1354
Ind. Eng. Chem. Res. 1993,32, 1354-1358
Hydrotreatment of Athabasca Bitumen Derived Gas Oil over Ni-Mo, Ni-W, and Co-Mo Catalysts? Rafael A. Diaz-Real3 Ranveer S. Mannt. and Inderjet S . Sambil Department of Chemical Engineering, University of Ottawa, Ottawa, Canada, KIN 6N5
The hydrotreatment of Athabasca bitumen derived heavy gas oil containing 4.08% S and 0.49% N was carried out in a trickle bed reactor over Ni-W, Ni-Mo, and Co-Mo catalysts supported on zeolite-alumina-silica at 623-698 K, LHSV of 1-4, gas flow rate of 890 m3~.2/m3,,i1(5000 sef/bbl), and pressure of 6.89 MPa. Analyses for viscosity, density, aniline point, ASTM mid boiling point distillation, C/Hratio, and percentage of N and S in the final product were carried out to characterize the product oil. The amounts of N and S removed indicated the hydrodenitrogenation and hydrodesulfurization activity of the catalysts. Results of zeolite-alumina-silica-supported catalysts are compared t o those obtained with commercially available Ni-Mo, Ni-W, and Co-Mo on y-alumina. Ni-Mo supported on zeolite-alumina-silica was most active and could remove as much as 99 5% S and 89% N present in the oil at 698 K. The data for HDN and HDS fitted the pseudo first order model. The kinetic model is described in detail. Introduction Crude oils are classified for commercial purposes into light ("API > 35), medium (20 IOAPI I35), and heavy oils ("API < 20). Heavy oils contain objectionable heteroatomic compounds, such as those containing sulfur and nitrogen which pose severe difficulties in further processing. Transforming crude oils into finished products requires several and sometimes complicated processes. Catalytic hydrotreatment is widely used to remove sulfur and nitrogen of distillates such as cat reformer feed, catalytic cracking feed, petroleum products such as kerosine,jet fuel, diesel and heating fluids, as well as heavy gas oils. The hydrotreating catalysts contain metals Co, Mo, Ni, and W as sulfides and are generally supported on silica-alumina or promoted alumina. Catalysts composed of Ni-Co-Mo or Ni-Mo supported on alumina are most efficient for nitrogen removal. Hydrodenitrogenation (HDN) is usually more difficult to achieve than hydrodesulfurization (HDS) for hydrocarbon streams. Thus, any hydrotreatment that reduces nitrogen to a satisfactory level will also achieve the same goal with sulfur. Since catalytic activity is generally proportional to the concentration of active sites, supports of materials having large areas are commonly employed. The preferred supports found in commercial sources are those made of silica, alumina, or a mixture of both. However, a further improvement has been the use of zeolites in the preparation of supports for catalysts. Zeolites are crystalline aluminosilicates whose carbonium ion activity has been attributed to strongly acidic protons in the crystalline lattice. Due to this reason zeolite-basedcatalysts have been found to be more active catalysts (Jacob, 1977;Venuto and Habib, 1979). They are more selective, and give more products, more paraffins and aromatics and less olefins (Scott, 1980; Satrina, 1982). The main objective of this study was to develop high-efficiency Ni-Mo, Ni-W, and Co-Mo catalysts using zeolite material and a composite of silicaalumina as support material. In this paper, we report the results of HDS and HDN of heavy gas oil derived from +Presented at the International Chemical Congress of the Pacific Basin Societies, Dec 15-19,1989,Honolulu, HI. 1Present address: Chemical Engineering Department, Universidad Iberoamericana, Mexico City, Mexico. 8 Present address: Bell Northern Research Laboratories, Nepean, Ontario, Canada. 0888-5885/93/2632-1354$04.00/0
