Catalytic Conversion of Canola Oil over Potassium-Impregnated

3332

Ind. Eng. Chem. Res. 1996, 35, 3332-3346

Catalytic Conversion of Canola Oil over Potassium-Impregnated HZSM-5 Catalysts: C2-C4 Olefin Production and Model Reaction Studies Sai P. R. Katikaneni, John D. Adjaye, Raphael O. Idem, and Narendra N. Bakhshi* Catalysis and Chemical Reaction Engineering Laboratory, Department of Chemical Engineering, 110 Science Place, University of Saskatchewan, Saskatoon, Canada S7N 5C9

The influence of catalyst acidity, reaction temperature, and canola oil space velocity on the conversion of canola oil was evaluated using a fixed-bed microreactor at atmospheric pressure at reaction temperatures and space velocities (WHSV) in the ranges 400-500 °C and 1.8-3.6 h-1, respectively, over potassium-impregnated HZSM-5 catalysts. These catalysts were thoroughly characterized using XRD, N2 adsorption measurements, 1H NMR, TPD of NH3, FT-IR, and model compound reactions. Also, conditions for the production of the maximum yield of C2-C4 olefins from canola oil were determined. The incorporation of potassium into HZSM-5 catalyst resulted in both the dilution and poisoning of Bronsted and total acid sites. These acidity changes only severely affected the acid catalyzed reactions, such as oligomerization and aromatization, and resulted in drastic modifications in product distribution. The maximum C2-C4 olefin yield of 25.8 wt % was obtained at 500 °C and 1.8 h-1 space velocity with catalyst K1 of relatively low Bronsted and total acidity. Introduction Paraffins, olefins, naphthenes, and aromatics are important groups of organic compounds. They are useful in many chemical and petrochemical industries for the manufacture of a wide range of products. The olefins are a particularly interesting group of hydrocarbons with unique chemistry characterized by relatively higher reactivity than paraffins, aromatics, or naphthenes. Consequently, they are used in various important industrial processes such as alkylation, polymerization, oligomerization, and aromatization reactions (Gary and Handwerk, 1984). Traditionally, hydrocarbons are produced from petroleum sources (Waddama, 1980). However, recent studies (Milne et al., 1990; Weisz et al., 1979; Campbell, 1983; Baker and Elliott, 1987; Boocock et al., 1992; Prasad et al., 1986a,b; Craig and Coxworth, 1987; Chantel et al., 1984; Katikaneni et al., 1995a; Adjaye and Bakhshi, 1995; Sharma and Bakhshi, 1991) have shown that they can also be obtained from other sources, such as from the catalytic conversion of various types of plant oils. According to these authors, such studies involve passing these oils at temperatures in the range 350-550 °C over cracking catalysts such as HZSM-5, Pt/ZSM-5, silica-alumina, H-Y, H-modernite, aluminum pillared clays, and hybrids of these catalysts. Also, in most of these studies, hydrocarbons in the gasoline boiling range were predominant in the organic liquid fraction, whereas the gaseous fraction contained mostly paraffinic hydrocarbon components. There is a considerable demand for lower olefins such as ethylene, propylene, 1- and 2-butenes, and isobutylene because of the growing interest in their application in the manufacture of desirable products such as polyethylene, polypropylene, methyl tert-butyl ether (MTBE), and ethyl tert-butyl ether (ETBE). Our interest is in the production of these olefins from plant oils. Considerable information exists in the literature on the catalyst characteristics and operating conditions required for the production of optimum amounts of C2* Author to whom all correspondence should be addressed.

S0888-5885(95)00740-8 CCC: $12.00

C4 olefins from petroleum sources. According to Olson et al. (1980), Haag et al. (1980), Chu and Chang (1984), Kaeding and Butter (1980), and Zatorski et al. (1985), these conditions are as follows: a relatively low catalyst acidity (especially the Bronsted acidity), a catalyst with relatively small pore sizes, short contact times between the feed and catalyst, and high reaction temperatures. On the other hand, the chemical characteristics of plant oils are markedly different from those of petroleum oil. Thus, it is highly desirable to determine the conditions for the production of optimum amounts of C2-C4 olefins from the catalytic conversion of plant oils. In the literature, studies on the modification of catalyst acidity have been reported mostly for cracking reactions involving feed from petroleum sources. These modifications have been achieved principally either by incorporating phosphorus, boron, or an alkali metal or by changing the silica/alumina ratio (Chang, 1983; Szostak, 1989; Bhatia, 1990; Chu et al., 1985; Rahman et al., 1988). Recently, studies have been reported on the modification of the acid sites of HZSM-5 catalyst by the incorporation of platinum (Katikaneni et al., 1996), mainly to evaluate the effects of catalyst acidity on the yields of various products from canola oil cracking. These studies showed that the impregnation of HZSM-5 catalyst with platinum resulted in an increase in the isomerization function of the catalyst. For example, there was an increase in the amount of isoparaffins formed using Pt/HZSM-5 catalysts compared with the use of nonimpregnated HZSM-5 catalyst. Also, it was shown in the study that impregnation of HZSM-5 catalyst with platinum only resulted in the dilution of the acid site density of the catalyst. On the other hand, it is well-known that impregnation with an alkali metal will result both in the dilution and poisoning of the catalyst acid sites. However, the consequence of alkali metal impregnation on the yield of C2-C4 olefins or yields of other products from the conversion of plant oils has not been reported so far in the literature. It is desirable to obtain a better understanding of the canola oil conversion process itself. Usually, this is done © 1996 American Chemical Society

Ind. Eng. Chem. Res., Vol. 35, No. 10, 1996 3333

in conjunction with studies of the characteristics of the catalyst. A number of workers (Katikaneni et al., 1995a,b; Prasad et al. 1986a,b) have proposed various reaction pathways for canola oil conversion, and these have been based primarily on the trends observed in the product distribution. In both the thermal and catalytic conversions, the intermediate reaction steps involved include the deoxygenation of oxygenated hydrocarbon intermediates such as alcohols, ketones, aldehydes, and organic acids (Chang and Wan, 1947; Nawar, 1969). In the case of zeolite catalysts, subsequent steps include isomerization reactions involving aromatic compounds (such as p-, m-, and o-xylenes), paraffins (such as n-butane), and olefins (such as butylene). In the literature (Adjaye and Bakhshi, 1995), the involvement of intermediate reactions in a proposed reaction pathway can be verified through the use of model compound reaction studies. Usually this involvement is confirmed if the variation in product distribution with operating conditions and catalyst characteristics exhibited by model reactions are consistent with that exhibited by the main reaction of interest (canola oil cracking) conducted under identical conditions (i.e., catalyst acidity, feed space velocity, and reaction temperature). In this work, we have studied the effects of the degree of acid site dilution and poisoning of HZSM-5 catalysts on the yields of C2-C4 olefins from canola oil conversion. We have also evaluated the optimum catalyst characteristics and reaction operating variables required for maximum production of lower olefins from canola oil. This work also involved the preparation of potassiumimpregnated HZSM-5 catalysts with potassium concentrations in the range 0-2 wt %. It also involved the thorough characterization of these catalysts using techniques such as X-ray diffraction (XRD), N2 adsorption measurements, 1H MAS solid state NMR, temperature programmed desorption (TPD) of NH3, FT-IR studies, and model compound reactions in order to obtain a better understanding of the relationship between catalyst performance and catalyst characteristics. The results have been used to elucidate the reaction pathway for canola oil cracking over potassium-impregnated HZSM-5 catalysts. These results are presented in this paper. Experimental Section Catalyst Preparation. The potassium-impregnated HZSM-5 catalysts (K/HZSM-5) used in this work were obtained by incorporating potassium into HZSM-5 catalyst. The nonimpregnated HZSM-5 catalyst was prepared according to the procedure reported by Chen et al. (1973). Altogether, four K/HZSM-5 catalysts of potassium concentrations in the range 0-2 wt % were prepared. Potassium was incorporated into the HZSM-5 catalyst by wet impregnation techniques using aqueous solutions of potassium carbonate (analytical grade and obtained from BDH Chemicals, Toronto, Canada) as the precursor. This impregnation step was carried out for 24 h. After this step, the catalysts were dried overnight at 100 °C and then calcined at 500 °C for 6 h in a muffle furnace. The resulting potassium-impregnated HZSM-5 catalysts were designated as K0 (0 wt % potassium), K1 (0.5 wt % potassium), K2 (1 wt % potassium), and K3 (2 wt % potassium) catalysts. Catalyst Characterization. Catalyst characterization involved the determination of both the physical and chemical characteristics of the catalysts.

