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Two-phase flow catalytic reactor, influence of hydrodynamics on selectivity
Pergamon
Chemical Engineerin 0 Science, VoL 50, No. 20, pp. 3303-3312, 1995 Copyright ~) 1995 Elsevier Science Ltd Printed in Great Britain. All rights reserved 0009-2509/95 $9.50 + 0.00
0009-2909(95)00141-7
T W O - P H A S E F L O W CATALYTIC REACTOR, I N F L U E N C E O F H Y D R O D Y N A M I C S O N SELECTIVITY t C. VERGEL Instituto Colombiano del Petroleo (ICP), A.A. 4185, Bucaramanga, Colombia J. P. EUZEN, P. TRAMBOUZE: and J. P. WAUQUIER Institut Fran~ais du P6trole, Centre d'Etudes et de D6veloppement Industriels, 69390 Solaize, France
(Received 28 October 1994; manuscript received and accepted in revisedform 10 April 1995) Abstract--In this paper, experimental results are presented, showing the influence of the hydrodynamic behavior of fixed-bed catalytic reactors with two-phase (gas + liquid) flow on the selectivityof a reaction. A mathematical model has been developed in order to represent up-flow data, for which the distribution of phases across the catalytic bed is near ideal. The influenceof maldistribution of the phases is noticeable for down-flow conditions. In this case, recommendations for the design of the reactor are given.
INTRODUCTION
An industrial chemical reaction was used for this study: the hydrogenation of a C4 olefinic cut, to convert selectively the small amount of butadiene present to butene-l, with minimum butane production. The selectivity of this reaction, to be preserved at very high butadiene conversion, appears to be highly sensitive to the flow characteristics. A model, including the complex chemical kinetics of the Langmuir-Hinshelwood type, with two parallel and two consecutive reactions, and the mass transfers between the phases (gas-liquid and liquid-solid), was developed to represent the experimental results. Discrepancies found between some experimental results and the model were interpreted on the basis of phase maldistribution and segregation in the catalyst bed. A more complex model, accounting for these segregation effects and, more precisely, local fluctuations of gas-liquid ratio, serves to interpret all the experimental results. Recommendations are proposed for the design of concurrent gas-liquid fixed-bed catalytic reactors, as these systems are very sensitive to flow regimes. KINETICEXPRESSIONS The aim of the selective hydrogenation of a C4 exsteam cracking cut is to remove any remaining butadiene after this compound has been extracted, with minimum effect on butenes and favoring the production of butene-1.
t Paper presented at ISCRE 13 Meeting in Baltimore, U.S.A., September 1994. : Corresponding author.
An intensive laboratory investigation (Boitiaux et al., 1985a, b) with a suitable catalyst led to the selection of the simplified reaction scheme: shown in Fig. 1.
At
1 , A2 2Xa 3~
4 ,
, A4
A3 Reaction 1: At + A5 --, A2
Reaction 2: AI + As --*A3 Reaction 3:
A2 " * A3
Reaction 4:A2 + As -, A4 where At = 1-3 butadiene A2 = butene-1 A3 = butene-2 A4 = normal butane A5 = hydrogen Fig. 1.
In the range of temperatures and hydrogen partial pressures used, experiments conducted in a batch reactor, with a catalyst powder finely dispersed in the liquid phase, showed that the kinetic equations of the process could be written as follows: dnl
Kt Ct
M d--'-~t---- -- rl -- r2 = -- (kl + k2)K1C1 + K2C2 C5
3303
(1)
3304 dn2 Mdt
-
C. VERGEL et al.
klK1C1 =
r 1 --
r 3 --
r 4
- (k3 + k 4 ) K 2 C 2
=
C 5
K1CI
+ K2C2
Parameters p and m depend on the catalyst, and must therefore be determined experimentally.
(2) REACTOR MODELING d/13
- - = r
Mdt
k 2 K 1 C 1 "4- k 3 K 2 C 2 C5
2 +r 3 =
(3)
KIC1 + K2C2
dn4 k4 K 2 C2 M d t = r4 = K1C1 + K2C2 C5.
(4)
The following constants are introduced for simplification:
R1 = k l K 1 C s ,
R2=k2KICs
~3 = kaK2Cs,
~4 = k4KECs.
The ratio of the two adsorption constants K1 and K 2 w a s found to approach 300. By defining m
and
p= -
~ 3 "~ ~ 4
-
~ 1 -~" ~:~2
it is shown that the combination of eqs (1) and (2), followed by an integration, helps to express the number of moles of butene-1 formed as a function of butadiene conversion X: - - X ) l/m --
n2 - - n 2 o = n 2 o [ ( 1
P
1/m
1 -
1]
nlo I-(1 - X) -- (1 - X)'/'-I
(5)
The simplest model that can be written for the different processes occurring inside the reactor must take account of the following mechanisms: (i) influence of mass transfer between gas phase and liquid phase, (ii) influence of mass transfer between liquid phase and catalyst surface, (iii) occurrence of the various chemical reactions inside the catalyst particles. As a first approximation, the gas and liquid concurrent flows can be considered as of the plug flow type. Moreover, because of the impregnation by the active metal limited to the external portion of the catalyst particles, the effects of intragranular diffusion are negligible. The reaction is assumed to be isothermal. The equations reflecting the molar balances in each phase in steady-state conditions can be written as follows.
Gas phase dNa, i d-~- + JC'aL.i = 0
(7)
The first term is the convective flux due to the gas flow. The second term representing the transfer flux between the gas phase and the liquid phase, can be expressed in terms of the film theory, as follows:
where X =
~P GL, i = ( K a ) L , i (C*,i -- CL, i ).
/110 - - n l /11o
Liquid phase
Since the volume of liquid phase remains virtually constant during hydrogenation, and since the molecular weights of the different C4 compounds are similar, the number of moles ni can be replaced either by the molar concentrations or by the weight percentages. In the case of a fixed bed reactor, assuming a plug flow through the bed and no mass transfer limitations, the material balance for butene-1 gives an expression similar to eq. (5), i,e.
Pl/m glo[(1 - X) -- (1 -- X)I:"].
(6)
This helps to calculate the selectivity S 1,2 of the transformation of butadiene to butene-1, defined as follows: $1,2
n 2 - - /120
-
-
=
/11o - - /11
-
g2 -- g20
-
dNL, i dz
JI/'GL, i -~- J~LS, i = O.
(9)
The first term is the convective flux due to the liquid flow. The third term representing the transfer flux between the liquid phase and the solid phase can also be expressed in terms of the film theory:
Jff Ls.i = (Ka)Ls.i(CL.i -- Cs.i).
(10)
It also equals the quantity of compound A~converted per unit volume of catalyst bed, and can be written:
g 2 - - g 2 o = g 2 o [ ( 1 - - X ) TM -- 1 ]
1-
(8)
with
glo -- gl
gl0 - - g l
X ~ . - 01o
According to eq. (6), the selectivity of the hydrogenation of butadiene to butene-1 is linked to the conversion X and influenced by three parameters: (i) the initial ratio of butene-1 and butadiene concentrations, g2o/glo, (ii) parameter m, (iii) parameter p.
