Ind. Eng. C h e m . Res. 1990,29, 1676-1681
1676
Verification of the Design Methods for Industrial Carbon Dioxide-Triethanolamine Absorbers: Laboratory Differential Simulation and Computational Methods Viclav Linek* and Ji*i Sinkule Department of Chemical Engineering, Institute of Chemical Technology, 166 28 Prague 6, Czechoslovakia
Miroslav Richter Chemical Works, 280 24 Kolin, Czechoslovakia
Jiri PospiLil Chemoprojekt, 1I 1 85 Prague 2, Czechoslovakia
Two design methods of the absorption column for COPremoval by a triethanolamine (TEA) solution a t elevated pressures are compared to existing industrial tower. The packing height was calculated from first principles (using physicochemical data and reaction kinetics from the literature) and by the differential simulation method. The latter method is based on the local absorption rate determined experimentally in a stirred cell for various compositions of the liquid and gas phases that are in contact in cross sections a t various heights in the industrial column. This method has the advantage that no physicochemical data are needed as the experiments are performed with the industrial absorption solution under conditions identical with those in the industrial column. The actual packing height is 24.5 m, the simulation method yielded 23.1 m, and calculation from first principles gave 29.0 m. The calculation does not take into account the presence of diethanolamine (ca. 3 wt 7%) in the industrial TEA solution.
Introduction The packing height in an absorption column can be calculated either on the basis of first principles or on the basis of the measurement of the rate of mass transfer in the given absorption system carried out in a laboratory apparatus under the conditions expected to be found in the industrial column. The design of absorption columns from first principles using physical data on the absorption system and processing data on the packing material taken from the literature is often a risky process. An industrial absorption solution is usually a mixture of substances for which the required data cannot be calculated even when the values are known for the individual components. The kinetic data for the accompanying chemical reactions often simply do not exist or are available for only a limited range of conditions. Empirical simulation methods are then practically the only feasible approach for the calculation of the column packing height. Differential Simulation of a Packed Column Absorber The most advanced method is that of differential simulation proposed by Danckwerts and Gillham (1966) and developed by other authors (Danckwerts and Alper, 1975; Laurent and Charpentier, 1983). The method is based on the experimentally determined mass-transfer rates of the absorbed component, N , in the laboratory model absorber, where the physical mass-transfer coefficients, k, and k?, are the same as in the industrial column. The concentrations and temperatures of the liquid and gas phases correspond to those in the column, and they are calculated from mass and enthalpy balances; see Figure 1. The packing height, H, for the required gas exhaustion can be found by numerical integration of the measured local absorption rates along the column
H
=
$iB':'!!!?! N
(1)
Obviously, empirical simulation methods do not require the knowledge of any physicochemical data on the absorption system; however, conditions must be maintained that ensure that absorption rates measured in the laboratory model can be transferred to the industrial equipment. If the absorption is accompanied by a slow chemical reaction that occurs not only in liquid film but also in the bulk of the liquid phase, an additional condition must be maintained (Laurent and Charpentier, 1983): the ratio of the interfacial area and the liquid holdup must be the same in both types of apparatus. When the absorption is accompanied by a fast chemical reaction, where all the absorbed gas reacts within the liquid film, this condition becomes irrelevant. This is true, for example, for the absorption of COPby an NaOH solution and hot potassium and ethanolamine solutions. Provided that the masstransfer resistance in the gas phase is neglibible compared to that in the liquid phase, then it is not necessary to adjust the same value for the coefficient of mass transfer in the gas phase, k and thus, a pure gas can be used for absorption in t8'e model apparatus. In such a case, the only parameter that must have the same value in both types of apparatus is kt. Systems with negligible resistance in the gas phase are encountered very frequently and include COz scrubbers working at elevated pressures. A stirred cell in which absorption occurs through the surface of a stirred liquid is most often used as a model absorber. The stirred cell has been used by a number of authors (Danckwerts and Gillham, 1966; Jhaveri and Sharma, 1969; Shaffer et al., 1974; Charpentier, 1976) to simulate the conditions in small packed columns: H I 1.92 m and d 5 0.3 m. These authors used such an arrangement in a stirred reactor, permitting adjustment of kt values attained in industrial packed columns, i.e., ca. 2 X lo-' m/s. Two stirrers were employed in the liquid phase; the lower stirrer at the bottom of the reactor stirs the bulk of the liquid while the second is just skimming the surface. In this arrangement, the k f value is very sensitive to adjustment of the surface level and requires optical control.
