Ind. Eng. Chem. Res. 1998, 37, 1761-1768
1761
Catalytic Cracking of Vacuum Gas Oil over the Modified Mordenites and Y-Type Zeolites Mixed with Alumina Kyong-Hwan Lee* and Baik-Hyon Ha Department of Chemical Engineering, College of Engineering, Hanyang University, 17 Haengdangdong, Sungdongku, Seoul 133-791, Korea
Youn-Woo Lee Environment Remediation Research Center, Korea Institute of Science and Technology, 39-1 Haweolkok-dong, Sungbuk-ku, Seoul 136-791, Korea
Catalytic cracking of vacuum gas oil has been studied over pure zeolites modified by steaming and acid treatment, and zeolite/alumina catalysts containing any of the former zeolite (35 wt %) and alumina (65 wt %) with a different pore size distribution. The conversion of vacuum gas oil (VGO) over zeolite/alumina catalysts with a wide pore size distribution is almost the same as that over pure zeolites, but less coke is formed over zeolite/alumina compared to that over pure zeolites. The catalysts containing zeolite Y show higher conversion of VGO than on the catalysts containing mordenite. In addition, the catalysts containing modified mordenite, resulting in the development of mesopores, have improved activity and selectivity for gasoline compared to those containing mordenite with a micropore structure. Selectivity of olefin is higher on the catalysts containing mordenites compared to the catalysts containing zeolite Y because of low-bimolecular hydrogen-transfer reaction rates. The distribution of paraffin and olefin in gasoline products does not depend on the addition of alumina in mixed catalysts but rely on mainly the zeolite type. The influence of shape selectivity on products distribution causes the catalysts containing zeolite Y with large kinetic diameter to produce more branched aromatics than those containing mordenites. The modified mordenite catalysts having mesopores in the 4-nm range yield more aromatics compared to the catalysts containing mordenite with mostly micropore. Catalysts containing zeolite Y give the highest yield of xylene whereas the catalysts containing mordenites produce a maximum yield of toluene. Introduction Cracking has long been carried out using solid acid catalysts such as zeolite Y. The use of zeolite incorporated into matrix in the 1960s led to greatly improved cracking catalysts for gasoline production (Venuto and Habib, 1979). Traditionally, commercial cracking catalysts are made from zeolite Y and silica-alumina. Since zeolites and matrix must be stable to high temperature, the zeolite fraction may be dealuminated by hydrothermal and chemical treatment to improve its stability. Cracking catalyst activity and selectivity are affected by changes in both matrix and zeolite (Biswas and Maxwell, 1990). In general, higher acid site density promotes hydrogen-transfer reaction and leads to higher yields of paraffin from the cracking of vacuum gas oil (VGO). Corma et al. (1996) have studied catalytic activity, selectivity, and hydrothermal stability of silicaalumina, ultrastable Y (USY) zeolite, and mesoporous MCM-41 aluminosilicate for VGO cracking. They found that the cracking activity of gas oil by MCM-41 was greatly improved due to the greater accessibility of large molecular species to the acid site in the pore. It is clearly demonstrated that MCM-41, which has pores in the 25-Å range, gave superior activity to USY zeolite and silica-alumina for gas oil hydrocracking, which is attributed to higher surface area, presence of mesopore, mild acidity, and stability (Corma et al., 1995). * Corresponding author. E-mail:
[email protected]. Tel: (+82-2) 958-5880. Fax: (+82-2) 958-5809.
