Active Site Accessibility of Resid Cracking Catalysts

Studies in Surface Science and Catalysis 134 M.L. Occelli and P. O'Conner (Editors) 9 2001 Elsevier Science B.V. All rights reserved

209

Active Site Accessibility of Resid Cracking Catalysts Yong Lu, Mingyuan He, Jiaqing Song, and Xingtian Shu Research Institute of Petroleum Processing, China Petrochemical Corporation, Beijing 100083, P. R. China It is well understood that pore structure is important in designing FCC catalysts. For residue cracking, the pore structure of catalyst matrix should be accessible for precracking the large hydrocarbon molecules so that the resulting smaller molecules can transfer into the zeolite channels and can be converted to value-added products over the zeolite sites. However, pore parameters such as pore volume and pore size distribution are insufficient for selecting an appropriate matrix. A method for testing the active site accessibility of resid cracking catalysts has been developed on currently widely used MAT equipped with the unstripped hydrocarbon burned off in place, with an online measurement of the CO2 and H20 that are produced. The effect of pore distribution of the FCC catalyst matrix on the active site accessibility of several catalysts has been tested tentatively in terms of bottoms conversion and strippable coke yield. The results showed that the yields of HCO and strippable coke increase when the matrix small pores increase, and decrease when the catalyst matrixes are highly accessible. These test observations are in agreement with commercial results. Key Words: resid cracking, catalyst, active site accessibility, strippability 1. INTRODUCTION Maintaining high conversion level and proper yield structure for a resid FCC unit (RFCC) is essential for maximizing the profitability of a refinery. The catalyst choice is certainly a key factor in the operation. Unlike the conventional FCC, where cracking reactions are usually not diffusionally limited, there are observable diffusion restrictions for large size molecules within zeolite channels and activity loss in RFCC units [ 1]. Recently, Khouw et al. [2] tested a vanadium contaminated equilibrium catalyst from a RFCC unit. The results have shown that the conversion level for a feedstock with high CCR content drops dramatically. Zhao [3] investigated the deactivation of FCC catalysts by sodium in commercial operations and found that the surface area of the matrix is more strongly affected than that of the zeolite. Mann et al [4] studied the mass transfer into catalyst particles using a special technique on SEM. The results illustrated the fact that the accessibility of a typical catalyst is far from perfect. Hence, for residue cracking, the pore structure of the catalyst matrix should be accessible

210 for precracking the large molecules so that the resulting smaller molecules can be converted to FCC products within the zeolite channels. However, pore parameters such as pore volume and pore size distribution are insufficient for selecting an appropriate matrix. To measure the catalyst accessibility, Akzo Nobel has developed a method based on the non-steady state diffusion of hydrocarbons into FCC catalyst particles, which measures a relative rate of mass transfer called the AAI (Akzo Nobel Accessibility Index) [5]. Also, the cracking of TIPB has been used as a test reaction to study the activity and accessibility of zeolites. The results showed that both seem to control the rate of reaction [6]. In the present paper, a method for testing the active site accessibility of resid cracking catalysts has been developed through modification of currently widely used MAT equipment. The effect of the pore size distribution of the catalyst matrix on the active site accessibility of several catalysts, has been tentatively tested in terms of bottoms cracking and strippable coke yield. 2. EXPERIMENTAL

