Catalytic Conversion of Canola Oil over Potassium-Impregnated HZSM-5 Catalysts: C 2 −C 4 Olefin Production and Model Reaction Studies

Catalytic Conversion of Canola Oil over Potassium-ImpregnatedHZSM-5 Catalysts: C2-C4 Olefin Production and Model ReactionStudies

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 SciencePlace, University of Saskatchewan, Saskatoon, Canada S7N 5C9

The influence of catalyst acidity, reaction temperature, and canola oil space velocity on theconversion of canola oil was evaluated using a fixed-bed microreactor at atmospheric pressureat reaction temperatures and space velocities (WHSV) in the ranges 400-500 °C and 1.8-3.6h-1, respectively, over potassium-impregnated HZSM-5 catalysts. These catalysts were thor-oughly 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 ofC2-C4 olefins from canola oil were determined. The incorporation of potassium into HZSM-5catalyst resulted in both the dilution and poisoning of Bronsted and total acid sites. Theseacidity changes only severely affected the acid catalyzed reactions, such as oligomerization andaromatization, and resulted in drastic modifications in product distribution. The maximumC2-C4 olefin yield of 25.8 wt % was obtained at 500 °C and 1.8 h-1 space velocity with catalystK1 of relatively low Bronsted and total acidity.

Introduction

Paraffins, olefins, naphthenes, and aromatics areimportant groups of organic compounds. They areuseful in many chemical and petrochemical industriesfor the manufacture of a wide range of products. Theolefins are a particularly interesting group of hydrocar-bons with unique chemistry characterized by relativelyhigher reactivity than paraffins, aromatics, or naph-thenes. Consequently, they are used in various impor-tant industrial processes such as alkylation, polymeri-zation, oligomerization, and aromatization reactions(Gary and Handwerk, 1984).Traditionally, hydrocarbons are produced from petro-

leum sources (Waddama, 1980). However, recent stud-ies (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; Adjayeand Bakhshi, 1995; Sharma and Bakhshi, 1991) haveshown that they can also be obtained from other sources,such as from the catalytic conversion of various typesof plant oils. According to these authors, such studiesinvolve passing these oils at temperatures in the range350-550 °C over cracking catalysts such as HZSM-5,Pt/ZSM-5, silica-alumina, H-Y, H-modernite, alumi-num pillared clays, and hybrids of these catalysts. Also,in most of these studies, hydrocarbons in the gasolineboiling range were predominant in the organic liquidfraction, whereas the gaseous fraction contained mostlyparaffinic hydrocarbon components.There is a considerable demand for lower olefins such

as ethylene, propylene, 1- and 2-butenes, and isobuty-lene because of the growing interest in their applicationin the manufacture of desirable products such aspolyethylene, polypropylene, methyl tert-butyl ether(MTBE), and ethyl tert-butyl ether (ETBE). Our inter-est is in the production of these olefins from plant oils.Considerable information exists in the literature on thecatalyst characteristics and operating conditions re-quired for the production of optimum amounts of C2-

C4 olefins from petroleum sources. According to Olsonet 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 catalystacidity (especially the Bronsted acidity), a catalyst withrelatively small pore sizes, short contact times betweenthe feed and catalyst, and high reaction temperatures.On the other hand, the chemical characteristics of plantoils are markedly different from those of petroleum oil.Thus, it is highly desirable to determine the conditionsfor the production of optimum amounts of C2-C4 olefinsfrom the catalytic conversion of plant oils.In the literature, studies on the modification of

catalyst acidity have been reported mostly for crackingreactions involving feed from petroleum sources. Thesemodifications have been achieved principally either byincorporating phosphorus, boron, or an alkali metal orby changing the silica/alumina ratio (Chang, 1983;Szostak, 1989; Bhatia, 1990; Chu et al., 1985; Rahmanet al., 1988). Recently, studies have been reported onthe modification of the acid sites of HZSM-5 catalystby the incorporation of platinum (Katikaneni et al.,1996), mainly to evaluate the effects of catalyst acidityon the yields of various products from canola oil crack-ing. These studies showed that the impregnation ofHZSM-5 catalyst with platinum resulted in an increasein the isomerization function of the catalyst. Forexample, there was an increase in the amount ofisoparaffins formed using Pt/HZSM-5 catalysts com-pared with the use of nonimpregnated HZSM-5 catalyst.Also, it was shown in the study that impregnation ofHZSM-5 catalyst with platinum only resulted in thedilution of the acid site density of the catalyst. On theother hand, it is well-known that impregnation with analkali metal will result both in the dilution and poison-ing of the catalyst acid sites. However, the consequenceof alkali metal impregnation on the yield of C2-C4olefins or yields of other products from the conversionof plant oils has not been reported so far in theliterature.It is desirable to obtain a better understanding of the

canola oil conversion process itself. Usually, this is done* Author to whom all correspondence should be addressed.

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

S0888-5885(95)00740-8 CCC: $12.00 © 1996 American Chemical Society

in conjunction with studies of the characteristics of thecatalyst. A number of workers (Katikaneni et al.,1995a,b; Prasad et al. 1986a,b) have proposed variousreaction pathways for canola oil conversion, and thesehave been based primarily on the trends observed inthe product distribution. In both the thermal andcatalytic conversions, the intermediate reaction stepsinvolved include the deoxygenation of oxygenated hy-drocarbon intermediates such as alcohols, ketones,aldehydes, and organic acids (Chang and Wan, 1947;Nawar, 1969). In the case of zeolite catalysts, subse-quent steps include isomerization reactions involvingaromatic compounds (such as p-, m-, and o-xylenes),paraffins (such as n-butane), and olefins (such asbutylene). In the literature (Adjaye and Bakhshi, 1995),the involvement of intermediate reactions in a proposedreaction pathway can be verified through the use ofmodel compound reaction studies. Usually this involve-ment is confirmed if the variation in product distributionwith operating conditions and catalyst characteristicsexhibited by model reactions are consistent with thatexhibited by the main reaction of interest (canola oilcracking) conducted under identical conditions (i.e.,catalyst acidity, feed space velocity, and reaction tem-perature).In this work, we have studied the effects of the degree

of acid site dilution and poisoning of HZSM-5 catalystson the yields of C2-C4 olefins from canola oil conversion.We have also evaluated the optimum catalyst charac-teristics and reaction operating variables required formaximum production of lower olefins from canola oil.This work also involved the preparation of potassium-impregnated HZSM-5 catalysts with potassium concen-trations in the range 0-2 wt %. It also involved thethorough characterization of these catalysts using tech-niques such as X-ray diffraction (XRD), N2 adsorptionmeasurements, 1H MAS solid state NMR, temperatureprogrammed desorption (TPD) of NH3, FT-IR studies,and model compound reactions in order to obtain abetter understanding of the relationship between cata-lyst performance and catalyst characteristics. Theresults have been used to elucidate the reaction pathwayfor canola oil cracking over potassium-impregnatedHZSM-5 catalysts. These results are presented in thispaper.

Experimental Section

Catalyst Preparation. The potassium-impregnatedHZSM-5 catalysts (K/HZSM-5) used in this work wereobtained by incorporating potassium into HZSM-5catalyst. The nonimpregnated HZSM-5 catalyst wasprepared according to the procedure reported by Chenet al. (1973). Altogether, four K/HZSM-5 catalysts ofpotassium concentrations in the range 0-2 wt % wereprepared. Potassiumwas incorporated into the HZSM-5catalyst by wet impregnation techniques using aqueoussolutions of potassium carbonate (analytical grade andobtained from BDH Chemicals, Toronto, Canada) as theprecursor. This impregnation step was carried out for24 h. After this step, the catalysts were dried overnightat 100 °C and then calcined at 500 °C for 6 h in a mufflefurnace. The resulting potassium-impregnated HZSM-5catalysts were designated as K0 (0 wt % potassium),K1 (0.5 wt % potassium), K2 (1 wt % potassium), andK3 (2 wt % potassium) catalysts.Catalyst Characterization. Catalyst characteriza-

tion involved the determination of both the physical andchemical characteristics of the catalysts.

