940
Ind. Eng. Chem. Res. 1989, 28, 940-947
the Annual American Institute of Chemical Engineers Meeting, New York). Enuiron. Progress. 1989, in press. Jacob, K. T.; Alcock, C. B. Thermodynamics of CuAIOz and CuA1,04 and Phase Equilibria in the System Cu20-Cu0-AlZO3. J. Amer. Ceram. Soc. 1975, 58(5-6), 192-195. Marcilly, C.; Courty, P.; Delmon, B. Preparation of Highly Dispersed Mixed Oxides and Oxide Solid Solutions by Pyrolysis of Amorphous Organic Precursors. J . Am. Ceram. SOC.1970,53(1), 56-57. Marqueen, T. J.; Carbone, D. J.; Ligammari, J. Coal Gasification Combined Cycle Systems-Technical Horizons. Proc. Am. Pouder Con/. 1986, 48, 235-241. Morgantown Energy Technology Center Hot Gas Cleanup Task Force, Chemistry of Hot Gas Cleanup in Coal Gasification and Combustion. Final Report MERC/SP 78/2, Feb 1978. Morimoto, N.; Gyobu, A. The Composition and Stability of Digenite. Am. Mineral. 1971, 56, 1889-1909. Morimoto, N.; Koto, K. Phase Relations of the Cu-S System a t Low Temperatures: Stability of Anilite. Am. Mineral. 1970, 55, 106-117. Morimoto, N.; Kullerud, G. Polymorphism in Digenite. Am. Mineral. 1963, 48, 110-123. Navrotsky, A.; Kleppa, 0. J. Thermodynamics of Formation of Simple Spinels. J . Inorg. Nucl. Chem. 1968, 30, 479-498. Patrick, V.; Gavalas, G. R. Structure and Reduction of Mixed Copper-Aluminum Oxide. J . Am. Ceram. SOC.1989, in press. Paulsson, H.; RosBn, E. A Study of the Formation of CuA120, from CuO and A1,0, by Solid State Reaction a t 1000 "C and 950 O C . Z. .4norg. Allg. Chem. 1973,401, 172-178. Puxley, D. C.; Kitchener, C.; Parkyns, N. D.; Komodromas, C. The Effect of Preparation Methods on the Structure, Stability, and
Metal Support Interactions in Nickel/Alumina Catalysts. In Preparation of Catalysts III; Poncelet, G., Ed.; Elsevier Science Publishers B.V.: Amsterdam, 1983. Robie, R. A,; Hemingway, B. S.; Fisher, J. R. Thermodynamic Properties of Minerals and Related Substances. U S . Geological Survey Bulletin 1452, 1984; US Government Printing Office, Washington, DC. Roseboom, E. H., Jr. An Investigation of the System Cu-S and Some Natural Copper Sulfides Between 25 'C and 700 "C. Econ. Geol. 1966, 61(4), 641-672. Ruth, L. A,; Squires, A. M.; Graff, R. A. Desulfurization of Fuels with Half-Calcined Dolomite: First Kinetic Data. Enuiron. Sei. Technol. 1972, 6(12), 1009-1014. Strohmeier, B. R.; Leyden, D. E.; Field, R. S.; Hercules, D. M. Surface Spectroscopic Characterization of Cu/A1203 Catalysts. :I. Catal. 1985, 94, 514-530. 'Tamhankar, S. S.; Bagajewicz, M.; Gavalas, G. R.; Sharma, P. K.; Flytzani-Stephanopoulos, M. Mixed-Oxide Sorbents for HighTemperature Removal of Hydrogen Sulfide. Ind. Eng. Chem. Process Des. Deu. 1986, 25, 429-437. Thomas, R.; deBeer, V.H. J.; Moulijn, J. A. A Temperature Programmed Reduction Study of r-A120, Supported Molybdenum and Tungsten Oxide. Bull. Soc. Chim. Belg. 1981, 90(12), 1349-1357. Worell, W. L.; Kaplan, H. I. Heterogeneous Kinetics at Elerbated Temperatures; Plenum Press: New York, 1979; p 113.
