Ind. Eng. Chem. Res. 1988,27, 1788-1792
1788
pyrolysis oils derived from wood are heated to the temperature levels necessary for HDO. This treatment involves the use of a supported ruthenium catalyst. I t was shown to be quite effective, as it permitted high HDO conversions to be reached for vacuum pyrolysis oils which, contrary to the other wood-derived liquids, are deemed to be impossible to hydrotreat. Registry No. Ru, 7440-18-8; Co, 7440-48-4; W, 7440-33-7.
Literature Cited Badilla-Olhbaum, R.; Pratt, K. C.; Trimm, D. L. “A Study of Nickel-Molybdate Coal-Hydrogenation Catalysts Using Model Feedstocks”. Fuel 1979, 58, 309-314. Baker, E. G.; Elliot, D. C. “Catalytic Hydrotreating of Biomass-Derived Oils”. Prepr. Pup.-Am. Chem. SOC.,Diu. Fuel Chem. 1987, 2, 257-263. Bizhanov, F. B.; Drozdova, R. B. “Studies of the Kinetics and Mechanism of Glucose Hydrogenation with Ruthenium Catalyst”. React. Kinet. Catal. Lett. 1982, 1-2, 35-39. Bredenberg, J.; Huuska, M.; Raty, J.; Korpio, M. ”Hydrogenolysis and Hydrocracking of the Carbon-Oxygen Bond”. J . Catal. 1982, 77, 242-247. Elloit, D. C. “Hydrodeoxygenation of Phenolic Components of Wood-Derived Oil”. Symposium on Heavy Oil and Residua Proceeding, Seattle, March 20-25 1983a. Elliot, D. C. Final Report, Vol. 4, IEA Co-operative Project D1, Biomass Liquefaction Project, 198313. Furimsky, E. “Catalytic Deoxygenation of Heavy Gas Oil”. Fuel 1978, Aug, 494-496. Furimsky, E. “Catalytic Removal of Sulfur Nitrogen and Oxygen from Heavy Gas Oil”. AZChE J. 1979, 2, 306-311. Furimsky, E. “Chemistry of Catalytic Hydrogenation”. Catal. Rev. Sci. Eng. 1983a, 3, 421-458. Furimsky, E. “The Mechanism of Catalytic Hydrogenation of Furan”. Appl. Catal. 198313, 6, 159-164. Furimsky, E. “Mechanism of Catalytic Hydrodeoxygenation of Tetrahydrofuran”. Znd. Eng. Chem. Prod. Res. Deu. 1 9 8 3 ~1, 31-34. Gagnon, J. “Hydrotraitement catalytique des huiles pyrolytiques du bois”. Master Thesis, Universit6 Laval, QuBbec, Canada, 1987. Johnson, D. K.; Ratcliff, M.; Black, S.; Posey, F.; Chum, H. L.; Gabeen, D. W.; Cowley, s.;Baldwin, R. “Liquids Fuels from Lignins”. Proceedings of the Biochemical Conversion Program Review Meeting, Golden, CO, 1986; pp 261-281.
