolite, NaY ion-exchange with Ru, and sulfided Ru(0.77)/zeolite, were characterized using X-ray diffraction (Rigaku X-ray

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diffractometer). All samples showed only the XRD features of the NaY zeolite. This result indicates that, under these experimental conditions, only small Ru particles are present. Similar results were reported by Moretti and Sachtler [14] for the characterization of Pt-Cu clusters in NaY and by Harvey and Matheson [9] for a zeolite containing ruthenium sulfide; the latter authors set an upper limit of 10 ,~ for the size of ruthenium sulfide. Also, no Ru containing particles could be identified in the TEM (Hitachi H800NA, 3.6~, and 200 Kev) pictures for the catalysts used in this study. Thus, XRD and TEM data lead to the conclusion that ruthenium sulfide was present in a well dispersed form. 2.2. Hydrotreating 2.2.1. Total remoual of N and S The Ru(0.77)/zeolite catalyst was tested for the hydrotreatment of the Illinois No. 6 naphtha at 275-400°C, 660 psig, and weight hourly space velocity (WHSV) of 3 g n a p h t h a / h / g catalyst. It was found that this catalyst exhibited significant HDN and HDS activities (Fig. 2). However, an interesting finding was that nitrogen compounds convert as easily as sulfur compounds using the Ru(0.77)/Zeolite catalyst (Fig. 3); this contrasts to most of the hydrotreatment catalysts where HDS is much more rapid than HDN. A similar result was reported [9] for a Y-zeolite supported ruthenium catalyst ( ~ 5 wt% Ru) for the conversion of model compounds since quinoline was removed more rapidly than dibenzothiophene. To determine the reaction order of the HDS and HDN of naphtha using the Ru(0.77)/Zeolite catalyst, the reaction temperature (350°C) and pressure (660 psig) were held constant while the flow rate of naphtha was varied. The results show that for coal-derived naphtha, the overall rates of HDN and HDS follow a first-order reaction (Fig. 4).

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2.2.2. Individual nitrogen Characterization of the Illinois No. 6 naphtha using a thermionic specific detector coupled with the capillary GC shows that the naphtha contains about 300-400 nitrogen compounds [13]. The major nitrogen class of the naphtha is anilines, particularly aniline and 1 to 4 carbon group(s)-substituted anilines. The next most abundant classes of nitrogen compounds are pyridines and quinolines. The total nitrogen analyses provide a measure of total removal of nitrogen from the naphtha. Also, the conversion of each individual nitrogen compound, on a mass basis, can be calculated from the TSD chromatogram as follows: X(i)% = [Area(i),f- Area(i),p]/Area(i),f. lO0

(1)

where i denotes individual compound, Area denotes the peak area, f denotes a feed component, and p denotes the same component after reaction. Among the three major nitrogen classes, pyridines are much easier to convert than anilines and quinolines; this agrees with the results using Ru alumina catalysts [7]. The removal of individual aniline compounds follows first-order kinetics only at higher flow rates (Figs. 5 and 6). The deviation from first-order

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1/WHSV, hr. Fig. 4. Fractional amount of heteroatoms (N and S) remaining in hydrotreated product as a function of hourly space velocity for Ru (0.07)/zeolite catalyst at 350°C and 660 psig. ( • ) HDN and ( 0 ) HDS.

S.-J. Liaw et al. /Applied Catalysis A: General 151 (1997) 423-435

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Fig. 5. Kinetic plots for the conversion of aniline class compounds using Ru (0.07)/zeolite catalyst at 350°C and 660 psig. ( • ) 2-methylaniline, (Q) 3-methylaniline, (zx) 4-methylaniline, and (O) aniline.

kinetics at lower flow rates may be due to the impact of H2S, H20 and/or NH 3 on the HDN kinetics. This phenomenon was also observed when MoS 2 and MozN catalysts were used to convert this naphtha [13]. At the lower flow rates the conversions are higher, and more HzS, H20 and NH 3 are generated. Therefore, the impact of HzS, H20 and/or NH 3 on HDN should be more important at lower flow rates. The individual quinoline compounds, in contrast, follow a first-order kinetics throughout the range of flow rates used in this study (Fig. 7). Thus, the adsorption of quinoline compounds is not impacted by HzS, H20 and/or NH 3 to the extent of the anilines.