Received for reuieu J u n e 29, 1988 Revised manuscript receiued February 24, 1989 Accepted March 26, 1989
Influence of Support on the Performance of Coal Liquid Hydrotreating Catalysts Robert L. McCormick, Julia A. King, Todd R. King, and Henry W. Haynes, Jr.* Department of Chemical Engineering, Uniuersity of Wyoming, P.O. Box 3295, University Station, Laramie, W.yoming 82071
A variety of supports for Co- or Ni-promoted molybdenum sulfide were used in hydrotreating coal liquid materials a t high-severity conditions. Both distillate- and residuum-containing feedstocks were employed in runs lasting u p to 500 h. T h e supports studied were alumina, titania, silicated alumina, titania-alumina, magnesia-alumina, chromia-alumina, activated carbon, and nitrided activated carbon. The lined out catalyst activities for both hydrogen uptake and hydrodenitrogenation correlated with either pore volume in 60-200-A-diameter pores or relative number of acid sites per unit mass. However, it was not possible to determine which independent variable had the predominant effect, on activity. T h e silicated alumina catalyst exhibited a markedly reduced coking tendency relative to all other materials studied. It is hypothesized that this is due to the presence of Bronsted acid sites on the surface of this catalyst in the sulfided state. Cobalt sintering was observed for the alumina and nitrided activated carbon supported catalysts. All catalysts exhibited some level of initial deactivation, but activity maintenance was excellent for several of the catalysts despite the high-severity conditions employed. The research described in this paper was undertaken with the goal of developing improved catalysts for hydrotreating coal liquids. In a general sense, this type of catalyst has three potential applications: as a coal liquefaction catalyst, as a catalyst used to produce hydrogendonor solvent for coal liquefaction, and as a catalyst for upgrading coal liquids in an initial refining step. Each of these applications may require a catalyst with somewhat different properties in order to achieve good activity maintenance and an optimum product slate. The conventional catalyst employed in these processes is molybdenum or tungsten sulfide promoted by Co or Ni and supported on a high surface area, porous alumina. * T o whom correspondence should be addressed.
When used in coal liquefaction applications, these catalysts experience rapid coking and a consequent decrease in activity (Thakur and Thomas, 1984; Stohl and Stephens, 1987). Metals deposition (Stanulonis et al., 1976; Stiegel et al., 1982) and active phase sintering (Freeman et al., 1986; Sajkowski et al., 1988) may also contribute to catalyst activity decline. Conventional catalysts catalyze hydrogenation as well as hydrocracking, hydroisomerization, hydrodesulfurization, hydrodenitrogenation, and hydrodeoxygenation (Gates et al., 1979). If the purpose of the hydrotreatment is to produce a low heteroatom refinery feedstock, then all of these reactions are desirable as long as hydrogen consumption is not excessive and light gas yields are low. But in coal liquefaction applications, cat-
O888-5885/89/2628-0940$01.50/0 8 1989 American Chemical Society
Ind. Eng. Chem. Res., Vol. 28, No. 7, 1989 941 alyst and hydrogen usage can be optimized if the catalyst is selective for hydrogenation. In particular, it is desirable to have a catalyst that is highly selective for the production of hydrogen-donor molecules. Other reactions may in fact be undesirable. It has been pointed out that solvent quality decreases with time on stream when conventional catalysts are employed (Ruberto, 1980; Delpuech et al., 1986) because cracking and isomerization reactions lead to degradation of hydrogen-donor molecules. There is also evidence that the presence of heteroatoms, in particular nitrogen (Bertolacini et al., 1979) and oxygen (Kamiya et al., 1978), may improve solvent quality. Thus, the hydrogenolysis reactions responsible for heteroatom removal may, in fact, be undesirable in coal liquefaction. On the other hand, heteroatom removal is desirable in upgrading coal liquids, while excessive hydrogenation of aromatic rings may be undesirable. Process economics could be improved using improved catalysts with better activity maintenance characteristics, leading to lower rates of catalyst addition or increased throughput. Improved catalysts might also allow milder reaction conditions as well as improve the selectivity to desired products and increase the efficiency of hydrogen use. In the work described here, catalyst properties have been varied widely by using different support materials. The results show that activity, coking tendency, and selectivity to the desired products can be significantly altered by changing the support material. Our approach in this research has been to explore a variety of catalyst supports and to determine how these materials affect activity, activity maintenance, and selectivity. It is well-known that molybdenum sulfide interacts strongly with many supports (Nag, 1987; Stampfl et al., 1987) and that this interaction can