I n d . Eng. Chem. Res. 1989,28, 1489-1493
energy balance for the reactor Vlpcp dTl/dt = FpcpTo + rFpcpT4 + Vl(-IIH)Rl (1 + r)FpcpTl - UIAl(Tl - Tcl) energy balance for the jacket vJ1pccp, d T c l / d t
=
+ UlAl(T1 - Tcl) Second Reactor: material balance for substance A v 2 dC~n/dt= (1 + r)FCA1 - (1 + ~)FCA~ - R2V2 energy balance for the reactor V2pcP dT,/dt = (1 + r)FpcpT1 + V2(-AH)R, (1 + r)FpcPT2- U Z A ~ T-ZTC2) energy balance for the jacket FclPcCpcTeo - Fc1PCpcTc1
vJ2pccp,
dTc2/dt
= Fc2PcCp,TCo
- Fc2PCpcTc2 + U2A2(T2 - Tc2)
reaction rate
Separator: Steady State F4= rF3 F3= F
F2= F3+ F4
Literature Cited Cheng, C.; Brosilow, C. Model Predictive Control of Unstable Systems. AIChE Annual Meeting, New York, Nov 1987.
1489
Cheung, T. F.; Luyben, W. L. PD Control Improves Reactor Stability. Hydrocarbon Process. 1979 (Sept), 215. Devia, N.; Luyben, W. Reactors: Size Versus Stability. Hydrocarbon Process. 1978 (June), 119. Edgar, T. F. Status of Design Methods For Multivariable Control. AIChE Symp. Ser. 1976, 72(159), 100. Friedland, B. Control System Design. A n Introduction To StateSpace Methods; McGraw-Hill: New York, 1986; p 230. Georgiou, A.; Georgakis, C.; Luyben, W. L. Design of Practical Multivariable Process Controllers. Tech. Research Progress Report No. 3; Chemical Process Modeling ahd Control Center, Lehigh University: Bethlehem, Sept 1986. Hwang, S. H.; Chang, H. C. A Theoretical Examination of Closedloop Properties and Tuning Methods of Single-Loop P I Controllers. Chem. Eng. Sci. 1987, 42(10), 2395-2415. Kuo, B. C. Automatic Control Systems, 5th ed.; Prentice-Hak New York, 1985. Postlethwaite, I.; MacFarlane, A. G. J. Lecture Notes in Control and Information Sciences; Springer: New York, 1979; p 12. Power, H. M. A New Result on Eigenvalue Assignment by Means of Dyadic Output Feedback. Ind. J. Control 1975,21, 149. Ray, H. W. Advanced Process Control; McGraw-Hill: New York, 1981. Tzafestas, S. G. Multivariable Control. New Concepts and Tools; D. Reidel: Dordrecht, The Netherlands, 1984. Wonham, W. M. On Pole Assignment in Multi-Input Controllable Linear Systems. IEEE Trans. Autom. Control 1967, AC-12,660. Yagle, A.; Levy, B. C. Multivariable Root Loci on the Real Axis. Int. J. Control 1982, 35(3), 491-507. Zafiriou, E.; Morari, M. Robust H,-Type IMC Controller Design Via the Structured Singular Value. ZFAC 1987.