~I/'LS, i "+- pcRs, i = 0
(11)
where R~,i is the reaction rate per unit mass of catalyst, with respect to compound A:
Rs, i = ~, vi.:j.
(12)
The mass transfer coefficients can be estimated from correlations proposed in the literature depending upon the type of flow (Vergel, 1993). After comparing the available correlations, we selected those listed in Table 1 for most of our calculations. EXPERIMENTAL
The pilot plant used to compare the performance obtained in up-flow and down-flow is shown in Fig. 2. Liquid feed (C4 cut), from a container pressurized
3305
Two-phase flow catalytic reactor Table 1 Flow direction
Correlation
Reference
Up-flow
(Ka)L.i (Ka)ts.i (Ka)L.i (Ka)Ls.i
Satterfield (1975) Mochizuki (1982) Charpentier (1976) Rao and Drinkenburg (1985)
Down-flow
i Nz
I
"~
~ ~ ~
FEED STORAGE TANK
_® ....
COOLING
REACTOR
i
STORAGE
Fig. 2. Scheme of the pilot-plant used for performing experiments.
with nitrogen, is mixed with hydrogen and sent to a heat exchanger, which helps to keep the feed temperature of the reactor constant (40°C). Two stainless reactors, with inside diameters of 5.5 and 10.5 cm, respectively, can be mounted on the unit. The catalyst bed heights are accordingly 108 and 156 cm for the small reactor, and 108 cm in the larger reactor. The catalyst volumes used thus range between 2560 and 9400 cm 3. The catalyst used is an industrial catalyst (Pd on alumina), specially designed for high selectivity; it is only impregnated at the bead surface. The fluid flow direction can be up-flow or down-flow, as desired. Inside each reactor, the catalyst is placed between two layers of inert alumina particles 4 to 6 mm in diameter. The catalyst arrangements inside the reactor are shown in Fig. 3. The reaction is carried out under pressure (6.5 bar) and at relatively low temperature (40-45°C), so that most of the C4 cut feedstock remains liquid, with only a small part in the vapor phase with hydrogen. The feed is characterized by: • very high initial content of the desired product, butene-1, • very low initial content of the product to be removed, butadiene.
The influence of feed flow rate and of initial hydrogen to butadiene ratio (H/BD) was investigated in order to vary the butadiene conversion, and hence selectivity, within a wide range (40-99.5%). A number of special features of the process must be pointed out: • The temperature in the reactor is virtually constant, the heat generated by the reaction being compensated by a slight vaporization of the C4 cut and heat losses. • Butadiene conversion is controlled by the quantity of hydrogen injected into the feed. • The liquid phase has rather unusual physical properties because it is near the critical point. Thus the density is about 560 kg/m 3 and the surface tension is about 10- 2 N/m. However, the liquid phase displays no tendency to foaming. The gas and liquid phases entering the reactor can be assumed to be in thermodynamic equilibrium. The corresponding calculations were thus carried out to determine the flow rates and compositions of the two phases at this point. In these calculations, the nitrogen (used for feed pressurization) dissolved in the C4 cut was also taken into account. Note that, even if most of
3306
C. VERGELet al.
I Alumina beads
"ATALYST
Mumina )eads
SMALL REACTOR 5.5 cm diameter Z.56 or 3.7 liter of catalyst
a)
b)
c)
d)
Fig. 3. Loading of the catalyst inside reactors. the hydrogen is consumed by the reaction, the gas phase flow rate remains approximately constant along the catalyst bed, because of the presence of nitrogen and of a slight vaporization of the liquid phase. The mixture leaving the reactor is cooled and stored in a tank. The reaction products are analyzed in line by gas chromatography. To guarantee optimal wetting of the catalyst particles, the reactor is filled with liquid in up-flow before each experiment. The operating conditions are listed in Table 2. A first series of experiments was conducted with the small reactor [Fig. 3(a)-I in up-flow, in order to adjust the kinetic model. The superficial liquid velocities (U)
were varied between 0.005 and 0.015 m/s. In these conditions, catalyst wetting and the radial distribution of the two phases can be assumed to be perfect. The difficulty of this type of experiment is commensurate with the accuracy required for the analyses. To obtain conversions and selectivities with an accuracy better than 10%, it is necessary to measure the butadiene, butene-1, butenes-2 and n-butane concentrations with'an accuracy better than 0.5%. This is why each experimental result is obtained from at least three analyses, which are averaged. The results of this first series of experiments are shown in Figs 4 and 5. Figure 4 shows the butene-I selectivity as a function of X, conversion of butadiene.
Two-phase flow catalytic reactor Table 2. Operating conditions Reactor inlet temperature Total pressure
40°C 6.5 bar absolute
Catalyst bed Reactor diameter Average diameter of catalyst particles Height of catalyst bed Density of bed Porosity Density of particles
5.5 and 10.5 cm 2.2 mm 108-156 cm 702 kg/m 3 0.33 1047 kg/m 3
Gas phase Density Superficial mass flow rate Superficial velocity Initial hydrogen/butadiene molar ratio
12-13 kg/m 3 0.17-0.53 kg/m 2 s 0.014-0.044 m/s 0.5-1.5
Liquid phase Density Superficial mass flow rate Superficial velocity Dynamic viscosity Surface tension
560 kg/m 3 2.9-8.9 kg/m 2 s 0.005-0.015 m/s 1.4 × 10-4 Pa s 9.7 x 10 -3 N/m
Average composition offeed lsobutane n-Butane Trans-butene-2 Butene-1 lsobutene Cis-butene-2 lsopentane Butadiene Other
% weight 27.06 9.8 19.41 12.66 13.8 12.6 2.4 0.77 1.5
On the top of Fig. 4, are also plotted experimental points showing the sum of the three selectivities (butene-1, butenes-2 and n-butane) vs the conversion. The sum of these three selectivities should be equal to
3307
one. It appears that the mean precision is around 6%, but with much greater errors for some experiments. Note that the dispersion of results, resulting from the analytical errors, is such that the influence of the linear liquid velocity cannot be positively identified. In Fig. 5, the selectivities of transformation of butadiene to butenes-2 and n-butane are plotted against the conversion X. The dispersion of experimental points is larger for butenes-2, due to the fact that the analytical measurements must take into account two butenes-2, accordingly cis and trans isomers. F r o m these data obtained in up-flow with the small diameter reactor, the mathematical model has been adjusted. By multi-variable optimization, the optimal values of constants ki and of the parameters m and p, were determined for the closest representation of the experimental results: • m = 97.2, • p = 0.995, • k3/k4 = 10. The curves on Figs 4 and 5 give the results of the model for two liquid linear velocities. The model indicates only a very slight variation of the selectivity vs conversion curve with the liquid linear velocity. O n Fig. 4 is also plotted a curve representing the variations of selectivity vs conversion according to eq. (6), which does not take into account any mass transfer limitations. The influence of mass transfers appears finally to be very limited and noticeable only at high conversion. On the other hand, the model shows that the conversion depends not only on the space velocity (VVH), but also on the liquid linear velocity (U) and the hydrogen to butadiene ratio (H/BD), as shown on
1,200
, ....... J,'
1,000
t
[]
0,800
0,600 o
z
-
Up Small U=0.013 m/s
£x
Up Small U=0.009 m/s
X
Up Small U=0.007 m/s
o
Up Small U=0 005 m/s
-
Model U=0 015 m/s
. . . . .