(A) Physical Property Characterization. (1) N2 Adsorption Measurements. The surface area, pore volume, pore size, and pore size distribution of all of the catalysts were determined using a Micromeritics adsorption equipment (Model ASAP 2000) which was equipped with a micropore analysis program using nitrogen (99.995% purity; obtained from Linde, Calgary, Canada) as the analysis gas. Prior to analysis, each catalyst was evacuated at 300 °C at a vacuum of 0.54 KPa for 10 h. HZSM-5 based catalysts contains mostly micropores. Thus, the surface areas measured were the micropore surface areas. Therefore, the DubininAstakhov program was used in evaluating the micropore surface areas, while the Hovarth-Kawazoe equation was used to estimate the median pore sizes of both HZSM-5 and K/HZSM-5 catalysts. (2) Powder X-ray Diffraction Measurements. Powder XRD measurements were performed to identify component phases as well as to determine the degree of crystallinity of the catalysts as a function of potassium concentration. The XRD measurements were made with a Phillips diffractometer using Fe KR radiation in the scanning angle (2θ) range of 10-90° at a scanning speed of 2 deg/min. (B) Chemical Property Characterization. The chemical property characterization techniques were employed principally for the determination of the acidity of the catalysts and are described below. (1) Temperature Programmed Desorption of Ammonia. TPD analysis was performed in order to determine the acid strengths and distributions on the catalysts. The analysis was conducted in a conventional flow system similar to the one described by Idem and Bakhshi (1994). The carrier gas used was N2 (99.995% purity; obtained from Linde) at a flow rate of 60 mL/ min, while a mixture of 1% NH3 in N2 (high purity and also obtained from Linde) was used as the adsorbing gas. About 0.5 g of the calcined catalyst sample was used for each experiment. NH3 adsorption was performed by flowing 1% NH3 in N2 gas over the catalyst for 1 h. Adsorption was carried out at 100 °C in order to eliminate physically adsorbed NH3. Temperature programmed desorption of NH3 started at 100 °C and ended at 650 °C at a temperature programming rate of 8 °C/min. (2) FT-IR Measurements. The FT-IR technique was employed to identify the nature of acid sites present on the catalyst samples. The IR measurements were made on powdered catalyst samples using a Biorad Infrared Spectrometer (Model FTS 40, Digilab Division). Two regions of the IR spectra were explored. These were the hydroxyl and pyridine regions. Consequently, fresh and pyridine chemisorbed catalyst samples were used to obtain the spectra for hydroxyl (3500-4000 cm-1) and pyridine (1400-1650 cm-1) regions, respectively. Pyridine chemisorbed samples were obtained by passing pyridine vapor over the catalysts at 150 °C for 1 h in the same flow system that was used previously for NH3 adsorption. After the pyridine adsorption, each sample was allowed to cool to room temperature and subsequently used for IR analysis. (3) Solid State NMR Studies. 1H MAS solid state NMR studies were carried out in order to determine the number of hydroxyl groups present on the catalyst. These experiments were performed using a Bruker AM 360 WB instrument equipped with a CP/MAS facility (available at the National Research Council Laboratory, Saskatoon, Canada). The instrument was operated at

3334 Ind. Eng. Chem. Res., Vol. 35, No. 10, 1996 Table 1. Composition of Canola Oil composition

chemical formula

wt %a

no. of carbon atoms

no. of carbon double bonds

oleic acid linoleic acid linolenic acid stearic acid palmitic acid

CH3(CH2)7CHdCH(CH2)7COOH CH3(CH2)4CHdCHCH2CHdCH(CH2)7COOH CH3CH2CHdCHCH2CHdCHCH2CHd(CH2)7 COOH CH3(CH2)16COOH CH3(CH2)14COOH

60 20 10 2 4

18 18 18 18 16

1 2 3 0 0

a Small quantities of eicosenic (20:1) and eruic (22:1) acids totalling 4 wt % (20, 22 represent the number of carbon atoms and 1 represents the number of carbon double bonds).

Figure 1. Scheme for the collection and analyses of various products from the conversion of canola oil.

a frequency of 360.13 MHz and a magnetic field of 7.05 T. Other operational parameters were a pulse width and frequency of 6.5 s and 4000 Hz, respectively. All operations were performed at room temperature, whereas all of the chemical shifts were measured relative to that of tetramethylsilane (TMS). Catalyst Performance Studies. (A) Equipment. The performance of the potassium-impregnated and nonimpregnated HZSM-5 catalysts was studied in a stainless steel (SS 316) fixed-bed (down flow) microreactor (11.5 mm id and 400 mm overall length). Details concerning the experimental rig used are given elsewhere (Katikaneni et al., 1995a,b). The furnace temperature was controlled by a series SR22 microprocessorbased autotuning PID temperature controller (Shimaden Co. Ltd., Tokyo, Japan) using a K-type thermocouple placed on the furnace side of the annulus between the furnace and the reactor. A separate thermocouple was used to monitor the temperature of the catalyst bed. This arrangement was capable of ensuring an accuracy of (1 °C for the catalyst bed temperature. (B) Canola Oil Conversion. Test runs were performed for each catalyst at atmospheric pressure at reaction temperatures in the range 400-500 °C and space velocities (WHSV) ranging from 1.8-3.6 h-1. The canola oil was obtained from CSP Foods, Saskatoon, Canada, and was of the degummed and refined variety. Earlier analysis showed that it consisted mainly of

unsaturated triglycerides having an average molecular formula of C59H94O5. The composition of canola oil is given in Table 1. Each experimental run required 2 g of the catalyst. A typical test run was conducted as follows: A plug of glass wool (about 0.5 g), which was used as screen as well as support for the catalyst, was placed on a stainless steel grid positioned centrally within the reactor. The catalyst was then loaded on the glass wool. The catalyst bed was heated to the desired reaction temperature in flowing argon gas. When this temperature was reached, the argon flow was stopped and the feed was pumped into the reactor at the desired canola oil space velocity. The product mixture leaving the reactor was condensed in a water-cooled heat exchanger, followed by an ice-cooled condenser to separate gaseous and liquid products for separate analysis. The gas product was collected over saturated brine, while the liquid product was collected in a glass trap positioned after the condenser. The scheme for the estimation of the amounts of coke, residual oil, organic liquid product (OLP), gas, and water formed for each canola oil upgrading run is given in Figure 1. A detailed description of the procedures is given elsewhere (Katikaneni et al., 1995a). (C) Model Compound Reactions. (1) Methanol Cracking and m-Xylene Isomerization Reactions. Methanol cracking and m-xylene isomerization were


used as model compound reactions to provide further evidence for the strengths of the various types of acid sites present on the catalysts. They were subsequently used for the elucidation of the reaction pathway for canola oil conversion. Both reactions were performed in the same fixed-bed reactor used for canola oil conversion at atmospheric pressure at reaction temperatures in the range 370-450 °C and space velocity (WHSV) of 1.8 h-1. The methanol and m-xylene used were analytical grade obtained from BDH Chemicals Ltd., Toronto, Canada, and Poole, England, respectively. (D) Analyses of Products. The gaseous products obtained from these reactions were analyzed with a Carle 500 GC using a combination of packed and capillary columns and both flame ionization and thermal conductivity detectors. The sample was injected into a 2 m long precolumn packed with OV-101 silicone oil. All components lighter than ethane passed rapidly through this column and were separated at 70 °C in a Porapak Q column in series with a molecular sieve 13X column (each 2.7 m in length) into CO, CO2, and C1 and C2 hydrocarbons. These gases were detected using a thermal conductivity detector (TCD). The C3+ components were back-flushed into the fused silica capillary column for which the oven temperature was programmed from 40 to 200 °C. The C3+ components were detected using a flame ionization detector (FID). The organic liquid product (OLP) was also analyzed with a Carle 500 GC which was equipped with a bonded nonpolar (methyl silicone) 50 m × 0.2 mm id column. The temperature in the GC oven was programmed from 40 to 200 °C, and detection was done using FID. Details concerning the analyses of these products are given elsewhere (Katikaneni et al., 1995a,b).

Figure 2. XRD spectra of HZSM-5 catalyst.

Results and Discussion Catalyst Characterization. (A) Powder XRD Measurements. The XRD pattern of the nonimpregnated HZSM-5 catalyst (K0) is shown in Figure 2 and is consistent with those reported in the literature (Chen et al., 1989; Campbell, 1983). The XRD patterns for the potassium-impregnated catalysts did not show any appreciable change from that given in Figure 2, implying that, up to 2 wt % potassium concentration, there was no loss of crystallinity by the catalysts due to potassium impregnation. Also, no new XRD lines were observed in the potassium-impregnated catalysts. (B) N2 Adsorption Measurements. The micropore surface areas of the potassium-impregnated and nonimpregnated HZSM-5 catalysts are given in Table 2 as a function of potassium concentration. It is seen that the micropore surface area decreased monotonically with increasing potassium concentration in the catalyst. As was mentioned earlier, HZSM-5 catalyst contains mostly micropores. Also, it is known that the micropore surface area in HZSM-5 is a measure of the inner channel surface area of the micropores. Thus, the decrease in micropore surface area with potassium concentration may be due to pore blockage of the micropores of the impregnated catalysts by potassium (most likely in the form of K2O). Pore blockage is also supported by the decrease in pore volume with increasing potassium concentration (see Table 2). These results were used in conjunction with results obtained from TPD of NH3 to evaluate the acid site densities of the respective catalysts. This is discussed below. (C) TPD of NH3. The TPD spectra for the four catalysts (K0, K1, K2, and K3) are given in Figure 3.

