ROLE OF CATALYTIC METALS IN HYDROCRACKING

literature Cited (1) Bird. K. B., Stewart. \V. E.: Lightfoot. E. N., !’Transport Phenomena.“ \Yiley, New York, 1960. (2) Corcoran, \V. H.. Opfell, 3 . B., Sage. B. H.. “Momentum Transfer in Fluids,‘’ Academic Press. New York. 1956. (3) Johnstone, R. E.. Thring. M. LY.: “Pilot Plants, Models, and Scale-Up Methods in Chemical Engineering.” McGraw-Hill. New York. 195’. (4) Marble, F. E.. in “Combustion and Propulsion.” 5th AG.4RDograph Colloquium, pp. 175-21 3, Pergamon Press, New York, 1963. (5) Regenscheit, B., Staub 24, 14 (1964).

(6) Soo, S. Id.>IND.ENG.CHEM.FUNDAMEKTALS 3, 75-80 (1964). ( 7 ) Sproull, \V. T., J . Air Pollution Control Assoc. 10, 307 (1960). (8) Sproull. \V. T., ,Vature 190, 976 (1961) ; 193, 464 (1962). (9) Streeter. V. L., “Handbook of Fluid Dynamics.” p. 2-14, McGraw-Hill. New York, 1961. ( I O ) Truesdell, C. .4., “The Kinematics of Vorticity,” Indiana University Press, Bloomington, Ind., 1954. (11) \Vhite, H . 3 . : “Industrial Electrostatic Precipitation,” Addison-\Yesley Publishing Co.: Reading, Mass., 1963. RECEIVED for review June 17, 1964 ACCEPTED December 16, 1964

ROLE OF CATALYTIC METALS

Downloaded by TEXAS A&M INTL UNIV on September 5, 2015 | http://pubs.acs.org Publication Date: April 1, 1965 | doi: 10.1021/i260014a010

IN HYDROCRACKING H A R O L D B E U T H E R A N D 0.A . L A R S O N Gulf Research and Development Go., Pittsburgh, Pa.

The principal role of metal on a hydrocracking catalyst is to keep the acidic sites active through the hydrogenation of coke precursors. A less important role is to fqrm olefin intermediates required in subsequent carbonium ion reactions. The rate of hydrocracking for on-stream or aged catalysts is related to metal crystallite size. The size of the metal crystallites controls the area of the support which can b e kept free from coke and made available for cracking and isomerization reactions. This property of the support to disperse metal is important in obtaining a maximum effective area of useful, active sites. When hydrocracking with base metal catalyst on a support, the reaction sequence of an adsorbed hydrocarbon varies with metallic character of the metal. The rate of rearrangement of the adsorbed species and their subsequent decomposition can be controlled by the degree of complexing of the metal with the support and the degree of sulfiding of the catalyst.

HE functions of a hydrocracking catalyst consist of hydroT g e n a t i o n , isomerization. and cracking. By controlling and optimizing these functions, improvements in hydrocracking processes and commercialization of these processes have been possible. Additional understanding of hydrocracking catalysts will aid considerably in the improvement and further commercialization of this very important operation. M‘hile hydrocracking catalysts contain both hydrogenation and acidic functions, the noble and nonnoble metals behave differently; therefore, these two divisions of catalysts are treated separately to simplify our interpretations. T h e role of metals in hydroisomerization and other catalytic reforming reactions is some\vhat similar; these reactions all require a dual-functional catalyst (6, 73, 78) with hydrogenation-dehydrogenation sites and acidic sites to effect the intended conversion. T h e interrelationships and functions of the two types of sires a r e complex and are the subject of extensive debate ; however, a simplified mechanism for hydrocracking can be based on the rather extensive literature of hydroisomerization and catalytic cracking \vith a few modifications. As in hydroisomerization (73),one of the roles of the metallic component is to convert paraffins and naphthenes to olefin intermediates. I t has been pointed out recently ( 7 ) that hydrocracking of paraffins a n d olefins proceeds through the same intermediate. Thus. naphthenes also can be expected to hydrocrack through the olefin intermediate. Recent experiments have shown (3, 77) that hydrocracking can occur even when the hydrogenation-dehydrogenation component is sepa-