Scherzer and McArthur (1988) have investigated a series of fluid catalytic cracking (FCC) catalysts with specific composition and physical properties for cracking high-nitrogen feedstocks. They found that a higher content of zeolite together with the presence of acid sites, high surface area, and a broader pore size distribution in the catalyst matrix was favorable for cracking high-nitrogen feedstocks. Smirniotis and Ruckenstein (1994) studied the cracking of n-octane, 2,2,4-trimethylpentane, and 1-octene over ZSM-5, β zeolite, Y, USY, and composites consisting of any two of the listed zeolites for investigation of the role of ZSM-5 and β zeolite as possible additives to the conventional cracking catalysts. One of the major findings of their work was that higher aromatic selectivities over ZSM-5 were observed than those over the β zeolite due to ZSM-5 low hydrogen-transfer capability. Guerzoni and Abbot (1994) investigated the cracking reaction of a commercial reduced crude over HY and HZSM-5 and reported variations in aromatic selectivities that are in agreement with the catalyst pore geometry even though the routes leading to the gasoline species were obviously distinct. The activity and selectivity can be improved markedly by control of pore size distribution in the matrix of a cracking catalyst. A matrix with proper pore size distribution increases the accessibility to active sites for cracking of high boiling hydrocarbons. In hydrogenation of high boiling coal liquids, bimodal catalysts with large pores are more active than monomodal catalysts with small pores (Tischer et al., 1985). According to Song et
S0888-5885(97)00699-4 CCC: $15.00 © 1998 American Chemical Society Published on Web 04/01/1998
1762 Ind. Eng. Chem. Res., Vol. 37, No. 5, 1998
al. (1991), the conversion of asphaltenes increases as the mean pore radius of NiMo/Al2O3 catalyst is increased up to 29 nm which is a limiting pore size to influence the conversion. There is no publication where various treated mordenites dispersed in aluminas of different pore size have been considered as an active catalyst for catalytic cracking. Mordenite, unlike other zeolites, can be modified by extracting aluminum oxide from the framework without deteriorating extensively the crystallinity of the material, which results in the formation of a mild acidic surface as well as being thermally stable. Moreover, mordenite has higher acid strength than zeolite Y. However, mordenite has one-dimensional pore channels, which are easily blocked around the pore mouth by coke formation (Guisnet and Magnoux, 1989). The purpose of the work described here is to achieve a fundamental understanding of the effect of modified mordenite with high thermal stability as an active catalyst in a matrix of alumina with different pore size distribution on conversion and selectivity of VGO cracking. In this study, zeolite Y, mordenite, zeolite Y/alumina, and mordenite/alumina were investigated for the cracking activity of VGO and the selectivity of the products such as gasoline, kerosene, and diesel. We will also show how the framework structure and pore size distribution can alter the product distribution in the gasoline. Experimental Section Preparation of Catalysts. NaM (granule type of Norton Zeolon 900) and NaY (Linde LZY 52) zeolites were used as starting materials for the catalyst preparation. Na-type zeolites were ion-exchanged with 1 N ammonium chloride solution at 80 °C. The NH4-type zeolites were washed with distilled water, dried at 120 °C overnight and calcined at 500 °C for 5 h to prepare H-type zeolites. H-type zeolites were then treated with 100% steam for 3 h and followed by calcining at 500 °C for the same hours. H-type mordenite and H-type zeolite Y exposed to steam were denoted as SM (SiO2/ Al2O3 ) 6.5), or mordenite with a micropore structure, and SY (SiO2/Al2O3 ) 3.0), respectively. Extraction of aluminum oxide from mordenite, which has a onedimensional pore structure but has a high stability, was accomplished by treatment with 6 N HCl at 90 °C for 4 h and followed by steaming at 500 °C for 3 h in order to construct secondary pores. Such dealuminated mordenite was treated with 0.01 N HCl at 90 °C for 2 h to remove aluminum from the pores and then washed, dried, and finally calcined at 500 °C for 3 h. It was classified as DM (SiO2/Al2O3 ) 15.5) or modified mordenite catalysts. The SM, DM, and SY preparations were crushed to smaller than 200 mesh and dried at 300 °C for 12 h under vacuum. To prevent precipitating aluminum hydroxide within the pore of zeolites, the dried samples were impregnated completely with a 50:50 mixture of n-hexane and isohexane and then added to an aluminum nitrate solution. A sufficient amount of the impregnated zeolites/hexane mixture was added to the aluminum nitrate solution so that the composite zeolite/ alumina contained a zeolite content of 35 wt %. Aluminum hydroxide was precipitated as a result of the reaction of aluminum nitrate with aqueous ammonia in which pHs are controlled at either 7.8 or 9.5. The precipitated aluminum hydroxide gel was aged for 20