Krishnamurthy, S.; Panvelker, S.; Shah, Y. T. “Hydrodeoxygenation of Dibenzofuran and Related Compounds”. AZChE J. 1981, 6, 994-1001. Pakdel, H.; Roy, C. “Chemical Characterization of Wood Oils Obtained in a Vacuum Pyrolysis Process Development Unit”. Prepr. Pap-Am. Chem. SOC., Diu. Fuel Chem. 1987,2, 203-214. Ratcliff, M.; Posey, F.; Chum, H. “Catalytic Hydrodeoxygenation and Dealkylation of a Lignin Model Compound”. Prepr. Pap.Am. Chem. SOC.,Diu. Fuel Chem. 1987,2,249-256. Roy, C.; Larouche, J. P.; de Caumia, B. Unpublished data, UniversiG Laval, QuBbec, Canada, 1987. Roy, C.; Lemieux, R.; de Caumia, B.; Pakdel, H. “Vacuum Pyrolysis of Biomass in a Multiple Hearth Furnace”. BiotechnoL Bioeng. 1985, 15,108-113. Soltes, E, J.; Lin, S. C. K. ‘Hydroprocessing of Biomass Tars for Liquid Engine Fuels”. In Progress in Biomass Conversion; Tillman, D. A,, Jahn, E. C., Eds.; Academic: New York, 1984; Vol. V. Ternan, M.; Brown, J. R. “Hydrotreating a Distillate Liquid Derived from Subbituminous Coal Using Sulphided CoO-Mo03-Al203 Catalyst”. Fuel 1982, 1110-1118. Train, P. M. “Chemical and Stochastic Modeling of Lignin Hydrodeoxygenation”. Ph.D. Thesis, University of Delaware, Newark, 1986. Vasilakos, N. P.; Sequera, C. E. “Catalytic Hydrogenolysis of Cellulose”. BiotechnoL Bioeng. Symp. 1983, 13, 65-79. Vasyunina, N. A.; Chepigos, S. V.; Balandin, A. A.; Barysheva, G. S. “Catalytic Hydrogenation of Wood and other Vegetal Materials”. Akad. Nauk SSSR. Otd. Khim. Nauk 1960,8,1419. Vasyunina, N. A.; Barysheva, G. S.; Balandin, A. A.; Chepigos, S. V.; Pogosov, Y. L. “Hydrolytic Hydrogenation of Cotton Cellulose”. Zh. Prikl. Khim. 1964, 12, 2725-2729. (Translation) Vasyunina, N. A.; Barysheva, G. S.; Balandin, A. A. “Catalytic Properties of Ruthenium in the Hydrogenation of Monosaccharides“. Zzu. Akud. Nuuk SSSR, Ser. Khim. 1969, 4, 848-854. (Translation) Weisser, 0.; Landa, S. Sulfide Catalysts, their Properties and Applications; Pergamon: New York, 1973. Wisniak, J.; Hershkowitz, M.; Leibowitz, R.; Stein, S. “Hydrogenation of Xylose to Xylitol”. Znd. Eng. Chem. Prod. Res. Deu. 1974, 1 , 75-79. Wisniak, J.; Simon, R. “Hydrogenation of Glucose, Fructose, and Their Mixtures”. Znd. Eng. Chem., Prod. Res. Dev. 1979,1,5&57. Received for review October 28, 1987 Revised manuscript receiued May 16, 1988 Accepted July 7, 1988
Hydrofining of Heavy Gas Oil on Zeolite-Alumina Supported Nickel-Mol ybdenum Catalyst R. S. Mann,* I. S. Sambi,? and K. C. Khulbe Department of Chemical Engineering, University of Ottawa, Ottawa, Canada K I N 6N5
The hydrotreatment of heavy gas oil derived from Athabasca bitumen was studied in a trickle bed reactor over Ni-Mo supported on zeolite-alumina-silica catalyst at 623-698 K (350-425 “C), LHSV 1-4, and 6.99 MPa. The effects of temperature and liquid flow rates on the product were investigated. The activities of this catalyst for hydrodenitrogenation (HDN) and hydrodesulfurization (HDS) are compared with a commercially available Ni-Mo on y-alumina catalyst. This catalyst was able to remove as much as 99% S and 86% N present in the oil a t 698 K (425 “ C ) . Activation energies for the HDS and HDN reactions were i0.8 and 25.1 kcal/mol, respectively. The use of zeolites in hydrocracking has been described in great detail (Bolton, 1976). A good hydrocracking catalyst should have a highly acidic cracking component along with a noble metal or a combination of noble metals as a hydrocracking component. Zeolites are highly acidic in nature. But pure zeolites due to its fine pores are not suitable for cracking. Hence, they are usually mixed with Present address: Energy Research Laboratories, CANMET, Energy Mines and Resources, Ottawa, Canada K1A OG1.
