I
MARVIN F. L. JOHNSON, WILLIAM E. KREGER, and HENRY ERICKSON Sinclair Research Laboratories, Inc., Harvey, 111.
Gas Oil Cracking by Silica-Alumina Bead Catalysts Pore Size and Particle Size Effects ,Experimental data on median pore size, pore volume, surface area, and product distribution for various silicaalumina bead catalysts are of value in predicting the performance of a catalyst. Reduction in particle size or increase in pore size and pore volume was found to improve cracking activity and selectivity.
ONE
of the factors affecting the performance of a catalyst is the rate a t which reactant and product molecules can diffuse within a particle, relative to the rate of reaction. This theory has been discussed by Wheeler (4, 75) and by Weisz (72). Wheeler points out that in the catalytic cracking of gas-oils by silica-alumina at about 500' C., particle size has little effect-Le., the surface of '/a inch pills is about 75% available. The present authors have examined this field in more detail. For a study of this type, bead catalysts are ideally suited, because of the absence of macropores which result from tabletting or extruding powders. There is thus only one pore structure to consider. Description of Catalysts
I is a commercial bead (BC), received from Socony-Vacuum Oil Corp., which contains about 10% alumina. Cat. A index (7) = 40. I1 was prepared from I by steam deactivation for 74 hours a t 594" C. and atmospheric pressure. I11 was prepared from a bead hydrogel, received through the courtesy of Socony-Vacuum Oil Corp. It was soaked 3 hours in a solution of 0.570 ammonium fluoride, drained, dried at 110" C., and calcined in a muffle furnace for 2 hours a t 565' C., plus 10 hours at 760" C. 7 0 A1203 = 10.2; 7 0 F = 0.73. IV is a Socony-Vacuum commercial bead, containing 0.1% chromium as a carbon-burning-rate promoter. A1203 = 10.4. Cat. A index (7) = 41. V was prepared in the same way as 111, but with a chrome bead hydrogel, with a 1% ammonium fluoride solution; it was calcined 3 hours a t 594" C. yo A1203 = 8.3; % Cr = 0.11; % F = 1.3.
VI was prepared by a "semiaerogel" technique, starting with the hydrogel beads used for 111. After air-drying a t 25" C . to 80.570 weight loss, the bead hydrogel was repeatedly washed with 100% methanol until the supernatant solution contained only about 1.5Y0 water. The material was then placed in a pressure vessel, and the temperature increased until the critical conditions were exceeded. Thereupon, the methanol pressure was slowly released to the atmosphere. The catalyst was cooled, then calcined 10 hours a t 594" C. in a muffle furnace. This technique is a modification of Kistler's Method (6) for preparing aerogels. For cracking studies, 4- to 5-mesh fractions of each catalyst were segregated. This is equivalent to an average particle size of 0.44 cm., which is larger than the average of the 4- to 10-mesh beads used commercially. Four of the catalysts were partially ground and the 28- to 35-mesh fraction was used in cracking studies; the average particle size is taken to be 0.057 cm. I was further ground to 100to 170-mesh, corresponding to an average particle size of 0.01 1 cm. I1 and IV were not studied after grinding. Pore Structure of Catalyst
For each catalyst, a complete adsorption-desorption isotherm for nitrogen a t -195.5' C. was determined (Figures 1 to 3). The technique used in these laboratories has been described (9). By application of the Barrett, Joyner, and Halenda ( 2 ) method to the desorption isotherms, pore distributions were calculated and plotted as cumulative pore volume against pore radius. These curves demonstrated that the pores were narrowly distributed about a mean; therefore, volume median pore radii and total pore volumes adequately describe thr pore structures of these catalysts. Table I lists median pore radii, total
Table I.
I BET area, sq. m./g. Total pore vol., cc./g. Median pore radius, A.
