Ind. Eng. C h e n . Res. 1988,27, 2149-2156
2149
Modeling of Carbon Dioxide Absorber Using Hot Carbonate Process Dipayan Sanya1,t Neeraj Vasishtha,* and Deoki N. Saraf* Department of Chemical Engineering, Indian Institute of Technology, Kanpur 208016, India
A mathematical model has been developed for the absorption of carbon dioxide from ammonia synthesis gas in amine-promoted hot potassium carbonate solution. The model uses vapor-liquid equilibrium data for the unpromoted carbonate-C02 system, which seems to be adequate at low fractional saturation of the liquid with carbon dioxide. In the temperature range of industrial operation, instantaneous reaction rate provides the appropriate enhancement factor for the chemical absorption. The calculated results have been found to be in close agreement with the observed values for two industrial units. The model has been used to calculate various flow, temperature, and concentration profiles in the absorber. Variation of exit C02 concentration with some of the operating parameters has also been examined. For a commercially useful ammonia synthesis process, economic removal of carbon dioxide has considerable importance since it acts as a diluent of the synthesis gas and is also thought to be a poison for the synthesis catalyst. Industrially, removal of carbon dioxide is largely brought about in two steps. The first step of the sequence is responsible for bulk removal of COz usually by absorption in hot potassium carbonate solution, promoted by organic amines. Among the processes employing hot potassium carbonate as the basic chemical solvent, the Benfield process has gained the broadest industrial acceptance for acid gas removal (Benson et al., 1954, 1956; Field et al., 1962; Parrish and Neilson, 1974; Bartoo and Ruzicka, 1983; Bartoo, 1984). Though the theory of chemical absorption was established in the late 1960s (Danckwerts and Sharma, 1966; Astarita, 1967; Ramm, 1968), adequate vapor-liquid equilibrium (VLE) data and the kinetics of reaction with concentrated hot potash solution, particularly for the high-temperature range usually employed by industry, are still not found in the open literature. In early 1970s, Uchida et al. (1972) published a rather simplistic “compartment-in-series” model for design and simulation of an adiabatic packed absorber, which was tested for the case of C02 absorption in hot potash solution. The thermodynamics and kinetics developed for the problem were anything but rigorous. Apart from this, a computer program used for the design of a hot carbonate absorption system was published by Wellman and Kate11 (1968) without any mathematical model to accompany. Availability of recent accurate thermodynamic and kinetic data (Chen et al., 1979; Astarita et al., 1983), as well as knowledge of the effects of rate promotion (Astarita et al., 1981, 1983) and other information related to this system, has paved the way for attempting simulation of the Benfield absorbers. Krishnamurthy and Taylor (1985,1986) have developed a nonequilibrium stage model for simulation and design of packed absorption columns.
difference between the partial pressure of carbon dioxide in the bulk of the gas and its equilibrium partial pressure at the gas-liquid interface, was believed to be sufficient for absorption without the conventional lowering of temperature. Since the carbonate-bicarbonate solution remains hot at all times, more concentrated solutions can be used than with cold absorption without precipitating the bicarbonate (Benson et al., 1954, 1956). A single stream absorber operating at about 25 atm of pressure and in the temperature range 105-120 “C can remove C02 to a level of 1% in the exit stream. For a higher degree of purification, the flow scheme known as a “split-flow absorber” design is used. The bulk of the hot, regenerated lean solution, 60-70% of the total, is introduced into the absorber at an intermediate point, while a small trim solution is cooled by 10-40 OC and pumped to the top of the tower. The lower temperature trim stream has a reduced equilibrium back pressure of carbon dioxide to permit attaining 0.1% high purity C02 in the purified gas. Since the bulk of the absorber still operates at high temperature, the rich solution stream is hot and regeneration efficiency is not impaired (Bartoo, 1984). When still higher purity is required, a “two-stage”design, i.e., a split-stream flow in the absorber and split-stream flow in the regenerator, is required. Two regenerated streams leave the regenerator; the major portion (6045% of the total) is removed from an upper level in a partially regenerated state for return to an intermediate level of the absorber. The minor stream, 15-40% of the total, is regenerated further in the bottom section of the regenerator by the total boil-up steam, removed from the sump, and cooled before being pumped to the top of the absorber. In this scheme, the C02 level in the purified gas is brought down to 0.05 % . When carbon dioxide is absorbed in an unpromoted solution of potassium carbonate and bicarbonate, it reacts according to the following overall reaction:
Benfield Process In order to exploit the advantage of the elevated pressure
Since potassium carbonate and bicarbonate are both strong electrolytes, it may be assumed that the metal is present only in the form of K+ ions, so reaction 1 may be more realistically represented in the ionic terms as
and temperature at which the impure synthesis gas is available, the use of hot, concentrated alkaline solutions under pressure was investigated by the US Bureau of Mines. In this system, the driving force, which is the ‘Present address: Larsen & Toubro, Powai, Bombay 400072, India. Present address: Department of Chemical Engineering, Rensselaer Polytechnic Institute, Troy, NY 12180.