Received for review August 16, 1988 Revised manuscript received April 20, 1989 Accepted June 13, 1989
Scaling Up of a Monolithic Catalyst Reactor with Two-Phase Flow Said Irandoust and Bengt Andersson* Department of Chemical Reaction Engineering, Chalmers University of .Technology, S-412 96 Gothenburg, Sweden
Erik Bengtsson and Mikael Siverstrom E K A Nobel AB, S-445 01 Surte, Sweden
Scaling up of the monolithic three-phase reactor has been studied in the hydrogenation step of anthraquinones for large-scale production of hydrogen peroxide. The influence of flow pattern (bubble flow and slug flow), flow distribution, pressure drop, and mass transfer has been investigated. It has been illustrated that the segmented gas-liquid flow (slug flow) gives the highest production rate with very small scale-up effects. 1. Introduction Multiphase reactors involving gas, liquid, and solid catalysts have important applications in industrial threephase processes. Due to the presence of three phases, the interfacial transport is of major importance for the reactor design. A novel multiphase reactor, the monolithic catalyst reactor, satisfies certain criteria of a desirable catalytic reactor, such as low pressure drop, high catalyst efficiency, uniform flow distribution, low axial dispersion, and simple reactor design. In this type of reactor, the catalyst consists of a large number of parallel straight channels in which the gas and liquid reactants flow cocurrently. A schematic design of the monolithic catalyst support is shown in Figure 1. Various aspects of this reactor have recently been reviewed by Irandoust and Andersson (1988a). Among several possible flow patterns in the monolith channels are slug flow and bubble flow (Figure 2). In both of these flows, the dispersed gas disturbs the laminar flow
in the liquid phase and forces the liquid to recirculate within the liquid plugs. This recirculation increases the radial mass transfer and decreases the axial dispersion. Hence, slug flow or bubble flow is desirable in monolith channels. The flow properties of Taylor flow in the monolithic catalyst support have been studied by Irandoust and Andersson (1989). The present study is concerned with the industrial application of monolithic catalyst reactors in hydrogen peroxide production. The formation of certain flow patterns, the pressure drop over the reactor, and the reactor efficiency in both bubble and slug flow regimes have been studied. Furthermore, scaling up of the monolithic reactor with slug flow has been investigated with a pilot plant reactor and an industrial reactor. 2. Experimental Section The monolithic support used in this investigation con-
0888-588518912628-1489$01.50/0 0 1989 American Chemical Society
1490 Ind. Eng. Chem. Res., Vol. 28, No. 10, 1989 Table I. Perforated Plates Tested for Distribution of the Liquid Phase perforation, % hole diam, mm Apitch, mm 27 22.7 22.7 18 8.2 5.7
I:]
Figure 1. Monolithic catalyst support. Reprinted with permission from Irandoust and Andersson (1988b). Copyright 1988 Pergamon
1.5
2.75 2 4 4.5 2.0 2.0
1.0 2.0 2.0
0.6 0.5
Table 11. Physical Properties of the Liquids Used in the Pressure Drop Measurements 103pl,kg/(m s) pl, kg/ms 103a1,N/m water 1 1000 72.4 ethanol 1.38 804 22.5 ethylene glycol 19.9 1110 47.7
. ..... '.. ..
. . . '. . . ,
. .. .
,
.
. . . . .. .
.,
. .
.
.. .. ..' ... .. . . .. . , _
......
.. .. . . I
.
... . . .... ,
.
'
'
I
t
Slug
Taylor
Annular
flow
flow
flow
Figure 2. Flow patterns in vertical upward flow. Reprinted with permission from Irandoust and Andersson (1988a). Copyright 1988 Marcel Dekker. Inc.
sisted of a mixture of various calcium aluminum silicates. The cell shape of the monolithic channels was sinusoidal with a cross-sectional area of 2.1 mm2per channel. A large surface area was obtained by wash-coating the surface of the carrier with siliceous gel. More details about the support can be found elsewhere (Berglin and Herrmann, 1982). Observation of the Flow Pattern. The flow pattern was observed for a reasonable range of flow rates of working solution and air at room conditions. Experiments were conducted in a 1.5-mm-internal diameter and 200mm-long circular glass tube, using a photographic technique. The inner wall of the glass tube was coated by siliceous gel in order to obtain a similar wetting as that for the monolith channel wall. The segmented downward gas-liquid flow was formed simply by supplying a proper amount of liquid from the top. The gas was then mixed in due to the gravity forces acting on the running liquid. Formation and Distribution of the Segmented Two-Phase Flow in the Monolithic Reactor. The experiments were carried out in a water-air system with a Plexiglas reactor of 25-cm diameter. The test section of the reactor consisted of two 20-cm monolith blocks stacked on top of each other. A uniform distribution of gas and liquid over the reador cross section is of major importance in all catalytic three-phase systems. In the present investigation, the slug flow was formed by showering the liquid with different types of perforated plates. The gas plugs (Taylor bubbles) were formed between liquid plugs by appropriate adjustment of the liquid loading. The perforated plates tested
Liquid
in
't' Gas in
Figure 3. Experimental setup for pressure drop measurements.
