I n d . Eng. Chem. Res. 1995,34, 1566-1571
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Large-Pore Catalysts for Hydroprocessing of Residual Oils Zhong-Shu Ying, B64e Gevert, and Jan-Erik Otterstedt Department of Engineering Chemistry, Chalmers University of Technology, 412 96 Goteborg, Sweden
Johan Sterte* Department of Chemical Technology, Luleii University of Technology, 97187 LulePL, Sweden
Hydroprocessing catalysts were prepared using fibrillar alumina and attapulgite as carrier materials and nickel and molybdenum as active substances. With fibrillar alumina as primary particles, a carrier material was obtained which combined a large surface area (138 mz/g) with a large average pore diameter (252 A). Due to side-by-side association of the fibers, attapulgite did not yield a catalyst carrier with as large pores as was expected. The catalysts were tested for hydroprocessing of a n atmospheric petroleum resid, and the results were compared with those obtained for a commercial catalyst with similar loading of nickel and molybdenum. The catalyst prepared using fibrillar alumina as carrier material was more active than the commercial catalyst for hydrodemetalization, equally active for hydrodesulfurization, and less active for hydrodenitrogenation. The catalyst prepared using attapulgite as carrier material was inferior to the other two catalysts in all respects. A continuously increased consumption of light petroleum products, a slowly declining average quality of the available crude oils, and an increased environmental consciousness are augmenting the need for hydroprocessing of larger quantities of different petroleum fractions, particularly heavy fractions. Heavy oil fractions are characterized by high contents of metal poisons, heteroatoms, and asphaltenes which make them difficult to upgrade, from both a process and an environmental point of view. Metals such as vanadium and nickel accumulate on catalyst surfaces and cause permanent deactivation of the catalysts. In order to protect and extend the lifetime of catalysts used for hydrodesulfurization (HDS)and catalytic cracking (FCC), metals can be removed upstream of these processes by catalytic hydrodemetalization (HDM). The metals are, t o a large extent, bound in porphyrin structures associated to the asphaltene fraction of the oil. Asphaltenes have formula weights in the range 5000-10000 daltons and are believed to be micelle clusters with the individual molecules arranged in graphite-like layers by means of interlayer forces. As hydrotreatment reactions generally are diffusion controlled, the pore size and the pore size distribution of a catalyst designed for HDM is of major importance. Richardson and Alley (1975) showed that asphaltenes in oil at 200-600 O F are excluded by catalysts with pore diameters smaller than 70 A. The pore size and the pore size distribution are also of major importance for the long-term activity of the HDM catalyst. Deactivation of HDM catalysts is primarily caused by deposition of vanadium sulfides at the pore mouths, eventually causing pore-mouth plugging. This deactivation process is slower on large-pore catalysts in comparison with small-pore ones. Thus, largepore catalysts can maintain an acceptable level of HDM activity even at relatively high metal loadings (Quann et al., 1988). Conventional hydrotreatment of oils, e.g., for removal of sulfur from gas oil fractions, uses catalysts with carriers of y-aluminum oxide having a relatively narrow
* Address correspondence t o this author. FAX: +4692091199. E-mail:
[email protected].