Figure 3. TPD spectra of HZSM-5 and K-impregnated HZSM-5 catalysts. Table 2. Characteristics of Potassium Impregnated and Nonimpregnated HZSM-5 Catalysts

catalyst HZSM-5 0.5 wt % K/HZSM-5 1.0 wt % K/HZSM-5 2.0 wt % K/HZSM-5

micropore pore catalyst surface volume, acid density, 2 identity area, m /g cm3/g mm2/(m2 g of catalyst) K0 K1

389 358

0.14 0.12

0.655 0.542

K2

325

0.10

0.400

K3

297

0.07

0.242

This figure shows that the nonimpregnated HZSM-5 catalyst (i.e., K0 containing 0 wt % potassium) as well as the impregnated HZSM-5 catalysts containing 0.5 and 1.0 wt % potassium (i.e., K1 and K2) exhibited two

3336 Ind. Eng. Chem. Res., Vol. 35, No. 10, 1996

well-resolved TPD peaks designated as the low-temperature peak (l) and the high-temperature peak (h). On the other hand, the impregnated catalyst containing 2.0 wt % potassium (K3) exhibited only the low TPD peak. The literature (Hidalgo et al., 1984) suggests that if the NH3 TPD peak temperature is greater than 350 °C, then this peak represents strong acid sites (i.e., h peak), whereas peaks with temperatures less than 350 °C represent weak acid sites (i.e., l peaks). Thus, the existence of TPD peaks at 240 and 485 °C for nonimpregnated HZSM-5 catalyst signifies the presence of both the weak and strong acid sites on the catalyst. It is well-known that, for both weak and strong acid sites, a decrease in the peak temperature represents a decrease in the strength of the acid site. Therefore the progressive shift to lower peak temperatures for both l and h peaks as the potassium concentration in the catalysts increased (see Figure 3) implies that there was a decrease in the strengths of the h and l acid sites with increasing potassium concentration. In the case of the catalyst containing 2 wt % potassium (i.e., K3), the absence of the h peak shows that there was a complete neutralization of the strong acid sites with 2 wt % potassium. Acid site density can be used to represent the concentration of the acid sites present on the surface area of the catalyst which is accessible to the reactants. Acid site density on the basis of surface area becomes important because catalytic cracking reactions are essentially a surface phenomenon which depend on the accessible acid sites present on the catalyst. In this work, acid site density of each catalyst was calculated as the ratio of the total TPD peaks area to the micropore surface area (i.e., inner channel area of the pores) of the catalysts as follows:

µ ) R/(micropore surface area)

(1)

where

µ ) total TPD peak area per unit micropore surface area, mm2/(m2 g of catalyst) R ) total area under the TPD peaks, mm2/g The results for the four catalysts are given in Table 2. It is seen from the table that the acid site density decreased as the potassium concentration in the catalysts increased. These results show that the total number of acid sites decreased with potassium concentration at a faster rate than the micropore surface area did. Thus, by combining the TPD and micropore surface area results, it can be concluded that the incorporation of potassium into HZSM-5 catalyst decreases not only the acid site strength but also the total acidity of the catalyst. This is consistent with the observations of Bhatia (1990) and Szostak (1989). (D) FT-IR Studies. In addition to determining the strength and total number of acid sites using the TPD of NH3, it was also necessary to determine the types of acid sites present (i.e., whether Lewis or Bronsted acid sites) as well as their distribution on the catalyst surface as a function of potassium concentration. According to the literature, the variation in the amounts of these various acid types has a tremendous effect on product distribution. As suggested by Vedrine et al. (1979), information regarding the types of acid sites present on the catalysts was obtained from the IR spectra of the samples in the pyridine region (frequency in the range

Figure 4. IR spectra of HZSM-5 and K-impregnated HZSM-5 catalysts in the hydroxyl region (K0, HZSM-5; K1, 1% potassium/ HZSM-5; K2, 2% potassium/HZSM-5).

Figure 5. IR spectra of HZSM-5 and K-impregnated HZSM-5 catalysts in the pyridine region (K0, HZSM-5; K1, 1% potassium/ HZSM-5; K2, 2% potassium/HZSM-5).

1425-1575 cm-1) as well as in the hydroxyl stretching region (frequency in the range 3500-4000 cm-1). The IR spectra of fresh and pyridine adsorbed samples showing these two regions are presented in Figures 4 and 5, respectively, for catalysts K0, K1, and K2. In Figure 4, the band at 3730 cm-1 is characteristic of a terminal hydroxyl group (OH) (i.e., adsorbed water), while the bands at 3670 and 3590 cm-1 are assigned to the lattice terminal Si-OH groups and acidic hydroxyl Al-OH groups, respectively. According to Vedrine et al. (1979), the terminal Si-OH and the acidic hydroxyl Al-OH groups are particularly important because of their ability to generate Bronsted acid sites which are known to be essential for all cracking reactions. It is therefore desirable to determine how these bands are distributed in the catalyst as a function of potassium concentration. In Figure 4, it is seen that these three bands are very prominent for the nonimpregnated HZSM-5 catalysts. However, as the potassium concentration increased in the catalyst, the IR spectra (Figure 4) showed a substantial decrease in the intensity of all three bands up to a potassium concentration of 0.5 wt % and a complete disappearance of the bands beyond this concentration. These observations confirm the earlier results from TPD analysis that the concentration of acidic hydroxyls (i.e., total acidity) in these catalysts decreased with an increase in potassium concentration. The IR spectra in the pyridine region (1410-1575 cm-1) for HZSM-5 (K0), K1, (0.5 wt % potassium), and K2 (1 wt % potassium) are presented in Figure 5. The

Ind. Eng. Chem. Res., Vol. 35, No. 10, 1996 3337 Table 3. Methanol Conversions and Yields of Various Products from the Methanol Cracking Reaction for given reactn temp 370 °C potassium concn, wt % methanol conversion, % total hydrocarbon yield, wt % dimethyl ether, wt % water, wt % methane, wt % ethane, wt % propane, wt % butanes, wt % ethylene, wt % propylene, wt % butenes, wt % C5-C10 aliphatics, wt % aromatics, wt % C2-C4 olefins, wt %

0 96.0 44.0 0.5 56.0 0.5 0.3 7.2 15.5 3.9 4.4 7.4 34.5 26.3 15.7

0.5 90.0 41.4 1.5 57.1 0.4 0.3 5.6 11.4 4.9 7.4 8.3 36.4 25.3 20.6

1.0 85.0 41.5 2.5 56.0 0.2 0.2 1.0 2.5 8.5 15.5 11.5 41.5 19.1 35.5

400 °C 2.0 80.0 40.5 3.0 56.5 0.2 0.2 1.3 2.8 8.3 14.6 10.8 47.6 14.2 33.7

HZSM-5 catalysts exhibited bands at frequencies of 1545, 1490, and 1445 cm-1. According to Rahman et al. (1988) and Borade and Clearfield (1994), the band at the frequency of 1545 cm-1 represents the Bronsted acid sites, whereas the one at 1490 cm-1 represents the presence of a mixture of both Bronsted and Lewis acid sites. On the other hand, the band at the frequency of 1445 cm-1, which was observed to be broader than the other bands, is characteristic of the Lewis acid sites. Figure 5 shows that as potassium concentration in the catalyst increased, the intensity of the band at 1545 cm-1 (Bronsted acid sites) decreased. On the other hand, an increase in the potassium concentration resulted in a shift in the frequency band for the Lewis sites to higher wavenumbers. This shift was attributed to the formation of Lewis acid sites, which were stronger than those present in nonimpregnated HZSM-5 catalyst. This means that the Bronsted acid sites are more susceptible to alkali metal poisoning than the Lewis acid sites. These results highlight the opposing effects which the incorporation of potassium into HZSM-5 has on the Bronsted and Lewis acid sites. These are the apparent increase in the strength of the Lewis acid sites while decreasing that of the Bronsted acid sites. (E) 1H MAS Solid State NMR. 1H MAS solid state NMR was used to provide quantitative information on both the number of structural hydroxyl groups and the total number of hydroxyl groups present on the catalyst as a function of potassium concentration. 1H MAS NMR spectra of K0 (nonimpregnated HZSM5) and K2 (1.0 wt % potassium) are presented in Figure 6. The figure shows a sharp peak at 6.17 ppm for catalyst K0, whereas the peaks for catalyst K2 (i.e., 1% potassium) were positioned at ca. 5.5 and 2.71 ppm, both with respect to the external tetramethylsilane. It is known (Thomas and Klinowski, 1985) that the chemical shift assigned to 6.1 ppm is due to structural hydroxyl groups, whereas that at 2.7 ppm is due to terminal OH groups and hydroxyl groups attached to extraframework aluminum. On the other hand, the peak at 5.5 represents structural hydroxyl groups which have been weakened due to their chemical environment. Since it is the structural OH group that is responsible for the generation of Bronsted acidity, it therefore implies that the nonimpregnated HZSM-5 catalyst contained mostly Bronsted acid sites, whereas potassium modification of the catalysts as in K2 (1 wt % potassium) changed the acid structure to a combination of Lewis acid and weak Bronsted acid sites. The area under the sharp line gives the number of respective hydroxyl groups present on the catalyst. It

0 100 43.5 0 56.5 1.2 0.8 10.5 14.4 6.5 6.4 6.3 25.5 28.4 19.2

0.5 92.0 42.9 1.0 56.1 0.8 0.5 7.1 10.5 8.5 7.5 8.0 29.7 27.4 24.0

1.0 90.5 43.2 1.8 55.0 0.5 0.4 5.0 3.0 11.2 17.2 9.6 29.9 23.2 38.0

450 °C 2.0 84.5 42.5 2.2 55.3 0.6 0.5 5.9 4.0 10.5 16.6 8.2 37.3 16.4 35.3

0 100 44.1 0 55.9 3.0 1.0 10.5 14.4 7.1 6.8 6.2 19.5 31.5 20.1

0.5 97.0 43.1 0.6 56.3 2.5 0.9 8.5 10.1 10.5 8.9 7.8 22.7 28.1 27.2

1.0 93.0 43.4 1.1 55.5 2.0 0.5 3.5 4.5 13.5 16.5 9.9 25.7 24.2 39.9

2.0 90.0 42.8 1.6 55.6 1.9 0.6 4.2 5.1 12.5 14.9 8.8 32.5 19.5 36.2

Figure 6. 1H MAS NMR spectra of HZSM-5 and 1% potassiumimpregnated HZSM-5 catalysts (K0, HZSM-5; K2, 2% potassium/ HZSM-5).