Athabasca tar sand bitumen in a trickle bed reactor using high-efficiency Ni-Mo, Ni-W, and Co-Mo catalysts supported on a zeolite-alumina-silica support developed in our laboratory. The results obtained in this study are compared to those obtained earlier with commercially available Ni-Mo, Ni-W, and Co-Mo on alumina (Mann et al., 1987)and Ni-Mo on zeolite-alumina-silica support (Mann et al., 1988). Experimental Section The hydrotreatment of heavy gas oil was carried out in a trickle bed reactor. The heavy oil used in this study (Table I) was a 618-797 K heavy gas oil fraction derived from the hydrocracking of Athabasca tar sand bitumen at about 723 K with hydrogen at a pressure of about 13.8 MPa and liquid hourly space velocity (LHSV) of 1.0. The apparatus, experimental technique, and procedure for the analysis of the products were the same as described earlier (Mann et al., 1982; 1987). Viscosity was measured by using a Cannon Fenske Routine calibrated viscometer No. 400. Aniline points were determined following the procedure as per ASTM D-611. Simulated distillation as per ASTM D-2087 was utilized for determining the mid boiling point distribution. The ainount of carbon, hydrogen, nitrogen, and sulfur in the product was determined with a Perkin Elmer elemental analyzer No. 240 B. Catalyst In an earlier study (Mann et al., 1987)three commercial (Harshaw) catalysts were used for hydrodenitrogenation and hydrodesulfurization studies. The characteristics of these catalysts and those prepared in our laboratory are given in Table 11. The support for the catalysts prepared in our laboratory was a mixture of silica, alumina, and rare earth exchanged Y-zeolite (Table 111). The alumina-silica gel was prepared by mixing required amountsof aluminium nitrate and sodium silicate solutions and precipitating it by adding NH40H as described earlier (Mann et al., 1988). This gel was then well mixed with zeolite (Table 111). Extrusions of 9-mm size were made and dried first at room temperature and then at 333 K for 2 h. It was finally calcined at 973 K, crushed, and screened to 70-80 mesh. The support particles containing 25 wt % zeolite, 65 wt %, alumina and 10 wt % silica were impregnated with required amounts of solutions of nickel 0 1993 American Chemical Society
Ind. Eng. Chem. Res., Vol. 32, No. 7,1993 1355 Table I. Properties of Heavy Gas Oil specific gravity 16/16 OC OAPI gravity viscosity at 25 OC, MPa-s asphaltene, wt % aniline point, "C elemental analysis, wt % C H N S
0.9884 11.7 246.6 1.0 47.8 87.20 8.23 0.49 4.08
ASTM D-2887 Distillation temp, 374.7 465.8 502.4 536.4 596.4 650.2 675.4 706.5 741.5 808.4
K vol% 10
20
30
40
50
60
70
80
90
100
Table 11. Catalyst Characteristics surf. pore area, vol, catalyst composition mz/g mL/g Harshaw 400 3% coo-15% Moosa 220 0.50 Harshaw 100 3.8% Ni0-16.8% Moosa 190 0.54 Harshaw 4303 6.0% Ni0-19% WOaa 152 0.54 uo 100 2.24% Ni0-5.37% MoOab 293 0.597 uo 200 1.13% Ni0-4.2% WOab 323 0.309 uo 300 1.15% coo-6.14% MoOab 273 0.419
pore size, nm 9.1 11.36 14.22 4.08 1.93 3.09
I, Supported on y-alumina. b Supported on zeolite-alumina-silica. Table 111. Details of Zeolite Material composition Si02 65 % &03 23 % NazO 2% rare earth 11% form powder (38-53 pm) surface area 550 m2/g manufacturer Sterm Chemicals, USA; Cat. No. 14-8910
nitrate, cobalt nitrate, ammonium molybdate, and ammonium tungstate using the vacuum impregnation technique (Sambi, 1986). The catalysts were sulfided prior to use by passing a 10% by volume mixture of H2S in Hz at a pressure of 380-415 kPa. Sulfiding was carried out for 5 h at 723 K. After the end of sulfiding process, H2 was passed through the reactor for 30 min. The deactivation of the catalyst was tested by carrying five replicate runs at standard operating conditions: temperature of 623 K, pressure of 6.89 MPa, and LHSV of 1.5 at a gas flow rate of 890 m3 of H2/m3of oil. No significant changes in catalyst activity were noticed after more than 2000 h of continuous operation (about 200 h of kinetic study). A slow stream of H2 (-0.2 m3 of HdhL was maintained over the catalyst even when experimental data were not collected to maintain the stability of the catalyst. The concentrations of Si, Al, Ni, Mo, Co, and W were analyzed by using argon plasma atomic emission spectroscopy. The concentrations of Si and Al were also checked by using the gravimetric method (Vogel, 1978). Surface area of the catalysts was measured by the BET method (Fateh and Willingham, 1955) and an Accusorb physical adsorption analyzer Model 2100E. Pure volume of the catalyst particles was measured by using the carbon tetrachloride adsorption method (Benesi et al., 1955).