1

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0.15

A 2-ethylaniline • 3--ethylaniline zx 4-cthylanilinc o 2,3-dimefftylanilincv 3.5MirnethylmtUine

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Fig. 6. Kinetic plots for the conversion of aniline class compounds using Ru (0.07)/zeolite catalyst at 350°C and 660 psig. ( • ) 2-ethylaniline, (Q) 3-ethylaniline, (zx) 4-ethylaniline, (©) 2,3-dimethylaniline, and ( v ) 3,5-dimethylaniline.

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0.01 0

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[ A Quinoline

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Fig. 7. Kinetics plot for the conversion of quinoline class compounds using Ru (0.07)/zeolite catalyst at 350°C and 660 psig. ( • ) quinoline, ( O ) 3-methylquinoline, and ( ~ ) 4-methylquinoline.

2.3. Comparison of support material The activity of ruthenium in a zeolite catalyst was compared with that of an alumina-supported catalyst. The data in Fig. 8 show that at the same loading of Ru (0.77 wt.%), the zeolite-supported catalyst is much more active for HDN than the alumina-supported catalyst. Tait and Hensley [15] reported that for HDN of a shale oil, a C o - M o / z e o l i t e catalyst was more active than a Co-Mo/alumina catalyst. For HDS, the alumina and zeolite catalysts containing Ru exhibited similar activity (Fig. 8). Mazur et al. [16] reported similar results showing that X and Y zeolite-supported C o - M o catalysts had HDS activities that were comparable to, or only slightly lower than, commercial catalysts for hydrotreating a vacuum distillate. I00 .

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0

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[IRu(0.77)/Alurnina [ ]

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Ru(0.77)/Zeolite j

Fig. 8. Comparison of supported material on hydrotreating of coal-derived naphtha at 300°C and WHSM = l. ( • ) Ru (0.07)/alumina, and ([]) Ru (0.07)/zeolite.

S.-J. Liaw et al. / Applied Catalysis A: General 151 (1997) 423-435

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0.2

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zx Ni-Mo/AI203

1

Fig. 9. Fractional amount of nitrogen rcmaining in hydrotreated product as a function of hourly space velocity at 350°C and 660 psig. ( • ) Ru (0.07)/zeolite, ( O ) Co-Mo/Al203, and (zx) Ni-Mo/A1203.

2.4. Comparison of a Ru(O. 77)/zeolite catalyst to commercial catalysts The two commercial catalysts were ground and the fraction passing through a 200-mesh sieve was used; therefore, the granular size of all catalysts were less than 74 )xm. In this set of experiments, the temperature and pressure were held constant at 350°C and 660 psig, respectively, while the WHSV of naphtha was varied. The flow of hydrogen was also varied to maintain the molar ratio of naphtha to hydrogen constant at 1:2.6. Plots of the logarithm of the fractional amount of nitrogen and sulfur

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Fig. 10. Fractional amount of sulfur remaining in hydrotreated product as a function of hourly space velocity at 350°C and 660 psig. ( • ) Ru (0.07)/zeolite, ( Q ) Co-Mo/AI203, and (/x) Ni-Mo/AI203.