rated from the acidic component by macroscopic distances T h e formation of olefins is believed to be a minor role of the metal component, however. Presumably aromatics can add a proton directly (76) to become carbonium ions; so it is not known whether the metallic component is involved in activating aromatics. Carbonium ions, once formed, can be assumed to follow the mechanism established for catalytic cracking ( 7 , 76). T h e highly reactive acidic surface then causes rearrangement and dissociation of the carbonium ion with the release of an olefin. A general rule for dual-functional catalysts is that reactions occurring on the acidic sites a r e rate-limiting. Since isomerization of carbonium ions by skeletal rearrangement has been reasoned to be the slow step (70), steps leading to scission must also be rate-limiting. This general type of behavior for dual-functional catalysis has been proposed for hydroisomerization. and it fits hydrocracking with some modifications. I n hydrocracking, the acidic sites must have stronger acidity than that apparently adequate for isomerization (75, 76). I n hydrocracking, another important difference is the competition for acidic sites. I n hydroisomerization of paraffins. although low temperatures are used. the low adsorptivity of paraffins tends to keep the catalyst surface clean. I n reforming reactions, high temperatures provide a favorable reversibility for what would otherwise be a catalyst surface covered Lvith strongly adsorbed aromatics. Hydrocracking, which combines a low temperature with strongly adsorbed aromatics. VOL. 4

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~~~~~

Table 1. Catalyst so.

(SzOnA1203

Supports) 1

2 3 4 5 6

P t , 7c 0 13 0 53 0 98

2 30 2 86 1 02

~

~

~

Preparation Procedure for Catalysts Calctnatzon Temp., OC.

Catalyst (A1203 Supports)

538 538 538 538 535 7 60

7 8 9

10 11

12

'Vz.

%

2 0 5 0 10 0 10 0 10 0 10 0

Calcinatzon Temp.,

c.

538 538 538 120 205 315


provides a physical situation that would predict a fairly complete surface coverage a t a given pressure of hydrocarbon. I t is desirable from a rate standpoint, of course, to have a high surface coverage with reacting hydrocarbons. However, the presence on the surface of strongly adsorbed hydrocarbons that react slowly or not a t all can lead to undesirable coke-forming reactions. Experimental

T h e platinum silica-alumina or nickel-alumina catalysts used in the experimental program were prepared by a standard impregnation method (77). For the platinum catalyst. a quantity of chloroplatinic acid corresponding to the correct weight of platinum o n the final calcined catalyst was added to American Cyanamid Triple A silica-alumina (257G alumina) to allow complete wetting. Nickel catalysts were prepared from nickel nitrate [Ni(N03)*.6H20] and Davison etaalumina by a similar procedure. All catalysts were dried a t 120' C. for 10 hours. Catalyst 10 was not calcined further. Table I summarizes variations in calcination temperature for the nickel catalysts studied. Several of the nickel-alumina catalysts were modified by a treatment with carbon monoxide. .4pproximately 50 cc. of the catalyst was placed in a glass tube and reduced in a stream of flowing hydrogen a t 426' C. and atmospheric pressure for 4 hours. The hydrogen rate was maintained a t 15 standard cu. feet per hour. T h e catalyst was then purged in high purity nitrogen for about 2 hours while cooling to 80' C. This is about the correct temperature for the highest rate of nickel carbonyl formation ( 9 ) ; while a t 80' C., carbon monoxide was passed over the catalyst for 24 hours. The gas stream containing nickel carbonyl was passed through a second glass reactor a t 450' C . , containing approximately 100 cc. of quartz chips of about I-mm. diameter. The nickel carbonyl \vas decomposed and formed a nickel mirror which was dissolved with 3.V HC1 and analyzed for nickel. Catalyst 9 was modified by a similar procedure, except that a 10% hydrogen sulfide-90% hydrogen (by volume) mixture \\'as used as the reducing gas for 2 hours. Pure hydrogen was used for a final 2-hour period to give a total of 4 hours for the reduction. Treatment with carbon monoxide was then carried out.

Table II.

Catalyst No.

1 2 3c 4 5

6

B E T Area, Sq. M . / G .