(A) Physical Property Characterization. (1) N2Adsorption Measurements. The surface area, porevolume, pore size, and pore size distribution of all ofthe catalysts were determined using a Micromeriticsadsorption equipment (Model ASAP 2000) which wasequipped with a micropore analysis program usingnitrogen (99.995% purity; obtained from Linde, Calgary,Canada) as the analysis gas. Prior to analysis, eachcatalyst was evacuated at 300 °C at a vacuum of 0.54KPa for 10 h. HZSM-5 based catalysts contains mostlymicropores. Thus, the surface areas measured were themicropore surface areas. Therefore, the Dubinin-Astakhov programwas used in evaluating the microporesurface areas, while the Hovarth-Kawazoe equationwas used to estimate the median pore sizes of bothHZSM-5 and K/HZSM-5 catalysts.(2) Powder X-ray Diffraction Measurements.

Powder XRD measurements were performed to identifycomponent phases as well as to determine the degreeof crystallinity of the catalysts as a function of potas-sium concentration. The XRD measurements weremade with a Phillips diffractometer using Fe KR radia-tion in the scanning angle (2θ) range of 10-90° at ascanning speed of 2 deg/min.(B) Chemical Property Characterization. The

chemical property characterization techniques wereemployed principally for the determination of the acidityof the catalysts and are described below.(1) Temperature Programmed Desorption of

Ammonia. TPD analysis was performed in order todetermine the acid strengths and distributions on thecatalysts. The analysis was conducted in a conventionalflow system similar to the one described by Idem andBakhshi (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 andalso obtained from Linde) was used as the adsorbinggas. About 0.5 g of the calcined catalyst sample wasused for each experiment. NH3 adsorption was per-formed by flowing 1% NH3 in N2 gas over the catalystfor 1 h. Adsorption was carried out at 100 °C in orderto eliminate physically adsorbed NH3. Temperatureprogrammed desorption of NH3 started at 100 °C andended at 650 °C at a temperature programming rate of8 °C/min.(2) FT-IR Measurements. The FT-IR technique

was employed to identify the nature of acid sites presenton the catalyst samples. The IR measurements weremade on powdered catalyst samples using a BioradInfrared Spectrometer (Model FTS 40, Digilab Division).Two regions of the IR spectra were explored. Thesewere the hydroxyl and pyridine regions. Consequently,fresh and pyridine chemisorbed catalyst samples wereused to obtain the spectra for hydroxyl (3500-4000cm-1) and pyridine (1400-1650 cm-1) regions, respec-tively. Pyridine chemisorbed samples were obtained bypassing pyridine vapor over the catalysts at 150 °C for1 h in the same flow system that was used previouslyfor NH3 adsorption. After the pyridine adsorption, eachsample was allowed to cool to room temperature andsubsequently used for IR analysis.(3) Solid State NMR Studies. 1H MAS solid state

NMR studies were carried out in order to determine thenumber of hydroxyl groups present on the catalyst.These experiments were performed using a Bruker AM360 WB instrument equipped with a CP/MAS facility(available at the National Research Council Laboratory,Saskatoon, Canada). The instrument was operated at

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

a frequency of 360.13 MHz and a magnetic field of 7.05T. Other operational parameters were a pulse widthand frequency of 6.5 s and 4000 Hz, respectively. Alloperations were performed at room temperature, whereasall of the chemical shifts were measured relative to thatof tetramethylsilane (TMS).Catalyst Performance Studies. (A) Equipment.

The performance of the potassium-impregnated andnonimpregnated HZSM-5 catalysts was studied in astainless steel (SS 316) fixed-bed (down flow) microre-actor (11.5 mm id and 400 mm overall length). Detailsconcerning the experimental rig used are given else-where (Katikaneni et al., 1995a,b). The furnace tem-perature was controlled by a series SR22 microprocessor-based autotuning PID temperature controller (ShimadenCo. Ltd., Tokyo, Japan) using a K-type thermocoupleplaced on the furnace side of the annulus between thefurnace and the reactor. A separate thermocouple wasused to monitor the temperature of the catalyst bed.This arrangement was capable of ensuring an accuracyof (1 °C for the catalyst bed temperature.(B) Canola Oil Conversion. Test runs were per-

formed for each catalyst at atmospheric pressure atreaction temperatures in the range 400-500 °C andspace velocities (WHSV) ranging from 1.8-3.6 h-1. Thecanola 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 molecularformula of C59H94O5. The composition of canola oil isgiven in Table 1.Each experimental run required 2 g of the catalyst.

A typical test run was conducted as follows: A plug ofglass wool (about 0.5 g), which was used as screen aswell as support for the catalyst, was placed on astainless steel grid positioned centrally within thereactor. The catalyst was then loaded on the glass wool.The catalyst bed was heated to the desired reactiontemperature in flowing argon gas. When this temper-ature was reached, the argon flow was stopped and thefeed was pumped into the reactor at the desired canolaoil space velocity. The product mixture leaving thereactor was condensed in a water-cooled heat exchanger,followed by an ice-cooled condenser to separate gaseousand liquid products for separate analysis. The gasproduct was collected over saturated brine, while theliquid product was collected in a glass trap positionedafter the condenser.The scheme for the estimation of the amounts of coke,

residual oil, organic liquid product (OLP), gas, andwater formed for each canola oil upgrading run is givenin Figure 1. A detailed description of the procedures isgiven elsewhere (Katikaneni et al., 1995a).(C) Model Compound Reactions. (1) Methanol

Cracking andm-Xylene Isomerization Reactions.Methanol cracking and m-xylene isomerization were

Table 1. Composition of Canola Oil

composition chemical formula wt %ano. of

carbon atomsno. of carbondouble bonds

oleic acid CH3(CH2)7CHdCH(CH2)7COOH 60 18 1linoleic acid CH3(CH2)4CHdCHCH2CHdCH(CH2)7COOH 20 18 2linolenic acid CH3CH2CHdCHCH2CHdCHCH2CHd(CH2)7 COOH 10 18 3stearic acid CH3(CH2)16COOH 2 18 0palmitic acid CH3(CH2)14COOH 4 16 0a 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.

3334 Ind. Eng. Chem. Res., Vol. 35, No. 10, 1996

used as model compound reactions to provide furtherevidence for the strengths of the various types of acidsites present on the catalysts. They were subsequentlyused for the elucidation of the reaction pathway forcanola oil conversion. Both reactions were performedin the same fixed-bed reactor used for canola oil conver-sion at atmospheric pressure at reaction temperaturesin the range 370-450 °C and space velocity (WHSV) of1.8 h-1. The methanol andm-xylene used were analyti-cal 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 aCarle 500 GC using a combination of packed andcapillary columns and both flame ionization and ther-mal conductivity detectors. The sample was injectedinto a 2 m long precolumn packed with OV-101 siliconeoil. All components lighter than ethane passed rapidlythrough this column and were separated at 70 °C in aPorapak Q column in series with a molecular sieve 13Xcolumn (each 2.7 m in length) into CO, CO2, and C1 andC2 hydrocarbons. These gases were detected using athermal conductivity detector (TCD). The C3

+ compo-nents were back-flushed into the fused silica capillarycolumn for which the oven temperature was pro-grammed from 40 to 200 °C. The C3

+ components weredetected using a flame ionization detector (FID). Theorganic liquid product (OLP) was also analyzed with aCarle 500 GC which was equipped with a bondednonpolar (methyl silicone) 50 m × 0.2 mm id column.The temperature in the GC oven was programmed from40 to 200 °C, and detection was done using FID. Detailsconcerning the analyses of these products are givenelsewhere (Katikaneni et al., 1995a,b).