0888-5885/88/2627-1788$01.50/0
other amorphous support materials. The rare earth exchanged Y-zeolite (faujasite) dispersed in the matrix of silica-alumina, synthetic or natural clays, is most commonly used as a cracking catalyst (Heinemann, 1981). However, use of such catalysts in hydrocracking is still quite limited. The main objective of this study was to develop a high-efficiency Ni-Mo catalyst using zeolite material and a composite of silica-alumina as support material. In this paper, we report the kinetics of hydrodesulfurization (HDS) and hydrodenitrogenation (HDN) 1988 American Chemical Society
Ind. Eng. Chem. Res., Vol. 27, No. 10, 1988 1789 Table I. Properties of Heavy Gas Oil specific gravity, 60/60 O F API gravity, OAPI viscosity (at 25 "C)
CP cst equiv Saybolt universal vis. (100 OF), s asphaltenes: wt % aniline pt, OC elemental anal.. wt % H N
S
0.9803 12.8 241.5 246.4 1140.0 1.0 52.7 86.44 10.12 0.47 2.97 0.70
C/H atomic ratio ASTM-2887 distillation 322, 340, 373, 392, 411, 429, 452, 477, 504 temp, "C 5, 10, 20, 30, 40, 50, 60, 70, 80 vol, % a
Pentane insoluble.
of heavy gas oil derived from the Athabasca tar sands bitumen in a trickle bed reactor using a high efficiency Ni-Mo on a zeolite-silica-alumina support catalyst developed in our laboratory. The results obtained in this study are compared to those obtained earlier with a commercially available Ni-Mo on alumina catalyst (Mann et al., 1987).
Experimental Section The hydrotreatment of heavy gas oil (Table I) was carried out in a trickle bed reactor. The experimental technique and analytical procedure used were similar to that described elsewhere (Mann et al., 1982) except that the exit gas flow rate was continuously monitored with a wet test meter and analyzed with an on-line gas chromatograph. Simulated distillation, as per ASTM-D2887, was performed for liquid product instead of ASTM-D-86 distillation. Kinematic viscosities were measured with Cannon Fenske routine viscometer (no. 400). a-Alumina was used to dilute the catalyst instead of kiselguhr used in previous study. The gas formation reported was a combination of n-C1-n-C5, i-C4,and i-C5. H2S and NH3 were not included in the gas formation. Catalyst Preparation. The catalyst used was prepared by mixing silica-alumina gel with zeolite powder. The silica-alumina was prepared by mixing required amounts of aluminum nitrate solution and sodium silicate. Gel was precipitated by adding ammonium hydroxide. Settled gel (after 1 day) was decanted, and the required amount of zeolite powder (type Y-zeolite, rare earth exchanged, supplied by Strem Chemical Co., 38-50-pm size) was added to the gel. The mixture was thoroughly stirred to disperse the powder uniformly. The gelzeolite powder mixture was filtered by using vacuum pump and washed with 2% by weight solution of ammonium acetate. After three washings, the filter cake was allowed to stay in the funnel, letting the air run through the cake. The gel-zeolite powder cake was well mixed, and a uniform extrudable paste was made. Extrusions of size 9 mm were made and dried first at room temperature and then dried at about 60 "C for 2 h. It was finally calcined at 700 "C. The calcined support was crushed and screened to 70-80 mesh. The particles were impregnated with nickel nitrate and ammonium molybdate solutions using the vacuum impregnation technique (Sambi, 1986). Ni-Mo-doped catalyst was calcined for 6 h after drying it at room temperature. The concentrations of Si, Al, Ni, and Mo were analyzed by using argon plasma atomic emission spectroscopy. The
Table 11. Catalyst Specification bulk catalyst density, pore vol, type n/mL mL/n 2.24% NiO0.7956 0.597 5.37% MoOB" (this study) 3.8% NiON.A. 0.54 16.8% MOO^^
BET surface pore size area, m2/n (av), A 293 40.8 152
113.6
'Support: Y-zeolite = 25%, SiOz = lo%, AlzOs = 65%. Prepared in our own laboratory. Support: y-alumina, commercial: Harshaw 100.