0.13

354 395 375 386 403 274

0.98 2.30 2.86 1.02

0 Actual rate constant = 1.46 hr.-l 515' C., 4 . 7 ~NHS/g. ~. ojcatalyst.

178

I&EC

Discussion

Nobie Metal Catalysts. T h e most commonly considered noble metal hydrocracking catalysts are platinum and palladium on silica-alumina. I n a previous study ( 7 7 ) , a number of experimental platinum catalysts were prepared and their surface properties and hydrocracking activity obtained. Table I 1 indicates that platinum crystallites on silica-alumina are between 31 a n d 466 A. in diameter for the range of platinum levels listed. If one assumes a random distribution of acidic sites and metal crystallites, their average distances apart Lan be calculated. T h e very strong acidic sites, those capable of retaining ammonia a t 515' C., are about 15 A. apart. If all the sites that retain ammonia a t 175' C . are included, the average distance between sites is about 10 A. (77). T h e density of all acid sites is 3.4 X 1020 per gram of catalyst and for very strong acidic sites is 1.3 X 1020 per gram of catalyst. T h e average distance between platinum crystallites for a 1% platinum on silicaalumina catalyst (50-A. crystals) is about 1500 A., two orders of magnitude greater than the distance between the acidic sites. These assumptions provide a physical picture of the catalyst surface where platinum crystallites are widely separated from each other and among a very high concentration of acidic sites. By using this picture with the dual-functional mechanism just described, the catalytic relationship of the metallic and acidic sites can be developed. No simple relationship exists between hydrocracking activity and number of acidic sites (77). If acidity is the most important correlating variable, acidity measurements on unaged catalysts must include many acidic sites that are active only during a brief period of operation and have no relationship to activity for a '.lined-out" catalyst. I t is postulated that the useful acidic sites are those in close proximity to metal crystal-

Surface Properties and Activities of Platinum Silica-Alumina Catalysts

Catalyst, 70 Pt 0 53

Hydrogen, carbon monoxide, nitrogen, and hydrogen sulfide were all dried by passing through Molecular Sieves prior to use. I n addition, the hydrogen was passed through a Deoxo unit, which contained a palladium catalyst to convert any oxygen to water. The hydrogen was then passed through a second drying train containing additional molecular sieves. Hydrocracking activity tests with platinum catalysts were made with a straight-run furnace oil which contained 0.047c by weight of sulfur, 18 p.p.m. of nitrogen, and 25.2% by volume of aromatics. Hydrogenation tests with nickel catalysts were made with pretreated light catalytic gas oil (205' to 350' C . range), which contained 0.00057c by weight of sulfur, 0.2 p.p.m. of nitrogen, and 36yc by volume of aromatics. Activity tests were made in a conventional, fixed-bed hydrogenation unit, approximately 1 inch in diameter a n d 10 inches in length (77). T h e conditions for hydrocracking were 316' C . , 750 p.s.i.g., and 2 LHSV with a hydrogen rate of 5000 scf. per barrel. Hydrogenation runs were made a t 260' to 316' C., 750 p.s.i.g., variable LHSV, and 10,000 scf. per barrel.

Platinum Area, .Cq. M . / G . Catalyst

0.10 0 23

0.40 0.60 0.69 0.05

Actual rate constant = 0.17 hr.-l

PROCESS D E S I G N A N D DEVELOPMENT

Au. Pt

Crystallite Size, dc, A .

Relative Rate Constant ,for Hvdrocrackine , Extrapolated Activity at zero-time 30 hours activity, hr. -l on-stream

31 54 57 89 .97 466 Acidity data f o r catalyst 3 ( 7 1 ) .

1.oa , . .

