Paradis, S.G., et al., Chem. Eng. Prog.. 67 (E), 57 (1971). Porter, F. W. B., et al., British patent 736782 (Sept 14, 1955). Ray, B. R.,et al., J. Phys. Chem., 61, 1296 (1957). Satterfield, C., “Mass Transfer in Heterogeneous Catalysis,” p 35, MIT Press, Cambridge, Mass., 1970. Schuit. G. C. A., Gates, B. C.. AIChEJ.. 19 (3), 417 (1973). Silvestri, A. J., Weisz, P. B., U S . patent 3716479 (Feb 13, 1973).
Van Driessen, R. P. et al., Oil Gas J., 82, 100 (May 18. 1964). Yen, T. F., Erdman, J. G., “Encyclopedia of X-rays and Gamma-rays,’’ G. L. Clark, Ed., p 65, Reinhold, New York, N.Y., 1963.
Receiued f o r review April 24, 1975 Accepted July 23,1975
A Kinematic Model for Catalytic Cracking in a Transfer Line Reactor John A. Paraskos, Yatlsh T. Shah,’
Joel D. McKlnney, and Norman L. Carr
Gulf Research and Development Company, Piitsburgh, Pennsylvania 15230
A kinematic mathematical model for an isothermal transfer line reactor for the fluid catalytic cracking of oil is presented. The model is based on the assumption that the catalyst particles are so small that both the catalyst and the fluid move at the same velocity within the reactor. Plug flow in the reactor is assumed. The model predicts that, for a given temperature, pressure, and catalyst, gas oil conversion and gasoline yield are functions of flowing space time only. The kinetic parameters for the model are evaluated from experimental data obtained in a pilot-scale transfer line reactor.
Introduction Isothermal kinetics of catalytic cracking of gas oil to gasoline in fixed and moving bed reactors have been investigated by Blanding (1953) and Weekman (1969), Weekman and Nace (1970), Nace et al. (19711, and Voltz et al. (1971, 1972). Voorhies (1945) examined carbon formation during catalytic cracking in fixed and fluid bed reactors. A significant recent advance in the fluid catalytic cracking field was the development of highly active, molecular sieve, zeolite cracking catalysts. With the use of these catalysts, it was soon recognized that substantial overcracking was occurring in the fluid-bed reactor of a conventional FCC process. Development efforts were then directed toward characterizing the cracking processes which took place within the transfer line. It was found that better utilization of the high-activity zeolite could be made by eliminating the fluid-bed reactor and extending the riser or transfer line to give the appropriate conversion and selectivities. The transfer line may be characterized as a co-current plug-flow tubular catalytic reactor. In most practical cases, it is operated under nonisothermal conditions. However, in the experimental work presented here, special precautions were taken to obtain near isothermal operation. When the catalyst particles are small, the residence times for the fluid and the catalyst may be assumed to be the same. Thus, under isothermal conditions, the transfer line reactor may be considered as a special case of the moving bed reactor. In this paper we first outline a model for the isothermal transfer line catalytic cracking wherein the catalyst particles and fluid flow a t the same velocity. The model is subsequently shown to correlate typical pilot-scale transfer line cracking data obtained with an equilibrium commercial zeolite catalyst. Department of Chemical/Petroleum Engineering, University of Pittsburgh, Pittsburgh, Pa. 15261.