powdered catalyst was first fused with lithium metaborate in graphite crucible. The melt was dissolved in nitric acid and was used for analysis. The concentrations of Si and A1 were also checked by using the gravimetric method (Vogel, 1978). The results compared well with the atomic emission spectroscopy results. Surface area of the catalyst was measured by the BET method (Faeth and Willingham, 1955). Pore volume of the catalyst particles was measured by using the carbon tetrachloride adsorption method (Benesi et al., 1955). A central composite experimental design (Box and Wilson, 1951; Box and Hunter, 1957) was used to locate an optimum catalyst support composition. Silica and REY-zeolite concentrations in the catalyst support were the only two variables. Ranges of 10-50% for silica and 10-40% for REY-zeolite were selected, and the rest was alumina. All other factors in catalyst preparation were controlled to get good reproducibility. This study indicated that a catalyst support containing 25% zeolite, 10% silica, and the rest alumina was best for HDS and HDN (Sambi, 1986). A catalyst with the above-mentioned composition was prepared and used for a kinetic study, the results of which are discussed in this paper. Physical and chemical properties (composition) of this catalyst are given in Table 11. Scanning electron microscope and X-ray analysis revealed a uniform distribution of Si, Al, Ni, and Mo atoms throughout the catalyst matrix. Operating Conditions. Eight milliliters of 70-80-mesh catalyst particles was diluted with equal volume of the same size a-alumina particles. The mixture was packed into the reactor tube by using a vibrator. The catalyst was conditioned for a period of 20 h at 648 K (375 "C), 6.99 MPa (1000 psig), LHSV of 1.5, and gas flow rate of 0.87 m3/L (5000 scf/bbl) prior to actual experimental run. After an initial conditioning period of 20 h, the kinetic study runs were carried out in order of increasing temperatures. The four selected temperatures were 623,648, 673, and 698 K, and the liquid flow rates were varied to obtain LHSV's of 4.0, 2.0, 1.333, and 1.0.
Results and Discussion In the beginning and at the end, as well as in between, temperature changes, the activity of the catalysts was checked by collecting samples at the standard preconditioning operating conditions. No significant catalyst deactivation was observed during the entire kinetic study (over 100-h operation). The results of the analysis of the samples obtained in this kinetic study are given in Tables I11 and IV. The gas formation values are given in percentage by weight of the oil feed. (a) Heat- and Mass-Transfer Effects. The catalyst had a particle size between 0.124 and 0.147 mm. The effectiveness factor for this size of catalyst used in similar application has been reported to be nearly one (Smith, 1981). Temperature inside the catalyst bed was within f l
1790 Ind. Eng. Chem. Res., Vol. 27, No. 10, 1988 Table 111. Results of Kinetic Study with Ni-Mo on Silica-Alumina-Zeolite Catalysta density, viscosity, aniline mid bp, sulfur, temp,OC LHSV g/mL CP Pt, "C "C wt % orig. oil 0.9803 241.5 52.7 429.1 3.05 50.4 411.8 1.04 350 4.0 0.9575 123.0 118.1 51.2 393.8 0.70 350 2.0 0.9517 51.7 0.9470 94.6 396.8 0.53 350 75.9 52.4 402.1 0.40 350 4/3 1.0 0.9437 95.7 51.6 402.1 0.72 375 4.0 0.9490 375 2.0 0.9431 78.7 52.5 403.0 0.39 0.9396 68.6 54.0 389.9 0.25 375 61.1 53.8 389.8 0.13 375 4/3 1.0 0.9379 400 4.0 0.9428 61.5 52.6 389.6 0.34 400 2.0 0.9349 41.7 52.0 385.8 0.13 0.9276 31.7 53.8 382.8 0.08 400 413 400 1.0 0.9256 26.2 51.3 378.3 0.04 425 4.0 0.9343 33.9 48.9 378.0 0.16 20.5 47.2 370.4 425 2.0 0.9247 0.04 0.9214 17.2 46.5 425 368.1 0.03 425 4/3 1.0 0.9216 16.9 46.0 388.5 0.03
nitrogen, PPm 4831 3650 3341 3094 2946 3214 2757 2235 2063 2280 1765 1403 1226 1927 987 760 738
formation gas, wt % 0.0864 0.1734 0.2553 0.3379 0.1929 0.2984 0.4198 0.5585 0.5837 0.9576 1.2815 1.4651 1.5824 1.9669 2.1747 2.6051
"Pressure = 6.99 MPa (1000 psig); gas flow rate = 0.89 m3/L (5000 scf/bbl). 140-
Table IV. Analysis of a Typical Exit Gas Stream Run Conditions temp 698 K (425 "C) pressure 6.99 MPa (1000 psig) LHSV 1.0 0.89 m3/L (5000 scfjbbl) gas flow rate methane ethane propane isobutane n-butane isopentane n-pentane
I
Gas Analysis, mol % 0.168 0.157 0.087 0.011 0.053 0.004 0.016
098r
Figure 2. Effect of temperature on viscosity of the product. ( 0 ) Silica-alumina-zeolite supported Ni-Mo; (0) commercial Ni-Mo on y-alumina.
