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Modelling and experimental studies of coking in a structured catalyst
Pergamon
ChemicalEngineeringScience, Vol. 51, No. 11, pp. 2939-2944, 1996 Copyright © 1996 Elsevier Science Ltd Printed in Great Britain. All righut reserved 0009-2509/96 $15.00 + 0.00
S0009-2509(96)00178-9 Modelling and Experimental Studies of Coldng in a Structured Catalyst O. H. J. Muhammad and E. K. T. Kam Petroleum Tecl'mology Department~ K u w a i t Institute for Scientific Research, P.O.Box 24885, Safat 13109, K U W A I T .
Abstract - A stochastic model has been developed to account for the catalyst decay by coking based on the
interaction of the geometries of the coke deposit and the pore structure. Coke is assumed to deposit randomly on the substrate macropores as well as in zeolite micropores resulting in both the active site coverage and pore blockage. The substrate pore structure is represented by a corrugated parallel bundle pore model. The mass balance equations are solved analytically to simulate conversion, coke content and surface area as a function of time-on-stream. Three different fouling mechanisms: parallel, series and triangular, are also investigated and from the simulated results, the series fouling mechanism is tbund to fit the experimental observations in a modified micro activity testing (MMAT) unit better. INTRODUCTION
The performance of catalysts is strongly influenced by diffusion of reactants and products in the catalyst substrate so the overall performance is influenced by the interaction between the intrinsic kinetics and the transport processes such as heat, mass and momentum transfer in catalyst supports (Cresswell, 1985; McGreavy et al. 1994). This is further complicated by the randomness of the pore networks and active unit cells which are very important to zeolite catalysts since they are mainly for resid processing to maximize the octane-barrel in fluid catalytic cracking (FCC) units. In particular, the rare-earth exchange/stabilization is highly suitable for the FCC's LPG/Gasoline quality requirements. However, the catalyst effectiveness also depends on the unit cell range and the matrix of the catalyst subswates, the feedstock characteristics and the accompanied deactivation contaminants such as vanadium, nickel, coke and sulphur. Most of these organo-fouling compounds found in petroleum feedstocks exist in various sizes and shapes but they are predominantly in relatively large molecular structures. Ideally, the catalyst pores should exert minimal diffusion resistance to enable the full utilization of all the catalyst active area but without compromising the catalyst strength. New types of catalysts are developed with specifically designed pore structures customized to the desired selectivities. This requires the optimization of the pore diameters, catalyst connectivity and access to closeend pores and hence, results numerous studies and some network models have been developed (Nicholson and Petropouios, 1968; Androutsopoulos and Mann, 1978; Mo and Wei, 1986; Burganos and Sotirchos. 1989, Hollewand et al., 1991; McGreavy et al. 1991 and Muhammad, 1992) to get insights in the diffusion, sorption and/or reaction, and deactivation in catalysts with random networks of pores. While such works have demonstrated the close agreement of theory with experiments, it has required considerable computational effort and the parameters used are confmed to catalysts but not the reactors. The present work is to develop a mathematical model to interpret experimental data and special attention has been directed to examine the catalyst deactivation due to coke deposition on the catalyst support. The application of the model to correlate the experimental results has shown that the timewise deactivation could be represented by either a substrate active site poisoning coking or a pore plugging mechanism due to one of the products. MODEL DEVELOPMENT
The zeolite contained within crystaUites is assumed to be uniformly distributed along the walls of silica alumina support pores. The diffusion, reaction and deactivation occur in both the support and zeolite structures as functions of time on stream. The deactivation is in the form of coke laydown in either structure through two possible mechanisms - active site poisoning and pore plugging. Other assumptions are: 1. mass transfer is by diffusion only without convection. 2. Dalton's Law of partial pressure applies, 3. the total pressure is constant and hence the molar density and total number of moles are constant 4. first order reaction is applied to the reactant as well as the active surface areas, 5. both the support and zeolite are catalytically active, and 6. the support and zeolite can be deactivated in different ways. Consider a single pore consisting of N elements in which the reaction is assumed as: A ~ Product. A mass balance for the reactant A in any pore element n, n < N, see in Figure 1, can be expressed as: n R2n D n d2CA
= ks Sn CA
(1)
For the pore elements n < N, the boundary conditions are: CA = CA, n-1
at x = 0
CA = CA.n
atx = L
(2) 2939
O.H.J. MUtlAMMAD and E. K. T. KAM
2940
ElementNumbe~ ~
~
Zeo4~ Mzroperes
SuppoctPore
._ ~(T'~_ _ .......
0
....
0
.............
t'-
t~l [Z/:l
I I
I I
I
CA!
f..-
Close end of
......