Results and Discussion The effect of temperature between 573 and 723 K and at LHSV of 1-4 on the catalytic hydrotreatment of heavy gas oil at 6.89 MPa was studied. Although a slightly different set of operating conditions was used with the
Table IV. Results of Kinetic Study with Catalysts. temp, K Ni-Mob Ni-Wb Co-Mob Ni-Moc Ni-WC (a) Densities, g/mL 573 0.946 0.963 0.971 623 0.953 0.947 0.959 0.9517 0.9705 648 0.9431 0.9521 673 0.939 0.934 0.948 0.9349 0.9501 698 0.9247 0.9484 723 0.911 0.907 0.941 (b)Viscosities, MPes 573 155 165 143 623 117 64.7 122 118.1 111.5 648 78.7 80.9 673 45 31.5 85 41.5 57.1 698 33.9 35.7 723 7.8 8.1 17 (c) C/H Ratio 0.63 0.69 573 0.68 623 0.67 0.62 0.66 0.62 0.57 648 0.62 0.58 673 0.64 0.62 0.65 0.63 0.60 698 0.64 0.61 723 0.66 0.61 0.65 (d) Aniline Point 573 50.00 49.30 48.00 52.00 51.10 45.55 51.20 623 50.70 648 52.50 52.60 50.40 673 50.5 50.5 52.00 51.62 698 49.20 49.45 723 38.10 40.80 40.5 (e) % N Removal 573 7 5 6 623 10 12 8 30.8 47.5 42.93 648 55.5 673 32 26 20 63.46 71.3 698 80.56 77.4 723 47 39 30 (0 % SRemoval 573 35 19 20.1 623 50 55 41 77.04 62.94 648 87.21 74.6 673 60 77 69.9 95.73 89.4 698 98.9 98.7 723 82 83 82
Co-Moc
0.9648 0.9573 0.9512 0.9377
97.6 80.8 54.7 35.3
0.63 0.65 0.65 0.64
51.40 52.8 51.55 49.95
55.34 55.56 69.7 77.01
62.91 72.23 75.68 92.4
Operating conditions: LHSV = 2, pressure = 6.89 MPa, gas flow rate 890 m2HJm90fi. Commercial Catalysts. Zeolite-supported catalysts.
three commercially available catalysts, a meaningful comparison can still be made. In the beginning and at the end as well as in between temperature changes, the activity of the catalyst was checked by collecting samples at the standard operating conditions. No significant catalyst deactivation was observed during the period of the entire kinetic study. The results of the analysis of the product samples are given in Table IV (Sambi, 1986; Dim-Real, 1988). (a) Reactor Performance. In kinetic study, the basic assumption is that the reactor operates essentially under the conditions of isothermal and plug flow. The performance of a laboratory trickle bed reactor can deviate substantially due to axial dispersion, channeling effects, and poor wetting efficiency. The axial dispersion is minimum (Mears, 1974) if Lld, (L= length of the ractor, d, = diameter of particle) is greater than 350. In this study since L/d, (L= 100 cm, d, = 0.02 cm) was 5000, it can be assumed that there was no significant axial dispersion or back mixing. The chances of wall flow being negligible and of a uniform liquid distribution are very good if dJd, (dt= diameter of reactor tube) is greater than 18 (Herskovitz and Smith, 1978). In the present study dJd, was 26. Hence, it can be assumed that the liquid distribution was very good with no adverse chan-
1356 Ind. Eng. Chem. Res., Vol. 32,No. 7, 1993
\-.