h at room temperature. Precipitated aluminum hydroxide containing zeolite Y (or mordenite) was then filtered, dried at 120 °C overnight, and then calcined at 600 °C for 3 h under a flowing air stream to form γ-alumina. These samples were denoted as SM/A (7.8 or 9.5), DM/A (7.8 or 9.5), and SY/A (7.8 or 9.5), in which A indicates alumina and 7.8 and 9.5 in parentheses are represented as the pH value during precipitation. Characterization of Catalysts. All samples were pretreated in a vacuum at 300 °C for 6 h before characterization by adsorption techniques. The adsorption/desorption isotherms of nitrogen on the zeolites and zeolite/aluminas were obtained at liquid nitrogen temperature using the Micromeritics ASAP-2000 adsorption system. From the observed isotherms, specific surface area using BET equation and pore volume were determined. The pore size distributions were obtained from the nitrogen desorption branches using the BJH method (Satterfield, 1980). Acidity of the samples was determined by temperature programmed desorption (TPD) of ammonia using an Altamira Instruments system equipped with a thermal conductivity detector. 100-200-mesh samples (0.3 g) were evacuated in a quartz cell at 500 °C for 1 h. Ammonia gas was then admitted to adsorb on the samples at 100 °C for 30 min and then dried at the same temperature for 1 h under a helium stream. For ammonia-TPD experiments, the temperature was linearly raised to 900 °C at a heating rate of 10 °C/min. The ammonia-TPD spectra were obtained under a helium flow of 30 cm3/min. The coke formed and accumulated on the catalysts during reaction was analyzed by an elemental analyzer (Heraeus Co., vario EL), by passing oxygen through a sample (about 3 mg) heated at 1200 °C. The gases evolved from coke on catalyst, such as carbon dioxide, sulfide dioxide, and water, were detected by a gas chromatograph (GC) with a thermal conductivity detector (TCD). Reaction Procedure and Analysis. The catalytic activity was measured in a modified microactivity tester (MAT) based on ASTM D-3907-87 (see Figure 1). Catalyst (1.7 g) was loaded in the microreactor and pretreated by nitrogen gas for 30 min at 500 °C. Table 1 shows that average molecular weight of the vacuum gas oil (VGO) used in this study as a reactant ranged from 360 to 380. The VGO (0.9 mL) was injected into the reactor for a 75-s period with a catalyst of 1.7 g. During the reaction period, liquid products were collected in an ice-cooled reservoir. Gaseous products were trapped by water displacement. At the end of the run, the reactor was purged with nitrogen for 20 min. The purged gases were also collected for analysis. Trapped liquid products were analyzed by gas chromatography (Shimadzu, GC-14A) with a capillary column (CBP-1, 25 m × 0.25 µm) and FID detector. Each peak of the products in the range of gasoline was identified using a mass spectrum analyzer (Shimadzu, GCMS-QP5000). The products in the gasoline were classified as paraffin (P), olefin (O), naphthene (N), and aromatic (A), and their fractions were defined as the percentage of each component. The gasoline fraction, kerosene + diesel (K + D) fraction, and VGO were fractionated by HTSIMDIST (Fisons) with a capillary column (7 m × 0.58 mm) coated by metal silicon. The fractions of gasoline and K + D were defined as the percentage of each component in the liquid products excluding gas and coke.
Ind. Eng. Chem. Res., Vol. 37, No. 5, 1998 1763
Figure 1. Schematic diagram of microactivity test (MAT) apparatus. (1) Microfeeder. (2) On/off valves. (3) Thermocouples. (4) Mass flow controllers. (5) Microreactor. (6) Liquid reservoir. (7) Gas reservoir. (8) Temperature controller. (9) Three-way valves. (10) Nitrogen gas. Table 1. Characterization Data for YuKong (YK Co., Korea) Vacuum Gas Oil description
VGO
Elemental Analysis (wt %) C H N
85.77 13.01 0.53
Molecular Weight number average (N) weight average (W) viscosity average (V) Z average (Z) Z/W API conradson carbon (%) IBP 5 10 50 90 95 EP
356 365 365 377 1.034 23.1 0.2
Distillation Distribution (vol %, °C) 287 342 362 440 508 523 545 Average Composition of Reaction Material (vol %)a
saturate aromatic polar a
66 33 1
HPLC analysis.