_ J •
L~ ~-~
)
I I )
CA0
.............
) I I
I I
CAn-2
CA2
I )
CAn-I
CAn
I I
CAN-2
CAN Intermediate Co~ent'e~ons
CA,N.t
l
I I
dCAN I
x~ F3gure 1. A pore schematic
x=L imviag N
elements.
and the analytical solution of Equation (1) becomes CA(X ) =
CA, n.I sinh[mn (L- x)] + CA, n sirdl(mn x) sinh(m n
(3)
L)
For the close-end pore element, n = N, the boundary conditions are as follows: CA dCA dx
= CA,N. I at X = O = 0
(4)
at x = L
and Equation (3) becomes,
CA(X ) =
CA, N. l cosh[m N (L- x)] cosh(m
n
(5)
L)
Concentration Profiles in the Intermediate and End Pore Elements To simulate the concentration profile in the complete pore by coupling all the intermediate concentrations together, the flux between two adjacent elements must be equal, i.e., Dn. | Rn2.1 d C A
= D n R2n dCA
dx x=L
(6)
dx x=O
Since catalysts are undergoing deactivation, any pore element may be either active or inactive at any specific time. Based on the conditions of the two adjacent pore elements, four different cases can exist as listed in Tables 1 and 2 for the close-end pore and intermediate elements respectively. The concentration expressions for each case are also shown. Table 1. The four activity cases for two adjacent pore elements with one end closed Case Number 0
•
O
Element Number Status (N-l)
Active
(N)
Active
(N-l)
Active
(N)
Inactive
(N-l) (N) (N-I) (N)
Mass Balance Expressions ^ . DN.!R2.l m~-I ] _ UA,N-2[ sinh(m N L)
(7a)
CA,N-I [DN.I R2.1 ruN-| coth(mN.l L) + D N R 2 mNtaah(m N L)] = 0 CA,N ,~_ CA,N_I cosh(mN.l L) = 0
(Tb)
Inactive Active
,,', t DN'I R ~ - I , ,,', , DN'I R~.I ~'A,N-2t""-'~J - ~A,N-lt L + DN R ~ m N utah ( m N L)] = 0
(7C)
Inactive Inactive
CA,N_2 - CA,N_I = 0
(7d)
2941
Coking in a structured catalyst Table 2. The four activity, cases for two adjacent pore elements Case Number
Element Number Status
0
(n-l)
Mass Balance Expressions r, r Dn'I R2n'i mn'l Dn R2n m n ~A,n-2t sinh(mn, t L) ] + [ s]i n h-( m n- L )
Active
(n)
Active
(n-l)
A~ive
Dn. 1 R2.1 ran. 1
0
O
On Rn2 m n
CA,n-2[sinh(mn.1L) ] + C A ' n [ ~
(n)
Inactive
(n-l)
Inactive
'
L
•
CA,hill
Active
(n-l)
Inactive
(n)
Inactive
(Sb)
r~ u2 D R2 CAn_2[~n-I AXn.I ] _ C A n _ l [ n-I n-i + D n R 2 n m n c o t h ( m n L ) ] + Dn R2 mn
(n)
]-
CA,n_l[Dn.l R2 1mn.lcoth(mn_lL) + D~-R-~2n ]=0
C
O
(Sa)
CA.,-I [Dn-1 R2n.l mn-1 eoth(mn.l L) + D n R 2 m n coth(m n L)] = 0
L
(8C)
=0
Dn 1 R2 Dn 1 R2 • Dn R2n t - n-~ ~ C r - n-~ + ]+ A,n-2 t - - - - - - ~ ' ~ l - A,n- 1t - - - - " - i ~ - L 2
8d)
."~A,nt,D.LR.~l -_ .
Rate o f Reaction in a Pore Element Due to the active status of the leading (first) element in a pore, two cases can occur: Case 1: the leading element of the pore ( n - l ) is active - r A = - D 1 z R~ [
(9)
m] CA, 1 - m I CA, 0 coth(m I L) ] Sild'l( m I L)
Case 2: the leading element of the pore ( n - l ) is inactive -r A = -D|n
R 2[
CA'! " CA'°-] L
(10)
Active Surface Areas in the Pores From Figure 2, the surface areas of the pore support and zeolite can be represented by :
R(t)
A°s =
(11)
R(O)
Totally Coked but Accessible Micropo(e PaPally Coked but Accessible Micropor,e Partially Coked but Inaccessible Micropqre / Coke
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Zeolite Micropores: Initial Pore Radius = r(O) Average Pore Radius at time, t = r(t) Figure 2. Schematic o f c o l d a g in s u p p o r t and zeolite.