0,400
Up Small U=0.015 m/s
O
MODEL U=0.005 m/s From Equation (6)
•
Sum of Selectivilies Mean of Sum of S
0,200
0,000
0.9
0.99 BUTADIENE CONVERSION (X)
Fig. 4. Comparison of experimental data with model, for butene-1 selectivity. Experimental errors are also indicated by the sum of selectivities.
0.999
3308
C. VERGELet al. 1,2000
I -
1,0000
.....
"
I
Selectivity BUTENES-2 U=0.015 m/s
-
Selectivity n-BUTANE U,0.015 m/s A
SelectivityBUTENES-2 Exp
0
Selectivityn-BUTANE Exp.
0,8000
0,6000
A
A 09 0,4000
/
f
f
0,2000
......
o. O
o_ca... . . .
73t~---
1--
"eL - ~- --
©- -
o .o_- qc-
----'13
0,0000
0
0.9
0.99
BUTADIENE CONVERSION (X)
Fig. 5. Comparison of up-flow experiments with model for butenes-2 and n-butane selectivities.
"'--r~''-''7 ~----.'7.'-" ;:-"-='-"'-:'~
×
°.~
°°
,°
0,8
#'
/ /, /'o"
,s,o, o ~°,,
-
~'"
U=0.015 m/s H/BD=I.2
,.('.."
U=0.005 m/s H/BD=I.2 U=0.005 m/s H/BD=l.15 ×
/;f f°°° 1, /
O O
......
U=0.015 m/s H/BD=1.25
-
,.,: -"
/"
:':::=
ooO..°
0,6
U=0015 m/s H/flD=l.29
X
U~0.015 m/s H/BD=l.17
O
U=0.0131 rn/s H/BD=l.26
•
U=0.009 m/s H/BD=l.26
,:.'."
U=0.009 m/s H/BD=I 21
,Z'"
•
U=0.O07rn/s H/BD=l.45 U=0.007 m/s H/BD=132
•
U=0.005 m/s HIBD=l.24
t
0,4 50,00
70,00
90,00
110,00
130,00
150,00
170,00
190,00
210,00
230,00
250,00
1NVH (s)
Fig. 6. Influence of space velocity and liquid linear velocity on conversion.
Fig. 6. On Fig. 6, the experimental points reported have been selected in order to show for the same space velocity the influence of the hydrogen-to-butadiene ratio. Other experiments were performed, again with the small reactor, but in down-flow and with linear liquid velocities also varying from 0.005 to 0.015 m/s. If the corresponding experimental points are plotted in Fig. 7, only the points corresponding to high linear velocities (U > 0.012 m/s) are seen to be represented by the model. At lower linear velocity, the loss of selectivity observed is too high to result from mass
transfer limitations. This loss of selectivity is probably due to a poor distribution of the phases over the packed bed. Two series of experiments in up-flow and downflow were then conducted with the larger reactor (inside diameter 10.5 c m - Fig. 3b). Given the liquidphase flow rate limitations, the linear liquid-phase velocities in this reactor could not exceed 0.009 m/s. Four comparative experiments were performed by inserting a Sulzer-type static mixer before the catalyst bed (Fig. 3c), in up-flow and in down-flow, with two linear liquid velocities (U = 0.005 and 0.009 m/s).
Two-phase flow catalytic reactor
3309
1,00 E3 ,3
0,80
Down Small U=0.0152 mls Down Small U=0.0131 m/s Down Small U=0 009 m/s
0,60
0.40
0,20
[]
+
Down Small U=O.007 m/s
©
Down Small U=0.005 rn/s
•
Up Large U=0.009-0.005 m/s
•
Down Large U=0.O09-O005 m/s
×
Up Large U=O.O09 m/s (+M)
m 0,00 "T,
X
Up Large U=0,005 m/s (+ M)
~ -0,20
•
Down Large U=0.005-0.009 m/s (. M)
•
Down Large U=0.005-0 009 m/s (+D)
-0,40
Model U=0 015 m/s
O -0,60
-0,80
A -1,00
0
0.9
0.99
0.999
Btn'AOIENEC,~VER$1ON(X) Fig. 7. Comparison of experimental results obtained with the small reactor (down-flow) and with the large reactor (up- or down-flow) ( + M = on line mixer ; + D = distributor).
Figure 7 shows that the experimental points corresponding to the up-flow experiments coincide with the model prediction. By contrast, the points obtained in up-flow without static mixer as well as in down-flow again indicate a loss of selectivity. Comparative experiments in down-flow were performed with and without mixer-distributor at the top of the catalyst bed [Fig. 3(d)]. The experimental points plotted in Fig. 7 show that, here also, the poor phase distribution is probably responsible for the deterioration of conversion selectivity.
INTERPRETATION
AND DISCUSSION OF RESULTS
Two important conclusions can be drawn from the results obtained. The selectivity of the selective hydrogenation reaction is highly sensitive to flow imperfections. The model that we have presented, taking into account all the mechanisms involved, can only account for a portion of the experimental points, namely, (i) those obtained in up-flow in the small reactor, (ii) those in down-flow in the small reactor with linear liquid velocities over 0.012 m/s, (iii) those obtained in the large reactor in up-flow, with the bottom of the reactor equipped with a static mixer. The problem of phase distribution at the inlet of the catalyst bed thus appears to be important. This is well known, but the result here is particularly demonstrative because of the types of phase and their physical properties, and also because of the specific sensitivity of the selectivity of the reaction with respect to conversion. In the small reactor with inside diameter 5.5 cm, the phase distribution is guaranteed by a layer of alumina particles (Fig. 3). This type of distributor
appears to be effective, except in down-flow for linear liquid velocities less than 0.012 m/s. In the large 10.5 cm diameter reactor, the layer of alumina particles, with or without a mixer-distributor, cannot guarantee a suitable distribution, either in up-flow or in down-flow. The only arrangement that proved satisfactory in up-flow is an in-line mixer of the Sulzer type [Fig. 3(c)]. This confirms the importance of efficient phase distribution and its decisive effect on selectivity. Note that, in down-flow, the same in-line mixer failed to provide improved results, compared with a distributor plate. The superiority of up-flow operation thus clearly emerges. In down-flow, it is known that phase segregation can be caused by the flow itself (Lutran et al., 1991). One is also led to consider the flow before it enters the reactor, i.e. in the feed lines and in the heat exchanger. After the hydrogen injection point in the liquid phase, the flow in small-diameter (1/4 in) and relatively long (several meters) tubes can take place in various regimes. In particular, this flow can be of the "slug" type. In these conditions, successive slugs of gas and liquid could only be destroyed if a device placed between the reactor inlet and the catalyst bed is capable of damping these sudden flow rate fluctuations. The in-line mixer used in up-flow appears to have these qualities, and to allow the re-mixing of the two phases, while guaranteeing good phase distribution across the cross-section of the bed. One is therefore led to consider that, not only the phase distribution defects across the cross-section of the catalyst bed are harmful, but that flow rate fluctuations must also be avoided. In general, it can be stated that any space-time fluctuation in the flow rates of the two phases can be