*
0888-5885/88/2627-2149$01.50/0
C02
+ K&03 + HzO + 2KHC03
C02
+ C 0 3 2 -+ H 2 0 I=t 2HC03-
(1)
(2)
In absorption, reaction 2 proceeds from left to right. The above reaction, being trimolecular, is evidently made up of a sequence of elementary steps. Carbon dioxide may undergo two direct reactions in an alkaline solution, namely, 0 1988 American Chemical Society
2150 Ind. Eng. Chem. Res., Vol. 27, No. 11, 1988
COP + H 2 0 ==HC03- + H+
(3)
COP + OH- + HCO3-
(4)
Reactions 3 and 4 are both followed by an instantaneous reaction, leading to the overall result represented by reaction 2: COS2-+ H+ + HC03-
C0,2-
+ H 2 0 + HC03- + OH-
-Z+dZ
(5) (6)
-2
Both of these reactions are accompanied with the following instantaneous dissociation reaction H20 + H+ + OH-
(7)
The first sequence, namely, reactions 3, 5, and 7, is known as the ”acidic”mechanism (Savage et al., 1980). The acidic mechanism’s contribution to the overall reaction rate is negligible unless the pH of the liquid solution is very low. Almost all the cases of industrial absorption are held at high pH (pH > 8), and hence the acidic mechanism is neglected altogether in the present study. Since reactions 6 and 7 are instantaneous reactions, reaction 4 is the rate-controlling step for absorption of C02in unpromoted hot potash solution. Promotion of mass-transfer rates, through additives to the K2C03solution, has been reported in the literature (Shrier and Danckwerts, 1969; Roberts and Danckwerts, 1962; Danckwerts and Sharma, 1966; McLachlan and Danckwerts, 1972; and McNeil and Danckwerts, 1967; Jeffreys and Bull, 1964; Astarita and Beek, 1962). The most recent theory has been discussed by Astarita et al. (1981). It is assumed that the role of additives is that of a catalyst in the reaction rather than a modifier of vapor-liquid equilibrium. For rate promotion with the help of amino compounds, the reaction which takes place at the gas-liquid interface is C02 + RR’NH =+RR’NCOOH (amine) (carbamate)
Figure 1. Differential section of a packed absorber.
Making overall and component mass balances on envelopes I and I11 and energy balances on envelopes I and I1 yields (Sanyal, 1986) dV/dz = -(NI + N2) (10)
_ dY1 - ND1dz
Nl(1 - Y1) V
(11)
(8)
Reversion of the carbamate to bicarbonate ion takes place in the bulk of the liquid as follows: RR’NCOOH
+ OH- + HC03- + RR’NH2
(9)
A t ambient temperature, reaction 8 is much faster than reaction 9 and the so-called “shuttle mechanism” is operative (Astarita et al., 1981), in which the carbamate diffuses into the bulk where it reacts with OH- ion to regenerate the amine which then diffuses back to the interface to react with more C02. However, at higher temperatures, in the range of industrial operation, the rate of reaction 9 increases significantly, and it follows reaction 8 immediately and locally; the system is better represented by the homogeneous catalysis mechanism. This mechanism may shoot up the overall reaction rate to an instantaneous regime. Model Development The material and energy balance equations are developed around a differential height of the packed absorber shown in Figure 1. Envelope I11 is an elemental volume in the differential packed height, dz, of the absorber, consisting of the gas and liquid films denoted by envelopes I and 11, respectively. The major assumptions are (1) steady-state conditions prevail, (2) pressure drop across the packed bed is negligible, and (3) C02and H 2 0 are the only components transported across the interface.