influence catalytic activity (Massoth et al., 1984). The present study investigates the possibility of exploiting this active phase-support interaction as a tool in catalyst design. The results are presented for coking tendency as well as for hydrodenitrogenation and hydrogenation. In a separate paper (McCormick et al., 1988a), we discuss the factors that influence the hydrogen-donor quality of the product. Experimental Section Catalyst Preparation. The commercial alumina-supported catalysts, Amocat 1A and Amocat lC, were obtained from the Amoco Oil Company. The silicated alumina-supported catalyst, Harshaw CoMo-0402, was supplied by the Harshaw/Filtrol Partnership. Several supports were also obtained commercially. These include titania (Ti-0720) and chromia-alumina (Cr-0103), which were supplied by Harshaw. The activated carbon was Nuchar WV-L, manufactured by Westvaco Chemical Division. A nitrided activated carbon was prepared from this support by heating to 800 "C in flowing ammonia for 4 h. The magnesia-alumina support (1:l Mg0-to-A1203 weight ratio) was prepared by dissolving aluminum and magnesium nitrates in boiling water and then adding ammonia to induce precipitation. A detailed procedure is given elsewhere (King, 1986). The titania-alumina support (62 wt % Ti02)was also prepared by coprecipitation. In this case, aluminum trichloride and titanium tetrachloride were dissolved in water and precipitated by adding ammonia. The precipitate was thoroughly washed to remove chloride. Analysis by Amoco indicated the presence of less than 0.05 w t % chloride. The detailed preparation procedure is provided elsewhere (King, 1987). Properties and calcination temperatures of all supports are listed in Table I.
Table I. Support Properties
support Cr203-A1203 (12 wt % Cr203) MgO-AlZO, (50 wt % MgO) Ti02-A1203 (62 wt % Ti02) carbon nitrided carbon
calcination temp, "C 450
PV,bcm3/g SBET,(I (>60-A m2/g diam) 75 0.27
RADC 0.038
500
140
0.21
0
500
290
0.07
0.075
821 799
0.61 0.22
0 0
SBET is BET surface area. tive acid site density.
PV is pore volume.
RAD is rela-
The loading of active metals onto the supports was accomplished by incipient wetness impregnations. Aqueous solutions of ammonium molybdate stabilized with hydrogen peroxide (Tsigdinos et al., 1981) and of cobalt and nickel nitrates were employed. In most cases, the molybdenum was added and then the catalyst was dried and calcined at 450 "C before adding the promoter. Promoter addition was followed by a second drying and calcining sequence. Carbon-supported catalysts were dried after molybdenum addition, and then promoter was added and the drying step repeated. All catalysts were sulfided with 10% H2S in H2 for 4 h at 450 "C prior to characterization. Optimum metal loadings were selected based on batch reactor screening studies which are reported elsewhere (Haynes, 1988). Catalyst Characterization. BET surface areas were obtained by nitrogen adsorption in a conventional volumetric apparatus. Mercury intrusion was used to measure pore size distributions down to 60-A pore diameter. Selected catalysts were analyzed by X-ray powder diffraction on a commercial instrument using Cu K a radiation. Metals analyses were performed by the Amoco Oil Company using inductively coupled plasma atomic emission spectroscopy. Catalyst relative acid site density (RAD) was measured by the temperature-programmed desorption of tert-butylamine using a method described in the literature (Mieville and Meyers, 1982; Nelson et al., 1983). In this temperature-programmed experiment, the gases desorbing from the catalyst surface are detected by a thermal conductivity detector (TCD). The relationship between the TCD response curve and the actual number of acid sites is unknown. The calculated RAD is the area under the high-temperature acid peak (0peak) divided by the BET surface area. Therefore, the RAD value is "relative" in the sense that it has no meaning for a single sample when viewed alone but it is a measure of acid site density when compared with RAD values for other samples. The apparatus and methods employed have been described previously (Baker et al., 1987). An interpretation of the 0-peak temperature maximum and its relation to acid site strength have been discussed by McCormick et al. (1988b). Used catalysts from the bench-scale trickle bed reactor (described below) were extracted in tetrahydrofuran to remove unreacted feed and liquid product. The catalysts were then dried in a vacuum oven before characterization using the techniques described above. A combustion analysis was performed to determine the coke deposition as weight percent carbon on catalyst. Bench-Scale Hydrotreating. A schematic diagram of the trickle bed reactor is presented in Figure 1. For deactivation studies, a charge of approximately 3 g of catalyst (10-30 mesh), diluted to 50% by volume with quartz chips (10-30 mesh), is employed. Pure hydrogen and liquid feed (described below) are passed over the fixed
942 Ind. Eng. Chem. Res., Vol. 28, No. 7, 1989
RD
A
Proportioning
6 cofscv io PR
~
J Compressor
BPR CO Cd
F FC
FI
