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2652 Ind. Eng. Chem. Res. 1992,31, 2652-2660
Modeling the Catalytic Oxidation of n-Butane to Maleic Anhydride
in a Circulating Fluidized Bed Reactor
Todd S. Pugsley,t Gregory S. Patience,*?* Franco Bermti,* and
Jamal Chaoukif Department of Chemical and Petroleum Engineering,
University of- Calgary, 2500 University Dr. NW, Calgary, Alberta,
Canada T2N lN4, and Departement de Genie Chimique, Ecole
Polytechnique de Montreal, C.P. 6079, Succursale A, Montreal,
Quebec, Canada H3C 3A7
A circulating fluidized bed (CFB) is a gas-solid contactor in
which solids are transported vertically in a riser by a
high-velocity gas stream. CFB reactors are becoming increasingly
attractive for conducting specific catalytic reactions. The
catalytic oxidation of n-butane to maleic anhydride (MAN) has been
simulated for a ao-m-high, 30-cm-diameter CFB to investigate the
reactor per- formance under various operating conditions. The
comprehensive simulation combines a core-annular hydrodynamic model
of the CFB with a fixed bed reaction kinetics model. Reactor
performance is evaluated in terms of n-butane conversions, product
yields, and selectivity. The simulation can easily accommodate
different reaction kinetics and therefore can be used to predict
the performance of a CFB reactor for any catalytic process.
1. Introduction Circulating fluidized bed (CFB) reactors are
typified by
riser superficial gas velocities normally in the range of 2-10
m/s, compared with less than 1 m/s in conventional bubbling beds.
At these gas velocities, solids throughput is very large while
reactor diameter is minimized and gas residence time is short.
Solids carried out of the riser are separated from the gas by
cyclones and are returned to the bottom of the riser by a vertical
standpipe.
The traditional uses of a CFB have been mainly in noncatalytic
proteases, such as the combustion of low-grade fossil fuels.
However, the unique design of a CFB also makes it attractive as a
catalytic reactor. Figure 1 is a schematic of a CFB catalytic
reactor for the partial oxi- dation of n-butane to maleic
anhydride. Solid catalyst particles are injected from the standpipe
and are trans- ported vertically in the high-velocity gas stream
containing n-butane. The reaction occurs along the riser length,
and the catalyst is rapidly deactivated. The solids and gas exit
from the riser via a smooth elbow and are separated in the cyclone.
The product stream containing maleic anhydride and other gases is
sent downstream for purification. The deactivated catalyst
particles are stripped of any carbo- naceous material by an inert
such as nitrogen or steam and sent on to the regenerator section
where contact with air oxidizes the catalyst surface once again.
The regenerated catalyst is gravity fed to the standpipe where it
moves to the base of the riser for reinjection into the riser
reaction zone.
Contractor and Sleight (1987) discussed the inherent advantages
of a CFB in the catalytic partial oxidation of n-butane to maleic
anhydride (MAN). These include separate catalyst reduction and
oxidation zones in the riser and standpipe, respectively, low
catalyst inventory, and uniform temperature throughout the riser
(elimination of hot spots). Since the catalyst may be regenerated
in a separate vessel, oxygen requirements in the feed are minimal,
thus leading to high selectivity to MAN and a more concentrated
product stream. Contractor and Chaouki (1990) reviewed other
potential uses for the CFB as a catalytic reactor.
* To whom correspondence should be addressed. University of
Calgary. t Ecole Polytechnique de Montreal. *Present address: E. I.
du Pont de Nemours & Co., Wilmington,
DE.
The interest in the partial oxidation of n-butane to maleic
anhydride is due to the importance of the reaction on an industrial
scale. For the first 40 years of maleic anhydride production, the
feedstock was benzene (DeMaio, 19801, but due to increasing
environmental concerns over the use of benzene and escalating
benzene costa, almost all US. plants have switched to n-butane.
Annual MAN production in the United States is approaching 19OOOO
metric tons and is increasing by approximately 6% per year to meet
increasing demand (Irving-Monshaw and Kislin, 1989). MAN is used
primarily in the production of unsaturated polyester resins, alkyd
resins, and specialty copolymers. I t is also used in the
manufacture of agri- cultural chemicals, as an additive in
lubricating oils, and as a building block for L-aspartic acid, used
in making NutraSweet. The high demand for MAN combined with the
interest in the circulating fluidized bed technology for its
production makes a comprehensive simulation for predicing reactor
performance a valuable tool for industry.
2. Kinetics 2.1. Reaction Kinetics Rate Equations. Intrinsic
rate
equations and reaction pathways for the partial oxidation of
n-butane to maleic anhydride (MAN) on a vanadium phosphorus oxide
(VPO) catalyst utilizing a packed bed experimental reactor have
been proposed by Centi et al. (1985). These expressions were later
cited in the work by Contractor and Sleight (1987) and Patience and
Chaouki (1990). The complete reaction pathway involves both series
and parallel reaction steps:
(1)
n-C4Hlo - C 0 2 (2) n-C4Hio - MAN - COZ The rate equations
proposed by Centi et al. (1985) are
(3)
r2 = rCOs = kZc,B
r3 = -rMAN = k$MAN(CoY/CB6)
(4)
(5)
where rMAN is the rate of MAN formation from n-butane, is the
rate of C02 formation from n-butane, and -rm
?&e rate of maleic anhydride decomposition to C02 and water.