Results and Discussion

Catalyst Characterization. (A) Powder XRDMeasurements. The XRD pattern of the nonimpreg-nated HZSM-5 catalyst (K0) is shown in Figure 2 andis consistent with those reported in the literature (Chenet al., 1989; Campbell, 1983). The XRD patterns for thepotassium-impregnated catalysts did not show anyappreciable change from that given in Figure 2, imply-ing that, up to 2 wt % potassium concentration, therewas no loss of crystallinity by the catalysts due topotassium impregnation. Also, no new XRD lines wereobserved in the potassium-impregnated catalysts.(B) N2 Adsorption Measurements. The micropore

surface areas of the potassium-impregnated and non-impregnated HZSM-5 catalysts are given in Table 2 asa function of potassium concentration. It is seen thatthe micropore surface area decreased monotonicallywith increasing potassium concentration in the catalyst.As was mentioned earlier, HZSM-5 catalyst containsmostly micropores. Also, it is known that the microporesurface area in HZSM-5 is a measure of the innerchannel surface area of the micropores. Thus, thedecrease in micropore surface area with potassiumconcentration may be due to pore blockage of themicropores of the impregnated catalysts by potassium(most likely in the form of K2O). Pore blockage is alsosupported by the decrease in pore volume with increas-ing potassium concentration (see Table 2). Theseresults were used in conjunction with results obtainedfrom TPD of NH3 to evaluate the acid site densities ofthe 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.

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

Figure 2. XRD spectra of HZSM-5 catalyst.

Figure 3. TPD spectra of HZSM-5 and K-impregnated HZSM-5catalysts.

Table 2. Characteristics of Potassium Impregnated andNonimpregnated HZSM-5 Catalysts

catalystcatalystidentity

microporesurface

area, m2/g

porevolume,cm3/g

acid density,mm2/(m2 g of catalyst)

HZSM-5 K0 389 0.14 0.6550.5 wt %

K/HZSM-5K1 358 0.12 0.542

1.0 wt %K/HZSM-5

K2 325 0.10 0.400

2.0 wt %K/HZSM-5

K3 297 0.07 0.242


well-resolved TPD peaks designated as the low-temper-ature peak (l) and the high-temperature peak (h). Onthe other hand, the impregnated catalyst containing 2.0wt % 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 °Crepresent weak acid sites (i.e., l peaks). Thus, theexistence of TPD peaks at 240 and 485 °C for nonim-pregnated HZSM-5 catalyst signifies the presence ofboth the weak and strong acid sites on the catalyst. Itis well-known that, for both weak and strong acid sites,a decrease in the peak temperature represents a de-crease in the strength of the acid site. Therefore theprogressive shift to lower peak temperatures for both land h peaks as the potassium concentration in thecatalysts increased (see Figure 3) implies that there wasa decrease in the strengths of the h and l acid sites withincreasing potassium concentration. In the case of thecatalyst containing 2 wt % potassium (i.e., K3), theabsence of the h peak shows that there was a completeneutralization of the strong acid sites with 2 wt %potassium.Acid site density can be used to represent the con-

centration of the acid sites present on the surface areaof the catalyst which is accessible to the reactants. Acidsite density on the basis of surface area becomesimportant because catalytic cracking reactions are es-sentially a surface phenomenon which depend on theaccessible acid sites present on the catalyst. In thiswork, acid site density of each catalyst was calculatedas the ratio of the total TPD peaks area to the microporesurface area (i.e., inner channel area of the pores) ofthe catalysts as follows:

where

The results for the four catalysts are given in Table 2.It is seen from the table that the acid site densitydecreased as the potassium concentration in the cata-lysts increased. These results show that the totalnumber of acid sites decreased with potassium concen-tration at a faster rate than the micropore surface areadid. Thus, by combining the TPD and micropore surfacearea results, it can be concluded that the incorporationof potassium into HZSM-5 catalyst decreases not onlythe acid site strength but also the total acidity of thecatalyst. This is consistent with the observations ofBhatia (1990) and Szostak (1989).(D) FT-IR Studies. In addition to determining the

strength and total number of acid sites using the TPDof NH3, it was also necessary to determine the types ofacid sites present (i.e., whether Lewis or Bronsted acidsites) as well as their distribution on the catalyst surfaceas a function of potassium concentration. According tothe literature, the variation in the amounts of thesevarious acid types has a tremendous effect on productdistribution. As suggested by Vedrine et al. (1979),information regarding the types of acid sites present onthe catalysts was obtained from the IR spectra of thesamples in the pyridine region (frequency in the range

1425-1575 cm-1) as well as in the hydroxyl stretchingregion (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 4and 5, respectively, for catalysts K0, K1, and K2. InFigure 4, the band at 3730 cm-1 is characteristic of aterminal hydroxyl group (OH) (i.e., adsorbed water),while the bands at 3670 and 3590 cm-1 are assigned tothe lattice terminal Si-OH groups and acidic hydroxylAl-OH groups, respectively. According to Vedrine etal. (1979), the terminal Si-OH and the acidic hydroxylAl-OH groups are particularly important because oftheir ability to generate Bronsted acid sites which areknown to be essential for all cracking reactions. It istherefore desirable to determine how these bands aredistributed in the catalyst as a function of potassiumconcentration. In Figure 4, it is seen that these threebands are very prominent for the nonimpregnatedHZSM-5 catalysts. However, as the potassium concen-tration increased in the catalyst, the IR spectra (Figure4) showed a substantial decrease in the intensity of allthree bands up to a potassium concentration of 0.5 wt% and a complete disappearance of the bands beyondthis concentration. These observations confirm theearlier results from TPD analysis that the concentrationof acidic hydroxyls (i.e., total acidity) in these catalystsdecreased 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), andK2 (1 wt % potassium) are presented in Figure 5. The

µ ) R/(micropore surface area) (1)

µ ) total TPD peak area per unit microporesurface area, mm2/(m2 g of catalyst)

R ) total area under the TPD peaks, mm2/g

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


HZSM-5 catalysts exhibited bands at frequencies of1545, 1490, and 1445 cm-1. According to Rahman etal. (1988) and Borade and Clearfield (1994), the bandat the frequency of 1545 cm-1 represents the Bronstedacid sites, whereas the one at 1490 cm-1 represents thepresence of a mixture of both Bronsted and Lewis acidsites. On the other hand, the band at the frequency of1445 cm-1, which was observed to be broader than theother 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 1545cm-1 (Bronsted acid sites) decreased. On the otherhand, an increase in the potassium concentration re-sulted in a shift in the frequency band for the Lewissites to higher wavenumbers. This shift was attributedto the formation of Lewis acid sites, which were strongerthan those present in nonimpregnated HZSM-5 catalyst.This means that the Bronsted acid sites are moresusceptible to alkali metal poisoning than the Lewis acidsites. These results highlight the opposing effects whichthe incorporation of potassium into HZSM-5 has on theBronsted and Lewis acid sites. These are the apparentincrease in the strength of the Lewis acid sites whiledecreasing that of the Bronsted acid sites.(E) 1H MAS Solid State NMR. 1H MAS solid state

NMR was used to provide quantitative information onboth the number of structural hydroxyl groups and thetotal number of hydroxyl groups present on the catalystas a function of potassium concentration.