Here it has been assumed that for amine-promoted carbonate solution the reaction rate tends to the instantaneous regime and hence the reaction takes place a t the same rate a t which the reactants are brought together by diffusion, i.e., N 1 = rS (18) Equations 10-17 form the model for the absorber. The boundary conditions have to be prescribed depending on the configuration of addition and withdrawal of various streams. However, before that, we need to develop procedures for computing various quantities involved in the model such as N,, N2, h, etc. In terms of the two film model of mass transfer, Nk = (K,ka)(Pk, - P*k)s (19) The wetted interfacial area, a, has been calculated from the correlation of Van Krevelen and Hoftijzer (1948):
The interfacial area calculated from above equation has been used for both mass and heat transfer. Several correlations are available for estimation of the gas-side mass-transfer coefficient (Van Krevelen and
Ind. Eng. Chem. Res., Vol. 27, No. 11, 1988 2151 Hoftijzer, 1948; Danckwerts, 1970; Uchida et al., 1972). The following correlation, proposed by Onda et al. (1968), has been used in the present model:
where the subscript g refers to properties in the gas phase. The liquid-side mass-transfer coefficient for C02 is obtained from the analysis of Van Krevelen and Hoftijzer (1948), which accounts for the chemical reaction through the enhancement factor, I. Since, for amine-promoted potassium carbonate solution, the rate of reaction can be considered to be instantaneous, Van Krevelen and Hoftijzer (1948) have suggested the use of the following expression for I, which is basically a Hatta number:
and
The kinetic constant, in the temperature range 0-110 "C and at the ionic strength of practical interest, has been well correlated (Astarita et al., 1983) as
+
log kII = 13.635 - 2895 0.081i
t
It may be assumed that there is no liquid-side resistance for the mass transfer of the solvent water vapor. Therefore, the overall mass-transfer coefficient for the solvent water vapor is (25)
For the C02-K2C03system, Astarita et al. (1983) correlated the Henry's law constant, H , with molarity m as follows: log ( H / H ") = 0.025m (26) where H " is Henry's law constant for COz in water and can be obtained from log [H"]-' = -4.3856
+ 867.4932/t
(27)
and m, the molarity of the solution, is the total KzC03 concentration. Before discussing vapor-liquid equilibria, we define x, the fractional saturation or the degree of saturation, as the ratio of the concentration of the absorbed component to the total concentration of the reactive component in both reacted or unreacted forms:
x = - bZ
2m where bz is the concentration of HCO
1.1. The heat-transfer coefficient, hk,can be estimated using the analogy between heat and mass transfer in view of the lack of experimental data (Uchida et al., 1972):
The overall heat-transfer coefficient, h, can be obtained by taking the molar average of component coefficients (Linch and Wilke, 1955; Heertjes and Ringens, 1956). For liquid density, the following relation obtained by linear regression of the data of Bocard and Maryland (1962) has been used. It is valid in the temperature range 70-130 “C and concentration range 20-40 wt Ti total equivalent K2C03: s = 1.0679 0 . 0 1 ~ 9.4 X 10-4tc (42)
+
where s is the specific gravity of the K2C03solution, w is the total equivalent weight percent of KzC03,and t, is the liquid temperature. The density of gas mixture is calculated by using ideal gas law. The liquid viscosity data of Bocard and Maryland (1962) for the KzC03-KHC0, solutions have been correlated by using nonlinear regression as p = 12.9275 - 0.1114t~+ 4.1184 X 10-4t~2 - 5.796 X l O - ’ t ~ ~(43) where p is the liquid viscosity and tF is the liquid temperature. This correlation is valid for 30 w t % KZCO* The gas mixture viscosity was obtained by using the Wilke correlation (Reid et al., 1977) and individual component viscosity data. The liquid diffusivity data of Savage et al. (1980) for the COz-KzC03 system have been correlated by using linear regression as log D = -3.0188 - 586.9729/t - 0.4437 (44) where D,the liquid diffusivity, has been assumed to be equal for all components. Mixture diffusivities for the gas phase have been calculated by using the correlation of Reid et al. (1977). The needed binary diffusivities were obtained by using the correlation proposed by Shemood and Pigford (1952).
,
1 I
I
‘bl
‘bg
\C
‘b3
Lb-’
Figure 2. Shooting method.