is seen that HZSM-5 catalyst possesses a higher number of Bronsted acid sites than the potassium-impregnated HZSM-5 catalysts. The NMR results are consistent with those obtained from FT-IR. These results indicate a marked decrease in Bronsted acidity with the incorporation of increasing amounts of potassium. (F) Model Compounds Reaction Studies. Methanol cracking and m-xylene isomerization reactions were used as model reactions to provide further evidence for the degree of acid site poisoning of the catalysts. It is well-established (Chang, 1983) that a high methanol conversion to produce mainly aromatic hydrocarbons in the methanol cracking reaction implies the presence of strong acid sites (especially the Bronsted acid sites) in the catalyst. On the other hand, low methanol conversions to produce dimethyl ether means the presence of weak Bronsted acid sites in the catalyst. In the case of the m-xylene isomerization reaction, a low m-xylene conversion to selectively produce p- or o-xylenes implies the presence of weak acid sites in the catalyst. Table 3 shows methanol conversions and product distributions obtained from methanol cracking over impregnated and nonimpregnated HZSM-5 catalysts. The table shows that methanol conversion and aromatic hydrocarbon yield decreased as the potassium concentration in the catalyst increased (i.e., acid site strength


tions required for the production of maximum amounts of C2-C4 olefins. This is discussed below using eqs 3-9,

canola oil f heavy oxygenated CxHy

(thermal) (3)

heavy oxygenated CxHy f heavy CxHy + H2O + (thermal and catalytic) (4) CO2 + CO heavy CxHy f paraffins + olefins (short and long chain) (catalytic and thermal) (5) Figure 7. Variation of the extent of isomerization with acid site density for m-xylene isomerization.

and density decreased) for all three reaction temperatures used in this work for methanol cracking reactions. On the other hand, the yield of dimethyl ether increased with potassium concentration for all three reaction temperatures. These results support our earlier characterization results, which showed that the acid sites (especially the Bronsted acid sites) diminished with increasing potassium concentration. In the case of m-xylene isomerization, the true isomerization process can be distinguished from side processes such as C2-C4 production by determining what fraction of the m-xylene converted goes to o- or p-xylene formation. This is referred to as the extent of isomerization (I) and is defined as

I ) (p- + o-xylenes produced (wt %))/ (total amt of products (wt %)) (2) Figure 7 shows the variation of the extent of isomerization with the catalyst acid site density. It is clear from this figure that the extent of isomerization increases as the acidity of the catalyst decreases. This is consistent with the literature (Nayak and Choudhary, 1982) and also confirms our earlier characterization results which showed that the incorporation of potassium into HZSM-5 not only dilutes the concentration but also weakens the strengths of the acid sites (especially that of Bronsted acid sites). Catalyst Performance Studies. Canola Oil Conversion and Product Distribution. The canola oil conversion, overall mass balances, and product distributions obtained from canola oil reactions over HZSM-5 and potassium-impregnated HZSM-5 catalysts are given in Table 4a-c as functions of reaction temperature, space velocity, and potassium concentration in the catalyst. These results involved the repetition of a number of runs in order to check for reproducibility. Reproducibility was less than (5 wt %. Table 4 shows that canola oil conversion decreased with increasing potassium concentration (Table 4a) and space velocity (Table 4c) as well as decreasing reaction temperature (Table 4a). Table 4a also shows that the variation of the yield of gas product with these parameters followed a trend similar to the one exhibited by canola oil conversion. On the other hand, there was a general decrease in the yield of OLP with increasing reaction temperature and potassium concentration in the catalyst. These results represent typical yields obtained from catalytic cracking reactions. Optimum Conditions for the Production of C2C4 Olefins. A major objective of this work was to determine the optimum catalyst and operating condi-

light olefins S C2-C10 olefins

(catalytic) (6)

step 4 C2-C10 olefins S aromatic CxHy + aliphatic CxHy

(catalytic) (7)

canola oil f coke n(aromatic CxHy) f coke

(thermal) (catalytic)

(8) (9)

which describe a typical reaction scheme for canola oil conversion (Katikaneni et al., 1995a,b) in conjunction with the effects of catalyst characteristics and reaction operating conditions on the yield of C2-C4 olefins. Here, the first step (eq 3) involves the initial thermal decomposition of canola oil molecules to give heavy oxygenated hydrocarbons (Chang and Wan, 1947; Alencar et al., 1983; Nawar, 1969; Idem et al., 1996), which subsequently undergo deoxygenation (eq 4) to give heavy hydrocarbons. In this work, the mixture of heavy hydrocarbons and heavy oxygenated hydrocarbons is regarded as “residual oil”. The next step involves the secondary cracking of heavy hydrocarbons (eq 5) to give short and long chain paraffins and olefins. In reaction 6, the light olefins typically in the C2-C4 range oligomerize to yield olefins in the C2-C10 range. Both aliphatic (mostly cyclic) and aromatic hydrocarbons are eventually produced in the pores of the zeolite catalysts (eq 7) as a result of cyclization and aromatization reactions of C2-C10 olefins. Coke is formed either due to the condensation of the canola oil molecules (eq 8) or by polymerization of large aromatic hydrocarbons inside the pores of the zeolite catalysts (eq 9) or both. These reactions are dependent on the catalyst acid sites (Prasad et al., 1986a,b; Katikaneni et al., 1995a,b). Also, some of the reaction steps are reversible (eqs 6 and 7). Thus, modifications in catalyst acid site density as well as changes in operating conditions (such as reaction temperature) are certain to have tremendous effect on both canola oil conversion and product distribution and, consequently, on C2-C4 olefins yields. The optimum values of these parameters that result in the production of optimum amounts of C2-C4 olefins are discussed below. (A) Catalyst Acidity. The variation of C2-C4 olefins with potassium concentration (i.e., catalyst acidity) is given in Figure 8. The figure shows that a maximum exists in the relationship between the yield of C2-C4 olefins and catalyst acidity (potassium concentration) for all reaction temperatures. These results can be explained using the reaction steps that involve the formation of C2-C4 olefins from residual oil (eqs 5 and 6) and depletion of these olefins to form aromatic hydrocarbons (eq 7). These results show that starting from the nonimpregnated HZSM-5 (high acidity) catalyst to catalyst K1 or K2 in some cases (i.e., decreasing

Ind. Eng. Chem. Res., Vol. 35, No. 10, 1996 3339 Table 4. Canola Oil Conversions and Yields of Various Products (a) Overall Mass Balances and Yields of Gaseous Products with catalyst K0 400 °C coke gas OLP water residue unacctd coversion methane ethylene ethane propylene propane isobutane n-butane isobutylene other C4d C5+ CO CO2 C2-C4 olefins

450 °C

3.0 25.0 63.0 3.0 3.0 3.0 96.9

4.0 34.0 56.0 1.0 1.0 4.0 98.9

0.3 0.5 0.5 0.9 9.9 1.0 6.2 0.6 2.1 2.7 0.4 4.1

0.3 2.2 0.6 4.4 10.6 2.1 6.1 1.4 1.9 3.2 0.7 0.6 9.9

with catalyst K1

500 °C

400 °C

450 °C

with catalyst K2

500 °C

400 °C

with catalyst K3

450 °C

500 °C

400 °C

450 °C

500 °C

8.0 27.0 54.0 5.0 8.0 2.0 91.8

7.0 42.0 35.0 6.0 6.0 0 94.0

13.0 11.0 44.5 9.0 20.5 2.0 79.1

10.0 21.0 44.5 8.0 15.5 1.0 84.3

9.0 38.0 36.0 5.0 10.0 2.0 89.8

Yields of Gaseous Products, wt % of Canola Oil Fed 1.1 0.2 0.3 1.2 0.6 0.3 5.2 1.2 2.3 5.0 1.8 2.2 1.9 0.3 0.6 1.6 0.7 0.6 7.8 2.9 6.2 12.4 4.2 5.6 16.4 3.8 4.1 4.1 1.1 4.2 2.7 3.1 2.6 1.8 1.0 2.6 5.7 1.7 3.7 5.8 2.4 3.1 3.0 2.5 3.7 3.6 2.1 2.6 1.2 2.6 4.8 0.8 2.8 4.4 3.9 0.7 0.8 2.2 0.8 0.2 0.5 0.7 0.9 0.5 1.0 1.1 1.9 2.4 3.0 0.7 1.1 17.2 6.6 14.8 25.8 8.8 13.3

0.9 5.0 1.4 10.2 6.5 2.7 4.3 3.5 3.6 0.9 1.3 1.7 18.7

0.3 0.8 0.4 2.6 0.8 0.4 1.4 1.0 1.6 0.8 0.6 0.7 6.0

0.7 1.6 1.0 4.7 1.3 0.3 2.4 1.9 2.8 1.1 0.7 2.5 11.0

3.3 3.4 3.4 7.9 2.4 0.2 3.8 3.2 4.7 1.3 1.0 3.5 19.2

Overall Mass Balances, wt % 6.0 5.0 5.0 10.0 21.0 30.0 45.0 18.0 57.0 58.0 39.0 51.0 5.0 3.0 4.0 8.0 10.0 4.0 5.0 11.0 1.0 0 2.0 2.0 89.9 96.0 96.9 88.8