1H MAS NMR spectra of K0 (nonimpregnated HZSM-5) and K2 (1.0 wt % potassium) are presented in Figure6. The figure shows a sharp peak at 6.17 ppm forcatalyst K0, whereas the peaks for catalyst K2 (i.e., 1%potassium) were positioned at ca. 5.5 and 2.71 ppm, bothwith respect to the external tetramethylsilane. It isknown (Thomas and Klinowski, 1985) that the chemicalshift assigned to 6.1 ppm is due to structural hydroxylgroups, whereas that at 2.7 ppm is due to terminal OHgroups and hydroxyl groups attached to extraframeworkaluminum. On the other hand, the peak at 5.5 repre-sents structural hydroxyl groups which have beenweakened due to their chemical environment. Since itis the structural OH group that is responsible for thegeneration of Bronsted acidity, it therefore implies thatthe nonimpregnated HZSM-5 catalyst contained mostlyBronsted acid sites, whereas potassium modification ofthe catalysts as in K2 (1 wt % potassium) changed theacid structure to a combination of Lewis acid and weakBronsted acid sites.The area under the sharp line gives the number of

respective hydroxyl groups present on the catalyst. It

is seen that HZSM-5 catalyst possesses a higher numberof Bronsted acid sites than the potassium-impregnatedHZSM-5 catalysts.The NMR results are consistent with those obtained

from FT-IR. These results indicate a marked decreasein Bronsted acidity with the incorporation of increasingamounts of potassium.(F) Model Compounds Reaction Studies. Metha-

nol cracking andm-xylene isomerization reactions wereused as model reactions to provide further evidence forthe degree of acid site poisoning of the catalysts. It iswell-established (Chang, 1983) that a high methanolconversion to produce mainly aromatic hydrocarbons inthe methanol cracking reaction implies the presence ofstrong acid sites (especially the Bronsted acid sites) inthe catalyst. On the other hand, low methanol conver-sions to produce dimethyl ether means the presence ofweak Bronsted acid sites in the catalyst. In the case ofthe m-xylene isomerization reaction, a low m-xyleneconversion to selectively produce p- or o-xylenes impliesthe presence of weak acid sites in the catalyst.Table 3 shows methanol conversions and product

distributions obtained from methanol cracking overimpregnated and nonimpregnated HZSM-5 catalysts.The table shows that methanol conversion and aromatichydrocarbon yield decreased as the potassium concen-tration in the catalyst increased (i.e., acid site strength

Table 3. Methanol Conversions and Yields of Various Products from the Methanol Cracking Reaction

for given reactn temp

370 °C 400 °C 450 °C

potassium concn, wt % 0 0.5 1.0 2.0 0 0.5 1.0 2.0 0 0.5 1.0 2.0methanol conversion, % 96.0 90.0 85.0 80.0 100 92.0 90.5 84.5 100 97.0 93.0 90.0total hydrocarbon yield, wt % 44.0 41.4 41.5 40.5 43.5 42.9 43.2 42.5 44.1 43.1 43.4 42.8dimethyl ether, wt % 0.5 1.5 2.5 3.0 0 1.0 1.8 2.2 0 0.6 1.1 1.6water, wt % 56.0 57.1 56.0 56.5 56.5 56.1 55.0 55.3 55.9 56.3 55.5 55.6methane, wt % 0.5 0.4 0.2 0.2 1.2 0.8 0.5 0.6 3.0 2.5 2.0 1.9ethane, wt % 0.3 0.3 0.2 0.2 0.8 0.5 0.4 0.5 1.0 0.9 0.5 0.6propane, wt % 7.2 5.6 1.0 1.3 10.5 7.1 5.0 5.9 10.5 8.5 3.5 4.2butanes, wt % 15.5 11.4 2.5 2.8 14.4 10.5 3.0 4.0 14.4 10.1 4.5 5.1ethylene, wt % 3.9 4.9 8.5 8.3 6.5 8.5 11.2 10.5 7.1 10.5 13.5 12.5propylene, wt % 4.4 7.4 15.5 14.6 6.4 7.5 17.2 16.6 6.8 8.9 16.5 14.9butenes, wt % 7.4 8.3 11.5 10.8 6.3 8.0 9.6 8.2 6.2 7.8 9.9 8.8C5-C10 aliphatics, wt % 34.5 36.4 41.5 47.6 25.5 29.7 29.9 37.3 19.5 22.7 25.7 32.5aromatics, wt % 26.3 25.3 19.1 14.2 28.4 27.4 23.2 16.4 31.5 28.1 24.2 19.5C2-C4 olefins, wt % 15.7 20.6 35.5 33.7 19.2 24.0 38.0 35.3 20.1 27.2 39.9 36.2

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


and density decreased) for all three reaction tempera-tures used in this work for methanol cracking reactions.On the other hand, the yield of dimethyl ether increasedwith potassium concentration for all three reactiontemperatures. These results support our earlier char-acterization results, which showed that the acid sites(especially the Bronsted acid sites) diminished withincreasing potassium concentration.In the case of m-xylene isomerization, the true

isomerization process can be distinguished from sideprocesses such as C2-C4 production by determiningwhat fraction of the m-xylene converted goes to o- orp-xylene formation. This is referred to as the extent ofisomerization (I) and is defined as

Figure 7 shows the variation of the extent of isomer-ization with the catalyst acid site density. It is clearfrom this figure that the extent of isomerization in-creases as the acidity of the catalyst decreases. This isconsistent with the literature (Nayak and Choudhary,1982) and also confirms our earlier characterizationresults which showed that the incorporation of potas-sium into HZSM-5 not only dilutes the concentrationbut also weakens the strengths of the acid sites (espe-cially that of Bronsted acid sites).Catalyst Performance Studies. Canola Oil Con-

version and Product Distribution. The canola oilconversion, overall mass balances, and product distribu-tions obtained from canola oil reactions over HZSM-5and potassium-impregnated HZSM-5 catalysts are givenin Table 4a-c as functions of reaction temperature,space velocity, and potassium concentration in thecatalyst. These results involved the repetition of anumber 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) andspace velocity (Table 4c) as well as decreasing reactiontemperature (Table 4a). Table 4a also shows that thevariation of the yield of gas product with these param-eters followed a trend similar to the one exhibited bycanola oil conversion. On the other hand, there was ageneral decrease in the yield of OLP with increasingreaction temperature and potassium concentration inthe catalyst. These results represent typical yieldsobtained from catalytic cracking reactions.Optimum Conditions for the Production of C2-

C4 Olefins. A major objective of this work was todetermine the optimum catalyst and operating condi-

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

which describe a typical reaction scheme for canola oilconversion (Katikaneni et al., 1995a,b) in conjunctionwith the effects of catalyst characteristics and reactionoperating conditions on the yield of C2-C4 olefins. Here,the first step (eq 3) involves the initial thermal decom-position of canola oil molecules to give heavy oxygenatedhydrocarbons (Chang and Wan, 1947; Alencar et al.,1983; Nawar, 1969; Idem et al., 1996), which subse-quently undergo deoxygenation (eq 4) to give heavyhydrocarbons. In this work, the mixture of heavyhydrocarbons and heavy oxygenated hydrocarbons isregarded as “residual oil”. The next step involves thesecondary cracking of heavy hydrocarbons (eq 5) to giveshort and long chain paraffins and olefins. In reaction6, the light olefins typically in the C2-C4 range oligo-merize to yield olefins in the C2-C10 range. Bothaliphatic (mostly cyclic) and aromatic hydrocarbons areeventually produced in the pores of the zeolite catalysts(eq 7) as a result of cyclization and aromatizationreactions of C2-C10 olefins. Coke is formed either dueto the condensation of the canola oil molecules (eq 8) orby polymerization of large aromatic hydrocarbons insidethe pores of the zeolite catalysts (eq 9) or both. Thesereactions 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 6and 7). Thus, modifications in catalyst acid site densityas well as changes in operating conditions (such asreaction temperature) are certain to have tremendouseffect on both canola oil conversion and product distri-bution and, consequently, on C2-C4 olefins yields. Theoptimum values of these parameters that result in theproduction of optimum amounts of C2-C4 olefins arediscussed below.(A) Catalyst Acidity. The variation of C2-C4 olefins

with potassium concentration (i.e., catalyst acidity) isgiven in Figure 8. The figure shows that a maximumexists in the relationship between the yield of C2-C4olefins and catalyst acidity (potassium concentration)for all reaction temperatures. These results can beexplained using the reaction steps that involve theformation of C2-C4 olefins from residual oil (eqs 5 and6) and depletion of these olefins to form aromatichydrocarbons (eq 7). These results show that startingfrom the nonimpregnated HZSM-5 (high acidity) cata-lyst to catalyst K1 or K2 in some cases (i.e., decreasing

Figure 7. Variation of the extent of isomerization with acid sitedensity for m-xylene isomerization.