Data on the specific heat of the K2C03solution published by Kohl and Risenfield (1979) have been correlated by using nonlinear regression as follows: Cl, = 0.9998 - 0.00914~0.1063 X 10-3w2+ 0.3058 X 1 0 % ~(45) ~ The specific heat of the gas mixture was calculated assuming the mixture to be ideal. The thermal conductivity of the gas mixture was obtained from individual component data of Chang (Reid et al., 1977) and the mixture rule of Wassiljewa (Reid et al., 1977). Absorption of C02 in K2C03solution is accompanied by two heat effects, the heat of solution and heat of reaction. The net heat of absorption has been correlated from the data of Bocard and Maryland (1962) by nonlinear regression as follows: AHl = 6510.8058 + 19.4580~+ 1.2727931~’0.0313384~~ + 3.955 X 10T4w4(46) Computational Method In a countercurrent absorber, the rich feed gas enters at the bottom and absorbing lean liquid enters at the top of the column, and these are usually completely specified. However, the lean gas leaves from the top and rich liquid leaves from the bottom, both of which are, at best, partially specified. Therefore, the state of neither end of the column is known completely, which leads to a two-point boundary value problem. Several methods are available for solving two-point boundary value problems. For the present work, the “shooting method” has been used (Fox, 1957). In this method, values for the unspecified conditions at the initial point of the interval (missing initial conditions) are assumed, and the differential equations are numerically integrated by using the fourth-order Runge-Kutta method to the terminal point by applying essentially an initial value algorithm (“shooting”at the target terminal points). If the computed terminal values satisfy the specified terminal conditions, the problem has been solved. If they do not, the differences between the computed and specified terminal conditions (the “miss distances”) are used to adjust the assumed initial conditions. If the differential equations and boundary conditions are linear, the adjustments need only be made once, but if the differential equations are nonlinear, the adjustment of the missing initial conditions is an iterative process (Roberts and Shipman, 1972; Fox, 1957; Keller, 1968). Figure 2 shows how the shooting method works. For the liquid flow rate at the bottom, two values, Lbl and LbZ, are assumed. The and Lu,dc,are used correspondingcalculated values, Ltl,calc to calculate the difference functions, L , - Ltl,calcand L, -. La,dc (where L, is the observed liquid flow rate at the top),
Ind. Eng. Chem. Res., Vol. 27, No. 11, 1988 2153 Table I. Design Data for the Absorbers case I case I1 ht of diameter ht of diameter packing, of packed packing, of packed m bed, m m bed. m 1st section 6.0 1.4 6.4 3.2004 2nd section 5.6 1.4 6.1 2.4384 3rd section 5.6 1.4 packing size 37 mm 37 mm Intalox saddles packing shape Intalox saddles packing material ceramic ceramic Table 11. Comparison of Calculated Results with the Observed Plant Data (Case 1) outlet oarameters inlet obsd calcd 70 dev liquid flow rate, kmol/h cold (343K) 2630 2630 hot (371K) 5308 5352 -0.82 343.0 392.0 390.0 0.50 liquid temp, K liquid composition, mole fraction 0.04013 0.0158 0.0159 -0.63 KZC03 0.02015 0.0663 0.0669 -0.90 KHC03 water 0.9320 0.9198 0.9095 1.12 7.193 X DEA 7.0053 x 10-3 7.070 X 10" -0.92 5.554 x 10-4 5.399 x 10-4 5.449 x 10-4 -0.93 v20.5 0.43 762.877 547.090 544.690 gas flow rate, kmol/h gas temp, K 408.0 382.2 gas composition, mole fraction 0.1659 0.001 0.00107 -7.0 COZ 0.001 84 0.00257 0.00258 -0.39 CHI 0.00263 co -0.55 0.00366 0.00368 0.5306 -0.58 0.7431 0.7388 HZ 0.1733 0.2413 -0.58 0.2427 NZ 0.0021 Ar -0.69 0.00292 0.0029 0.1237 41.40 0.0058 0.0099 H20
which are plotted as A and B. Line AB is extended to intersect the x axis at C, which provides the next guess for L b , say Lm For the next iteration, L b 2 and Lb3are the two values used to get the new guess in the same fashion. The remaining seven variables, x l b , X2b, x3b, x4b, X5b, x 6 b ) and tb, are treated in the same manner. A computer program has been written in Fortran language for the DEC 1090 system using the above procedure. A step size of 1/20 of the packed height for each section was found to be adequate. A convergence tolerance of the order of 0.001 was used.