338 0.490 28
pore volumes, and surface areas (BET) derived from the isotherms. Commercial materials I and IV are similar to the bead catalyst used in sintering studies (77). Upon steam treatment of I to obtain 11, the pore size is increased and area is decreased, with but little decrease in pore volume. This result is similar to that previously observed (7,s). VI has not only large pores but also a large pore volume. This is a consequence of conducting the preparation to decrease shrinkage during drying of the hydrogel ( 6 ) . Without preliminary drying, pore size and pore volume would have been even larger, but the beads would have been too fragile. The pore structure of V is very similar to that of VI. Addition of ammonium fluoride to the hydrogel has apparently produced a finished catalyst with large pores and a large pore volume. A mechanism involving the inhibition of shrinkage during drying is suggested as an explanation, but nothing further is .known. 111, similarly prepared with less fluoride, has a larger median pore size but a lower pore volume and surface area than V. Part of the differences might be associated with differences in calcination temperature, as 111 after 565' C. calcination had smaller pores and a higher area than after 760' C. calcination. The pore structures determined for the whole beads are also considered to represent the structures of the ground materials. A moderate amount of grinding does not alter pore structure appreciably (77). Cracking Experiments
Apparatus and Procedure. The equipment used for the cracking studies
Pore Structure Data I1 I11 IV 269 0.470 31.5
154 0.615 71.5
371 0.486 25.5
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V
VI
266 0.860 52
310 0.904 51
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RELATIVE PRESSURE, P/Po
RELATIVE PRESSURE, P/Po
Figure 1.
Nitrogen adsorption-desorption isotherms
is typical of those in general use (3). Catalyst (42.6 grams) plus 100 ml. of 10to 20-mesh tabular alumina is charged to a I-inch (inside diameter) reactor. The catalyst is mixed with 25-ml. portions of alumina in such a way that the bottom portion contains 21.6 grams of catalyst, the next 12 grams, the next 6 grams, and finally 3 grams. T h e reactor is placed in a n aluminum-bronze block, held at 482" C. by Nichrome windings. After purging with nitrogen while the catalyst is attaining the desired temperature, gasoil at a constant measured rate is pumped into the top of the reactor, where it passes through a helical plug preheater before coming in contact with the catalyst. The properties of the gas-oil used are shown in Table 11. The liquid effluent is collected in a flask held at 15" C., while the gaseous product passes through a 15" C. condenser, and is collected over
Table II.
Gas-Oil Properties
Source
East Texas
200-cc. distillation
I.B.P. 50%
500' 590
E.P.
700
Gravity, OAPI
% s. Aniline No., F.
284
34 0.3
175
F.
Figure 2.
Nitrogen adsorption-desorption isotherms
water. At the end of a 30-minute process period, the catalyst is purged with a measured amount of nitrogen (900 ml,), which is added to the gaseous product. Regeneration is then started by passing air through the catalyst, slowly at first, but at increased rates until about 6.5 liters per hour is reached. A4fterabout 30 minutes. the block temperature is increased to 620" C.. and maintained until regeneration is complete, as evidenced by a negative test for carbon dioxide in the effluent; this may require 3 to 5 hours at 620" C. Regeneration gases are passed through a copper oxide conversion catalyst at 450" to 500' C., to convert carbon monoxide to carbon dioxide, through a dryer, and through U-tubes containing Ascarite absorbent. The weight of carbon on catalyst produced during a test is therefore determined by the change in weight of the Ascarite. The weight of gas produced is obtained from a measurement of its volume and density, corrected for the nitrogen used in purging. Mass spectrometric analyses have established that, although the gas contains some C j and Ce hydrocarbons, the liquid contains some CBand Cd. Consequently, the weight of gas as measured corresponds closely to the weight of Cd and lighter products, and can be considered as a convenient measure of conversion to gas.
INDUSTRIAL AND ENGINEERING CHEMISTRY
The weighed synthetic crude is distilled in a Cat-A microstill (7) until the overhead reaches 410" I?. ; distillation time ranges between 45 and 60 minutes. Upon cooling, the bottoms are weighed, reported as unconverted gas-oil; the weight of overhead, obtained by difference, is reported as gasoline. I n nearly all cases the sum of gas, coke, gasoline, and unconverted oil was within 1% of the weight of oil charged; tests in which the losses exceeded 2y0were discarded. T h e percentage total conversion and conversion to coke, gas, and gasoline were calculated on the basis of total products recovered. Effects of "flush" activity were eliminated by discarding the results of initial tests on a fresh catalyst whenever latcr tests a t the same space velocity showed a decrease in activity. Precision. The usual procedure of estimating standard deviations from a mean could not be applied to these data because of the varying space velocity. The following procedure was therefore adopted. For the 11 catalysts studied 60 runs were made, with a minimum of four per catalyst. Plots were made of percentage total conversion us. WHSV-l, percent gasoline, percent gas, and percent coke: smooth curves were drawn visually, ex-
P R E P A R I N G CATALYSTS
IN T H E L A B O R A T O R Y
RELATIVE PRESSURE, P/Fb
Figure 3.