Results and Discussion Characterization of Catalyst. Figure 2 shows the adsorption and desorption isotherms of nitrogen on zeolites (SM, DM, and SY) and zeolite/aluminas (SM/A (7.8 or 9.5), DM/A (7.8 or 9.5), and SY/A (7.8 or 9.5)) measured at liquid nitrogen temperature. The isotherms indicate that SY has more adsorption capacity of nitrogen than SM or DM by a factor of about 50%. When SM was treated by HCl and steam, the adsorption capacity of nitrogen was increased about 10%.
Figure 2. N2 isotherms over modified zeolites and modified zeolite/aluminas. Table 2. Surface Areas and Pore Volumes over Modified Zeolites and Modified Zeolite/Aluminas surface area (m2/g) catalyst (SiO2/Al2O3 ratio) BET micro- external SM (6.5) DM (15.5) SY (3) SM/A (7.8) SM/A (9.5) DM/A (7.8) DM/A (9.5) SY/A (7.8) SY/A (9.5)
500 512 822 323 319 313 317 345 353
485 437 735 118 108 111 119 117 73
15 75 87 205 211 202 198 228 280
pore volume (cm3/g) total micro- meso0.2170 0.2793 0.3976 0.2947 0.3126 0.2951 0.3421 0.3461 0.4121
0.1858 0.1686 0.2835 0.0527 0.0483 0.0498 0.0535 0.0523 0.0334
0.0312 0.1107 0.1141 0.2420 0.2643 0.2453 0.2886 0.2938 0.3787
The isotherms of the zeolites show a hysteresis between adsorption and desorption and could be classified as isotherm type H4 by IUPAC for micropore structure (Gregg and Sing, 1982). DM shows more hysteresis than SM, and zeolite/aluminas exhibit more pronounced hysteresis compared to all the zeolites. The zeolite/aluminas prepared at pH 7.8 have type H2 isotherm of ink-bottle form, whereas zeolite/aluminas prepared at pH 9.5 have type H3 of slit-shaped form or platelike particles (Gregg and Sing, 1982; Lee and Ha, 1996a,b). The specific surface area and pore volume of zeolites (SM, DM, and SY) and zeolite/aluminas (SM/A (7.8 or 9.5), DM/A (7.8 or 9.5), and SY/A (7.8 or 9.5)) obtained from the isotherms are summarized in Table 2. Both data of surface area and pore volume shown in Table 2 indicate that pure zeolites (SM, DM, and SY) have mostly micropores while zeolite/aluminas have mainly
1764 Ind. Eng. Chem. Res., Vol. 37, No. 5, 1998
Figure 3. Distribution of micropore volume and mesopore volume over modified zeolites and modified zeolite/aluminas.
mesopores. The mesopore volume of the zeolite/aluminas prepared at pH 9.5 is somewhat larger compared to values observed for zeolite/aluminas prepared at pH 7.8. Figure 3 shows the distribution of micropore volume and mesopore volume over the zeolites and the zeolite/ aluminas. Zeolites (SM, DM, and SY) lie within a microdominant region below the diagonal dotted line. Among them, particularly SM has a structure in which the mesopore volume is very small. When SM is treated by HCl and steam, the mesopore volume is increased sharply to a mesopore volume similar to SY, whereas its micropore volume is decreased slightly. It is also seen from Figure 3 that when alumina containing zeolites are prepared, the mesopore volume is produced by the alumina while the micropore volume is decreased substantially due to alumina formation inside zeolite pores. The mesopore volumes of zeolite/aluminas prepared at pH 9.5 are greater than those at pH 7.8, suggesting that zeolite/aluminas prepared at pH 9.5 are expected to be better catalysts for cracking of large molecular reactants such as VGO. The pore size distributions obtained from nitrogen desorption branches on the zeolite (SM, DM, and SY) and zeolite/aluminas (SM/A (7.8 or 9.5), DM/A (7.8 or 9.5), and SY/A (7.8 or 9.5)) are shown in Figure 4. It is seen from Figure 4a, mesopores of about 4 nm were formed in steamed zeolites. One of many interesting things is that mesopores in the 4-nm range were developed further by a factor of 2 in DM compared to that in SM. Unlike pure zeolites, zeolite/aluminas prepared at both pH 7.8 and 9.5 consist of up to three peaks. For the composite catalysts, the mesopores of 5 and 6.5 nm originate from alumina. Zeolite/aluminas prepared at pH 9.5 have a much wider pore size distribution compared to those of zeolite/alumina prepared at pH 7.8. In particular, zeolite/aluminas prepared at pH 7.8 show substantial development of mesopore with a diameter of 5 nm as well as further development of mesopore with a diameter of 4 nm due to the formation of pseudoboehmite (Huang et al., 1989). A distinct change was observed in zeolite/aluminas prepared at pH 9.5 which exhibit a greater formation of mesopore with a diameter of 6.5 nm. These large pores are due to the presence of bayerite in the precipitation (Huang et al., 1989). Therefore, it is supposed that zeolite/aluminas prepared at pH 9.5 may be
Figure 4. Pore size distributions obtained by desorption isotherm branches over modified zeolites and modified zeolite/aluminas.
