Fluid Catalytic Cracking II - American Chemical Society

Chapter 9

Translation of Laboratory Fluid Cracking Catalyst Characterization Tests to Riser Reactors Downloaded by NORTH CAROLINA STATE UNIV on August 6, 2012 | http://pubs.acs.org Publication Date: January 23, 1991 | doi: 10.1021/bk-1991-0452.ch009

A. V. Sapre and T. M . Leib Paulsboro Research Laboratory, Mobil Research and Development Corporation, Paulsboro, NJ 08066

A theory has been developed which translates observed coke -conversion selectivity, or dynamic activity, from widely-used MAT or fixed fluidized bed laboratory catalyst characterization tests to steady state risers. The analysis accounts for nonsteady state reactor operation and poor gas-phase hydrodynamics typical of small fluid bed reactors as well as the nonisothermal nature of the MAT test. Variations in catalyst type (e.g. REY versus USY) are accounted for by postulating different coke deactivation rates, activation energies and heats of reaction. For accurate translation, these parameters must be determined from independent experiments. This work provides conclusive evidence that transient catalyst characterization tests can result in erroneous catalyst ranking. For example, USY catalysts show higher activity than REY catalysts in the laboratory tests but lower activity in a steady state riser. Although emphasis in this paper is placed mainly on the coke-conversion selectivity, the analysis is also extended to yields of other FCC products.

While most catalyst vendors rely on fixed bed microactivity (MAT) tests, fixed fluid bed (FFB) reactor experiments are widely used within Mobil to characterize FCC catalysts. The amount of catalyst used is constant for each test, and products are collected for a known period of time. In MAT experiments, catalyst bed is fixed while in FFB test the catalyst bed is fluidized. As products are collected over the decay cycle of the catalyst, the resulting conversion and coke yields are strongly influenced by catalyst deactivation. Systematic differences exist between the measured conversion or catalyst activity and coke yields for the MAT and FFB tests. The magnitude of these differences varies depending on the type of catalyst being tested (REY or USY). Experimental data in Figure 1 clearly show that FFB conversion is higher than MAT conversion for USY catalysts. On the other hand, FFB conversion is lower than MAT conversion for REY catalysts. Furthermore, the quantitative 0097-6156/9iy0452-0144$06.25A) © 1991 American Chemical Society

In Fluid Catalytic Cracking II; Occelli, M.; ACS Symposium Series; American Chemical Society: Washington, DC, 1991.

Translation of FCC Tests to Riser Reactors

Downloaded by NORTH CAROLINA STATE UNIV on August 6, 2012 | http://pubs.acs.org Publication Date: January 23, 1991 | doi: 10.1021/bk-1991-0452.ch009

SAPRE AND LEIB



146

FLUID CATALYTIC CRACKING H: CONCEPTS IN CATALYST DESIGN

translation of the catalyst performance from these laboratory tests to steady state riser is a strong function of catalyst type. Therefore, we recognized that a fundamental approach relating laboratory catalyst screening tests to riser performance is needed. In this paper we summarize a theory which quantitatively accounts for these differences and give insights into interpreting laboratory test data. Potential pitfalls exist in ranking catalysts based solely on correlations of laboratory tests (MAT or FFB) to riser performance when catalysts decay at significantly different rates. Weekman first pointed out the erroneous conversion ranking of decaying catalysts in fixed bed and moving bed isothermal reactors (1-3). Phenomena such as axial dispersion in the FFB reactor, the nonisothermal nature of the MAT test, and feedstock differences further complicate the catalyst characterization. In addition, differences between REY, USY and RE-USY catalyst types exist due to differences in coke deactivation rates, heats of reaction, activation energies and intrinsic activities. In this paper, we will first illustrate the mathematical models used to describe the coke-conversion selectivity for FFB, MAT and riser reactors. The models also include matrix and zeolite contributions. Intrinsic activity parameters estimated from a small isothermal riser will then be used to predict the FFB and MAT data. The inverse problem of predicting riser performance from FFB and MAT data is straightforward based on the proposed theory. A parametric study is performed to show the sensitivity to changes in coke selectivity and heat of reaction which are affected by catalyst type. We will highlight the quantitative differences in observed conversion and coke-conversion selectivity of various reactors. Theory Coke-Conversion Kinetics. Coke formation kinetics and gas oil conversion are represented by the following irreversible reactions (cracking and coking): cracking >

gas o i l

\

products

(1)