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2942
O.H.J. MUHAMMADand E. K. T. KAM
in the pore support, and St
=
S.o p N ~Y. YAt Rij(0)C Vg i=lj=l
(12)
in the zeolite respectively.
Rate of Coking in a Pore Element The coking reaction is assumed first order with respect to coking precursors and support active areas, then dR - d'~ = ( ~-op kop CA + Xc, k a Cp ) Aos
(13)
Application of Coking Rate to MMA T Reactor Simulation The behaviour of the modified micro activity test reactor is assumed to be similar to that of a CSTR or hackmixed reactor, then the overall coke contents, Vc(t), can be expressed as, vo(t) = vdt ) + Vm(t)
where
vdt)
p
-
v,(t)
(14)
vg
(15)
E Y:g R~(0)L i=lj=l p N Vg i~lj~l M'lj Sij L d z
and
Vm(t) =
pN
(16)
Z E x R~j(0)L i=lj=l
The loss in zeolite volume due to coking is, p N
E E v d t ) vg vz(t) = i=Ij=l
(17)
p N E E ~ R~(0) L i=lj=l
EXPERIMENTAL
INVESTIGATION,
RESULTS
AND DISCUSSIONS
Experimental work on a commercial zenlitic catalyst, Super-D, to disproportionate cumene was investigated in a modified micro activity testing (MMAT) unit. Since the deactivation is due to the coke laydown on the support macropores and the zeolite micropores, the changes in catalyst structure have been monitored through surface area measurement and scanning electronic microscope image analysis (Muhammad, 1995). The analysis shows the effect of catalyst to feed ratio on the feed conversion (Fig. 3), catalyst surface areas (Fig. 4) and catalyst coke content (Fig. 5). The influences to all the three variables were substantial initially and then greatly reduced thereafter.
"
" Symkotcat/F*~RiUo -s,oo
,~.
! ....
! ...... - s -
NO
,,.
iiiiiiiiii i/ii
l [
0
zo
40 6o so too tzq Time On Strum (mill.) ]
Figure 3. Effect ofcahdyst/feed ratio on conversion with time.
l [
0
20
4o 60 so 100 Time On Strum (mln.)
Flgu~ E ' " ~• - ~ c ~ y e t / ] h ~ l ratio on surface area with time.
1oo
I
I
12q ]
| [
[
o.¢g~, . . 4 . . . . ~ . ~ . , . . . . ~. . . . 0 20 40 6o so ,00 lzc
Time On Stream (mlu.) Figure ~ Effect of catalyst/feed ratio on coke content (%) with time.
Coking in a structured catalyst
2943
The overall coke laydown process simulated from the developed model is shown in Figures 6a to 6d. Initially, the catalyst pore elements are free from coke deposits and they are accessible pending on the connectivity in the pore networks, Figure 6a. As coking proceeds, the thickness of the coke increases from zero to 50(0 which is equivalent to 4.5% of coke content, Figure 6b. When the thickness of the coke layer reaches 1000A, the coke content becomes 6.8%, Figure 6c. From the figure, some totally blocked micropores and a number of partially closed macropores are clearly indicated. The coking rate reduces drastically as the coke layer gets thicker but this does not increase the coke content very much. When the catalyst is fully deactivated, the coke content reaches just over 8% as illustrated in Figure 6d which shows all the macro- and micro- pores are totally blocked. Some uncontaminated pore elements still exist, but they are inaccessible. This observation is similar to the experimental results from the disproportionation of cumene over the zeolite catalyst and catalyst deactivation by coke laydown in the MMAT reactor as shown in Figures 3 to 5. Clearly, at high conversion levels, the coking rates indicated by the slopes of the coke content with time-on-stream were substantial. Whereas at low conversions, reversed trends were observed. These are typical characteristics in series fouling (Viner and Wojciechowski, 1984; Mandani and Kam, 1995). The corrugated parallel bundle pore model is then applied to simulate the coke content in a micro-scale fluidized-bed reactor in which the series fouling mechanism is assumed. After undertaking several simulation experiments to determine the best fitted values of coking reaction parameters, the optimized values were obtained and listed in Table 3 for catalyst deactivation due to active site poisoning and pore plugging, A comparison of the coke content with time-on-stream from the experimental data and simulated results from the model is given in Figure 7. Both the simulated results from active site poisoning and pore plugging are compared well with the experimental data initially. However, as the process continues, the pore plugging deactivation matches better than the active site poisoning. This is consistent with the overall coke laydown visual display in Figures 6a to d.
g Fresh, uncokedpor~
Catalyst with 6.8% coke deposition.
b. Catalyst with 4.5% coke deposition,
d Catalyst with 8% coke deposition.