LdS,v=2
LHSV
:: 0 9 4 -
\
I
a
i
t m w 3
092-
3 g o L - . - - i 300
1
350
400
TEMPERATURE,
450
"C
Figure 1. Effect of temperature on density of the product. ( 0 ) Silica-aluminazeolite supported Ni-Mo; (0) commercial Ni-Mo on y-alumina.
>
I
20-
"C. Since the reactor used in this study had an internal diameter of only 0.52 cm, temperature gradient in the radial direction of the catalyst bed was considered insignificant. Thus, the heat- and mass-transfer effects were considered to be minimum. (b) Effect of Temperature. The effect of an increase in temperature from 623 to 698 K (350-425 "C) can be seen from Table 111. The results indicate that both density and viscosity of the product oil samples decreased with an increase in temperature. Similarly, higher S and N removal and lower midboiling point values are obtained with an increase in temperature. Aniline point values show an increase and then a decrease with an increase in reactor
l o ' 3do
1
I
350
400
_-L 450
TEMPERATURE, "C
Figure 3. Effect of temperature on hydrodesulfurization activity of comthe catalyst. ( 0 )Silica-alumina-zeolite supported Ni-Mo; (0) mercial Ni-Mo on y-alumina.
temperature within the selected temperature range. The activity of our catalyst is compared with the results obtained with a commercially available Ni-Mo on y-alumina catalyst in Figures 1-4. The physical properties of the two catalysts are given in Table 11. Ni-Mo supported on silica-alumina-zeolite catalyst gave products of lower
Ind. Eng. Chem. Res., Vol. 27, No. 10, 1988 1791 1OOr
Table V. Rate Constants for HDS and HDN" HDN 2nd-order HDS 1.5th-order temp, rate const, h-' rate const, h-' "C (wt %)-1 (wt % ) 4 5 350 1.5067 2.158 375 2.9971 4.200 400 6.5309 8.571 425 13.1237 12.584 "Pressure = 6.99 MPa (1000 psig); gas flow rate = 0.89 m3/L (5000 scf/bbl).
density at all temperatures compared to commercial NiMo catalyst supported on alumina (Figure l). Viscosities of the products showed a similar trend (Figure 2). These results indicate that hydrogenolysis of heteromolecules was more pronounced. The hydrodesulfurization (HDS) and hydrodenitrogenation (HDN) activities of this catalyst were much higher as compared to the commercial catalyst (Figures 3 and 4). Both HDS and HDN activities increased with increasing temperature. At 698 K (425 "C) and LHSV of 1, percentage removal of sulfur and nitrogen was 99 % and 84.7 % , respectively. A typical gas analysis indicating various hydrocarbon gases in an exit gas stream is shown in Table IV. (c) Effect of Liquid Flow Rate. The effect of the change in feed rates of heavy oil on the properties of the product was studied at 623-698 K (350-425 "C) and 6.99 MPa by varying the flow rates of the oil to correspond to LHSV values of 1.0-4.0. The effect of liquid feed rates on physical properties like density, viscosity, etc., is given in Table 111. The densities of the product oil decreased with a decrease in LHSV. The aniline point did not change significantly with LHSV. The HDN and HDS activities of the catalyst increased with decreasing LHSV. It can be seen that, even with a high LHSV of 4, this catalyst was able to remove 95% S and 60% N from the heavy gas oil. (d) Kinetic Model. The power law rate equation is r = dC/dt = -K,(C)m (1) where C is the concentration of S and N in weight percent and m is the order of the rate. The ideal plug flow model equation is
where Km" is the intrinsic reaction rate constant, z is the dimensionless length parameter, t is the void fraction of the catalyst bed, q is the catalyst effectiveness factor, K,' is the mth order reaction rate constant based on the total volume of catalyst, and C is the concentration. By use of eq 2 and if the effectiveness factor is combined with the rate constant, eq 1 can be integrated to the following: (3) and (4) Table VI. Results of Kinetic Studs with a-Alumina" density, viscosity, aniline temp, "C e/mL CP Pt, "C 193 46.8 350 0.9798 375 0.9799 186 46.8 157 46.1 400 0.9766 131 45.7 425 0.9765
I
1
350
400
I 450
TEMPERATURE, "C
Figure 4. Effect of temperature on hydrodenitrogenation activity of the catalyst. ( 0 )Silica-alumina-zeolite supported Ni-Mo; (0) commercial Ni-Mo on yalumina.