preferred for the catalytic cracking of heavy oil. These points are discussed in detail later in this paper. Figure 5 shows the TPD spectra of ammonia from samples pretreated at 500 or 600 °C. The zeolites (SM, DM, and SY) have two desorption peaks at about 230 and 650 °C. It would appear that the order for acid site density for the weak acid sites (occurring at 230 °C) is SY (or SY/A) > SM (or SM/A) > DM (or DM/A). The desorption peak at 650 °C in modified zeolites is remarkably small compared to the peak at 230 °C. However, the desorption peaks at 650 °C is observed to be enlarged in zeolite/aluminas which is attributed to ammonia desorption from very strong acid sites existing on the aluminas. The amount of strong acid site of SY/A is larger than that of SM/A or DM/A as shown in Figure 5. Catalytic Cracking of Vacuum Gas Oil. Conversion of VGO on the zeolites and zeolite/aluminas at 500 °C is shown in Figure 6. Conversion is defined as the weight percentage of feedstock converted to kerosene, diesel, gasoline, lighter components, and coke. For pure zeolites, the conversion of VGO shows the order SY > DM > SM. This can be explained by taking into account total acid amount, pore structure, and pore size distribution. First, higher conversion of zeolite Y is expected due to higher total acid site density of zeolite Y compared to that of mordenite. In addition to total acid site density, the structure of zeolite Y is more favorable than that of mordenite because zeolite Y has threedimensional channels with supercage and a larger pore
Ind. Eng. Chem. Res., Vol. 37, No. 5, 1998 1765
Figure 7. Conversions obtained from cracking of vacuum gas oil based on same amount of zeolite over modified zeolites and modified zeolite/aluminas. (Reaction temperature: 500 °C. WHSV: 24 h-1.)
Figure 5. Ammonia-TPD spectra over modified zeolites and modified zeolite/aluminas.
Figure 6. Conversions obtained from cracking of vacuum gas oil on modified zeolites and modified zeolite/aluminas. (Reaction temperature: 500 °C. WHSV: 24 h-1.)