/ coke

Based on Voorhies time-on-stream theory (4), catalytic coke is a function of catalyst contact time: AcÂtc"

(2)

and the coke yield on fresh feed is given by:

coke = C C R + Ac[C/0]

(3)

Typically less than 100% of the feed CCR goes to coke depending on the feedstock and operating conditions. In Equation 3, however, for simplicity we assume that 100% of the feed CCR is deposited as coke on catalyst.


9. SAPRE AND LEIB


147

The reaction rate per unit volume for gas oil cracking is given by: r = r *(t)

(4)

0

where r is the initial cracking rate, and Φ(Χ) is the catalyst deactivation function. We assume that the catalyst deactivates due to coke formation, and the cracking activity declines in direct proportion to the rate of coke formation. The deactivation function therefore may be written as: 0


*(t) = dAc/dt

(5)

Similar dependence of catalyst deactivation on coke or catalyst residence time is suggested by Corella et al. (5-7). The authors give details on possible mechanisms of catalyst deactivation by coke, and also suggest, based on their data, that the deactivation order η may not be a constant. For our analysis, however, we will assume that η is constant and a function of catalyst type. Further theoretical treatment of catalyst decay is given by Wojciechowski (8,9). The effect of feedstock quality and oil partial pressure on coking is given by: n

Ac = A t ( a k C ) P c

+

A

A

b

(6)

oiI

Blanding (10) first proposed the second order cracking kinetics for FCC. Krambeck (11) theoretically demonstrated that conversion in systems with a large number of parallel reactions can be approximated by simple second order kinetics. More recently, Ho and Aris (12) have developed a further mathematical treatment of this concept. An inhibition term was incorporated into the second order cracking kinetics for gas oil conversion to account for competitive adsorption. The initial cracking rate is then given by: r

k

c

2/

o = Pb A (1+k C ) HA

(7)

A

Differences between various zeolites and matrix components of the catalyst have been accounted for by fitting coke deactivation rates (exponent n), coking and cracking activities (A and k), and adsorption coefficients (k and k ). Similar ideas on deactivation of a composite cracking catalyst have been presented by Dean and Dadyburjor (13). Coke on catalyst then becomes the sum of the coke on the matrix and the coke on the zeolite: A

HA

(8)

Ac = ZACjZj

Here Z is the mass fraction of zeolite or matrix components. {

And the deactivation functions become: dACi/dt = î(t)(a k C )P i, +

Ai

A

b

(9)

0

*j(t) = dACj/dt

(10)

Similarly, the cracking reaction will be composed of the reactions occurring on the


148

FLUID CATALYTIC CRACKING II: CONCEPTS IN CATALYST DESIGN

matrix and zeolite: r = lrj

(11)

where z

1 +

12

n= Mi iC V( W>A)

(>

A


Although the theory includes the impact of matrix contribution, we will restrict our discussions primarily to low matrix activity catalysts, and highlight the differences in zeolite type. FFB Model. Use of FFB units in the industry for catalyst characterization is primarily due to excellent temperature control and resulting isothermal reactor temperature. The operating conditions of the FFB activity unit used in our study are given in Table I. The material balance using a pseudo steady state assumption (at any given instant the vapor is in contact with catalyst of almost uniform activity) gives: 2

2

* D d C / d z - d(u C )/dz = 2rj(C ,t) A

A

A

(13)

A

where r is given by Equation 12. In the above equation, the hydrodynamics of FFB reactors are represented by an axial dispersion model. The degree of bubbling, and hence, the bypassing are reflected in the dispersion coefficient. The superficial gas velocity u changes along the reactor due to molar expansion. Assuming u varies linearly with conversion: ;

u = u (1+«y X)

(14)

X = 1-uC /(u C )

(15)