pore radius distribution in the range 40-70 8, and surface areas of 200-300 m2/g. The active phase on these catalysts consist of metal sulfides such as molybdenum disulfide and cobalt sulfide. These catalysts are excellent for HDS of relatively light fractions but less effective and rapidly deactivated when used for heavier oils. Optimal pore sizes for maximum HDM activity have been suggested by many investigators (Quann et al., 1988). Hardin et al. (1981) found a maximum activity for Ni and V removal from Athabasca bitumen with catalysts characterized by median pore diameters of 200 A. Increases in pore size above this resulted in a decrease in activity. This phenomenon is easily rationalized on the basis of catalyst surface area. As the pore size increases, the surface area decreases and so does the number of active sites. There is a compromise between the increase in activity due t o an increase in surface area and the decrease due to diffusional resistances; hence an optimum in pore size exists. This correlation between surface area and pore size assumes that the primary particles building up the catalysts have the same geometrical shape and are packed in the same manner. If the shape or the packing of the primary particles is altered, a new correlation between surface area and pore size is obtained. The primary particles in conventional HDS and HDM catalysts are generally corpuscular. A more favorable correlation between surface area and pore size, Le., catalyst carriers having large surface areas in combination with relatively large pores, can be prepared by using materials with rodlike primary particles for building up the structure instead of spherical or corpuscular ones. One such material is attapulgite, which is a crystalline magnesium-aluminum silicate in which the primary particles are fibers with a length of 5002000 nm and a diameter of 50-100 nm. HDM catalysts prepared using acid-leached attapulgite in combination with corpuscular alumina as a carrier material have been structurally characterized by Tokarz (1988), and their resistance to deactivation has been investigated by Aktius (1987). By using proper synthesis conditions and reactant mixtures, aluminum oxide hydroxide of the boehmite
0888-5885/95/2634-1566$09.QQ~Q 0 1995 American Chemical Society
Ind. Eng. Chem. Res., Vol. 34, No. 5, 1995 1567 Table l. Characterization of Catalysts catalysta surface area (m2/g) pore volumed (cmVg) average pore diame (A) bulk density (g/cm3) NiO (wt %) M O z 0 3 (wt %) P205f (wt %)
Ab 138 0.87 252 0.61 3.5 18.1
Bb 72 0.44 244 0.57 3.2 18.2
C" 158 0.39 99
0.75 3.6 18.8 5.7
Catalyst A fibrillar alumina as carrier material. Catalyst B: attapulgite as carrier. Catalyst C: commercial reference catalyst. Cylindrical pellets with d = 1 mm. Trilobed pellets with d = 1.5 mm. Estimated as the liquid volume adsorbed a t a relative pressure of 0.995. e Calculated assuming cylindrical pores. f Value provided by catalyst manufacturer.
type can be crystallized from basic aluminum salt solutions in the form of fibrillar particles. By varying the synthesis conditions, particularly temperature and time, the size of these fibrils can be varied in a relatively wide range (Sterte and Otterstedt, 1986). Upon calcination, fibrillar boehmite transforms to y-aluminum oxide maintaining the fibrillar shape of the primary particles. The present paper reports a preliminary study on the HDM performance of catalysts prepared using attapulgite and fibrillar alumina as carrier materials. The HDM activity and selectivity is evaluated and compared with that of a commercial hydrotreatment catalyst containing corpuscular alumina as carrier material.
Experimental Section Hydroprocessing Catalysts. Fibrillar boehmite was prepared by treating an aluminum chlorohydrate solution in a Teflon-coated stainless steel autoclave a t 160 "C for 24 h (Sterte and Otterstedt, 1986). The aluminum chlorohydrate solution (Locron L, Hoechst) had an OWAl molar ratio of 2.5 and was used at a concentration of 0.8 M (Al). The attapulgite used for preparation of one of the catalyst carriers was Attagel 50 from Engelhard Co. The attapulgite was leached with six parts of HCI(9 w t %) per part of dry attapulgite under vigorous stirring a t 60 "C for 2 h. The attapulgite was then washed repeatedly with distilled water to remove excess HC1. Catalyst carriers were prepared by first mixing the carrier materials with water to a suitable viscosity and then extruding in a radial extruder with a 1mm sieve. Aluminum chlorohydrate corresponding to 3 wt % aluminum oxide in the dry carrier was added to the attapulgite carrier as a binder. After extrudation the catalyst supports were dried at 110 "C overnight and then calcined at 510 "C for 3 h. The catalyst carriers were impregnated by pore volume filling first with an aqueous solution of ammonium heptamolybdate and then with one of nickel nitrate. Between and after these impregnations the samples were dried a t 110 "C for 12 h and calcined at 480 "C for 3 h. The catalysts produced from the fibrillar boehmite carrier and from the attapulgite carrier are designated catalyst A and catalyst B, respectively. A commercial catalyst, Shell 424 (catalyst C) in the form of 1.5 mm trilobed pellets, was used as a reference catalyst in this study. This choice of reference catalyst was motivated by the fact that Shell 424 has a specific surface area comparable to that of catalyst A. Properties of the three catalysts are given in Table 1.
