Table 111. Effect of Oxygen Concentration
Run No.
T,, C.
29 30 31 32
432 432 377 377
6.5% C catalyst AT,,,, C. 68 8 32 3
Run Conditions Air; 400 cc./min. Air/K2= 1 / 4 ; 1000 cc./min. Air; 400 cc./min. Air/Nz = 1/4; 1000 cc./min.
Table IV. Effect of Repeat Oxidation 7% C catalyst 400 cc./min. air, 447” C.
Product Gas Concn., P.p.m. Reaction Step 1st oxidation 2nd oxidation 3rd oxidation
AT,, C. 23 4 4
COZ 800 590 600
CO
> 1000 40 22
and runs 31 us. 32 reveals that a decrease in 0 2 concentration and/or an increase in flow rate causes a drastic drop in AT,. If AT, is proportional to the Oz concentration, we would expect AT,’s of 15” and 6” for the two lower concentration runs, instead of 8”and 3” obtained, respectively. The additional lowering may be attributed to the higher flow rate used in the latter runs, permitting faster removal of the heat (this would be expected to be a small effect). Hence, it may be deduced from these runs that the reaction of the active species responsible for the AT, shows a strong dependence on oxygen concentration. The effect of repeat oxidations on the same sample is shown in Table IV. Each oxidation was carried out for 800 seconds, with a 15-minute N Zpurge between oxidations. The product concentrations were obtained 150 seconds after start of each oxidation. Again, the initial AT, is markedly lower in subsequent oxidations while reaction is still proceeding appreciably, as evidenced by the CO, produced. The 4” values obtained on subsequent oxidations may represent the true A T , effect due to the carbon oxidation reaction alone, whereas the large initial AT, must be due to a highly reactive species originally present
on the coke. The abnormally high CO value obtained initially should be noted here-the active species gives significantly more CO than normal coke oxidation. Effluent water concentrations were also high initially, dropping off rapidly with continued reaction (Massoth, 1967). I n summary, the collective results lead us to conclude that the AT, effect does not arise from either adsorption of oxygen nor seem to be due to the carbon oxidation reaction per se. That is not to say that a AT, will not arise from the carbon oxidation, but only that it will be rather small under our experimental conditions compared to AT,'^ caused by oxidation of the active species. The active species is probably formed during the coking process; it is labile (decomposed in vacuum or NP a t high temperature), reactive towards 0 2 (high adsorption), and forms CO as a major product. I t may involve a strongly adsorbed surface hydrocarbon containing a higher ratio of hydrogen than the normal coke to account for its higher reactivity, since the A T m seems to relate to the H / C ratio rather than the carbon content. Weisz and Goodwin (1966) occasionally had observed initially rapid transient burning rates which they attributed to volatile hydrocarbon formed from “thermal desorption or cracking and product desorption a t the temperature of the burning zone.” Earlier, Haldeman and Botty (1959) had also observed faster initial burning rates, but made no mention of a temperature rise associated with this higher reactivity. Literature Cited Haldeman, R. G., Botty, M. C., J . Phys. Chem. 63, 489 (1959). Haywood, D. O., Trapnell, B. M. W., “Chemisorption,” 2nd ed., p. 200, Buttenvorths, London, 1964. Massoth, F. E., IND. ENG. CHEM. PROCESS DESIGN DEVELOP. 6, 200 (1967). Weisz, P. B., Goodwin, R. B., J . Catalysis 6, 227 (1966). Zhorov, Yu. M., Panchenkov, G. M., Levinter, M. E,, Morozov, B.F., Russ. J . Phys. Chem. 40, 290 (1966).
RECEIVED for review July 10, 1968 ACCEPTED March 25, 1969
KINETICS AND DYNAMICS OF CATALYTIC CRACKING SELECTIVITY IN FIXED-BED REACTORS VERN
W.
WEEKMAN,
JR.