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Ni-Mo-Alumina

[..o.G.Ds] Fig. 11. First-order rate constant for HDN and HDS (based on three grams of catalyst). ( • ) HDN and ([]) HDS.

remaining in the naphtha as a function of space time (reciprocal of WHSV) for the zeolite and commercial catalysts are shown in Figs. 9 and 10, respectively. The overall HDN of naphtha is a first-order reaction for all three catalysts (Fig. 9). The first order rate constant, kapparent (based on 3 g of catalyst used), of HDN can be obtained from the slope of the lines shown in Fig. 9. For HDN of coal-derived naphtha at 350°C, the Ru(0.77)/zeolite catalyst, on a total weight basis, is approximately 1.5 times more active than the N i - M o catalyst and 2.4 times more active than the C o - M o catalyst (Fig. 11). For the zeolite catalyst, the plot of HDS (Fig. 10) is similar to the HDN shown in Fig. 9 so that first-order kinetics apply for the HDS of naphtha. However, for the commercial catalysts, while the data can be plotted in a straight line, the line extrapolates to

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S.-J. Liaw et al. /Applied Catalysis A: General 151 (1997) 423-435

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an initial sulfur concentration (zero space time) that is smaller than the actual concentration of sulfur in the feed (Fig. 10). A similar result has been reported for the hydrotreatment of naphtha using Mo2N and MoS 2 catalysts; a two-lump kinetic model was proposed to account for the HDS and HDN of naphtha with these two catalysts [17]. To compare the activity of HDS among these three catalysts, first-order kinetics was assumed for the commercial catalyst in the range of the space time used in this study. The C o - M o catalyst is as active as the N i - M o catalyst and they are approximately two times more active catalyst for HDS of coal-derived naphtha than the Ru(0.77)/zeolite (Fig. 11). Another approach is to make a comparison of the activity on the basis of the weight of active metal. When this is done, the Ru(0.77)/zeolite is much more active than the commercial catalysts for both HDS and HDN of naphtha (Fig. 12). This is due to the low loading, 0.77 wt.%, of Ru on zeolite compared to the 13.8 wt.% and 16.4 wt.% of metals present in the C o - M o and N i - M o catalysts, respectively. Although the Ru(0.77)/zeolite is a potential substitute for the commercial hydrotreating catalyst because of its high activity for HDN for the hydrotreatment of naphtha, the price of Ru metal is a major concern. Today, the price of materials used for the preparation of a C o - M o catalyst and a Ru zeolite catalyst are as follows 1:Co(NO3)2' $5.5/lb; (NHa)6MOTO24, $5.00/lb; RuC13 solution, $242.3/lh, Y-zeolite, $0.6/lb; and A120 3, $0.6/lb. Based on this information, a comparison on the basis of the price of the catalyst for the HDN and HDS of naphtha can be made. The first-order rate constant for HDN, normalized to the price of the catalyst, is approximately I. 1 times larger for the Ru(0.77)/zeolite catalyst than for the C o - M o catalyst (Fig. I Quotes from Engelhard Corp. and Climax Molybdenum Co., August 5th, 1996 published trading price.

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S.-J. Liaw et aL / Applied Catalysis A: General 151 (1997) 423-435

13). However, for HDS, based on the price of the catalyst, the C o - M o catalyst is more active than the Ru(0.77)/zeolite catalyst (Fig. 13).

3. Conclusion For hydrotreating naphtha, the Ru (0.77 wt.%) zeolite catalyst showed high activities for HDS and HDN. For the Ru(0.77)/zeolite catalyst, the rates for HDN and HDS of the naphtha are essentially equal, which is contrary to conventional hydrotreating catalysts. Nitrogen compounds convert as easily as sulfur compounds over Ru-zeolite catalyst, which is in contrast to commercial catalysts, and different kinetics indicate different mechanisms between R u zeolite and commercial catalysts. Although information on commercial catalysts is extensive, that on Ru-zeolite is rather limited. Thus, hydrotreatment of the model pairs such as pyridine-quinoline, thiophene-benzothiophene, furan-benzofuran, and pyrrole-indole can be used to study the fundamental difference in the hydrogen transfer from the catalyst to the reactant molecule as well as in the activation of the C - N and C - S bonds. A comparison between the activity for HDN and HDS of a naphtha for a Ru(0.77)/zeolite and a commercial C o - M o catalyst was made on the basis of 1 g of active metal and on the price of the catalyst. This comparison shows that the Ru(0.77)/zeolite is more active for HDN than the commercial C o - M o catalyst. However, for HDS, the commercial catalyst is more active than the zeolite catalyst except for a comparison based on the weight of active metal. Based on these results, it appears that a process variation, where the easily convened nitrogen and sulfur compounds would be convened using the conventional catalyst while deep HDN would be accomplished with the more active Ru/zeolite catalyst, is reasonable. This would permit the sulfur and some nitrogen to be removed with the cheaper, current commercial catalyst and while the Ru/zeolite catalyst can be used to reduce the nitrogen content to a low level.