The kinetic parameters determined by Centi et al. (1985) are
summarized in Table I.
0888-5885/92/2631-2652$03.00/0 0 1992 American Chemical
Society
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Ind. Eng. Chem. Res., Vol. 91, No. 12,1992 2653
proportional to the number of catalytic sites st i l l active.
If q(t) corresponds to the number of deactivated sites at solids
residence time t., and qo to the maximum number of sites for a
totally deactivated catalyst, then the rate of maleic anhydride
(MAN) formation, for example, h o m e s
RISER
/-MALEIC ANHYDRIDE
7 OFT GAS
REGENERATOR
AIR
STANDPIPE
BUTANE
Figure 1. Schematic of a eirculatiag fluidized bed an a
catalytic reactur.
Table I. ginetia Param&- for Itate Quntions of Centi et
340 6.230 9.040 0.966 KB = 2616 mol/L; (1 - @ = 0.2298; 7 =
0.6345; 6 = 1.151. Unlike the resulta of other studies (Wohlfaiut
and
Hoffman, 1980, Sharma et al, 1991), the catalyst employed by
Centi and mworkera was very selective toward maleic anhydride. The
vanadium present in the VPO catalyst is the main oxygen carrier,
and during the reaction with the hydrocarbon in the gas phase,
certain crystalline phases in the solid catalyst are transformed
via a reduction pro- ceea. This is written in terms of the valence
change as V+5 - V+. Contractor (1988) found the average oxidation
state of the catalyst to he about 4.1, with a difference of 0.2 in
the average oxidation state between the oxidized and reduced state
of the catalyst. This indicates that most of the active catalyst
exists in the V4 oxidation state and that the catalyst surface
provides most of the oxygen needed for the reaction. As discussed
by Centi et al. (1988). the kinetic model of Centi et al. (1985) is
indicative of a fresh catalyst poeaessing large amounts of surface
oxygen and readily forming the V+ crystalline phase. Therefore, the
model of Centi et al. (1985) is appropriate to use for the modeling
of the CFB reador because of the continuous reiniection of
reeenerated &e.. fresh) catalvst from the
~
vertical standpiie. 22. Catalyst Deactivation Model. The
reduction of
the VPO catalvst is an examole of oarallel deactivation. where
the rea&& form prdrxta &d, at the same time; deactivate
the catalyst. The rate of reaction then becomes
This parallel deactivation is repreaentsd by Smith (1981) as a
simple fmt-order reaction at the catalyst surfam, and thus the rate
of catalyst deactivation, rd, is writtan as
Integration of eq 7 gives
Subetitution of eq 8 into eq 6 provides the rate of reaction
incorporating catalyst deactivation:
b ( t 3 = MAN exp[-at.l (9) The solids residence time used in
the catalyst deactivation model is calculated on the basis of the
upward solids v e locity in the core region. Aa a result, it is
assumed that the catalyst deactivation in the core and annular
regions is the same. This is a large simplifying assumption because
the solids present in the annular region will have been in the
riser for a longer period of time due to solids recircu- lation
down the riser walL Furthermore, the solids present at any point in
the riser will have a varied history due to solids crossflow
between the core and annular regions. A more rigorous formulation
would incorporate the solids residence time distribution (RTD), and
more work is re- quired in this area to improve the simulation. The
purpose of this simplified deactivation model was to acknowledge
the deactivation, which is known to be very fast, and to
investigate the importance of the deactivation on reactor
performance.
3. Riser Hydrodynamics 3.1. Riser Flow Structure. The flow
pattam of the
solids in the riser of a CFB is extremely complex and is
dependent on many variables. The column diameter, height, and exit
configuration, particle properties, and gas characteristics all
affect the flow structure in the riser. Visual observation in a
riser of circular cross section in- dicates the existence of a lean
suspension of solids in gas flowing upward in the center of the
riser, with a denser downilow of solids next to the wall. Weinsteii
et aL (1986) and Bader et al. (1988) both observed decreasing
radial voidage profdes in their experimental units, with the most
dramatic decrease occurring very close to the wall. The radial
location of this sudden change in voidage indicates the interface
between the dilute upward flowing region and the dense downward
flowing region. The CFB riser is also characterized by a dense,
turbulent portion at ita base where the solids are introduced into
the riser from the standpipe, which becomes leaner as the flow of
solids d e velops and the particles are accelerated to their
steady- state upward velocity in the riser. 3.2.
Berruti-Kalogerakis Model. Bermti and Ka-
lcgerakis (1989) modeled the CFB riser as behg charac- terized
by a ooreannular type of flow stmctwe, with solids moving upward in
the lean core region and downward in the dense annular regior.. The
model assumes that the solids move dmnward ii the dense annular
region at the
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2654 Ind. Eng. Chem. Res., Vol. 31, No. 12,1992
terminal velocity of a single particle, and that the voidage in
the dense annular region is the same as that of a bub- bling bed of
the solids at minimum fluidization. This model requires the solids
circulation rate, the axial su- perficial gas velocity, and the
average axial riser voidage profile as inputs.