2.1. Catalysts Two catalysts, designated as M-1 and M-2, were prepared with equal zeolite contents and with matrixes having different pore structures. For this purpose, A-1 (7-A1203, SA: 275m2/g, PV: 1.22ml/g, Crystallinity: 85%) and A-2 (7-A1203, SA: 266m:/g, PV: 0.286ml/g, Crystallinity: 85%) were used as the matrix for M-1 and M-2 respectively. The pores in A-1 with diameter of >80A share ca. 88% of the total pore volume, but the pores in A-2 with diameter of <80A share ca. 85% of the total pore volume. The typical preparation is as follows. The catalyst M-1 was prepared by mixing 6.0 g (dry base) of A-I, 10.5 g (dry base) of kaolin clay, 14.1 g of A1 sol-gel (21.34 wt% of A1203) and 46 ml of distilled water, and stirring the mixtures for 10 minutes. The pH of the mixture was adjusted to ca. 3.0 by slowly adding 2.0g of HC1 (19% by weight) and stirred continuously for 0.5 hour. 6.0g of USY and 3.0g of REUSY were added to the resulting slurry, followed by vigorously stirring for 0.5 hour. The final slurries were dried and calcined at 523 K in air for 3 hours. Four commercial fresh catalysts (A, B, C, D) and three corresponding commercial equilibrium catalysts (designated as E-A, E-C and E-D), were also employed in this study. The fresh catalysts were deactivated at 1073 K for 4 hours with 100% steam in a fixed bed reactor and the equilibrium catalysts were calcined at 823 K for 2 hours under air before testing. 2.2. Characterization Nitrogen adsorption-desorption analyses were performed for all catalysts to measure the surface area and pore structure. Table 1 shows that the matrix surface areas of the catalysts are in the range of 33-6 lm2/g. The micropore volumes are between 2.2-6.1 x 10.2ml/g. In-situ IR analyses for adsorbed pyridine were conducted for measuring the quantity of

211 Bronsted and Lewis acid sites. Table 2 shows that the concentrations of the acid sites measured by IR spectra of pyridine adsorbed at temperatures 473 K (623K) ranks as follows" Bronsted acid site, A>C-B (A-B>C); Lewis acid site, B>C-A (A>C>B). On the whole, there are no significant differences in the acidity of the tested catalysts. Table 1 The Pore structure of the catalysts Total surface Matrix surface Catalyst . area, m2/g area, m 2 / g M-1 185 61 M-2 187 57 A 122 63 B 122 45 C 182 57 D 125 46 E-A 95 48 E-C 101 33 E-D 96 35 .

.

.

.

.

.

.

.

.

.

.

.

Total pore volume, cm3/g 0.253 0.127 0.180 0.156 0.139 0.135 0.161 0.128 0.100 .

.

Microp0re volume, cm3/g 0.052 0.061 0.028 0.036 0.059 0.037 0.022 0.032 0.028

Probable pore size, A 108 76 121 40/109 39/82 39/83 112 40/109 40/88

,

Table 2 The acid sites concentrations of the commercial FCC catalysts ' Catalyst A B C D

473 K B acid, ~tmpl/g 35.59 28.81 28.81 n.d.*

L acid, gmol/g 26.19 30.95 26.19 n.d.

6 2 3 I( B acid, lamol/g 11.86 11.86 10.17 n.d.

" " L acid, ~tm.ol/g 15.48 13.10 14.23 n.d.

* not detected. 2.3. Reactor and Procedures The cracking experiments were carried out in a downflow MAT (microactivity test) unit designed for handling heavy feeds. The unit consists of a fixed-bed reactor made of stainless steel, a syringe pump, a hot box, a liquid product collector, a liquid displacement column and a wet test meter. Catalyst loading was always 5 grams. The amount of heavy oil feedstock (the properties were summarized in Table 3) injected was 1.56 grams. Cracking was performed at a reactor temperature of 773 K and a time on stream of 70 seconds. The stripping time was between 30 and 900 seconds, using high purity N 2 as stripping gas with a flow rate of 30 ml/min. The liquid products were characterized by high-temperature gas chromatograph distillation (GCD). Gas compositions were analyzed by GC.

212 Table 3 The properties of the heavy oil feedstock Density, (g/cm 3, 293 K) Viscosity, (mm2/s) 353 K 373 K Solidifying Point, (K) Aniline Point, (K) Acidity, (mgKOH/g) Basic Nitrogen, (ppm) CCR, (wt%) Composition, (wt%) Parraffin Aromatics Resins Distillation, (K) 5% 50% 90% ,