I ) (p- + o-xylenes produced (wt %))/(total amt of products (wt %)) (2)

canola oil f heavy oxygenated CxHy (thermal)(3)

heavy oxygenated CxHy f heavy CxHy + H2O +CO2 + CO (thermal and catalytic) (4)

heavy CxHy f

paraffins + olefins (short and long chain)(catalytic and thermal) (5)

light olefins S C2-C10 olefins (catalytic) (6)

step 4 C2-C10 olefins S

aromatic CxHy + aliphatic CxHy (catalytic) (7)

canola oil f coke (thermal) (8)

n(aromatic CxHy) f coke (catalytic) (9)


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

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

Table 4. Canola Oil Conversions and Yields of Various Products

(a) Overall Mass Balances and Yields of Gaseous Products

with catalyst K0 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

Overall Mass Balances, wt %coke 3.0 4.0 5.0 6.0 5.0 5.0 10.0 8.0 7.0 13.0 10.0 9.0gas 25.0 34.0 50.0 21.0 30.0 45.0 18.0 27.0 42.0 11.0 21.0 38.0OLP 63.0 56.0 41.0 57.0 58.0 39.0 51.0 54.0 35.0 44.5 44.5 36.0water 3.0 1.0 0 5.0 3.0 4.0 8.0 5.0 6.0 9.0 8.0 5.0residue 3.0 1.0 0 10.0 4.0 5.0 11.0 8.0 6.0 20.5 15.5 10.0unacctd 3.0 4.0 4.0 1.0 0 2.0 2.0 2.0 0 2.0 1.0 2.0coversion 96.9 98.9 100 89.9 96.0 96.9 88.8 91.8 94.0 79.1 84.3 89.8

Yields of Gaseous Products, wt % of Canola Oil Fedmethane 0.3 0.3 1.1 0.2 0.3 1.2 0.6 0.3 0.9 0.3 0.7 3.3ethylene 0.5 2.2 5.2 1.2 2.3 5.0 1.8 2.2 5.0 0.8 1.6 3.4ethane 0.5 0.6 1.9 0.3 0.6 1.6 0.7 0.6 1.4 0.4 1.0 3.4propylene 0.9 4.4 7.8 2.9 6.2 12.4 4.2 5.6 10.2 2.6 4.7 7.9propane 9.9 10.6 16.4 3.8 4.1 4.1 1.1 4.2 6.5 0.8 1.3 2.4isobutane 1.0 2.1 2.7 3.1 2.6 1.8 1.0 2.6 2.7 0.4 0.3 0.2n-butane 6.2 6.1 5.7 1.7 3.7 5.8 2.4 3.1 4.3 1.4 2.4 3.8isobutylene 0.6 1.4 3.0 2.5 3.7 3.6 2.1 2.6 3.5 1.0 1.9 3.2other C4d 2.1 1.9 1.2 - 2.6 4.8 0.8 2.8 3.6 1.6 2.8 4.7C5

+ 2.7 3.2 4.4 3.9 0.7 0.8 2.2 0.8 0.9 0.8 1.1 1.3CO - 0.7 0.2 0.5 0.7 0.9 0.5 1.0 1.3 0.6 0.7 1.0CO2 0.4 0.6 1.1 1.9 2.4 3.0 0.7 1.1 1.7 0.7 2.5 3.5C2-C4 olefins 4.1 9.9 17.2 6.6 14.8 25.8 8.8 13.3 18.7 6.0 11.0 19.2

(b) Yields of Liquid Products, wt % Canola Oil Fed

with catalyst K0 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

methanol 1.3 0.8 0 1.5 2.3 1.1 2.0 1.9 0.7 3.3 2.3 1.3acetone 1.5 0.4 0 1.1 0.6 0.2 1.3 1.1 0.2 0.4 0.2 0aliphatic hydrocarbons 1.0 0.5 0.1 3.5 4.8 2.4 15.7 8.0 3.7 24.7 21.8 14.6benzene 6.4 5.8 4.3 3.0 4.5 3.6 2.0 4.0 3.5 0.7 1.7 1.5toluene 18.6 18.1 14.9 10.0 12.9 9.5 4.1 11.8 8.2 0.8 1.6 3.1xylenes 15.1 16.1 14.4 11.5 12.6 8.7 4.6 11.9 7.6 0.6 1.7 2.0ethylbenzene 4.6 3.4 2.0 4.2 3.7 2.1 1.6 3.1 1.8 0.7 0.7 0.5propylbenzene 3.9 3.4 2.3 5.5 3.6 1.6 2.9 3.5 1.7 0.4 0.4 0.6C9 + aromatics 3.2 2.7 1.3 6.0 2.7 2.1 8.1 1.7 0.9 1.6 1.2 2.1total aromatics 51.8 49.5 39.2 40.2 40.1 29.0 23.3 36.1 24.7 4.9 7.5 9.8unidentified 7.2 4.7 1.7 12.1 9.6 7.9 8.7 7.1 3.3 11.2 12.6 10.3

(c) Yields of Gaseous and Liquid Products

400 °C 450 °C 500 °C

components 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

Yields of Gaseous Products (wt % of Canola Oil Fed)methane 0.4 0.6 0.2 0.3 1.1 0.9ethylene 1.3 1.8 2.1 2.2 4.0 5.0ethane 0.3 0.7 0.4 0.6 1.6 1.4propylene 3.7 4.2 6.0 5.6 10.9 10.2propane 1.6 1.1 2.5 4.2 7.3 6.5isobutane 1.3 1.0 1.9 2.6 1.2 2.7n-butane 1.7 2.4 3.3 3.1 5.2 4.3isobutylene 1.4 2.1 2.4 2.6 4.3 3.5other C4d 1.7 0.8 2.9 2.8 3.4 3.6C5

+ 0.6 2.2 0.6 0.8 1.0 0.9CO 0.6 0.5 0.8 1.0 1.1 1.3CO2 1.4 0.7 2.1 1.1 2.6 1.7C2-C4 8.0 8.8 11.2 13.3 23.0 18.7

Yields of Liquid Products (wt % of Canola Oil Fed)methanol 2.7 2.0 2.4 1.9 1.2 0.7acetone 1.5 1.3 1.1 1.1 0.5 0.2aliphatic hydrocarbons 16.9 15.7 9.9 8.0 4.0 3.7benzene 1.7 2.0 4.0 4.0 2.8 3.5toluene 3.6 4.1 9.1 11.8 2.8 8.2xylenes 3.7 4.6 9.0 11.9 7.0 7.6ethylbenzene 1.6 1.6 3.2 3.1 1.6 1.8propylbenzene 2.7 2.9 3.5 3.5 1.3 1.7C9 + aromatics 8.2 8.1 2.3 1.7 4.0 0.9total aromatics 21.4 23.3 31.2 36.1 23.4 24.7unidentified 9.5 8.7 12.4 7.1 8.9 3.3


aromatization reactions increases as catalyst aciditydecreases, hence, the net increase in C2-C4 olefins withpotassium concentration (i.e., decreasing catalyst acid-ity). On the other hand, the decrease in the yield ofC2-C4 olefins with further increase in potassium con-centration (i.e., decrease in catalyst acidity) observedfor catalyst K3 (and K2 in some cases) shows that thelow consumption of C2-C4 olefins for oligomerization(eq 6) and aromatic hydrocarbon formation (eq 7) doesnot compensate for the low formation of C2-C4 olefinsfrom residual oil cracking (eq 5). The net result is thedecrease in C2-C4 olefin yield for catalyst K3 (and K2in some cases). These results thus establish the levelof acidity required for the production of optimumamounts of C2-C4 olefins.(B) Effect of Reaction Temperature. The varia-

tion of the yield of C2-C4 olefins with reaction temper-ature is given in Figure 9. This figure shows that C2-C4 olefins yield increases with reaction temperature forall catalysts. This result is typical of the highly endot-hermic thermal cracking process which results in anincrease in C-C bond scission of C5