Back Pressure Regulator Cutoff Valve Check Valve Filter Flow Controller Flow Indicator
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!/
l l
&
Vent, WTM or Gas Bag
co MV PCO PG PR RD SCV
Metering Valve Pressure Cutoff Switch Pressure Gauge Pressure Regulator Rupture Disk Surge Check Valve
High Pressure Accumulators
Figure 1. Bench-scale h y d r o t r e a t i n g a p p a r a t u s . Table 11. Nominal Conditions for Deactivation Studies P = 13.9 MPa (2000 psig) WHSV = 3 h-' T = 440 "C (825 O F ) H2t r e a t r a t e = 0.8-0.9 L ( S T P ) / c m 3 (4500-5000 SCF/bbl)
bed of catalyst in concurrent downflow. The liquid product is collected in high-pressure accumulators. Gas from the high-pressure accumulators leaves through a back-pressure regulator, passes through the wet test meter, and is collected in butyl rubber gas bags for analysis. The unit is operated continuously, but the balance period data are typically collected over a period of 4 h. Every effort was made to maintain the reactor isothermally. This necessitated the fabrication of a jacket surrounding the reactor which allowed air to be injected to cool hot spots in the bed. With this arrangement, temperature profiles with a 4-6 "C variation were typically observed. An effective isothermal temperature was calculated for each balance period by integrations along the experimental temperature profile. All deactivation studies were conducted at the nominal conditions listed in Table 11. Prior to the start of a run, the catalyst was presulfided according to the following procedure. With the reactor pressure a t 2.7 MPa (400 psig) and a hydrogen flow rate of 2.5 L (STP)/h, a feed of 5 wt % CS2in cyclohexane was passed over the catalyst a t 6 cm3/h. The reactor temperature was then raised from ambient to 340 "C at 120 "C/h and held at 340 "C for 30 min. The temperature was then raised to 440 "C a t 120 "C/h, and this temperature was held for 2 h. After this treatment, the reactor was allowed to cool to ambient temperature under flowing hydrogen and sulfiding feed. The objective of these hydrotreating studies was to provide a basis for comparison of the different catalyst preparations in terms of their relative activity and activity maintenance characteristics. To accomplish this task, good reproducibility and the ability to distinguish between catalysts performing at different activity levels are essential. Experience with this reactor configuration and operating procedure has demonstrated that these require-
ments are met. (Hydrodesulfurization levels approached 100% for many of the catalysts tested; hence, no comparison in terms of HDS activities is attempted.) The reactor is thus an excellent catalyst screening tool. However, it is widely recognized that laboratory trickle bed reactors of the type employed in this investigation cannot be relied upon to provide meaningful kinetics data. This is due primarily to poor contacting efficiencies which arise because of the unavoidably low liquid mass velocities associated with short laboratory reactors (Satterfield, 1975; Koros, 1976). Accordingly, values of rate constants calculated from these data should be considered meaningful only when compared among themselves. While the unreliable nature of any kinetics investigation in emphasized, it was necessary to have kinetics models in order to make corrections for s m d daily deviations from the target temperature. Since the deviations were typically of the order of f2-3 "C, very crude models were sufficient. For this purpose, a brief process variable study was conducted and an analysis of these data was performed along the lines similar to the procedure described by Baker et al. (1987). Details are provided in Haynes (1988). Deactivation run durations varied from 200 to over 500 h on stream. Several researchers have shown that the equilibrium coke level is established in less than 100 h on stream in coal liquid hydrotreating (Crynes and Seapan, 1983; Stohl and Stephens, 1987; Furimsky, 1982a; Cillo et al., 1985). Therefore, coke levels should be roughly comparable for used catalysts from each of these runs even though the times on stream may differ. Feedstock and Product Characterization. Several different coal-derived feedstocks were employed during the course of this investigation. A majority of the deactivation runs employed a mildly hydrogenated creosote oil supplied by Amoco. The preparation of this material, identified as FSN-30, is described by Mahoney et al. (1982). The HCO-1 feedstock was provided by the USDOE/PETC, who had earlier obtained the material from Amoco. It is believed that the origins of HCO-1 and FSN-30 are similar. Deactivation runs were also conducted on a feedstock that
Ind. Eng. Chem. Res., Vol. 28, No. 7, 1989 943 Table 111. Analytical Results for Trickle Bed Reactor Feedstocks FSN-30 HCO-1 FSN-51 HCO1-R' 88.54 91.43 91.51 92.1 wt % c 6.81 6.84 6.83 6.9 wt 70 H 0.20 0.17 0.51 0.17 wt 7 0 s 1.25 0.68 0.68 0.57 wt 70 N 2.36' 0.85b 1.45c 0.26b w t 70 0 wt '70 asphaltene 0 60 12 wt 70 preasphaltene 0 20 4 specific gravity 1.1161 1.1132 1.1232
LEGEND
\
OMixture of 20 wt 70 FSN-51 in HCO-1. bBy difference. 'By analysis.