Prior to testing, the catalysts were sulfided with a mixture of 10 vol % hydrogen sulfide in hydrogen a t 400 "C and a total pressure of 0.4 MPa for 5 h. The amount of hydrogen sulfide used was in 3-fold excess of that theoretically required for complete sulfidation. Hydroprocessing. The hydrotreatment process used in this study was described in detail in an earlier publication (Haglund et al., 1991). The hydroprocessing experiments were performed in a down-flow trickle bed reactor with an inner diameter of 24 mm and a height of 565 mm. The reactor temperature was maintained by three individually controlled heating coils making it possible to keep the temperature difference over the catalyst bed within 2 "C. The catalysts were loaded in the middle of the reactor, and the void space above and below the catalyst was filled with glass beads. A constant catalyst volume of 31 mL was used in all experiments. The hydrogen flow was controlled by a thermal flow controller calibrated for 100% hydrogen, whereas the feed oil was fed by a displacement pump which was controlled by the stroke length of the pump. The reactor was designed for a maximum pressure of 200 bar and a maximum temperature of 450 "C. A hydrogen pressure of 100 bar and a hydrogen to oil ( W O ) ratio of 350 m3(NTP)/m3were used in all runs. Variables in the hydroprocessing experiments were reactor temperature and liquid hourly space velocity (LHSV). Reactor temperatures in the range 370-430 "C and LHSVs in the range 0.5-3.0 h-l were used. The times required in order to reach steady state after changing reaction conditions were 9, 6, 3, and 1.5 h a t LHSVs of 0.5,0.75,1.5, and 3.0 h-l, respectively. These times were experimentally determined in a separate run. The catalysts were deactivated with the feed oil at a LHSV of 1.0 h-l until the HDS conversion reached a steady value (2-3 days). The duration of an experiment, using one catalyst at different process conditions, was 8 days. The deactivation, after the initial deactivation period, was found t o be negligible. The feed oil used in all hydroprocessing experiments was an atmospheric residue from Scanraff AB (Lysekil, Sweden) containing a mixture originating from Oman, Mexico, and Sovjet export blend. Characterization of Catalysts. Nitrogen adsorption-desorption isotherms were determined using a Digisorb 2600 surface-area pore-volume analyzer (Micromeritics Instrument Corporation). The samples were first outgassed at 200 "C for 3 h, and the isotherms were recorded at liquid nitrogen temperature. Surface areas were calculated using the BET equation, and pore volumes were estimated to be the liquid volume adsorbed at a relative pressure of 0.995. Pore-size distributions were calculated from the desorption branches of the isotherms assuming cylindrical pores. Elemental analysis of the catalysts was carried out by atomic absorption spectroscopy (AAS)employing LiB02 fusion (Medlin et al., 1969). Scanning electron microscopy (SEMI was performed using a JEOL Model JSM-5200 electron microscope. Characterization of Feedstock and Hydroprocessed Products. Boiling point distributions were analyzed by simulated distillation using a HP 5880 gas chromatograph equipped with a 0.7 m OV-101 column. The principle of this method is described in ASTM D 2887-73. The amounts of carbon, nitrogen, and hydrogen were determined using a LECO CHN-600 elementary ana-
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lyzer. Sulfur was determined with a LECO SC-132 sulfur analyzer. Rasic nitrogen was determined by titration with perchloric acid using crystal violet as an indicator. The metal contents in the feed oils were analyzed spectrophotometrically. Oil samples were first mixed with sulfuric acid and sulfonated by treatment at 120-160 "C on a hot plate (see ASTM D 1548-83).The sulfonated samples were placed in a muffle furnace with a temperature of 150 "C and the temperature was slowly increased to 525 "C. A slow flow of air was passed through the furnace to aid in reducing the coke t o inorganic ash. After cooling, the ash was dissolved in a mixture of sulhric and nitric acid. Nickel was analyzed, after dilution, by AAS. Part of the solution was further treated according to ASTM D 1548-83 in order t o convert the vanadium to its phosphotungstovanadic acid complex which was analyzed spectrophotometrically. Asphaltene contents were determined as the fraction of the oil which was insoluble in n-hexane.
