554
Ind. Eng. Chem. Prod. Res.
Dev. 1986, 25, 554-562
Pilot Plant for the Fluid Catalytic Cracking Process: Determination of the Kinetic Parameters ob Deactivation of the Catalyst Jos6 Corella,' Ana FernBndez, and Jos6 M. Vldal Department of Chemical Engineering, University of Zaragoza, 50 009 Zaragoza, Spain
A detailed description is given of a fluid catalytic cracking (FCC) pilot plant with a vertical pneumatic transport reactor that operates with contact times for the catalyst ranging from 0.5 to 6 s. It can operate with or without continuous catalyst regeneration, with fresh catalysts or catalyst in equilibrium, and with different feedstocks. Special attention is devoted to the estimation in such pilot plant of the kinetic parameters of cracking and of deactivation of the catalyst for a given feed-catalyst system. These parameters are necessary in order to elaborate control models and optimize commercial FCC units.
Introduction The fluid catalytic cracking (FCC) process can be applied to many different types of feedstocks. There are also many types of catalyst on the market that can be used in the process. In order to determine the most suitable catalyst for a given feed and predict the conversion and product distribution that will be obtained in a commercial unit, two types of experiments are usually carried out: microactivity tests (MAT) and pilot-plant experiments. MAT determinations are carried out on a very small scale with a small fixed-bed reactor or, sometimes, with a fluidized bed, discontinuous for the catalyst. The time on stream of the catalyst ranges from 50 to 100 s, which is a much longer time than that existing in a commercial unit (10 s approximately). For this reason, the MAT results are merely for guidance purposes, and we believe that they are sometimes far from the reality of the commercial unit. On the other hand, the very small amount of gas oil used in the MAT does not allow the determination of the product distribution obtained in the experiment. To avoid all these problems it is highly advisable to use an FCC pilot plant that reproduces the experimental conditions of the commercial units. In the pilot plant, the reactor, its continuous operation, the catalyst/oil (C/O) ratio, and the gas oil and catalyst contact times can be very similar to those found in the refinery. In addition, with a pilot plant operating in a continuous fashion, the yields of liquid and gaseous products are high enough to carry out complete analysis of the products obtained from cracking. Therefore, experimentation in a pilot plant is highly advisable in order to predict the results of cracking using a new catalyst and/or feed in a commercial unit. However, an FCC pilot plant has at least two drawbacks. First, commercialized pilot plants are quite expensive. Also, if one wishes to design and build a pilot plant, there is in our opinion little information in the available literature describing the details in order to build it reliably. Table I shows a summary of the information that we have found in the available literature concerning characteristics of the FCC pilot plants. Among these references only Paraskos et al. (1976) and Shah et al. (1977) show the evaluation of the kinetic parameters (with the pilot plant of the Gulf Co.), but they do not show any scheme of the pilot plant. The works by Pine et al. (1984) and Cimbalo et al. (1972) are also interesting, although they do not describe the ARC0 type circulating pilot plant they use. The former work studies the selectivity in the catalytic 0196-4321/86/ 1225-0554$01.50/0
cracking toward the different families of hydrocarbons, while the latter carries out studies on deactivation by deposition of metals on FCC catalysts. In view of the lack of information in the open literature, it was decided to investigate the design, construction, and operation of an FCC pilot plant. In this work, a detailed description will be given of an FCC pilot plant that operates correctly. Some results will be shown, and special attention will be devoted to the determination of the kinetic parameters of deactivation of a commercial catalyst. Before the design and construction of the FCC pilot plant were started, some experiments were carried out in a small installation with a fluidized-bed reactor which was discontinuous for the solid and was fed with oil pulses of 4 s (modified MAT). In this experimentation (Corella et al., l985,1986a,b), the importance of the first seconds on stream of the catalyst became evident. On the other hand, previous fluid-dynamic experiments in cold flow models and cracking experiments in a small installation with a continuous fluidized bed were also performed. Since the experience acquired in these two installations was found to be very important for the correct design and operation of the pilot plant, these two installations are shown in Appendix I.
Description of an FCC Pilot Plant with a Riser Reactor After the previous experiments indicated above and in Appendix I, the small pilot plant with a riser reador shown in Figure 1was built. This plant is in good working order at the present time. In this installation the cracking reactor is a 304-L stainless steel tube of 2.75-m length and 10-mm i.d. The catalyst and the previously vaporized gas oil are continuously fed at the bottom of this reactor. This installation can work with or without continuous regeneration of the deactivated catalyst and allows working with catalysts that are initially either fresh or in equilibrium. Catalysts obtained from commercial units are initially in equilibrium. In this case, the catalyst is fed from a hopper using a screw feeder of 0.035-m diameter, powered by a 368-W (0.50-hp) engine. The turning speed of the screw is adjusted by means of two power reducers. This enables us to obtain catalyst flow rates between 0.16 X and 4.2 X kg/s. The catalyst is fed with a prefixed flow rate to the preheater, which is a fluidized bed of 0.080-m i.d. and 0.20-m bed height. The fluidizing gas is N2 previously 0 1986 American Chemical Society
Ind. Eng. Chem. Prod. Res. Dev., Vol. 25, No. 4, 1986 555
Table I. Some Previously Published Data on FCC Pilot Plants
ARC0 Petroleum Products Co." circulating catalyst mass flow rate (C), kg/s oil mass flow rate (O), kg/s C/O ratio reactor type catalyst mean residence time, s riser length
1.6 X 104-2.5 X 1.6 X 10-"1.6 X 2-30 fluidized bedd riser (Utube$ 30-360d 12.1e
Amoco Oil C O . ~ 6.3 X 10-3-2.5 X 6.3 X 104-2.6 X 3-16 fluidized bed
lo4
6-240
Gulf Research and Development C O . ~ pilot plant of this work 1.6 X 10-4-4.2 X 1.2 X 104-2.9 X 6.8-9 2-10 riser riser 0.14-7.6 7.3
0.25-6 2.75
"Wachtel et al. (1972); Humes (1983); Wagner et al. (1984). bHerring et al. (196); Wollaston et al. (1975). CParaskoset al. (1976); Shah et al. (1977). dWachtel et al. eHumes; Wagner et al. TG8
by several independent resistances of 500 W each, with automatic temperature control, which supply in each zone the heat required to maintain the temperature difference between the lower and upper part of the riser within &5
K.
