2404
Ind. Eng. Chem. Res. 1991,30,2404-2410
iented Technical Meeting; International Flame Research Foundation: Amsterdam, 1989. Jodal, M.; Nielsen, C.; Hulgaard, T.; Dam-Johansen, K. Pilot-Scale Experiments with Ammonia and Urea as Reductants in Selective Non-Catalytic Reduction of Nitric Oxide; Twenty-Third Symposium (International)on Combustion; The Combustion Institute: Pittsburgh, PA, 1990;pp 237-243. Kilpinen, P. Kinetic Modelling of Gas Phase Nitrogen Chemistry in Combustion. Thesis of Techn. Lic., Combustion Chemistry Research Group, Abo Akademi University, Report 90-13, 1990. Kilpinen, P.; Hupa, M. Homogeneous N20 Chemistry at Fluidized Bed Combustion Conditions: A Kinetic Modelling Study. Combust. Flame 1991,85,94-104. Leckner, B.; Golriz, M. R.; Zhang, W.; Anderson, B.-A,; Johnsson, F. Boundary Layers-First Measurements in the 12 MW CFB
Research Plant a t Chalmers University. The Eleuenth Znternational Conference on Fluidized Bed Combustion; ASME: Montreal, 1991;pp 771-776. Lyon, R. K. US Patent 3900554,1975. Miller, J. A.; Bowman, C. T. Mechanism and Modelling of Nitrogen Chemistry in Combustion. Prog. Energy Combust. Sci. 1989,15, 287-338. Mjornell, M.; Leckner, B.; Karlsson, M.; Lyngfelt, A. Emission Control with Additives in CFB Coal Combustion. The Eleventh International Conference on Fluidized Bed Combustion; ASME Montreal, 1991;pp 655-664.
Received for review February 13, 1991 Revised manuscript receiued July 10,1991 Accepted July 29,1991
Some Experimental Liquid Saturation Results in Fixed-Bed Reactors Operated under Elevated Pressure in Cocurrent Upflow and Downflow of the Gas and the Liquid Faigal Larachi, Andre Laurent,* Gabriel Wild, and No61 Midoux Laboratoire des Sciences du Ggnie Chimique, CNRS-ENSIC-INPL, BP451-I, Rue Grandville, 54001 Nancy Cgdex, France
The effect of pressure (0.3 IP/MF'a 5 5.1) on the total liquid saturation (measured by RTD method) of a fixed bed operated with cocurrent gas and liquid in both upflow and downflow is investigated. For upward flows, the liquid saturation is greater than for downward flows regardless of the operating pressure. However, in the pulsing flow regime and a t high gas velocities, the same asymptotic value is observed for both flow directions. The liquid saturation increases with pressure, mass flow rates being constant, but decreases when the liquid viscosity is decreased, independently of the flow direction and the operating pressure. For low gas velocities (uG < 1cm/s), the total liquid saturation no longer depends on the pressure. For higher gas velocities (uG > 1-2 cm/s) and nonfoaming liquids, the drift flux can provide an acceptable estimation technique of the liquid saturation if experiments under high pressure could not be conducted.
Introduction Historically, refining industries and petrochemical processes have widely used trickle-bed reactors, i.e., fixed beds with gas and liquid flowing cocurrently downward throughout a catalyst bed. Nevertheless, in many processes, such reactors, compared with fixed beds with an upward flow of gas and liquid (flooded-bed reactors), present drawbacks from the point of view of the heat transfer and the wetting efficiency. Indeed, with an upward flow of the two phases (both gas and liquid), under similar conditions, because of the high liquid saturations encountered for fixed beds with this flow configuration, risks due to hot spots appearance may be avoided. Moreover, the following other advantages of the floodedbed reactors over trickle beds have been described: better fractional wetting (Goto and Mabuchi, 1984), higher conversions (Snider and Perona, 1974),higher selectivity and catalyst life and less residue production as in selective hydrogenation of diolefin compounds (Ragaini and Tine, 1984), and better residence time and liquid distribution throughout the catalyst bed (Montagna and Shah, 1975). Actually, only a few papers in the literature deal with the experimentaldetermination of the liquid saturation under high-pressure conditions in fixed-bed reactors with twophase gas-liquid upward and downward flow. This paper presents original data on the liquid saturation (measured from tracer injection in the liquid phase) up to 5.1 MPa in a fixed bed operated in both upflow and downflow. We describe here successively the influence of the operating pressure, the gas and liquid throughputs, the flow direction, and the nature of the gas phase on the liquid
saturation. The drift flux concept, initially proposed in the modeling of bubble columns and fluidized beds, is applied here to high-pressure fixed-bed reactors; the drift flux used can be seen as the superficial velocity of the gas phase relative to a frame represented by both fluid phases flowing into the section offered to the gas. A more detailed description of this drift flux is given in the discussion of the results. Even though the type of reactor investigated here is quite different from the bubble column or from the three-phase fluidized bed, the drift flux concept proves to be a potent tool also in the investigation of cocurrent fixed-bed reactors. By using simple correlations of the drift flux obtained at atmospheric pressure, it is possible to get rough estimates of the liquid saturation at high pressures without carrying out high-pressure experiments, as long as the liquid does not exhibit foaming behavior and the gas superficial velocities are larger than 1-2 cm/s.