Results and Discussion The model presented above has been tested for carbon dioxide absorber in two existing ammonia plants. The plants use hot potassium carbonate solution with diethanolamine (DEA) promoter for absorption of carbon dioxide. Case 1. The split flow packed absorber has three sections. The lean solution from the regenerator is split into two parts. One part enters the absorption tower at the top after being cooled, whereas the other part enters after the first section of the packing. Table I shows the design parameters for the absorption tower. Table I1 shows the values of the input and output variables such as flow rates of liquid and gas streams, the composition of each stream, temperature, etc. Also shown in Table 11are the calculated values for the same outlet variables for comparison. Except for the mole fraction of water vapor in the gas phase, all the calculated liquid and gas compositions match the plant values very well, the maximum deviation being 7% for carbon dioxide. Even this 7 % deviation may be due to measurement error, since there is a higher uncer-
I Plant data
B Liquid temperature C Gos flow rate 0 Gas temperature
c
..-
h
v
3500
L
IO0
Y
a8
L
3
0
L
a4 c
5006
I 0.2
I
1
0.4 0.6 Bed height ,fractional
0.8
1
40
Figure 3. Temperature and flow profiles in the absorber.
tainty in measurement for the small concentration range. If we calculate the percent COzremoved, then the observed value is 99.4, which is to be compared with the calculated value of 99.36. The calculated value of the mole fraction of water vapor in the gas was 41.4% lower than the plant data. Again, if we consider the percent removal of water vapor in the column, then the observed value is 92%) which must be compared with the calculated value of 95.3% , and now the match looks more reasonable. The flow rates and temperatures of both gas and liquid streams match well. The temperature, flow, and concentration profiles have been calculated by using the model along the packed tower height from the bottom to the top. Figures 3 and 4 show these profiles along with plant data at the top and bottom. The step change in liquid flow rate and liquid temperature seen in Figure 3 is caused by the addition of the split stream of the lean absorbent at the end of the first section of packing. While the gas flow rate is not affected by the liquid addition, the gas temperature shows a slight increase. Since the bed height is measured from the bottom up, the addition of liquid shows a step decrease when the x coordinate changes from 0 to 1. Gas flow rate decreases along the height because of absorption of C 0 2and water vapor. The gas temperature decreases as it moves up the column since heat is being transferred to the liquid because of absorption and temperature differences. For the same reasons, liquid temperature increases as it moves down the column. The concentration profiles for various components are shown in Figure 4. The concentrations of C02 and water vapor in the gas phase decrease as the gas moves up the column. It may be noted (curve A in Figure 4) that the bottom two sections of the tower, which constitutes about 2 / 3 rds of the bed height, remove nearly 94% of the COZ
2154 Ind. Eng. Chem. Res., Vol. 27, No. 11, 1988 0.IS[
I
I
1
Table IV. Comparison of Calculated Results with Observed Plant Outlet Data (Case 2) outlet parameter obsd calcd % dev liquid flow rate, kmol/h 33655 33020 1.86 liquid temp, K 393.8 383.7 2.60 liquid composition, mole fraction
I
Plant data A Conc o f 8 " C " D " E '' F "
0 12
C02 in gas
m
H20 in gas C o p in liquid H20 in l i q u i d K CO3 in l i q u i d
a
K?-ICO3 in l i q u i d
0
0
V 0
C
K2C03
o_
-
KHC03
0
H20 DEA
e
L
v2°5
gas flow rate, kmol/h gas temp, K gas composition, mole fraction COZ H2
N2
co CHI Ar HzO
h
0
0.2
0.4 0.6 Bed height ,tractional
0.016813 0.050 13 0.896 69 0.00692 0.00692 2295.753 353.0