5.0 50.0 41.0 0 0 4.0 100

(b) Yields of Liquid Products, wt % Canola Oil Fed with catalyst K0 methanol acetone aliphatic hydrocarbons benzene toluene xylenes ethylbenzene propylbenzene C9 + aromatics total aromatics unidentified

with catalyst K1

with catalyst K2

with catalyst K3

400 °C

450 °C

500 °C

400 °C

450 °C

500 °C

400 °C

450 °C

500 °C

400 °C

450 °C

500 °C

1.3 1.5 1.0 6.4 18.6 15.1 4.6 3.9 3.2 51.8 7.2

0.8 0.4 0.5 5.8 18.1 16.1 3.4 3.4 2.7 49.5 4.7

0 0 0.1 4.3 14.9 14.4 2.0 2.3 1.3 39.2 1.7

1.5 1.1 3.5 3.0 10.0 11.5 4.2 5.5 6.0 40.2 12.1

2.3 0.6 4.8 4.5 12.9 12.6 3.7 3.6 2.7 40.1 9.6

1.1 0.2 2.4 3.6 9.5 8.7 2.1 1.6 2.1 29.0 7.9

2.0 1.3 15.7 2.0 4.1 4.6 1.6 2.9 8.1 23.3 8.7

1.9 1.1 8.0 4.0 11.8 11.9 3.1 3.5 1.7 36.1 7.1

0.7 0.2 3.7 3.5 8.2 7.6 1.8 1.7 0.9 24.7 3.3

3.3 0.4 24.7 0.7 0.8 0.6 0.7 0.4 1.6 4.9 11.2

2.3 0.2 21.8 1.7 1.6 1.7 0.7 0.4 1.2 7.5 12.6

1.3 0 14.6 1.5 3.1 2.0 0.5 0.6 2.1 9.8 10.3

(c) Yields of Gaseous and Liquid Products 400 °C components methane ethylene ethane propylene propane isobutane n-butane isobutylene other C4d C5+ CO CO2 C2-C4 methanol acetone aliphatic hydrocarbons benzene toluene xylenes ethylbenzene propylbenzene C9 + aromatics total aromatics unidentified

450 °C

500 °C

WHSV ) 3.6 h-1 WHSV ) 1.8 h-1 WHSV ) 3.6 h-1 WHSV ) 1.8 h-1 WHSV ) 3.6 h-1 WHSV ) 1.8 h-1 0.4 1.3 0.3 3.7 1.6 1.3 1.7 1.4 1.7 0.6 0.6 1.4 8.0

Yields of Gaseous Products (wt % of Canola Oil Fed) 0.6 0.2 0.3 1.8 2.1 2.2 0.7 0.4 0.6 4.2 6.0 5.6 1.1 2.5 4.2 1.0 1.9 2.6 2.4 3.3 3.1 2.1 2.4 2.6 0.8 2.9 2.8 2.2 0.6 0.8 0.5 0.8 1.0 0.7 2.1 1.1 8.8 11.2 13.3

1.1 4.0 1.6 10.9 7.3 1.2 5.2 4.3 3.4 1.0 1.1 2.6 23.0

0.9 5.0 1.4 10.2 6.5 2.7 4.3 3.5 3.6 0.9 1.3 1.7 18.7

2.7 1.5 16.9 1.7 3.6 3.7 1.6 2.7 8.2 21.4 9.5

Yields of Liquid Products (wt % of Canola Oil Fed) 2.0 2.4 1.9 1.3 1.1 1.1 15.7 9.9 8.0 2.0 4.0 4.0 4.1 9.1 11.8 4.6 9.0 11.9 1.6 3.2 3.1 2.9 3.5 3.5 8.1 2.3 1.7 23.3 31.2 36.1 8.7 12.4 7.1

1.2 0.5 4.0 2.8 2.8 7.0 1.6 1.3 4.0 23.4 8.9

0.7 0.2 3.7 3.5 8.2 7.6 1.8 1.7 0.9 24.7 3.3

catalyst acidity), the formation of C2-C4 olefins from cracking of residual oil (eq 5) occurs to a greater extent than their oligomerization to C2-C10 olefins (eq 6) and

the subsequent aromatization reactions to produce aromatic hydrocarbons (eq 7). The difference between the extents of residual oil cracking and oligomerization/


Figure 8. Variation of the yields of C2-C4 olefins with potassium concentration for canola oil conversion.

Figure 9. Variation of the yields of C2-C4 olefins with reaction temperature for canola oil conversion.

aromatization reactions increases as catalyst acidity decreases, hence, the net increase in C2-C4 olefins with potassium concentration (i.e., decreasing catalyst acidity). On the other hand, the decrease in the yield of C2-C4 olefins with further increase in potassium concentration (i.e., decrease in catalyst acidity) observed for catalyst K3 (and K2 in some cases) shows that the low consumption of C2-C4 olefins for oligomerization (eq 6) and aromatic hydrocarbon formation (eq 7) does not compensate for the low formation of C2-C4 olefins from residual oil cracking (eq 5). The net result is the decrease in C2-C4 olefin yield for catalyst K3 (and K2 in some cases). These results thus establish the level of acidity required for the production of optimum amounts of C2-C4 olefins. (B) Effect of Reaction Temperature. The variation of the yield of C2-C4 olefins with reaction temperature is given in Figure 9. This figure shows that C2C4 olefins yield increases with reaction temperature for all catalysts. This result is typical of the highly endothermic thermal cracking process which results in an increase in C-C bond scission of C5+ aliphatic hydrocarbons and the consequent increase in the amounts of C2-C4 olefins formed. (C) Effect of Canola Oil Space Velocity. The effect of canola oil space velocity on the combined C2-

C4 olefins yield was evaluated only for catalyst K2 (the catalyst with lowest acid site density in which both weak and strong acid sites are present; see Figure 3). These results are given in Table 4c, which shows that the yield of C2-C4 olefins for runs at high canola oil space velocity (3.6 h-1) was higher than the yield at low space velocity (1.8 h-1) at the reaction temperature of 500 °C. The reverse was the case at lower reaction temperatures. It thus appears from these results that the maximum yield of C2-C4 olefins is obtained from catalyst K2 at 500 °C and 3.6 h-1 space velocity. Therefore, favorable conditions for maximum C2-C4 olefin production are high reaction temperatures, high canola oil space velocities, and catalysts with relatively low Bronsted and total acidity. However, the overall results show that the maximum C2-C4 olefin yield of 25.8 wt % was obtained with catalyst K1 at 500 °C and 1.8 h-1 space velocity. Reaction Pathway for Canola Oil Conversion. It was of interest to postulate a reaction pathway for canola oil cracking over impregnated and nonimpregnated HZSM-5 catalysts. This was performed by using model compound reactions such as methanol cracking and m-xylene isomerization. The model reactions were used to show the nature of involvement of certain reaction intermediates (oxygenates and aromatics) in the reaction pathway for the catalytic conversion of canola oil. This was achieved by comparing the ways in which the conversions and product distributions of model reactions responded to changes in catalyst acidity and reaction temperature with those for canola oil. (A) Methanol Cracking. (1) Effect of Catalyst Acidity. Table 3 shows the conversions as well as the product distributions obtained from methanol cracking reactions over the four catalysts as a function of potassium concentration and reaction temperature. As was shown earlier, methanol conversion decreased as the amount of potassium in the catalyst increased for all reaction temperatures. Also, as was shown earlier by TPD, FT-IR, and NMR results, the role of the incorporation of potassium into HZSM-5 catalyst is that of decreasing both the total and Bronsted acid sites originally present in the nonimpregnated HZSM-5 catalysts. Thus, the decrease in methanol conversion with potassium concentration can be attributed to the decrease in the total and Bronsted acidity of the catalyst. This is consistent with the results of Chang (1983) and Itoh et al. (1984). On the other hand, although the variation in potassium concentration (Bronsted acidity) had a drastic effect on methanol conversion, its effect on the selectivity for hydrocarbon production was negligible. This is seen in Table 3, where the conversion decreased from 100 to 80 wt %, whereas the yield of total hydrocarbons was always within the range 40.5-44.1 wt % for all catalysts and reaction temperatures used. It is wellknown that hydrocarbon formation is highly influenced by the catalyst Bronsted acidity and its shape selectivity characteristics. However, these results for hydrocarbon selectivity suggest that the total yield of hydrocarbons was not affected significantly by catalyst Bronsted acidity. Figure 10 shows the typical effect of catalyst acidity (potassium concentration) on the yields of aromatic and C5-C10 aliphatic hydrocarbons for reactions conducted at 370 °C. The figure shows that while the yield of aromatic hydrocarbons increased with acid site density, that for aliphatic hydrocarbons followed the reverse trend. These results, in conjunction with TPD, FT-IR,


Figure 10. Variation of the yields of aromatic and aliphatic hydrocarbons with catalyst acidity for the methanol cracking reaction.

Figure 11. Variation of the yields of C2-C4 olefins with catalyst acid site density for the methanol cracking reaction.

and NMR results, show conclusively that, unlike in the formation of aliphatic hydrocarbons, high-strength Bronsted acid sites are required for the formation of aromatic hydrocarbons. Figure 11 illustrates the variation of C2-C4 olefins with catalyst acid site density. The combined yield of C2-C4 olefins gives an indication of the extent of both the initial conversion of dimethyl ether to C2-C4 olefins, the oligomerization of C2-C4 olefins to C2-C10 olefins, and the subsequent aromatization of these olefins to produce aromatic hydrocarbons. These steps for methanol cracking are given in eqs 10-13, which describe a typical reaction scheme for methanol cracking to produce hydrocarbons over zeolite catalysts (Chang, 1983).