Products
011
Figure 1. Pilot plant with riser reactor.
heated in a coil at the bottom of the bed. The catalyst preheater is surrounded by a 3800-W oven with an automatic temperature control which heats the catalyst up 973 K (700 "C). The catalyst leaves the preheater for the reactor through a thermally isolated standpipe of 10-mm id., 0.63-m length, and 60" slant with respect to the horizontal. In this standpipe there is a valve and a N2aeration. At the bottom of the riser the vaporized gas oil disperses and transports the catalyst. Alternatively, a dispersing stream of gas oil-steam or gas oil-N2 mixtures can also be used. Cracking in the riser is carried out isothermally in order to facilitate the subsequent data analysis. The temperature is measured by means of six chromel-alumel thermocouples located at different riser heights. Since cracking is an endothermic process, it is necessary to heat the reactor in order to maintain isothermal operation in the riser. The heating must be carried out with variable intensity along the reactor. For this reason, the riser is surrounded
The products and the deactivated catalyst are separated in two cyclones at the exit of the cracking reactor. The gaseous products are then cooled in two tubular heat exchangers and in a cold trap with ice, as shown in Figure 1,and collected for their subsequent analysis. The deactivated catalyst passes to the stripper, which is a fluidized bed of 0.040-m i.d. and 0.10-m bed height. The catalyst is stripped with N2 or steam to remove the hydrocarbons retained or adsorbed on the catalyst. The stripper is surrounded by two ovens of 2000 W each and operates at about 723 K (450 "C). On leaving the stripper, the catalyst can pass either to a hopper, where it is kept for later analysis, or to the regenerator. When fresh catalyst is used, the catalyst passes to the lower fluidized bed (Figure 1) where it is always regenerated with air at 953 K (680 "C). After several cycles in the plant, the catalyst arrives at a state of equilibrium. When a catalyst that is initially in equilibrium is used, it passes from the stripper to the final collection hopper. In this case, the lower fluidized bed (Figure 1)acts simply as a catalyst preheater. The stripper and the regenerator are connected by a vertical standpipe of 12-mm i.d. There is a slide valve and an aeration with N2 at the bottom of the standpipe. The aeration in the two standpipes facilitates the pass of the descending packed bed and also prevents the contact between the air of the regeneratorlpreheater and the hydrocarbons in the stripper or in the riser. The installation is provided with thermocouples in the gas oil preheater, at six points in the riser, in the stripper, and in the regenerator/preheater. There are also several pressure probes in the installation: at the gas inlet and outlet of the regenerator/preheater, in both standpipes, at the bottom of the riser, and in the stripper (Figure 1). The standpipe from the regenerator/preheater to the reactor is the most delicate part of the installation. In the first experiments two important irregularities in its operation were detected: (1)the existence of obstructions which often made it necessary to stop the experiment; (2) the passage of part of the gas oil from the bottom of the riser to the regenerator/preheater. In order to solve these problems it was necessary to study the circulation of the catalyst-gas mixtures in isolated standpipes (Cdmara, 1984). We concluded that in order to obtain correct circulation it was necessary (1) to place a valve in the standpipe which allows the formation of a descending packed-fluidized bed of catalysts, (2) to provide an aeration (with N2in our case) which favors the descending circulation of the catalyst packed-fluidized bed, and (3) to
556
Ind. Eng. Chem. Prod. Res. Dev., Vol. 25, No. 4, 1986
control at all times the pressure difference (AP)between the two ends of the standpipe to avoid emptying the standpipe. This pressure difference should be of about 0.020 m of water. Since there is a gas volume increase in the riser due to cracking, the space time (7)is, in principle, different from the average residence time of the gas (EJ. However, given the fact that the gas conversion obtained here is going to be small, and to avoid complications, both times will be considered identical in our case (7 = fJ. This is equivalent to saying that the factor of expansion of the gas is equal to zero. Other authors such as Paraskos et al. (1976) or Lewis and Wrende (1984) have also used this assumption in their data analysis in installations similar to the one used in this work. In this respect, Blanding (1953), Weekman (1968), and Paraskos et al. (1976) have already indicated that the use of second-order kinetics for the cracking reaction of gas oil to gasoline accounts satisfactorily for the effects from the real variation of volume, thus avoiding the use of the factor of expansion which would complicate the analysis of some data. The time on stream of the catalyst (t,) in the riser is equal to the contact time between the catalyst and the reacting gas. Assuming piston flow for the catalyst, the time on stream will be the same for all the particles and will be given by
fa=
w/c
(1)
where W is the catalyst weight in the riser and C the catalyst mass feed rate. In this work we shall use the following approximation:
fa = fg
(=7)
(2)
Although the first equality of this equation has often been the subject of controversy, it has been verified experimentally in an FCC circulating cold model by Santos and Dantas (1983) and has also been used by Paraskos et al. (1976), Shah et al. (1977), and Lewis and Wrende (1984).