0888-5885/90/2629-16765O2.50,IO. 0 1990 American Chemical Society T~
Ind. Eng. Chem. Res., Vol. 29, No. 8, 1990 1677
t
i
I
3
k? value
In the industrial-----column
stirred cell gas
'YL Liquid
packed absorption column
Figure 1. Simulation of the local absorption rate in an industrial column using a stirred cell.
Thus, the apparatus was made of glass and cannot be used to simulate C 0 2 absorption in hot solutions at elevated pressures. We used a stirred reactor as the model absorber with only one stirrer located in the bulk liquid phase, eliminating the need for precise adjustment of the relative positions of the stirrer and the liquid surface. It is assumed that the interfacial area equals the cross section of the reactor. The intensity of stirring is restricted to stirrer rate per minute (rpm) that does not disturb the surface, which would increase the interfacial area. As a result, the maximum rZp value that can be attained in the reactor is limited. This value is usually lower than that attained in packed columns. Thus, the following procedure was employed: first the dependence of k," (physical absorption) on the stirrer rpm, f , was measured in a nonreactive system and modeled as kf' 0: f". Extrapolation of this dependence yielded the value of the fictive stirring rpm, f p , at which a k," value would be attained in the laboratory reactor, corresponding to the value in the industrial column. Then the dependence of the absorption rate for the industrial absorption system (chemisorption) on the stirrer rpm was measured and modeled as N a f '". Extrapolation of this dependence to the fictive stirring rpm yielded the absorption rate corresponding to the conditions in the industrial column. This procedure was repeated for all simulated cross sections in the column. A scheme of this procedure is given in Figure 2.
Design of a Packed Column from First Principles The height of packing is calculated from eq 1, but the local absorption rates are not determined experimentally but are calculated from model equations of mass transfer. The input data for the calculation are the flow rate and composition of the gas at the inlet; the flow rate, composition, and temperature of the liquid at the inlet; the required exit gas composition and data characterizing the packing material used (i.e., the wetted packing surface area and the coefficient of mass transfer in the gas and liquid for physical absorption under the hydrodynamic conditions considered); the cross section of the apparatus; and the overall pressure. The following assumptions are included: (a) The only transferred components are C02 (A) and H20. The partial pressure of water is in equilibrium with the liquid at the given composition and temperature. The other gas components, e.g., H2, N2 (I), are inert, and their chemical nature is important only for enthalpy balancing.
1 1
I
1.
10. f
[s-41
I
100
Figure 2. Illustration of the use of the simulation empirical method: (top) dependence of the mass-transfer coefficient of argon absorbed into the TEA solution in the stirred reactor, k?, on the stirring rate, f ; (bottom) dependence of the mass flux of COz into the TEA solution in the stirred reactor, N , on the stirring rate, j , for the temperature and phase composition a t some points in the industrial column (see Table I).