Brief Literature Survey In a recent paper (Larachi et al., 1991), the few references concerned with two-phase pressure drop and liquid saturation measurements in pressurized trickle-bed reactors have been cited: one has to mention the excellent experimental work of the team of Twente University (Wammes et al., 1990,1991). Unfortunately, there is even less research dealing with the hydrodynamics of fixed beds with upward two-phase flow under high pressure, and to our knowledge, only one published work has compared liquid saturation results in fixed beds with both upflow and downflow (Turpin and Huntington, 1967). These authors showed that, at least for nearly atmospheric con-
0888-5885/91/ 2630-2404$02.50/0 0 1991 American Chemical Society
Ind. Eng. Chem. Res., Vol. 30, No. 11, 1991 2405 ditions, liquid saturations (determined volumetrically) obtained with cocurrent upflow were always greater than those measured in cocurrent downflow, everything else being equal. The system they investigated was air/water in three different columns (5.1-, 10.2-, and 15.2-cm i.d.) packed with 7.6- and 8.23-mm alumina cylinders. Saada (1975) used also the air/water system in a 4.52-cm-i.d. and 40-cm-height column packed with 0.514-, 0.974-, and 2.064-mm glass ballotini in cocurrent upflow. Liquid saturations were determined at 1.36 MPa by use of a y-ray technique in the following range of gas and liquid superficial velocities: 0.004 m/s i UG 5 1.8 m/s and 0.001 m/s 5 uL 5 0.4 m/s. The author proposed dimensionless correlations for each flow regime, including inertial and viscous effects of both phases in terms of Reynolds numbers. Very recently van Gelder (1988) and van Gelder and Westerterp (1990) published results on cocurrent upflow liquid holdup (fraction of the total geometrical volume occupied by the liquid) and axial dispersion at very low gas and liquid superficial velocities (uG i 0.0145 m/s, 0.02 x m/s) for the system hym/s 5 u L i 0.151 X drogen/methanol; the measurements were carried out in a stainless steel column of 65-mm i.d. and 50-cm height, packed with 3.8 mm X 4.8 mm glass cylinders. The values of the liquid holdup at pressures between 0.2 and 1.2 MPa were obtained by using the imperfect pulse method. Following these authors, the liquid holdup seems to be independent of the mass flow rate of the gas, but seems to decrease when the gas velocity is increased. They correlated empirically their holdup results in terms of superficial gas and liquid velocities and suggested many equations. An extensive experimentalwork concerned with liquid holdup atmospheric measurements (determinedwith four different techniques: drainage, weighing, RTD, and y-rays) in cocurrent gas-liquid upflow reactors was presented by Yang (1989) and Yang et al. (1989); the authors studied the influence of several parameters such as fluid throughputs (uc5 0.15 m/s, 0.001 m/s i u L 50.03 m/s), particle size, column diameter, liquid viscosity, and foaming effect. They proposed to correlate their results with a simplified drift flux model (assuming a negligible slip velocity between the gas and the liquid). Their one-parameter correlation includes only the influence of the gas and the liquid superficial velocities. Recently, Gutsche and Gutsche et al. (1990) measured the hydrodynamic characteristics of cocurrent upflow reactors with large variations of the particle diameter (glass beads diameter 1.0, 1.4,3.1,6.0 mm) in two different columns (column diameter 5.0 and 10.0 cm). The technique used by these authors is the RTD method in the liquid phase. The correlation proposed by Yang (1989) describes correctly their data as long as they are concerned with moderate fluid throughputs. In this work, the pressure influence on the liquid saturation is emphasized: more than 180 data were obtained with several gas-liquid systems (water/ helium, water/ nitrogen, water/argon, water/carbon dioxide, and ethylene glycol/nitrogen) at pressures up to 5.1 MPa with upward and downward cocurrent gas-liquid flow through a packed bed.