The average axial voidage profile can be obtained from an
experimentally determined pressure drop profile ac- cording to
1 d P eaVg = 1 - - -
P& dx Equation 10 assumes that the axial pressure drop
profile is due solely to the weight of solids present at any axial
location, ignoring frictional effects at the wall. Arena et al.
(1986) compared the average axial voidage profile predicted from
pressure drop measurement with the profile obtained from trapped
solids weighing along the riser length. Pressure drop measurements
were found to un- derpredict the average voidage in the riser, and
this dis- crepancy was attributed to the frictional effects of the
wall. They indicated that the small diameter column used (4.1-cm
i.d.) may have enhanced the frictional effects.
Berruti and Kalogerakis (1989) derived the following expression
for the voidage in the core region of the riser:
A solids balance over a volume element of the riser resulted in
the following expression for the core radius, r,:
Equations 11 and 12 may be solved simultaneously to obtain the
values of core voidage and core radius at any axial position along
the riser. The limitation of the Berruti and Kalogerakis model is
the requirement of an experi- mentally determined axial pressure
drop profile. There- fore, the model can only be applied to
experimental unita for which the pressure drop profile has been
obtained.
3.3. Predictive Model of Wong et al. (1992). Wong et al. (1992)
have recently developed a predictive model for the average axial
riser voidage profile which is used in conjunction with the
Berruti-Kalogerakis model to de- scribe the internal flow structure
of a CFB. Given riser geometry, solids and gas physical properties,
solids circu- lation rate, and superficial gas velocity, the
average axial voidage profile may be calculated along with the core
voidage and core radius, core gas and solids velocity, and solids
interchange coefficienta between the core and an- nular regions,
all functions of axial location along the riser.
This model considers the CFB riser divided into three sections:
an acceleration zone corresponding to the dense region at the base
of the riser, a fully developed flow zone immediately above the
acceleration zone, and a decelera- tion zone for risers equipped
with an abrupt exit geometry. If the riser has a smooth exit to the
cyclone, the fully developed zone is assumed to extend all the way
to the outlet.
3.3.1. Acceleration Zone. In the dense region at the base of the
riser of a CFB, a fraction of the total measured pressure drop
between the entrance and the start of the fully developed flow zone
can be attributed to the accel- eration of the particles to their
steady-state velocity. If
this acceleration component is not separated from the pressure
drop data, erroneous voidages will be predicted in the acceleration
zone. The voidage predicted from pressure drop data if the
acceleration component is not separated is referred to aa the
apparent voidage, ewp The voidage predicted once the acceleration
component is separated is called the actual voidage. The
acceleration zone is modeled assuming a predominantly upward
flowing gas-solids suspension. Wong et al. (1992) derived the
following expression for the aparent solids holdup in the
acceleration zone:
where 1 - e, r=1- - (14) 1 - 0
The first term on the right-hand side of eq 13 representa the
effect of the acceleration of particles on the observed preaeure
drop, while the gecond term on the rightrhand side represents the
effect of the weight of solids.
In order to evaluate the constant r, the apparent voidage at the
bottom of the riser, q,, must be known. Wong and co-workers
proposed an empirical correlation for eb de- veloped from a
regression analysis on a large pool of published experimental data.
For Geldart group A solids the correlation is
-0.0 2 6 2 8 ~ ~ - 0 . 0 7 9 4 R ~ -0.12016 (15) eb = 0.714( '.)
Ps uo P
This value is substituted into eq 13 with x = 0. The solids
velocity gradient in the acceleration zone is assumed to be
proportional to the difference between the steady-state solids
velocity in the fully developed flow zone, U,,, and the average
solids velocity in the acceleration zone, U,:
(16)
Assuming U, is zero at the riser inlet ( x = 0), and taking U,
as 0.99U8, at the end of the acceleration zone (where x = LA, then
the proportionality constant k is determined from the integrated
form of eq 16 with the substituted limits:
k = -h(O.Ol/Lacc) (17) Thus, if the length of the acceleration
zone, Lam, is known, a value of the proportionality constant, k,
may be deter- mined. Wong et al. (1992) used a modified version of
the correlation of Enick et al. (1987) to calculate the length of
the acceleration zone:
(18) 3.36. Developed Flow Zone. Immediately above the
acceleration zone, a fully developed flow zone is assumed where
the solids and gas have been accelerated to their steady-etate
velocity, and the average solids holdup re- mains essentially
constant.
Patience et al. (1993) have found that in the fully de- veloped
flow regions of circulating fluidized beds, the slip factor,
defined as the ratio of the interstitial gas velocity to the solids
velocity, is approximately equal to 2:
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Ind. Eng. Chem. Res., Vol. 31, No. 12, 1992 2655
Knowing that the average solids velocity in the fully de-
veloped zone is given by
(20)
eq 19 may be rearranged to solve for the average voidage in the
fully developed flow zone:
Equations 13 and 21 will fully describe the average axial
voidage profdes in CFB risers equipped with a smooth exit
configuration. Typically, average voidage values in the range of
0.85-0.98 are observed in CFB risers. As shown in Figure 1, the CFB
catalytic reactor for the partial ox- idation of n-butane is
equipped with a smooth exit, and therefore this simulation of the
CFB catalytic reactor does not incorporate the deceleration
zone.