3310
C. VERGEL et al.
Table 3. Comparison of results for a reactor divided into two identical portions operating in parallel Mixer redistributor
FL FG
Catalyst volume
Butadiene conversion
Butene-1 selectivity
1
0
FL = 0.5 FG = 0.5
4 4.5 5
0.9496 0.9768 0.9901
0.4258 0.2955 0.1471
2
0
FL = 0.60 FG = 0.40
4 4.5 5
0.8933 0.9140 0.9281
0.1431 -- 0.0399 -- 0.2096
3
1
FL = 0.60 FG = 0.40
4 4.5 5
0.9400 0.9662 0.9814
0.3772 0.2375 0.0921
4
2
FL = 0.60 F6 = 0.40
4 4.5 5
0.9460 0.9727 0.9868
0.4043 0.2705 0.1221
5
0
FL = 0.55 FG = 0.45
4.5
0.9562
0.1928
6
0
FL = 0.65 FG = 0.35
4.5
0.8654
- 0.3384
Case
detrimental to reactor performance. This is especially significant for selective hydrogenation. The selectivity of the reaction in fact depends on the conversion, and, as one approaches total conversion of butadiene, selectivity drops rapidly and even becomes negative. As stated above, conversion is controlled by the accurate adjustment of the hydrogen flow rate. Any excess hydrogen over stoichiometry causes excessive conversion and hence poor selectivity. Any fluctuation (space-time) of the flow rates of the phases in the catalyst bed can lead to a situation where some portions of the bed may contain excess hydrogen, with a concomitant drop in selectivity. This situation can be modeled approximately by considering the reactor as split into two identical and parallel portions, each receiving a fraction of the liquid and gas phases. This phase distribution takes place while complying with the same pressure drop in each portion of the reactor [Fig. 8(a)]. Thus, for overall operating conditions corresponding to down-flow, and with a linear liquid velocity U = 0.015 m/s, one can compare the results given by the model for various fractions of liquid and gas passing through one of the two portions of the reactor. Listed in Table 3 are the results that could be obtained by placing redistributors inside the catalyst bed, capable of rehomogenizing the phases, before feeding the portions of the catalyst bed placed underneath [Fig. 8(b)]. Thus, up to two mixers-distributors were considered in a reactor, in which the poor phase distribution corresponds to FL and FG ratios. Case 1 is the ideal situation with only one bed of catalyst and a good initial phase distribution. Case 2 corresponds to an unbalanced initial phase distribution, but meeting the requirement of identical pressure drops in each portion of the bed. Case 3 is identical to case 2, but the catalyst bed has been split up in two
successive beds separated by an intermediate m i x e r redistributor. Finally, cases 5 and 6 are identical to case 2, with no intermediate mixers-redistributors, but with different FL and FG ratios in order to show the influence of these maldistribution ratios.
MIXER DISTRIBUTC
a) Model for simulating a non uniform distribution
b) Reactor with mixer-distributor between beds
Fig. 8. Models for simulating nonuniform distribution: (a) model for simulating a non uniform distribution; (b) reactor with mixer-distributor between beds.
Two-phase flow catalytic reactor
3311
0,7 0,6 0,5 0,4
•
One Bed-Case 1
-"
One Bed- Case 2 Two Beds- Case 3
0,3
g
-~ ~
Three Beds-Case4
~I o,2
MODEL U=0.015 m/s
t.U 0,1
×
One Bed- Case 6
o
x
One Bed- Case 7
-0,1 -0,2 -0,3 -0,4
0.9
0.99
ntrr~lmE CON'~FISION(X} Fig. 9. Effects of mixers- redistributors between beds (see Table 3). A glance at Fig. 9 presenting graphically all the data of Table 3 calls for the following comments. Slight variations in the phases distribution in the catalyst bed cause significant variations in conversion and selectivity. If one tries to maintain the conversion equal to the conversion obtained with a balanced distribution by increasing the catalyst volume (or feeding more hydrogen), the loss of selectivity will be substantial. The presence of two, or even a single mixer-distributor significantly improves conversion and selectivity simultaneously. A balanced distribution of the liquid and gas phases over the entire cross-section of the catalyst bed thus appears to be crucial for obtaining good performance.
g~ H/BD k~ Ki R (Ka)L,i (Ka)Ls.~ M ni N ~ff
CONCLUSION
By comparing the results obtained for the selective hydrogenation of butadiene to butene-1 with a model which accounts for the chemical kinetics, as well as the mass transfer mechanisms, we reached a number of practical conclusions. The selective hydrogenation reaction is extremely sensitive to any phases maldistribution in the catalyst bed. This could even represent a test-reaction to investigate problems of poor distribution in a fixed bed. These maldistributions can result from flow rate fluctuations, both in terms of space and time. The installation of phase mixers-redistributors in the catalyst bed can substantially enhance selectivity performance.
r~ St.2 t U VVH x X y z
Greek letters
NOTATION
Ci F
concentration of compounds Ai, mol/m 3 fraction of flow rate of a phase, dimensionless
mass per cent of compound A~, dimensionless initial hydrogen-to-butadiene molar ratio, dimensionless rate constant of reaction j, m3/s kg cata adsorption coefficient of compound A~, m3/mol grouped constant, 1/s gas-liquid mass transfer coefficient liquid phase side, 1/s liquid-solid mass transfer coefficient, 1/s mass of catalyst used, kg number of moles of compound A~ in liquid phase, mol molar flow rate per unit cross-sectional area of reactor, mol/m 2 s molar flux per unit volume of bed, mol/s m 3 intrinsic rate of reaction j, mol/s kg cata selectivity of conversion of butadiene to butene-1, dimensionless time, s liquid linear velocity m/s liquid space velocity, 1/s molar fraction in liquid phase, dimensionlesss butadiene conversion, dimensionless molar fraction in gas phase, dimensionless axial coordinate in catalyst bed, m
vi.~ Pc
liquid-vapor equilibrium coefficient, dimensionless stoichiometric coefficient of compound Ai in reaction j, dimensionless density of catalyst bed, kg/m 3
3312
C. VERGELet al.
Subscripts G L S i j 0
gas phase liquid phase catalyst solid c o m p o u n d Ai reaction j initial conditions at reactor inlet
Exponent •
value considered at physical equilibrium REFERENCES
Boitiaux, J. P., Cosyns, J., Derrien, M. and Leger, G., 1985a, Newest hydrogenation catalysts. Hydrocarbon Process. 64, 51-59. Boitiaux, J. P., Cosyns, J., Derrien, M. and Leger, G., 1985b, Proper design of butadiene selective hydrogenation process for maximum 1-butene yield by using comprehensive
kinetic modeling. Paper 1453 presented at the A.I.Ch.E. Spring National meeting, Houston. Charpentier, J. C., 1976, Recent progress in two-phase gasqiquid mass transfer in packed beds. Chem. Engng J, 11, 161-181. Lutran, G., Ng, K. and Delikat, E., 1991, Liquid distribution in trickle beds. An experimental study using computer assisted tomography. Ind. Engng Chem. Res. 30, 1270-1280. Mochizuki, S., 1982, Empirical expressions of liquid-solid mass transfer in cocurrent gas-liquid upflow fixed beds. Chem. Engng Sci. 37(9), 1422-1424. Rao, V. G. and Drinkenburg, A. A. H., 1985, Solid-liquid mass transfer in packed beds with cocurrent gas-liquid flow downflow. A.I.Ch.E.J. 31, 1059-1068. Satterfield, C. N., 1975, Trickle-bed reactors. A.I.Ch.E.J. 21, 209-228. Vergel, C., 1993, Les r6acteurs catalytiques fi lit fixe avec 6coulement de gazet de liquide--Comparaison sur le plan th6orique et exp6rimental. Thesis, Institut National Polytechnique de Lorraine.