2CH3OH S CH3OCH3 + H2O

(10)

CH3OCH3 S C2-C4 olefins + CO2 + CO + H2O (11) C2-C4 olefins S C2-C10 olefins

(12)

C2-C10 olefin S aromatic CxHy + aliphatic CxHy (13) A high yield of C2-C4 olefins implies that while dimethyl ether conversion was relatively high, the extents of oligomerization and aromatization were low, and consequently, only a small quantity of the C2-C4 olefins was used up. The converse is also true. How-

ever, Figure 11 shows that the yield of C2-C4 olefins increased initially with acid site density (i.e., decreasing potassium concentration) up to 1 wt % potassium, and then decreased beyond this concentration. The decrease in the yield of C2-C4 olefins with potassium concentration for the catalyst containing greater than 1 wt % potassium can be explained as follows. It is known that there are low extents of oligomerization and aromatization reactions over catalysts with low acid site density (example, K3). Therefore a large fraction of C2-C4 olefins already formed is retained in the case of cracking reactions over such catalysts. However, this retention cannot compensate for the low conversions of dimethyl ether to produce C2-C4 olefins over catalysts containing weak acid sites. These results concerning methanol conversions and product distributions can be combined and used to provide additional reasons why methanol conversions decreased with decreasing catalyst acidity. This is discussed using eqs 10-13. The forward reactions (eqs 10-13) are known to be catalyzed by high catalyst acid site density. It is also known that, for each of these reversible reaction steps, the forward reaction is catalyzed by acid sites to a greater extent than the corresponding backward reaction. Also, the equilibrium for methanol conversions with zeolite catalysts is in favor of aromatic hydrocarbon formation (large methanol conversion) under typical reaction conditions (eq 13). It thus appears from the product distribution in Table 3 that there is also an indirect effect that a decrease in catalyst acidity has on methanol conversion. This indirect effect involves a shift in the equilibrium in reactions 10-13 in favor of a smaller methanol conversion. This can be explained as follows: Table 3 shows that dimethyl ether conversion to olefins increased with catalyst acid site density. It also shows that the subsequent conversion of these olefins to aromatics increased with acid site density, as shown previously. Thus, for a catalyst with high acid site density, the large conversion of C2-C10 olefins to aromatics will require large amounts of olefins to be supplied to replenish the ones used up by the aromatization reactions. The results in Table 3 shows that these C2-C10 olefins will have to be supplied by the conversion of dimethyl ether. Similarly, the large conversions of dimethyl ether to C2-C4 olefins and subsequent oligomerization of these light olefins to produce olefins in the C2-C10 range will induce larger conversions from methanol to supplement the dimethyl ether depleted by the succeeding reaction (eq 11). The net result of these inductive processes is that, for a catalyst with high acid site density and for a system of reversible reactions in series where the forward and corresponding backward step are not catalyzed by acid sites to the same extents, there is a shift in the reaction equilibrium in the direction that will favor the production of a large amount of aromatic hydrocarbons (i.e., a large extent of methanol conversion). Conversely, for catalysts with high potassium concentration (low acid site density), the shift in equilibrium will be in the direction that will favor the production of a small amount of aromatic hydrocarbons (i.e., a small extent of methanol conversion). (2) Reaction Temperature. Table 3 shows that methanol conversion increased with reaction temperature for all four catalysts. This is typical of cracking reactions, as was shown earlier in the case of canola oil conversion.

3342 Ind. Eng. Chem. Res., Vol. 35, No. 10, 1996 Table 5. m-Xylene Conversions and Yields of Various Products for the m-Xylene Isomerization Reactiona potassium concn in HZSM-5 catalyst, wt % m-xylene conversion, wt % C2-C4 hydrocarbons, % benzene, % toluene, % p-xylene, % m-xylene, % o-xylene, % trimethylbenzenes, % a

Figure 12. Variation of the yields of aromatic hydrocarbons with reaction temperature for the methanol cracking reaction.

Figure 13. Variation of the yields of aliphatic hydrocarbons with reaction temperature for the methanol cracking reaction.

The effect of reaction temperature on the yields of aromatic and C5-C10 aliphatic hydrocarbons are presented in Figures 12 and 13, respectively. It is seen from Figure 12 that the aromatic hydrocarbon yield increased with an increase in reaction temperature for all of the catalysts. On the other hand, Figure 13 shows that the variation of the yield of C5-C10 aliphatic hydrocarbons with reaction temperature showed a reverse trend to that exhibited by the yield of aromatic hydrocarbons. These results imply that an increase in temperature results in large extents of cyclization and aromatization reactions and, consequently, a decrease in the yield of C5-C10 aliphatic hydrocarbons. It is seen in Table 3 that the yield of C2-C4 olefins increased with reaction temperature. It was mentioned earlier that, at high reaction temperatures, large amounts of C2-C4 olefins will oligomerize to produce C2-C10 olefins (eq 12), which in turn will aromatize to produce large amounts of aromatic hydrocarbons (eq 13). Generally, this should result in a decrease in the C2C4 olefins. However, high temperatures are also known to favor the C-C scission of C5+ aliphatic hydrocarbons as well as dealkylation of alkylated aromatic and aliphatic hydrocarbons to produce mostly the chemically stable C2-C4 olefins. The result is an increase in the yield of C2-C4 olefins and a consequent decrease in the yield of C5-C10 aliphatic hydrocarbons. (B) Isomerization of m-Xylene. Effects of Catalyst Acidity. The m-xylene conversions and the product yields for the isomerization of xylenes are presented in Table 5 as a function of potassium concentration (i.e., catalyst acidity). The products obtained were o-xylene, p-xylene, toluene, benzene, trimethylbenzenes, and small amounts of C2-C4 hydrocarbons. The presence of C2-C4 hydrocarbons in the m-xylene isomerization

0.0

0.5

1.0

2.0

59.5

47.4

40.9

31.4

0.2 0.4 14.4 21.0 40.5 16.5 7.0

0.2 0.1 10.4 24.8 52.6 7.3 4.6

0.1 0 5.3 25.2 59.1 6.9 3.4

0 0 1.5 23.3 68.6 5.4 1.2

Reactions conducted at 400 °C and 1.8 h-1 space velocity.

products supports our earlier assertion regarding the involvement of the dealkylation reactions of alkylated aromatic hydrocarbons during methanol cracking. Table 5 shows that the conversion of m-xylene increased as the catalyst acid site density increased (i.e., with decreasing potassium concentration in the catalyst). However, an increase in potassium concentration in the catalyst was detrimental to m-xylene conversion; its effect on p-xylene selectivity was highly favorable up to a potassium concentration of 1 wt %. Beyond this potassium concentration, the selectivity to p-xylene started to decrease. On the other hand, the yield of o-xylene decreased progressively with increasing potassium loading. The selectivity for either o- or p-xylene formation can be seen more clearly in Table 6, which shows the variation of the p-xylene/o-xylene ratio (i.e., product selectivity ratio) with potassium concentration (i.e., catalyst acidity). This table shows that the pxylene/o-xylene ratio increased as the catalyst acidity decreased. This means that selectivity for p-xylene decreased with increasing catalyst acidity. These observations are consistent with the findings of Nayak and Choudhary (1982). Apart from C2-C4 hydrocarbons, the other side products obtained from the m-xylene reaction were benzene, toluene, and trimethylbenzene. The concentrations of these side products decreased as the potassium loading increased (i.e., as the catalyst acidity decreased). The type of side products obtained gives an indication of the type of reactions that took place during the conversion of m-xylene over the zeolite catalysts. Generally, these are isomerization, disproportionation, dealkylation, and hydrogen transfer. For example, the presence of trimethylbenzene (TMB) and toluene indicates principally the involvement of disproportionation (eq 14) and dealkylation (eq 15) reactions of m-xylene.

2C8H10 w C7H8 + C9H12 2C8H10 w 2C7H8 + C2H4

(disproportionation) (14) (dealkylation)

(15)

Both reactions 14 and 15 involve the hydrogen transfer reaction. On the other hand, the presence of p- and o-xylene indicates the involvement of the isomerization reaction. The typical product selectivity ratios shown in Table 6 indicate the extents of various side reactions. These are the toluene/o- + p-xylenes ratios, indicating the extent of the dealkylation of m-xylene, and TMB/p+ o-xylenes ratios, indicating the extent of m-xylene disproportionation. Since these ratios were less than 0.38 and were decreasing with decreasing catalyst acidity, it can be concluded that weak acid sites are required for the isomerization of m-xylene while the

Ind. Eng. Chem. Res., Vol. 35, No. 10, 1996 3343 Table 6. Product Selectivity Ratios for Various Products from the m-Xylene Isomerization Reactiona potassium concn in HZSM-5 catalyst, wt %

0.0

0.5

1.0

Product Selectivity Ratio, wt % p-xylene/o-xylene 1.27 3.39 toluene/p- + o-xylenes 0.38 0.32 trimethylbenzene/ 0.19 0.14 p- + o-xylenes toluene/trimethylbenzenes 2.06 2.26 a

Reactions conducted at 400 °C and 1.8 (WHSV).