F i g u r e 6. Different coke levels in the pores u n d e r a parallel coifing mechanism.
8O T a b l e 3. F i t t e d v a l u e s o f d e a c t i v a t i o n parameters from experimental data. Parameters
Active Site Poisoning
Pore length, [micron]
14
Zeolite coke unit size. [A]
1.2
Support coke unit size, [A]
23
Main reaction rate. [m/s]
6x10 s
Initial Zeolite fractional activity Series coking rate constant, [m ~/kmol/s] o P.S.D. (uniform distribution), [A]
0.86 1.18x l(J ~ 60 - 3200
Symbols
i •
Pore Plugging
Pa~c~ar~
....
Exper~eeud D~ ActiveSitePoisoningSimula~on
40
60
14 1.13 23
Conversiol
[%1
6x10 z 0.86 1.32x 10 ~ 60 - 3200
2oi 20
80
100
Time-on-Stream (rain) Figure 7. Conversion comparison.
!
2944
O.H.J. MUHAMMADand E. K. 1".KAM
CONCLUSIONS The experimental results have shown a rapid initial drop in the conversion of the feed, rapid coke build up within the catalyst substrate macropores and zeolite micropores, and this is followed by a period of much slower rate of deactivation. These are the classical series fouling characteristics. Similar observations are also obtained from model simulation in terms of coke laydown. The coke laydown by pore plugging process is compared better than that from the active site poisoning. Hence it becomes clear that the products from the disproportionation of cumene over the Super-D zeolite catalyst, caused the deactivation through a pore plugging mechanism. NOTATION A os
fraction of active surface area in x pore eletmmt
R
C(x)
point concentration in a pore element
-r
C &l
~lrafion
S
between pore elentents n m d n- 1
radius of pore element n A
s
faction rate of teactmt A
sp~if~ ~tiv¢ sm'fa~~ of zooli~
Cp
meam woduct conceatratimt in s pore eleme~tl
S
total active surface area hi a pore element n
D
diffusion coefficient in a pore elemem n
t
time-on-stremn
d
sizc of coke trait in ~colitc mioroporc
vo(t) specificcote contentof the catalyst
k
pmdlclcok~ n~ comamt
V
specific support volume
k~
series coking n ~ constant
V~(t)
specific coke c~ntm~ of J
k
surface rcsctioa r l ~ comamt
v(t)
volumeof cokein a zcolit~mi~opo~
micropore
L
length of e pore clengnt
V(t)
specific coke c o n ~ t of J support pot©
M
mean depth of coke units in a pore eknmnt
v(t)
volume of coke in a pore element
M'
mean depth of coke mits in zcofite in a pore element
m ~
reaction modulus for pore element n
V(t) v (t)
specificvolume of zeoliteloss volume of zeolitelossin a pore el~mcm
N
total number of elements in a pore
x
lensth coordimte of a pore element
n
nlh pore clement
P
total manber of pores in s catalyst particle
R(0)
initial radius of e pore elemem
R(t)
average radius of a poge element at time t
Gnmk Symbols re~fion switch fimc~m, (0 or I )
REFERENCES Androutsopoulos, G.P. and Mann, R., 1978, Chem. Eng. Sci., 33, 673. Burganos, V.N. and Sotirchos, S.V., 1989, Chem. Eng. Sci., 44, 2629. Cmsswell, D.L., 1985, Applied Catalysis, 15, 103. Hollewand, M.P., C~addcm, L.F. and Kenny, C.N., 1991, I.Chem.E. Research Event, Cambridge, England, p.999. Mandani, F. Kam, E.K.T. and Hughes, R~, 1995, 2nd Int. Conf. on Catalysts in Petroleum & Petrochemical Industries, Kuwait. McCacavy, C., Kam, E.K.T., Andradc Jr., LS. and Rajagopal, IC, 1991, The 1st. Anglo4apancsc Sym. on Catalysis, Tokyo, Japan. McGrcavy, C., Draper, L. and Kam, E.K.T., 1994 Chem. Eng. Sci., 49, 5413. Mo, W.T. and Wei, J., 1986, Chem. Eng. Sci., 41,703. Muhammad, O.HJ., 1992, Ph. D. Thesis, University of Manchester, UMIST. Muhammad, O.H3., 1995, 2nd International Conference on Catalysts in Petroleum & Petrochemical Industries, Kuwait. Nicholson, D. and Pctropoulos, J.I-L, 1968, J. Phys. d., I, 1379. Vincr, M. and Wojcicchowski, B., 1984, Caned. J Chem. Eng., 62, 870.