Mobil Research and Development Gorp., Paulsboro, N . J . 08066
IN the catalytic cracking of
petroleum a critical factor in evaluating a particular catalyst, charge stock, or processing scheme is the selectivity of the reaction for the production of gasoline. Most exploratory work takes place in fixed-bed reactors which are cheap and easy to operate on a laboratory scale; commercial units, however, are either moving- or fluid-bed reactor systems. One of the most significant characteristics of the catalytic cracking reaction is the rapid decline in the activity of the catalyst
a t reaction conditions. Thus a fixed-bed reactor represents an unsteady-state process, while a moving- or fluid-bed reactor represents a steady-state condition with respect to the reactor effluent. The products from a fixed-bed laboratory reactor are commonly allowed to accumulate over the course of the reaction and the results are then reported in terms of time-average yields. The present paper relates intrinsic selectivity behavior and the inherent catalyst decay to the time-averaged yields. VOL. 8 N O . 3 J U L Y 1 9 6 9
385
~~~~~
The effect of time-averaging the product from a fixed-bed catalytic cracking reactor has been compared to a kinetic-catalyst aging model of the conversion and gasoline selectivity. Time-averaging has a profound smoothing effect on the selectivity behavior of rapidly decaying catalysts. Selectivity behavior of catalysts in fixed beds may rank differently compared to fluid- or moving-bed reactors, depending on the rate of decay.
While a significant amount of wotk has been reported on the kinetics of catalytic cracking conversion (Andrews, 1959; Blanding, 1953; Voorhies, 1945; Weekman, 1968), little has been published on the kinetics of the selectivity behavior. I n a previous paper (Weekman and Nace, 1969) intrinsic or steady-state gasoline selectivity behavior of catalytic cracking was described for fixed-, fluid-, and moving-bed reactors. The effect of various types of catalyst decay on the selectivity of fixed-bed reactions has been treated theoretically by Froment and Bischoff (1961, 1962), and Masamune and Smith (1966). Sargara, Masamune, and Smith (1967) have treated catalyst decay in terms of time-dependent effectiveness factors for isothermal and nonisothermal cases. Carberry and Gorring (1966) have shown the consequences of progressive pore mouth poisoning and Olson (1968) has extended this work and showed the effects of an adsorbing guard chamber. Sada and Wen (1966) have analyzed the effects of pore mouth poisoning on selectivity behavior, and Chu (1968) has included a Langmuir-Hinshelwood kinetic model to study adsorptive catalyst fouling. Ozawa and Bischoff (1968) employed a linear decay law based on carbon formed in describing the kinetics of n-hexadecane cracking on silica-alumina catalyst. Recently Szepe and Levenspiel (1968) in a general study of catalyst deactivation have shown that a simple mth-order decay law describes many cases of catalyst fouling. Selectivity Model
I n studying the selectivity behavior of the unsteadystate fixed-bed catalytic cracking reactor, a previously developed model (Weekman and Nace, 1969) was employed. This model utilizes a simplified description of the product yield distribution in terms of three lumped components: the original charge material, the gasoline fraction (Cs-41OoF.), and the remaining Cd's, dry gas, and coke. This model proved capable of describing the selectivity behavior observed from catalytically cracking gas oils in moving-bed reactors. The simplified reaction scheme employed is:
c1 ko U l C 2 + U s C a c2 cs --f
The normalized bed depth and catalyst time are given by x and 0, the liquid hourly space velocity by S , and the modified reaction rate constants by KO, K,, Kz. Catalyst decay is represented by a. The characteristic B is the ratio of vapor to catalyst residence time ( t u / t,) and is essentially zero for typical fixed-bed catalyst runs. Thus if little catalyst decay takes place during the transit of the vapor, we can drop the time derivatives. More rigorously we may integrate Equations 2 along the characteristic de/dx = B Y 0, which is equivalent to writing Equations 2 as follows:
Weekman and Nace (1969) showed that a simple firstorder catalyst decay expression adequately represented a wide range of experimental data. Thus for a fixedbed reactor the catalyst decay function is given as: @ = ,-at, (4) where 01 is the catalyst decay velocity constant and t, represents the time that the catalyst has been exposed to oil. The apparent cause of the loss of catalyst activity is the adsorption of polynuclear aromatic compounds and olefins which polymerize with time to form a coke deposit. Thus this time-dependent decay function represents the summation of all the various complex reactions which proceed with time to cause a decline in catalyst activity. Weekman and Nace (1969) also showed that this decay constant is not unique and that other functional forms are also adequate. For a fixed-bed reactor the unsteadystate conversion behavior as measured a t the bed exit can be determined from Equation 3a as follows:
k2
--f
where CI represents the gas oil charged while Cz represents the C6-410"F. gasoline fraction and C3 t h e butanes, dry gas, and coke. The a, and the a2 coefficients represent the mass of Cp and C3 produced per mass of CI converted, respectively. For an isothermal, vapor phase, plug flow reactor with negligible interparticle diffusion resistance, Weekman and Nace (1969) showed that the following ,equations described the fraction of the charge stock remaining unconverted, y l , and the gasoline produced, y2. 386
I & E C PROCESS D E S I G N A N D DEVELOPMENT
c=l-y1=
A e-" 1 + Ae-Ao
(5)
where A represents the extent of reaction group, KO/ S , and A the extent of decay group at, and 0 the normalized time. If both the gas oil and the gasoline cracking activity of the catalyst decay a t the same rate-Le., 01 = 02the instantaneous gasoline yield, y 2 , may be expressed in terms of the instantaneous fraction unconverted, y ~ ; for all x by simultaneous solution of Equations 3a and 3b we obtain (Weekman and Nace, 1969):
K2
Kz
.
Kz
>* n
1.0
s = 2.0 V/(VXHR)
-I
PARA METERS(^) Kg = 2 2 . 9
MODEL
w W
5
K , = 18.1 K 2 = 1.7
*8
-I
=
d
0
42.1
In
where
Ein(u) =
s
- 5
U
*z
e" -du u
.6
2 c
U
Thus Equations 5 and 6 represent the product yields prior to any time-averaging of the product.
$
.4
$
Effect of Time Averaging
I t is common practice to allow the products from a fixed-bed laboratory reactor to accumulate with time. This both simplifies the operation of the reactor and reduces the analytical problems, since a single set of tests can now be made on the final mixed product. T o include the effect of the time-averaging step on the previously developed mathematical model, it is necessary to integrate the instantaneous results a t the reactor exit over time. Thus the time-averaged conversion and time-averaged gasoline yield are:
(74
;2
0 0
.2 .4 .6 .8 FRACTIONAL DISTANCE I N BED, X
(I) M C C O ,
900°F,
1.0
D U R A B E A D 5 CATALYST
Figure 1. Instantaneous gasoline profiles in a decaying fixed-bed reactor a t various times
'
CALC. FROM EQUATIONS MODEL PARAMETERS Kg * 2 2 . 9
K,
0
5
18.1
6,7
t, = 7.5 MINUTES
s =
2 . 0 v/(v)(HR)
The solution for time-average conversion can be obtained as follows: w
f -I It is not possible to obtain the time-averaged gasoline yield in terms of ordinary functions; however, standard numerical integration techniques were employed. Thus, substitution of Equation 6 into Equation 7b gives the desired time-averaged gasoline yields. Equation 6 is a function of 8 through substitution of Equation 5 . Equations 3a and 6 can be used to develop profiles of gasoline yield using reaction rate constants determined from a Mid-Continent gas oil and a zeolitic catalyst (Weekman and Nace, 1969). The results of this computation are shown in Figure 1; when the catalyst is freshi.e., 8 = 0-gasoline yield increases and then decreases as more gasoline cracking takes place. As time proceeds, however, the catalyst loses its activity and the point of maximum gasoline yield shifts further and further down the bed. Thus a t the exit of the bed the gasoline yield first increases a n d , then decreases as time goes on. This effect is shown in Figure 2 along with the resultant time-averaged gasoline yield. Both the time-averaged and the instantaneous gasoline yields show a maximum with time; however, the maximum timeaveraged yield is lower because of the smoothing effect of the time-averaging process. This smoothing of the yields leads to the pronounced effect of space velocity and catalyst residence time on the gasoline selectivity behavior. If there is no decay in catalyst activity, the timeaveraged and instantaneous yields become identical. As the averaging time goes to zero, the time-average and instantaneous yields also become identical. Thus the rate
8U
.4
u)
z 0 c
y
cz U.