h-1

2.0

3.65 0.17 0.11

4.31 0.05 0.04

1.55

1.25

space velocity

strong acid sites are required for the disproportionation and dealkylation reactions. Similar conclusions were reached by Nayak and Choudhary (1982) in the case of m-xylene isomerization. (C) Application of Model Reactions to Canola Oil Conversion. The application of model reactions to canola oil conversions are discussed using Table 4ac, which shows canola oil conversions and product distributions for canola oil cracking reactions for all catalysts. (1) Effects of Catalyst Acidity. It is observed from the table that canola oil conversion decreased with increasing potassium concentration (i.e., decreasing acidity) in the catalyst. It is also observed that low canola oil conversions were obtained only under conditions where high yields of residual oil were obtained in the product (see Table 4a). This indicates that although it was possible for the initial cracking of canola oil to residual oil (a mixture of heavy hydrocarbons and heavy oxygenated hydrocarbons) to occur in thermal runs (Chang and Wan, 1947; Alencar et al., 1983; Nawar, 1969; Idem et al., 1996) as well as with catalysts of low acidity (example, catalyst K3), high acidity appears to be essential for subsequent deoxygenation, oligomerization, and aromatization reactions. According to Prasad et al. (1986a,b) and Katikaneni et al. (1995a,b), the cracking reactions of hydrocarbons or unsaturated triglyceride molecules are favored by catalysts with high strength and density of acid sites, such as zeolites and silica-alumina. Our results show that the neutralization of these acid sites with potassium resulted in a decrease in the extent of the cracking of residual oil molecules and not of the long and bulky triglyceride molecules (canola oil). This was evidenced by the presence of some high molecular weight materials having boiling points higher than 350 °C in the product and not of the triglyceride molecules themselves. Thermal cracking of plant oils was observed to occur by Chang and Wan (1947), Alencar et al. (1983), and Idem et al. (1996) at temperatures g300 °C. These workers carried out their investigations using batch and fixed-bed reactors in the temperature range 300-500 °C. In the present study, the canola oil feed entered the reactor from the top of the reactor where the temperature was essentially the same as the reaction temperature (400-500 °C). There was also a long preheating zone (≈160 mm). This was the space inside the reactor between the top of the reactor and the top of the catalyst bed. On the basis of the total volume of the reactor and also, based on canola oil vapor at standard temperature and pressure (25 °C and 1 atm), the calculated residence times of canola oil vapor through the reactor are ca. 13.6 and 27.2 min for canola oil WHSV values of 3.6 and 1.8 h-1, respectively. Also on the basis of uncracked canola oil vapor and the

Figure 14. Variation of the turnover number for canola oil conversion with catalyst acid site density.

volume of the preheating zone, the times it woud take for the feed vapor to reach the catalyst bed are 5.5 and 10.9 min for canola oil WHSV values of 3.6 and 1.8 h-1, respectively. However, it was observed in all of our experimental runs that the gas product appeared in less than 1 min from the start of the feed pump. This observation therefore makes the calculated residence times meaningless. These much shorter experimentally observed residence times are attributed to the initial thermal decomposition of each canola oil molecule into a number of relatively small molecules in the preheating zone and the subsequent vaporization of the decomposed molecules before catalytic reactions. Also, decomposition must happen as soon as canola oil enters the preheating zone in order to achieve the short experimentally observed residence times. This result of the initial thermal cracking is in contrast with the proposal of Weisz et al. (1979) which indicated that triglyceride molecules have to enter the pores of the zeolite catalyst before they undergo cracking reactions. Thus, our present results show that triglyceride molecules initially undergo thermal cracking before subsequent diffusion of cracked products into the pores of the zeolite catalyst for catalytic cracking reactions. It is highly desirable to verify whether catalyst performance for canola oil conversion (i.e., the extent of cracking of residual oil) can be attributed exclusively to catalyst acidity. This verification can be performed by evaluating the relationship between the turnover number for canola oil conversion and catalyst acid site density. In the literature, turnover number is defined in terms of the rate of reaction per active site. However, because of the complexity of the reaction/catalyst system used in the present study, we have decided to define turnover number in terms of canola oil conversion as follows:

TON ) (canola oil converted (wt %))/ (catalyst acid site density (mm2/m2)) ) (wt % canola oil converted)/(mm2/m2) ) turnover number for canola oil conversion (16) The relationship between TON and acid site density is shown in Figure 14. This figure shows that the turnover number for canola oil conversion decreases with an increase in the acid site density. Thus, although an


Figure 17. Variation of the yields of C2-C4 olefins with catalyst acidity for canola oil conversion. Figure 15. Variation of the yields of aromatic hydrocarbons with catalyst acidity for canola oil conversion.

Figure 16. Variation of the yields of aliphatic hydrocarbons with catalyst acidity for canola oil conversion.

increase in catalyst acidity results in an increase in canola oil conversion (see Table 4a), the effectiveness of these acid sites decreases as the acid site density increases. Also, as can be seen in Table 4a, there was some level of canola oil conversion even with the catalyst whose acid sites were almost completely neutralized by potassium impregnation (i.e., catalyst K3). These results show that canola oil conversion does not depend entirely on catalyst acidity but also on other factors such as temperature. Figures 15 and 16 show respectively the variation of total aromatic and aliphatic hydrocarbon yields with catalyst acidity. It is seen from Figure 15 that there was a drastic reduction in total aromatic hydrocarbon yield with a decrease in catalyst acidity (i.e., increase in potassium concentration). On the other hand, Figure 16 shows that total aliphatic hydrocarbons increased drastically with decreasing catalyst acidity. As in the case of methanol cracking, these results, in conjunction with TPD, FT-IR, and NMR results, show that highstrength Bronsted acid sites are required for the formation of aromatic hydrocarbons. These results establish a strong similarity between methanol cracking and canola oil conversion. In can be seen from eq 7 that there is a competition for the consumption of C2-C10 olefin intermediates between the formation of aromatic hydrocarbons and that of aliphatic hydrocarbons. It appears from the results that weak acid sites favor the

formation of aliphatic hydrocarbons in preference to aromatic hydrocarbon formation. Figure 17 illustrates the variation of C2-C4 olefins with catalyst acid site density. As in the case of methanol cracking, the yield of C2-C4 olefins gives an indication of the extent of both the initial conversion of heavy hydrocarbons and heavy oxygenated hydrocarbons to C2-C4 olefins and subsequent oligomerization and aromatization of these olefins to produce aromatic hydrocarbons. In the case of canola oil cracking, a high yield of C2-C4 olefins implies that while the conversion of heavy hydrocarbons and heavy oxygenated hydrocarbons was relatively high, the extents of oligomerization and aromatization were low. Consequently, only a small quantity of the C2-C4 olefins was used up. The converse is also true. However, Figure 17 shows that the yield of C2-C4 olefins increased with acid site density (i.e., decreasing potassium concentration) up to 0.5 wt % potassium (catalyst K1) at 400 °C and 1 wt % potassium (catalyst K2) at 450 and 500 °C and then decreased slightly beyond these potassium concentrations at the specified temperatures. The decrease in yield of C2-C4 olefins beyond these potassium concentrations is explained as follows: for catalysts with high potassium concentrations (i.e., low catalyst acid site densities) the high yield of C2-C4 olefins which is due to the decrease in oligomerization and aromatization reactions does not compensate for the decrease in the production of C2-C4 olefins from the low conversions of heavy hydrocarbons and heavy oxygenated hydrocarbons. Similar results were obtained for methanol cracking reactions. (2) Reaction Temperature. Figures 15 and 16 are used to respectively illustrate the variation of aromatic and C5-C10 aliphatic hydrocarbons with reaction temperature. It is seen from Figure 15 that the total yield of aromatic hydrocarbons increased with reaction temperature for all catalysts. This implies an increase in cyclization and aromatization reactions with reaction temperature. On the other hand, Figure 16 shows that the total yield of aliphatic hydrocarbons decreased with a decrease in reaction temperature, signifying a decrease in the cyclization and aromatization reactions. These results also highlight the competition which exists for the consumption of the available C2-C10 olefins with regard to reaction temperature. The competition is for the formation of either aromatic or aliphatic hydrocarbons. This scenario was also observed for methanol cracking. Figure 16 shows the variation of the yield of C2-C4 olefins with reaction temperature. It is seen from this


figure that the yield of C2-C4 olefins increased with reaction temperature for all catalysts, as was shown earlier. At high reaction temperatures large amounts of C2-C4 olefins will oligomerize and then aromatize to produce large amounts of aromatic hydrocarbons (eqs 7 and 13). Generally, this should lead to a decrease in the C2-C4 olefins. However, high temperatures are also known (Mcketta, 1992) to favor the C-C scission of C5+ aliphatic hydrocarbons as well as the disproportionation and dealkylation of some aromatic hydrocarbons (see m-xylene reactions in Table 5), to produce mostly the chemically stable C2-C4 olefins. The result is an increase in the yield of C2-C4 olefins and a consequent decrease in the yield of C5-C10 aliphatic hydrocarbons. Similar results were obtained for methanol cracking. Parts b and c of Table 4 show the various product yields in the organic liquid product (OLP). The tables show that the yields of toluene and trimethylbenzene (represented by C9+ aromatic hydrocarbons) were large, implying that there were large extents of dealkylation and disproportionation reactions involving a number of aromatic hydrocarbon compounds such as xylene. It is also seen in the tables that the yields of toluene and the C9+ aromatic hydrocarbons decreased tremendously with an increase in potassium concentration (i.e., decreasing catalyst acidity), showing that dealkylation and disproportionation reactions decreased as the acid site density of the catalyst decreased. This corroborates the results obtained earlier for m-xylene isomerization reactions. Again, this shows that reactions similar to those of m-xylene are involved in canola oil conversion over HZSM-5 and K/HZSM-5 catalysts. Results of model compound reactions compared with those from canola oil conversion have shown clearly that cracking, deoxygenation, isomerization, aromatization, dealkylation, disproportionation, and oligomerization reactions are involved in canola oil conversion over potassium-impregnated catalysts. They have also shown that the effect of modifying the acidity of the catalysts on the canola oil reaction scheme is essentially that of changing both the overall equilibrium conversion and equilibrium conversions of individual acid catalyzed reaction steps. Conclusions (1) The incorporation of potassium into HZSM-5 catalyst affected the strength and density of both the total and Bronsted acid sites, thus dramatically altering the product distribution. (2) Thermal cracking of canola oil was the initial step, and it was found to be independent of catalyst acidity. On the other hand, subsequent reaction steps such as deoxygenation of oxygenated hydrocarbons, secondary cracking, oligomerization, aromatization, disproportionation, dealkylation, and coke formation were a strong function of catalyst acidity. (3) Within the reaction temperature range used in this study, the highest yields of aromatic hydrocarbons were obtained with catalysts K0, which contained strong acid sites. (4) Lowering the acidity changed the product distribution from aromatics to aliphatics. Thus, catalyst K3, containing weak acid sites, resulted in the production of the highest yields of aliphatic hydrocarbons. (5) The maximum C2-C4 olefin yield of 25.8 wt % was obtained at 500 °C at a space velocity (WHSV) of 1.8 h-1 with catalyst K1 containing relatively low Bronsted and total acidity.