1.4 1.6 2.0 0,89

1 Ob 1.9 2.9 2.2 2.1 0,082

175' C., 72.5 cc. NHa/g. of catalyst.


lites, for only these are kept free from coke a n d remain active. O n e of the principal roles of the metal in hydrocracking is to keep the acidic sites “clean” and active through the hydrogenation of coke precursors; and this effective area is related to the geometry of the metal crystallites. T h e state of the catalyst surface away from platinum crystallites is probably similar to that which occurs on a silica-alumina catalyst during catalytic cracking a t comparable temperatures. I n these areas, the reactive hydrocarbons condense a n d form coke. With time on-stream, the area of the coked or deactivated surface will increase rapidly and approach the area of the support less a n annulus around each metal crystallite. T h e catalyst will show a rapid decline in activity or loss of catalyst “edge.” ‘ l h e equilibrium catalyst activity will then be related to the number of acidic sites the metal crystallites can maintain coke-free. \2’ith this model of the hydrocracking catalyst surface, a number of observations made in hydrocracking operations can be explained. Initial hydrocracking rates are not strongly dependent on measured metal a r e a ; rates a t later on-stream times are closely related to metal area. Reactions occurring on acic‘ic sites are the slow steps and are rate-controlling (70) ; yet the rate of hydrocracking is linear with metal area. Initial hydrocracking activity is lost rapidly even in the absence of poisons. In the presence of poisons, activity declines more rapidly, and the activity level after a given time on-stream is lower. T h e temperature response for aged catalysts is considerably less than for fresh catalysts. Data shohving the relationship of metal area to activity are given in Figure 1. Relatively little difference exists in initial (6 hours) activities of the catalysts, even though the metal areas vary from 0.05 to 0.60 sq. meter of platinum per gram of catalyst. After 30 hours on-stream dependency of activity on metal area is greater. I t is believed that the activity of all catalysts on comparable silica-aluminas of constant area would be approximately the same a t zero time. Hydrocracking catalysts are apparently similar to reforming catalysts, in that higher metal areas afford greater stability. I t has been established (78) that reactions occurring on acidic sites are rate-controlling ; in hydrocracking the cracking rate is reduced much faster than hydrogenation rate upon aging. This is expected, since the platinum has very high activity for hydrogenation-dehydrogenation reactions. Observations that the hydrocracking rate is approximately linear with the metal area (Figure 2) suggest a n important relationship between metal area and the number of hydrocracking sites. T h e correlation of high temperature hydrocracking to dehydrogenation activity rather than isomerization activity developed by Myers and Munns (74) may be related to this concept. However: these authors measured a high temperature hydrocracking activity. I t has been recently pointed out (77) that this “hydrogenolysis” occurs over metal sites exclusively. I t seems more likely that the linearity of low temperature hydrocracking with metal area that we have observed is d u e to an indirect relation of metal area and acidic sites. An activity “edge” is lost during hydrocracking, whether poisons are present or not. At any time, activity is related to the number of available and noncoked acidic sites. At zero time. all acidic sites are available for cracking. As the catalyst ages. sites most remote from the hydrogenation component soon become deactivated with coke. Although the two types of sites can function when separated by macroscopic distances, this activity is short-lived (3, 77). Two stages of this coke build-up with time are shown schematically in Figure 3.

T h e loss of temperature response of a n aged catalyst has been shown in long-time pilot plant hydrocracking operations, where fresh catalysts show a much better activity response to a change in temperature than catalysts aged for many months. An explanation for this change has been postulated based on the availability of useful sites. At an early age: all moderately strong sites can be made useful by an increase in temperature. At severely aged conditions where coke covers large areas of the catalyst, most of the moderately strong sites are assumed also to be covered with coke, leaving only a relatively few near the

60

-

*

I

I

I

I

I

16

20

24

I

I

\

o

12

HOURS

Figure 1.

’;.

30

28

ON-STREAM

Effect of metal area and time on aging rate

-

,

m 2 / p ( P L A T I N U M AREA) 0 ; o

~~

0:o

0:3

0;o

0169

4

v) +

q g 2 4r e a

1 6

0 I-

D:

v)

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a 10

20

W T OF M E T A L I A V E R A G E

IL

Figure 2.