Operation Intervals in the Pilot Plant Our installation operates satisfactorily in a certain range of conditions determined by the following: Maximum Gas Oil Flow Rate Given by the Pump. The maximum gas oil flow rate given by the pump determines the maximum velocity of the gas oil in the riser and, therefore, the shortest space time that can be obtained. In this work two pumps have been used: a peristaltic pump of 70 W, which provides flow rates of up to 1.1 X lo4 m3/s, and a diaphragm pump of 250 W, which was able to feed 2.8 X lo* m3/s. The latter was the one used in the experiments carried out with space times of 0.5 s. Length and Diameter of the Riser. For a given gas oil flow rate, the diameter of the riser determines the superficial velocity of the gas, and its length determines the space time (7). The riser in our plant had a length of 2.75 m and an internal diameter of 0.010 m (0.78 x m2 transversal area). Since the maximum gas oil flow rate given by the pump is 8 X lo4 m3/s (gas, 1 bar, 773 K), the maximum superficial velocity of the gas that can be reached at the riser inlet is (8 X 1Oe4/O.78X lop4)= 10.25 m/s. Therefore, the minimum space time for the gan oil is 2.75 m/10.25 m s-l = 0.27 s. Maximum Catalyst Mass Flow Rate Given by the Screw Feeder. For a given gas oil mass flow rate, the maximum mass flow rate of catalyst supplied by the screw feeder determines the largest C/ 0 ratio that can be reached. The 370-W (0.50-hp) feeder installed in our pilot
0
1
2
3
4
5
-6 ( 5 ) Figure 2. Limits of operation of our pilot plant with riser reactor.
plant gave a maximum catalyst mass flow rate of 4.2 X k/s. With these operational limits as well as the choking velocity (Appendix I) in mind, the possible operating conditions of our pilot plant are those shown in Figure 2. It can be observed that the operating range in this pilot plant is not too wide. In order to obtain longer space times it would be necessary to increase the reactor length. However, according to the conversion and yield shown below, this pilot plant adequately meets the usual requirements and the different types of problems to be solved in an FCC pilot plant.
Description of an Experiment Each experiment starts with the introduction of Nz in the three main apparatus of the installation (reactor, preheater, and stripper). At the same time, the installation is progressively heated, avoiding the formation of temperature profiles along the riser. Once the desired working temperature in every zone of the installation has been reached, the catalyst is fed and circulated throughout the plant until a stationary state is achieved. At this moment the gas oil is introduced and the Nz flow to the riser is stopped. If the fed catalyst is not in equilibrium, i.e., if it is new or fresh, the stripper exit is connected with the regenerator and air is fed to the regenerator, previously heated to 680 "C. In each run, the catalyst and gas oil mass flow rates are selected in order to achieve the desired operating conditions (C/Oand 7). A short period of time is allowed to stabilize the temperature and the gas oil and catalyst flows. After this stabilization period, samples are taken of the gas and liquid products and of the deactivated catalyst. The time required for collecting such samples is about 10 min. Once sampling has finished, about 20 min are allowed before sampling again. Due to the limitations imposed by our timetable, 6-10 samples were collected in each experiment, the total duration of the experiment being 6-8 h. The gas and liquid products were analyzed by gas chromatography and mass spectroscopy, and the deactivated catalyst was analyzed by thermogravimetry, according to the procedures and experimental conditions detailed in Appendix 11. These analyses provided the following information: yield, in weight percent of the feed of the gases (Hz, H2S, CHI, ..., n-butane), light naphtha (C5,-466 K), heavy naphtha (466-494 K), and LCCO + HCCO (494-633 K); percentage of aromatics, naphthenes, and paraffins in the liquid products (light and heavy naphtha); gasoline content of the liquid sample (RG), calculated as the gasoline weight fraction in the liquid sample and taking as gasoline the cut between C5and 483 K (this value ( R G ) is obtained directly from the gas chro-
Ind. Eng. Chem. Prod. Res. Dev., Vol. 25, No. 4, 1986 557
Table 11. Characterization of the Gas Oil Used gravity ('API), 298 K density (283 K), kg/m3 density (773 K), kg/m3 viscosity (773 K), CP molecular weight four point, K aniline point, K Conradson carbon residue, wt % (10% residue) hydrogen, wt % sulfur, wt % nitrogen, ppm Ransbotton carbon, wt 90 (1090residue) distillation curve IP % 50 % 95 % EP % K (UOP) freeze point, K flash point, K color cetane index viscosity (313 K), CS corrosion, 3 h at 373 K metals Ni, ppm v , PPm analysis type, wt % paraffins naphthenes aromatics
37.9 860.5 3.4 0.0087 221 258 335 0.01 13.36 1.32 86 0.05
0
m0
L
5
6
7
8
9
1
0
02
0
1
MZ-7P
OIL G 0 V
05/
04
08
06
10
MZ-7P
L
O;L G O V
500 'C
500 'C
!& i%)
, 08
06 04
1B
02
3 0
-10
-10
0 1
2
3
L
5
c/o
41.9 23.9 34.2
01
and yield to gasolines (4c), defined as RG(wt of sample) =
3
0.6
m0
d'G
2
Figure 3. Variation of the gas oil conversion (a) and of the yield to T = 2 gasolines (b) a t the riser exit: (+) T = 0.5 s; ( 0 )T = 1 s; (0) S; ( X ) T = 3 s; (A)T = 4 s.