Acknowledgment The financial support of the Natural Sciences and Engineering Research Council of Canada (NSERC) is gratefully acknowledged. Literature Cited Adjaye, J. D.; Bakhshi, N. N. Catalytic conversion of a biomassderived oil to fuels and chemicals I: Model compound studies and reaction pathways. Biomass Bioenergy, 1995, 8, 131-149. Alencar, J. W.; Alves, P. B.; Craveir, A. A. Pyrolysis of tropical vegetable oils. J. Agric. Food. Chem. 1983, 31, 1268-1270. Baker, E. G.; Elliott, D. C. Catalytic upgrading of biomass pyrolytic oils. In Research in thermochemical biomass conversion; Bridgwater, A. V., Ed.; ACS Division of Fuel Chemistry, Elsevier Applied Science: London, 1987; Vol. 32, pp 883-895. Bhatia, S. Zeolite Catalysis: Principles and Applications; CRC Press: Boca Raton, FL, 1990. Boocock, D. G. B.; Konar, S. K.; Mackay, A.; Cheung, P. T. C.; Liu, J. Fuels and chemicals from sewage sludge. 2. The production of alkanes and alkenes by the pyrolysis of triglycerides over activated alumina. Fuel 1992, 71, 1291-1297. Borade, R. B.; Clearfield, A. A comparative study of acidic properties of SAPO-5, -11, -34 and -37 molecular sieves. J. Mol. Catal. 1994, 88, 249-266. Campbell, I. M. Biomass, catalysts, and liquified fuels; Holt, Rinehart and Winston: London, 1983. Chang, C. D. Hydrocarbons from methanol. Chemical Industries; Dekker: New York, 1983; Vol. 10. Chang, C.; Wan, S. China’s motor fuel from tung oil. Ind. Eng. Chem. 1947, 39, 1543-1548. Chantal, P.; Kaliaguine, S.; Grandmaison, J. L.; Mahay, A. Production of hydrocarbons from aspen poplar pyrolytic oils over HZSM-5. Appl. Catal. 1984, 10, 317-32. Chen, N. Y.; Miale, J. N.; Reagan, W. J. Preparation of zeolites: Example 5. U.S. Patent 4112056, 1973. Chen, N. Y.; Garwood, W. E.; Dwyer, F. G. Shape selective catalysts in industrial applications; Dekker: New York, 1989. Chu, C. T.; Chang, C. D. Methanol conversion to olefins over ZSM5, II. Olefin distribution. J. Catal. 1984, 86, 297-300. Chu, C. T.; Kuehl, G. H.; Lago, R. M.; Chang, C. D. Isomorphous substitution in zeolite frameworks, II. Catalytic properties of [B]ZSM-5. J. Catal. 1985, 93, 451-458. Craig, W.; Coxworth, E. Conversion of vegetable oils to conventional liquid fuel extenders. In Proceedings of the Sixth Canadian Bioenergy R&D Seminar; Granger, C., Ed.; Elsevier Applied Science: London, 1987; pp 407-411. Gary, J. H.; Handwerk, G. E. Petroleum refining technology and economics; Dekker: New York, 1984. Haag, W. O.; Rodewald, P. G.; Weisz, P. B. Catalytic production of aromatics and olefins from plant materials. ACS Symp. Ser. 1980, 25, 650-656. Hidalgo, C. V.; Itoh, H.; Hattori, T.; Niwa, M.; Murakami, Y. Measurement of acidity of various zeolites by temperatureprogrammed desorption of ammonia. J. Catal. 1984, 85, 362369. Idem, R. O.; Bakhshi, N. N. Production of hydrogen from methanol. 1. Catalyst characterization studies. Ind. Eng. Chem. Res. 1994, 33, 2047-2055. Idem, R. O.; Katikaneni, S. P. R.; Bakhshi, N. N. Thermal cracking of canola oil: Reaction products in the presence and absence of steam. Energy Fuels 1996, in press. Itoh, H.; Hidalgo, C. V.; Hattori, T.; Niwa, M.; Murakami, M. Role of acid property of various zeolites in the methanol conversion to hydrocarbons. J. Catal. 1984, 85, 521-526. Kaeding, W. W.; Butter, S. A. Production of chemicals from methanol. I. Low molecular weight olefins. J. Catal. 1980, 61, 155-164. Katikaneni, S. P. R.; Adjaye, J. D.; Bakhshi, N. N. Catalytic conversion of canola oil to fuels and chemicals over various cracking catalysts. Can. J. Chem. Eng. 1995a, 73, 484-498. Katikaneni, S. P. R.; Adjaye, J. D.; Bakhshi, N. N. Studies on the Catalytic Conversion of Canola Oil to Hydrocarbons: Influence of Hybrid Catalysts and Steam. Energy Fuels 1995b, 9, 599609. Katikaneni, S. P. R.; Adjaye, J. D.; Bakhshi, N. N. Conversion of canola oil to various hydrocarbons over Pt/HZSM-5 bifunctional catalyst. Can. J. Chem. Eng. 1996, submitted for publication.

3346 Ind. Eng. Chem. Res., Vol. 35, No. 10, 1996 McKetta, J. J. Petroleum Processing Handbook; Dekker: New York, 1992. Milne, T. A.; Evans, R. J.; Nagle, N. Catalytic conversion of micro algae and vegetable oils to premium gasoline with shapeselective zeolites. Biomass 1990, 21, 219-232. Nawar, W. W. Thermal degradation of lipids. A review. J. Agric. Food. Chem. 1969, 17, 18-21. Nayak, V. S.; Choudhary, V. R. Isomerization of m-xylene on HZSM-5: Part I. Influence on catalytic activity and selectivity of Si/Al ratio, degree of cation exchange, deammonation conditions, and poisoning of stronger acid sites. Appl. Catal. 1982, 4, 333-352. Olson, D. H.; Haag, W. O.; Lago, R. M. Chemical and Physical Properties of the ZSM-5 Substitutional Series. J. Catal. 1980, 61, 390-396. Prasad, Y. S.; Bakhshi, N. N.; Mathews, J. F.; Eager, R. C. Catalytic Conversion of Canola Oil to Fuels and Chemical Feedstocks, Part I. Effect of Process Conditions on the Performance of HZSM-5 Catalyst. Can. J. Chem. Eng. 1986a, 64, 278284. Prasad, Y. S.; Bakhshi, N. N.; Mathews, J. F.; Eager, R. L. Catalytic conversion of canola oil to fuels and chemical feedstocks, part II. Effect of co-feeding steam of the performance of HZSM-5 catalysts. Can. J. Chem. Eng. 1986b, 64, 285-292. Rahman, A.; Lemay, G.; Adnot, A.; Kaliaguine, S. Spectroscopic and Catalytic Study of P-modified ZSM-5. J. Catal. 1988, 112, 453-463. Sharma, R. K.; Bakhshi, N. N. Catalytic Conversion of Biomass Derived Oils to Fuels and Chemicals over HZSM-5: Effect of Co-feeding Steam. Fuel Process. Technol. 1991, 27, 113-130.

Szostak, R. Molecular Sieves: Principles of Synthesis and Identification; Van Nostrand Reinhold: New York, 1989. Thomas, J. M.; Klinowski, J. The study of aluminosilicate and related catalysts by high resolution solid state NMR spectroscopy. Adv. Catal. 1985, 33, 200-374. Vedrine, J. C.; Auroux, A.; Bolis, V.; Dejaifve, P.; Naccache, C.; Wierzchowski, P.; Derouane, E. G.; Nagy, J. R.; Gilson, J. P.; Jan, H. C.; van Hooff; Berg, J. P. V. D.; Wolthuizen, J. Infrared, Microcalorimetric, and Electron Spin Resonance Investigations of the Acidic Properties of the HZSM-5 Zeolite. J. Catal. 1979, 59, 248-262. Waddama, R. D. Petroleum chemical products and their applications. Chemicals from Petroleum; Gulf: London, 1980. Weisz, P. B.; Haag, W. O.; Rodewald, P. G. Catalytic Production of High-Grade Fuel (Gasoline) from Biomass Compounds by Shape Selective Catalysts. Science 1979, 206, 57-58. Zatorski, L. W.; Wierzchowski, P. T.; Cichowlas, A. A. Conversion of Methanol to Lower Olefins-Application Studies. Acta Phys. Chem. 1985, 31, 515-523.

Received for review December 6, 1995 Revised manuscript received May 16, 1996 Accepted June 10, 1996X IE950740U

X Abstract published in Advance ACS Abstracts, August 15, 1996.

Catalytic Conversion of Canola Oil over Potassium-Impregnated

Recommend Documents