Chemical Engineerin 0 Science, VoL 50, No. 20, pp. 3303-3312, 1995 Copyright ~) 1995 Elsevier Science Ltd Printed in Great Britain. All rights reserved 0009-2509/95 $9.50 + 0.00
0009-2909(95)00141-7
T W O - P H A S E F L O W CATALYTIC REACTOR, I N F L U E N C E O F H Y D R O D Y N A M I C S O N SELECTIVITY t C. VERGEL Instituto Colombiano del Petroleo (ICP), A.A. 4185, Bucaramanga, Colombia J. P. EUZEN, P. TRAMBOUZE: and J. P. WAUQUIER Institut Fran~ais du P6trole, Centre d'Etudes et de D6veloppement Industriels, 69390 Solaize, France
(Received 28 October 1994; manuscript received and accepted in revisedform 10 April 1995) Abstract--In this paper, experimental results are presented, showing the influence of the hydrodynamic behavior of fixed-bed catalytic reactors with two-phase (gas + liquid) flow on the selectivityof a reaction. A mathematical model has been developed in order to represent up-flow data, for which the distribution of phases across the catalytic bed is near ideal. The influenceof maldistribution of the phases is noticeable for down-flow conditions. In this case, recommendations for the design of the reactor are given.
INTRODUCTION
An industrial chemical reaction was used for this study: the hydrogenation of a C4 olefinic cut, to convert selectively the small amount of butadiene present to butene-l, with minimum butane production. The selectivity of this reaction, to be preserved at very high butadiene conversion, appears to be highly sensitive to the flow characteristics. A model, including the complex chemical kinetics of the Langmuir-Hinshelwood type, with two parallel and two consecutive reactions, and the mass transfers between the phases (gas-liquid and liquid-solid), was developed to represent the experimental results. Discrepancies found between some experimental results and the model were interpreted on the basis of phase maldistribution and segregation in the catalyst bed. A more complex model, accounting for these segregation effects and, more precisely, local fluctuations of gas-liquid ratio, serves to interpret all the experimental results. Recommendations are proposed for the design of concurrent gas-liquid fixed-bed catalytic reactors, as these systems are very sensitive to flow regimes. KINETICEXPRESSIONS The aim of the selective hydrogenation of a C4 exsteam cracking cut is to remove any remaining butadiene after this compound has been extracted, with minimum effect on butenes and favoring the production of butene-1.
t Paper presented at ISCRE 13 Meeting in Baltimore, U.S.A., September 1994. : Corresponding author.
An intensive laboratory investigation (Boitiaux et al., 1985a, b) with a suitable catalyst led to the selection of the simplified reaction scheme: shown in Fig. 1.
At
1 , A2 2Xa 3~
4 ,
, A4
A3 Reaction 1: At + A5 --, A2
Reaction 2: AI + As --*A3 Reaction 3:
A2 " * A3
Reaction 4:A2 + As -, A4 where At = 1-3 butadiene A2 = butene-1 A3 = butene-2 A4 = normal butane A5 = hydrogen Fig. 1.
In the range of temperatures and hydrogen partial pressures used, experiments conducted in a batch reactor, with a catalyst powder finely dispersed in the liquid phase, showed that the kinetic equations of the process could be written as follows: dnl
Kt Ct
M d--'-~t---- -- rl -- r2 = -- (kl + k2)K1C1 + K2C2 C5
3303
(1)
3304 dn2 Mdt
-
C. VERGEL et al.
klK1C1 =
r 1 --
r 3 --
r 4
- (k3 + k 4 ) K 2 C 2
=
C 5
K1CI
+ K2C2
Parameters p and m depend on the catalyst, and must therefore be determined experimentally.
(2) REACTOR MODELING d/13
- - = r
Mdt
k 2 K 1 C 1 "4- k 3 K 2 C 2 C5
2 +r 3 =
(3)
KIC1 + K2C2
dn4 k4 K 2 C2 M d t = r4 = K1C1 + K2C2 C5.
(4)
The following constants are introduced for simplification:
R1 = k l K 1 C s ,
R2=k2KICs
~3 = kaK2Cs,
~4 = k4KECs.
The ratio of the two adsorption constants K1 and K 2 w a s found to approach 300. By defining m
and
p= -
~ 3 "~ ~ 4
-
~ 1 -~" ~:~2
it is shown that the combination of eqs (1) and (2), followed by an integration, helps to express the number of moles of butene-1 formed as a function of butadiene conversion X: - - X ) l/m --
n2 - - n 2 o = n 2 o [ ( 1
P
1/m
1 -
1]
nlo I-(1 - X) -- (1 - X)'/'-I
(5)
The simplest model that can be written for the different processes occurring inside the reactor must take account of the following mechanisms: (i) influence of mass transfer between gas phase and liquid phase, (ii) influence of mass transfer between liquid phase and catalyst surface, (iii) occurrence of the various chemical reactions inside the catalyst particles. As a first approximation, the gas and liquid concurrent flows can be considered as of the plug flow type. Moreover, because of the impregnation by the active metal limited to the external portion of the catalyst particles, the effects of intragranular diffusion are negligible. The reaction is assumed to be isothermal. The equations reflecting the molar balances in each phase in steady-state conditions can be written as follows.
Gas phase dNa, i d-~- + JC'aL.i = 0
(7)
The first term is the convective flux due to the gas flow. The second term representing the transfer flux between the gas phase and the liquid phase, can be expressed in terms of the film theory, as follows:
where X =
~P GL, i = ( K a ) L , i (C*,i -- CL, i ).
/110 - - n l /11o
Liquid phase
Since the volume of liquid phase remains virtually constant during hydrogenation, and since the molecular weights of the different C4 compounds are similar, the number of moles ni can be replaced either by the molar concentrations or by the weight percentages. In the case of a fixed bed reactor, assuming a plug flow through the bed and no mass transfer limitations, the material balance for butene-1 gives an expression similar to eq. (5), i,e.
Pl/m glo[(1 - X) -- (1 -- X)I:"].
(6)
This helps to calculate the selectivity S 1,2 of the transformation of butadiene to butene-1, defined as follows: $1,2
n 2 - - /120
-
-
=
/11o - - /11
-
g2 -- g20
-
dNL, i dz
JI/'GL, i -~- J~LS, i = O.