(b) A countercurrent arrangement and piston flow of both phases is considered. (c) The absorption is isobaric and either isothermal or adiabatic. (d) The film model of mass transfer is used. (e) The accompanying reversible chemical reaction between C 0 2 (A) and triethanolamine (TEA, B) follows the stoichiometric equation k
A+B--?'P (2) The chemical reaction that accelerates the interfacial transfer occurs only in the liquid film. In the bulk of the liquid, its components are in equilibrium. The rate of the reversible reaction (2) is given by Astarita et al. (1983) as = MBI([AI - [AI+) (3) where [A] is the local concentration of COz in the liquid film and [A]+ is the concentration of COP that would equilibrate the local liquid composition ([B] and [PI). The rate constant of the reaction, k2, was taken from Hikita et al. (1977) as log k, = 10.72 - 2688/T (4) The local equilibrium concentration, [A]+,was calculated from the ionic equilibria given by Jou et al. (1985). This procedure yields the equilibrium partial pressures of C02 over the TEA solution dependent on the temperature, concentration, and C 0 2content in the liquid solution. The C 0 2 pressure was recalculated to the equilibrium molar concentration by using the solubility of COz in water according to Danckwerts and Sharma (1966) log HAW = 7.30572 - 1140/T (5)
1678
Ind. Eng. Chem. Res., Vol. 29, No. 8, 1990 The resultant relationship derived by using the diffusivity of COPin water at 25 "C, DM = 1.99 X m2/s, has the form
Interface
I
liquid
DM([BO],T) = -~
[
~
I
I
'dl
b'
I1
I I bulk of
90s
I I
,
6.677 X 10-'2(1 - 0.2185[Bo])T
I
' u =o liquid
'
IfiLm
x
Pd [BO1,T)
]
2'3
(16)
Similarly, the data of Hikita et al. (1980) were employed to correlate the diffusivities of TEA measured at 25 "C in the form DB(25 "C)= 7.7
X
10-"(1
+ 0.1537[B0]'.2726)-2
(17)
to yield the following relationship for different temperatures:
I
gas film
pl([Bo],25"C)
bulk of liquid
X=\
Figure 3. Sketch of concentration profiles of the rcacting components in the liquid film.
The correction for the presence of TEA was estimated from the data of Sada et al. (1977) on the solubility of N20. The data were modeled as follows: In (HAw/HA) = 4.1277 X - 0.19532[B0] + 2.9327 X 10-2[B0]2- 9.4887 X 10-3[B0]3(6) The concentration distribution of the reacting components in the liquid film (see Figure 3) was obtained by solving the following set of differential equations:
DM d2[A]/dx2= r
(7)
D1B d2[B]/dx2 = r
(8)
DlP d2[P]/dx2= -r with boundary conditions atx=O - D M d[A]/dx = kg(PA - f f ~ [ A ] = ) N d[B]/dx = d[P]/dx = 0 atx=1 [AI = [ A I b [B] = [Bib [PI = [ p l b The thickness of the liquid film is given as
1 = DM/k?
where K = HAk?/k,. The local absorption rate is then given as
(10)
(21) The values of the physical mass-transfer coefficients are given below.
(11)
(12)
(13)
which replaces eq 9. The diffusion coefficient for COPin aqueous solutions was estimated from the values published by Sada et al. (1978) for N20 diffusivities at 25 "C. These values were linearized by using the relationship DINZ0(25 "C) = 1.78 X - 3.89 X 10-'O[B0] (14) For different temperatures, the relationship proposed by Hikita et al. (1980) was used D 1 ~ , 2 /T3 /= constant
The acceleration of absorption by the reaction is given by the overall enhancement factor, E,, which is the ratio of absorption rate with and without the reaction for the same driving force E, = E(l + K)/(1 + K E ) (20)
(9)
The concentrations in the bulk of the liquid, [A]b, [B]b, and [P]b, are the equilibrium values with respect to the CO, content in the liquid solution. These values were calculated from the ionic equilibria (Jou et al., 1985). The set of equations was simplified by introducing the assumption that the diffusivities of TEA and of reaction product P are identical. It then holds that
P I + [PI = [BO1
The viscosities used are given below. The solution of the ordinary differential equation set (71, (8), and (101413) yields the enhancement factor, E , defined as the ratio of the rate of absorption with reaction to the rate of physical absorption at the same concentration of C02 at the phase boundary [A],
(15)
Industrial Absorption Column The correctness of the two design methods was tested by simulating the design of an existing industrial packed absorption column employed in the Chemical Works in Litvinov, Czechoslovakia, for scrubbing of C02at elevated pressures with a 30% TEA solution containing also 3% diethanolamine (DEA) and traces of monoethanolamine (MEA). Column parameters: cross section, S = 3.57 m2;packing, 5 layers of 25-mm metal Pall rings, each with a volume of 17.5 m3; height of packing, H = 24.5 m; overall pressure in the column, p = 2.6 MPa. Input gas: flow rate, 65-150 N m3/h; temperature, 70 "C; C 0 2 content, 8.65 vol 9'0; residue practically only hydrogen. Output gas: flow rate, not measured; temperature, 47 "C, C 0 2 content, 0.12 vol TO;pressure drop, 6.5 kPa. Input liquid: flow rate, 356 m3/h; temperature, 47 "C; TEA concentration, 27.5 wt % ([BO] = 1.902 kmol/m3); C 0 2 content, 1.5 N m3/m3 of solution (0.035 kmol of CO,/kmol of TEA). Output liquid: the flow rate along the column does not change; temperature, 58 "C.