512 K 556 K 583 K 590 K 11.5 258 389 -0.5 49 3.30
matography results); gas oil conversion at the reactor exit (X,), defined as mo - (1 - R G ) ( wof ~ sample) = (3)
x,
1
(4)
These last two definitions are necessary for the kinetic analysis of the data since according to Abbot and Wojciechowski (1985) depending on how X,and are determined, the kinetic parameters calculaJed from them may differ considerably. Once the conversion and the yields to gasolines and to the different gases have been obtained for each sample, the values obtained for all the samples were examined again, discarding those with deviations higher than 10% from the average. In steady-state operation, the results obtained at the riser exit should not depend on the operation time of the plant but only, for a given feed-catalyst system, on C/O, T , and T. Typical Results. Most experiments were carried out at 773 K by using a gas oil with the characteristics shown in Table 11. The results reported in this work were obtained with a commercial catalyst, MZ-7P of Akzo-Chemie, in equilibrium. The main properties of this catalyst are catalyst type, zeolite in a Si02-A1203 matrix; D , 66.1 pm; pp, 1.8 X kg/m3; A1203, 33 wt %; pore vorume, 0.25 cm3/g; and surface area, 140.0 m2/g. The experiments were carried out with space times for the gas oil of 0.5, 1, 2, 3, and 4 s, C/O ratios varying between 1.3 and 8.6, and temperatures varying between 753 and 793 K. As an example of the results obtained, the variations at 773 K of X , with the C/O ratio and of 4G with X,are shown in Figure 3. These representations and results are typical for this FCC process. In figure 3b all the experimental points have been fitted to the same curve (con-
6
7
8
9
10
0
02
04
06
08
XJ-I
1-1 bl
Figure 4. Coke concentration on the catalyst (a) and yield to coke T = 2 s; ( X ) T = (b) at the riser exit: (+) T = 0.5 s; ( 0 )T = 1 s; (0) 3 s; (A)T = 4 5.
tinuous line). A better fitting could be obtained by using the selective deactivation model developed by Corella et al. (1986b). However, it is out of the scope of the present work to go deeply into this subject and, therefore, only a reference is made to such a possibility. The coke concentration on the catalyst at the riser exit vs. the C/O ratio and the yield to coke vs. X , are shown in Figure 4. These results are also expected for this FCC process, and they are similar to those reported, for instance, by Voorhies (1947) or Masologites and Beckberger (1973).
Calculation of the Parameters of Catalyst Deactivation Special attention has been devoted in this work to the precise determination of the parameters of catalyst deactivation. For this purpose, use has been made of the equation (5) In this equation there are two parameters, s/ and d , and a' represents the catalyst activity, averaged for all the reactions and reactants in the cracking process and for all the acidities or strengths of the active sites on the catalyst surface. It is a simplified average activity which corresponds to model V of Corella et al. (1985). According to eq 5, for a given feedstmk-catalyst system, a' depends only on the temperature and on the time on stream, t,, of the catalyst. Equation 5 needs to be used in an integrated form, but its integration demands previogs knowledge of the "observable" deactivation order, d. For this reason and in order to calculate this deactivation order, use has also been made of the equation
5 = At,-@ (6) which, as is well known (see, for example, Corella et al. (1985)), i s an integrated form of eq 5, for d # 1, being
558
Ind. Eng. Chem. Prod. Res. Dev., Vol. 25, No. 4, 1986
a- = -P + l
i
(7)
P
and
02 O3 1/ca-1,
MZ-7P
OIL G.0.V 500'C
A=[&] The network describing the main process is Weekman's (1968) lumped model 0.04
k3
L
kG
/P
k
p
t
1
u
)
/P
05 0.7
1
2
3
6
5
1
9
7; (SI
described by the kinetic equations
Figure 5. Calculation of the observable order of deactivation (2) with eq 16.