are listed in Table I. The flow distribution, pressure drop over the plates, and the superficial velocity in the monolith channels were measured during the test of the perforated plates. Pressure Drop. The pressure drop in the monolith channels was determined for different flow systems in a precision-drilled glass capillary tubing. The flow systems selected were water-air, ethanol-air, and ethylene glycol-air, with different physical properties (Table 11). The experimental setup for the pressure-drop study is shown in Figure 3. The equipment consisted of a vertical glass tube (2.2- or 1.1-mm internal diameter and 328 mm long) connected to a water-filled manometer at the lowest point. The gas and liquid slugs were formed in the inlet cell. A photocell enclosing the glass tube measured the length of the gas and liquid slugs. The static pressure drop was measured by switching the valve placed at the bottom. The pressure drop in the pilot plant and the full-scale reactor was measured indirectly by measuring the flow rate due to gravity. Reaction Conditions. The physical properties of the reaction mixture used were as follows. The hydrogen solubility and diffusivity in the liquid were 2.5 x 10" mol/(m3 Pa) and 1 X lo4 m2/s, respectively. The dynamic viscosity of the solution was 1.45 X lov3kg/(m s). Moreover, the density was 925 kg/m3 and the surface tension N/m. All data between hydrogen and liquid was 20 X presented above were measured at reaction conditions. The hydrogenations were carried out at a total pressure of 4 X IO5 Pa and a temperature of 50 "C. The gas-liquid
Ind. Eng. Chem. Res., Vol. 28, No. 10, 1989 1491 DOWNWARD
SLUG
r - 4
UPWARD
FLOW
BUBBLE
FLOW
L'""
GAS IN
-MONOLITH
m
-GAS
IN
LiQ
LIO.OUT
Figure 4. Schematic design of experimental setups for monolithic catalyst reactors (pilot plant).
flow through the monolith channels was either upward bubble flow or downward slug flow. The hydrogenations were performed in both the pilot plant reactor and the industrial reactor, and the scale-up volume ratio was 1:20. The experimental setups are shown schematically in Figure 4.
3. Theory Flow Description. As indicated earlier, several flow patterns have been observed in the vertical channels. These are frequently classified as bubble, slug, Taylor, and annular flows. Experimental flow pattern studies (Bilicki and Kestin, 1987; Moujaes and Dougall, 1987; Troniewski and Spisak, 1987; Brauner and Barnea, 1986; Khatib and Richardson, 1984; Barnea et al., 1982; Barnea et al., 1980; Taitel et al., 1980; Oliver and Hoon, 1968) have shown that any particular flow pattern in the vertical channels depends on the properties of the fluids, the flow rates, and the channel size. In monolithic catalyst reactors involving two-phase flow, the mass transfer is strongly dependent on the flow pattern prevailing in the monolith channels. In bubble flow, the gas flows as small bubbles in the continuous liquid phase. These bubbles have a diameter much smaller than that of the channels. The introduction of bubbles into the liquid phase causes an increased radial dispersion and mass transfer. In slug flow, the gas is located as large bubbles (Taylor bubbles) occupying the cross section of the channel and separated by liquid slugs. These gas bubbles have a length greater than the channel diameter. In this case, the radial mass transfer is increased due to an internal recirculation in the liquid slugs (Irandoust and Andersson, 1989). Between the Taylor bubbles and the channel wall, there is a thin liquid film. The thinness of this film favors mass transfer and reduces the axial dispersion. The transition from bubble flow to slug flow has been subjected to many investigations (Bilicki and Kestin, 1987; Barnea et al., 1982; Taitel et al., 1980). Although most of these studies are based on experiments with large-diameter tubes, they give a qualitative understanding of the phenomena involved in the transition between the flow patterns. Taitel et al. (1980) and recently Bilicki and Kestin (1987) have considered transition criteria for the transition from bubble flow to slug flow, among others. In bubble flow, the gas bubbles move and collide randomly if the bubbles are big enough to be deformed. The bubble void fraction at which this agglomeration happens is around 0.25-0.30. The coalescence of small bubbles into larger gas
slugs can also occur due to an acceleration of a certain bubbule in a wake behind the bubble ahead (Bilicki and Kestin, 1987). This wake is formed by a relative slip between the flowing bubbles and the adjacent liquid. However, the large gas slugs may be broken up if the liquid rate increases enough to cause large velocity gradients. Pressure Drop. The pressure drop in the two-phase flow is highly dependent on the type of flow obtained in the monolith channels. Several expressions for predicting the pressure drop in a slug flow have been developed in the literature (Orell and Rembrand, 1986; Khatib and Richardson, 1984; Satterfield and Ozel, 1977; Oliver and Hoon, 1968; Oliver and Wright, 1964). The total pressure drop in the Taylor flow can be analyzed on the basis of pressure gradients over the liquid slugs. This concept is attributable to the fact that the pressure gradient along an individual Taylor bubble is negligible. Thus, the total pressure drop is the sum of the frictional and hydrostatic pressure drops of the liquid slugs: Utot