Nitrogen adsorption-desorption isotherms
tending to the origins. For each experimental point, the difference between the observed value and the corresponding value on the smoothed curve was taken as the deviation, A. The estimated standard deviation (c)for each type of plot is then given by :
where N = number of points = 60, and P = number of plots = 11. The factor of 2 in the denominator was used because the curves were approximately second degree, restricted to going through the origin. The following values of 2 F were calculated in this fashion: 70 70
total conversion
gasoline
% gas yo coke
i l 9 10.9 3~0.9 2t0.2
These values indicate a significantly greater precision compared to the conventional cracking tests, such as Cat-A. I t is felt that this improvement was the result of careful attention to detail, particularly in the distillation step. These values mean that an observed difference greater than 2 u may be said to be real at the 95y0 probability level. I n no case was the deviation of an experimental point from the curve greater than 2.5 U ;
for only eight of the 60 runs was it as great as 2 U . The ideal method of establishing standard deviations would involve the use of a family of curves fitting the same general expression, for each type of plot. However, the mechanism of cracking is not sufficiently known to permit the assumption of a particular form for the equation. Since such an assumption would introduce another uncertainty, the degree of improvement in the analysis would not be warranted. I t is entirely possible, however, that the “true” standard deviations could be somewhat dif. ferent from the above. I t is difficult to believe the difference would be great, in view of the usual observation of the agreement between two duplicate runs. Results For each sample of whole bead and ground catalyst a series of cracking runs a t varying space velocities was made. For each series, l/WHSV was plotted against total conversion. In Figure 4,W refers to 4- to 5-mesh beads, G to 28- to 35mesh. By definition, total conversion is conversion to coke plus material boiling below 410’ F. This definition is arbitrary, but represents conventional practice. The space velocity required to obtain a given conversion may be consid-
ered a measure of relative activity. For present purposes, the 40% conversion level was chosen, because it is in a region sensitive to change and is close to many actual data points. T o perform the interpolation, pseudo-first-order plots of 1/WHSV us. -log (1 -a) were used, where 01 = fractional conversion, because this type of plot is more nearly linear than those of Figure 4. Values so determined are included in Table 111. In actual practice the activity level of a catalyst is less important than its selcctivity. Because the desired product of catalytic cracking is gasoline, a catalyst that produces more gasoline at a given level of total conversion is considered more selective. Selectivity is defined here as the gasoline yield a t 4@y0total conversion. T o compare selectivity, total conversion is plotted against conversion to gasoline, gas, and coke (Figures 5, 6, and 7). Values a t the 4@y0conversion level were obtained by interpolation, and are listed in Table 111. The reduction in particle size of catalyst I from 4-5 mesh to 28-35 mesh increased activity almost twofold and increased selectivity. Further grinding produced no detectable effect upon either activity or selectivity. This effect on selectivity is somewhat startling, but there can be no doubt of its occurrence. Similar effects have been reported for other systems (73). The increase from 27.3 to 31.5y0 gasoline is beyond the limits of experimental error. Four experimental points were obtained for I-W, six for I-G; in each case two were near the interpolation level of 40% conversion. The curves for I-W were also somewhat substantiated by the data for IV-W, which had the same Cat. A activity index. Catalyst 11, produced by a steam treatment of I, has a lower activity per unit weight, but slightly higher activity per unit area than I-W; the selectivity was apparently improved, but within experimental error. These small changes are consistent with the fact that the pore size increase was relatively small. Reducing the particle size of the large-pore catalysts, 111, V, and V I caused an appreciable increase in activity and a small improvement in selectivity. Although the observed selectivity increases (0.6 to 0,7y0 gasoline) might individually be considered within limits of error, observation of the same small increase for three catalysts makes it appear to be real, and to be due to pore structure. Pore structure thus appears to alter the effect of particle size. I t is more difficult to assess the effect of pore size alone, because of differences in intrinsic activity of the catalysts. The activities of the 28- to 35-mesh materials may be considered representative of intrinsic activities, as further grinding of I proVOL. 49, NO. 2
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t-
o o 10
I
20
WT%
WHSV
Figure 4.
Effect of space velocity on total conversion
20
Figure
5.
4.0
30
5
0
6
0
70
CONVERSION
Conversion to gasoline
0-JW
I
'Q- SG
1
1
0
i
p
3
0
WT.%
Figure 6.
