Ind. Eng. Chem. Res. 199433,667-675
667
Simultaneous SO%/NOSeparation from Flue Gas in a Contained Liquid Membrane Permeator S. Majumdar,? A. Sengupta? J. S. Cha,t and K. K. Sirkarst Department of Chemistry and Chemical Engineering, Stevens Institute of Technology, Castle Point, Hoboken, New Jersey 07030
Separation of SO2 and NO from a simulated gas mixture using the hollow fiber contained liquid membrane (HFCLM) technique has been explored. Separation of SO2 from a gas mixture containing SOz, COZ,02,and Nz has been studied with pure water and aqueous 1N NaHS03solution. Combined separation of SO2 and NO from a simulated flue gas mixture of SOZ, NO, COZ, 0 2 , and Nz has been achieved utilizing aqueous FeZ+EDTA or Fe3+EDTAsolutions. In both cases, sweep and vacuum modes of operation have been investigated. The HFCLM technique has been applied successfully for flue gas cleanup a t 24 "C; it was also found to be quite effective a t 70 "C. A simplified multicomponent gas permeation model based on solution diffusion through the liquid membrane appears to describe the SO2 separation as well as C02 separation through the water membrane satisfactorily for the sweep mode of permeator operation.
Introduction Abatement technologiesfor oxides of sulfur and nitrogen present in flue and stack gases from coal-fired boilers are becoming increasingly important. A typical untreated flue gas contains 0.1-0.5% S02,0.01-0.05% NO, 10-1576 COZ, 10-15% H20, 1-596 02, and 70-75% Nz. Scrubbing it with an aqueous limestone slurry to remove SO2is a widely used treatment process (Walker et al., 1985; Drummond and Gyorke, 1986). These scrubbing solutions are, however, not very effective in removing NO. Further, the process is expensive and produces an enormous amount of sludge. Advanced separation technologies are being developed to effect SOz-NO, removal from flue gas (Walker et al., 19851, to improve separation efficiency, and to reduce the treatment cost. Membrane gas separation is one of the technologies under development. Membrane processes are simple, are often less expensive, and are modular in nature for ease of scaleup or flexibility. However, polymeric membranes do not currently exhibit the requisite combination of high flux and selectivities. Moreover, due to the small concentration of SO2 and NO, in flue gases, the membrane permeation flux, governed directly by the transmembrane partial pressure driving force, would naturally be low. Liquid membranes may be, therefore, better candidates since diffusivities of gaseous species through liquids are generally orders of magnitude larger than those through polymeric solids. Furthermore, facilitating agents can be added to liquid membranes to enhance the flux of certain species significantly generating high selectivity between the permeant(s) of interest (here SO2 and NO), and the other species (COz,02,and N2). In addition, the membrane process not only removes SO2 and NO, from the flue gas but also can concentrate simultaneously these species in the product side for easier disposal and/or possible further use. Liquid membranes have been studied in immobilized liquidmembrane (ILM) form (Ward and Robb, 1967;Otto
* To whom correspondence should be addressed. + Current address: Department of Chemical Engineering, Chemistry and Environmental Science, New Jersey Institute of Technology, Newark, NJ 07102. t Current address: Separations Products Division, Hoechst Celanese Corporation, Charlotte, NC 28273.
0888-5885/94/2633-0667$04.50/0
and Quinn, 1971; Bhave and Sirkar, 1986; Hughes et al., 1986; Meldon et al., 1986; Way and Noble, 1992) for gas separations. The first study on SO2 separation via an ILM was carried out by Ward and Neulander (1970). However, ILMs are generally unstable, have low operating life, and cannot be regenerated in situ (Kimuraand Walmet, 1980; Matson et al., 1983; Sengupta and Sirkar, 1986). A different liquid membrane structure called the hollow fiber contained liquid membrane (HFCLM) has been used recently for separation of gas mixtures (Majumdar et al., 1988, 1989; Guha et al., 1991, 1992) and for liquid separations (Sengupta et al., 1988a,b; Basu and Sirkar, 1991);it eliminates most of the deficiencies of traditional ILMs. Guha et al. (1992) have demonstrated that the HFCLM system can be operated successfully in different operating modes: sweep gas, sweep liquid, permeate side vacuum, and conventional permeation. We explore here S02-NO, separation from flue gas using different liquid membranes in hollow fiber contained liquid membrane permeators using either permeate side vacuum or a sweep gas. Sengupta et al. (1990) have studied the permeation behavior of S02, COZ,02,and NZpresent in a typical flue gas through various liquid membranes. Excellent permeabilities of SO2 and very high selectivities between SOzCO2 (75-200) and SOz-NZ (900-5000) were obtained using pure water and aqueous solutions of various inorganic salts, e.g., NaHS03, Na2SO3, Fe2+EDTA,and Fe3+EDTA. These values are considerably higher than the polymeric membrane selectivities (SOz-COZ selectivity of 10 and s 0 2 - N ~ selectivity of 250) obtained by Baker et al. (1988). In the present studies, water and aqueous solutions of Fe2+EDTA, Fe3+EDTA, and NaHSOa have been used as liquid membranes in HFCLM permeators. For a pure water membrane, the following principal ionization reaction occurs with SO2 (Roberts and Friedlander, 1980): SO2+ 2H20 = HSO;
+ H30+
On the other hand, the permeation of NO through an aqueous solution of Fe2+EDTA is facilitated by the complex formed by the reaction (complexation of NO with Fe3+EDTA is, however, questionable),
NO + Fe2+EDTA= Fe2+EDTA.N0 Separation studies were carried out to remove SO2 from 0 1994 American Chemical Society