3.3.3. Solids Interchange Coefficients. The model of Wong et al.
(1992) uses the empirical expression of Bolton and Davidson (1988),
based on a turbulent diffusion mechanism, to calculate the solids
interchange coefficient between the core and the annular
region:
1 - 2.0Re-lI8 1 + s / 1 2 K, = 0.1a1/2Uo
The determination of the annulus-to-core solids inter- change
coefficient is based on an analogy with gas-liquid countercurrent
annular flow. It has been suggested (Senior and Brereton, 1990;
Takeuchi and Hirama, 1990) that the diffuse interface between the
core and annular regions of a CFB is comparable to that of a
gas-liquid annular flow reported in the literature. The expression
for the solids annulus-to-core solids interchange coefficient
is
where pmm and Ucom are the combined phase core density and the
combined phase core velocity, respectively, and defined by Wong et
al. (1992). The quantity fi, is the smooth pipe friction factor.
The proportionality constant, 4, is determined for the case of a
negligible solids wall layer thickness such that r, = R and the net
solids interchange is mro. The detailed procedure for calculating 4
is outlined by Senior and Brereton (1990).
The simulation presented here predicts core-to-annulus solids
interchange coefficients in the range of 1-3 m/s, while the
predicted annulus-to-core solids interchange coefficients vary
between 0.2 and 1 m/s. The largest values of the solids interchange
coefficients occur in the dense, turbulent bed at the riser base
and then decrease as the fully developed flow zone forms.
3.3.4. Gas Interchange Coefficients. In this work, the Higbie
penetration theory (Higbie, 1935) is used to calculate the gas
interchange coefficient. The theory was originaUy developed to
describe contact between liquid and gas occurring for a short
period of time. Thus a steady- state concentration gradient would
not have time to de- velop as the gas moves into the liquid. This
results in the following expression for the gas mass-transfer
coefficient:
k, = (4Dmt/~t)'/~ (24) Eaending this to the CFB, the crossflow
of gas is assumed to occur in the coreteannulus direction only. The
voidage in eq 24 is the annular voidage, and the diffusivity Dm
is
Table 11. Summary of the Major Assumptions Used for the Computer
Simulation the CFB catalytic reactor operates isothermally chemical
reaction occura in both the core and annular regions catalyst
deactivation i a the same in both core and annular
gas input to the annular region is due solely to crossflow of
gas
the chemical reaction is kinetically controlled rieer diameter =
0.3 m; riser height = 20 m riser is equipped with a smooth exit to
the cyclone operating conditions G, = -800 kg/(m2 8) Uo = 4-6 m/s
CB = 1-50 mol % T = 573 K
D, = 75 pm p, = 1500 kg/m3
Ut = 0.05 m/s
regions at the same riser axial location
from the core
W O catalyst physical properties (Geldart A)
e d = 0.5
that of n-butane in air. The predicted values of the gas
interchange coefficients are of the order of m/s, while data
supplied by G. S. Patience calculated from an ex- perimental gas
RTD indicate values for the core-bannulus gas interchange
coefficient in the range of 0.015-0.094 m/s. Thus there is
agreement between the simulation and the experimental data.
4. Computer Simulation The major assumptions incorporated into
the computer
simulation are outlined in Table 11. The CFB catalytic reactor
is assumed to operate isothermally at T = 573 K. Solution of the
riser energy balance assuming a typical value of the solids-to-wall
heat-transfer coefficient of 150 W/(m2 K) indicates a maximum
temperature variation of 3 "C. Contractor and Sleight (1987) listed
isothermal operation as one of the inherent advantages of the CFB
catalytic reactor for the production of maleic anhydride.
The reaction occurs in both the core and annular regions, the
proportion of which is dependent on the local solid catalyst
concentration as calculated by the hydrodynamic model. The
simulation of the reaction requires solution of the mam balance for
each species, i, reacted or formed. In the core region, where the
gas is flowing, the steady-state mass balance includes a net
convective input term plus the reaction term which is equal to the
gas crossflow to the annulus. The reaction term ri in eqs 25 and
26, is positive for products formed and negative for reactants
consumed. The core mass balance is written as follows (Patience,
1990):
In the annular region, there is no convective input term. The
gas input to the annulus is due to crossflow from the core. The
annular region mass balance is thus
2 k g r c ~ s ( 1 - cann)ri + (Ci ,c - Ci ,a) = 0 (26) (R2 -
r:)
The units of the rate of reaction term, ri, in eqs 25 and 26 are
mol/ (gcat.s). In order to incorporate the reaction rate expression
into the mass balance equations, the pseudo- homogeneous system was
assumed. Such a reactor as- sumes no concentration gradients within
a volume element and is valid for reactions which are kinetically
controlled. The half-life of this reaction waa estimated assuming
that the reaction occurs in a constant-volume batch reactor. This
calculation results in a half-life of approximately 50
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2656 Ind. Eng. Chem. Res., Vol. 31, No. 12, 1992
ms, indicative of an extremely fast reaction which would not
normally be kinetically controlled. However, for a mean particle
diameter of 75 fim, typical of the VPO catalyst used in the
reactor, the Thiele modulus is of the order of and the
effectiveness factor is 1. Thus, in- traparticle resistance is
negligible. Furthermore, the high slip velocities typical of
circulating fluidized bed operation result in excellent mass
transfer between the gas phase and the catalyst surface. Fbh (1986)
indicated, in a qualitative manner, that the slip velocity is at a
maximum in the circulating fluidized bed regime. Therefore, mass
transfer or interparticle resistance will also be negligible. The
negligible intraparticle and interparticle resistances in- dicate
the reaction is kinetically controlled even though it is very fast,
and thus the pseudohomogeneous assump- tion is valid for the
reactor.