0.8916 9.629 6.312 320 366.5 0.81 404 0.28 63.6 29.2 7.2 616 715 791

The amount of coke formed on the catalysts was determined by in-situ combustion, using a CO2 IR Carbon Detector. An on-line CO oxidation reactor (873 K) packed with cobalt oxide catalyst was added between the reactor and CO2 IR Carbon Detector to ensure the analyzed product was CO 2. Hydrogen content of coke (designated as H) was determined by combustion. Before the combustion of the coke was started, two weighed glass tubes packed with Mg(C104) 2 particles, were connected to the reactor outlet and the CO oxidation reactor outlet respectively. The product HzO formed during coke combustion could be completely adsorbed by Mg(CIO4)z and its amount determined by weighing. By this way, H (hydrogen content of coke) was calculated. Table 4 shows how reproducible the H measurements are. Table 4 The replicate measurements of the hydrogen content of coke (/-/) over catalyst A ,

,

,m

Test number

Hydrogen content of coke (/-/) Commercial spent A Spent A in MAT test 1 8.4 7.6 2 8.2 7.8 3 8.3 7.4 4 8.4 7.5 a After a 15rain-Nitrogen-Stripping

a

....

To calculate the yield of strippable coke, an intrinsic parameter for the strippability of FCC

213 catalysts, the coupled equations (Equation 1 and Equation 2) given below was set up on the basis of the mass balance of C and H in coke. Hx Y = Hsc x Ysc + Hrc x Yrc Y= Ysc + Yrc

1) 2)

where Y, Ysc and Yrc represent yields of total coke, strippable coke, and reaction coke (catalytic coke + contaminative coke + feed CCR coke) respectively; H, Hsc and Hrc represent the hydrogen content of corresponding cokes respectively. According to literatures [7,8], Hsc and Hrc were assigned values of 12 and 5 respectively. 3. RESULTS AND DISCUSSION 3.1. Strippability of tested catalyst Table 5 shows the yield of strippable coke with different stripping time lengths. The yield of strippable coke closely increases in the following sequences: M-l, M-2; A, B-C, D; E-A, E-C, E-D. If the yield of strippable coke characterizes the strippability of an FCC catalyst, then the strippability of the tested catalysts can be rank-ordered. The results can be approximately correlated with the total pore volume and average pore size, but not with the total matrix surface areas as shown in Table 1. In the catalyst group of A, B, C, and D, the catalyst A has the largest matrix surface area, but gives the lowest yield of strippable coke.

Table 5 The yield of strippable coke after stripping for different time lengths Catalyst Yield of Strippable Coke After Stripping For Different Time Length, wt% 30 s 480 s 900 s M1 12.72 4.27 2.96 M2 11.52 3.90 2.90 A 9.05 2.39 1.52 B 9.36 4.15 2.83 C 10.60 5.37 1.89 D 13.42 5.73 2.89 E-A -1.90 1.15 E-C -2.89 1.58 E-D ~ 3.97 2.34 3.2. Effect of pore size distribution of catalyst matrix on reduction of strippable coke Fig. 1 shows the relationship between the yield of strippable coke and the pore diameter distribution of the tested catalysts. In the tested catalysts, the specific surface areas in pores in 20-50 A range increase in following orders: MI
214 pores in the 50-200A and 200-1000A range, the order (except for B) is reversed. The yield of strippable coke generally increases when the area in the small pore (20-50A range) increases, and generally decreases when large pores (>50A) in the catalyst matrix increase. Yanik et al. [9] reported that the initial soft coke increased with the increase of zeolite content and the content of small pores in the matrixes. Their results also show that a catalyst based on a moderate zeolite content and a highly accessible large pore matrix system can reduce the quantity of adsorbed hydrocarbons and hence the soft delta coke.

20 r 18

7

t

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~. lO o

8

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0

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B

C

D

E-A E-C E-D

Catalyst ,n I~

SA in 20-50 A range pores + stripping for 900s

stripping for 480s

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o

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Catalyst n

SA in 50-200A range pores

i

i "-- sb'ipping for 480s

SA in 200-1000A range pores stripping for 900s

Fig. 1. Relationship between the yield of strippable coke and the pore size distribution of catalyst matrixes

215 Fig. 1 also shows that although catalyst B has fewer small pores than catalyst C, its strippability is poorer. Table 2 shows that catalyst B possesses more Lewis acid sites than catalyst C, which may outweigh the effect of pore structure. Pavel [10] has noted that the acidity of a catalyst is an important factor in its strippability. 3.3. Effect of pore size distribution of catalyst matrixes on heavy oil cracking ability Table 6 shows the tested catalysts activity and selectivity for heavy oil cracking. The heavy oil cracking ability of the tested catalysts decreases in the following order: M-I
a