z
.2
I l l
INSTANTANEOUS GASOLINE Y,
0
I
-', I \
-.1
.2 .4 .e .8 FRACTIONAL TIME OF RUN, e
1.0
Figure 2. Instantaneous and time-averaged gasoline yield in fixed-bed cracking
of catalyst decay and the length of the averaging time affect the degree of smoothing and the resultant difference between instantaneous and time-averaged yields. Using the previously developed rate constants (Weekman and Nace, 19691, it is possible to show the parametric effect of the important process variables on the gasoline selectivity behavior of^ time-averaged fixedbed data. From Figure 3 we see that as the run time of the fixed-bed experiment is increased, we get an increasing departure from the instantaneous yield behavior. Thus a t long catalyst residence times, the timeaveraged selectivity behavior is substantially poorer than would be observed in a steady-state reacting system such as a fluid- or moving-bed reactor. VOL. 8 N O . 3 JULY 1969
387
CALC. FROM EQUATIONS
z-I
KO
8u
=
.8
K2
v
. i c
z
I
0
cy)
1.7
I
a LL
I
I
42.7
2 c 0
f
d
.6
f
n
.4
W
W
u
.4
0
6
U
a W >
(z
W
u
>
u
.2
W
.2
W
I
f
I-
I-
-
-
o 0
.2
E ,
.4
.6
.e
T I M E AVERAGED CONVERSION
CALC. FROM EQUATIONS
.2
0
1.0
Figure 3. Effect of catalyst residence time on timeaveraged gasoline selectivity
.4
.6
6,7, 8
.8
r , T I M E AVERAGED WT. FRACTION CONVERTED
1.0
Figure 4. Effect of space velocity on time-averaged gasoline selectivity Figure 4 reveals the effect of space velocity on the time-averaged selectivity behavior. As space velocity is decreased, a t the same conversion level, time-averaged gasoline yield becomes increasingly poorer compared to the instantaneous yield behavior. As space velocity is increased, the catalyst residence time (fixed-bed run time) decreases in order to achieve a given conversion, which in turn reduces the smoothing effect of time-averaging. Another parameter commonly employed in catalytic cracking evaluations is the catalyst-oil ratio, which is defined as the volume of catalyst divided by the total volume of oil pumped over the catalyst during the run. The catalyst-oil ratio ( p ) is related to the total catalyst residence time (tJ and the space velocity (s) by
388
=
LL
n
I.I
I .7
K2 U,
z
2
.,1I-
.8
v
= 42.7
6,7, 8
KO = 22.9 K , = 18.1
U
u
INSTANTANEOUS YIELD
.6
CALC. FROM EQUATIONS
1.0
m
9
2
-zJ
22.9 18.1
K, =
c
W
MODEL PARAMETERS
0 +m
6,7, 8
1.0
I & E C PROCESS DESIGN AND DeVELOPMENT
Figure 5. Effect of catalyst-oil ratio on time-averaged gasoline selectivity For a given temperature, any two variables of the three possible (space velocity, catalyst-oil ratio, and catalyst residence time) are sufficient to define the extent of cracking. Figure 5 shows that when catalyst-oil ratio is employed as a parameter on the selectivity plot, an interesting two-valued or “looping” behavior is observed. For a given catalyst-oil ratio, it is possible to observe two different time-averaged gasoline yields a t a fixed conversion level. From Equation 9 we see that the same catalyst-oil ratio can be achieved by a short catalyst residence time and a high space velocity or by a low space velocity and a long catalyst residence time. On Figure 5 the upper part of the selectivity envelope is achieved at short catalyst residence times and high space velocities and thus little smoothing from the time averaging effects is possible and the time-averaged yield almost equals the instantaneous yield. However, on the bottom side of the envelope, the catalyst-oil ratio is achieved by low space velocities and longer catalyst residence times, which allows considerable time-averaged smoothing. The time-averaged selectivity behavior is always less than or equal to that obtained from a steady-state reactor such as a moving- or fluid-bed reactor (Figures 3, 4, and 5). Since Equation 6 shows that the instantaneous selectivity behavior is independent of catalyst-oil ratio, space velocity, or catalyst residence time, it is clear that the time-averaging process by itself introduces substantial structure into the experimental data. If there were no gasoline cracking-i.e., K Z = 0-all the parametric structure shown on Figures 3, 4, and 5 would collapse to a single straight line. This can be demonstrated by setting K z = 0 in Equation 3b and dividing it by Equation 3a. The resultant expression is:
With the boundary condition a t solution is:
t
= 0 of yz = 0, the
For the effect on the time-averaging process we can substitute Equation 11 and 7b to give:
0.9
'KO .8
Thus without gasoline cracking, the instantaneous and time-averaged yield structure would be identical and the selectivity behavior could be described by the straight line with slope Kl/Ko. The instantaneous selectivity behavior is independent of the rate of catalyst decay, 01 (Equation 6). For timeaveraged selectivity behavior, however, the rate of catalyst decay becomes important. The effect of catalyst decay is demonstrated on Figure 6 with the time-averaged selectivity behavior approaching that of the instantaneous yield as the catalyst decay velocity, C Y , approaches zero.