i
OJ
Table IV. Catalysts Selected for Study in the Continuous Unit designation composition Amocat 1A 3% COO,"16% MoO,"/alumina CoMo/Cr203-A1203 2.8% COO,9.0% MoO,/chromia (12%")-alumina 3.8% COO, 18.9% MoO,/activated C CoMo/AC(HM) 0.9% COO, 4.4% Mo03/activated C CoMo/AC(LM) 3% NiO," 16% Mo03"/alumina Amocat 1C CoMo/NAC 5.8% COO, 16.4% MoO,/nitrided activated C 3% COO: 15% MoO,"/silica (5%")-alumina Harshaw CoMo CoMo/Mg0-A1203 2.1% COO, 15.0% Mo03/magnesia (36.3%)-alumina (37.2%) NiMo/Ti02-A1203 3.1% NiO, 14.1% MoO,/titania (44.7%)-alumina (31.2%) CoMo/Ti02 2.3% COO, 9.6% MoO,/titania ~
~
I
"
'
"
'
~~~~
Nominal values supplied by manufacturer, all other values by analysis.
consisted of a mixture of HCO-1 with 20% by weight coal liquid vacuum bottoms (FSN-51) from run 247 at Wilsonville. The residuum spiked material is identified as HCO-1R. Analytical results for these feedstocks are presented in Table 111. These data indicate that FSN-30 is virtually 100% soluble in cyclohexane. The HCO1-R feedstock contains significant quantities of cyclohexane (asphaltene) and toluene (preasphaltene) insolubles but is quite similar to FSN-30 in elemental composition. Product gases were analyzed by gas chromatography employing an FID and an 80-100-mesh Chromosorb 102 column. Spot checks on a TCD gas chromatograph indicated that no carbon oxides were present. Prior to each determination, the GC was calibrated using a standard gas mixture. Elemental analyses were performed at the Amoco Research Center in Naperville, IL.
Results Based on the results of batch reactor screening studies (Haynes, 1988), 10 catalysts were selected for investigation in the continuous unit. Descriptions of these catalysts are Table V. Properties of Fresh Sulfided Catalysts CoMo/ Amocat Cr2O3CoMo/ 1A Ah03 AC(HM) 571 167 53 BET surface area, m2/g 0.75 total pore vol, cm3/g 0.28 0.25 0.75 pore vol (>60-Adiam), cm3/g 0.12 0.08 0.43 pore vol (60-200-A diam), cm3/g T
Figure 7. Relative deactivation curves for hydrodenitrogenation on Amocat 1A. Comparison of distillate (FSN-30) feedstock versus residuum-spiked feedstock (HCO1-R).