volume while mordenite has one-dimensional channels. This indicates that the accessibility of the reactant in zeolite Y is better than mordenite so that the large molecular reactant can be broken easily at the active sites within the pore structure. When SM is dealuminated by HCl and steam treatment (DM), one might expect the reduction of conversion of VGO since the acid amount of DM is decreased sharply as shown in Figure 4. However, our result in Figure 6 show that the conversion of VGO is enhanced by about 20%. This is probably due to the enlargement of mesopore volume caused by structural aluminum extraction. Therefore, two opposing factors (number of acid sites and diffusion of reactant) which affect VGO conversion may compete
with each other when SM is treated by HCl and steam. In our results, the diffusion of the reactant appears to play a major role, causing the enhancement of the conversion because the fast diffusion of the reactant overwhelms the reduction of the conversion due to decreased acid site density. For the alumina mixed with zeolites, the catalysts precipitated at pH 9.5 show higher conversion than those precipitated at pH 7.5 because the zeolite/aluminas precipitated at pH 9.5 have a wider pore size distribution with favorable large mesopores of 6.5 nm. This result suggests that the development of large mesopore structure allows better accessibility of the larger molecular weight reactants to zeolite active sites and might reduce blockage of pore openings by coke deposition. To investigate the effect of alumina contained in zeolite/aluminas, Figure 7 shows the conversion of VGO based on the same amount of zeolite for pure zeolites and zeolite/aluminas. The catalysts consisting of 35 wt % zeolite and 65 wt % alumina reveal conversion values of about the same as the those of pure zeolites as shown in Figure 6. However, the effect of alumina present in the zeolite/alumina on the conversion of VGO is clearly shown in Figure 7. For example, the conversion per unit weight of zeolite for SM is about 20%/g; however, that for SM/A (7.8) is about 50%/g. Therefore, the improved activity of 30%/g between SM and SM/A (7.8) can be attributed to the presence of alumina in catalysts. This indicates that the conversion of VGO was improved by the cracking of large molecules at active sites on mesopores of alumina included in zeolite/aluminas. A further enhancement in activity per zeolite weight (15%/ g) is observed in SM/A (9.5) which has a pore structure with large mesopore diameter around 6.5 nm compared to SM/A (7.8). Similar trends are observed over DM (DM/A) or SY (SY/A). As a conclusion, zeolite/aluminas prepared at pH 9.5 had higher cracking activity for VGO than zeolite/aluminas prepared at pH 7.8 due to the larger mesopore size and wider mesopore size distribution (see Figure 4). The elemental analysis for the formed coke on catalysts is presented in Table 3. Higher amounts of coke
1766 Ind. Eng. Chem. Res., Vol. 37, No. 5, 1998 Table 3. Carbon, Nitrogen, Sulfur, and Hydrogen Amount Obtained from the Formed Coke on Modified Zeolites and Modified Zeolite/Aluminas after Cracking of Vacuum Gas Oil (Reaction Temperature, 500 °C; WHSV, 24 h-1) catalyst
carbona
nitrogena
sulfura
hydrogena
totala
SM SM/A (7.8) SM/A (9.5) DM DM/A (7.8) DM/A (9.5) SY SY/A (7.8) SY/A (9.5)
1.98 1.13 0.89 1.04 0.93 1.04 4.17 2.36 3.36
0.06
0.08 0.12 0.03 0.22 0.42 0.22 0.03 0.02 0.04
1.47 1.68 1.19 1.09 0.94 0.81 1.92 1.86 1.59
3.59 2.93 2.13 2.57 2.49 2.31 6.63 4.24 5.01
a
0.02 0.22 0.20 0.22 0.51 0.02
Elemental analysis.
Figure 9. Fractions of gasoline and kerosene + diesel obtained from cracking of vacuum gas oil on modified zeolites and modified zeolite/aluminas. (Reaction temperature: 500 °C. WHSV: 24 h-1.)
Figure 8. HT-SIMDIST spectra of commercial oil (YuKong Co., Korea).
are formed over pure zeolite than over the zeolite/ alumina (35 wt % zeolite) because the coke forms easily in the micropores of zeolite as large polymerized molecules. It is well-known that a higher density of the aluminum atom in zeolite leads to the higher hydrogentransfer rate through the inductive effect (Biswas and Maxwell, 1990; Corma et al., 1996). Therefore, SY which has a high density of acid site compared to SM produced a greater coke amount by means of an intermolecular hydrogen-transfer reaction between adsorbed alkyl aromatic species. Furthermore, dealuminated mordenite (DM) produces less coke due to the low aluminum density and wider micropores compared to SM. It is also worthy note that more coke forms over pure zeolite than over a zeolite/alumina catalyst. The liquid product distribution as gasoline, kerosene + diesel (K + D), and VGO obtained from cracking of VGO on catalysts was determined on the basis of HTSIMDIST spectra for commercial oil produced from Yukong Co., Korea (see Figure 8). The product fractions classified as gasoline and K + D are presented in Figure 9. All the catalysts used in this study produced more gasoline than K + D. Our results confirmed that the catalysts containing zeolites produce more gasoline than K + D. Therefore, zeolite is an essential component for effective gasoline production in the cracking process. Furthermore, catalysts containing zeolite Y yield a higher fraction of gasoline and K + D than catalysts containing mordenite. This probably occurs because the
Figure 10. Distribution of paraffin and olefin in the gasoline obtained from cracking of vacuum gas oil on modified zeolites and modified zeolite/aluminas.