\
0.975 125.0
2
%
-
%~
--..I\
.:\ --\
,,
~
\-. 'x \~
-
95.0
2. c .Ln
0 u ,? 6 5 . 0 -
>
-..,
=Ni-W zed. o =Co-Mo zeol.
0.9001
573.0
I
598.0
I
I
623.0
6480
I
673.0
1
698.0
I
_Iv
7230
5730
5980
6230
Temperoture, K
Figure 1. Effect of temperature on density. Operating conditions: LHSV = 2, pressure = 6.89 MPa, gas flow rate = 890 ms~Jm30u.
neling effects. Wetting efficiencies in the trickle bed reactor are generally low. By diluting the catalyst bed with an equal volume of inert material (a-alumina), the wetting efficiencywas increased to a point where complete catalyst wetting can be assumed (van Klinken and van Dongen, 1980,de Bruijn, 1976). (b) Heat- and Mass-Transfer Effects. The catalyst had a particle size between 0.175 and 0.246 mm. The effectivenessfactor for this size of catalyst used in similar applicationshas been reported to be nearly 1 (Smith, 1981). Since the reactor used in this study had an internal diameter of about 0.52 cm, the temperature gradient in the radial direction of the catalyst bed was considered insignificant. The temperature and partial pressure gradients between the flowing fluid and the external surface of the catalyst were evaluated (Yang and Hougen, 1950). The maximum temperature difference of 1 K and the maximum pressure gradient of 10 Pa suggested that the external heat and mass transfer were negligible. (c) Effect of Temperature. Higher hydrogenolysis activity was obtained as the temperature increased. This is indicated by the decrease in density and viscosity of the product oil. Zeolite-supported catalysts gave similar or lower density and viscosity product oil than those produced with commercial catalysts at comparable temperatures. Table IV and Figures 1 and 2 show the effect of temperature on density and viscosity of product oil. The experiments indicated that higher hydrogenolysis activity was obtained as temperature increased. The density and viscosity of the product oil decreased with increasing process temperatures for both the commercial and the zeolite-supported catalysts. The effect of temperature on the C/H ratio was slight. While the insignificant increase in C/H ratio over zeolitealuminaailica-supportedcatalyst can be due to the slight cracking of the heavy oil and the removal of N and S, the decrease in C/H ratio can be attributed to some intake of Hz by oil as reflected by the decrease in its density. Both HDN and HDS activities increased with increasing temperatures. The activities shown by the zeolite-supported catalysts are nearly twice that observed with commercial catalysts, denoting a much higher efficiency of N and S removal. Table IV and Figures 3 and 4 show the effect of temperature on the percentage removal of nitrogen and sulfur from gas oil. Zeolite-
648.0
6730
6980
7230
Temperature, K
Figure 2. Effect of temperature on viscosity. Operating conditions: LHSV = 2, pressure = 6.89 MPa, gas flow rate = 890 m%J msou.
~
75.0
-
8
= Ni-W zeal. o =Co-Mo zeal. Ni-Mo 2801. + =Ni-W cornrn. x =Co-Mocornrn. 0 =Ni-Mo cornrn.
A
-
C
0 .-c C 0
01
50.0-
.-C
c L
01
_-
U
e
D >
r
25.0
-
'
0.0
573.0
I
,,'
.*
I
598.0
I
623.0
I
,'
I
648.0
,
..+"
_-
,-
I
673.0
I
698.0
72
Temperature, K
Figure 3. Effect of temperature on hydrodenitrcgenation. Operating conditions: LHSV = 2, pressure = 6.89 MPa, gas flow rate = 890 m3~Jm3,fi.