(9)
The first term is the convective flux due to the liquid flow. The third term representing the transfer flux between the liquid phase and the solid phase can also be expressed in terms of the film theory:
Jff Ls.i = (Ka)Ls.i(CL.i -- Cs.i).
(10)
It also equals the quantity of compound A~converted per unit volume of catalyst bed, and can be written:
g 2 - - g 2 o = g 2 o [ ( 1 - - X ) TM -- 1 ]
1-
(8)
with
glo -- gl
gl0 - - g l
X ~ . - 01o
According to eq. (6), the selectivity of the hydrogenation of butadiene to butene-1 is linked to the conversion X and influenced by three parameters: (i) the initial ratio of butene-1 and butadiene concentrations, g2o/glo, (ii) parameter m, (iii) parameter p.
~I/'LS, i "+- pcRs, i = 0
(11)
where R~,i is the reaction rate per unit mass of catalyst, with respect to compound A:
Rs, i = ~, vi.:j.
(12)
The mass transfer coefficients can be estimated from correlations proposed in the literature depending upon the type of flow (Vergel, 1993). After comparing the available correlations, we selected those listed in Table 1 for most of our calculations. EXPERIMENTAL
The pilot plant used to compare the performance obtained in up-flow and down-flow is shown in Fig. 2. Liquid feed (C4 cut), from a container pressurized
3305
Two-phase flow catalytic reactor Table 1 Flow direction
Correlation
Reference
Up-flow
(Ka)L.i (Ka)ts.i (Ka)L.i (Ka)Ls.i
Satterfield (1975) Mochizuki (1982) Charpentier (1976) Rao and Drinkenburg (1985)
Down-flow
i Nz
I
"~
~ ~ ~
FEED STORAGE TANK
_® ....
COOLING
REACTOR
i
STORAGE
Fig. 2. Scheme of the pilot-plant used for performing experiments.
with nitrogen, is mixed with hydrogen and sent to a heat exchanger, which helps to keep the feed temperature of the reactor constant (40°C). Two stainless reactors, with inside diameters of 5.5 and 10.5 cm, respectively, can be mounted on the unit. The catalyst bed heights are accordingly 108 and 156 cm for the small reactor, and 108 cm in the larger reactor. The catalyst volumes used thus range between 2560 and 9400 cm 3. The catalyst used is an industrial catalyst (Pd on alumina), specially designed for high selectivity; it is only impregnated at the bead surface. The fluid flow direction can be up-flow or down-flow, as desired. Inside each reactor, the catalyst is placed between two layers of inert alumina particles 4 to 6 mm in diameter. The catalyst arrangements inside the reactor are shown in Fig. 3. The reaction is carried out under pressure (6.5 bar) and at relatively low temperature (40-45°C), so that most of the C4 cut feedstock remains liquid, with only a small part in the vapor phase with hydrogen. The feed is characterized by: • very high initial content of the desired product, butene-1, • very low initial content of the product to be removed, butadiene.
The influence of feed flow rate and of initial hydrogen to butadiene ratio (H/BD) was investigated in order to vary the butadiene conversion, and hence selectivity, within a wide range (40-99.5%). A number of special features of the process must be pointed out: • The temperature in the reactor is virtually constant, the heat generated by the reaction being compensated by a slight vaporization of the C4 cut and heat losses. • Butadiene conversion is controlled by the quantity of hydrogen injected into the feed. • The liquid phase has rather unusual physical properties because it is near the critical point. Thus the density is about 560 kg/m 3 and the surface tension is about 10- 2 N/m. However, the liquid phase displays no tendency to foaming. The gas and liquid phases entering the reactor can be assumed to be in thermodynamic equilibrium. The corresponding calculations were thus carried out to determine the flow rates and compositions of the two phases at this point. In these calculations, the nitrogen (used for feed pressurization) dissolved in the C4 cut was also taken into account. Note that, even if most of
3306
C. VERGELet al.
I Alumina beads
"ATALYST
Mumina )eads
SMALL REACTOR 5.5 cm diameter Z.56 or 3.7 liter of catalyst
a)
b)
c)
d)
Fig. 3. Loading of the catalyst inside reactors. the hydrogen is consumed by the reaction, the gas phase flow rate remains approximately constant along the catalyst bed, because of the presence of nitrogen and of a slight vaporization of the liquid phase. The mixture leaving the reactor is cooled and stored in a tank. The reaction products are analyzed in line by gas chromatography. To guarantee optimal wetting of the catalyst particles, the reactor is filled with liquid in up-flow before each experiment. The operating conditions are listed in Table 2. A first series of experiments was conducted with the small reactor [Fig. 3(a)-I in up-flow, in order to adjust the kinetic model. The superficial liquid velocities (U)
were varied between 0.005 and 0.015 m/s. In these conditions, catalyst wetting and the radial distribution of the two phases can be assumed to be perfect. The difficulty of this type of experiment is commensurate with the accuracy required for the analyses. To obtain conversions and selectivities with an accuracy better than 10%, it is necessary to measure the butadiene, butene-1, butenes-2 and n-butane concentrations with'an accuracy better than 0.5%. This is why each experimental result is obtained from at least three analyses, which are averaged. The results of this first series of experiments are shown in Figs 4 and 5. Figure 4 shows the butene-I selectivity as a function of X, conversion of butadiene.
Two-phase flow catalytic reactor Table 2. Operating conditions Reactor inlet temperature Total pressure
40°C 6.5 bar absolute
Catalyst bed Reactor diameter Average diameter of catalyst particles Height of catalyst bed Density of bed Porosity Density of particles
5.5 and 10.5 cm 2.2 mm 108-156 cm 702 kg/m 3 0.33 1047 kg/m 3
Gas phase Density Superficial mass flow rate Superficial velocity Initial hydrogen/butadiene molar ratio
12-13 kg/m 3 0.17-0.53 kg/m 2 s 0.014-0.044 m/s 0.5-1.5
Liquid phase Density Superficial mass flow rate Superficial velocity Dynamic viscosity Surface tension
560 kg/m 3 2.9-8.9 kg/m 2 s 0.005-0.015 m/s 1.4 × 10-4 Pa s 9.7 x 10 -3 N/m
Average composition offeed lsobutane n-Butane Trans-butene-2 Butene-1 lsobutene Cis-butene-2 lsopentane Butadiene Other
% weight 27.06 9.8 19.41 12.66 13.8 12.6 2.4 0.77 1.5
On the top of Fig. 4, are also plotted experimental points showing the sum of the three selectivities (butene-1, butenes-2 and n-butane) vs the conversion. The sum of these three selectivities should be equal to
3307
one. It appears that the mean precision is around 6%, but with much greater errors for some experiments. Note that the dispersion of results, resulting from the analytical errors, is such that the influence of the linear liquid velocity cannot be positively identified. In Fig. 5, the selectivities of transformation of butadiene to butenes-2 and n-butane are plotted against the conversion X. The dispersion of experimental points is larger for butenes-2, due to the fact that the analytical measurements must take into account two butenes-2, accordingly cis and trans isomers. F r o m these data obtained in up-flow with the small diameter reactor, the mathematical model has been adjusted. By multi-variable optimization, the optimal values of constants ki and of the parameters m and p, were determined for the closest representation of the experimental results: • m = 97.2, • p = 0.995, • k3/k4 = 10. The curves on Figs 4 and 5 give the results of the model for two liquid linear velocities. The model indicates only a very slight variation of the selectivity vs conversion curve with the liquid linear velocity. O n Fig. 4 is also plotted a curve representing the variations of selectivity vs conversion according to eq. (6), which does not take into account any mass transfer limitations. The influence of mass transfers appears finally to be very limited and noticeable only at high conversion. On the other hand, the model shows that the conversion depends not only on the space velocity (VVH), but also on the liquid linear velocity (U) and the hydrogen to butadiene ratio (H/BD), as shown on
1,200
, ....... J,'
1,000
t
[]
0,800
0,600 o
z
-
Up Small U=0.013 m/s
£x
Up Small U=0.009 m/s
X
Up Small U=0.007 m/s
o
Up Small U=0 005 m/s
-
Model U=0 015 m/s
. . . . .