Mathematical Model A mathematical model of the transfer line reactor can be developed from a consideration of the physical and chemical rate processes which occur within such a reactor. In order to understand the features which this fundamental mathematical model encompasses, let us briefly consider the features of the process. At the inlet of a transfer line reactor, hot freshly regenerated, fluidized FCC catalyst is brought into intimate contact with a mixture of preheated, partially vaporized gas oil and gaseous inerts. The inerts can consist of steam or nitrogen or a combination of the two. The high fluid velocity and the presence of small catalyst particles ensure a high degree of turbulence. For the purpose of model development, we have assumed that no radial temperature or concentration gradients exist in the transfer line. Furthermore, the catalyst particles are so small that they move essentially a t the same velocity as that of the fluid. In addition, the reactor is assumed to be a t constant pressure along its length for purposes of model development. Since the feed is preheated, a portion of the gas oil can be in the vapor phase. At the instant of contact between the gas oil and hot catalyst, further vaporization of the feed takes place. In the development of this model, we assume that this vaporization is instantaneous and only two phases, catalyst and the vapor, at the same temperature exist throughout the reactor space. Gasoline is produced within the reactor and can be cracked further to lighter-boiling materials along the reaction path. Carbon is deposited upon the catalyst, thus contributing to a decrease in catalyst activity along the reactor length. The following assumptions are made in the derivation of the mathematical model. (1) The catalyst particles have a uniform size in a given differential element. The flowing void fraction of fluid for all practical purposes is close to unity. (2) As shown by Weekman and co-workers, the kinetics Ind. Eng. Chem., Process Des. Dev., Vol. 15, No. 1, 1976
165
of the cracking reaction can be modeled by a three-component system, namely, gas oil, gasoline, C4-and-lighter gases plus coke. The reaction paths used in this study are as follows gas oil
ki
gasoline
al(gaso1ine)
+ where
az(C4 and lighter hydrocarbons plus coke)
(1)
a3(C4 and lighter hydrocarbons plus coke)
(2)
kz
Following the study of Weekman and Nace (1970) we assume a2 = a3 = 1. Both reactions 1 and 2 are irreversible and the first reaction is a second order with respect to gas oil weight fraction while the second reaction is a first order with respect to gasoline weight fraction. (3) We assume that the catalyst deactivation functions @ I = @2 = e-av, where g is the flowing space time. Mass Balance With the above assumptions, the steady-state mass balances for gas oil and gasoline in the dimensionless forms can be expressed as
(3) and (4)
where R1 = klOP”,q =
(%)
VRPvi
(z) G
Pvi
Here k1° and kZ0 are the kinetic parameters for reactions 1 and 2 at the flowing space time 7 = 0. a1 is the mass of gasoline produced per mass of gas oil reacted. The ratio pvlpvi can be rigorously correlated to the weight fractions of various components as v -P= Pvi
(6)
However, in the present study for simplicity we assume pv = pvi. In eq 5, VR is the volume of empty reaction, G,, Go, and GF are the mass flow rates of catalyst, oil, and the fluid, respectively, and y, Y G are the instantaneous weight fractions of gas oil and gasoline, respectively. The light gases, &-and-lighter, plus coke is given by an overall material balance on the oil feed as Yc
= Yi - Y
(7)
-YG
A solution to eq 3 can be obtained in a straightforward manner using the conditions y = 1a t p = 0. Thus
-
-
-
In limiting conditions g 0, l l y - 1 Rlv and as 7 m, l l y -1 R1/Qi.Using ( 8 ) , a solution to (4) is in principle possible. However, due to the complexity of the equation, 166
this solution will not be in a closed form. Alternatively, one can combine (3) and (4) and express Y G in terms of y. This solution is given by Weekman and Nace (1970) as
+
Ind. Eng. Chem., Process Des. Dev., Vol. 15, No. 1, 1976