4
0
5
0
6
0
7
0
W T % CONVERSION
CONVERSION
Figure 7.
Conversion to gas
Table 111. Catalyst designationsb Particle size Screen size, mesh Av. diameter, cm. At 40% total conversion (wt.) WHSV, g./g./hr. Gasoline, % Gas, % Coke, % WHSV X 100 Area
-
I-W 4-5 0.44 2.3 27.3 10.9 1.8
..
Effectiveness factor, f 0.55 At 900' F., atmospheric pressure. W whole; G ground.
286
Conversion to coke
Gas-Oil Cracking"
I-G
I-G'
11-W
111-W
111-G
IV-W
V-W
V-G
VI-W
VI-G
28-35 0.057
100-170 0.011
4-5 0.44
4-5 0.44
28-35 0.057
4-5 0.44
4-5 0.44
28-35 0.057
4--5 0.44
28-35 0.057
4.16 31.5 7.7 0.8 1.23
4.18 31.5 7.7 0.8 1.23
1.94 27.8 10.7 1.5
1.80 31.4 7.7 0.9
...
2.98 32.0 7.3 0.75 1.93
2.17 26.6 11.5 1.9
3.82 32.5 6.8 0.67
5.28 33.2 6.2 0.56 1.98
1.82 28.9 9.9 1.2
...
2.28 30.5 8.6 0.9 0.74
0.98
1.0
..
0.60
0.98
...
0.71
0.99
0.79
0.99
INDUSTRIAL AND ENGINEERING CHEMISTRY
*.
...
..
PREPARING CATALYSTS IN T H E LABORATORY duced no additional effect; these values are included in Table 111, corrected to a unit area basis. The selectivities for the 28- to 35-mesh materials increase as intrinsic activity increases; the estimated standard deviation from a plot of percent gasoline us. intrinsic activity for the four ground catalysts is 0.5%, equal to the estimated standard deviation of the measurement itself. As a t least part of the differences in selectivity is due to inherent differences in the catalytic surfaces, we cannot determine how much improvement in selectivity results from pore enlargement alone. The fact that the selectivity of I-G is better than that of VI-G, presumably because of intrinsic activity differences, while I-W is poorer than VIW, suggests that the pore structure of the latter has a beneficial effect.
Discussion The process of catalytic cracking to produce gasoline as the desired product may be considered an example of Wheeler’s (74) Type I11 selectivity: A 4 - 1 - B - k p C where B is the desired product. Although not all of the gas and coke formation are necessarily derived from secondary cracking of gasoline fractions, a sizable portion of these products is probably a product of gasoline cracking. Wheeler points out that the selectivity to the desired intermediate should improve as the diffusion modulus decreases (smaller particles and/or higher D,). This is essentially observed here. The fact that differences in selectivity are found among the 28- to 35-mesh material proves the existence of intrinsic differences in selectivity, or the ratio kl/kz, as diffusional effects have been shown not to be controlling in this range of particle sizes. I t may be concluded that the differences in selectivity among the whole beads are due in part to intrinsic differences, and in part to diffusional effects which disappear upon reduction of particle size. The extent of improvement in selectivity upon grinding is therefore a measure of the role of diffusion. As expected, the improvement is greatest for small-pore I, much less for 111, V, and VI. In considering the effects of pore structure and particle size, the concept of “effectiveness factor” ( 5 )is used, variously , 17 (72). This designated E A ( 5 ) , f ( 7 4 )or is defined as the ratio of the observed reaction rate to that which would be ob-
Table IV.