supported Co-Mo showed the best results at most operating conditions for HDN, whereas zeolite-supported NiMo gave the best results for HDS. Close to 100% desulfurization of the product was observed with the zeolite-supported catalysts. (d) Effect of Liquid Feed Rates. The effect of the change of the feed rates of heavy oil on the products for zeolite-supported catalysts was studied at 6.89 MPa and at temperatures of 623-723 K. The flow rate was varied between 8 and 32mL/h (LHSV = 1.0-4.0). Figure 5 shows the effect of LHSV on density. In all cases, the density of the product oil decreased with decreasing LHSV. The viscosity of the products also decreased with decrease in LHSV. The C/H ratio did not vary significantly with changes in LHSV. HDN increased with decreasing flow rates. The zeolite-supported Co-Mo catalyst was most effective for HDN at most operating conditions. HDS increased with decreasing flow rates. Zeolite-supported
Ind. Eng. Chem. Res., Vol. 32,No. 7,1993 1357 Table V. Rate Constants for HDS and HDN Studies. catalysts ks,h-' kN,h-1 commercial Ni-Mo 1.74 1.10 Ni-W 3.0 0.5 Co-Mo 2.08 0.86 zeolite Ni-Mo 2.16 1.50 Ni-W 2.41 0.52 Co-Mo 1.8 0.87
100.0
75.0-
Operating conditions: pressure m3~Jm30il, temperature = 648 K.
Table VI. Activation Energies (E)for HDS and HDN E, kcal/mol
zeol. zeol.
= Ni-W = Co-Mo A = Ni-Mo
o
catalyst Ni-Mo Ni-W Co-Mo
ZeOl.
+
=Ni-W wmm. x = Co-Mo comm. o-Ni-Mo cornm.
00'
1
I
598.0
5730
6230
648.0
6980
6730
7230
Temperature, K
Figure 4. Effect of temperature on hydrodesulfurization. Operating conditions: LHSV = 2,pressure = 6.89 MPa, gas flow rate = 890 m%IJmS/oil.
HDS 20.80 21.05 21.67
0.9251
1 o
-
Ni-W zeol. GO-MO ZWi.
I
A =Ni-Mo zeol.
0.900' 0.5
I
1.0
1.5
I
2.0
2.5
I
3.0
3.5
I
4.0
I
4.5
LHSV
Figure 5. Effect of LHSV on density. Operating conditions: temperature = 673 K, pressure = 6.89 MPa, gas flow rate = 890 m%IJmsOil.
Ni catalysts (Ni-Mo and Ni-W) showed higher HDS activities as compared with zeolite-supported Co-Mo catalyst . (e) Hydrotreatment without Catalyst. Some researchers have observed HDS and HDN occurring without a catalyst (Miki et al., 1983). The proportions of these reactions occurring without a catalyst depend on the feed type and the operating conditions. In order to evaluate the extent of thermal HDS and HDN for the present study, a set of experimental runs were carried out using the inert material (a-alumina) utilized to dilute the catalyst bed. An operating procedure similar to that in the case of the kinetic study was followed. A maximum of 18.3% HDN and 13.8% HDS occurred without a catalyst as compared with 87.7% HDN with the zeolite-supported Co-Mo catalyst and 99.5% with zeolite-supported Ni-W and NiMo catalysts. The gas formation rates were much lower as the cracking activity of a-alumina was less.