In this work, the simulation was run for a 20-m-high riser of
30-cm diameter. The riser is divided into 100 equal elements of 20
cm each. The length of the acceleration zone is calculated, and the
hydrodynamic model is invoked to obtain the average axial voidage
profile. From the Berruti-Kalogerakis model, the core voidage and
core ra- dius are obtained for each element. Forward finite dif-
ferentiation is used to solve the core and annulus mass balances of
eqs 25 and 26. For a given inlet concentration of reactants to the
first volume element, the reaction rates may be calculated from eqs
3-5 using the kinetic data of Table 1. From the stoichiometry of
the chemical reactions, the rates of water formation and oxygen
depletion are calculated. It should be noted that since the
majority of gas flows in the core region, the reactant
concentration in the annular region in the first element is taken
to be 0. This sets up the concentration gradient required for the
mass transfer of gas from the core to the annulus.
Taking the gas velocity in the core (Vo in eq 25) as constant
with height within each of the volume elements, then the axial
concentration gradient of species i in the element dci/dx may be
determined. If the partial deriv- ative is put in discrete form,
and with the inlet concen- trations to the first element known, the
outlet concen- tration of species i from each element may be found.
This concentration becomes the input to the next element where the
mass balances of eqs 25 and 26 are performed again. These results
then become the input to the next element, and the procedure is
continued to the reactor exit,
By keeping track of the molar flows of each species reacted or
formed, the moles of gas generated by the re- action will be known.
This is converted to a volumetric flow of gas via the molar volume
at reaction conditions. In this way, the increase in gas velocity
in the riser due to gas generation may be calculated and the
hydrodynamic model updated at the end of each interval, in order to
evaluate the new flow structure corresponding to increasing values
of the gas superficial velocity.
The n-butane conversion, yields of maleic anhydride and COP, and
maleic anhydride selectivity are calculated at each 20-cm interval.
The conversion, X, is defined as
(27) moles of butane converted
moles of butane fed X =
Maleic anhydride selectivity, S, is defined here as moles of
maleic anhydride formed
moles of butane converted S = (28)
Finally, the product yield of MAN or COz, Y, is defined as
moles of product formed moles of butane fed Y = = SX (29)
5.0
5.0 mol% c g 20.0 mol% cB
3.5 - C, = 800 kg/m s Uoi = 6 m/s 0 m I. 3.0 -
T = 573 K
>
- 2 .5 -
b?
0 2 4 6 8 10 12 14 16 18 20
Riser Height. m
Figure 2. Single pasa butane conversion for various butane in
air mixtures.
5. Simulation Results and Discussion 5.1. Effect of Feed
Concentration. A mixture of
n-butane in air can form a potentially explosive mixture. The
experiments of Wohlfahrt and Hoffman (1980) used a 1 mol % n-butane
concentration while Sharma et al. (1991) limited their n-butane
concentrations to less than 3 mol %. However, since the majority of
the oxygen re- quired for the reaction to proceed is supplied by
the cat- alyst surface oxygen, and because catalyst regeneration
occurs in a separate section, the CFB reactor may be operated in
the absence of, or with small amounts of, gas-phase oxygen. As
indicated by Contractor et al. (1988), this allows for operation
with higher n-butane feed con- centrations. The simulation was
performed for feed con- centrations of up to 50 mol % n-butane in
air, and as indicated in Figure 2, conversion decreased with
increasing n-butane feed concentration. Figure 2 also indicates
that a limit is reached whereby an increase in the n-butane feed
concentration has only a marginal effect on the conversion.
Repeated runs of the simulation have indicated that for n-butane
feed concentrations greater than 25 mol % , the riser conversion
profile is essentially unchanged.
Figure 3 presents the rate of n-butane consumption to form both
maleic anhydride and COP Although the n- butane concentration
increases by a factor of 20, the re- action rate for 20 mol %
n-butane is only 1.5 times the reaction rate at 1 mol % . The
expected increase in reac- tion rate at higher n-butane feed
concentrations is lessened by the inhibition term found in the
denominator of eq 3. Conversion is defined in eq 27 as the moles of
n-butane converted per mole of n-butane fed. For an increase in
n-butane feed concentration, the rate of reaction does not increase
proportionally, and thus conversion is decreased.
With the competing series and parallel reactions, the conversion
includes formation of both desired and unde- sired products.
Figures 4 and 5 compare the yields of maleic anhydride and COP
along the riser length for two different &/butane mixtures.