30

X10-3

t

D I A M E T E R OF C R Y S T A L S , m I d c

Rate of hydrocracking vs. metal area P t - S i O ~ - A b O ~catalysts

I

I

I

8

Completely c o k e d w r f o c e

Partially

t

>

I-- I

coked

surface

\

-> Io a

TIME

-

Figure 3. Models of catalyst surface at early and steadystate activities VOL. 4

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1965

179


perimeter of crystallites available for activation by an increase in temperature. Base Metal Catalysts. The dependence of hydrocracking rates and selectivity on the geometry of the metallic phase is believed to be pertinent for base metal as well as noble metal catalysts. With base metals the amount and area of the metal are generally greater, but the importance of metal geometry should be the same. A more critical item is the fact that base metals react more readily with the support and with sulfur. These reactions reduce their metallic characteristics and lower hydrogenation activity. COMPLEXATIOS O F NICKEL WITH THE SUPPORT. Nickel is a typical base metal that when distended on an acidic support is an effective hydrocracking catalyst. Its role in hydrocracking has been considered to be largely that of a hydrogenation component that "protects" acidic sites. The degree with which nickel interacts with the support (2, 77) to produce nickel-alumina complexes that have low activity in hydrogenation reactions can be illustrated by a series of nickel-eta-alumina catalysts which were tested for the hydrogenation of aromatics in a pretreated furnace oil a t 315' C. and 750 p.s.i.g. Pseudo-first-order rate constants for these tests are given in Table 111. All aromatic hydrogenation activity associated with these catalysts presumably results from the metallic nickel produced during prereduction of these catalysts, for kvhen this phase of the metal is removed by CO extraction, the rate of hydrogenation is practically immeasurable. Thus, no hydrogenation activity can be attributed to the complexed phase of the nickel which remains on the catalyst after metallic nickel extraction. The complexation of nickel with the support (-7, 79) is favored by high temperature calcination (Table 111). MODIFICATION O F NICKELBY SULFUR. Sulfur also complexes with nickel to reduce its hydrogenation activity. Table I\: shows the decrease in rate constant on hydrogenating the aromatics in a FCC furnace oil as the amount of sulfur on the catalyst is increased. Relatively small quantities of sulfur on the reduced catalyst (0.02%) drxtically reduce hydrogenation rate. A relatively small amount of metallic nickel (3% of the nickel) is responsible for the activity. The role of sulfur in reducing aromatic hydrogenation can be explained by the formation of sulfides of nickel. Thermodynamic data (72) are available for predicting when bulk sulfiding of nickel occurs. Presumably, Ni& would be predicted for sulfiding under the conditions used. However, it is unlikely that surface equilibrium and bulk equilibrium are the same. I t can be postulated that stoichiometrically correct nickel sulfides have no activity for aromatic hydrogenation, and that activity which exists in the sulfided state results from regions of isolated nickel areas on the surface that have metallic properties. Any surface region that contains a high ratio of sulfur to nickel atoms, although the ratio does not correspond to Ni3S2, is not likely to display high hydrogenation activity. In the dynamic sense, bulk Ni3S2 may have apparent hydrogenation activity as sulfur atoms are continuously removed from the surface as hydrogen sulfide to form transitory metallic regions. Experiments showed that H2S can reduce complexed nickel to produce areas of metallic nickel which serve as the active hydrogenation component. Only 3% of the nickel of a 10% nickel on eta-alumina catalyst could be removed by CO after a standard hydrogen reduction. After reduction with a mixture of hydrogen and H2S followed by hydrogen alone, 19% of the nickel \vas removed Lvith carbon monoxide. 180

l & E C PROCESS D E S I G N A N D D E V E L O P M E N 1

With sulfided nickel catalysts, two other effects are noted which are different from catalysts that do not contain sulfur: A higher temperature is required to effect a given conversion; and a higher iso-normal paraffin ratio exists in the hydrocracked products. The higher temperature requirement of the sulfided catalyst is explained if one assumes that a smaller number of acidic sites are available on an equilibrium sulfided catalyst. Since hydrogenation activity is reduced when the metal component is complexed with sulfur, coke develops over a larger area of the catalyst surface, and less area is available for cracking reactions. Thus, temperature must be increased to compensate for the decreased number of acidic sites. The same difference in number of available acidic sites explains the difference in iso-normal ratio of paraffins in the product from operations using a sulfided and nonsulfided catalyst. In a previous study (7) it was pointed out that the high ratio of iso-to-normal light paraffins (C, and C,) in the product from hydrocracking was the result of a coincidental branching of the olefin broken off in the hydrocracking step. These olefins are rapidly hydrogenated, but isomerization toward equilibrium is prevented by competitive adsorption processes. These concepts fit the observed iso-normal product distributions in hydrocracking with catalysts of variable hydrogenation activity. In fact, in hydrocracking, iso-normal ratio in product paraffins was related to the hydrogenation activity of the metal and to the operating conditions which lead to a maximum number of available acidic sites on the surface. The best metal hydrogenation components produce iso-normal ratios approaching equilibrium, for enough acidic sites and isomerization activity are available. Only with less active hydrogenation catalysts were high iso-normal ratios obtained ( 4 ) . Part of the variance in iso-normal ratio with metal sulfiding also can be explained from a kinetic standpoint. Rate models which include differences in rate constants for hydrogenation of is0 and normal olefins as well as differences in adsorption coefficients of is0 and normal olefins and paraffins probably con-