HDN 25.10 18.90 16.56
(f) Kinetic Analysis of HDS and HDN Data. Data from both HDS and HDN studies (Table IV) fitted a pseudo first order reaction model quite satisfactorily (Frost and Cottingham, 1975)
Nf In-=N,
'
6.89 MPa, gas flow rate 890
kN
LHSV+IN
where ks and k~ are the rate constants for pseudo first order HDS and HDN reactions; Sfand Nf and S, and N , are weight percent of sulfur and nitrogen in the feed ( f ) and product oil (p), respectively. Is and INare the factors to account for the reactions that occur due to thermal (noncatalytic) effects only. This effect has been previously reported (Miki et al., 1983;Mann et al., 1987). Similar results were also reported with the alumina-supported commercial catalysts (Mann et al., 1983). Values of kinetic parameters ks and k~ were evaluated from experimental data by plotting In SdS, and In Nf/N, against (LHSV)-l at four different temperatures for all the catalysts. ks and k~ values thus determined at 648 K are given in Table V. The apparent activation energies were obtained from Arrhenius plots of log k vs 1/Tfor the zeolite-supported catalysts. These are given in Table VI. (e) Discussion. It is well documented (Lo Jacono and Schiavello, 1976;Jacobs, 1977;Venuto and Habib, 1979) that gasolinesproduced by zeolite-based catalysts generally have more isoparaffins and aromatics and less olefins than those obtained with silica-alumina alone. However, whether they show higher octane number or not depends upon the specific member of the group paraffins, olefins, naphthenes, and aromatics present and their amounts in the product. The octane number is a function of the feedstock composition and process variables. Results obtained by a variety of methods (de Lasa, 1984;Sinfelt, 1986)indicated that the density of the Bronsted acid sites was much higher for the Y-type zeolites in hydroprocessing than that for amorphous silica-aluminas. The higher selectivity to aromatics and isoparaffins indicates that hydrogen-transfer reactions occur more readily with zeolites as catalysts (Jacobs, 1977). This seems to explain the much better results obtained with the zeolite-supported catalysts. Gasoline and oil products produced over zeolite catalysts are commonly lower in sulfur and nitrogen content and have better storage capability and lead susceptibility. For
1358 Ind. Eng. Chem. Res., Vol. 32, No. 7, 1993
both types of support, the Ni-Mo proved in general to be the best catalyst out of the three studied. The effect of the support on the product composition is of great importance. This factor must be considered in the performance of any hydroprocessing or hydrotreating catalyst. It is known (Kellet et al., 1980) that the most selective catalysts for HDS are those made with Co-Mo compounds and for HDN are those made with Ni-Mo compounds. However, the use of one or the other catalyst will depend very much on the desired product. In our study catalysts based on Ni were found to have higher activity for HDN or HDS than those based on Co-Mo. Zeolite-supported Ni-Mo catalyst was found to be the best for HDS. However,zeolite-supported Ni-W was slightly less active. For HDN zeolite-supported Ni-Mo catalyst gave the best results, although at some flow rates and temperatures its HDN activity was similar to or a little less than the other two zeolite-supported catalysts. It can also be observed that most of the cracking activity of the big molecules and most of the sulfur and nitrogen removal were by virtue of the catalysts and not because of the thermal effects alone.
Conclusions The hydrotreatment of heavy gas oil was investigated over zeolite-silica-alumina-supported and commercial silica-alumina-supported catalysts in a trickle bed reactor at temperatures 573-723 K, LHSV of 1-4, and pressure of 6.89 MPa. The presence of zeolite material in silicaalumina material positively enhanced the hydrotreatment characteristics of the catalysts for a trickle bed operation. They produced liquids which have lower densities, viscosities, and lower amounts of S and N at comparable temperatures. It is alsoworth noting that the metal loading of the catalysts on the zeolite support is about one-fourth of that required in the case of commercial catalysts (see Table 11). A highly efficient catalyst containing 2.24 wt 5% NiO and 5.35 wt % Moos on a zeolite-silica-alumina support (10 wt 5% silica, 25 wt % rare earth exchanged Y-zeolite, and 65 wt % alumina) has been produced. While at a pressure of 6.89 MPa, a temperature of 425 OC, and a LHSV of 2 it could remove 99% S and 80% N, at a LHSV of 1,it removed nearly 99.5% S and 89 % N present in the gas oil. The activity and efficiency of the other two zeolite-supported catalysts (Ni-W, Co-Mo) were comparably higher with respect than that shown by the commercial ones.
Acknowledgment The authors are thankful to National Science and Engineering Research Council of Canada for financial aid (A-1125) and to CANMET. Literature Cited Benesi, H. A.; Bonner, R. V.; Lee, C. F. Determination of Pore Volume of Solid Catalysts. Anal. Chem. 1955,27,1963-1965.