0

Ao

Where conversion X is given by A

0

Ao

Danckwert's boundary conditions give: @ z = 0 u(C -C ) = -eD dC /dz

(16)

@ z = LdC /dz = 0

(17)

Ao

A

A

A

A

Equations 13-17 expressed in dimensionless form give: 1

2

2

Pe" d y/dx - d(v y)/dx =

falê)

(18)

v=1+«y X A o

(19)

X = 1-vy

(20) _1

@ x = 0 v(1 -y) = -Pe dy/dx

(21)

@x = 1 dy/dx = 0

(22)


9. SAPRE AND LEIB

149


Here, 1

2

fj = G> e \- ; n

6>j = kiAiniZjM C /(S (1 +k C )) A

m

n

1

»?i = t r ;

Ao

v

HA|

Ai

Pe = L u / ( e D ) 0

y=C /C ;

x = z/L

v = u/u ;

$ = t/t

A

A o

0

A

c

S = (t [C/0]) •1 Downloaded by NORTH CAROLINA STATE UNIV on August 6, 2012 | http://pubs.acs.org Publication Date: January 23, 1991 | doi: 10.1021/bk-1991-0452.ch009

v

c

Solution of Equations 18 to 22 gives the instantaneous outlet concentration y, which when integrated over the run length gives the average conversion (X): 1 X = 1-

vy d0

(23)

o The crackability (Cr), or the second order conversion is given by: Cr = X/(1-Xl

(24)

A parameter k representing coke-conversion selectivity, is defined as: c

l^ = coke/Cr

(25)

The use of k in analyzing data from pilot units was proposed by Krambeck in the early 1970's and has been used in Mobil since then. More recently, the same concept has been published in the open literature, and the reciprocal of k is defined as UOP "dynamic activity" (14). The dynamic activity is now popularly used in the FCC literature, and is even used to correlate catalyst performance with fundamental catalyst parameters such as unit cell size (15). In this paper, however, we will use the Mobil defined k parameter. c

c

c

Table I. FFB, MAT and Riser Operating Conditions

FFB

Reactor T.-F C/0,wt/wt t ,s Poil, psia c

Gas Oil C ,wt% CCR, wt % M A

A

850 2 300 10 Light East Texas 10 .01 233

MAT (Davison) Fixed Bed 930 3 75 15

18 .40 400


150


MAT Model. The operating conditions of Davison's MAT unit (16) are given in Table I. These conditions correspond to testing prior to the change in November 1988. The material and energy balances, using pseudo steady state and adiabatic plug flow assumptions give: -d(u C )/dz = Lr|(C ,T,t) A

w

c

c pc

d T / a t

+

m

(26)

A

c

A pA

L a T / d z

88

-M V£AHjrj(C ,T,t) A

A

(27)


The inlet conditions are:

@z =0 C

A

=C

A o

,T =T

(28)

0

The partial differential Equation 27 can be simplified to an ordinary differential equation in the independent variable z, by the following approximation: For typical FCC catalysts and feedstocks, c - 3 x Cp , and for the MAT test, [C/O] is 3 (Table I). Hence, if we assume that the catalyst and oil temperatures are identical (no heat transfer resistance between oil and catalyst), then as a first approximation, the change in the energy content of the oil and of the catalyst are roughly the same over the length of the run. Thus, the two terms on the left hand side of Equation 27 are approximately the same magnitude. Therefore, the time derivative of T can be lumped with the distance derivative. The right hand side of Equation 27 is divided by 2, and it becomes: p A

C

m c L d T / d z = -M VE(AH/2)rj(C ,T,t) A

pA

A

A

(29)

This approximation of the energy balance may lead to inaccurate prediction of the temperature profile as a function of time on stream. Thus, at short contact times it may underpredict the temperature, while at long contact times it may overpredict the actual temperature profiles. However, on the average, the model predictions are good, as seen from the comparisons with data in subsequent sections. In our opinion, the above simplifications are reasonable, especially since the reproducibility of the MAT test is somewhat poor. Normalizing Equations 27-29 gives:

2

-d(vy)/dx = y Zfj(0)

(30)

ds/dx = y E0jfi(0)exp(-e|/s)

(31)

@x = 0 y = s= 1

(32)

2


9.