ChemicalEngineeringScience, Vol. 51, No. 11, pp. 2939-2944, 1996 Copyright © 1996 Elsevier Science Ltd Printed in Great Britain. All righut reserved 0009-2509/96 $15.00 + 0.00
S0009-2509(96)00178-9 Modelling and Experimental Studies of Coldng in a Structured Catalyst O. H. J. Muhammad and E. K. T. Kam Petroleum Tecl'mology Department~ K u w a i t Institute for Scientific Research, P.O.Box 24885, Safat 13109, K U W A I T .
Abstract - A stochastic model has been developed to account for the catalyst decay by coking based on the
interaction of the geometries of the coke deposit and the pore structure. Coke is assumed to deposit randomly on the substrate macropores as well as in zeolite micropores resulting in both the active site coverage and pore blockage. The substrate pore structure is represented by a corrugated parallel bundle pore model. The mass balance equations are solved analytically to simulate conversion, coke content and surface area as a function of time-on-stream. Three different fouling mechanisms: parallel, series and triangular, are also investigated and from the simulated results, the series fouling mechanism is tbund to fit the experimental observations in a modified micro activity testing (MMAT) unit better. INTRODUCTION
The performance of catalysts is strongly influenced by diffusion of reactants and products in the catalyst substrate so the overall performance is influenced by the interaction between the intrinsic kinetics and the transport processes such as heat, mass and momentum transfer in catalyst supports (Cresswell, 1985; McGreavy et al. 1994). This is further complicated by the randomness of the pore networks and active unit cells which are very important to zeolite catalysts since they are mainly for resid processing to maximize the octane-barrel in fluid catalytic cracking (FCC) units. In particular, the rare-earth exchange/stabilization is highly suitable for the FCC's LPG/Gasoline quality requirements. However, the catalyst effectiveness also depends on the unit cell range and the matrix of the catalyst subswates, the feedstock characteristics and the accompanied deactivation contaminants such as vanadium, nickel, coke and sulphur. Most of these organo-fouling compounds found in petroleum feedstocks exist in various sizes and shapes but they are predominantly in relatively large molecular structures. Ideally, the catalyst pores should exert minimal diffusion resistance to enable the full utilization of all the catalyst active area but without compromising the catalyst strength. New types of catalysts are developed with specifically designed pore structures customized to the desired selectivities. This requires the optimization of the pore diameters, catalyst connectivity and access to closeend pores and hence, results numerous studies and some network models have been developed (Nicholson and Petropouios, 1968; Androutsopoulos and Mann, 1978; Mo and Wei, 1986; Burganos and Sotirchos. 1989, Hollewand et al., 1991; McGreavy et al. 1991 and Muhammad, 1992) to get insights in the diffusion, sorption and/or reaction, and deactivation in catalysts with random networks of pores. While such works have demonstrated the close agreement of theory with experiments, it has required considerable computational effort and the parameters used are confmed to catalysts but not the reactors. The present work is to develop a mathematical model to interpret experimental data and special attention has been directed to examine the catalyst deactivation due to coke deposition on the catalyst support. The application of the model to correlate the experimental results has shown that the timewise deactivation could be represented by either a substrate active site poisoning coking or a pore plugging mechanism due to one of the products. MODEL DEVELOPMENT
The zeolite contained within crystaUites is assumed to be uniformly distributed along the walls of silica alumina support pores. The diffusion, reaction and deactivation occur in both the support and zeolite structures as functions of time on stream. The deactivation is in the form of coke laydown in either structure through two possible mechanisms - active site poisoning and pore plugging. Other assumptions are: 1. mass transfer is by diffusion only without convection. 2. Dalton's Law of partial pressure applies, 3. the total pressure is constant and hence the molar density and total number of moles are constant 4. first order reaction is applied to the reactant as well as the active surface areas, 5. both the support and zeolite are catalytically active, and 6. the support and zeolite can be deactivated in different ways. Consider a single pore consisting of N elements in which the reaction is assumed as: A ~ Product. A mass balance for the reactant A in any pore element n, n < N, see in Figure 1, can be expressed as: n R2n D n d2CA
= ks Sn CA
(1)
For the pore elements n < N, the boundary conditions are: CA = CA, n-1
at x = 0
CA = CA.n
atx = L
(2) 2939
O.H.J. MUtlAMMAD and E. K. T. KAM
2940
ElementNumbe~ ~
~
Zeo4~ Mzroperes
SuppoctPore
._ ~(T'~_ _ .......
0
....
0
.............
t'-
t~l [Z/:l
I I
I I
I
CA!
f..-
Close end of
......