+ aliphatic hydro-carbons and the consequent increase in the amounts ofC2-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 (thecatalyst with lowest acid site density in which both weakand strong acid sites are present; see Figure 3). Theseresults are given in Table 4c, which shows that the yieldof 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. Thereverse 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 at500 °C and 3.6 h-1 space velocity. Therefore, favorableconditions for maximum C2-C4 olefin production arehigh reaction temperatures, high canola oil space veloci-ties, and catalysts with relatively low Bronsted and totalacidity. However, the overall results show that themaximum C2-C4 olefin yield of 25.8 wt % was obtainedwith 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 forcanola oil cracking over impregnated and nonimpreg-nated HZSM-5 catalysts. This was performed by usingmodel compound reactions such as methanol crackingandm-xylene isomerization. The model reactions wereused to show the nature of involvement of certainreaction intermediates (oxygenates and aromatics) inthe reaction pathway for the catalytic conversion ofcanola oil. This was achieved by comparing the waysin which the conversions and product distributions ofmodel reactions responded to changes in catalyst acidityand reaction temperature with those for canola oil.(A) Methanol Cracking. (1) Effect of Catalyst

Acidity. Table 3 shows the conversions as well as theproduct distributions obtained from methanol crackingreactions over the four catalysts as a function of potas-sium concentration and reaction temperature. As wasshown earlier, methanol conversion decreased as theamount of potassium in the catalyst increased for allreaction temperatures. Also, as was shown earlier byTPD, FT-IR, and NMR results, the role of the incorpora-tion of potassium into HZSM-5 catalyst is that ofdecreasing both the total and Bronsted acid sitesoriginally present in the nonimpregnated HZSM-5catalysts. Thus, the decrease in methanol conversionwith potassium concentration can be attributed to thedecrease in the total and Bronsted acidity of thecatalyst. This is consistent with the results of Chang(1983) and Itoh et al. (1984).On the other hand, although the variation in potas-

sium concentration (Bronsted acidity) had a drasticeffect on methanol conversion, its effect on the selectiv-ity for hydrocarbon production was negligible. This isseen in Table 3, where the conversion decreased from100 to 80 wt %, whereas the yield of total hydrocarbonswas always within the range 40.5-44.1 wt % for allcatalysts and reaction temperatures used. It is well-known that hydrocarbon formation is highly influencedby the catalyst Bronsted acidity and its shape selectivitycharacteristics. However, these results for hydrocarbonselectivity suggest that the total yield of hydrocarbonswas not affected significantly by catalyst Bronstedacidity.Figure 10 shows the typical effect of catalyst acidity

(potassium concentration) on the yields of aromatic andC5-C10 aliphatic hydrocarbons for reactions conductedat 370 °C. The figure shows that while the yield ofaromatic hydrocarbons increased with acid site density,that for aliphatic hydrocarbons followed the reversetrend. These results, in conjunction with TPD, FT-IR,

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

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


and NMR results, show conclusively that, unlike in theformation of aliphatic hydrocarbons, high-strength Bron-sted acid sites are required for the formation of aromatichydrocarbons.Figure 11 illustrates the variation of C2-C4 olefins

with catalyst acid site density. The combined yield ofC2-C4 olefins gives an indication of the extent of boththe 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 toproduce aromatic hydrocarbons. These steps for metha-nol cracking are given in eqs 10-13, which describe atypical reaction scheme for methanol cracking to pro-duce hydrocarbons over zeolite catalysts (Chang, 1983).

A high yield of C2-C4 olefins implies that whiledimethyl ether conversion was relatively high, theextents of oligomerization and aromatization were low,and consequently, only a small quantity of the C2-C4olefins was used up. The converse is also true. How-

ever, Figure 11 shows that the yield of C2-C4 olefinsincreased initially with acid site density (i.e., decreasingpotassium concentration) up to 1 wt % potassium, andthen decreased beyond this concentration. The decreasein the yield of C2-C4 olefins with potassium concentra-tion for the catalyst containing greater than 1 wt %potassium can be explained as follows. It is known thatthere are low extents of oligomerization and aromati-zation reactions over catalysts with low acid site density(example, K3). Therefore a large fraction of C2-C4olefins already formed is retained in the case of crackingreactions over such catalysts. However, this retentioncannot compensate for the low conversions of dimethylether to produce C2-C4 olefins over catalysts containingweak acid sites.These results concerning methanol conversions and

product distributions can be combined and used toprovide additional reasons why methanol conversionsdecreased with decreasing catalyst acidity. This isdiscussed using eqs 10-13.The forward reactions (eqs 10-13) are known to be

catalyzed by high catalyst acid site density. It is alsoknown that, for each of these reversible reaction steps,the forward reaction is catalyzed by acid sites to agreater extent than the corresponding backward reac-tion. Also, the equilibrium for methanol conversionswith zeolite catalysts is in favor of aromatic hydrocarbonformation (large methanol conversion) under typicalreaction conditions (eq 13). It thus appears from theproduct distribution in Table 3 that there is also anindirect effect that a decrease in catalyst acidity has onmethanol conversion. This indirect effect involves ashift in the equilibrium in reactions 10-13 in favor ofa smaller methanol conversion. This can be explainedas follows: Table 3 shows that dimethyl ether conver-sion to olefins increased with catalyst acid site density.It also shows that the subsequent conversion of theseolefins to aromatics increased with acid site density, asshown previously. Thus, for a catalyst with high acidsite density, the large conversion of C2-C10 olefins toaromatics will require large amounts of olefins to besupplied to replenish the ones used up by the aroma-tization reactions. The results in Table 3 shows thatthese C2-C10 olefins will have to be supplied by theconversion of dimethyl ether. Similarly, the largeconversions of dimethyl ether to C2-C4 olefins andsubsequent oligomerization of these light olefins toproduce olefins in the C2-C10 range will induce largerconversions from methanol to supplement the dimethylether depleted by the succeeding reaction (eq 11). Thenet result of these inductive processes is that, for acatalyst with high acid site density and for a system ofreversible reactions in series where the forward andcorresponding backward step are not catalyzed by acidsites to the same extents, there is a shift in the reactionequilibrium in the direction that will favor the produc-tion of a large amount of aromatic hydrocarbons (i.e., alarge extent of methanol conversion). Conversely, forcatalysts with high potassium concentration (low acidsite density), the shift in equilibrium will be in thedirection that will favor the production of a smallamount of aromatic hydrocarbons (i.e., a small extentof methanol conversion).(2) Reaction Temperature. Table 3 shows that

methanol conversion increased with reaction tempera-ture for all four catalysts. This is typical of crackingreactions, as was shown earlier in the case of canola oilconversion.

Figure 10. Variation of the yields of aromatic and aliphatichydrocarbons with catalyst acidity for the methanol crackingreaction.