Ind. Eng. Chem. Res., Vol. 29, NO. 8, 1990 1679 14
The following temperature dependencies of the viscosity and water vapor pressure (above the solution freed of all dissolved COP)were obtained for the industrial TEA solution: In = -10.448 + 740.82/(T - 146.26) (22) In pw= 25.75671 - 5309.9/T
(23)
The specific geometric surface area of the metal Pall rings employed (manufactured by the Chemical Works) was found to be a, = 204 m-l. The physical mass-transfer coefficient was calculated from the correlation of Onda et al. (1968a)
A mean column temperature of 53 "C was considered. The wetted fraction of the packing surface, aw/a, = 0.785, was calculated from the correlation of Onda et al. (1968a) for a well-wetted packing surface (a = a, = 0.071 kg/s2)
Figure 4. Scheme of the apparatus: 1, solution storage tank; 2, desorption column; 3, liquid volumetric pipet; 4, tempering vessel; 5, stirred cell; 6, volumetric cylinder with piston; 7; contact U manometer; 8, gas storage tank; 9, pressure bottle; 10, motor function control; 11, motor; 12, vacuum; 13, separating valve; 14, precise manometer.
= 1 - exp(-l.45(~,/a)~/~Re~'Fr~~~~~We~~) (25) drops due to absorption while the pressure in the gas tank remains constant. As a result, electric contact in U maThe value kf = 2.18 X m/s was found from eq 24. nometer 7 is disconnected, motor ll is switched on, and The mass-transfer coefficient in the gas phase, k, = 2.85 piston 6 is pushed into the cylinder. After equalizing the X lo* mol/(m2 s Pa), was calculated from the correlation gas pressure in the cell and the gas tank, the contact is of Onda et al. (196813) restored and the motor is stopped. The diameter of the piston was 10 mm for physical absorption in contrast to 30 mm used for much faster chemisorption. The absorption rate is determined from the movement of piston 6, the position of which, 2, is recorded as a for a mixture of 8 vol % CO,, 92 vol 70 H, at 50 "C, and function of time, t. 2.6 X lo6 Pa (M,= 5.4 X loT3kg/mol* p = 5.3 kg/m3; pg(H2)= 9.5 x IO* Pa s, DgA = 2.18 x IOk k2/s, mg= 4.36 Evaluation of Physical Absorption kg/s). The resistance to the transport of C02 in the gas phase (l/HAkg)= 39.7 s/m) is only 0.87% of the resistance The physical mass-transfer coefficient of COPabsorbed in the liquid phase (l/k," = 1/2.18 X lo4 = 587 s/m) and into the TEA solution cannot be measured directly. Consequently,an inert gas, argon, was used instead, as the can thus be neglected. The accompanying reaction between TEA and COBis sufficiently fast so that the gas is diffusivities of both gases are almost the same. For example, for water D M = 2 X lo4 and DIAr= 2.02 X lo* mz/s entirely consumed within the liquid film. As no carbamate is formed, its subsequent hydrolysis in the bulk of liquid (Vivian and King, 1964; Baird and Davidson, 1962). does not occur. Thus, the ratio of the interfacial area and The mass-transfer coefficient, k t , for argon was evaluthe liquid holdup is not decisive for the absorption of C02 ated by using the method that does not require knowledge into a TEA solution. Then the coefficient kf is the only of the solubility of argon in the liquid. The solution of the parameter that must be simultaneouslymade equal for the balance equation for argon in an ideally stirred batch reindustrial column and the laboratory reactor. actor, a,/a,