Results and Discussion Characterization of Catalysts and Feedstocks. Table 1 shows the surface area, pore volume, and average pore diameter, calculated assuming cylindrical pores, for the three catalysts studied. In Figure 1, the desorption pore-size distributions for the different catalysts are shown. The fibrillar alumina catalyst, catalyst A, has a surface area of 138 m2/g, slightly lower than that of the reference catalyst, and a pore volume of 0.87 cm3/g,more than twice that of catal st C. The average pore diameter of catalyst A is 252 and the pore-size distribution is relatively narrow with a tail toward smaller pores. The attapulgite catalyst, catalyst B, has a pore volume of 0.44cm3/g, i.e., similar to that of the reference catalyst, and a surface area of only 72 m2/g. The average pore diameter of catalyst B is 244 A, and the pore-size distribution is very broad. The reference catalyst, catalyst C,has a surface area of 158 m2/g,a pore volume of 0.39 cm3/g,and an average pore diameter of 99 A. The pore-size distribution determined for this catalyst is narrow with a majority of the pores in the range 60-120 A. Figure 2 shows scanning electron micrographs of the three catalysts taken at a magnification of 50000x The
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primary particles in catalyst A are straight fibrils with an approximate length of 600 nm while those of catalyst B are somewhat curved attapulgite fibers with an average length of about 1 pm. As expected, the primary particles of catalyst C are corpuscular with a size of 2050 nm. Above it was reasoned that it should be possible t o prepare catalysts which combine a large surface area with large pores by using rodlike primary particles. A further requirement, in addition to the shape, for such a preparation to be successful is that the particles are essentially randomly packed in the final catalyst. From the scanning electron micrograph of catalyst A (Figure 2a) it is clear that the alumina fibrils building up this material are more or less randomly packed and thus fulfill this requirement. This is, however, not the case for catalyst B in which the attapulgite fibers to a large extent are aggregated in a side-by-side manner rather than randomly packed (Figure 2b). Moreover, scanning electron micrographs taken at lower magnifications (not shown here) show that the attapulgite contains flakelike particles, probably of other clay minerals, as a contamination. These two factors most likely explain why catalyst B does not possess the large pore volume and average pore size that would be expected considering the shape of the primary particles. The NiO and Moz03 contents are similar for the three catalysts (see Table 11, 3.4 f 0.2 and 18.4 f 0.4 w t %, respectively. The effects of differences in metals loading of the catalysts should therefore be negligible when comparing catalysts prepared using the different carriers. The feed oil used in all the hydroprocessing experiments (see the Experimental Section)was analyzed with respect to elemental composition and boiling point distribution. The results of these analyses are given in Table 2. Hydroprocessing. Table 3 shows the simulated distillation results for samples of oil hydroprocessed over the three catalysts a t temperatures of 370 and 430 "C and LHSV's of 0.5 and 3.0 h-l. The simulated distillation characteristics for all samples treated at the lower temperature are almost identical to those of the feed oil (see Table 2), indicating that the extent of hydrocracking is very limited a t this reaction temperature
Ind. Eng. Chem. Res., Vol. 34,No. 5,1995 1569 Table 2. Characterization of Atmospheric Resid Used as Feed Oil elemental analysis (wt %) C H N basic N S WC metals Cppm) V Ni simulated distillation (wt %) 360 "C density (g/cm3)
83.1 10.9 0.43 0.16 2.1 1.6 44 23
0 5 95 0.925
Table 3. Simulated Distillation of Hykmprocessed Oil (wt %) for LHSV of 0.5 and 3.0 h-l catalyst B catalyst C fraction catalyst A reactor 3.0 0.5 3.0 0.5 3.0 0.5 ("C) temp("C) 370 360 91 92 94 96 96 96 430 360 14 73 76 84 21 83
Figure 2. Scanning electron micrographs of (a, top) catalyst A, (b, middle) catalyst B, and (c, bottom) catalyst C, taken at a magnification of 50000x . (Reproducedat 74% of the original size.)
independent of LHSV. At a processing temperature of 430 "C, on the other hand, a significant reduction in molecular weight is taking place particularly at the lower LHSV. At this temperature and a LHSV of 0.5 h-l, catalyst A converts 63 wt % of the resid to compounds boiling below 215 "C. The corresponding values for catalysts C and B are 54 and 15 wt %, respectively, indicating that the order of activity of the three catalysts for hydrocracking is catalyst A > catalyst C >> catalyst B.