_ J •
L~ ~-~
)
I I )
CA0
.............
) I I
I I
CAn-2
CA2
I )
CAn-I
CAn
I I
CAN-2
CAN Intermediate Co~ent'e~ons
CA,N.t
l
I I
dCAN I
x~ F3gure 1. A pore schematic
x=L imviag N
elements.
and the analytical solution of Equation (1) becomes CA(X ) =
CA, n.I sinh[mn (L- x)] + CA, n sirdl(mn x) sinh(m n
(3)
L)
For the close-end pore element, n = N, the boundary conditions are as follows: CA dCA dx
= CA,N. I at X = O = 0
(4)
at x = L
and Equation (3) becomes,
CA(X ) =
CA, N. l cosh[m N (L- x)] cosh(m
n
(5)
L)
Concentration Profiles in the Intermediate and End Pore Elements To simulate the concentration profile in the complete pore by coupling all the intermediate concentrations together, the flux between two adjacent elements must be equal, i.e., Dn. | Rn2.1 d C A
= D n R2n dCA
dx x=L
(6)
dx x=O
Since catalysts are undergoing deactivation, any pore element may be either active or inactive at any specific time. Based on the conditions of the two adjacent pore elements, four different cases can exist as listed in Tables 1 and 2 for the close-end pore and intermediate elements respectively. The concentration expressions for each case are also shown. Table 1. The four activity cases for two adjacent pore elements with one end closed Case Number 0
•
O
Element Number Status (N-l)
Active
(N)
Active
(N-l)
Active
(N)
Inactive
(N-l) (N) (N-I) (N)
Mass Balance Expressions ^ . DN.!R2.l m~-I ] _ UA,N-2[ sinh(m N L)
(7a)
CA,N-I [DN.I R2.1 ruN-| coth(mN.l L) + D N R 2 mNtaah(m N L)] = 0 CA,N ,~_ CA,N_I cosh(mN.l L) = 0
(Tb)
Inactive Active
,,', t DN'I R ~ - I , ,,', , DN'I R~.I ~'A,N-2t""-'~J - ~A,N-lt L + DN R ~ m N utah ( m N L)] = 0
(7C)
Inactive Inactive
CA,N_2 - CA,N_I = 0
(7d)
2941
Coking in a structured catalyst Table 2. The four activity, cases for two adjacent pore elements Case Number
Element Number Status
0
(n-l)
Mass Balance Expressions r, r Dn'I R2n'i mn'l Dn R2n m n ~A,n-2t sinh(mn, t L) ] + [ s]i n h-( m n- L )
Active
(n)
Active
(n-l)
A~ive
Dn. 1 R2.1 ran. 1
0
O
On Rn2 m n
CA,n-2[sinh(mn.1L) ] + C A ' n [ ~
(n)
Inactive
(n-l)
Inactive
'
L
•
CA,hill
Active
(n-l)
Inactive
(n)
Inactive
(Sb)
r~ u2 D R2 CAn_2[~n-I AXn.I ] _ C A n _ l [ n-I n-i + D n R 2 n m n c o t h ( m n L ) ] + Dn R2 mn
(n)
]-
CA,n_l[Dn.l R2 1mn.lcoth(mn_lL) + D~-R-~2n ]=0
C
O
(Sa)
CA.,-I [Dn-1 R2n.l mn-1 eoth(mn.l L) + D n R 2 m n coth(m n L)] = 0
L
(8C)
=0
Dn 1 R2 Dn 1 R2 • Dn R2n t - n-~ ~ C r - n-~ + ]+ A,n-2 t - - - - - - ~ ' ~ l - A,n- 1t - - - - " - i ~ - L 2
8d)
."~A,nt,D.LR.~l -_ .
Rate o f Reaction in a Pore Element Due to the active status of the leading (first) element in a pore, two cases can occur: Case 1: the leading element of the pore ( n - l ) is active - r A = - D 1 z R~ [
(9)
m] CA, 1 - m I CA, 0 coth(m I L) ] Sild'l( m I L)
Case 2: the leading element of the pore ( n - l ) is inactive -r A = -D|n
R 2[
CA'! " CA'°-] L
(10)
Active Surface Areas in the Pores From Figure 2, the surface areas of the pore support and zeolite can be represented by :
R(t)
A°s =
(11)
R(O)
Totally Coked but Accessible Micropo(e PaPally Coked but Accessible Micropor,e Partially Coked but Inaccessible Micropqre / Coke
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Zeolite Micropores: Initial Pore Radius = r(O) Average Pore Radius at time, t = r(t) Figure 2. Schematic o f c o l d a g in s u p p o r t and zeolite.