Experimental Section The experiments are performed in a stainless steel column of 0.4-m height and 23-mm inner diameter. The reactor is packed with nonporous polypropylene cylinders (PE707T from Multibase) loaded with talc (loading ratio 70%). The characteristics of the packing and the critical surface tensions are listed in Table I. In order to study the influence of the surface state of the extrudates on the
Table I. Some Characteristics of the Packing Used 4: d,," mm* a,, m-l e, ?% p8,b kgm-3 a,, N m-l 0.910
3.37
1204' 1262d
38.5' 37.4d
1705
0.020' 0.071d
a Averaged from 90 observations with light microscope. Measured by helium pycnometry. ' Hydrophobic. Hydrophilized.
1 h
7 probe2 +
Signal processing
manometers
Figure 1. Flowsheet of the experimental plant.
liquid saturation, the reactor is first packed with fresh (hydrophobic) particles; then the same particles are treated chemically to render them hydrophilic. The surface chemical treatment is very close to that followed by Linek et al. (1974): the particles are immersed in a sulfochromic solution (810 cm3 of H2S04in 1000 cm3 of water and 170 g/L K2Cr207)and put into an autoclave at 70 "C during 15 min to allow the formation of a thin hydrophilic layer and to increase also the critical wetting surface tension up to 0.071 N/m (Linek et al., 1977). A detailed description of the installation is given elsewhere (Larachi et al., 1991). Figure 1shows a simplified flowsheet of the installation. It has been designed to withstand a maximum operating pressure of 10.0 MPa. The apparatus allows the measurement of the total liquid saturation (fraction of the porous volume held by the liquid) under high-pressure conditions (0.3 IP/MPa i 5.1) and ambient temperature (293-298 K), by using a tracer injection of salt solutions into the liquid phase (residence time distribution method). Two conductimetric probes located in the packed volume at each extremity of the column are used for this purpose (see Figure 1). The distance between the two cells is 360 mm. The liquid feed is provided from a reciprocating proportioning pump of a maximum volumetric liquid flow rate of 198 L/min. The gas is supplied from a rack of gas cylinders; its throughput is measured after reduction at atmospheric pressure by means of a membrane displacement device. A cyclonic separator situated under the bottom of the reactor is used
2406 Ind. Eng. Chem. Res., Vol. 30, No. 11, 1991 Table 11. Experimental Conditions Investigated in This Study direction packing system 11 tt treated nontreated H,O/N, X X X X H;O/He X X H,O/Ar X X X X HzO/COz X X X ETG/N2 Table 111. Physical P r o m r t i e s of t h e Fluids Used” surf. viscosity, tension, liauids densitv, k d m 3 Pa-s N/m lo00 0.001 0.072 water ethylene glycol 1107 0.0189 0.047 compressibility viscosity x 106,bPes gases factorb helium z =1 1.95 1.75 1.54 X 10-3P+ 5.31 X lO*p nitrogen z = 1 2.23 1.97 X 10-3P+ 11.7 X 1O”p argon z =1 (1-5.14)X 10-3P- 1.47 + 0.19 X 10-3P+ 88.5 X 1 O ” p carbon dioxide‘ 29.95 X 104P2 All the properties are given or evaluated at 293 K. Estimated to 30 bar. from the table data of l’Air Liquide (1976).