Acknowledgements This work was supported by the DOE contract #DE-AC22-90PC91058 and the Commonwealth of Kentucky. The authors also thank the personnel at Wilsonville liquefaction facility for providing the naphtha sample.

References [1] H. Fmmkin, R.F. Sullivan and B.E. Strangeland, in R.F. Sullivan (Ed.), Upgrading Coal Liquids, Am. Chem. Soc. Symp. Ser., 156, 1981, 75. [2] V.E. Smith, C.Y. Cha, N.W. Merriam, J. Faky and F. Guffy, Proceedings of the Third Annual Oil Shale, Tar Sand and Mild Gasification Contractor's Review Meeting, DOE/METC-88/6098, 1988, p. 166.

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[3] U.R. Gaeser, R. Holighaus, K.D. Dohms and J. Langhoff, Am. Chem. Soc., Div. Fuel Chem. Prepr., 33 (1988) 339. [4] R.J. Parker, P. Mohammed and J. Wilson, Am. Chem. Soc., Div. Fuel Chem. Prepr., 33 (1988) 135. [5] L. Xu, R.A. Keogh, C.S. Huang, R.L. Spicer, D.E. Sparks, S. Lambert, G.A. Thomas and B.H. Davis, Am. Chem. Soc., Div. Fuel Chem. Prepr., 36 (1991) 1909. [6] T.A. Pecoraro and R.R. Chianelli, J. Catal., 67 (1981) 430. [7] S.-J. Liaw, A. Raje, R. Lin and B.H. Davis, Am. Chem. Soc., Div. Fuel Chem. Prepr., 39 (1994) 636. [8] S.-J. Liaw, A. Raje and B.H. Davis, unpublished results. [9] T.G. Harvey and T.W. Matheson, J. Catal., 101 (1986) 253. [10] S. GSbiSltis, M. Breysse, M. Cattenot, T. Decamp, M. Lacroix, J.L. Portefaix and M. Vrinat, in M.L. Occelli and R. Anthony (Eds.), Advances in Hydrotreating Catalysts, 1, Elsevier, Amsterdam, 1989, p. 243. [11] S. GiSbiiltis, M. Lacroix, T. Decamp, M. Vrinat and M. Breysse, Bull. Soc. Chim. Belg., 100 (1991) 907. [12] S..-J. Liaw, A. Raje, X.X. Bi, P.E. Eklund, U.M. Graham and B.H. Davis, Energy and Fuels, 9 (1995) 921. [13] S.-J. Liaw, A. Raje, K.V.R. Chary and B.H. Davis, Appl. Catal. A: General, 123 (1995) 251. [14] G. Moretti and W.M.H. Sachtler, J. Catal., 15 (1989) 205. [15] A.M. Tait and A.L. Hensley, Am. Chem. Soc., Div. Petr. Chem. Prepr., 26 (1981) 924. [16] K. Mazur, J. Wrzyszcz and M. Ihnatowicz, Nafta (Katowice, Poland), 36 (1980) 165. [17] A. Raje, S.-J. Liaw, K.V.R. Chary and B.H. Davis, Appl. Catal. A: General, 123 (1995) 229.