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

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)


The effect of reaction temperature on the yields ofaromatic and C5-C10 aliphatic hydrocarbons are pre-sented in Figures 12 and 13, respectively. It is seenfrom Figure 12 that the aromatic hydrocarbon yieldincreased with an increase in reaction temperature forall of the catalysts. On the other hand, Figure 13 showsthat the variation of the yield of C5-C10 aliphatichydrocarbons with reaction temperature showed areverse trend to that exhibited by the yield of aromatichydrocarbons. These results imply that an increase intemperature results in large extents of cyclization andaromatization reactions and, consequently, a decreasein 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 mentionedearlier that, at high reaction temperatures, largeamounts of C2-C4 olefins will oligomerize to produceC2-C10 olefins (eq 12), which in turn will aromatize toproduce large amounts of aromatic hydrocarbons (eq 13).Generally, this should result in a decrease in the C2-C4 olefins. However, high temperatures are also knownto favor the C-C scission of C5

+ aliphatic hydrocarbonsas well as dealkylation of alkylated aromatic andaliphatic hydrocarbons to produce mostly the chemicallystable C2-C4 olefins. The result is an increase in theyield of C2-C4 olefins and a consequent decrease in theyield of C5-C10 aliphatic hydrocarbons.(B) Isomerization of m-Xylene. Effects of Cata-

lyst Acidity. The m-xylene conversions and the prod-uct yields for the isomerization of xylenes are presentedin Table 5 as a function of potassium concentration (i.e.,catalyst acidity). The products obtained were o-xylene,p-xylene, toluene, benzene, trimethylbenzenes, andsmall amounts of C2-C4 hydrocarbons. The presenceof C2-C4 hydrocarbons in the m-xylene isomerization

products supports our earlier assertion regarding theinvolvement of the dealkylation reactions of alkylatedaromatic hydrocarbons during methanol cracking.Table 5 shows that the conversion of m-xylene in-

creased as the catalyst acid site density increased (i.e.,with decreasing potassium concentration in the cata-lyst). However, an increase in potassium concentrationin the catalyst was detrimental tom-xylene conversion;its effect on p-xylene selectivity was highly favorableup to a potassium concentration of 1 wt %. Beyond thispotassium concentration, the selectivity to p-xylenestarted to decrease. On the other hand, the yield ofo-xylene decreased progressively with increasing potas-sium loading. The selectivity for either o- or p-xyleneformation can be seen more clearly in Table 6, whichshows 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 p-xylene/o-xylene ratio increased as the catalyst aciditydecreased. This means that selectivity for p-xylenedecreased with increasing catalyst acidity. These ob-servations are consistent with the findings of Nayak andChoudhary (1982).Apart from C2-C4 hydrocarbons, the other side

products obtained from the m-xylene reaction werebenzene, toluene, and trimethylbenzene. The concen-trations of these side products decreased as the potas-sium loading increased (i.e., as the catalyst aciditydecreased). The type of side products obtained gives anindication of the type of reactions that took place duringthe conversion of m-xylene over the zeolite catalysts.Generally, these are isomerization, disproportionation,dealkylation, and hydrogen transfer. For example, thepresence of trimethylbenzene (TMB) and toluene indi-cates principally the involvement of disproportionation(eq 14) and dealkylation (eq 15) reactions of m-xylene.

Both reactions 14 and 15 involve the hydrogen transferreaction. On the other hand, the presence of p- ando-xylene indicates the involvement of the isomerizationreaction. The typical product selectivity ratios shownin Table 6 indicate the extents of various side reactions.These are the toluene/o- + p-xylenes ratios, indicatingthe extent of the dealkylation of m-xylene, and TMB/p-+ o-xylenes ratios, indicating the extent of m-xylenedisproportionation. Since these ratios were less than0.38 and were decreasing with decreasing catalystacidity, it can be concluded that weak acid sites arerequired for the isomerization of m-xylene while the

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

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

Table 5. m-Xylene Conversions and Yields of VariousProducts for the m-Xylene Isomerization Reactiona

potassium 0.0 0.5 1.0 2.0concn in HZSM-5catalyst, wt %

m-xyleneconversion, wt %

59.5 47.4 40.9 31.4

C2-C4 hydrocarbons, % 0.2 0.2 0.1 0benzene, % 0.4 0.1 0 0toluene, % 14.4 10.4 5.3 1.5p-xylene, % 21.0 24.8 25.2 23.3m-xylene, % 40.5 52.6 59.1 68.6o-xylene, % 16.5 7.3 6.9 5.4trimethylbenzenes, % 7.0 4.6 3.4 1.2a Reactions conducted at 400 °C and 1.8 h-1 space velocity.

2C8H10 w C7H8 + C9H12 (disproportionation)(14)

2C8H10 w 2C7H8 + C2H4 (dealkylation) (15)


strong acid sites are required for the disproportionationand dealkylation reactions. Similar conclusions werereached by Nayak and Choudhary (1982) in the case ofm-xylene isomerization.(C) Application of Model Reactions to Canola

Oil Conversion. The application of model reactionsto canola oil conversions are discussed using Table 4a-c, which shows canola oil conversions and productdistributions for canola oil cracking reactions for allcatalysts.(1) Effects of Catalyst Acidity. It is observed from

the table that canola oil conversion decreased withincreasing potassium concentration (i.e., decreasingacidity) in the catalyst. It is also observed that lowcanola oil conversions were obtained only under condi-tions where high yields of residual oil were obtained inthe product (see Table 4a). This indicates that althoughit was possible for the initial cracking of canola oil toresidual oil (a mixture of heavy hydrocarbons and heavyoxygenated 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 lowacidity (example, catalyst K3), high acidity appears tobe essential for subsequent deoxygenation, oligomer-ization, and aromatization reactions. According toPrasad et al. (1986a,b) and Katikaneni et al. (1995a,b),the cracking reactions of hydrocarbons or unsaturatedtriglyceride molecules are favored by catalysts with highstrength and density of acid sites, such as zeolites andsilica-alumina. Our results show that the neutraliza-tion of these acid sites with potassium resulted in adecrease in the extent of the cracking of residual oilmolecules and not of the long and bulky triglyceridemolecules (canola oil). This was evidenced by thepresence of some high molecular weight materialshaving boiling points higher than 350 °C in the productand not of the triglyceride molecules themselves.Thermal cracking of plant oils was observed to occur

by Chang and Wan (1947), Alencar et al. (1983), andIdem et al. (1996) at temperatures g300 °C. Theseworkers carried out their investigations using batch andfixed-bed reactors in the temperature range 300-500°C. In the present study, the canola oil feed enteredthe reactor from the top of the reactor where thetemperature was essentially the same as the reactiontemperature (400-500 °C). There was also a longpreheating zone (≈160 mm). This was the space insidethe reactor between the top of the reactor and the topof the catalyst bed. On the basis of the total volume ofthe reactor and also, based on canola oil vapor atstandard temperature and pressure (25 °C and 1 atm),the calculated residence times of canola oil vaporthrough the reactor are ca. 13.6 and 27.2 min for canolaoil WHSV values of 3.6 and 1.8 h-1, respectively. Alsoon the basis of uncracked canola oil vapor and the

volume of the preheating zone, the times it woud takefor the feed vapor to reach the catalyst bed are 5.5 and10.9 min for canola oil WHSV values of 3.6 and 1.8 h-1,respectively. However, it was observed in all of ourexperimental runs that the gas product appeared in lessthan 1 min from the start of the feed pump. Thisobservation therefore makes the calculated residencetimes meaningless. These much shorter experimentallyobserved residence times are attributed to the initialthermal decomposition of each canola oil molecule intoa number of relatively small molecules in the preheatingzone and the subsequent vaporization of the decomposedmolecules before catalytic reactions. Also, decomposi-tion must happen as soon as canola oil enters thepreheating zone in order to achieve the short experi-mentally observed residence times. This result of theinitial thermal cracking is in contrast with the proposalof Weisz et al. (1979) which indicated that triglyceridemolecules have to enter the pores of the zeolite catalystbefore they undergo cracking reactions. Thus, ourpresent results show that triglyceride molecules initiallyundergo thermal cracking before subsequent diffusionof cracked products into the pores of the zeolite catalystfor catalytic cracking reactions.It is highly desirable to verify whether catalyst

performance for canola oil conversion (i.e., the extentof cracking of residual oil) can be attributed exclusivelyto catalyst acidity. This verification can be performedby evaluating the relationship between the turnovernumber for canola oil conversion and catalyst acid sitedensity. In the literature, turnover number is definedin terms of the rate of reaction per active site. However,because of the complexity of the reaction/catalyst systemused in the present study, we have decided to defineturnover number in terms of canola oil conversion asfollows:

The relationship between TON and acid site density isshown in Figure 14. This figure shows that the turnovernumber for canola oil conversion decreases with anincrease in the acid site density. Thus, although an

Table 6. Product Selectivity Ratios for Various Productsfrom the m-Xylene Isomerization Reactiona

potassiumconcn in HZSM-5catalyst, wt %

0.0 0.5 1.0 2.0

Product Selectivity Ratio, wt %p-xylene/o-xylene 1.27 3.39 3.65 4.31toluene/p- + o-xylenes 0.38 0.32 0.17 0.05trimethylbenzene/p- + o-xylenes

0.19 0.14 0.11 0.04

toluene/trimethylbenzenes 2.06 2.26 1.55 1.25a Reactions conducted at 400 °C and 1.8 h-1 space velocity

(WHSV).