Table 111.

Variation in Activity with Nickel Content and Calcination Temperature First-Order Rate Constant Calcinafor Aromatic yo Nickel tion Hydrogenation %a Nickel Catalyst (on EtaTemp., at 315' C . , Present No. Alumina) "C. Hr.-l as Metal 7 2 538 0 14 ... 8 5 538 0.20 ... 9 10 538 2.7 3.0 10 10 120 >12 ... 11 10 205 10.0 12 10 31 5 5.8 62.0 8B 9.7 538 0.0 0.0 a Ohtained by carbon monoxide extractions. ~~~

Table IV.

Change in Nickel Alumina Catalyst Activity with Sulfur Sulfur on Rate Constant Catalyst, s1.VI at 750 P.S.I.G., Ratio 288' C., Hr.-I W t . 70 0 0 2.5 0.014a 0.0022 1.3 0.0045 0 .025a 1.o 0.26 1 .44b 0.06 a Resulting from sulfur in charge stock. Resulting from suifding catalyst with 7.7 mole yo H2S-92.3 mole yc Ha at 3.5" C. ana' 15p.s 1.g.


tribute in determining the paraffin iso-normal ratio in the product from hydrocracking. If Langmuir-Hinshelwood or Rideal-Eley kinetic treatments are employed, it is likely that large differences in iso-normal light paraffin ratios can be ascribed to so-called empirical adsorption coefficients. In one study (5) of the hydroisomerization of n-pentane, it was sho\vn that the adsorption coefficient of isopentane was 1 7 times that of n-pentane on the catalytic sites over which isomerization occurs. T h a t this was not an unambiguous determination of a relative property of the iso- and normal paraffins seems particularly likely, since olefins would have been present in some concentration, and their unmeasured variance with pressure would have been reflected in apparent adsorption coefficients for paraffins. Sonetheless, a preferential adsorption for the is0 form, whether olefin or paraffin, may be likely over acidic catalysts. In a more recent study (8), it was shown that the ratio of iso- to n-pentane in the product from the hydroisomerization of 2-pentene can vary drastically with total pressure and hydrogen pressure. Complex rate models that incorporate relatively different adsorption coefficients for the iso- and normal forms presumably could predict the variable iso-normal pentane ratios that are obtained in these similar processes. T h e existence of different adsorptivity at a constant carbon number is likely, in addition to the \vel1 established competition a t variable carbon numbers in low temperature processes over acidic supports such as hydroisomerization and hydrocracking. Conclusions

In hydrocracking the metallic component of the catalyst has a profound effect on both activity and selectivity. Its geometry. its tendency to combine with the support and sulfur, and its basic activity all influence the hydrogenation function as well as the cracking of the catalyst.