668 Ind.
Eng. Chem. Res., Vol. 33, No. 3,1994 CONTAINED UPUlD UEMBRANE
MICROPOROUS SUBSTRATE HOLLOW FIBER W&U )
PERUEATE SIDE
I 4> >4Ji Feed
Feed
E J J
reed
PhU4
".O""r"
SIeeP
rml
weep
Sweep Mode of
Operation
111
Vacuum Mode
LlPUlD MEMBWNE PERMEATOR
of
Operation
THIS END MAY BE C W S E D FOR VACUUM RUNS
F i r e 1. Hollow fiber mntained liquid membrane structure and two modes of operation.
Fimre 3. Material balance in CLM permeator with sweep ga8 or with vacuum.
a mixture of S0+2O&rNz and to eliminate SO2 and NO simultaneously from a mixture of S O r N O - C O r O r Nz. Most experiments were conducted at 25 "C, but a few werealsocarriedout at 70 oCconsistentwiththeconditions existing in current aqueous solution-based scrubbers.
simplifyingassumptions are made. The main hypothesis here is that the overall mass-transfer resistance ean be expressed as the s u m of individual resistances in series as shown in Figure 2. The overall mass-transfer coefficient based on the fiber outside diameter, K , for a permeant can therefore he expressed as
The Contained Liquid Membrane: Concept and Operation In the HFCLM technique for flue gas separation illustrated in Figure 1, a thin aqueous liquid layer functioning as the membrane is kept between two sets of microporous hydrophobic hollow fibers tightly packed in a permeator shell (Majumdar et al., 1988). One set of fibers carries the flue gas, and the other set carries the sweep and/or the permeated gases. The aqueous membrane liquid is connected to an external membrane liquid reservoir which automatically replenishes any lost liquid. This structure eliminates the problems of operational instability,membranedryingorflding, andlow operating life encountered inconventionalILMprocesses (Majumdar et al., 1988). For flue gas purification, it is desirable to withdraw SO? through the liquid membrane as a concentrated permeate streamasthefluegasiscleaned.Inaneconomicevaluation of membrane gas separation processes for SO2 removal, Walker et al. (1985) have indicated that operation of a permeator using a condensable vapor as a sweep gas on the permeate side was quite economical. Operating the permeate side under vacuum was found by them to be economically only slightly inferior to the sweep mode of operation. In this study, therefore, both the sweep mode and the vacuum mode were adopted (Figure 1). For simplicity, helium was chosen as the sweep gas instead of a condensable vapor. Mass-Transfer Rates in HFCLM: Fluxes and Removal Firsborder estimates of the rate of mass transfer in HFCLMpermeatonrcan bedeveloped theoreticallyifsome
4 4 1 do 1 h k f , km ~ L M ~ . Q dik:
_ 1= _ do_1 +--+-+--+-1 1
K
di k i
(1)
The permeation flux, N,at any point is given by
N = [K/RTI(Px -PY)
(2)
Although eq 1is applicable a t any c r w section inside the permeator, we assume that it can be extended to the whole module with average values for the individualCoefficients. In general, the transfer resistances of the microporous hollowfiberwallsarepracticallynegligible. Thesubstratetransfer coefficients are typically an order of magnitude higher than the liquid membrane transfer coefficient (which is defined as the ratio of the diffusivity of the permeant to the effective liquid membrane thickness). Further, if feed and sweep (or permeate, for vacuum runs) flow rates are known, one can estimate the film-transfer coefficientsusing availablecorrelations. In membrane gas permeation, the gas film resistance is usually neglected. The CLM permeator module is schematically shown in Figure 3 along with the relevant quantities. For sweep runs, the component (e.g., SOZ)material balance can be expressed as
Lprr - L r x w = "fYf
(3)
Thefractionalremovalofaspeeies,F,anditsexpehental overall mass-transfer coefficient, Kexpt,are expressed respectively by
F = (Lprx,- LrxW)/Lprf and
(4)
Ind. Eng. Chem. Res., Vol. 33, No. 3. 1994 669
RT = [K,,,JRTIATAPLM
(5)
For runs in sweep mode, RT and Apm can be expressed as follows:
RT = V~YI A~LM =
(6)
(Prrr- PIYI) - (P& - P A ) I n [ ( P ~ t - p t y I ) / ( P ~-wpaw)]
(7)
Sincepressures, flow rates, and compositionsareall known experimentally, one can calculate ICexpt. Note that Kelp, can be calculated in this fashion only for sweep runs and not for vacuum runs since for vacuum runs ywvalues are generally unknown.
i
MICROPOR0U'HOLLOW FIRERS
Model for Multicomponent Gas Permeation in HFCLM Permeator The general model for multicomponent gas permeation in a HFCLM permeator without any chemical reaction is available in Majumdar et el. (1989) when permeation through liquid membrane controls. Following their prccedure, the permeation equations and the boundary conditions for the present case can be written in dimensionless form as n
dL*ldS = C Q j * [ P x j - p*yj]
(8)
1-1 I
dV*/dS = z Q j * [ P r ~ , p*yjl
(9)
1=1
" dxi/ds = [Qi*(Pxi - p*yi) - x i C Q j * P x j - p*yj)I/L* 1-1
i = 1,n (10)
n
dyildS = [Q;*(P*x;
- p*yi) - ~ i z Q j * ( P-~p*~j)l/V* j
CELL
Figure 4. SO?permeation in a unit cell of the CLM.