where kl and kz are the first- and second-order rate constants and subscripts "i" and indicate the inlet and outlet concentrations of S and N, respectively. The integrated equation for the mth order can be written as (5) Nitrogen and sulfur removal data were first tried with first- and second-order models. Reaction orders of m = 1.1,1.2, 1.3, ...,1.9 were then tried for both HDS and HDN. The best fit for HDN was obtained for m = 2, and for HDS, 1.5 order gave the best fit. The k values are listed in Table V. The activation energies (using Arrhenius equation) for hydrodesulfurization and hydrodenitrogenation were 20.81 and 25.12 kcal/mol, respectively. (e) Hydrotreatment without Catalyst. To study the extent of thermal activity, only the inert material (a-alumina) was packed in the reactor. This was preconditioned exactly in the same way as the catalyst, and then 16 runs with exactly the same operating conditions as used for the kinetic study with the catalyst were carried out and the products analyzed. The results of the four runs with LHSV of 1.0 are given in Table VI. It can be seen that the effect of an increase in temperature on the nitrogen and sulfur removal was insignificant. A maximum of 18% HDN and 14% HDS occurred with a-alumina. The gas formation did show an increase with increasing temperature. The densities of the products at all temperatures were very close to the density of the original oil, although a slight decrease in density with increasing temperature was indicated. There was a slight decrease in viscosity as the mid bp, "C 414.5 418.2 413.1 412.2
sulfur, wt%
2.78 2.72 2.78 2.73
"Pressure = 6.99 MPa (1000 psig); LHSV = 1.0; gas flow rate = 0.89 m3/L (5000 scf/bbl).
nitrogen, mm 3950 4113 4165 3973
gas formation, wt%
0.1785 0.2627 0.3379 0.7013
1792
I n d . Eng. Chem. Res. 1988,27, 1792-1797
temperature was increased. The midboiling point values showed only a very small decrease with increasing temperatures. The aniline point values for all these runs were within 44-47 "C, which was lower than for the original oil (52.7 "C). It appears from these observations that mostly the long-chain paraffinic components are cracked, giving lower viscosity values and lower aniline points. The resulting increased aromaticity could thus explain the negligible change in density and midboiling points. It can be concluded from a comparison of these results with those obtained with the catalyst that most of the sulfur and nitrogen removal was mainly due to the catalytic activity and not the thermal effects. Also the catalyst was able to successfully break the condensed aromatic molecule to yield more paraffinic product oil. Conclusions The hydrotreatment of heavy oil was investigated over a highly efficient catalyst containing 2.24 wt % NiO and 5.37 wt % Mooz supported on a support (10 w t % silica, 25 w t % rare earth exchanged Y-zeolite, and 65 wt % alumina) in a trickle bed reactor at 623-698 K (350-425 "C), LHSV of 1-4, and 6.99 MPa. At 698 K (425 "C) and a LHSV of 2, it removed 99% S and 80% N present in the heavy oil as compared to 86% S and 61.4% N removed by a commercial Ni-Mo on alumina catalyst at 723 K (450 "C)and at the same pressure and LHSV. A maximum gas formation of about 0.7 wt % of feed oil was observed at 698 K (425 " C ) and LHSV of 1. The kinetic study suggested orders of 1.5 and 2.0 in the power law model for S and N removal, respectively. Activation energies for the
HDS and HDN were found to be 20.8 and 25.1 kcal/mol, respectively.