Figure 3 shows the removal of vanadium when hydroprocessing the feed oil over the three catalysts at temperatures in the range 370-430 "C and using liquid hourly space velocities of 0.5-3.0 h-l. Independent of reaction conditions, catalyst A is the most effective and catalyst B the least effective catalyst for removal of V whereas the reference catalyst, catalyst C, shows an intermediate activity. At high reactor temperatures and low LHSV's, the differences between the catalysts are less pronounced than at lower temperatures and higher LHSV's. At the highest temperature (430"C)and the lowest LHSV (0.5 h-l) the three catalysts show approximately the same vanadium removal. In a separate study we investigated the thermal decomposition of vanadium ethylporphyrin and found that this reaction starts at about 370 "C and proceeds at an appreciable rate at temperatures above 400 "C. Using a reactor pressure of 5 MPa, feeding 16 cm3 h-l of an oil containing 40 ppm vanadium ethylporphyrin, and substituting the catalyst with glass beads, a thermal conversion of the porphyrin of 18% at 370 "C and 62% at 420 "C was measured. The value of this experiment is limited as the vanadium in the feed oil is bound in a variety of porphyrin as well as non-porphyrin structures. It does, however, indicate that at least part of the remarkably high activity, particularly of catalyst B, at high temperatures and long residence times can be explained by thermal rather than catalytic reactions. The high activity of the fibrillar alumina catalyst (A) is believed to be a consequence of a relatively high surface area in pores large enough to eliminate the diffusional constraints encountered in catalysts with smaller pores, such as catalyst C. Figure 4 shows the removal of nickel when hydroprocessing the feed oil over the three catalysts at temperatures in the range 370-430 "C and using LHSV's of 0.5-3.0 h-l. The trends observed for nickel removal are similar to those observed for vanadium removal: catalyst A is the most active and catalyst B the least active catalyst. This was expected as nickel is bound in the same type of structures in the feed oil as vanadium.
1670 Ind. Eng. Chem. Res., Vol. 34,No. 5, 1995
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Figure 5 shows the removal of sulfur when hydroprocessing the feed oil over the three catalysts at temperatures in the range 370-430 "C and using LHSVs of 0.5-3.0 h-l. The sulfur reduction over catalyst B is significantly poorer than that over catalysts A and C. The latter two catalysts show essentially the same activity for sulfur removal at all the experimental conditions studied. This may seem somewhat surprising considering the differences in activity for hydrodemetalization between the two catalysts. Sulfur is, however, more evenly distributed in the different boiling fractions of the resid in comparison with the metals which are primarily associated with the asphaltene
fraction. Furthermore, the hydroprocessing experiments were performed using a constant catalyst loading on a volume basis: 31 cm3 of catalyst was used in all experiments. A consequence of the lower bulk density of catalyst A compared with catalyst C is that 32%more of catalyst C was used in the experiments on a weight basis. A considerably larger catalyst area, about 50% more, was thus used in the experiments with catalyst C . The physical shape of catalyst C, 1.5 mm trilobed pellets is, moreover, advantageous in comparison with that of catalyst A, 1 mm cylindrical pellets, from a diffusion point of view as the external surface area is larger and the average diffusion distance is shorter in
Ind. Eng. Chem. Res., Vol. 34, No. 5, 1995 1571 100
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the former catalyst. Taken together, these factors may explain the relatively high activity of catalyst C for HDS. Figure 6 shows the removal of basic nitrogen over the three catalysts when hydroprocessing the feed oil at temperatures in the range 370-430 "C at LHSVs of 0.5-3.0 h-l. The removal of basic nitrogen is considerably better over catalyst C than over catalysts A and B. This may partly be explained in the same manner as the relatively high HDS activity of catalyst C, i.e., a more even distribution of nitrogen between the fractions of the oil in combination with a higher catalyst loading of catalyst C on a weight basis. Another factor that most likely contributes to the higher activity of catalyst C for removal of basic nitrogen is that this catalyst , has been reported to increase the contains P z O ~which HDN activity of hydroprocessing catalysts (Eijsbouts et al., 1991).