Experimental Section A scheme of the apparatus is depicted in Figure 4. The equipment was made of stainless steel and can be used for measurements at absolute pressures up to 0.8 MPa and at 150 "C. The reactor is a cylindrical vessel with a diameter of 0.08 m, liquid volume of 250 cm3, and total reactor volume of 900 cm3. The turbine stirrer (with a diameter of 0.03 m) with six blades is driven by a magnetic drive. Tank 8, which is a cylindrical vessel with a volume of 3 L, serves as a constant-pressure source, with which the pressure in the reactor is compared through a contact U manometer. The closing or opening of the circuit between the electrodes through the motion of the manometer liquid controls the movement of piston 6. The experiment begins by filling gas tank 8 and cylinder 6 with gas and evacuated cell 5 with absorption liquid. After a constant temperature is attained (after about 2 h), valve 13 is opended and gas from tank 8 is introduced into cell 5. When the pressure and temperature are equalized in the whole apparatus (the liquid in the cell is not agitated to keep absorption minimal), valve 13 is closed again and the stirrer is switched on. The pressure in the stirred cell
k,"A([Ar], - [Ar]) = V, d[Ar]/dt
(27)
with initial condition [Ar], = 0 yields the dependence of argon concentration in the liquid on the duration of the experiment [Ar] = [Ar],(l - exp(-k,"At/ VI))
(28)
The amount of gas absorbed is directly proportional to the position of the piston and thus eq 28 can be rewritten in the form 2 = Z,,[l
- exp(-kfAt/VJ]
(29)
Z,, is the piston position for a saturated solution. The coefficient, k,", was evaluated by fitting eq 29 with 2 and T values from the experiment. The dependence of k," on the stirrer rpm was measured at 53 "C and an overall pressure of 495 kPa. The results plotted in Figure 2 are correlated by the following equation:
kf = 9.721 X 10-6f0.7069
(30)
The value of the fictive stirrer rpm at which k," in the stirred reactor would equal that in the industrial absorp-
1680 Ind. Eng. Chem. Res., Vol. 29, No. 8, 1990 Table I. Values of Characteristic Quantities at Simulated Cross Sections in the Industrial Column Y , mol 103~, COz/mol mol m-2 i p A , kPa TEA T , "C p w , kPa lo4& s-l 3.14 0.0354 47.0 47.3 9.72 3.663 2.69 0.0438 8.66 48.0 10.1 6.061 25.6 4.44 0.0699 48.8 10.6 7.452 5.47 0.0942 41.3 49.8 11.1 8.709 6.39 0.1289 63.2 50.9 11.7 9.084 0.1649 85.6 6.66 0.2013 108.0 1.23 51.9 12.3 9.853 54.0 13.7 8.766 0.2713 150.0 6.43 55.5 14.8 7.504 177.0 5.50 0.3194 58.1 16.8 4.909 0.4072 227.0 3.60 0.4041 225.0 58.0
tion column (kf = 2.18
X
m/s) is thus f, = 81.4 s-l.