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2942
O.H.J. MUHAMMADand E. K. T. KAM
in the pore support, and St
=
S.o p N ~Y. YAt Rij(0)C Vg i=lj=l
(12)
in the zeolite respectively.
Rate of Coking in a Pore Element The coking reaction is assumed first order with respect to coking precursors and support active areas, then dR - d'~ = ( ~-op kop CA + Xc, k a Cp ) Aos
(13)
Application of Coking Rate to MMA T Reactor Simulation The behaviour of the modified micro activity test reactor is assumed to be similar to that of a CSTR or hackmixed reactor, then the overall coke contents, Vc(t), can be expressed as, vo(t) = vdt ) + Vm(t)
where
vdt)
p
-
v,(t)
(14)
vg
(15)
E Y:g R~(0)L i=lj=l p N Vg i~lj~l M'lj Sij L d z
and
Vm(t) =
pN
(16)
Z E x R~j(0)L i=lj=l
The loss in zeolite volume due to coking is, p N
E E v d t ) vg vz(t) = i=Ij=l
(17)
p N E E ~ R~(0) L i=lj=l
EXPERIMENTAL
INVESTIGATION,
RESULTS
AND DISCUSSIONS
Experimental work on a commercial zenlitic catalyst, Super-D, to disproportionate cumene was investigated in a modified micro activity testing (MMAT) unit. Since the deactivation is due to the coke laydown on the support macropores and the zeolite micropores, the changes in catalyst structure have been monitored through surface area measurement and scanning electronic microscope image analysis (Muhammad, 1995). The analysis shows the effect of catalyst to feed ratio on the feed conversion (Fig. 3), catalyst surface areas (Fig. 4) and catalyst coke content (Fig. 5). The influences to all the three variables were substantial initially and then greatly reduced thereafter.
"
" Symkotcat/F*~RiUo -s,oo
,~.
! ....
! ...... - s -
NO
,,.
iiiiiiiiii i/ii
l [
0
zo
40 6o so too tzq Time On Strum (mill.) ]
Figure 3. Effect ofcahdyst/feed ratio on conversion with time.
l [
0
20
4o 60 so 100 Time On Strum (mln.)
Flgu~ E ' " ~• - ~ c ~ y e t / ] h ~ l ratio on surface area with time.
1oo
I
I
12q ]
| [
[
o.¢g~, . . 4 . . . . ~ . ~ . , . . . . ~. . . . 0 20 40 6o so ,00 lzc
Time On Stream (mlu.) Figure ~ Effect of catalyst/feed ratio on coke content (%) with time.
Coking in a structured catalyst
2943
The overall coke laydown process simulated from the developed model is shown in Figures 6a to 6d. Initially, the catalyst pore elements are free from coke deposits and they are accessible pending on the connectivity in the pore networks, Figure 6a. As coking proceeds, the thickness of the coke increases from zero to 50(0 which is equivalent to 4.5% of coke content, Figure 6b. When the thickness of the coke layer reaches 1000A, the coke content becomes 6.8%, Figure 6c. From the figure, some totally blocked micropores and a number of partially closed macropores are clearly indicated. The coking rate reduces drastically as the coke layer gets thicker but this does not increase the coke content very much. When the catalyst is fully deactivated, the coke content reaches just over 8% as illustrated in Figure 6d which shows all the macro- and micro- pores are totally blocked. Some uncontaminated pore elements still exist, but they are inaccessible. This observation is similar to the experimental results from the disproportionation of cumene over the zeolite catalyst and catalyst deactivation by coke laydown in the MMAT reactor as shown in Figures 3 to 5. Clearly, at high conversion levels, the coking rates indicated by the slopes of the coke content with time-on-stream were substantial. Whereas at low conversions, reversed trends were observed. These are typical characteristics in series fouling (Viner and Wojciechowski, 1984; Mandani and Kam, 1995). The corrugated parallel bundle pore model is then applied to simulate the coke content in a micro-scale fluidized-bed reactor in which the series fouling mechanism is assumed. After undertaking several simulation experiments to determine the best fitted values of coking reaction parameters, the optimized values were obtained and listed in Table 3 for catalyst deactivation due to active site poisoning and pore plugging, A comparison of the coke content with time-on-stream from the experimental data and simulated results from the model is given in Figure 7. Both the simulated results from active site poisoning and pore plugging are compared well with the experimental data initially. However, as the process continues, the pore plugging deactivation matches better than the active site poisoning. This is consistent with the overall coke laydown visual display in Figures 6a to d.
g Fresh, uncokedpor~
Catalyst with 6.8% coke deposition.
b. Catalyst with 4.5% coke deposition,
d Catalyst with 8% coke deposition.