nonlinear second-order differential equations over curved boundaries of the unit cell (Majumdar et al., 1992). Note alsothat the localfeedandsweepcompositions willchange along the module (in the r-direction). Besides, there are other possible feed-sweep fiber arrangements (Majumdar et al., 1989). The exact numerical solution appears to be extremely complex. For the immediate objective of modeling the transport behavior, an effective membrane thickness is considered. The best way of estimating the effective liquid membrane thickness is by independent HFCLM experiments. The permeation of a pure nonreacting gas such as N2 or C02 through a pure water membrane is normally carried out. The liquid membrane thickness can be calculated from the measured permeation rate (Majumdar et al., 1988). The membrane thickness estimated this way may be used in calculations involving regular separation experiments using the same module, since the module property remains constant from run to run.
1=1
i = 1,n (11) d P l d S = &*L*/P
(12)
dp*ldS
(13)
-&*V*/p*
v* = v; p =pw*;y i = yiw i = 1,n (14) S = 1 L* = LI*; p* = pr*; xi = xif i = 1,n (15)
s = 0-
-
The above equations have been solved simultaneously for five components (e.g., S02, C02, N2, 02, and He). The solution procedure is similar to that described earlier (Majumdar et al., 1988; 1989). Determination of Effective Membrane Thickness One possible fiber arrangement inside the CLM permeator shell and a small section of the CLM between the fibers are shown in Figure 4. A t any crms section of the module, there would be hundreds of such unit cells (Majumdar et al., 1988). A true permeant concentration profile can only be obtained by solving the diffusion reaction equations inside such unit cells. Even if only one principal reaction is considered (in reality, there would be a number of side reactions, (Roberts, 1979)), one has to solve a boundary value problem for two simultaneous
Experimental Section Permeator Modules. The HFCLM permeator modules used for flue gas cleanup experiments were made using Celgard X-10 microporous hydrophobic polypropylene hollow fibers obtained from Hoecbst Celanese (Charlotte, NC). Fabrication of a permeator involved preparation of a fiber bundle from two sets of fibers of known length and number, insertion of the bundle in a shell, and finally potting the different ends of each fiber set with a resin mixture to form a tube sheet. The fabrication procedure was essentially similar to that described in the literature (Majumdar, 1986; Guha, 1989). The geometrical characteristics of hollow fiber modules used in this study are provided inTable 1. Tbenumber of fibers in the feed side was always equal to that in the permeate side. Experimental Apparatus. The apparatus used for HFCLM studies consisted of three different segments: feed gas line, permeate or sweep gas line, and membrane liquid line. The setup is described schematically in Figure 5. The required feed gas composition was achieved by mixing different gas mixtures. The pressure of the feed gas was measured after it was humidified in a stainless steel vessel. The feedgascomposition after humidification was measured under steady flow conditions before the line was connected to the module. The purified feed gas stream from the permeator was passed successively
670 Ind. Eng. Chem. Res., Vol. 33, No. 3, 1994 Table 1. Characteristics fiber diameter effective insidel module length, outside, no. cm cmxW C 157.5 100/150 E 43.2 100/150 F 44.5 100/150 G 63.5 240/290
of Permeator Modules no. of fibers feed/ sweep 300/300 300/300 300/300 120/120
mass- effective arean transfer membrane area, thickness, vortAe, cm* cm cmZ/cma 0.0258 48.5 2227 610 0.0123 48.5 628 0.0719 48.5 694 0.0175 37.5
HEWUM
rd&, rclb
PRESSURE
FEED
.
0 Area per volume indicates the ratio of active membrane surface area to equipment volume.
OUT
through a water (or liquid) separator to remove liquid entrained by the gas stream and a dryer. Helium was used as a sweep gas in the sweep mode of operation. It was also fed to the module after humidification. However, humidification of the gas streams was avoided altogether for experiments with water membranes. The exit pressures of the feed and the sweep gas streams were maintained close to the atmospheric. The membrane liquid storage vessel connected to the module shell side was pressurized by a helium gas cylinder. This pressure was always kept higher than those of the feed or the sweep gas stream. Both gas streams were dried using Nafion membrane dryers (Permapure Products, Farmingdale, NJ) before analysis. The concentration of NO was determined by infrared analyzers (Mine Safety Appliances, Pittsburgh, PA) when applicable. The concentrations of other components were measured periodically either by a Hewlett Packard Gas Chromatograph (Model 5890) using a twocolumn system based on a 6 f t X l/s in. chromosorb 108 column (Chrompack, Inc., Bridgewater, NJ) and a 10 f t X l / 8 in. molecular sieve column (Alltech Associates, Deerfield, IL) or by a Varian 3700 Gas Chromatograph equipped with a 6 f t X l / 4 in. CTR column (Alltech Associates, Deerfield, IL) using a thermal conductivity detector. For purification studies in the vacuum mode, an oilless diaphragm vacuum pump (KNF Neuberger, Inc., Princeton, NJ) was used to create and maintain the vacuum in the permeate fiber set. One can apply vacuum from both ends of the module: the feed-permeate flow pattern inside the permeator is partly cocurrent and partly countercurrent. Or, one can pull vacuum from one side keeping the other side closed: the flow pattern is either completely countercurrent or completely cocurrent. All three variations of the vacuum mode were utilized in the present work.