151


SAPRE AND LEIB

Here, s = T/T

0i = AH|/(2Cp T )

0

A

ei = Ei/(RT )

o

f^f/expf-ej/s)

0


The other parameters are the same as for the FFB model. Similar to the FFB model, Equations 30-32 give the instantaneous conversion; to obtain time averaged values, integration is required for X (Equation 23) and T:

T =

(33)

T 60

Forissier and Bernard (17), also give a model for the MAT test. However, since they did not attempt to translate product selectivities from the MAT test to a steady state riser, it is somewhat less complex. Riser Model. The riser is represented by a steady state adiabatic plug flow model with no slip between oil and catalyst (18). The material and energy balances give: -d(u C )/dz = Lrj(C ,T) A

(34)

A

(m Cp +m c )LdT/dz = -M VZAHjrj(C ,T) c

C

A

pA

A

A

(35)

The inlet conditions are identical to Equation 28. Equations 28,34 and 35 can be written in dimensionless form similar to Equations 30-32. The definition of the dimensionless variables is identical to those of the MAT model, except for p{ 0j = AH^CpATÎ+mcCpc/tmACpA))) Since the riser operates at steady state no time averaging is required. Numerical Method. Both the isothermal FFB and the adiabatic MAT models are very stiff due to the coke deactivation terms \. The spline orthogonal collocation technique was used to solve the above models (19). Typically, the distance x was divided into two regions (0 oo can be easily derived, and also will not be shown here. Equations (36) were solved numerically using the spline collocation technique discussed before. The product selectivities in various reactors for the following relative rates, k /k = 0.5; k^/k-i = 0.7; k^-j = 0.1 k /k = 0.05; kg/k^ = 0.035 are shown in Figure 7. The product yields in Figure 7 correspond to steady state riser (Pe -» »), isothermal MAT (Pe - * » ) , and FFB (Pe - 1) reactors. The yields of desirable products, gasoline and LCO are always lower in transient units due to time averaging. The axial dispersion causes further decline in product selectivity. In Figure 7, the maximum in gasoline yield occurs much earlier in the FFB reactor compared to MAT, or steady state riser reactors. There are sufficient data in the literature to verify that overcracking of gasoline occurs earlier in laboratory fluid bed reactors than in risers [23,241. Thus, if laboratory reactors are used to rank catalysts based on product selectivities, one should consider the above differences. Furthermore, at conversions above those which give maximum gasoline yields, there is a sharp increase in the light hydrocarbon yields and increased gasoline octanes. The product qualities from the laboratory transient reactors, like octane for example, should therefore be interpreted cautiously. At the same conversion level, Research Octane numbers of the gasoline obtained from a transient laboratory reactor could be higher than those of steady state risers. The effect of time averaging on yields in transient tests can be minimized by shortening the duration of the test. Also, a fixed bed test is superior to an FFB activity test in that backmixing is minimized. Furthermore, an isothermal fixed bed test would be easier to interpret than the adiabatic MAT test. This work shows that from the point of view of a catalyst characterization test, a small steady state riser will give the most direct information for catalyst performance in a commercial riser. f

4

3

1

1



160


Crackability

Figure 6: Predicted coke-conversion selectivity as a function of catalyst activity (crackability) for USY catalysts.

Fractional Concentration

Average Conversion, %

Figure 7: Predicted product selectivities as a function of conversion for three different laboratory catalyst characterization test units.


9.

SAPRE AND LEIB


161

Summary


F C C catalyst testing prior to use in commercial reactors is essential for assuring acceptable performance. Purely correlative relations for ranking catalysts based on laboratory tests, however, can be erroneous because of the complex interaction of the hydrodynamics in the test equipment with the cracking kinetics. This paper shows how the catalyst activity, coke-conversion selectivity and other product selectivities can be translated from transient laboratory tests to steady state risers. Mathematical models are described which allow this translation from FFB and MAT tests. The model predictions are in good agreement with experimental data on identical catalysts run in the FFB, MAT and a laboratory riser. Legend of Symbols A