Evaluation of the Chemisorption The rate of chemisorption decreases during the experiment as a result of increasing COz content in the batch of TEA solution. The instantaneous absorption rate was calculated from the measured pairs of values of 2 and t
The overall change in the position of the piston from the beginning of the experiment corresponds to a certain increase in the COz content in the batch, y , for which it holds that S,Z (32) The following values of the pertinent quantities were emm2; TEA concenployed: interfacial area, A = 5 X tration, [BO]= 1.902 kmol/m3; piston cross section, S, = 7.0686 x m3; VA, the molar volume of COz under the conditions in the cylinder, m3/mol (calculated from the van der Waals equation with a = 0.364 Pa m6/mo12and b = 42.7 X lo* m3/mol). To adjust a given value of the partial pressure of COz in the laboratory reactor, PA,it is necessary to know the water vapor pressure above the TEA solution, pw,as the overall pressure in the reactor, p , is given by the sum of the partial pressures of the two components, p = pA + pw. The water vapor pressure over the industrial TEA solution was calculated from eq 23. The initial COz content in the batch, yo, is chosen so that the value of y, for which the N should be determined, lies between the initial and final COz content in the solution in the experiment. The final C02 content was determined gravimetrically (Matous et al., 1969) at the end of each experiment. The industrial column was simulated at cross sections in nine heights within the packing. Table I lists the P A and y values corresponding to these heights. The temperature distribution along the column was considered to be linear with respect to the amount of CO, absorbed. The first and last lines in the table are the values at the top and bottom of the column and were not simulated. The measurement of the absorption rate for one point, i.e., for a single set of conditions @, T, y), was repeated for several stirring rpm. Some of the experimental results are plotted in Figure 2. The data were fitted by the following relationship: N = Bif0.4529 (33) Constants Bi are listed in Table I. The rates of absorption for fictive stirring rpm ( 4 1 . 4 s-l) were found from eq 33
I
I
r
2'
J--
/
/+ 1
0'
L -_--LA 03
02 [mol
m2 / m d
OL
TEA]
Figure 5. Dependence of the absorption rate of COz on the COz loading, obtained by a simulation empirical method for nine cross sections in the industrial column.
and are given in Table I; f, is the rpm at which the same hydrodynamic conditions would be attained in the stirred cell as in the industrial column. Thus, the N values obtained in this way represent the mass-transfer rates in the nine cross sections of the industrial column. Figure 5 depicts the dependence of this local rate on the local COz loading of the liquid.
Calculation of the Packing Height The height of the column packing was calculated by numerical integration of the empirical mass-transfer rates from Table I by using a modified relationship of (1) (34) yielding a value of 23.1 m, i.e., only 6% less than the actual packing height in the industrial column (24.5 m). Calculation from first principles yields a packing height of 29.0 m for adiabatic conditions and 29.7 m for isothermal conditions (for a mean temperature of 53 "C), i.e., by 18% and 21% greater values, respectively.
Conclusions The method of laboratory differential simulation yields a very good estimate of the packing height in an industrial column. The computational method based on first principles overestimates the column height by 20%. This difference is probably a result of the fact that the computational method does not take into account that the industrial solution contains not only TEA but also 3 wt 7'0 of the more reactive DEA. Acknowledgment We thank Raschig A. G., FRG, for support of this work.
Nomenclature A = interfacial area, m2 [A]+ = concentration of C02that would equilibrate the local liquid composition ([B] and [PI), kmol/m3 a, = specific geometric surface of the packing, m-l a, = wetted specific surface of the packing, m-l Bi = regression constant in eq 33, Table I [Bo] = equivalent concentration of TEA, kmol/m3 D = molecular diffusivity, m2/s d, = characteristic dimension of the packing, m E = enhancement factor defined by eq 19 E, = enhancement factor defined by eq 20 Fr = Froude number f = stirring rate, l / s f,= fictive value o f f for which the k p value in the stirred reactor would equal that in the packed column, l / s
Ind. Eng. Chem. Res., Vol. 29, No. 8, 1990 1681 g = gravitational acceleration, m/s2 H = packing height, m HA= solubility of component A, Pa m3/mol [i] = concentration of component i in the liquid phase: i = A (C02), B (TEA), P (reaction products), Ar (argon),
kmol/m3 k p = physical mass-transfer coefficient in the liquid phase, m/s k2 = rate constant of the accompanying reaction, m3/(s kmol) k, = mass-transfer coefficient in the gas phase, mol/(m2 s Pa) I = thickness of the liquid film, m M = molecular weight, kg/mol m = mass flow rate, kg/s N = absorption rate, mol/(m2 s) p = overall pressure, Pa pi = partial pressure of component i, Pa R = universal gas constant, J/(mol K) R e = Reynolds number r = rate of reaction, kmol/(m3s) S = column cross section, m2 = piston cross section, m2 = absolute temperature, K t = time, s VI = liquid volume, m3 VA = molar volume of component A, m3/mol VI = volume flow rate of the liquid, m3/s We = Weber number x = distance from the interface, m y = COPcontent in the liquid, mol of C02/mol of TEA 2 = piston position, m
?