pore structure of zeolite Y, which is three-dimensional, has a great advantage for accessing the reactant as well as higher acid site density of zeolite Y compared to mordenite. Similar tendency in the cracking activity and the selectivity of gasoline was shown in Figures 6 and 9, respectively. The distribution of paraffin (P) and olefin (O) in gasoline products is presented in Figure 10. The catalysts containing zeolite Y such as SY, SY/A (7.8), and SY/A (9.5) produced more paraffin and less olefin than those containing SM (SM/A (7.8) and SM/A (9.5)) or DM (DM/A (7.8) and DM/A (9.5)). This occurs because the high acid site density of SY contributes to a high hydrogen-transfer rate at two adjacent active sites by means of a bimolecular Eley-Rideal mechanism (Corma and Martinez-Triguero, 1994). Hence, enhancement of the hydrogen-transfer reaction over Y-type zeolite leads to higher paraffin yield and lower olefin
Ind. Eng. Chem. Res., Vol. 37, No. 5, 1998 1767
Figure 11. Fractions of benzene, toluene, xylene, and trimethylbenzene in the aromatic obtained from cracking of vacuum gas oil on modified zeolites and modified zeolite/aluminas.
yield. In the present study, the ratio of olefin to paraffin (O/P) in gasoline products changes dramatically from 0.08 in catalysts containing Y-type zeolite to approximately 3 in catalysts containing mordenite. It is interesting to note that the distribution of paraffin and olefin between pure zeolite and zeolite/alumina containing alumina at the same zeolite type is not differed largely. In other words, the distribution of paraffin and olefin is not influenced by the addition of alumina into the pure zeolite. Therefore, the distribution of paraffin and olefin in gasoline products seems to mainly depend on the zeolite type. The product distribution of benzene, toluene, xylene, and trimethylbenzene in gasoline products over the catalysts used in this study is presented in Figure 11. One of the interesting findings is that the maximum yield of xylene is observed in the catalysts containing SY consisting of the three-dimensional pore structure with large kinetic diameter (8.0 Å) whereas the catalysts containing mordenites (6.2 Å) give a maximum yield of toluene among the aromatics. It is also seen from Figure 11 that the catalysts containing SY produce a higher amount of xylene and trimethylbenzene (two branched aromatics or higher) compared to the catalysts containing mordenite, while the catalysts containing SM produce relatively more benzene (no branched aromatic). Among pure mordenites, DM yields more branched aromatics compared to SM due to the further development of 4-nm mesopore as shown from Figure 11b. All these points can be explained by the influence of difference of pore structure between SY and SM (or DM). Therefore the pore structure of zeolite included in zeolite/alumina play an important role in product distribution of benzene, toluene, xylene, and trimethylbenzene in the gasoline fraction. Conclusions Catalytic cracking of vacuum gas oil (VGO) over modified zeolites (SM, DM, and SY) and zeolite/aluminas (SM/A, DM/A, and SY/A) prepared at pH 7.8 or 9.5 was investigated at 500 °C in a microactivity tester.