0,400
Up Small U=0.015 m/s
O
MODEL U=0.005 m/s From Equation (6)
•
Sum of Selectivilies Mean of Sum of S
0,200
0,000
0.9
0.99 BUTADIENE CONVERSION (X)
Fig. 4. Comparison of experimental data with model, for butene-1 selectivity. Experimental errors are also indicated by the sum of selectivities.
0.999
3308
C. VERGELet al. 1,2000
I -
1,0000
.....
"
I
Selectivity BUTENES-2 U=0.015 m/s
-
Selectivity n-BUTANE U,0.015 m/s A
SelectivityBUTENES-2 Exp
0
Selectivityn-BUTANE Exp.
0,8000
0,6000
A
A 09 0,4000
/
f
f
0,2000
......
o. O
o_ca... . . .
73t~---
1--
"eL - ~- --
©- -
o .o_- qc-
----'13
0,0000
0
0.9
0.99
BUTADIENE CONVERSION (X)
Fig. 5. Comparison of up-flow experiments with model for butenes-2 and n-butane selectivities.
"'--r~''-''7 ~----.'7.'-" ;:-"-='-"'-:'~
×
°.~
°°
,°
0,8
#'
/ /, /'o"
,s,o, o ~°,,
-
~'"
U=0.015 m/s H/BD=I.2
,.('.."
U=0.005 m/s H/BD=I.2 U=0.005 m/s H/BD=l.15 ×
/;f f°°° 1, /
O O
......
U=0.015 m/s H/BD=1.25
-
,.,: -"
/"
:':::=
ooO..°
0,6
U=0015 m/s H/flD=l.29
X
U~0.015 m/s H/BD=l.17
O
U=0.0131 rn/s H/BD=l.26
•
U=0.009 m/s H/BD=l.26
,:.'."
U=0.009 m/s H/BD=I 21
,Z'"
•
U=0.O07rn/s H/BD=l.45 U=0.007 m/s H/BD=132
•
U=0.005 m/s HIBD=l.24
t
0,4 50,00
70,00
90,00
110,00
130,00
150,00
170,00
190,00
210,00
230,00
250,00
1NVH (s)
Fig. 6. Influence of space velocity and liquid linear velocity on conversion.
Fig. 6. On Fig. 6, the experimental points reported have been selected in order to show for the same space velocity the influence of the hydrogen-to-butadiene ratio. Other experiments were performed, again with the small reactor, but in down-flow and with linear liquid velocities also varying from 0.005 to 0.015 m/s. If the corresponding experimental points are plotted in Fig. 7, only the points corresponding to high linear velocities (U > 0.012 m/s) are seen to be represented by the model. At lower linear velocity, the loss of selectivity observed is too high to result from mass
transfer limitations. This loss of selectivity is probably due to a poor distribution of the phases over the packed bed. Two series of experiments in up-flow and downflow were then conducted with the larger reactor (inside diameter 10.5 c m - Fig. 3b). Given the liquidphase flow rate limitations, the linear liquid-phase velocities in this reactor could not exceed 0.009 m/s. Four comparative experiments were performed by inserting a Sulzer-type static mixer before the catalyst bed (Fig. 3c), in up-flow and in down-flow, with two linear liquid velocities (U = 0.005 and 0.009 m/s).
Two-phase flow catalytic reactor
3309
1,00 E3 ,3
0,80
Down Small U=0.0152 mls Down Small U=0.0131 m/s Down Small U=0 009 m/s
0,60
0.40
0,20
[]
+
Down Small U=O.007 m/s
©
Down Small U=0.005 rn/s
•
Up Large U=0.009-0.005 m/s
•
Down Large U=0.O09-O005 m/s
×
Up Large U=O.O09 m/s (+M)
m 0,00 "T,
X
Up Large U=0,005 m/s (+ M)
~ -0,20
•
Down Large U=0.005-0.009 m/s (. M)
•
Down Large U=0.005-0 009 m/s (+D)
-0,40
Model U=0 015 m/s
O -0,60
-0,80
A -1,00
0
0.9
0.99
0.999
Btn'AOIENEC,~VER$1ON(X) Fig. 7. Comparison of experimental results obtained with the small reactor (down-flow) and with the large reactor (up- or down-flow) ( + M = on line mixer ; + D = distributor).
Figure 7 shows that the experimental points corresponding to the up-flow experiments coincide with the model prediction. By contrast, the points obtained in up-flow without static mixer as well as in down-flow again indicate a loss of selectivity. Comparative experiments in down-flow were performed with and without mixer-distributor at the top of the catalyst bed [Fig. 3(d)]. The experimental points plotted in Fig. 7 show that, here also, the poor phase distribution is probably responsible for the deterioration of conversion selectivity.
INTERPRETATION
AND DISCUSSION OF RESULTS
Two important conclusions can be drawn from the results obtained. The selectivity of the selective hydrogenation reaction is highly sensitive to flow imperfections. The model that we have presented, taking into account all the mechanisms involved, can only account for a portion of the experimental points, namely, (i) those obtained in up-flow in the small reactor, (ii) those in down-flow in the small reactor with linear liquid velocities over 0.012 m/s, (iii) those obtained in the large reactor in up-flow, with the bottom of the reactor equipped with a static mixer. The problem of phase distribution at the inlet of the catalyst bed thus appears to be important. This is well known, but the result here is particularly demonstrative because of the types of phase and their physical properties, and also because of the specific sensitivity of the selectivity of the reaction with respect to conversion. In the small reactor with inside diameter 5.5 cm, the phase distribution is guaranteed by a layer of alumina particles (Fig. 3). This type of distributor
appears to be effective, except in down-flow for linear liquid velocities less than 0.012 m/s. In the large 10.5 cm diameter reactor, the layer of alumina particles, with or without a mixer-distributor, cannot guarantee a suitable distribution, either in up-flow or in down-flow. The only arrangement that proved satisfactory in up-flow is an in-line mixer of the Sulzer type [Fig. 3(c)]. This confirms the importance of efficient phase distribution and its decisive effect on selectivity. Note that, in down-flow, the same in-line mixer failed to provide improved results, compared with a distributor plate. The superiority of up-flow operation thus clearly emerges. In down-flow, it is known that phase segregation can be caused by the flow itself (Lutran et al., 1991). One is also led to consider the flow before it enters the reactor, i.e. in the feed lines and in the heat exchanger. After the hydrogen injection point in the liquid phase, the flow in small-diameter (1/4 in) and relatively long (several meters) tubes can take place in various regimes. In particular, this flow can be of the "slug" type. In these conditions, successive slugs of gas and liquid could only be destroyed if a device placed between the reactor inlet and the catalyst bed is capable of damping these sudden flow rate fluctuations. The in-line mixer used in up-flow appears to have these qualities, and to allow the re-mixing of the two phases, while guaranteeing good phase distribution across the cross-section of the bed. One is therefore led to consider that, not only the phase distribution defects across the cross-section of the catalyst bed are harmful, but that flow rate fluctuations must also be avoided. In general, it can be stated that any space-time fluctuation in the flow rates of the two phases can be