served if no diffusional effects were operating. Values of f are measured directly by comparing rates of reaction for two particle sizes of the same catalyst (5, 72). I n the present case, WHSV a t 40% conversion is used as the measure of reaction rate. As shown in Table 111, f is essentially unity for the 28- to 35-mesh materials, confirmed by further grinding of I to 100-170 mesh. For the 4- to 5mesh beads, f is 0.55 for I, increasing to 0.79 for VI, as pore size and pore volume increase. f is calculated from data on a single particle size of catalyst, by using the diffusivity modulus, originally defined by Thiele (70). Wheeler (74) and Weisz (72) have defined the moduli in an analogous fashion, and have derived corresponding relationships between f and their moduli. According to Wheeler, the modulus is: h = R/3 ( k / D e ) l / 2 ,where R = radius of spherical particles, De = effective diffusion coefficients in the catalyst particle, and k = intrinsic rate constant per unit of catalyst particle volume. Therefore, from values o f f determined by the effect of particle size on activity, values of h can be calculated; as R and k are known, De can be calculated. k is obtained from the WHSV a t conversion for the 28- to 35-mesh catalyst, assuming a first-order relationship between conversion and space velocity. Values so calculated are shown in Table IV, designated as D’*. The direct method of calculating De from pore structure data and the nature of the vapor has been outlined by Wheeler (74). If it be assumed that pore sizes are much smaller than the mean free path, the so-called Knudsen diffusion coefficients can be used for a given pore size: D, = 9.7 X lo3 r ( T / M ) l / * , where r = pore radius (cm.), T = temperature, and M = molecular weight. However, for D in a given pore the Wheeler expression was used, which reduces to Dk in small pores, and to the bulk coefficient, DB, in large pores. For this purpose, DB for the gas-oil a t 900’ F. and 1 atm. was calculated by the method of Hirschfelder, Bird, and Spotz (76) to be 0.035 sq. cm. per second. From DB and Dh, individual values of D, the coefficients within the pore, were calculated and converted to D, by multiplication by the fractional porosity, and by 0.5 to approximate the effect of variation of orientation of the pores. The values of D, so obtained are also given in
Effective Diffusion Coefficients for Gas-Oil at
900’ F., 1 Atm.
(Sq. cm./sec.)
Catalyst D* De‘
I
I1
I11
IV
V
VI
0.0012 0.00071
0.0013
0.0030 0.00060
0.00105
0.0026 0.0016
0.0026 0.0011
...
...
Table IV, compared with those calculated from effectiveness factors. The comparison made in Table I V between the results of the two independent methods of calculating D, shows a fair degree of agreement. The worst discrepancy is with catalyst 111, for which pore structure data would lead one to expect the highest value of D,, whereas reaction data indicate it to be lowest; the effect of particle size decrease is more than that predicted from the pore structure. For I, V, and VI, the ratio of De’/D, might be interpreted as a “tortuosity factor’’ (4), which has been observed in these laboratories to have this order of magnitude. I t is concluded that whether D, is calculated from reaction data (73) or from pore structure data ( 7 4 , reasonable results can be obtained for high-area catalysts.
Acknowledgment The authors thank W. E. Stewart for advice and assistance in interpretation of the data, T. H. Misdom for help in determining the adsorption isotherms, and R. D. Duncan for preparation of the figures. In particular, they are indebted to J. W. Teter (deceased) for encouragement and interest in this and related work in the field of catalysis.
Literature Cited (1) Alexander, J., A P I Proceedings 27
(111), 51-6 (1947).
(2) Barrett, E. P., Joyner, L. G., Halenda, P. P., J . Am. Chem. SOC.73,
373 (1951): (3) Chem. Processing 16 (May 1953) (flow diagram). (4) Hoogschagen, J., IND. END. CHEM. 47, 906 (1955). (5) Hougen, 0. A,, Watson, K. M., “Chemical Process Principles,” ~. Wiley, New York, 1947. (6) Kistler, S. S., J . Phys. Chem. 36, 52 (1932). (7) Ries, H. E,, Advances in Catalysis 4,87149 (1952). (8) Ries, H. E,, Johnson, M. F. L., Melik, J . S., Discussions Faraday SOC.1950, pp. 303-5. (9) Ries, H. E., Van Nordstrand, R. A., Johnson, M. F. L., Bauermeister, H. 0.. J . Am. Chem. SOC.67, 1242 (i945j. Thiele, F. W., IND.END. CHEM.31, 916 (1939). Van Nordstrand, R. A., Kreger, W. E., Ries, H. E., J . Phys. Colloid Chem. 55, 621 (1951). Weisz, P. B., Prater, C. D., Advances in Catalysis 6, 143-96 (1954). Weisz, P. B., Swegler, E. W., J . Phys. Chem. 59, 823 (1955). Wheeler, A., Advances in Catalysis 3, 250-326 (1950). Wheeler, A., “Catalysis,” vol. 11, pp. 105-65, Reinhold, New York, 1955. Wilke, C. R., Lee, C. Y., IND.END. CHEM.47, 1253 (1955). RECEIVED for review May 11, 1956 ACCEPTED November 13, 1956 VOL. 49, NO. 2
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