Figure 4 indicates that for 1 mol % n-butane in the feed, the yield
of C02 is actually greater than the yield of maleic anhydride. When
the feed concentration is increased to 5 mol % as shown in Figure
5, the yields of both maleic anhydride and COa decrease, but the
yield of the desired product, maleic anhydride, is now greater than
the yield of COO. This is obviously a more
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Ind. Eng. Chem. Res., Vol. 31, No. 12, 1992 2657
: 0.020 m
0.010
0.100 I 1.0 mol% cB
G. = 800 kg/m2s UOi = 6 m / s
0.080 T = 573 K 20'o molX c B
e.
nv) 0.070 E L E 0.060 v
6 0.050 n 5 0.040 .- - YI c
0.030 /-
G, = 800 kg/m2s Uoi = 6 m/s
cg = 1.0 mol%
2.4 T = 573 K
0.0 I I 1 I I I I I I I 0 2 4 6 8 10 12 14 16 18 20
Riser Height. m
Figure 4. Single pass product yields for 1 mol % butane in air
mixture.
profitable situation. These results indicate that although
higher conversions are attained at relatively lower feed
concentrations of n-butane, such an operating condition is not
advantageous as the C02 yield exceeds that of the maleic anhydride.
The ability to operate a CFB reactor in the absence of gas-phase
oxygen now becomes impor- tant, as the resulting higher n-butane
concentrations will result in improved yields of maleic
anhydride.
Figure 6 shows the selectivity ,to maleic anhydride (MAN) along
the riser defined in this work as the moles of MAN produced per
mole of n-butane converted. Lower concentrations of n-butane in the
feed result in lower se- lectivity. At concentrations of 20 mol %
or greater, se- lectivities of 80% are realized. Experimental
graphical data at 360 "C presented by Contractor et al. (1988) in-
dicate selectivites of 75-80% for n-butane concentrations in the
range of 12-50%. Thus the results of this work at 300 OC agree
favorably with that data. The initial drop in selectivity seen a t
the lower concentrations in Figure 6 may be attributed to the
relatively greater initial oxygen
GI = 800 kg/m2 s Uoi = 6 m/s cg = 5.0 mol%
1.2 T = 573 K 1.4 t
; 0.8 c MAN Yield Y
.-
0.6
0.4
0.2 /
0.0 I I I I 4 1 I I I 0 2 4 6 8 10 12 14 16 18 20
Riser Height. m
Figure 5. Single pass product yields for 5 mol % butane in air
mixture.
& 100.0
90.0 ......... ...................... .. ............... ... ..
.................. .. ... ... ............. ,...., ,. _ _ _ _ _ _ _
_ _ _ _ _ _ _ _ _ _ _ _ _ _ _ _ _ _ _ _ . _ _ _ _ _ _ _ _ _ _ _ _ -
- -
_ _ - - - - ............. .........
80.0 .___.__- -- - - - - - - -
70.0 ' x - - - - _ - - _ _ _ _ _ _ _ _ _ _ - - - - - . i
2 bp 30.0
10.0
- 1.0 mol% cB - - - 5.0 mol% cB
20.0 mol% cg 50.0 mol% cg
.__--
..... ... .
G, = 800 kg/m2s Uo, = 6 m/s T = 573 K
n o 1 -.- 0 2 4 6 8 10 12 14 16 18 20
Riser Height, m
Figure 6. Maleic anhydride selectivity for various butane in air
mixtures.
concentration in the feed. For a 1 mol % n-butane con-
centration in an air/butane mixture, the corresponding oxygen
concentration is 20.8 mol %. Therefore, the initial reaction rate
for COz formation is favored, and a drop in selectivity is
observed. For 20 mol % n-butane in the feed, the initial oxygen
concentration is 16.8 mol % and a flatter selectivity profile is
seen. Once again, the advantage of operating the CFB reactor in the
absence of gas-phase oxygen and increased n-butane concentrations
is reinforced as the higher feed concentrations result in
significantly improved selectivities. These selectivity results
agree with the previous results of the product yields. Low yields
at low n-butane concentrations correspond to low maleic anhydride
selectivities.
The yield is defined in eq 29 as the product of the se-
lectivity and the conversion. It has been shown in Figures 2 and 6
that, at high n-butane feed concentrations, con- version and
selectivity profiles are constant. Therefore, the yield will also
be constant. It was previously mentioned that approximately 25 mol
% n-butane in the feed is the
-
2658 Ind. Eng. Chem. Res., Vol. 31, No. 12, 1992
0 . . .... ... .- r 3.0
0.40 1 I W 0.30 k$. I 0.20 \ '. \*..
0.00 1 ' I I I I I I I I 1 0 2 4 6 8 10 12 14 16 18 20
Riser Height, m
Figure 7. Effect of solids circulation rate on butane conversion
and solids holdup.
critical concentration. Operation above this concentration will
not significantly improve product yields or selectivity. Though the
conversions and yields are very low, throughput from CFB reactors
is very large and maleic anhydride production can be significantly
higher than that from a fixed bed operation. Furthermore, there is
the potential for enhanced reaction rates due to the cyclic
operation between the riser and the regenerator (Matros, 1989; Lang
et al., 1989). The cyclic operation has not been included in the
simulation at this time due to lack of re- action kinetics for the
process. Contractor et al. (1988) state that it is possible to
improve this conversion by ad- dition of small amounts of gas phase
oxygen in the riser. The potential for secondary air injection
along the CFB riser would be valuable for this purpose.