Acknowledgment

T h e authors are indebted to H. H. Tobin, who worked out the technique of extracting metallic nickel from catalysts with carbon monoxide. literature Cited (1) Akimov, V. M., Slinkin, A. A . , Kretalova, L. D., Rubinshstein, A. M.: Zzuest. Akad Xauk S S S R . Otdel Khim. .Vauk 1960, p. 624. (2) Archibald, R.C., Greensfelder, B. S., Holzman, G., Rowe, D. H., Znd. Eng. Chem. 52, 745 (1960). unpublished work. (3) Beuther, H., Larson, 0. .4., (4) Beuther. H., McKinely, J. B., Flinn, R. .4.,Division of Petroleum Chemistry, ACS, Preprints, Vol. 6, No. 3, p. A-75, Chicago, 1961. (5) Carr. N. L.. Znd. Enp. Chem. 52. 391 11960). ’ (6) Ciapetta, F: G., Zbtd:, 45, 159 (i953): (7) Flinn, R . .4,, Larson, 0. A,, Beuther, H., Zbid., 52, 153 (1960). (8) Frye, C. G., Baraer, B. D., Brennan, H. M., Coley, J. R., ‘Gutberlet, L. C., TND.ENG.CHEM.PRODUCT RES. DEVELOP: 2, 40 (1963). (9) Goldberger, W. M., Othmer, D. F., IND.EKG.CHEM.PROC. RES.DEVELOP. 2, 202 (1963). (10) Keulemans, A. I. M., Voge, H. H., J . Phys. Chem. 63, 476 (1959). (11) Larson, 0. iz., MacIver, D. S., Tobin, H. H., Flinn: R. A., IND.ENG.CHEM.PROC.RES.DEVELOP. 1 , 300 (1962). (12) McKinley, J. B., “Catalysis,” Vol. VII, p. 451, Reinhold, New York. 1958. (13) Mills: G, A , , Heinemann, H., Milliken, T. H., Oblad, A. G., Znd. Eng. Chem. 45, 136 (1953). (14) Myers, C. G., Munns, G. W., Jr., Zbid., 50, 1727 (19%). (15) Pines, H., Haag, W. O., J . A m . Chem. Soc. 82, 2471 (1960). (16) Voge, H. H., “Catalysis,” Vol. VI, p. 407, Reinhold, New York. 1958. 117) it‘eis7. P. B.. Aduan. Catalvrir ,- 13. 137 (19621 (18j Weisz, P. B.: Prater, C : D., Zbid., 9, 5y5 (7557). (19) Yoshitomi, S., Morita, Y., Yamamoto, K., Bull. Japan Petrol. Znst. 5 , 27 (1963). ->

~

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RECEIVED for review May 12, 1964 ACCEPTED September 14, 1964 Division of Industrial and Engineering Chemistry, 47th Meetings

ACS, Philadelphia, Pa., April 1964.

CHARACTERIZING CRACKING CATALYSTS B Y T H E KINETICS OF CUMENE CRACKING W.

F. PANSING AND J.

B. MALLOY

Research and Development Department, American Oil Co., Whiting, Znd.

The kinetics of cumene cracking was studied in a differential reactor and the kinetic parameters-an adsorption equilibrium constant and a rate constant for decomposition of a surface complex-were used to characterize differences between various catalysts. Silica-alumina i s twice as effective as silica-magnesia for forming surface complexes, and almost seven times as effective for cracking them. Steaming and commercial use have no effect on the adsorption step for either silica-alumina or silica-magnesia catalysts. Steaming decreased the rate constant per unit area 50% for silica-alumina but increased i t 275% for silicamagnesia. Silica-alumina catalysts crack gas oils in proportion to their ability to crack cumene, but silicamagnesia cracks gas oils about eight times faster than expected from its cumene cracking rate. Cumene cracking is more complicated than generally assumed. Because of side reactions of propylene, cracking rates inferred from propylene appearance are low by a factor of 3 a t 550” F. and 1 atm., and by a factor of 1.25 even as high as 800” F.

catalysis, a catalyst and a reactant form a Thus a catalyst can be characterized by its ability to form the surface complex. and by the reactivity of the complex. We have measured these characteristics for silica-alumina and silicamagnesia cracking catalysts by studying the kinetics of cumene cracking. N HETEROGENEOUS

I surface complex ivhich decomposes into products.

Cumene cracking is frequently investigated as a reaction characteristic of cracking catalysts because it appears to involve only decomposition into benzene and propylene. Prater and Lago ( 6 ) showed that most earlier studies to establish the kinetics of the reaction were inconclusive, except for the work of Weisz and Prater ( 9 ) . T h e principal objections were the use of integral-reactor techniques insensitive to the actual VOL. 4

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ROLE OF CATALYTIC METALS IN HYDROCRACKING

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