+ +
to separate the gas from the liquid. The column can be operated in upflow and in downflow. An RTI 800 data acquisition unit connected to a personal computer is used for the monitoring of the transient conductivity variations of the liquid after the tracer injections. The operating pressure is adjusted manually by means of back-pressure regulators until the gas and liquid flow rates become stationary. The different fluids experimented as well as the pressure ranges investigated are summarized in Table 11. Aqueous sodium chloride solution (40g/L) is used as tracer for the systems water/gases, whereas 40 g/L sodium perchlorate in ethylene glycol is used for the system ethylene glycol/ nitrogen. The physical properties of the liquids and the gases used in this study are listed in Table 111. The main advantage of this determination technique of the total liquid saturation 0 compared with other techniques such as weighing or drainage lies in the fact that it is used in the actual flow conditions and provides an on-line value of 8. The total liquid saturation can be calculated from the space time 7 and the liquid mass flow rate with the equation 0 = L7/cpLZ (1) The space time is determined by time domain nonlinear fitting of a least-squares objective function F using the axially dispersed plug flow model (PD model) with open boundaries:
F = X ( y ( ~ ) . ~-p x t ~ ( u ) e s P E -( t u ) du)’ k
(2)
Here y and x refer to the reduced outlet and inlet signals, while E refers to the impulse response of the liquid phase in the vessel. For the cocurrent gas-liquid downflow, the experiments are conducted essentially near the transition between the trickling and the pulsing flow regimes. The liquid saturations measured in upward flow lie in the bubbly dispersed flow regime as well as in the pulsing flow and the transition between them.
Analysis and Discussion of the Results The total liquid saturation j3 is measured in both upflow and downflow for ETG/N,/PPphil, water/Nz/PPphob, and water/Nz/PPphil,respectively. For the other systems,
P,MPa 0.3-5.1 0.35-2.1 0.3-2.1 0.3-2.1 0.3-5.1
L,kg m-2 s-l
C, kg m-2 0-1.5 0-0.42 0-1.62
6.6 13.2 6.6 13.2 6.6 13.2 6.6 13.2 6.6 13.2
0-1.70 0-1.5
a O9
9
P MPa
P
0.8
I A
11-11
0.7
2 1-11
I
0.6
51-i1 3
0.5 0.4
- - .-c
035-tt
I ’ I-tt 1 G:kdm2.s A
0
2
c
03
0.3
0
0.9
0.6
1 -
i
1.2
1.5
1.8
b
P P:MPa
uG:m/s
0.5 -
0
004
0 08
0.12
0.16
0.2
Figure 2. Influence of operating pressure on liquid saturation (L = 12.8 kg m-2 s-l) (a) vs C in cocurrent upflow of water and argon (open symbols) and cocurrent downflow of water and nitrogen (closed symbols); (b) vs uc in cocurrent upflow of ETG and nitrogen.
only the cocurrent gas-liquid upflow was examined (Table 11). Influence of the Operating Pressure and of the Flow Rates. Figure 2a shows the variation of the liquid saturation versus the gas mass flow rate at various pressures in upflow and downflow for the systems water/Ar (0.35 IP / W a 5 2.1 and water/Nz (0.3 IP/MPa I 5 . 1 ) , respectively. At a given pressure and liquid mass flow rate, the liquid saturation decreases when the gas mass flow rate is increased. An abrupt decrease is observed at low gas flow rates, followed by a slight decrease at higher gas flow rates. Such trends are confirmed regardless of the flow direction and the pressure. On the other hand, due to the increase of the gas density, the liquid saturation becomes much larger under high pressure at given fluid mass flow rates. In addition, since the gas velocity decreases (Gbeing constant), the mean residence time of the liquid phase increases necessarily because of the lesser shear at the gas-liquid interface, leading thus to an increase of the liquid fraction in the vessel; this confirms observations made in cocurrent downflow by Wammes et al. (1990, 1991). These authors reported a decrease of the liquid saturation with pressure when the gas superficial velocity is increased. As shown in Figure 2b, at s m d gas velocities (less than 1cm/s), the liquid saturation no longer depends upon the pressure. The same trend has been noticed for the trickle-bed reactors (Larachi et al., 1991). Therefore, for scaling-up purposes it does not seem necessary to carry
Ind. Eng. Chem. Res., Vol. 30, No. 11, 1991 2407 0.8 -
I - P
p
I 0 9 L ,'
PMPa
-t, 'c
1 G:kgJmZ.s
0.4 -
04 0
04
08
12
0
16
Figure 3. Influence of packing wetting on liquid saturation for water and nitrogen flow through treated (closed symbols) and untreated (open symbols) particles (L= 12.8 kg m-2 8% PMR
' 1151 Tt5l
10
I 0.1 o.ooo1
1103
UG: ny's Y
0.01
1
Figure 4. Effect of flow direction on liquid saturation vs UG for both upward and downward flows at 0.3 and 5.1 MPa of the system water and nitrogen.