Results and Discussion Before presenting the results of separation runs in HFCLM permeators, we first report the effective liquid membrane thicknesses of different permeators. We then present the observed SO2 removal performances in different permeators using water as the liquid membrane under the sweep as well as the vacuum mode. We have also compared the experimental results with those from numerical simulations of our simplified model of HFCLM permeation through water membrane in the sweep mode. We have next illustrated how the permeator performance changes for SO2 removal if 1 N NaHSO3 solution is used as the liquid membrane. The HFCLM permeator performances for simultaneous removal of SO2 and NO through aqueous membranes containing a chelating agent at 24 and 70 "C are presented and discussed next. Finally, we have touched upon a number of HFCLM features that introduce significant complexities in HFCLM permeation.
GAS
DRYER
HUYlDlFlER
ANALYZER
Figure 5. Experimental setup for gas separation using HFCLM.
Effective Membrane Thicknesses of the Modules. To determine the effective membrane thickness experimentally, pure C02 and/or pure Nz permeation studies through water as a liquid membrane were carried out. As the permeabilities of these gases through water are wellknown, the membrane thickness can be calculated easily from the measured permeation rates. The effective liquid membrane thicknesses of different modules are available in Table 1. Modules C and F show somewhat higher values of membrane thicknesses. Module E shows a much more favorable value, and the number is close to what can be expected from a purely theoretical prediction (Majumdar et al., 1988). SO2 Separationin HFCLM Permeators and Model Predictions. Separation experiments were carried out first in permeator module C using water as the liquid membrane and a feed gas having 5000 ppm SO2,1.8% 0 2 , 12% C02, and balance N2. The results of SO2 removal through water CLM for various combinations of feed/ sweep flow rates in permeator C are shown in Table 2. Most of the experiments were repeated and found to be highly reproducible. The actual transfer rates were higher for higher feed flow rates, although the corresponding percent removal rates were lower. Higher sweep flow rates for a given feed flow rate appear to achieve a higher SO2 removal rate due to lower values of p y and possible higher facilitation due to a reduced py. The results of numerical simulations of the model for SO2 removal in permeator C are also indicated in Table 2 in terms of the SO2 flux and the SO2 concentration at the sweep gas outlet. The model results appear to predict the trend of the experimental data; further, the absolute values of the predicted quantities, the SO2 concentration in the sweep gas, and the SO2 flux are reasonably close to experimental values in spite of the complexity of the physicochemical configuration. The model employed here for multicomponent gas mixture separation in a HFCLM permeator is based on the solution-diffusion mechanism in the liquid membrane and is valid for gadliquid systems in which the permeability coefficients of gases do not change with the applied pressure. Strictly, the model is somewhat inadequate for SO2 due to facilitation and chemical reaction in water especially at low SO2 concentrations. However, for water membranes, the separation behavior appears to have been predicted reasonably using an effective permeability for S02. Sengupta et al. (1990) have found the permeability of SO2 to be 1.5 X 10" (cm3 (STP) cm)/(s cm2 cmHg) for a model flue gas mixture containing 4100-5400 ppm S02. We have considered the same value of the permeability for SO2 in our numerical calculations. The permeabilities of other gas species were obtained by multiplying the
Ind. Eng. Chem. Res., Vol. 33, No. 3, 1994 671 Table 2. SO2 TransDort acrosa Water Membrane in a Long- HFCLM Permeator' SO2 flux x 106, cmS/(s cm2) percent flow rates feedhweep, APLM, Kexpt X 10'9 cmS/min cmHg cmS/(s cm2 cmHg) SO2 removal expt simulation 38.5142.9 75.0/45.1 75.0/139.5 146.3/142.9
0.1686 0.1328 0.1579 0.1775
0.643 1.54 1.62 2.68
75.3 72.6 91.1 86.9
1.09 2.04 2.56 4.76
SO2 ppm in sweep outlet
1.38 2.29 2.71 5.24
expt
simulation
3378 6037 2449 4448
4300 6775 2622 4903
a Nominal feed concentration (dry): 5000 ppm SOz, 12 % Con, 1.8%0 2 , balance Nz. Module: C. Temperature: 24 O C . Basis for simulation: SO2permeability, 1.5 X 106 (cmS(STP) cm)/(s cm2 cmHg);COZpermeability, 2.1 X lo-' (cmWTP) cm)/(s cm2 cmHg); Nz permeability, 5.57 (cm3(STP)cm)/(s cm2cmHg); He permeability, 9.74 X (cmYSTP) cm)/(s x 10-9 (cmS(STP) cm)/(s cm2cmHg); 02 permeability, 9.01 X cm2 cmHg).