Conclusions Catalyst carriers combining large surface areas with large average pore sizes can be prepared by using rodlike primary particles in the preparation of the carriers provided that these particles pack in a random manner. This appears to be the case for fibrillar alumina prepared by autoclave treatment of aluminum chlorohydrate but not for the clay mineral attapulgite. The attapulgite fibers tend to associate in a side-by-side fashion to a much larger extent than the alumina fibrils. For HDM of a resid petroleum oil, a catalyst prepared using fibrillar alumina as carrier material was considerably more active than a commercial catalyst with similar surface area and loading of cobalt and molybdenum. For HDS the activities of the two catalysts were similar whereas the activity for HDN was higher for the commercial catalyst. A catalyst prepared using attapulgite, a natural clay mineral, as carrier material was inferior t o both the catalyst prepared from alumina and the commercial catalyst in all respects. Encouraged by these results we are currently investigating properties and catalytic performance of a series of catalysts prepared using different fibrillar aluminas as carrier materials. The size and shape of the fibrils in fibrillar boehmite used as a precursor for preparation of fibrillar alumina are varied within relatively wide ranges by varying the conditions in the hydrothermal synthesis. We are also studying the long-term deactiva-
tion by metal poisoning on hydroprocessing catalysts prepared from fibrillar alumina.
Acknowledgment The authors wish to thank the Swedish National Board for Industrial and Technical Development (NUTEK) and the OK Petroleum Environmental Foundation for financial support of this project.
Literature Cited Aktius, L. A New Improved HDM Catalyst for Heavy Crude Oils. International Conference on Heavy Crude Tar Sands, Third; UNITARKJNDP Inf. Cent. Heavy Crude Tar Sands: New York, 1985; Vol. 4, pp 1868-82. Eijsbouts, S.; van Gestel, J. N. M.; van Veen, J . A. R.; de Beer, V. H. J.; Prins, R. The Effect of Phosphate on the Hydrogenation Activity and Selectivity of Alumina-Supported Sulfided Mo, Ni, and Ni-Mo Catalysts. J. Catal. 1991,131,412. Haglund, R.; Otterstedt, J.-E.; Sterte, J. Upgrading of Hydropyrolysis Coal Tar by Hydroprocessing. Erdoel-Erdgas-Kohle 1991, 107,232. Hardin, A. H.; Ternan, M.; Packwood, R. H. The Effects of Pore Size in M003-C00-&03 Hydroprocessing Catalysts, CANMET Report 81-4E, Energy, Mines and Resources, Canada, 1981. Medlin, J. H.; Suhr, N. H.; Bodkin, J. B. Atomic Absorption Analysis of Silicates Employing LiBOz Fusion. At. Absorpt. Newsl. 1969,8, 25. Quann, R. J.; Ware, R. A.; Hung, Chi-Wen; Wei, J . Catalytic Hydrodemetallation of Petroleum. Adu. Chem. Eng. 1988,14, 95. Richardson, R. R.; Alley, S. K. Consideration of Catalyst Pore Structure and Asphaltenic Sulfur in the Desulfurization of Resids. In Hydrocracking and Hydrotreating; ACS Symposium Series 20; American Chemical Society: Washington, DC, 1975; p 136. Sterte, J.; Otterstedt, J.-E. A Study on the Preparation and Properties of Fibrillar Boehmite. Mater. Res. Bull. 1986,21, 1159. Tokarz, M. A Clay-Based Support with High Porosity for HDMCatalyst. In Studies in Surface Science and Catalysis; Unger, K. K., et al., Eds.; Elsevier Science Publishers: Amsterdam, 1988; Vol. 39, pp 559-569.
Received for review September 12, 1994 Accepted March 6, 1995 * 139405394 ~~~~~~~
* Abstract published 1995.
in Advance A C S Abstracts, April 15,