out experiments under high pressure to predict liquid saturation at low gas superficial velocities. A determination at atmospheric pressure at the same value of U G is sufficient to predict what will happen when the reactor is operated under high pressure. As far as we know, we are the first to report this phenomenon. Influence of the Wettability of the Packing. Figure 3 illustrates the influence of hydrophobic and hydrophilic behavior of the packing on the liquid saturation for water/N2 a t 0.3 MPa upward and 1.1 MPa downward, respectively. As can be seen, there is no apparent effect of the packing wetting phenomenon on the liquid saturation. In fact, it seems that the original hydrophobic particles are progressively hydrophilized by the presence of impurities in the installation (such as rust). Indeed, after several days of operation we emptied the reactor and tested the particles; these had become less hydrophobic and were covered by a thin layer of fouling. Linek et al. (1974) showed that mass-transfer performances (interfacial gasliquid area and volumetric mass transfer coefficients) in cross-flow packed-bed reactors with polyethylene and polypropylene hydrophilized plastics were greatly improved in comparison with nontreated particles. Considering our results, we are convinced that previous treatment of the packing is not a workable mean of improving the quality of industrial packings. Influence of the Flow Direction. As illustrated by Figure 4 for water/nitrogen at 0.3 and 5.1 MPa, the total liquid saturation obtained in upflow is always larger than that measured under the same conditions in downflow regardless of the operating pressure. Turpin and Huntington (1967) have already noticed such a behavior at atmospheric pressure for the system air/water. The figure shows also that, at moderate pressures (S0.3MPa) and high gas superficial velocities, the liquid saturation be-
0.4
0.6
0.8
1
1.2
1.4
Figure 5. Liquid viscosity effect on liquid saturation for water, nitrogen, and ethylene glycol; nitrogen in cocurrent downflow, P = 2.1 MPa (square symbols), and cocurrent upflow, P = 0.3 MPa (diamond symbols).
j Yo
' - p
0.2
comes independent of the flow direction. Such an asymptotic behavior was pointed out by Yang (19891, who showed the hydrodynamic equivalence between the pulsing flow regimes (which is essentially governed by the fluid inertia; (Larachi et al., 1990, 1991)) observed for both configurations (upflow and downflow): these authors compared two-phase upflow pressure drop measurements obtained in the pulse flow with the correlation established by Ellman et al. (1988) for the high interaction regime in trickle beds. However, the same trend is not observed under high-pressure conditions where the two curves are distinct. It appears that the flow regimes at 0.3 and 5.1 MPa are not the same. At small gas velocities and a pressure of 5.1 MPa, the flow regime is bubble flow in upflow and trickling flow in downflow; however, at 0.3 MPa and higher gas superficial velocities (110 cms-'), the flow regime will certainly be pulsing flow in both flow directions. This behavior is not surprising: in pulsing flow, the hydrodynamic characteristics are mainly governed by inertia