Table 3. Separation of SO, with Water Membrane in a Short HFCLM Permeatop mode (run time in days)
flow rates feed/sweep, cm3/min
SO2 flux x 106, cm3/(s cm2)
SO2 removal
sweep (2) sweep (2) sweep (2)
60.U166.2 60.1/288.5 103.6/288.5
5.21 5.26 6.95
65.4 66.0 50.6
a
percent
K X 104 cm3/(scm2 cmHg) expt simulation 1.03 1.88 1.05 1.84 1.17 1.86
SO2 ppm in sweep outlet
expt 1182 688 908
simulation 1795 1037 1765
Nominal feed concentration (dry): 5000 ppm S02, 12% COz, 1.8% 02, balance N2. Module: F. Temperature: 24 'C.
Table 4. Separation of SO2 with Water Membrane in Modules E and F under Sweep Mode. flow rates feedlaweep,* cm3/min 72.7/114.9 72.71225.8 72.21290.3 144.8/223.9 201.31223.9
flux '08, Cm3/(sCm2) E F 7.21 1.69 8.35 1.10 0.97 8.89 3.57 14.49 5.07 15.74
APLM, cmHg
E 0.2096 0.2126 0.2092 0.2719 0.3223
F 0.0616 0.0381 c 0.0873 0.1280
KexptX 105, cm3/(s cm2 cmHg) E F 3.44 2.75 3.93 2.89 4.25 C 4.09 5.33 3.96 4.88
percent SO2 removal E EandF 72.6 84.1 89.8 69.4 58.4
90.2 95.5 100 88.1 77.8
*
a Nominal feed concentration (dry): 5000 ppm S02,12% COZ,1.8%0 2 , balance Nz. Temperature: 24 O C . Sweep flow rates are reported as total flow rates through two modules. No SO2 was detected in the purified stream by the GC column therefore, A ~ L and M Kex,t calculations were not possible.
diffusivity and solubility values for water (Perry and Chilton, 1973) and are indicated at the bottom of Table 2. In the present HFCLM permeator, gases flow through the hollow fiber lumina. The porous hollow fiber walls (called here the substrate) are also filled with stagnant gas phases (Majumdar et al., 1988). The overall masstransfer resistance can be expressed as the sum of the individual resistances (Figure 2). The substrate-transfer coefficients, which do not depend on the lumen flow rates, were obtained by first calculating the mean free path of SO2 under the given conditions. I t was found that the slip flow regime prevails. Using the appropriate equations (Rangarajan et al., 19841, the corresponding transfer coefficient was calculated to be 6.50 X le3cm3/(s cm2 cmHg). Further, gas-phase mass-transfer coefficients calculated using various correlations (Seider and Tate, 1936;Skelland, 1974) are at least 2 or 3 orders of magnitude higher than the Kexptobserved in Table 2. These justify our use of permeation through the liquid membrane as the controlling resistance for SO2 permeation in the numerical model. Is the long permeator C with large membrane surface area optimal for removing SO2from the gas streams flowing at the rates shown in Table 2? Why does the value of Kexptincrease in Table 2 with increasing feed gas flow rate when permeation through the liquid membrane controls the SO2 removal? The answers are obtained if we consider the high values of the permeability of SO2 through water as measured by Sengupta et al. (1990) from the ILM studies. In a long HFCLM module, most of the SO2 quickly disappears from the feed gas at the entry region of the module due to this high SO2 permeability. For the rest of the permeator length, the driving force for SO2 transfer (Apso2)is quite small. Therefore, the amount of SO2 transferred is
drastically reduced in the rest of the permeator length. However, Kexpt is determined over the whole permeator length with overall APSO~LM, and is, therefore, low. With a higher feed gas flow rate, more SO2 can permeate through a longer length of permeator for essentially similar apso,,^^ leading to higher Kexpt.For , 3 0 2 removal, a short module is, therefore, likely to show a much higher value of Kexpt. Subsequent separation experiments for S02-C02-N2-02 mixtures were, therefore, carried out using much shorter permeators. Table 3 illustrates the experimental SO2 separation results in the sweep gas mode using a much shorter permeator (module F) and water as the liquid membrane. The value of the SO2 mass-transfer coefficient, Keipt, is much higher (of the order of 1 X 1 V cm3/(s cm2 cmHg). Correspondingly,the fractional SO2removal is significantly lower than those in Table 2 even though the feed gas flow rates are comparable to some of the runs in Table 2. Table 3 also provides the results from numerical simulations of the model. The model results agree with the general trend of the observed permeator behavior; the numerical results are satisfactory even though we have made a number of major assumptions. The short permeator F of Table 3 does not quite remove SO2 to the extent needed for flue gas desulfurization. We have, therefore, studied two small permeators in series (E and F) for increased SO2removal. The separation behavior obtained by varying the feed and the sweep gas flow rates with water membranes in two short permeators connected in series is presented in Table 4. Note that these modules are substantially smaller than module C used in our earlier experiments. For example, the membrane surface area of module E is only 27.4%of that of module C. The combined membrane area of modules E and F is about 55 % of that of module C. For each module we have calculated the SO2 flux and the value of Kexpt separately. As Table 4 shows,
672 Ind. Eng. Chem. Res., Vol. 33, No. 3, 1994
-:
4000
-
I
,
I
I
a
3500
f: 3000
cn
U
2500
.* 3
1.1; 0"
I/]
2
Table 5. Comparison of Experimental Separation Data of COz with Numerical Simulation Results.