de Bruijn, A. Testing of HDS Catalyst in Small Trickle Phase Reactom. Sixth International Congress on Catalysis; The Chemical Society: London, 1976;pp 951-964. de Lasa, H. Engineering Aspects of Catalytic Cracking. In Zeolites: Science and Technology; Ribeiro, F., Rodrigues, A., Deane, L., Naccache, C., Eds.; NATO Advanced Study on Zeolite-Science and Technology; Martinus Nijhoff: The Hague, 1984;pp 491513. Diaz-Real, R. A. Hydrotreatment of Athabasca Bitumen Derived Heavy Gas Oil over Ni-W and Co-Mo Catalysts. M.A.Sc. Thesis, University of Ottawa, Ottawa, Canada, 1988. Fateh, P. A,; Willingham, C. B. “The Assembly, Calibration and Operation of a Gas Adsorption Apparatus for Measurement of Surface Area, Pore Volume Distribution and Density of Finely Divided Solids”; A Technical Bulletin of the Department of Physical Chemistry, Mellon Institute of Industrial Research Pittsburgh, PA, September 1953. Frost, C. M.; Cottingham, P. L. “Hydrodesulfurizationof Venezuelan Residual Fuel Oils”;Report No. 7557;US.Department of Interior, Bureau of Mines: Washington, DC, 1975. Herskovitz, M.; Smith, J. M. Liquid Distribution in Trickle Bed Reactors. Part 11. Tracer Studies. AIChE J. 1978,24,450-454. Jacobs, P. A. CarboniogenicActivity ofzeolites; Elsevier: New York, 1977. Kellett, T. F.; Trevino, C. A.; Sartor, A. F. Decision Tree Helps Select Hydrotreating Catalyst. Oil Gas J. 1980,78,244-246. Le Jacono, M.; Schiavelto, M. Preparation of Catalysts, Scientific Basis for the Preparation of Heterogeneous Catalysts; Divide la Catalyse, SOC.Chim. Belgique, Elsevier: Amsterdam. 1976;pp 473-488. Mann, R. S.; Sambi, I. S.; Khulbe, K. C. Catalytic Hydrofining of Heavy Gas Oil. Znd. Eng. Chem. Res. 1987,26,410-414. Mann, R. S.;Sambi, I. S.; Khulbe, K. C. Hydrofining of Heavy Gas Oil on Zeolite-AluminaSupported Nickel Molybdenum Catalyst. Znd. Eng. Chem. Res. 1988,27,1788-1792. Mears, D. E.The Role of Axial Dispersion in Trickle Flow Reactors. Chem. Eng. Sci. 1971,26,1361-1366. Miki, Y.; Yamadaya, S.; Oba, M.; Sugimoto, Y. Role of Catalyst in Hydrocracking of Heavy Oil. J. Catal. 1983,83,371-383. Sambi, I. S.Hydrotreatment of Athabasca Bitumen Derived Heavy Gas Oil over Modified ZeoliteSupported Catalysts. Ph.D. Thesis, University of Ottawa, Ottawa, Canada, 1986. Sambi, I. S.; Khulbe, K. C.; Mann, R. S. Catalytic Hydrotreatment of Heavy Gas Oil. Ind. Eng. Chem. Prod. Res. Deu. 1982,21, 575-580. Satrina, M. J., Ed. “Hydroprocessing Catalyst for Heavy Oil and Coal”; Noyes Data Corporation: Park Ridge, NJ, 1982. Scott, J., Ed. Zeolite Technologyand Applications,Recent Advances. Chem. Technol. Reu. 1980,No. 170. Sinfelt, J. H. Influence of Technology on Catalytic Science. Ind. Eng. Chem. Fundam. 1986,25,2-9. Smith, J. M. Chemical Engineering Kinetics, 3rd e& McGraw Hill: New York, 1981. van Klinken, J.; van Dongen, R. H. Catalyst Dilution for Improved Performance of Laboratory Trickle Flow Reactor. Chem. Eng. J. 1980,35,59-66. Venuto, P. B.; Habib, E. T., Jr. Fluid Catalytic CrackingwithZeolite Catalysts; Marcel Dekker: New York, 1979. Vogel, A. Textbook of Quantitative Inorganic Analysis, 4th ed.; Longman: London, 1978. Yang, K. H.; Hougen, 0.A. Determination ofMechanism of Catalyzed Gaseous Reaction. Chem. Eng. Prog. 1950,46, 146. Receiued for review September 15,1992 Revised manuscript receiued April 1, 1993 Accepted April 8, 1993