The conversion of VGO on the mixed catalyst is practically the same as that over pure zeolites due to the development of a bimodal mesopore structure and a wide mesopore size distribution even though the mixed catalysts contain only 35 wt % zeolite. Furthermore, the amount of coke formed over the mixed catalyst is lower than that on pure zeolites. The catalysts containing SM or DM show lower activity than those containing SY. However, the catalysts containing DM, which has a higher fraction of mesopores compared to those existing in SM, show higher activity and higher selectivity for gasoline. Especially, mixed catalysts contain alumina with a wide mesopore size distribution (such as SM(DM or SY)/A (9.5)) yield higher activity than those with alumina having a narrow mesopore size distribution, such as SM(DM or SY)/A (7.8). The selectivity of olefin in the gasoline fraction is higher on the catalysts containing mordenite than on those containing zeolite Y and is a reverse correlation with the selectivity of paraffin by means of a hydrogentransfer reaction. It is found that there is no distinct difference in the selectivity of paraffin and/or olefin in the gasoline product when alumina is introduced to the pure zeolite. The distribution of paraffin and olefin in gasoline products seems to mainly depend on the zeolite type. For the aromatics fraction obtained on various catalysts, the highest yield of xylene was obtained in the catalysts containing SY whereas the maximum yield of toluene was obtained in the catalysts containing mordenites due to the difference of pore structure of zeolite included in zeolite/alumina catalysts. The catalysts containing mordenite with a developed mesopore (DM) produced a higher fraction of branched aromatics compared to the catalysts containing mordenite with a high proportion of micropore (SM). Acknowledgment The financial support of this work by the Research Center for Catalytic Technology (Pohang University of Science and Technology) is gratefully acknowledged. Literature Cited Biswas, J.; Maxwell, I. E. Recent Process- and Catalyst-Related Developments in Fluid Catalytic Cracking. Appl. Catal. 1990, 63, 197-258. Corma, A.; Martinez-Triguero, J. Kinetics of Gasoil Cracking and Catalyst Decay on SAPO-37 and USY Molecular Sieves. Appl. Catal. A 1994, 118, 153-162. Corma, A.; Martinez, A.; Martinez-Soria, V.; Monton, J. B. Hydrocracking of Vacuum Gas Oil on the Novel Mesoporous MCM-41 Aluminosilicate Catalyst. J. Catal. 1995, 153, 25-31. Corma, A.; Grande, M. S.; Gonzalez-Alfaro, V.; Orchilles, A. V. Cracking Activity and Hydrothermal Stability of MCM-41 and its Comparison with Amorphous Silica-Alumina and a USY Zeolite. J. Catal. 1996, 159, 375-382. Gregg, S. J.; Sing, K. S. W. Adsorption, Surface Area and Porosity 2nd ed.; Academic Press: New York, 1982. Guerzoni, F. N.; Abbot, J. Catalytic Cracking of a Gippsland Reduced Crude on Zeolite Catalysts. J. Catal. 1994, 147, 393403. Guisnet, M.; Magnoux, P. Coking and Deactivation of ZeolitesInfluence of the Pore Structure. Appl. Catal. 1989, 54, 1-27. Huang, Y.; White, A.; Walpole, A.; Trimm, D. L. Control of Porosity and Surface Area in Alumina-Effect of Preparation Conditions. Appl. Catal. 1989, 56, 177-186 Lee, K. H.; Ha, B. H. Characteristics of Mordenite/Alumina and Their Cracking Activity for n-Decane. Hwahak Konghak 1996a, 34 (3), 363-368.
1768 Ind. Eng. Chem. Res., Vol. 37, No. 5, 1998 Lee, K. H.; Ha, B. H. Preparation and Characteristic for γ-AluminaEffect of Precipitation pH and Calcination Temperature. Hwahak Konghak 1996b, 34 (1), 28-35. Satterfield, C. N. Heterogeneous Catalysis in Practice; McGrawHill Book Company: New York, 1980. Scherzer, J.; McArther, D. Catalytic Cracking of High-Nitrogen Petroleum Feedstocks; Effect of Catalyst Composition and Properties. Ind. Eng. Chem. Res. 1988, 27, 1571-1576. Smirniotis, P. G.; Ruckenstein, E. Comparison of the Performance of ZSM-5, β Zeolite, Y, USY, and Their Composites in the Catalytic Cracking of n-Octane, 2,2,4-Trimethylpentane, and 1-Octene. Ind. Eng. Chem. Res. 1994, 33, 800-813. Song, C.; Nihomatsu, T.; Nomura, M. Effect of Pore Structure of Ni-Mo/Al2O3 Catalysts in Hydrocracking of Coal Derived and
Oil Sand Derived Asphaltenes. Ind. Eng. Chem. Res. 1991, 30, 1726-1734. Tischer, R. E.; Narain, N. K.; Stiegel. G. J. S.; Cillo, D. L. Largepore Ni-Mo/Al2O3 Catalysts for Coal-Liquids Upgrading. J. Catal. 1985, 95, 406-413. Venuto, P. B.; Habib, E. T., Jr. Fluid Catalytic Cracking with Zeolite Catalysts; Marcel Dekker: New York, 1979.
Received for review September 30, 1997 Revised manuscript received February 5, 1998 Accepted February 7, 1998 IE970699Q