1500
1
SO2 composition
COZcomposition insweep outlet i n feedbutlet ppm stream, % data simulation data simulation 2500 2.76 4.13 2423 6167 6738 5.39 5.90 4848 4636 3.64 3.54
Feed: 5000 ppm S02,12% C02,1.8% 0 2 , balance Nz. Sweep: helium. Module: C. Membrane: water. Temperature: 24 "C.
A , 0 : Module E & F 1
0
flow percent so2 feed/sweep, removal cmVmin exDt 73.71143.9 94.6 73.7145.1 75.5 43.8145.1 99.8
50
I
,
,
100
150
200
Feed Flow Rate
1
IO 250
(cc/min)
Figure 6. Performance in vacuum mode of operation: single permeator and two permeators in series configuration.
very high removal of SO2 was obtained. As pointed out already, the percent removal of SO2 is much higher in the first permeator compared to that in the second one even though the available membrane areas in both are similar. This shows that the second permeator is being grossly underutilized here. The first small permeator was able to reduce the SO2 composition to a sufficiently low level which, in turn, drastically decreased the d_rivingforce for SO2 permeation in the second permeator. The performance of each permeator in this series combination can be simulated if we know the sweep gas flow rate in each permeator, the feed SO2 concentration to permeator F, and the SO2 concentration in the exiting sweep stream in each permeator. Pure helium sweep gas was applied to each permeator separately in these experiments to increase the partial pressure driving force of SO2 across the contained liquid membrane and thereby increase the SO2 removal efficiency. Unfortunately, the experiments were conducted such that the sweep gas flow rate through each individual permeator was not measured. Rather, the overall sweep gas flow rate used for both permeators was determined. Further, the feed SO2 concentration to permeator F was unknown. As a result, numerical simulation of the experimental results was not possible. SO2 separation experiments under the vacuum mode of operation were made with a single short module (module E) as well as with two short modules (E and F) in a series configuration. The separation behavior for both cases is shown in Figure 6 where SO2 concentration in the purified stream and the total percent SO2 removal are plotted against the feed gas mixture flow rate. The removal rate increases considerably with a decrease in the feed gas flow rate. The performance of two permeators in series (modules E and F combined) is obviously better than a single permeator (module E). However, we do not see a dramatic increase in performance when two modules are used instead of one due to the reasons given earlier (i.e., the driving force in the second permeator is considerably lower). With two permeators as much as 93% of feed SO2 was removed when 46 cm3(STP)/min feed was introduced through modules E and F. In general, higher fractional removal of SO2 was achieved for a lower feed flow rate. In flue gas desulfurization based on a SO2 selective membrane, it is desirable not to remove too much COZ into the permeate stream. Table 5 illustrates the experimental data obtained in permeator C on the extent of
Table 6. Separation of SO2 with 1 N NaHSOl Solution a s a Liquid Membrane. flow rates mode feed/ vacuum, SOzflux K. t~ 106, percent (run time sweep, in.Hg X 10'3,cm3/ cmT/(scm2 SO2 in days) cm3/min (cmHg) cmHg) removal (s cmz) sweep (2) 41.5/78.3 6.61 5.71 98.0 vacuum (2) 41.5/ 27.8 (70.6) 5.60 83.1 Nominal feed concentration (dry): 6764 ppm S02,13.2% C02, 1.8% 0 2 , balance N2. Module: G. Temperature: 24 "C.
COz removed as SO2 is being removed to a greater extent, on a percent removal basis. Table 5 also illustrates the predictions from the simplified model of the extent of C02 removal. The agreement between the numerical simulation results and the experimental data appears to be satisfactory. The experimental data obtained with 1 N NaHS03 solution as a liquid membrane are presented in Table 6. These experiments were carried out at room temperature in module G having larger diameter hollow fibers (240 pm i.d.1. Excellent fractional removal of SO2 was obtained in both sweep and vacuum modes. These particular experiments with 1 N NaHS03 solution were stopped when steady state was achieved. Therefore, we do not know the effect of prolonged exposure of SO2 on the CLM, if any, under such conditions. Combined SOz-NO Separation in HFCLM Permeators. The separation experiments for simultaneous separation of SO2 and NO from a feed mixture of SOzNO-COZ-OZ-NZ at 24 "C were carried out in short permeators (module E and module G, respectively) both in sweep and vacuum modes with 0.01 M Fe2+EDTA solution as a liquid membrane. The steady-state separation results for both modules are reported in Table 7. Excellent separation results were achieved in both modes. With a high sweep flow rate (3 times that of the feed gas flow rate), greater than 95% of the feed SO2 was removed. Comparing the results of both modes, it seems that the vacuum mode is more effective in removing NO. Due to the evaporation of water at a high rate, the permeate side partial pressure of NO is probably reduced more in the vacuum mode than in the sweep mode. This will lead to better facilitation and higher flux. Since the NO-complexing chelate concentration in water as well as the NO concentration is very low, the multicomponent permeation model may also be applied to the SO2 removal results observed in Table 7. For the sweep runs using modules E and G, the simulated SO2flux values (7.42 X lo4, 7.50 X 10-6,6.53 X 10-6) are quite close to the experimental values (6.48 X lo4, 7.18 X 10-8,5.89 X 10-6). This again demonstrates the usefulness of the multicomponent permeation model. Combined S02-NO separation runs were then carried out in both sweep and vacuum modes in module F at a higher temperature of 70 "C with a solution of 0.01 M Fe2+EDTA chelate as a liquid membrane. The steady-
Ind. Eng. Chem. Res., Vol. 33, No. 3, 1994 673 Table 7. Simultaneous SeDaration of SO2 and NO with 0.01 M FeZ+EDTASolution as a Liauid M e m b r a n e flow rates mode feed/sweep, flux x IO6,cm3/(s cm2) ICexp$ x lo5,cm3/(s cm2cmHg) module (run time in days) cm3/min in.Hg (cmHg) SO2 NO so2 NO E sweep (3) 41.5178.3 6.48 0.284 3.86 1.51 E sweep (2) 41.5/118.2 7.18 0.350 5.10 2.16 E vacuum (1) 41.41 27.5 (69.9) 4.89 0.306 E vacuum (1) 60.U 27.5 (69.9) 10.66 0.559 G sweep (3) 41.5/78.3 5.89 0.209 2.71 0.73 G vacuum (2) 41.5/ 27.5 (69.9) 5.87 0.313 ~~
percent removal SO2 NO 84.6 51.1 99.4 63.1 70.0 73.3b 98.6 75.Y 87.4 42.8 87.2 64.1
0 Nominal feed concentration (dry): 6764 ppm S02,490ppm NO, 13.2% Con, 1.8% 0 2 , balance N2. * Nominal feed concentration: 6188 ppm S02,370 ppm NO, 1.8% 02,balance N2. Nominal feed concentration: 6588 ppm S02,450ppm NO, 1.8% 0 2 , balance N2.