F i g u r e 6. Different coke levels in the pores u n d e r a parallel coifing mechanism.
8O T a b l e 3. F i t t e d v a l u e s o f d e a c t i v a t i o n parameters from experimental data. Parameters
Active Site Poisoning
Pore length, [micron]
14
Zeolite coke unit size. [A]
1.2
Support coke unit size, [A]
23
Main reaction rate. [m/s]
6x10 s
Initial Zeolite fractional activity Series coking rate constant, [m ~/kmol/s] o P.S.D. (uniform distribution), [A]
0.86 1.18x l(J ~ 60 - 3200
Symbols
i •
Pore Plugging
Pa~c~ar~
....
Exper~eeud D~ ActiveSitePoisoningSimula~on
40
60
14 1.13 23
Conversiol
[%1
6x10 z 0.86 1.32x 10 ~ 60 - 3200
2oi 20
80
100
Time-on-Stream (rain) Figure 7. Conversion comparison.
!
2944
O.H.J. MUHAMMADand E. K. 1".KAM
CONCLUSIONS The experimental results have shown a rapid initial drop in the conversion of the feed, rapid coke build up within the catalyst substrate macropores and zeolite micropores, and this is followed by a period of much slower rate of deactivation. These are the classical series fouling characteristics. Similar observations are also obtained from model simulation in terms of coke laydown. The coke laydown by pore plugging process is compared better than that from the active site poisoning. Hence it becomes clear that the products from the disproportionation of cumene over the Super-D zeolite catalyst, caused the deactivation through a pore plugging mechanism. NOTATION A os
fraction of active surface area in x pore eletmmt
R
C(x)
point concentration in a pore element
-r
C &l
~lrafion
S
between pore elentents n m d n- 1
radius of pore element n A
s
faction rate of teactmt A
sp~if~ ~tiv¢ sm'fa~~ of zooli~
Cp
meam woduct conceatratimt in s pore eleme~tl
S
total active surface area hi a pore element n
D
diffusion coefficient in a pore elemem n
t
time-on-stremn
d
sizc of coke trait in ~colitc mioroporc
vo(t) specificcote contentof the catalyst
k
pmdlclcok~ n~ comamt
V
specific support volume
k~
series coking n ~ constant
V~(t)
specific coke c~ntm~ of J
k
surface rcsctioa r l ~ comamt
v(t)
volumeof cokein a zcolit~mi~opo~
micropore
L
length of e pore clengnt
V(t)
specific coke c o n ~ t of J support pot©
M
mean depth of coke units in a pore eknmnt
v(t)
volume of coke in a pore element
M'
mean depth of coke mits in zcofite in a pore element
m ~
reaction modulus for pore element n
V(t) v (t)
specificvolume of zeoliteloss volume of zeolitelossin a pore el~mcm
N
total number of elements in a pore
x
lensth coordimte of a pore element
n
nlh pore clement
P
total manber of pores in s catalyst particle
R(0)
initial radius of e pore elemem
R(t)
average radius of a poge element at time t
Gnmk Symbols re~fion switch fimc~m, (0 or I )
REFERENCES Androutsopoulos, G.P. and Mann, R., 1978, Chem. Eng. Sci., 33, 673. Burganos, V.N. and Sotirchos, S.V., 1989, Chem. Eng. Sci., 44, 2629. Cmsswell, D.L., 1985, Applied Catalysis, 15, 103. Hollewand, M.P., C~addcm, L.F. and Kenny, C.N., 1991, I.Chem.E. Research Event, Cambridge, England, p.999. Mandani, F. Kam, E.K.T. and Hughes, R~, 1995, 2nd Int. Conf. on Catalysts in Petroleum & Petrochemical Industries, Kuwait. McCacavy, C., Kam, E.K.T., Andradc Jr., LS. and Rajagopal, IC, 1991, The 1st. Anglo4apancsc Sym. on Catalysis, Tokyo, Japan. McGrcavy, C., Draper, L. and Kam, E.K.T., 1994 Chem. Eng. Sci., 49, 5413. Mo, W.T. and Wei, J., 1986, Chem. Eng. Sci., 41,703. Muhammad, O.HJ., 1992, Ph. D. Thesis, University of Manchester, UMIST. Muhammad, O.H3., 1995, 2nd International Conference on Catalysts in Petroleum & Petrochemical Industries, Kuwait. Nicholson, D. and Pctropoulos, J.I-L, 1968, J. Phys. d., I, 1379. Vincr, M. and Wojcicchowski, B., 1984, Caned. J Chem. Eng., 62, 870.