where the reaction of gas-oil cracking is of second order while the cracking of gasoline to gases is of first order. Krishnaswamy and Kittrel(1978) have also used eq 10 and 11 for the modeling of catalytic cracking. In a riser reactor the gas and the catalyst circulate in an ascending direction at high velocities, with a flow pattern approaching the plug flow. With this consideration, the mass balance for the gas oil in a differential mass element of solid, d W, becomes
0 d X = [ i o (1- X)'Z]dW
(13)
where eq 10 has been taken into account. If, as in our case, the riser is isothermal, the simplified deactivation function (5/ or A ) is constant throughout the reactor and a' depends only on the time on stream of the catalyst. As is shown in detail in Appendix 111,substituting eq 6 into eq 13 and integrating with W = 0 and X = 0 give the following expression:
used for data analysis and the final and integrated kinetic equations are also very different. For instance, Paraskos et al. (1976) obtained another set of equations because they started from the assumption that the value of d is equal to 1. Thus, theif equations can easily be obtained from eq 13 and 5 with d = 1. However, Corella et al. (1985) have already shown that if the conversion values given by Paraskos et al. (1976) are fitted without a prejudgment of the deactivation order, very different results are obtained. Thus, from the data given by Paraskos et al. (1976) and eq 14, Corella et al. (1985) found d = 3 for t , C 3 s. The observable deactivation order mentioned corresponds to eq 5 or to model V of Corella et al. (1985) and is different from the intrinsic deactivation order for each active site, (a ), when the real catalyst surface with active sites of di?ferent strength is considered or when model I11 of Corella et al. (1985) is used. So, Corella and Mengndez (1986) have shown that for the same FCC process and for the same catalyst as the one used here, the intrinsic deactivation order is equal to one, although the observable order is three. Their findings are not in disagreement with the fact that d = 3. Calculation of 5, and k o an-d Confirmation of = 3. The integration of eq 5 for d = 3 yields
a
On the other hand, from eq 1 and 2
a' = 1/[1
+ 25/W/C]1/'
(17)
Substituting this equation into eq 13 and integrating for W = 0 and X = 0, as shown in Appendix 111, give This equation has also been used by Paraskos et al. (1976), Lewis and Wrende (1984),and Shah et al. (1977),who have also studied the FCC process in a pilot plant with a riser reactor. This equation allows eq 14 to be written as
Therefore, representing the left-hand side of this equation vs. (log T ) , one obtains a straight line whose slope yields the value=of P or, from eq 7, the observable order of deactivation, d. This representation is shown in Figure 5 with the data obtained in our riser. The slope of the straight line of Figure 5 is I-@)= 4.457. Therefore, from eq 7, d N 3. This value of d confirms that found by Corella et a]. (1985) for the same FCC process and catalyst in an experimental device very different from the one used here. We believe that it is important to know the value of d since depenting on this value (1) the subsequent determination of $ can yield very different results, since both parameters are closely related by eq 5, and (2) the method
or, in combination with eq 15
Representing the left-hand side of this equation vs. [ (1
- X , ) / X , ]rC/O,the experimental confirmation of this
equation is obtained, as shown in Figure 6. Therefore, for these low values of time on stream, eq 17 is verified, i.e., d = 3. From the slope and the intercept of the straight line in Figure 6, the values of the parameters & and ko can be obtained. For the feedstock and the MZ-7P catalyst in equilibrium used in this work, at 773 K these values are 4 = 4.47 s - ~ ko = 0.23 s-' To verify the accuracy of these results, the ioand & values obtained have been substituted in eq 19 for all the
Ind. Eng. Chem. Prod. Res. Dev., Vol. 25, No. 4, 1986 559 Table 111. Values of the Kinetics Constants Obtained with Commercial FCC Catalysts in This Work and by Various Authors authors T,K A,, s-1 J. or Ad, s-l AI, s-1 I,, s-l Nace et al. (1971) Weekman (1968) Paraskos et al. (1976) Corella et al. (1983, 1985) this work
2.8 x 10-3-1.1 x 10-2 6.3 x 10-3 0.36-0.7
755 755 783-811 793 773
0.19
0.23
5.2 x 10-3-1.1 x 1.2 x 10-2 0.13-0.21 1.58 4.47
MZ-7P O I L G.O.V.
2.2 x 10-3-9.3 x 10-3 5 x 10-3 0.27-0.52 0.12 0.18
3.3 x 10-4-8.3 x 10-4 5 x 10-4 0.03-0.52 6.6 X 3.2 X lo-*
L
0.3 -
(m)
2
0.2 -
0.1
10-2
-
RISER
1
V 6
4
0
12
16
6 Figure 6. Confirmation of d = 3 and calculation of k=oand 4 with eq 19.
0.8
0
20
10
60
80
I00
P, c m H20
Figure 8. Cold model 1 and pressure-position diagram (parameter: superficial gas velocity in the riser, U-).
c
'
0.6 ' X S
calculated (-1
RISER
0.4
0.2
Pt 02
0
0.8
0.6
04
1.0
X S
experimental
(-1
0
Figure 7. Verification of results.
P,cm H20
values of space time and C/O ratio used. The values calculated for the conversion, X, , are plotted in Figure 7 against those found experimenay, X It can be seen that these values are in good agreemen3or all the operating conditions studied. Calculation of kl,k2,and is. Dividing eq 11by eq 10 gives
The experimental data were fitted to this equation by means of an iterative fourth-order Runge-K$ta method. Thg valyes found for the kinetic constants k,, k2,and Eo (=kl k3) at 773 K were -
+
hl = 0.18 s-l
-
k2 = 0.032 s-l
-
k3 = 0.053 s-l
The kinetic constants of cracking and deactivation obtained by different authors and those found in this work are shown in Table 111. There is a remarkable similarity between the cracking kinetic constants obtained by Paraskos et al. (1976) and Corella et d. (1983) and those reported in this work. The values found by Nace et al.