Table 8. Simultaneous Separation of 902 and NO at 70 flow rates fluxx 108 cm3/(s cm2) mode feedlsweep, vacuum (run time in days) cm3/min in.Ha (cmHa) so2 NO sweep (2) 41.5/78.3 5.18 0.253 vacuum (4) 41.5/ 27.8 (70.6) 6.62 0.320
ICelpt X 105, cmg/(s cm2 cmHg)
a Nominal feed concentration (dry): 6764 ppm S02, 490 ppm NO, 13.2% C02, 1.8% Fe2+EDTAin water.
*t
0 :
NO'
480
P
2 :
0.0 o.2 0
20
40
60
80
100
120
140
160
Elapsed Time ( h r )
Figure 7. Simultaneous separation of SO2 and NO at 70 OC with 0.04 M Fe3+EDTA solution as a liquid membrane under vacuum mode. Nominal feed concentration (dry): 5000 ppm S02,500 ppm NO, 12% C02, 1.8% 0 2 , balance Nz.
state performances are shown in Table 8. Although these preliminary findings show a small decrease in performance at the high temperature compared to those in Table 7, the results are very encouraging. Here, a Fe2+EDTAchelate solution of only 0.01 M concentration was utilized as a liquid membrane. Higher chelate concentration may be used (because the solubility of chelate is higher at 70 "C), and that would increase the NO removal considerably. Note that the feed gas contains about 1.8% 0 2 . At high temperature, the Fe2+EDTA chelate solution probably degrades in the presence of oxygen. With that idea in mind, another experiment was carried out with a solution of 0.04 M Fe3+EDTAchelate solution at 70 "C. The performances are shown in Figure 7. A continuous run was made for 6 days with a feed gas mixture containing SO2, NO, COZ, Nz, and 0 2 under the vacuum mode of operation. The results demonstrate the absence of liquid membrane degradation in the presence of 02when Fe3+EDTAsolution instead of Fe2+EDTAsolution is used as a membrane. The flux of SO2 dropped quite a bit from its initial value, but that of NO remained constant for the entire period. During the experiment, the feed gas mixture was humidified at the room temperature before it was introduced
so2 1.68 02,
NO 0.92
percent removal so2 NO 70.0 46.9 89.0 59.4
balance N2. Module: F. Membrane: 0.01 M
into the module kept in a constant temperature bath at 70 "C. The gas mixture inside the module, not being prehumidified at 70 "C, naturally picked up moisture from the membrane liquid solution and thereby increased the solution concentration. This should decrease the removal rate of both SO2 and NO due to any salting out effect. However, we do not see any substantial change in the NO removal rate since the reaction rate of NO and the chelate increases a t the higher chelate concentration produced by evaporation. In actual practice, the problem of increasing the chelate concentration can be easily avoided by very slow recirculation of the membrane liquid outside and addition of moisture. Gas Flow Pressure Drop. Gas flow pressure drop is important for flue gas cleanup application as the gas stream is available essentially at atmospheric pressure. The typical pressure drop calculated for a module with 300, 100-pm-i.d. fibers ranged from 9.7 to 97 in. of water per foot of the fiber length at the flow rates of 20-200 cm3/ min. The pressure drop ranged from only 0.3 to 3 in. of water per foot of the fiber length when 240-pm-id. fibers were used for calculation under identical conditions. In fact, by using a 6 in. long commercial module containing a single set of 1800 hollow fibers of 240-pm i.d., the gasphase pressure drop was experimentally found to vary from 1 to 9 in. of water for a gas flow rate variation of 1000-5000 cm3/min. Based on these data, we see that for 10 in. of water allowable pressure drop, a maximum flue gas flow rate of 3 cm3/min in each fiber can be handled. A module made of 400-pm-i.d. hollow fibers is expected to provide much lower pressure drops. Comments on Factors Affecting HFCLM Separation. Gas separation in a HFCLM permeator is affected by a number of factors. These are as follows: the feed/ sweep (or vacuum) flow patterns, extent of sweep gas permeation into the feed side, position of the membrane replacement port, longitudinal diffusion in the liquid membrane phase, and amount of water condensation in the gas lines, etc. As the membrane liquid remains at equilibrium with the gas at every permeator location, there would be gradients in SO2 and HSOC ion concentrations in the membrane phase along the length of the module. For countercurrent operation, the permeant concentration in the membrane liquid phase is highest at the feed gas inlet side of the module. Within the fiber bundle, the radial gradient is orders of magnitude larger than the longitudinal
674 Ind. Eng. Chem. Res., Vol. 33, No. 3, 1994
gradient, thus the longitudinal gradient may not be disturbed. But, if the HFCLM module has a large pool of liquid outside the fiber bundle (as in the case of module C),its longitudinal gradient would be smeared easily. This may affect the concentration profiles in the membrane liquid within the fiber bundle, reducing the SO2 transfer rate. To minimize this problem, HFCLM modules without any extraneous membrane liquid around the fiber bundle are desirable. Modules E, F, and G are built in that fashion and, therefore, may benefit in terms of higher Kexpt. During any experiment, water accumulates in the liquid separators provided at the feed and sweep outlet lines. This water may act as a SO2 sink by continuously absorbing some gas species. The extent of this effect is not known. It will also depend on any fiber leaks.