,

,, ,,

Specific Coke b 23.40 20.60 2.16 2.20 2.46 2.78 2.58 2.67 5.06 5.24 5.33

216 3.4. Commercial data Table 7 shows the commercial performance of catalysts A and C. Clearly, when resid FCCU is using catalyst A (with more accessible pore system), a high yield of higher-value FCC products is achieved with a slight reduction of HCO yield and with a remarkable decrease in hydrogen content of the generated coke.

Table 7 The Commercial results of FCC catalysts Catalyst A Density of Feed, g/cm 3 0.9043 E-cat. Property Conversion, wt% 56 Ni on catalyst, ppm 3885 V on catalyst, ppm 8700 Cat. Addition, tons/day 1.35 Yield structure, wt% Dry gas 3.97 LPG 9.03 Gasoline 43.11 LCO 29.81 HCO 4.60 Coke 9.48 Hydrogen content of coke, wt% 8.89

C 0.9108 57 3500 8700 1.40 4.31 10.22 43.13 28.08 4.92 9.34 9.90

4. CONCLUSIONS A method for testing the active site accessibility of resid cracking catalysts has been developed through modification of currently widely used MAT equipped with the unstripped hydrocarbon burned off in place, with an online measurement of the CO2 and H20 that are produced. The bottoms conversion and the strippable coke yield, which are characteristics of the active site accessibility of the FCC catalyst, can be determined by the proposed method. Test results show that the pore size distribution of the catalyst matrix is a key factor in improving active site accessibility of an FCC catalyst. With the increase of the surface area or pore volume in the matrix small pores, the strippable coke yield and the bottoms yield increases. However, the catalyst with a highly accessible large pore matrix exhibits good performance in reducing the strippable coke yield and in enhancing bottoms conversion. Test results are well in agreement with commercial results of the catalysts with respect to bottoms cracking and hydrogen content of coke. However, for optimizing the active site accessibility of FCC catalysts, a better understanding of the interactions between the catalyst pore parameters and heavy oil molecules is required. Moreover, the development of novel porous materials is right now the

217 subject of further thinking and experimentation. REFERENCES

1. P. O'Connor and A. P. Humphries, Am. Chem. Soc. Div. Petrol. Chem. Preprints, No. 38(3) (1993) 598. 2. F. H. H. Khouw, M. J. R. C. Nieskens, M. J. H. Borley, and K. H. W. Roebschlaeger, NPRA 1990 Annual Meeting, Paper No. AM-90-42 (1990). 3. X. Zhao and W. Cheng, in: P. O'Connor and T. W. Takatsuka (eds.), ACS Symposium Series 634, American Chemical Society, Washington, D C, Chap 11 (1995) 159. 4. R. Mann, K. Khalaf, and A. Lamy, in: P O'Connor and T. W. Takatsuka (eds.), ACS Symposium Series 634, ACS, Washington, D C, Chap 3 (1995) 42. 5. A. Humphries, R. P. Fletcher and J. R. Pearce, NPRA 1999 Annual Meeting, Paper No. AM-99-63 (1999). 6. S. Falabella, E. Aguiar, M. L. Mur-Valle, E. V. Sobrinho and D. Cardoso, in: Bonnevoit, S. Kaliaguine (eds.), Zeolite: A Refined Tool for Designing Catalytic Sites, Elsevier, Amsterdam, (1995). 7. H. C. Kliesch, et al., Paper presented at the Ketjen catalysts Symposium, (1987). 8. J. L. Mauleon and W. S. Letzseh, Paper Presented at the 5th Katalistiks Annual FCC Symposium, Chap 7 (1984). 9. S. J. Yanik and P. O'Connor, NPRA 1995 Annual Meeting, Paper No. AM-95-35 (1995). 10. S. K. Pavel and F. J. Elvin, NPRA 1998 Annual Meeting, Paper No. AM-98-42 (1998).