3310
C. VERGEL et al.
Table 3. Comparison of results for a reactor divided into two identical portions operating in parallel Mixer redistributor
FL FG
Catalyst volume
Butadiene conversion
Butene-1 selectivity
1
0
FL = 0.5 FG = 0.5
4 4.5 5
0.9496 0.9768 0.9901
0.4258 0.2955 0.1471
2
0
FL = 0.60 FG = 0.40
4 4.5 5
0.8933 0.9140 0.9281
0.1431 -- 0.0399 -- 0.2096
3
1
FL = 0.60 FG = 0.40
4 4.5 5
0.9400 0.9662 0.9814
0.3772 0.2375 0.0921
4
2
FL = 0.60 F6 = 0.40
4 4.5 5
0.9460 0.9727 0.9868
0.4043 0.2705 0.1221
5
0
FL = 0.55 FG = 0.45
4.5
0.9562
0.1928
6
0
FL = 0.65 FG = 0.35
4.5
0.8654
- 0.3384
Case
detrimental to reactor performance. This is especially significant for selective hydrogenation. The selectivity of the reaction in fact depends on the conversion, and, as one approaches total conversion of butadiene, selectivity drops rapidly and even becomes negative. As stated above, conversion is controlled by the accurate adjustment of the hydrogen flow rate. Any excess hydrogen over stoichiometry causes excessive conversion and hence poor selectivity. Any fluctuation (space-time) of the flow rates of the phases in the catalyst bed can lead to a situation where some portions of the bed may contain excess hydrogen, with a concomitant drop in selectivity. This situation can be modeled approximately by considering the reactor as split into two identical and parallel portions, each receiving a fraction of the liquid and gas phases. This phase distribution takes place while complying with the same pressure drop in each portion of the reactor [Fig. 8(a)]. Thus, for overall operating conditions corresponding to down-flow, and with a linear liquid velocity U = 0.015 m/s, one can compare the results given by the model for various fractions of liquid and gas passing through one of the two portions of the reactor. Listed in Table 3 are the results that could be obtained by placing redistributors inside the catalyst bed, capable of rehomogenizing the phases, before feeding the portions of the catalyst bed placed underneath [Fig. 8(b)]. Thus, up to two mixers-distributors were considered in a reactor, in which the poor phase distribution corresponds to FL and FG ratios. Case 1 is the ideal situation with only one bed of catalyst and a good initial phase distribution. Case 2 corresponds to an unbalanced initial phase distribution, but meeting the requirement of identical pressure drops in each portion of the bed. Case 3 is identical to case 2, but the catalyst bed has been split up in two
successive beds separated by an intermediate m i x e r redistributor. Finally, cases 5 and 6 are identical to case 2, with no intermediate mixers-redistributors, but with different FL and FG ratios in order to show the influence of these maldistribution ratios.
MIXER DISTRIBUTC
a) Model for simulating a non uniform distribution
b) Reactor with mixer-distributor between beds
Fig. 8. Models for simulating nonuniform distribution: (a) model for simulating a non uniform distribution; (b) reactor with mixer-distributor between beds.
Two-phase flow catalytic reactor
3311
0,7 0,6 0,5 0,4
•
One Bed-Case 1
-"
One Bed- Case 2 Two Beds- Case 3
0,3
g
-~ ~
Three Beds-Case4
~I o,2
MODEL U=0.015 m/s
t.U 0,1
×
One Bed- Case 6
o
x
One Bed- Case 7
-0,1 -0,2 -0,3 -0,4
0.9
0.99
ntrr~lmE CON'~FISION(X} Fig. 9. Effects of mixers- redistributors between beds (see Table 3). A glance at Fig. 9 presenting graphically all the data of Table 3 calls for the following comments. Slight variations in the phases distribution in the catalyst bed cause significant variations in conversion and selectivity. If one tries to maintain the conversion equal to the conversion obtained with a balanced distribution by increasing the catalyst volume (or feeding more hydrogen), the loss of selectivity will be substantial. The presence of two, or even a single mixer-distributor significantly improves conversion and selectivity simultaneously. A balanced distribution of the liquid and gas phases over the entire cross-section of the catalyst bed thus appears to be crucial for obtaining good performance.
g~ H/BD k~ Ki R (Ka)L,i (Ka)Ls.~ M ni N ~ff
CONCLUSION
By comparing the results obtained for the selective hydrogenation of butadiene to butene-1 with a model which accounts for the chemical kinetics, as well as the mass transfer mechanisms, we reached a number of practical conclusions. The selective hydrogenation reaction is extremely sensitive to any phases maldistribution in the catalyst bed. This could even represent a test-reaction to investigate problems of poor distribution in a fixed bed. These maldistributions can result from flow rate fluctuations, both in terms of space and time. The installation of phase mixers-redistributors in the catalyst bed can substantially enhance selectivity performance.
r~ St.2 t U VVH x X y z
Greek letters
NOTATION
Ci F
concentration of compounds Ai, mol/m 3 fraction of flow rate of a phase, dimensionless
mass per cent of compound A~, dimensionless initial hydrogen-to-butadiene molar ratio, dimensionless rate constant of reaction j, m3/s kg cata adsorption coefficient of compound A~, m3/mol grouped constant, 1/s gas-liquid mass transfer coefficient liquid phase side, 1/s liquid-solid mass transfer coefficient, 1/s mass of catalyst used, kg number of moles of compound A~ in liquid phase, mol molar flow rate per unit cross-sectional area of reactor, mol/m 2 s molar flux per unit volume of bed, mol/s m 3 intrinsic rate of reaction j, mol/s kg cata selectivity of conversion of butadiene to butene-1, dimensionless time, s liquid linear velocity m/s liquid space velocity, 1/s molar fraction in liquid phase, dimensionlesss butadiene conversion, dimensionless molar fraction in gas phase, dimensionless axial coordinate in catalyst bed, m
vi.~ Pc
liquid-vapor equilibrium coefficient, dimensionless stoichiometric coefficient of compound Ai in reaction j, dimensionless density of catalyst bed, kg/m 3
3312
C. VERGELet al.
Subscripts G L S i j 0
gas phase liquid phase catalyst solid c o m p o u n d Ai reaction j initial conditions at reactor inlet
Exponent •
value considered at physical equilibrium REFERENCES
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