Greek L e t t e r s I.C = viscosity, Pa s p = density, kg/m3 u = surface tension,
kg/s2
uc = critical surface tension, kg/s2 Subscripts
1 = input stream 2 = output stream A = absorbed gas component, C02 b = bulk of the liquid B = TEA g = gas phase 1 = liquid phase P = reaction products s = liquid interface w = water Registry No. C02, 124-38-9; triethanolamine, 102-71-6.
Literature Cited Astarita, G.; Savage, D. W.; Bisio, A. Gas Treating with Chemical Solvents; John Wiley: New York, 1983. Baird, M. H.; Davidson, J. F. Gas Absorption by Large Rising Bubbles. Chem. Eng. Sci. 1962, 17, 87-96. Charpentier, J. C. Recent Progress in Two Phase Gas-Liquid Mass Transfer in Packed Beds. Chem. Eng. J. 1976,11,161-181. Danckwerts, P. V.; Alper, E. Design of Gas Absorbers IIILaboratory Point Model of a Packed Column Absorber. Trans. Inst. Chem. Eng. 1975,53,34-40. Danckwerts, P. V.; Gillham, A. J. The Design of Gas Aborbers IMethods for Predicting Rates of Absorption with Chemical Reaction in Packed Columns, and Tests with 11/2in. Raschig Rings. Trans. Inst. Chem. Eng. 1966,42,T42-T54. Danckwerts, P. V.; Sharma, M. M. The Absorption of Carbon Dioxide into Solutions of Alkalis and Amines. Chem. Eng. 1966,202, CE244-CE280. Hikita, H.; Assai, S.; Ishikawa, H.; Honda, M. The Kinetics of Reactions of Carbon Dioxide with Monoethanolamine, Diethanolamine and Triethanolamine by a Rapid Mixing Method. Chem. Eng. J. 1977,13,7-12. Hikita, H.; Ishikawa, H.; Uku, K.; Murakami, T. Diffusivity of Mono, Di-, and Triethanolamines in Aqueous Solutions. J. Chem. Eng. Data 1980,25,324-325. Jhaveri, A. S.; Sharma, M. M. Absorption with Fast Chemical Reaction. Chem. Eng. Sci. 1969,24,189-191. Jou, F. Y.; Otto, F. D.; Mather, A. E. Equilibria of H2S and C02 in Triethanolamine Solutions. Can. J. Chem. Eng. 1985, 63, 122-125. Laurent, A,; Charpentier, J. C. The Use of Experimental Laboratory-Scale Models in Predicting the Performance of Gas-Liquid Reactors, Inst. Chem. Eng. 1983,23,265-275. Matoui, J.; Sobr, J.; Novak, P.; Pick, J. Solubility of Carbon Dioxide in Water a t Pressures up to 40 atm. Collect. Czech. Chem. Commun. 1969,34,3982-3997. Onda, K.; Sada, E.; Takeuchi, H. Gas Absorption with Chemical Reaction in Packed Columns. J. Chem. Eng. Jpn. 1968a,1,6246. Onda, K.; Takeuchi, H.; Okumoto, Y. Mass Transfer Coefficients Between Gas and Liquid Phases in Packed Columns. J. Chem. Eng. Jpn. 1968b,l,56-62. Sada, E.; Kumazawa, H.; Butt, M. A. Solubilities of Gases in Aqueous Solutions of Amine. J. Chem. Eng. Data 1977,22,277-278. Sada, E.;Kumazawa, H.; Butt, M. A. Solubility and Diffusivity of Gases in Aqueous Solutions of Amines. J. Chem. Eng. Data 1978, 23, 161-163. Shaffer, D. L.; Jones, J. H.; Daubert, T. E. Simultaneous Absorption and Chemical Reaction of Butenes. Znd. Eng. Chem. Process Des. Dev. 1974,13,14-19. Vivian, J. E.; King, J. C. Diffusivities of Slightly Soluble Gases in Water. AIChE J. 1964,10,220-221. Received for review June 28, 1989 Revised manuscript received March 12, 1990 Accepted March 28, 1990