Experimental Section Apparatus. The experimental apparatus used in the present study is a continuous, isothermal wall, transfer line catalytic cracking reactor. Oil feed is continuously drawn from a constant temperature reservoir and discharged to the low-temperature section of the feed oil preheater. A small pump furnishes distilled water (for dispersion steam) to the regulator stand-pipe outlet, and then to the oil-catalyst transfer line heater shell where it is vaporized. The oil flows through the low- and high-temperature preheater sections into the oil-catalyst section to meet hot regenerated catalyst and dispersion steam. Oil, steam, and catalyst enter the reactor together with bleed nitrogen flowing cocurrently through the reactor where the cracking reactions take place. The reaction products are carried into small disengaging chamber and a cyclone, where the entrained catalyst is separated from the product fluid. Hydrocarbons which have been adsorbed on the catalyst are removed in the spent catalyst stripper by steam (again vaporized distilled water). The stripped hydrocarbon and steam pass into the reactor disengaging space where they follow the same path as the reactor vapors. The vapors from the reactor cyclone pass from a quench bomb and accumulator to the fractionator system where the products are first separated into two streams, a heavy gas oil and a mixture of gases, gasoline, and water. The gas oil flows through a level control valve from the fractionator reboiler to the furnace oil tower where a furnace.oi1 cut and heavy gas oil cut are made. The gases, gasoline, and water are taken overhead into a water-cooled condenser where all the water and most of the gasoline are condensed. The condensate flows into a sight glass where as gasoline-water interface is maintained. Water and gasoline are continuously withdrawn from the bottom and top of the sight glass. The water is collected but the gasoline is passed through a water separator, an activated alumina drier, and then to a stabilizer. Light gases from the fractionator overhead condenser, with vent gases from the interface sight glass and water receiver, are passed through a calcium chloride drying tower to join the gasoline stream entering the stabilizer. Overhead gases from the stabilizer are metered, sampled, and chromatographed. Spent catalyst from the reactor flows through a standpipe, through a slide valve, and into the stripper. The catalyst level and flow into the stripper are controlled by the slide valve. An overflow in the stripper maintains the catalyst level at about the bottom of the disengaging chamber. Catalyst spills from the stripper overflow line into the catalyst hopper, where it flows, through a standpipe and slide valve into a lift unit which effectively permits control of the catalyst circulation rate. In the lift unit, the catalyst is picked up by air or nitrogen and carried up the transfer line into the regenerator. The catalyst flow can be diverted just ahead of the regenerator to obtain a circulation rate. Here gases, excess air, and steam leave the top of the regenerator and pass in series through a cyclone, filter, surge bomb, double pipe steam heater, a pressure control valve, condenser, water accumulator, and a wet test meter. Regenerated catalyst flows by gravity through a slide valve
Table I. A Summary of Isothermal Pilot Plant Reactor Data0
_-__ Wt fraction gas oil remaining at reactor outlet, Y , (lb of gas oil)/(lb of gas oil
Wt Wt fraction fraction Catalystgasoline gas oil to-fluid Flowing yield, y G , Reactor ratio, Contact space (lb of wall in feed, gasoline)/ temper(Ib of gas G,/GF, time, time, q , Run oil)/ (lb of (lb of cat)/ 7, (lb of cat.)/ (lb of gas ature, no. fluid). (lb of fluid) sec (lb of oil), sec_ _ _ feed) l /_ (Y_ - 1) oil feed) T W , “F _~ ____ _ ____ 1 0.694 4.73 0.26 1.762 0.599 0.67 0.313 949 5.09 0.25 0.772 2 0.694 1.835 0.564 0.344 950 4.30 0.24 0.80 3 0.500 2.06 0.556 0.348 951 4 0.534 4.27 0.14 1.12 0.609 0.645 0.303 950 5.24 0.60 5 0.688 4.57 0.414 1.418 0.384 949 6 0.897 6.74 0.72 5.41 0.430 1.325 0.423 947 7 0.889 6.88 1.39 10.71 0.340 1.940 0.473 948 8 0.894 6.88 1.42 10.94 0.355 1.816 0.463 949 9 0.522 3.65 0.14 0.98 1.01 0.498 0.371 996 7.16 7.62 10 0.939 58.10 0.228 3.38 0.533 1000 6.89 5.97 44.30 11 0.930 0.236 3.24 0.526 1003 12 0.911 7.93 5.55 48.30 0.225 3.44 0.523 1005 13 0.926 7.68 5.25 43.50 0.276 2.63 0.524 951 14 0.960 7.54 7.62 59.80 0.269 2.72 0.530 952 a All runs were made with a blend Mid Continent gas oil as the FCC charge and an equilibrium commercial zeolite catalyst with a reactor pressure of 30 psig. The gasoline fraction is debutanized and has a cut point of 430°F.