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

TON ) (canola oil converted (wt %))/(catalyst acid site density (mm2/m2))

) (wt % canola oil converted)/(mm2/m2)) turnover number for canola oil conversion

(16)


increase in catalyst acidity results in an increase incanola oil conversion (see Table 4a), the effectivenessof these acid sites decreases as the acid site densityincreases.Also, as can be seen in Table 4a, there was some level

of canola oil conversion even with the catalyst whoseacid sites were almost completely neutralized by potas-sium impregnation (i.e., catalyst K3). These resultsshow that canola oil conversion does not depend entirelyon catalyst acidity but also on other factors such astemperature.Figures 15 and 16 show respectively the variation of

total aromatic and aliphatic hydrocarbon yields withcatalyst acidity. It is seen from Figure 15 that therewas a drastic reduction in total aromatic hydrocarbonyield with a decrease in catalyst acidity (i.e., increasein potassium concentration). On the other hand, Figure16 shows that total aliphatic hydrocarbons increaseddrastically with decreasing catalyst acidity. As in thecase of methanol cracking, these results, in conjunctionwith TPD, FT-IR, and NMR results, show that high-strength Bronsted acid sites are required for the forma-tion of aromatic hydrocarbons. These results establisha strong similarity between methanol cracking andcanola oil conversion. In can be seen from eq 7 thatthere is a competition for the consumption of C2-C10olefin intermediates between the formation of aromatichydrocarbons and that of aliphatic hydrocarbons. Itappears from the results that weak acid sites favor the

formation of aliphatic hydrocarbons in preference toaromatic hydrocarbon formation.Figure 17 illustrates the variation of C2-C4 olefins

with catalyst acid site density. As in the case ofmethanol cracking, the yield of C2-C4 olefins gives anindication of the extent of both the initial conversion ofheavy hydrocarbons and heavy oxygenated hydrocar-bons to C2-C4 olefins and subsequent oligomerizationand aromatization of these olefins to produce aromatichydrocarbons. In the case of canola oil cracking, a highyield of C2-C4 olefins implies that while the conversionof heavy hydrocarbons and heavy oxygenated hydrocar-bons was relatively high, the extents of oligomerizationand aromatization were low. Consequently, only asmall quantity of the C2-C4 olefins was used up. Theconverse is also true. However, Figure 17 shows thatthe yield of C2-C4 olefins increased with acid sitedensity (i.e., decreasing potassium concentration) up to0.5 wt % potassium (catalyst K1) at 400 °C and 1 wt %potassium (catalyst K2) at 450 and 500 °C and thendecreased slightly beyond these potassium concentra-tions at the specified temperatures. The decrease inyield of C2-C4 olefins beyond these potassium concen-trations is explained as follows: for catalysts with highpotassium concentrations (i.e., low catalyst acid sitedensities) the high yield of C2-C4 olefins which is dueto the decrease in oligomerization and aromatizationreactions does not compensate for the decrease in theproduction of C2-C4 olefins from the low conversionsof heavy hydrocarbons and heavy oxygenated hydrocar-bons. Similar results were obtained for methanolcracking reactions.(2) Reaction Temperature. Figures 15 and 16 are

used to respectively illustrate the variation of aromaticand C5-C10 aliphatic hydrocarbons with reaction tem-perature. It is seen from Figure 15 that the total yieldof aromatic hydrocarbons increased with reaction tem-perature for all catalysts. This implies an increase incyclization and aromatization reactions with reactiontemperature. On the other hand, Figure 16 shows thatthe total yield of aliphatic hydrocarbons decreased witha decrease in reaction temperature, signifying a de-crease in the cyclization and aromatization reactions.These results also highlight the competition whichexists for the consumption of the available C2-C10olefins with regard to reaction temperature. The com-petition is for the formation of either aromatic oraliphatic hydrocarbons. This scenario was also observedfor methanol cracking.Figure 16 shows the variation of the yield of C2-C4

olefins with reaction temperature. It is seen from this

Figure 15. Variation of the yields of aromatic hydrocarbons withcatalyst acidity for canola oil conversion.

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

Figure 17. Variation of the yields of C2-C4 olefins with catalystacidity for canola oil conversion.


figure that the yield of C2-C4 olefins increased withreaction temperature for all catalysts, as was shownearlier. At high reaction temperatures large amountsof C2-C4 olefins will oligomerize and then aromatize toproduce large amounts of aromatic hydrocarbons (eqs7 and 13). Generally, this should lead to a decrease inthe C2-C4 olefins. However, high temperatures are alsoknown (Mcketta, 1992) to favor the C-C scission of C5

+

aliphatic hydrocarbons as well as the disproportionationand dealkylation of some aromatic hydrocarbons (seem-xylene reactions in Table 5), to produce mostly thechemically stable C2-C4 olefins. The result is anincrease in the yield of C2-C4 olefins and a consequentdecrease 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 tablesshow that the yields of toluene and trimethylbenzene(represented by C9

+ aromatic hydrocarbons) were large,implying that there were large extents of dealkylationand disproportionation reactions involving a number ofaromatic hydrocarbon compounds such as xylene. It isalso seen in the tables that the yields of toluene andthe C9

+ aromatic hydrocarbons decreased tremendouslywith an increase in potassium concentration (i.e., de-creasing catalyst acidity), showing that dealkylation anddisproportionation reactions decreased as the acid sitedensity of the catalyst decreased. This corroborates theresults obtained earlier for m-xylene isomerizationreactions. Again, this shows that reactions similar tothose of m-xylene are involved in canola oil conversionover HZSM-5 and K/HZSM-5 catalysts.Results of model compound reactions compared with

those from canola oil conversion have shown clearly thatcracking, deoxygenation, isomerization, aromatization,dealkylation, disproportionation, and oligomerizationreactions are involved in canola oil conversion overpotassium-impregnated catalysts. They have also shownthat the effect of modifying the acidity of the catalystson the canola oil reaction scheme is essentially that ofchanging both the overall equilibrium conversion andequilibrium conversions of individual acid catalyzedreaction steps.

Conclusions

(1) The incorporation of potassium into HZSM-5catalyst affected the strength and density of both thetotal and Bronsted acid sites, thus dramatically alteringthe 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 asdeoxygenation of oxygenated hydrocarbons, secondarycracking, oligomerization, aromatization, disproportion-ation, dealkylation, and coke formation were a strongfunction of catalyst acidity.(3) Within the reaction temperature range used in this

study, the highest yields of aromatic hydrocarbons wereobtained with catalysts K0, which contained strong acidsites.(4) Lowering the acidity changed the product distri-

bution from aromatics to aliphatics. Thus, catalyst K3,containing weak acid sites, resulted in the productionof 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 1.8h-1 with catalyst K1 containing relatively low Bronstedand total acidity.

Acknowledgment

The financial support of the Natural Sciences andEngineering Research Council of Canada (NSERC) isgratefully acknowledged.

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Received for review December 6, 1995Revised manuscript received May 16, 1996

Accepted June 10, 1996X

IE950740U

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


Documents

Catalytic Conversion of Canola Oil over Potassium-Impregnated HZSM-5 Catalysts: C 2 −C 4 Olefin Production and Model Reaction Studies