.2 80%-the cracking of gasoline greatly reduces the selectivity. Figure 8 shows experimental data drawn from the same experimental series and plotted with catalyst-oil ratio (p) as a parameter. Again the broken lines represent the prediction of the selectivity model and the solid line represents the instantaneous or steady-state selectivity envelope. We see the interesting two-valued behavior of the constant catalyst-oil curves as predicted by the mathematical model and confirmed by the experimental data. If there were no gasoline cracking-Le., K2 = 0-all of
Figure 6. Effect of decay velocity constant, time-averaged selectivity behavior
cy,
on
Catalyst residence time, t, = 20 minutes
n 2
wh
1.0
EXPERIMENTAL
164.8
CAT. R E S . TIME
d. =
18.2
0
.8 v)
s
-
0 0 -
A-
n .6
I
210.7
I
KI
z
W
I KO
W
5 IO 20 40
MIN
MIN MIN MIN
.4
F z
0
F u U
.2
e
0
1 2
0 .2 .4 .6 6 , WT. FRACTION T I M E AVERAGED
.8 1.0 CONVERSION
Figure 7. Comparison of model and experimental data for constant residence time Catalyst. 30- to 60-mesh pure zealite Stock. Mid-Continent gas oil Temperature. 900" F.
_.-. Calculated from
model
these data would collapse to a single straight line. Thus substantial cracking of the gasoline is clearly taking place. Figure 9 presents experimental data from cracking a Mid-Continent gas oil over a commercial zeolite fluid cracking catalyst a t 900" and 1000°F. in the contained fluidized-bed reactor. The dashed lines represent a prediction of the selectivity model using constants derived from the data by the nonlinear estimation method. As before, the solid lines represent the steady-state selectivity behavior predicted by the model. As shown in a previous paper (Weekman and Nace, 1969), increasing reaction temperature decreases the selectivity of the reaction for gasoline production, and the VOL. 8 NO. 3 JULY 1 9 6 9
389
4
w>
1.0 KI
W
z
1
s:
I!
K O = 210.8
=
164.8
EXPERIMENTAL C A T / 01L RATIO
.e
4
u
. 1875
0-
.375
I
CATALYST A CATALYST B
0.9 0.7
LL
20 IO
0.1 0. I
I l l !
- 1.50
n
Y a
4-
K I / K O- K 2/K0
.6
1
1
INSTANTANEOUS Y I E L D
W
2 W
I i=
*4
z
0 Iv 4 d
.2 Y I E L D , t c = 2 0 MIN
IL
c: 3-
0
Ix"
-€ ,
0
.2 .4 .6 .e I .o WT. FRACTION TIME AVERAGED CONVERSION
Figure 8. Comparison of model and experimental data for constant catalyst-oil ratio
0 f
.2 OR
7,WT.
.4 .6 .8 FRACTION CONVERSION
1.0
Figure 10. Effect of time-averaging on catalyst selectivity evaluation
Catalyst. 30-to 60-mesh pure zeolite Stock. Mid-Continent gas ail Temperature. 900" F.
-...Calculated n 1
w-
from model
.7
> K O = 20.2 K I = 18.6
W
f 1
4