into the oil-catalyst transfer line where hot regenerated catalyst and dispersion steam meet with the oil to be cracked. The reactor is supplied with eight heaters, one on the disengaging space and seven on the reactor proper. Thermocouples along the reactor length provide for temperature indication and control. Thermocouples within the disengaging space provide for reactor outlet temperature measurement. Procedure and Results In order to obtain the values of key parameters such as specific reaction rate constants, activation energies and catalyst activity decay parameter for the mathematical model, a series of runs was conducted in the above described pilot plant reactor using a blend of mid continent gas oils as the FCC charge and an equilibrium zeolitic FCC cracking catalyst. A summary of the experimental results is given in Table I. The reactor was essentially an isothermal wall unit and pressures were maintained at about 30 psig. Although bulk and individual phase temperatures within the reactor could not be measured, it is assumed that the wall to bulk slurry phase heat transfer coefficient was quite large (Zenz and Othmer, 1969; Chu et al., 1953). Measurements of the temperature entering the disengaging space, for the lowest and highest contact times runs at an average wall temperature of 950°F only indicated an apparent temperature difference of f2OF. The ratio of contact times for these two runs is about 50, implying that the bulk temperatures line out rapidly in this system. Heat transfer coefficients from the catalyst particles to the fluid phase are difficult to estimate because of the complex physical and chemical phenomena occurring within the transfer line, especially near the inlet. The order of magnitude calculation based on the correlation of Heertjes and McKibbins (1956), however, indicate that the heat transfer rates under the present experimental conditions are most probably very large, particularly due to the small particle size and high surface area of zeolitic catalyst. The average catalyst particle diameter in the present experiment was 0.000197 ft.
Thus, in order to simplify the analysis of data, the reaction temperature has been chosen to be equal to the average reactor wall temperature. The pilot plant runs were grouped in accordance with the constant wall temperature. For each run a t a given level of average reactor temperature (f4OF), the dimensional group known as the “flowing space time 7” was calculated. Thus, for a given average reactor temperature of 950 f 4 O F , a profile of data points could be generated. Since the independent variable we have called the “flowing space time,” or 7, is analogous to space time or reciprocal space velocity in a fixed-bed process, it remains the key variable for any correlation of transfer line data. Most of the experimental runs made a t low contact times are subject to greater experimental error than are the high contact time runs. It will be noted that the quantity 9 is directly proportional to reactor length a t a fixed catalyst-to-oil ratio and average reactor temperature. Again, examination of the data in Table I will show that there is a range of catalyst-to-oil and catalyst-to-total fluid ratios within any group of data taken at the same average reactor temperature. However, lumping the catalyst-to-oil ratio and the contact time together by multiplication results in a readily understood relationship with real kinetic significance. When 7 is used to plot the pilot plant data related to gas oil disappearance, the performance curves shown in Figure 1 result. Figure 2 shows similar plots for the fractional gasoline yield. At low space time, significantly higher gasoline yields are observed at higher temperatures. However, with increased space time, the differences in gasoline yield a t varying temperature levels are difficult to distinguish. The fact that gasoline selectivity is independent of contact time or cat/oil ratio a t large values of flowing space time bears out the assumption of Weekman (1969) and Weekman and Nace (1970) that all cracking sites decay at the same rate. Very similar results for the fractional yields of light gases plus coke versus flowing space time were observed. The results for individual component weight fractions for those components included in the light gases plus coke fraction versus flowing space time are shown in Figures 3 to 11. These correlations are not as good as the ones shown in Ind. Eng. Chem., Process Des. Dev., Vol. 15, No. 1, 1976
167
9
iI
0 O7 .1
-
25
5
10
100
50
bL
Figure 7. Butylene yield, wt% (symbols as in Figure 3).
Kzr3 25
7
1
c.
5
IO
100
50
vl
Figure 8. Isobutane yield, wt% (symbols as in Figure 3). Y)
Figure 1. Temperature-flowing space time effects in gas oil cracking.
501