=
ep, + @,t
(1)
The symbols introduced can be found in the Nomenclature section. A t low velocity and narrow channels, the frictional pressure drop is viscosity-dominant and can be calculated from (Bird et al., 1960)
Chemical Reactions and Mass Transfer. Hydrogen peroxide is generally produced from cyclic reduction and oxidation of anthraquinones (Berglin and Schoon, 1983). The most complicated step in this process is the hydrogenation of the working solution in a three-phase system. The influence of mass-transfer resistance in the observed reaction rate is of major interest in most liquid-phase hydrogenations. In the present investigation,the reactions occur in a narrow zone of the external surface of the catalytic channel walls (wash-coated surface). This will result in an easy calculation of the reactant concentrations at the catalyst wall. The modeling of mass transfer in slug flow is given by Irandoust and Andersson (1988b). The transport of hydrogen from the Taylor bubble to the reaction zone was considered to occur partly through the thin liquid film surrounding the cylindrical part of the Taylor bubble and partly through the ends of the Taylor bubbles into the liquid plugs. The former stands for the major part of the total mass transfer according to Hatziantoniou et al. (1986) and Irandoust and Andersson (1989). Thus, the very high interfacial surface area between phases in combination with the very thin liquid film will result in high performance of the monolith catalyst with slug flow. For the bubble flow, gas bubbles in the continuous liquid flow can be approximated by rigid spheres. The shape regimes of the bubbles are discussed by Clift et al. (1978). The mass transfer of hydrogen from the gas bubbles into the liquid phase can be estimated with dimensionless correlations (Clift et al., 1978). The dissolved hydrogen in the liquid will diffuse into the reaction zone in the catalytic wall. The mass transfer occurring at this step is complicated due to the unknown size distribution of the gas bubbles. 4. Results and Discussion Flow Visualization. The results of the flow pattern observations obtained in the glass capillary are summarued in Table 111. It was concluded that the Taylor flow is
1492 Ind. Eng. Chem. Res., Vol. 28, No. 10, 1989 Table 111. Flow Patterns Observed in the Glass Capillary liquid obsd loading, holdup, total superficial flow m3/(m2s) 9% velocitv. m / s Dattern 0.04 85 0.27 annular 0.10 75 0.36 Taylor 0.19 45 0.42 Taylor 0.24 43 0.42 Taylor 0.33 38 0.53 Taylor 0.35 20 0.44 Taylor 0.43 0 0.43
I1-
1.50
Eq.12) d t = i i mrr
-From 3
/"/o
0.501
I
d! -115 mrr
Q' +
Table IV. Total Superficial Velocity in the Monolith Channels gas total superficial liquid loading, m3/(m2s) holdup, 70 velocity, m/s 0.35 20 0.44 0.29 30 0.42 0.27 33 0.41 0.23 43 0.40 0.17 51 0.35
0.4
1.2
0.8 qt0t
x 106
2.0
Figure 5. Frictional pressure drop in a capillary: water-air slug flow. Eq.12) d t = i 2 mrr d t = 1.15 mrr
-From +
0
easily formed within a wide range of liquid loading (0.10-0.35 m3/(m2s)). The length of the gas and liquid plugs in the slug flow appeared to be appropriate. For liquid loadings exceeding 0.4 m3/(m2s), no gas was mixed in. On the other hand, the segmented gas-liquid flow tends to be unstable for liquid loading less than 0.1 m3/(m2s). In the industrial reactor, slug flow was obtained for a liquid loading between 0.12 and 0.33 m3/m2s. Segmented Two-Phase Flow in the Monolithic Reactor. In order to distribute the liquid phase uniformly over the cross section of the monolith block, the shower method with different types of perforated plates was used (Table I). A low pressure drop over the perforated plates used is desirable in order to minimize the required liquid pump energy. However, the liquid was poorly distributed when the pressure drop over the perforated plate fell below 100-200 mm of water. This resulted in a much thicker spray of the liquid over the monolith channels. A liquid loading of 0.2 m3/(m2s) necessitates a perforation below 10% to achieve the required pressure drop. Considering a certain perforation in a plate, a more uniform flow could be obtained for smaller holes. The same is also true for tighter triangular centers. However, there is a lowest limit of about 0.6 mm for the hole diameter to avoid problems of plugging. The superficial velocities obtained in the monolith channels for different gas and liquid loadings are listed in Table IV. It is shown that the total superficial velocity is about 0.4 m/s, independent of gas and liquid holdup