and the flow direction does not play any role. At lower velocities, that is, in trickling (downflow) or bubble flow (upflow), gravity (and buoyancy) play an important role. Influence of the Liquid Viscosity. The effect of the liquid viscosity on the total liquid saturation is illustrated by Figure 5, which compares the liquid saturation obtained with water/N2 and ETG/N2 at 2.1 MPa (downflow) and 0.3 MPa (upflow). Regardless of the flow direction, the liquid saturation is an increasing function of the liquid viscosity. Yang et al. (1989) did not mention any significant effect of this physical parameter on the liquid saturation for cocurrent upflow. However, the most viscous liquid they used was a 4 mPa-s liquid. Influence of the Nature of the Gas Phase. Figure 6a shows the total liquid saturation versus G obtained in cocurrent upflow at 2.1 MPa with four different gases (helium, nitrogen, argon, and carbon dioxide) and water as liquid. As in Figure 2a, the higher the gas molar weight (or the gas density at a given pressure), the larger is the liquid fraction at given pressure and fluid mass flow rates. The influence of the gas dynamic viscosity on the liquid saturation is not clear since it varies very slightly with pressure (at least far from the critical conditions). As shown in Figure 6b, when helium and nitrogen are used under different pressures but with same densities (nitrogen at 0.3 MPa, helium at 2.1 MPa), the liquid saturation seems to be independent of the nature of the gas. I t is therefore highly probable that the gas phase participates only by its inertia and that gas viscosity does not play any role in the hydrodynamics of fixed beds. In order to confirm this statement, measurements should be carried out at different temperatures with different gases chosen
2408 Ind. Eng. Chem. Res., Vol. 30, No. 11, 1991 0.9 r
B
a
Table IV. Drift Flux Correlation of the Total Liquid Saturation Results' A Darticle diam for aas/liauid syst d, = 1.4 mm,* glass beads propylene carbonate/nitrogen (11) 0.53 d, = 2.0 mm,* glass beads water/nitrogen (11) 0.41 propylene carbonate/helium (11) 0.43 propylene carbonate/nitrogen (11) 0.40 ethanol/nitrogen (11) 0.41 water + 40% sucrose/nitrogen (14) 0.47 ethylene glycol/nitrogen (11) 0.51 d, = 3.3 mm, nonporous polypropylene extrudate 0.41 water/helium (17 ) water/nitrogen (11) 0.41 water/nitrogen (1t) 0.42 water/argon (tt) 0.49 water/carbon dioxide (11) 0.41 0.45 ethylene glycol/nitrogen (4 1) ethylene glycol/nitrogen ( 7 1) 0.54 d, = 2.8 mm: porous cylinders 0.77 cyclohexane/air (t 1) d, = 1.0 mm: glass beads water/nitrogen (t t) 0.36 d, = 2.0 mm: glass beads 0.80 water/nitrogen ( 7 t fluidization)
Gas
0.8
1
G:kg/mZ.s
0.5
0 0.9
0.4
0.8
1.2
rp
2
1.6
b
PMPa
06 O": -
05
\
+
8
He-2 1
C
N2-03
,
b
1.04
1.0 0.96 0.95 0.95 0.99 0.94 0.92 1.0 0.94 0.97 0.94 0.91 1.0 1.07 0.81 1.0
1 -
O b is dimensionless; the dimension of A is (md)'-*. *From
Larachi et al. (1991). eFrom Yang et al. (1989). dFrom Gutsche (1990). eFrom Nacef et al. (1989).
_r
G kg/m2 s
04
0
02
04
06
08
I
Figure 6. Influence of nature of gas phase on water saturation in (a) cocurrent upflow at P = 2.1 MPa; (b) cocurrent upflow at P = 0.3 MPa of nitrogen and 2.1 MPa of helium.