5.2. Effect of CFB Operating Conditions on Reactor Performance.
The use of the predictive hydrodynamic model in this simulation
allows the effect of operating parameters such as solids
circulation rate and superficial gas velocity to be carefully
studied. These variables will affect the local solids catalyst
holdup in the riser and hence the conversions and yields.
The n-butane conversion for solids circulation rates of 400,
600, and 800 kg/(m2.s) is presented along with the corresponding
actual solids holdup for each circulation rate in Figure 7. The
inlet gas superficial velocity is main- tained at 6 m/s, and
n-butane concentration in the feed is 5 mol % . An increased solids
circulation rate results in a greater solids holdup. This improves
the contact between gas and solids, and therefore a higher
conversion is achieved. It should be observed that, at the inlet of
the riser, solids holdup is increased from about 28% solids at the
low solids circulation rate to roughly 35% solids at the highest
solids circulation rate. This increase in solids holdup is
important because the catalyst will be moat active when it reenters
the riser after regeneration in the standpipe.
In Figure 8, the n-butane conversion along the riser is
presented for inlet superficial gas velocities of 4,5, and 6 m/s
along with the corresponding solids holdups. The solids circulation
rate and n-butane feed concentration are 800 kg/ (m2-s) and 5 mol %
, respectively. The effect of increased superficial gas velocity is
a dilution of the bed and thus a decrease in solids holdup. As a
result, a de-
0.00 1 I I I I I 1 I I I 0 2 4 6 8 10 12 14 16 18 20
Riser Height. m Figure 8. Effect of inlet superficial gas
velocity on butane conver- sion and solids holdup.
50% deactivation 70% deoctivotion
G, = 800 kg/m s Uo, = 6 m/s
I cg = 5.0 mol% 2.0 c C 0 .- E l T = 5 7 3 K
0.0 0 2 4 6 8 10 12 14 16 18 20
Riser Height, m
Figure 9. Effect of catalyst deactivation on single pass butane
conversion.
crease in the conversion of n-butane is seen at higher gas
velocities. Again, the dense bed at the base of the riser exhibita
the greatest solids holdup.
Figures 7 and 8 confirm that it is beneficial to operate at the
highest possible solids loadings and low gas veloc- ities. The high
initial reaction rates and the resulting gas generation in the
dense bed will tend to increase the gas superficial velocity.
5.3. Results of Catalyst Deactivation Model. The preceding
results were obtained assuming a catalyst deactivation
corresponding to a value of 20% at the outlet of the riser. This
section investigates the importance of catalyst deactivation as it
is modeled here. Values of the decay constant a in eq 9 were chosen
such that the catalyat was 20%) 50%, and 70% deactivated upon
exiting to the cyclone. Feed concentration and riser operating
conditions were held constant. As shown in Figure 9, increased
catalyst deactivation reduces conversion along the riser. For
approximately the first 3 m of riser length, the con- version is
unaffected, but beyond this point the conversion
-
Ind. Eng. Chem. Res., Vol. 31, No. 12, 1992 2659
a rapid initial rise in reaction rate in the annular region.
After 8-10 m, the rate of reaction levels off as reactants are
depleted and mass transfer between the core and the annulus is
lessened.
0.050 r 1
G, = 800 kg/m2s Uoi = 6 m/s
cg = 5.0 mol%
T = 573 K
0.045
E 0.040
c
I
0.000 I I I I I I I I I 0 2 4 6 8 10 12 14 16 18 20
Riser Height. m
Figure 10. Contribution to reaction by the core and annular
regions for a 5 mol % butane in air mixture.
profiles begin to separate. At the riser inlet, the catalyst has
just been regenerated and exposure to reaction con- ditions does
not show a noticeable effect. Reaction occurs rapidly at this
level, however, and after 3 m the trans- formation of the catalyst
crystalline phases at the expense of the surface oxygen layers
becomes significant. The higher degree of deactivation, resulting
in reduced con- version, wil l consequently reduce the yields of
both maleic anhydride and COP The results of this simple catalyst
deactivation model indicate that reduction of the catalyst activity
as the reaction proceeds has only a slight effect on the reactor
performance. However, the results will underpredict the effect of
deactivation on conversion due to the assumptions involved in the
deactivation model. Further work to include the solids RTD in the
simulation will produce more realistic results.
5.4. Influence of the Core and Annular Regions. It is
interesting to investigate the separate contributions of the core
and the annulus to the reaction. While the ma- jority of the gas
flows upward in the solids-lean core region, the annular region
will have a higher concentration of catalyst. Furthermore, although
the annulus is very dense, it is also very thin, which will affect
the reaction rate per unit volume of bed.
Figure 10 separates the contribution to the reaction by the core
and the annular regions. The model assumes that, initially, all of
the gas fed to the riser is in the core region. The high initial
concentration of reactants combined with the large solids holdup in
the core at the base of the riser results in a high initial
reaction rate in the core and es- sentially zero reaction rate in
the annulus. The solids holdup in the core quickly decreases
however, as illustrated earlier in Figures 7 and 8, and some of the
upward flowing gas is transferred to the annulus. Also, the rapid
initial reaction rate depletes a large portion of the reactants.