Acknowledgment We are grateful to the Natural Science and Engineering Research Council of Canada for financial aid (A-1125)and to CANMET. Registry No. Ni, 7440-02-0;Mo, 7439-98-7. Literature Cited Benesi, H. A.; Bounar, R. V.; Lu,C. F. Anal. Chem. 1955,27,1963. Bolton, A. P.Experimental Methods i n Catalyst Research; Anderson, R. B., Ed.; Academic: New York, 1976;p 33. Box, G. E. P.; Hunter, J. S. An. Math. Statis. 1957,28, 195. Box, G.E. P.; Wilson, K. B. J . R. Stat. SOC.1951,13, 1. Faeth, P.A.; Willingham, C. B. "The Assembly, Calibration, and Operation of Gas Adsorption". Technical Bulletin of Physical Chemistry, Carnegie Mellon Institute of Research, Pittsburg, Sept
1955. Heinemann, M. CataE. Rev. Sci. Eng. 1981,23(1982),315-328. Mann, R.S.;Sambi, I. S.; Khulbe, K. C. Ind. Eng. Chem. Prod. Res. Deu. 1982,21 (4),575. Mann. R. S.: Sambi., I. S.:. Khulbe. K. C. Ind. Enp. - Chem. Res. 1987, 26,'410.
'
Sambi, I. S. Ph.D. Thesis, University of Ottawa, Ottawa, Canada,
1986. Smith, J. M. Chemical Engineering Kinetics, 3rd ed.; McGraw-Hill: New York, 1981. Vogel, A. Textbook of Quantitative Inorganic Analysis, 4th ed.; Longman: London, 1978.
Received for review September 30, 1987 Revised manuscript received April 12, 1988 Accepted May 23, 1988
Measurement of Brernsted Acid and Lewis Acid Strength Distributions of Solid Acid Catalysts Using Chemisorption Isotherms of Hammett Indicators Kenji Hashimoto,* T a k a o Masuda, and Hideki Sasaki Department of Chemical Engineering, Kyoto University, Kyoto 606, Japan
A new method was developed that discriminates Brernsted acid and Lewis acid sites and measures the distributions of both kinds of sites. It uses the chemisorption isotherms of Hammett indicators. Three different samples were prepared: an unpoisoned sample and two poisoned samples. In one of the poisoned samples, only the Brernsted acid sites were poisoned with 2,6-dimethylpyridine, whereas in the other poisoned sample both types of acid sites were completely poisoned with ammonia. The amount of indicator chemisorbed on each kind of acid site was calculated from differences in the amounts of indicators adsorbed on the three samples. Using the two kinds of chemisorption isotherms obtained, we calculated the acid strength distribution curves of the Brernsted acid and Lewis acid sites for the range of acid strengths (Ho)from -3 to the strongest strength using our recently reported indicator adsorption method. The acidic properties of six silica-alumina catalysts were measured in the range of acid strength from -3 to about -15 by use of the proposed method. The Brernsted acid sites on the silica-alumina catalysts had a relatively wide distribution of strengths, whereas most of the Lewis acid sites were distributed in the narrow range of acid strength from -10 to -13. Acid sites of different strengths are distributed over the surfaces of such solid acid catalysts as silica-alumina and zeolite. These acid sites are classified as Brernsted acid sites which donate protons to the reactants and as Lewis acid sites which accept a lone pair of electrons from the reactants. The functions of these sites differ. Therefore, it is necessary to discriminate between them and to 0888-5885/88/2627-1792$01.50/0
evaluate the distribution curves of the two kinds of acids. Methods have been presented for measuring the acidic properties of the Brernsted acid and Lewis acid sites (Benesi and Winquist, 1978; Tanabe, 1970). The infrared absorption spectra of pyridine on silica-alumina and on zeolite have been studied (Parry, 1963; Ward, 1967). Pyridine chemisorbed on Brernsted acid sites can be dis0 1988 American Chemical Society