in such a way that the gas density is the same in all cases. If that statement proves true, it is therefore possible to simulate the hydrodynamics of a hydrogenation reactor operated at high pressure (20 MPa, for example) by measurements with other gases at lower pressures (helium at 10 MPa or argon at 1.0 MPa if compressibility effects neglected, in the example considered). On the other hand, if one compares results obtained with different gases at the same pressure and given mass flow rates, hydrogen will lead to smaller values of the liquid saturation than any other gas. Drift Flux Concept and Its Application to Cocurrent Gas-Liquid High-pressureFixed-Bed Reactors. The so-called drift flux concept was first used by Wallis (1969) to characterize the hydrodynamics of bubble columns. It was extended by Darton and Harrison (1975) to three-phase cocurrent upflow fluidized beds to correlate their holdup results. Later, it was extensively used by our team (Saberian-Broudjenni et al., 1984, 1987; Wild et al., 1987,1989 Nacef et al., 1988) and also by other teams (e.g., Nikov et al., 1990). Various expressions of the drift flux are found in literature; we choose here the same definition as given by Saberian-Broudjenni et al. (1987): (3)
This is equivalent to p=1+
JD-F
G
- G/PG
-PG+ - PLL
(4)
The drift flux JD.F thus defined can be seen as the velocity of the gas phase relative to a frame represented by both fluid phases flowing into the section offered to the gas. As will be seen later, the drift flux can provide an interesting means to estimate liquid saturations in fixed-
bed reactors, regardless of the operating pressure and the liquid superficial velocity as well as the direction of the flow. Saberian-Broudjenniet al. (1984,1987), Wild et al. (1987, 1989), and Nacef et al. (1988) showed that, under certain conditions, the drift flux can represent an interesting scale-up criterion for gas-liquid cocurrent upflow fluidized beds: provided the liquid is coalescing and the flow regime is the heterogeneous flow regime, the drift flux does not depend on the liquid velocity, or on the viscosity, or on the particle diameter or density and is only a function of the gas flow rate for a given apparatus. As shown by Nacef et al. (1988), the drift flux concept can represent an efficient way to account for the nature of the distributor used and for scale effects. In general, empirical correlations like J D - F = AuGbpfovide a very accurate way to describe the liquid saturation (or holdup) results. The exponent 6 is always close to 1, whereas the factor A is a characteristic of the quality of the distribution of both gas and liquid and is in the order of 0.5-0.9. In fixed beds, A must be necessarily smaller than in fluidized beds. Results obtained by Nacef et al. (1988) in fluidized beds, by Yang et al. (1989) and Gutsche (1990) in fixed beds in cocurrent upflow, and by ourselves for cocurrent upflow and downflow through fixed beds confirms the lower value of A for fixed beds (see Table IV). This fact can be ascribed to the better distribution of the fluids occurring in porous media than in the classical distributors used in fluidization (grids, perforated toruses, or plates) and to the higher mechanical energy dissipation rate (pressure drop) in fmed beds. Table IV includes some of our data (about 800) previously published on cocurrent downflow liquid saturation in a 1.0-m-height pressurized trickle bed (Larachi et al., 1991). As can be seen, the drift flux does not depend on the gas and liquid properties (density, liquid surface tension, and viscosity) or on the particle diameter. In fact, even if there is an influence of the fluid properties on the drift flux, the experimental accuracy is not sufficient to display it. The values of A and 6 are respectively encompassed within 0.40-0.54 and 0.91-1.07 for our data. Figure 7a (water/nitrogen in cocunent downflow at 2.1 MPa through 2-mm glass beads) shows that J D - F is only
[nd. Eng. Chem. Res., Vol. 30, No. 11, 1991 2409 0.I
I
J D-Fm/s
f
uG ds
L:kghn2.s A
0.1
5.8
m$PAA
0 7.1
7.8
0.01
0 9.6 A
0.001
A 15.7
x M.3 24.5
t 1
0.001 0.01
I
__
uG:m/s 0.1
1
b
ID-Fds
1 0.001 0.01
UG:ds
1
0.1
-
Figure 7. Variation of drift flux vs gas superficial velocity ( L 13 kg mV2s-l) (a) at 2.1 MPa and different liquid flow rates for water and nitrogen in cocurrent downflow (glass beads, 2 mm), results of Larachi et al. (1991); (b) in cocurrent upflow of water and argon (0.35 5 P/MPa I2.1),this study, and cocurrent downflow of PC and nitrogen (0.3IPIMPa I5.1,glass beads, 2 mm), results of Larachi et al. (1991).
affected by the gas superficial velocity; provided the latter is larger than 1-2 cm/s (for UG < 1cm/s, the data are not plotted), the drift flux does not depend on the liquid velocity. As illustrated in Figure 7b (propylene carbonate/ nitrogen in downflow through 2-mm glass beads, P from 0.3 to 5.1 MPa, and water/argon in cocurrent upflow, P from 0.35 to 2.1 MPa), the drift flux does not depend on pressure. Such a trend was also noticed when low surface tension (such as ethanol) or high viscosity (such as ethylene glycol) liquids were used. The conclusions relative to downflow are still valid for the two-phase upflow. If the drift flux approach is a useful tool to estimate the total liquid saturation for nonfoaming liquids, this is no longer true with coalescence-inhibiting liquids as shown in Figure 8 for the system water 1% ethanol/nitrogen: at a given pressure, the drift flux is a function of the liquid superficial velocity, regardless of the flow regime occurring in the reactor. The same conclusion may be drawn for the pressure effect: in this case (coalescence-inhibitingliquid), at a given liquid superficial velocity, the drift flux depends on the operating pressure. This is not surprising, since the liquid saturations encountered are smaller than for a nonfoaming liquid, and the drift flux is rather weak (JD.F < 2 cm/s). Hence, for scale-up purposes, the designer can obtain a rough estimate of the liquid saturation under vigorous conditions (high pressure) without carrying out experiments there, as long as the liquid is not coalescence inhibiting. It is possible to establish something like a calibration equation of the type JDF= AuG* using experiments conducted at atmospheric pressure; provided the flow regimes are similar, the liquid saturation at high pressure can then be calculated from (4).