Figure 9. Cold model 2 and pressure-position diagram (parameter: superficial gas velocity at the bottom of the standpipe stripper-regenerator, Ua).
(1971) and Weekman (1968) are much smaller. We believe this is due to the fact that the kinetic parametkrs in these last two works were obtained for other feedstocks and catalysts and from experiments carried out at relatively long times on stream for the FCC process (several minutes), whereas in the first three studies times on stream were of a few seconds. Regarding the deactivation kinetic constant or the simplified deactivation function, $, it is observed that the values found by Corella et d. (1983,1985) and those found in this work are greater than those given by the other authors. We think this is mainly due to the fact that in this-work the determination of $ is based on a value of 3 for d while the other authors quoted above used a firstorder deactivation equation. To t h k respect, Corella et al. (1985) have clearly shown that the J/ value from the data of Paraskos et al. (1976) changes considerably and becomes close to the valye obtained in this work if these data are adjusted with d = 3 instead of 1.
560
Ind. Eng. Chem. Prod. Res. Dev., Vol. 25, No. 4, 1986
Table IV. Main Dimensions of the Cold Models Used model 1 model 2 vessel or pipe (Figure 8) (Figure 9) riser 2.72 3.20 length, m id., m 0.02 0.011 upper fluidized bed, stripper height, m 0.55 0.70 0.12 0.19 i.d. top, m 0.05 0.10 i.d. bottom, m first standpipe (stripper-regenerator) 1.55 length, m 0.02 id., m regenerator 0.60 0.60 height, m 0.12 0.12 i.d. top, m 0.06 0.06 i.d. bottom, m second standpipe (regenerator-riser) 0.65 0.27 length, m 0.02 0.006 i.d., m
Figure 10. Pilot plant with fluidized-bed reactor: (1) gas oil feed pump; (2) reactor; (3) liquid products collection vessel; (4) catalyst preheater; (5) catalyst screw feeder; (6) gas collecting bell. 1 .o
MZ-IP OIL G 0 V. -
500°C
0.L 0.2
0
02
OL
06
08
1
X,(-l Figure 11. Yield to gasolines at the continuous fluidized-bed reactor fs = 0.5 min. exit: ( 0 )f. = 1 min; (X) fs = 0.78 min; (0)
Since in commercial FCC units the times on stream are about 10 s, the values of the parameters to be used in the modeling and control of these units must be thwe obtained in the first few seconds. The determination of the kinetic parameters of the FCC process should, therefore, be made
in installations like the one described in this work so that they can be applied or extrapolated to the commercial units with acceptable accuracy.
Acknowledgment We express our gratitude to the CAICYT (Madrid) and to the U.S.-Spanish Joint Committee for Scientific and Technological Cooperation for their financial support (Projects No. 1984/82 and CCA 8411/083, respectively). We also express our gratitude for the grant awarded to A.F. by the Direccibn General of Energy of the Spanish Ministry of Industry and Energy. We are also grateful to the EMP Refinery in Puertollano, Ciudad Real, Spain, for the catalysts and gas oils used in this work. Appendix I: Previous Studies A. Circulating Cold Models of the Plant. Two cold models of an FCC plant without chemical reaction were built in order to study the circulation between vessels of the solids and gases. These models are shown in Figures 8 and 9. The two cold models were made of transparent glass and methyl methacrylate in order to facilitate the observation of the gas and solid flows inside. Air was used in all the experiments. The numbers shown on the lefthand side of Figures 8 and 9 indicate the location of the manometers. A summary of the most important dimensions of each cold flow model is given in Table IV. The pressure vs. position diagrams for the different cold flow models used are also shown in Figures 8 and 9. They were obtained by varying the air and/or catalyst flow rates. Some important data were obtained from these diagrams. Thus, from the representation of AP/L in the riser against the superficial air velocity, the choking velocity, Uch(the minimum gas velocity at which there is no flooding in the riser), was calculated. In general, the choking velocity increases slightly with the catalyst flow rate. The catalyst used in our experiments was the MZ-7P of Akzo-Chemie (a, = 66.1 pm and pp = 1.8 X kg/m3). The Uchfor this solid, obtained experimentally with air a t room temperature and with a catalyst flow rate of 2 X kg/s, was found to be equal to 0.76 m/s. The calculation of Uchwith gas oil at the cracking temperature was carried out by using the method proposed by Zenz (1964). The results obtained for the catalyst flow rates used in the experimentation were as follows (gas oil, 773 K): Uch(C = 1.82 X kg/s) = 0.68 m/s (C = 1.4 X kg/s) = 0.57 m/s; (C = 1 X kg/s) = 0.50 m/s (C = 8.5 X kg/s) = 0.45 m/s. The terminal velocities of catalyst particles, Ut, were obtained by using the methods indicated by Kunii and Levenspiel (1969) with the following results: Ut (air; 293 K) = 0.17 m/s; (N2,723 K) = 0.46 m/s; (gas oil; 773 K) = 0.38 