in a wide flow range. In order to check the validity of this observation, the Navier-Stokes equations were solved with a finite difference program (PHOENICS by CHAM Ltd) based on body-fitted coordinates. The obtained total superficial velocity was 0.45 m/s, which is in good agreement with the experimental value. The liquid flowed downward due to gravity with a velocity of about 0.42 m/s. When liquid is added to the channel at a lower rate, gas will be entrained to make the total velocity 0.42 m/s. This means that slug flow will be formed in all channels that are hit by the spray. Pressure Drop. The pressure drop in Taylor flow arises from both frictional and hydrostatic effects. The frictional pressure drop is calculated by substracting the hydrostatic term from the total pressure drop (eq 1). Figures 5-7 illustrate the effect of the total flow rate on the frictional pressure drop for various two-phase systems in the slug flow regime. The frictional pressure drop was
1.6
"/s
+
qtor
Y
lo6, m3/s
Figure 6. Frictional pressure drop in a capillary: ethanol-air slug flow. 6r -From +
Eq ( 2 )
dt-22mm
o d, = l l 5 m m
q t o t x 10'.
m3is
Figure 7. Frictional pressure drop in a capillary: ethylene glycol-air slug flow.
adjusted for the liquid holdup (el) and the tube length. It should be noted that the observed pressure-drop values include the meniscus effects arising from the surface tension source. This contribution is not included in eq 2. Figures 5-7 indicate that the adjusted frictional pressure drop can satisfactorily be predicted by eq 2 within the accuracy of the measurements. This finding suggests that only the liquid controls the pressure-drop characteristics for the fluids tested. Under such conditions, for a fixed total flow rate, bubble flow is expected to give rise to a higher pressure drop in monoliths than slug flow does. In the pilot plant reactor, the pressure drop fits the theoretically expected value. The flow rate due to gravity was calculated to be 0.45 m/s and observed as 0.42 m/s. In the industrial reactor, the reaction rate was maintained up to a liquid flow loading of 0.33 m3/(m2s), which in-
Ind. Eng. Chem. Res., Vol. 28, No. 10, 1989 1493 dicates a similar total flow rate as in the pilot plant reador. Mass Transfer. The importance of flow type on the mass-transfer properties of gas-liquid-solid systems was studied by measuring the production capacity of the reactor system given as kg of H202 produced/(h m3 of bed). The production capacity observed in the case of downward slug flow was 30-5070 higher than that of upward bubble flow at the same total superficial velocity. For the case of slung flow, the reactor capacity observed was nearly constant for both pilot plant and industrial reactors, indicating the absence of scale-up effects. Since the concentration of anthraquinones is much higher than the hydrogen concentration, the higher mass transfer for slug flow compared to bubble flow is due to higher gas-solid mass transfer occurring directly from the Taylor bubble into the monolith wall. The very thin liquid film in this flow type gives a negligible mass-transfer resistance. This finding can also be due to uniform distribution of the gas and liquid phases for the slug flow type. The problems of nonuniformity are more pronounced in upward bubble flow and will lead to low efficiency of the monolithic catalyst and scale-up problems. Moreover, the low pressure drop in downward slug flow results in a higher and more even presure in the reactor. This will favor the rate of hydrogenation and selectivity. Furthermore, the production capacity in the case of slug flow was found to be essentially constant when the liquid flow loading was changed from 0.12 up to 0.33 m3/(m2s), which was the maximum possible liquid loading. This observation may be due to nearly constant total superficial velocity as observed earlier. No influence of the gas and liquid plug lengths was detected. 5. Conclusions Scaling up of monolithic three-phase reactors with downward slug flow is very straightforward. The same flow behavior was obtained in a single capillary, a pilot plant reactor, and an industrial reactor, and almost the same reaction rate/volume reactor was obtained in the industrial as in the pilot plant reactor. The only crucial part in the scaling up is the inlet flow distribution. Sieve plates, with holes somewhat smaller than the monolith channels and a certain pressure drop above 200 mm water across the plate, give a very good flow distribution. Upward bubble flow was more difficult to scale up. The radial distribution of gas bubbles in the continuous liquid flow was always uneven. This uneven distribution and the lower mass transfer in the bubble flow gave a 30-5070 lower conversion compared to slug flow in the largest reactor. The flow in the narrow monolith channels is viscositydominant, which gives an easily predictable flow behavior. Also annular flow and bubble flow could easily be avoided.