9A

c

A

CCR Cr C/O D f f

A

*i k k

A c HA ko L m M n Pe r s k

k

A

s

t

v

«c T u V

V w

c

X

Intrinsic coking kinetic constant Heat capacity at constant pressure Volumetric concentration of gas oil Total carbon in aromatics Conradson carbon residue Crackability Catalyst/oil ratio, wt/wt Axial dispersion coefficient Dimensionless reaction parameter for FAI Dimensionless reaction parameter for MAT and riser Dimensionless constants in Equation (36) Ratio of cracking to coking rate constants Intrinsic coking reaction constant Coke-conversion selectivity Adsorption constant Exponential factor of k Reactor height Mass flow rate Oil molecular weight Voorhies parameter Axial Peclet number Reaction rate per reactor volume Dimensionless temperature Space velocity Time Catalyst contact time Temperature Superficial vapor velocity Dimensionless vapor superficial velocity Reactor volume Catalyst holdup in reactor Dimensionless axial distance


162


X y y y YQ y z

Conversion Dimensionless volumetric concentration Fraction of oil in feed, mol/mol Dimensionless concentration of LCO Dimensionless concentration of gasoline Dimensionless concentration of gas Axial distance; fraction of zeolite in catalyst, wt/wt

A

B

D


Greek Letters Axial molar expansion coefficient Dimensionless heat of reaction Dimensionless activation energy Heat of reaction ACoke on catalyst Dimensionless time Dimensionless reaction parameter Dimensionless reaction parameter Spline point Coke deactivation function

a 0 e AH Ac e (j 17 7 4 Subscripts A c 1 0

Oil Catalyst 1=zeolite, 2=matrix feed

Appendix Estimates of Model Parameters. The reactor models for FFB, MAT and riser include important features for translating the MAT and FFB data to steady state riser performance. A series of key parameters specific to a given zeolite and matrix component are needed for a given catalyst. Such key parameters are intrinsic cracking anc| coking activities (kj, Aj), activation energies and heats of reaction (Ej, AHj), coke deactivation rate (exponents n,-), and axial dispersion in the FFB unit (D ). Other feedstock dependent parameters include the inhibition constants ( k ) , the coking constants (K ), and the axial molar expansion factor (a). It is the objective of the FFB or the MAT test to determine cracking and coking activities of the catalyst (kj, Aj). Therefore, all other parameters should be known from independent experiments or estimated prior to determining these intrinsic catalyst parameters. These other parameter estimates are described below. It can be shown, that the molar expansion parameter a is a function of the charge stock and product molecular weights: A

HAl

Ai

M is the molecular weight of the charge stock, and M is the average molecular weight of the product. The value of a varied from 2 to 6 in this study. 0


9. SAPRE AND LEIB

163


An axial Pe number of 5 was estimated for the FFB unit, corresponding to a D of -30 cm /s. This estimate was based on dispersion data reported by Zenz (25) on Geldart's Group A powders. The coke deactivation exponent n, is typically estimated from riser pilot plant experiments at varying catalyst contact time for different catalyst types. A value of n of 0.2 was found for REY catalyst data base. For USY and RE-USY catalysts n was estimated to be 0.4. Heats of reactions were estimated from heats of formations and chemical compositions of feed and product using standard procedures. For REY catalysts, we estimated approximately 130 Btu/lb heat of reaction. The heat of reaction was close to 200 Btu/lb for USY catalysts. These values are in close agreement with reported data (21J. The activation energies for different catalyst types were estimated from our extensive pilot plant data base, and found to be a weak function of catalyst type. The adsorption constants and other kinetic parameters used in these simulations were fitted to a large in-house data base. Typical parameter values are reported in Tables III and V. The kinetic parameters (kj and Aj) are a strong function of catalyst used, whereas the adsorption parameters were found to be relatively insensitive. One could estimate these parameters even from a limited data base as illustrated below for Catalyst D.