+
Conclusions A comparative study of the total liquid saturation in a fixed-bed reactor with a cocurrent upflow and downflow
*416
O
DAkt
.0
11.6
A
Am
Y O
m
A+
021
M A 0
-0.01
, J D-F ds
A0
o.ooo1 -0.005
0
O.Oo5
0.01
0.015
0.02
Figure 8. Variation of drift flux vs gas superficial velocity for different liquid flow rates, P = 5.1 MPa (closed symbols) and different operating pressures, L = 13.2 kg m-2 s-l (open symbols) for coalescence-inhibiting system water + 1% ethanol and nitrogen in cocurrent downflow (glass beads, 2 mm), results of Larachi et al. (1991).
of gas and liquid under high pressure is presented. The influence of the following parameters was examined pressure (0.3 5 P/MPa I 5.1), flow direction, fluid throughputs, nature of the gas phase, and wettability of the packing. A successful attempt was made to apply the drift flux concept to fixed beds in cocurrent upflow and downflow under high pressure. The following conclusions were drawn: For upward flow, the liquid saturation is greater than for downward flow regardless of the operating pressure. However, in the pulsing flow regime and at high gas velocities, the same asymptotic value is observed for both flow directions. The influence of pressure on the liquid saturation is exclusively due to the inertia of the gas phase (gas density). Nevertheless, high-temperature hydrodynamic measurements will be necessary to clarify whether the dynamic gas viscosity is to be included or not in dimensionless liquid saturation correlations. The liquid saturation increases with pressure, mass flow rates being constant, but decreases when the liquid viscosity is decreased, independently of the flow direction and the operating pressure. For gas velocities UG > 1-2 cm/s and nonfoaming liquids, the drift flux can provide an acceptable estimation technique of the liquid saturation if experiments under high pressure could not be conducted. This approach seems to agree also with results of other literature sources obtained under atmospheric pressure. One must be careful in using these results: all data presented here were obtained in a small-scale column; experiments in larger reactors are needed to confirm these conclusions. Acknowledgment We are grateful for the financial support for this work received from the Institut Fransais du PBtrole and the Algerian MinistGre aux UniversitBs. Nomenclature a,: specific area of the bed, m-l d,: volume equivalent diameter, mm d : particle diameter, mm f i t ) : impulse response of the liquid phase, s-l G: gas mass superficial flow rate, kg mw2s-l JD-F: drift flux, m s-l L: liquid mass superficial flow rate, kg m-2 s-l P: operating pressure, bar, MPa u: superficial velocity, m s-l t : time, s x : reduced inlet signal, s-l y: reduced outlet signal, s-l
2410 Ind. Eng. Chem. Res., Vol. 30,No. 11, 1991 z: compressibility factor 2 packed height, m 8: total liquid saturation, volume of liquid/porous volume c: bed porosity, % 9,: shape factor p: density, kg m-3 u,: critical surface tension, N m-l 7: space time, s Subscripts G: gas L: liquid s: solid c: critical Acronyms ETG: ethylene glycol PC: propylene carbonate PPphil: hydrophilized polypropylene PPphob: hydrophobic polypropylene RTD: residence time distribution tt: cocurrent upflow 11: cocurrent downflow
Registry No. H 2 0 , 7732-18-5; He, 7440-59-7;N2, 7727-37-9; Ar, 7440-37-1;COz, 124-38-9;ethylene glycol, 107-21-1.
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Accepted June 20,1991