Conclusions An innovative hollow fiber contained liquid membrane permeator was investigated for removal of SO2 and NO from flue gas. The excellent aqueous liquid membranes selected via ILM studies (Sengupta et al., 1990)were found to perform efficiently when used as liquid membranes in HFCLM permeators of small dimensions. For SO2removal alone, water and solutions of sodium bisulfite showed excellent results. Depending on the permeator length, which varied between 40 and 160 cm, and the gas flow rates, more than 60-95% of the SO2 was easily removed from a feed flue gas containing 5000 ppm SO2 by an aqueous membrane into the permeate which either had a sweep gas or was subjected to vacuum. The SO2 mass-transfer coefficient at 24 "C in a small permeator was found to be around 1 X lo4 cm3/(s cm2 cmHg). A simplified permeation model of nonfacilitated multicomponent gas separation in the HFCLM permeator describes satisfactorily the removal of both SO2 and CO2 from the feed gas into the sweep gas stream. Using an aqueous solution of 0.01 M Fe2+EDTAand a feed flue gas containing both SO2 and NO in the presence of 02,70-90 '3 of feed SO1 and 50-75 % of feed NO were simultaneously removed at 24 OC in a small HFCLM permeator. The HFCLM purification run at 70 "C with the same feed mixture and liquid membrane showed a performance only slightly inferior to the run carried out at 24 "C. The problem of flux reduction with Fe2+EDTA chelate solution in the presence of 0 2 may be avoided by using Fe3+EDTAchelate solution in the first place. Much more experimentation at 70 OC is desirable to draw firmer conclusions. Further research is required to optimize the design and operating conditions of this system for future applications.
Acknowledgment This work was supported by Contract DE-AC2287PC79853 of the Pittsburgh Energy Technology Center (PETC) of the U S . Department of Energy. Hoechst Celanese Corporation (SPD, Charlotte, NC) provided the hollow fibers used in this study. The suggestions made by Dr. Soung Kim and Dr. Richard Walker of PETC are very much appreciated.
Nomenclature AT = total membrane permeation area, m2 d = liquid membrane thickness, m di = fiber inside diameter, m do = fiber outside diameter, m F = fractional removal of a species, defined in eq 4 k = mass-transfer coefficient, m/s
K = overall mass-transfer coefficient, m/s 1 = active length of the module, m L = feed side gas flow rate, mol/s L , = reference flow rate, mol/s L* = dimensionless feed side gas flow rate, L / L , N = flux across the liquid membrane, mol/(m2 s) p = sweep side pressure, Pa p* = dimensionless sweep side pressure, p / P , P = feed side pressure, Pa P, = reference pressure, Pa PC = dimensionless feed side pressure, P/P, Qi = permeability of species i, (mol m)/(m2 s Pa) Q, = reference permeability, (mol m)/(m2 s Pa) Qi* = dimensionless permeability, QJQ, R = universal gas constant, (m3Pa)/(mol K) RT = total permeation rate, mol/s S = dimensionless area, do(Qr/d)(P,/Lr)l T = absolute temperature, K V = sweep side gas flow rate per fiber, mol/s x = mole fraction in the feed side gas mixture y = mole fraction in the sweep side gas mixture Greek Symbols
fi = dimensionlessparameter, 128RTL,2~r/[r2d+do(Qr/d)P,31 p = gas viscosity, Pa s p, = reference viscosity, Pa s p* = dimensionless gas viscosity, p / p , Superscripts g = gas phase m = membrane s = substrate * = pertaining to a dimensionless quantity
Subscripts expt = experimental value f = feed inlet end F = feed side i = species i LM = logarithmic mean r = pertaining to a reference parameter S = sweep side w = feed outlet end
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Received for review April 15, 1993 Revised manuscript received September 24, 1993 Accepted December 20,1993O Abstract published in Advance ACS Abstracts, February 1, 1994. @