m/s. Finally, the U,, determined with air and at room temperature was 2.1 X m/s. At high temperatures, the U,, calculated with the correlation of Corella and Otero m/s; (with N2,923 (1971) was (with N2,723 K) 2.4 X K) 2.5 X m/s; (with the gas oil used in this work, 773 m/s. K) 8.2 X These two cold models showed that the most difficult zones to control were the standpipes in which the flow of the solid-gas mixture is descending. In these standpipes obstructions can be easily encountered. The valve and the aeration both at the bottom of each standpipe are key factors for the correct operation of the continuous installation (CBmara, 1984). B. Pilot Plant with Cracking Fluidized-Bed Reactor. Initially, a small pilot plant based on a fluidizedbed cracking reactor was mounted (Figure 10). This installation lacked the regeneration of the catalyst, a cir-
Ind. Eng. Chem. Prod. Res. Dev., Vol. 25, No. 4, 1986 561
cumstance that made it necessary to work with previously equilibrated catalysts taken from commercial FCC units. The preheater of the catalyst in equilibriumwas a fluidized bed of 0.080-m diameter and 0.2-m bed height, and it used Nzas fluidizing gas. It was heated by means of two ovens of 2000 W each, which allowed one to reach quite easily 973 K in the catalyst bed. The cracking reactor was a fluidized bed of 0.040-m i.d. The height of the bed was variable from one experiment to another, between 0.050 and 0.10 m, which allowed a variation of the catalyst average residence time. The gas oil was heated before entering the reaction zone by means of a 2000-W oven. The cracking products passed through a cyclone located inside the reactor. Subsequently, they were cooled and condensed in a series of heat exchangers and collected for analysis. The deactivated catalyst from the reactor passed to two hoppers for its analysis. Some of the results obtained at 773 K in this installation are shown in Figure 11. It may be observed that, under the operating conditions studied, the conversions obtained were very high and the yield to gasolines was in the over-cracking zone. Also the average residence time of the solid (30-300 s) was very high and larger than that in the commercial riser reactors. This plant, therefore, did not allow one to know the behavior of the catalyst in its first second on stream. Finally, this installation lacked a stripper for the catalyst. Because of these drawbacks, it was decided to make the installation described in this work.
Appendix 11: Analysis of the Products The analysis methods for the liquid, gases, and deactivated catalyst used in the experimentation are the following. Gases. The analysis of gases is carried out by gaseous chromatography. The complete characterization of the gaseous sample implies the use of three different chromatographic columns. The first, of the molecular sieve 5D type, is used for the analysis of hydrogen. The carrier gas is nitrogen. The second, of the Carbosieve S I1 type, separates N2,CO, C2H2,CzH4,and C2H6. It uses helium as carrier. Liquids. The liquid samples collected are analyzed in two different ways: (1)by simulated distillation for obtaining the values of conversion of the gas oil, yield to gasolines, and content in gasolines of the liquid sample; and (2) by mass spectrometry for obtaining the sample's composition in naphthenes, paraffins, and aromatics. The simulated distillation by gaseous chromatography is carried out according to ASTM D2887-73 standards, applicable to fractions of petroleum whose final boiling point does not exceed 811 K. The columns used consist of a 10% CV W 982 in P-AW 80/lOO. The carrier gas is nitrogen. The analysis by mass spectrometry includes the separation of the sample taken into fractions of bp 523 K. The fraction of lower boiling point is analyzed directly by mass spectrometry by means of the ASTM D2789-71 standard. In the fraction of greater boiling point a separation is carried out of aromatic and nonaromatic compounds by chromatography in column by means of the ASTM D2549 standard. Both fractions are separately analyzed by mass spectrometry by means of the ASTM D3239-81 and D2786 standards, respectively. Coke in the Deactivated Catalyst. The samples of deactivated catalyst are analyzed by thermogravimetry for determining their coke content. The determinations are carried out in a Perkin-Elmer thermobalance, Model kg at the TGS-2, sensitive to changes in weight of 2 X
maximum sensitivity. As each author uses different experimental conditions for determining the quantity of coke in the catalyst, we believe that these conditions should be set out in detail in each case: In our case, a determined quantity (5 X lO"-l X lo4 kg) of catalyst is heated in an atmosphere of Nzup to 973 K. In this period, the catalyst loses all the volatiles absorbed in its surface. After this, air is introduced at 973 K, and one records the loss of weight of the burning coke. The concentration in coke (C,) of the deactivated catalyst is easily obtained.