1
1
i
+HCCI-R
-EZENO FSN-33
N i
-
I
L ! i
I
d^
4c:
c c-v
803 2 wTS
GI. PER
1220 3 liT
ISCO C
CATAL~S-
Figure 6. Relative deactivation curves for hydrogen uptake on Amocat 1A. Comparison of distillate (FSN-30) feedstock versus residuum-spiked feedstock (HCO1-R). Table VII. Properties of Used Catalysts from HCO1-R Runs Harshaw Amocat 1A CoMo 750 cum hours on catalyst 219 0.36 pore vol (>60-8, diam), cm3/g 0.15 0.16 pore vol (60-200-A diam), cm3/g 100 av micropore diam, A a 0 4000 av macropore diam, 8, 1000 0.018 RAD, m-2 0.035 19.4 wt % c 9.08 2.6 1.8 wt % H 1.53 0.63 fc
Discussion Because two different feedstocks were employed, the rate constants in Figures 4 and 5 cannot be directly compared to rate constants in Figures 2 and 3. To aid in the interpretation of these results, a relative rate constant was calculated. The relative rate constant, defined as the rate constant at any time on stream divided by the rate constant on day 1 is plotted in Figures 6 (hydrogenation) and 7 (hydrodenitrogenation) for the Amocat 1A deactivation runs with the distillate and residuum feedstocks. It is apparent that the 20% residuum material produces no significant enhancement in the rate of deactivation for hydrogenation. This is probably because of the relatively low level of benzene (toluene) insolubles present in the HCO1-R feedstock. The residuum material may be responsible for a slight enhancement of the deactivation rate for hydrodenitrogenation, however.
While the effects on activity maintenance are minimal, it is apparent from a comparison of data in Tables VI and VI1 that increased coke laydown has occurred in runs using the residuum-containing feed material. The spent Amocat 1A catalyst from the FSN-30 run contained 11.0 wt % C compared to 19.4 wt '70C on the spent catalyst from the HCO1-R run. Correspondingly, the spent CoMo-0402 catalyst from the FSN-30 run contained only 6.1 wt % C compared to 9.1 wt 70 C on the used catayst from the HCO1-R run. It might be argued that the higher coke level on the Amocat 1A catalyst of Table VI1 is due to the longer run duration (750 h) by comparison to the coke level on the Amocat 1A catalyst of Table VI (507 h). However, it was pointed out earlier that the equilibrium coke level in coal liquid hydrotreating is usually established in less than 100 h on stream. In support of this contention, the used Amocat 1A catalyst from a recently completed 245-h duration run with the HCO1-R feedstock analyzed 17.2 w t 70C (Papaioannou, 1988), in reasonably good agreement with the coke loading observed in the much longer duration run of Table VII. Freeman and co-workers (1986) observed Co sintering in catalysts from the H-Coal process. These researchers noted that the growth of cobalt sulfide crystals paralleled catalyst activity loss, and they concluded that a significant fraction of the long-term catalyst activity loss must be due to cobalt sintering. Sajkowski and co-workers (1988) observed Co sintering in coal liquefaction catalysts. They concluded that Co sintering was probably not the dominant mode of deactivation. In this work, cobalt sintering was observed by XRD for Amocat 1A and nitrided activated carbon supported catalysts. We have made no attempt to quantify the loss of Co dispersion. However, a support effect is evident in the sense that Co sintering was not observed for many of the catalysts studied. This indicates either the absence of sintering or a slower rate of sintering in these catalysts. Cobalt sulfide crystallites must grow to be larger than about 50 8,to be detected by XRD. From the deactivation curves of Figures 2 and 3, it is evident that most of the catalyst has reached a plateau in activity at about 500 cumulative weights of oil per weight of catalyst. To reduce day-to-day fluctuations, "lined out" activities were calculated for each run by taking the average of the balance period closest to 500 cumulative weights of oil per weight of catalyst with balance periods on either side. These results are compiled in Table VIII. Several correlations between lined out activities and measurable properties of the catalyst, such as molybdenum loading and surface area, were explored. Workers at
946 Ind. Eng. Chem. Res., Vol. 28, No. 7, 1989 N
Table VIII. “Lined Out” Activities (FSN-30 Feedstock) k N , h-’ catalyst k H , h-’ Amocat 1A 2.21 3.14 CoMo/Crz03-A1203 0.86 1.34 1.29 1.80 CoMo/AC(HM) 0.69 0.94 CoMo/AC(LM) 2.27 3.49 Amocat 1C 1.52 CoMo/NAC 1.18 1.84 2.87 Harshaw CoMo 0.72 1.04 CoMo/MgO-A1203 2.07 3.74 NiMo/TiOz-Al20, 1.10 0.87 CoMo/Ti02
i
I
31