With a liquid flow above 0.1 m/s and a gas holdup above 25%, slug flow was always obtained.
Acknowledgment The National Energy Adminstration of Sweden provided financial support.
Nomenclature d, = tube diameter, m L = tube length, m APfr= frictional pressure drop, Pa AP,,= hydrostatic pressure drop, Pa APtot = total pressure drop, Pa qtot = total flow rate, m3/s v1 = liquid velocity, m/s Greek Letters c1
= liquid holdup
dynamic viscosity, liquid phase, kg/(s m) density, liquid phase, kg/m3 u1 = surface tension, liquid-air, N/m
pl = p1 =
Registry No. HzOz, 7722-84-1.
Literature Cited Barnea, D.; Shoham, 0.;Taitel, Y. Int. j . Multiphase Flow 1980,6, 387-397. Barnea, D.; Shoham, 0.; Taitel, Y. Chem. Eng. Sci. 1982, 37, 741-744. Berglin, T.; Herrmann, W. Sweden Patent 0102934 A3, 1982. Berglin, T.; Schoon, N. H. Ind. Eng. Chem. Process Des. Dev. 1983, 22, 150-153. Bilicki, Z.; Kestin, J. Int. J. Multiphase Flow 1987, 13, 283-294. Bird, R. B.; Stewart, W. E.; Lightfoot, E. N. Transport Phenomena; John Wiley & Sons: New York, 1960. Brauner, N.; Barnea, D. Chem. Eng. Sci. 1986, 41, 159-163. Clift, R.; Grace, J. R.; Weber, M. E. Bubbles, Drops, and Particles; Academic Press: New York, 1978. Hatziantoniou, V.; Andersson, B.; Schoon, N. H. Ind. Eng. Chem. Process Des. Dev. 1986,25, 964-970. Irandoust, S.; Andersson, B. Catal. Rev.-Sci. Eng. 1988a, 30, 341-392. Irandoust, S.; Andersson, B. Chem. Eng. Sci. 1988b,43,1983-1988. Irandoust, S.; Andersson, B. Comput. Chem. Eng. 1989,13,519-526. Khatib, Z.; Richardson, J. F. Chem. Eng. Res. Des. 1984,62,139-154. Moujaes, S.; Dougall, R. S. Can. J. Chem. Eng. 1987, 65, 705-715. Oliver, D. R.; Hoon, A. Y. Trans. Inst. Chem. Eng. 1968,46, T106T115. Oliver, D. R.; Wright S. J. Br. Chem. Eng. 1964, 9, 590-596. Orell, A.; Rembrand, R. Ind. Eng. Chem. Fundam. 1986,25,196-206. Satterfield, C. N.; Ozel, F. Ind. Eng. Chem. Fundam. 1977,115,6147. Taitel, Y.; Bornea, D.; Dukler, A. E. AIChE J. 1980, 26, 345-354. Troniewski, L.; Spisak, W. Int. J. Multiphase Flow 1987, 13, 257-260.
Received for reuiew November 29, 1988 Revised manuscript received June 5, 1989 Accepted June 28, 1989