These fadors combine to give the steep decline in reaction rate in
the core seen in Figure 10. The gas concentration gradient between
the core and the annulus is initially very large because all of the
gas is concentrated in the core. Thus, with crossflow of gas
modeled by a mass-transfer mechanism, the annulus gas concentration
rises quickly. This fact, combined with the constant solids holdup
cor- responding to minimum fluidization voidage as assumed in the
development of the hydrodynamic model, leads to
6. Conclusions The hydrodynamics of a circulating fluidized bed
and
a fixed bed intrinsic kinetics model have been combined to form
a comprehensive computer simulation of the partial oxidation of
n-butane to maleic anhydride. The model allows investigation of the
effect of changes in op- erating conditions or reactor geometry,
important in re- actor design and process control. Key results
presented indicate decreased conversion with increased n-butane
feed concentration, but improved selectivity at the higher n-
butane concentrations. For high solids circulation rates,
conversion is increased due to increased solids holdup, and thus
better gasaolids contacting. Increasing the gas su- perficial
velocity decreased conversion due to the dilution of the bed or
decreased solids holdup. Catalyst deacti- vation had a only a
slight effect on reactor performance, however; more work is needed
to make the deactivation model more rigorous.
Acknowledgment
The work reported in this paper has been supported by a Natural
Sciences and Engineering Research Council of Canada operating grant
and scholarship.
Nomenclature a = deactivation parameter (8-l) cB = n-butane
concentration (mol/L) c , = oxygen concentration (mol/L) cMAN =
maleic anhydride concentration (mol/L) q C = concentration of the
ith species in the core (mol/L) ci,, = concentration of the ith
species in the annulus (mol/L) Dm = molecular diffusivity of
n-butane in air (m2/s) D, = mean particle diameter (m) D, = riser
diameter (m) fi0 = smooth pipe interfacial friction factor g =
gravitational acceleration constant (m/s2) G, = overall solids
circulation rate (kg/(m2 8 ) ) k = proportionality constant in the
acceleration zone (m-l) kl = kinetic constant for maleic anhydride
formation (mol'-*
k z = kinetic constant for COz formation (moll-8 LB/(g a)) k3 =
kinetic constant for MAN decomposition (mol" L1+/(g
kd = kinetic constant for catalyst surface deactivation (s-l) k,
= gas mass-transfer coefficient from core to annulus (m/s) K,, =
annulus-to-core solids interchange coefficient (m/s) K,, =
core-to-annulus solids interchange coefficient (m/s) KB =
equilibrium constant in Centi kinetics (mol/L) Lac, = length of the
acceleration zone (m) P = total pressure (Pa) qo = number of
deactivated catalytic sites for a totally
deactivated catalyst q ( t ) = number of deactivated catalytic
sites at any time t rl, r m = rate of maleic anhydride formation
from n-butane
(mol/(g 9)) r2, rCOl = rate of C02 formation from n-butane
(mol/(g a)) r,, -rmN = rate of maleic anhydride decomposition
(mol/(g
8 ) ) r, = rate of the ith reaction, where i = 1, 2, 3 (mol/(g
8)) r, = core radius (m) rd = rate of catalyst deactivation (mol/(g
8 ) ) R = riser radius (m) Re = Reynolds number = U o p p p / p Re,
= particle Reynolds number = fJ$gDp/bg
L"/(g 5 ) )
8))
-
2660 Ind. Eng. Chem. Res., Vol. 31, No. 12, 1992
S = product selectivity, moles of product formed per moles
S = Stokes number = pJl,2Uo/18pgUt t , = solids residence time
(s) T = temperature (K) U,,, = combined phase core velocity (m/s)
Uo = riser superficial gas velocity (m/s) Ua = average solids
velocity (m/s) U,, = core solids velocity (m/s) U,, = solids
velocity at the end of the acceleration zone (m/s) Ut = terminal
settling velocity of a single solids particle (m/s) x = axial
location in the riser (m) X = n-butane conversion, moles converted
per moles fed Y = product yield, moles of product formed per moles
fed Greek Letters a, 8, 6 , = kinetic constants, exponents in Centi
rate ex-
pressions r = constant in acceleration zone tan = annular
voidage fapp = apparent axial voidage in the acceleration zone taVg
= average axial voidage t b = apparent voidage at riser bottom tC =
core voidage emf = voidage at minimum fluidization conditions to =
actual voidage at riser bottom t, = voidage at the end of the
acceleration zone { = constant in K,, expression pb = bulk density
of solids (kg/m3) pcom = combined phase core density (kg/m3) pg =
gas density (kg/m3) pa = solid particle density (kg/m3) cp = slip
factor 4(t) = fraction of the total number of catalytic sites
deacti-
vated Registry No. Butane, 106-97-8; maleic anhydride,
108-31-6;
converted
vanadium phosphorus oxide, 65506-75-4.
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Received for review March 23, 1992 Revised manuscript received
August 10, 1992
Accepted September 4,1992