0.029 84 0.045 979 0.916 57 0.0069387 O.OOO66922 2295.118 353.9470
-43.6 9.0 -2.2 -0.3 2.3 0.03
0.001 9697 0.73422 0.237 94 0.004 530 14 0.0031515 0.003 053 0.015 1324
0.001 9612 0.681 569 0.221 660 0.004 2678 0.00296679 0.002 8464 0.014 901
0.4 7.7 6.8 5.8 6.2 7.2 1.5
-1.9
0.8 10-2,
I
I
I
i
I
1
I
I
I
I
1
I
I
Figure 4. Concentration profiles in the absorber. Table 111. Inlet Plant Data (Case 2)
stream
parameter total flow rate, mol/h composition, mole fraction K2C03
KHCO:,
DEA vZ05
H20
coz
lean liquid stream
semilean liquid
7410.56
25 458.26
0.041 957 0.028 073 0.007 353 0.000 7036 0.921 927 0.00
0.026 3125 0.051 192 0.006 818 0.000 6592 0.915018 0.00
353
393.4
H2 N2
co CHI Ar temp, K
gas 3081.92
0.050 303 0.215 77 0.548 35 0.177 48 0.003 42 0.003 75 0.002 2793 380
and the top section is meant only to fine tune the C02 concentration from 0.01 to 0.001 at the exit. This low split type of absorber is designed to reduce the exit C02 concentration to a fairly low level by finally contacting cooled absorbent in the top section. Since C02reads with K2C03, there is no free carbon dioxide in the liquid phase. The concentration of K&03 in the solution decreases as it moves down the column, since it reacts with the dissolved COD As the reaction product is KHC03, its concentration increases as the liquid travels toward the bottom. Case 2. The design parameters are included in Table I. This is a split flow type of absorber with a two-stage regenerator. A cooled lean solution coming from the bottom of the regenerator enters at the top of the absorber, whereas a semilean hot solution from the middle of the regenerator enters in the middle of the absorber. The inlet operating data for the lean and semilean liquid streams and the gas stream are presented in Table 111. The calculated temperatures, flow rates, and compositions of the outgoing streams for this case are included in Table IV along with observed values for comparison. The calculated total flow rates and temperatures for both liquid and gas streams match well with the observed values. Except for the mole fraction of KzC03in the liquid phase, all the calculated liquid and gas compositions also
I
10
20
30
I
40
T o t a l pressure ,atm
Figure 5. Exit C02 concentration versus total pressure.
match reasonably well. For K2C03,the calculated concentration is 43.6% higher than the plant value, and for KHC03 it is 9% lower. For a total K+ balance, the calculated outlet rate of 3472.89 kmol/h matches almost exactly the inlet rate of 3472.96 kmol/h. This shows that the measurement of outlet concentrations of K2C03and KHC03 may be in error. The temperature, flow, and concentration profiles along the length of the absorber are similar to those for case 1 and hence omitted. Sensitivity Analysis. From a plant operator's point of view, it is important to know how sensitive the absorption process is with respect to certain operating variables. In the present case, where the objective is to remove COzfrom the synthesis gas, one would like to find how the exit stream C02concentration would vary if the operating pressure of the tower was to change from its design value. Figure 5 shows this result for case 1. A decrease of 5 atm in total pressure could significantly lower the C02removal efficiency. Figure 6 shows how exit C 0 2 concentration would change if the fraction of lean absorbent fed at the top of the column is varied from its design value. This is
Ind. Eng. Chem. Res., Vol. 27, No. 11, 1988 2155 close agreement between the calculated values and experimental results for two sets of plant data justifies the assumptions made. o Plant data
$ 11
0
I
c
I
C
.-
-e
I
I
i
I
I
i
1
-
I 0
;lo E
-2 -
o A B
Plant data Temperature of cooler lean stream Temperature of hot split stream
-.
.-C0
I
0 I L
-
al C
u N
0 0
50
70
110
90 Temperature
130
,OC
Figure 7. Exit COz concentration versus temperature of lean
streams.