Appendix I11 A. Derivation of Equation 14. Substituting eq 6 in eq 13 gives
0 dX =
[go(l - X)2Ait;a]
dW
(111-1)
And substituting now eq 2, we have
:)"I
[
0 d X = R0(l - X)'A(
dW
(111-2)
Since the reactor used is isothermal, the deactivation functien (4 or A) is constant throughout the reactor. On the other hand, for a given experiment, the catalyst and gas oil flows (C and 0, respectively) are kept constant. With these considerations, eq 111-2 can be written as
which upon integration yields
X,
LOA
-=--
CB
0 (1 - 0)
(1 - X,)
W'-a
(111-4)
This equation can be rearranged to give
The above expression can be linearized by taking logarithms:
(111-6) And using eq 2, this equation becomes eq 14. B. D-erivation of Equation 18. The substitution of eq 17 (d = 3) in eq 13 gives 1 0 d X = io(1- x)2 d W (111-7) 1
[
Rearranging gives
--d X (1 - X)2
+
-
- -Eo O
[
1
1
dW
(111-8)
+2 3 . y 2
For a given experiment, 0, C , and $ remain constant (isothermal reactor). Thus integrating yields
or equivalently
562
Ind. Eng. Chem. Prod. Res. Dev., Vol. 25, No. 4, 1986
This equation can also be written as
(&
g+
1 y= 1
fg
+ 2 4 -W
(111-11)
t, f, uch
C
U,,
ut
which transforms into
W X
x, 4G
After simplification and rearrangement, the above expression becomes
x, _4 _0 -- 2ko-w (1 -&I (1- X,) k,c 0 x, 2 =
~
(111-13)
This equation can be written as
which is eq 18. Nomenclature activity of the catalyst, averaged for all the reactions of the network and strengths of active sites parameter of eq 6 mass flow rate of catalyst fed to the riser, kg/s coke concentration on the catalyst gases (IC,) average particle diameter order of deactivation (average or observable) gasoline, lump formed by hydrocarbons between C5 and 483 K rate constant of the cracking reaction of the gas oil, S-1
kinetic constants of formation of gasoline starting from gas oil; of formation of gases starting from gasoline; and of formation of gases starting from gas oil, respectively, s-l weight of gas oil fed in a At (a0 At), kg gas oil, group of hydroczrbons with boiling point greater than 483 K mass flow rate of gas oil, kg/s weight fraction of gasoline in the liquid sample rate of formation of gasoline, of gases, and of cracking of gas oil, respectively above values at zero time temperature time, s
$.
2 7
average time of residence of the gas, s time on stream of the catalyst, s average t , for all catalyst particles, s chocking velocity, m/s minimum fluidization velocity, m/s terminal velocity of particles, m/s weight of the catalyst in the reactor, kg conversion of gas oil, wt % conversion of gas oil at the reactor outlet, wt % yield to gasoline yield to coke simplified function of deactivation, s-l density of the catalyst particles exponent in eq 6 space time of the gas oil
Literature Cited Abbot, J.; Wojciechowski, B. W. Ind. Eng. Chem. Prod. Res. Dev. 1985, 24, 501. Blanding, F. H. Ind. Eng. Chem. 1953, 4 5 , 1191. Cimara, R. Ph.D. Thesis, University of Zaragoza, Spain, 1984. Cimbalo, R. N.; Foster, R. L.; Wachtel, S. J. Oil Gas J. 1972, 70, 112. Corella, J.; Otero, A. M. An. Qulm. 1971, 72, 1221. Corella, J.; MenBndez, M.; Chem. Eng. Sci. 1986, 4 7 , 1817. Corella, J.; Bilbao, R.; Molina, J. A.; Artigas, A. Proc. Inter-Am. Congr. Chem. Eng. 1983, 1 , 133. Corella, J.; Bilbao, R.; Molina, A.; Artigas. A. Ind. Eng. Chem. Process Des. Dev. 1985, 24,625. Corella, J.; Bilbao, R.; Gonzilez, J.; MonzBn, A. An. Quim. 1986a, in press. Corella, J.; Molina, J. A.; MenBndez, M. Ind. Eng. Chem. Process Des. Dev. 1988b, submitted for publication. Herring, W. M.; Hinman, J. E.; Shields, S. E. Chem. Eng. Prog. 1963, 59, 38. Humes. W. H. Chem. Eng. Prog. 1983, 79(Feb), 51. Krishnaswamy, S.; Klttrell. J. R. Ind. Eng. Chem. Process Des. Dev. 1978, 17, 200. Kunii, D.; Levenspiel, 0. Fluidization Engineering;Wiley: New York, 1969; p 76. Lewis, C. Y.; Wrende, D. E. Presented at Katalistiks, 5th Annual Fluid Catalytic Cracking Symposium, Vienna, May 1984. Masologites. G. P.; Beckberger, L. H. Oil Gas J. 1973, 79(Nov), 49. Nace, D. M.; Voltz, S. E.; Weekman, V. M. Ind. Eng. Chem. Process Des. Dev. 1971, IO, 530. Paraskos, J. A.; Shah, Y. T.; McKinney, J. D.; Carr, N. L. Ind. Eng. Chem. Process Des. Dev. 1978, 15. 165. Pine, L. A.; Maher, P. J.; Wachter. W. A. J. Catal. 1984, 85, 466. Santos, V. A.; Dantas, C. C. Chem.-Ztg. 1983, 32, 289. Shah, Y. T.;Huling, G. P.; Paraskos, J. A.; McKinney, J. D. Ind. Eng. Chem. Process Des. Dev. 1977, 16, 69. Voorhles, A. Ind. Eng. Chem. 1947, 37,318. Wachtel, S.J.; Baillie, L. A.; Foster, R. L.;Jacobs, H. E, Oil Gas J . 1972, 70, 104. Wagner, M. C.; Humes, W. H.; Magnabosco, L. M. Plantloper. Prog. 1984, 3 , 222. Weekman, V. W. Ind. Eng. Chem. Process Des. Dev. 1968, 7, 90. Wollaston, G.; Haflin, W. S.; Ford, W. D.; D'Souza. G. J. Hydrocarbon Process. 1975, 54, 93. Zenz, F. A. Ind. Eng. Chem. Fundam. 1964, 3 ,65.
Received for review April 23, 1985 Revised manuscript received February 25, 1986 Accepted May 3, 1986