A


2

Table VI. Commercial Equilibrium REY Catalyst (Catalyst D) FFB Test Conversion wt% 63 63 64 64 64 65 66

MAT Test Conversion wt% 66 68 69 71 71 72 72

Coke wt% 0.43 0.44 0.47 0.49 0.50 0.51 0.55

Coke wt% 1.56 1.51 1.59 1.51 1.51 1.85 1.59

A set of MAT and FFB data on a low metals equilibrium commercial REY catalyst with a low matrix activity (Catalyst D) is summarized in Table VI. The range of activities indicates varying steam deactivation levels in the commercial unit. These data could be used to estimate the following four parameters: intrinsic cracking and coking activities (kj, Aj) and feedstock parameters ( k , k ). The feedstock differences between Light East Texas Gas Oil (LETGO) and West Texas Gas Oil used in the FFB and MAT tests, respectively, allows the determination of k and k . A modified Marquardt algorithm was used for parameter estimation. The fit was meaningful at a 95% confidence level, and the correlation coefficient was 0.93. These parameters are summarized in Table V. These parameter values were in close agreement with other REY catalysts, for example Catalyst A in Table III. HA

A

H A

A


164


Acknowledgments Useful discussions with F. J . Krambeck are sincerely appreciated. We thank D. M. Nace for obtaining the experimental data. Also, D. H. Anderson's effort in programming the spline orthogonal collocation technique is appreciated.


Literature Cited 1. 2. 3. 4. 5. 6. 7. 8. 9. 10. 11. 12. 13. 14. 15. 16. 17. 18. 19. 20. 21. 22. 23. 24. 25.

Weekman, Jr. V. W. I&EC Proc. Des. & Dev. 1968, 7, (1), 90. Weekman, Jr. V. W. I&EC Proc. Des. & Dev. 1969, 8, (3), 385. Weekman, Jr. V. W. and Nace, D. M. AlChE J. 1970, 16, (3), 397. Voorhies, A. I&EC 1945, 37, 318. Corella, J. and Menendez, M. C E S , 1986,41.. 1817. Corella, J., Bilbao, K., Molina, J. A. and Artigas, A. I&EC Proc. Res. Dev., 1985, 24, 625. Corella, J., Fernandez, A. and Vidal, J. M. I&EC Prod. Res. Dev., 1986, 25, 554. Wojciechowski, B. W. Can. J. Chem. Eng. 1968, 46, 48. Wojciechowski, B. W. Cat. Rev. Sci. Eng. 1974, 9(1), 79. Blanding, F. H. I&EC. 1953,45, 1186. Krambeck, F. J . Int. Chem. E. Symp. Ser.,(ISCRE-8), 1984, 733. Ho, T. C. and Aris, R. AlChE J.. 1987,33, (6), 1030. Dean, J . W. and Dadyburjor, D. B. Ind. Eng. Chem. Res., 1989, 28, 271. Moh, R. W. Oil & Gas Journal. 1987, January 26, 73. Rajagopalon, K. and Peters A. W. Journal of Catalysis, 1982, 106, 410. Davison Catalagram, 1988 Nov. 1. Forissier, M. and Bernard, J. R. AlChE Spring Meeting, 1989, Houston, April 2-6, Paper 84d. Yen, L. C., AlChE Spring Meeting, 1989, Houston, April 2-6, Paper 84b. Villadsen, J. and Michelsen, M. L. Solution of Differential Equation Models by Polynomial Approximations. Prentice Hall, Inc., Englewood Cliffs, NJ, 1978. Cronkright, W. A., Butler, M. M. and Harter, D. A. Ketjen Catalyst Symposium, Kurhaus, The Netherlands, 1986, May 25-28. Leuenberger, E. L. and Wilbert, L. J. Oil & Gas Journal. 1987, May 25, 38. Desai, P. H. Oil & Gas Journal. 1986, September 22,42. Creighton, J. E. and Young, G . W. ACS National Meeting, Philadelphia, 1984. Gross, B., Nace, D. M. and Voltz, S. E. I&EC Proc. Des. & Dev., 1974,13, (3), 199. Zenz, F. A. and Othmer, D. F. Fluidization and Fluid Particle Systems, Reinhold Publishing Co., New York, 1960, 300.

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Fluid Catalytic Cracking II - American Chemical Society

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