Amoco (Bertolacini et al., 1979; Kim et al., 1979) found that micropore volume in a critical pore size range was extremely important in coal liquefaction catalyst activity. In particular, their results indicated that the micropore volume should be narrowly distributed about 120-A-diameter pores in order to prevent plugging of the high surface area regions of the catalyst by coke. Relative acid site density (RAD) is another potential correlating variable. Because the rate constants in Table VI11 are not on a surface area basis, the RAD values were converted to a mass basis. The product of RAD and BET surface area is a measure of the relative acidity per gram. Bivariate correlation coefficients between the lined out activities and each of these potential correlating variables were calculated, and the results are presented in Table IX. The critical value of the correlation coefficient is 0.632 at the 95% confidence level and 0.765 at the 99% confidence level. Based on the values in Table IX, it appears that activities for both hydrogenation and hydrodenitrogenation are highly correlated (>99% confidence) with the pore volume in the preferred range (60-200-A diameter) and with the relative acidity per unit mass. The correlation could not be improved significantly by a multiple linear regression on both independent variables because, as it turns out, the two independent variables are correlated (rI2 = 0.8131. On cursory examination, it appears surprising that a physical property of the catalyst, pore volume in the preferred range, should correlate with a surface property, number of acid sites. A possible explanation is that the tert-butylamine probe molecule may be sterically hindered from contacting acid sites in small pores. As discussed previously, measurements on the carbon-supported catalysts indicate that molecular size pores are present. Acid sites in molecular sieves would not be expected to contribute to the acidity determination, and perhaps this is why the RAD values reported for the carbon-supported catalysts are low, Table V. On the other hand, we might expect acidity to play a significant role in the hydrodenitrogenation reaction since many nitrogen-containing compounds found in coal liquids are organic bases. Unfortunately, the data do not allow a preference to be made between these two correlating variables. In order to compare the coking tendencies of the various catalysts, we have found it convenient to calculate. the fraction of a monolayer coverage, f,, assuming that the coke is deposited as graphite. This is obviously an oversimplification, but it does serve to quantify coke deposition on a unit surface area basis. Values off, are compiled in
I
i
’
I H a r s h o w CoMo-0402
11
Table IX. Bivariate Correlation Coefficients”
kH kN
a t 500 wt oil/wt catalyst a t 500 wt oil/wt catalyst
wt % Moo3 0.532 0.494
SBET m2/g -0.060 -0.227
lreritl= 0.632 a t 95% confidence level; lrcritl= 0.765 a t 99% confidence level.
P V in 60-200-A-diam pores, cm3/g 0.815 0.836
RAD, m-2 0.386 0.583
RADSBET 0.843 0.928
Ind. Eng. Chem. Res., Vol. 28, No. 7, 1989 947 for a specific application. It also seems likely that the support properties can be used to produce catalysts with resistance to coke formation and active-phase sintering.
Conclusions Of the variety of catalysts tested, three commercial catalysts (Amocat 1 A and 1C and Harshaw CoMo-0402) and one laboratory preparation (NiMo/ titania-alumina) exhibited excellent activity and activity maintenance characteristics despite the high-severity conditions employed. Spiking the feedstock with 20 wt% coal-derived residuum did not increase the deactivation rate significantly, but higher coke levels were observed on the used catalysts. Amocat 1A and a nitrided activated carbon supported catalyst exhibited cobalt sintering to form Co&+ It was observed that lined out activities could be correlated with either pore volume in the preferred pore range (60200-A diameter) or with the relative acidity per unit mass. A silicated alumina-supported catalyst (Harshaw CoMo0402) exhibited a reduced coking tendency relative to the other catalysts studied. It is speculated that this may be due to the presence of Bronsted acid sites known to be associated with this type of support. The results indicate that the choice of catalyst support can have a significant influence on catalyst properties and on catalyst activity, activity maintenance, and selectivity in coal liquids hydrotreating. Acknowledgment This work was jointly sponsored by the US Department of Energy (Grant DE-FG22-84PC70812) and the Amoco Oil Company. We are grateful for both the financial support and the helpful consultations provided by these organizations. Registry No. Cr203, 1308-38-9; MgO, 1309-48-4; TiOz, 13463-67-7; C, 7440-44-0; Co, 7440-48-4; Ni, 7440-02-0; Mo, 7439-98-7.
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