a less sensitive parameter compared to the total pressure, but an increase in this stream will result in a higher purity of synthesis gas. Changing the temperature of cold lean stream fed at the top of the tower seems to have very little effect on COz removal (Figure 7, curve A). This is because, in the low-temperature range, VLE is not significantly affected by temperature variation. Also the amine promotion is negligible at low temperatures. However, an increase in temperature of the split stream lean absorbent adversely affects the gas purity, as seen from curve B in the same figure. Conclusions The model of C 0 2 absorber, developed in the present study, adequately accounts for the effect of the presence of amines in the hot carbonate solution. While amine seems to play the role of a catalyst and pushes the C0,KzCO3 reaction to the instantaneous regime, the promotional effect in modeling the VLE curve is negligible. A
Nomenclature a = specific wetted area available for mass transfer, m2/m3 al = concentration of C02 in liquid, kmol/m3 a d = specific surface area of dry packing, m2/m3 B = chemical base b l , b2 = concentrations of C032-and HC03- ions bj = concentration of jth component in liquid C = total amount of amines added to the solution C, = constant c , k = average specific heat of kth component in the gas phase for the temperature range 70-135 OC, kcal/ (kmol K) C p =~ molar specific heat of liquid, kcal/(kmol K) Cl,, = specific heat of liquid, kcal/(kg K) ,C, = molar specific heat of gas, kcal/(kmol K) C R = average concentration of active component, kg/m3 C ' = constant D = liquid diffusivity, m2/h Dgk = diffusivity of kth absorbed component in the gas, m2/h ds = equivalent packing size, m2 G = mass velocity of inert gas, kg/(m2 h) g = acceleration due to gravity, m/h2 H = Henry's law constant in the reactive solution, (atm m3)/ kmol H o = value of H in pure solvent AHl = heat of absorption and reaction of COzwith hot potash, kcal/kg AHz = heat of vaporization of water, kcal/kg h = heat-transfer coefficient in the gas-liquid interface, kcal/(m2 h K) hk = heat-transfer coefficient of kth component, kcal/(m2 h K) I = enhancement factor Ii = ionic strength of solution [K+] = potassium ion concentration K,, = first ionization constant of carbonic acid Kgk = overall gas-phase mass-transfer coefficient of the kth component, kmol/(m2 atm h) K p = protonation constant of base B kgk = gas-phase mass-transfer coefficient of kth component, kmol/(m2 h atm) kk = liquid-phasemass-transfer coefficient of kth component, kmol/(m2 h (kmol/m3)) kII = second-order reaction rate constant 4 = molar flow rate of liquid, kmol/h L = liquid mass flow rate, kg/h m = molarity of the liquid solution Nk = moles of kth component transferred per unit time per unit length, kmol/(h m) Pko = partial pressure of the kth component in the bulk of the gas, atm Pk? = partial pressure of the kth component in the gas phase if it were in equilibrium with the bulk liquid phase, atm q = exponent in VLE relationship R = universal gas constant r = rate of reaction, kmol/(h m2) S = area of cross section of the tower, m2 s = specific gravity of the liquid T = gas temperature, K t = liquid temperature, K tc = liquid temperature, "C t~ = liquid temperature, OF V = molar flow rate of gas, kmol/h w = total equivalent weight percent of KzC03 x = fractional saturation of the liquid x, = fractional conversion of K2C03to KHC03 x k = mole fraction of the kth component in the liquid phase in unreacted form
Ind. Eng. Chem. Res. 1988,27, 2156-2161
2156
yk = mole fraction of the kth component in the gas phase z = spatial variable along the height of the column, fractional
Greek Symbols A, = thermal conductivity of gas, kcal/(m h K) 1.1 = liquid viscosity, kg/(m h) 1.1~= gas viscosity, kg/(m h) 1.1,+, p a = viscosities of the ith and jth
components in the gas,
kg/(m h) p = density of the liquid, kg/m3 pg = density of the gas, kg/m3 Registry No. NH,, 7664-41-7; COz, 124-38-9;K2C03,584-08-7.
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Received for review May 29, 1987 Revised manuscript received June 30, 1988 Accepted August 3, 1988
GENERAL RESEARCH Use of Ferrous Chelates of SH-Containing Amino Acids and Peptides for the Removal of NO, and SO2 from Flue Gas Shih-Ger Chang,* David Littlejohn, and David K. Liu Applied Science Division, Lawrence Berkeley Laboratory, University of California, Berkeley, California 94720
The use of ferrous complexes of SH-containing amino acids and peptides for the removal of NO and SOz in wet flue gas clean-up systems is reported. The ferrous chelates investigated in the present study include those of cysteine, N-acetylcysteine, penicillamine, N-acetylpenicillamine, glutathione, and cysteinylglycine. Compared to conventional chelates such as EDTA, these thioamino acids/ peptides not only can stabilize ferrous ion in alkaline solutions to promote the absorption of NO but are also capable of rapidly reducing any ferric ions formed during the scrubbing process back to ferrous ions so that continual absorption of NO can be achieved. In the case of ferrous cysteine and ferrous penicillamine, most of the absorbed NO is reduced to N2. The disulfide form of several of the thioamino acidslpeptides produced upon oxidation can be conveniently reduced by SO2and H2S to regenerate the starting materials, thus making possible the recycling of the reagents. The wet limestone system is currently the most widely used flue gas desulfurization scrubber in the utility industry. This system is very efficient in SOz removal; 0888-5885/88/2627-2156$01.50/0
however, it removes only a little NO,. This is because most of the NO, in flue gas is NO, which is only slightly soluble in aqueous solution. Work has been conducted at several 0 1988 American Chemical Society
