Absorption of Carbon Dioxide in Aqueous Morpholine Solutions

Oct 14, 2013 - Mass transfer coefficients for CO2 absorption by aqueous 4 M, 5 M, and 6 M morpholine (MOR) solution at 30 °C under carbon loaded cond...
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Absorption of Carbon Dioxide in Aqueous Morpholine Solutions Kun Liu, Krishna M. Jinka, Joseph E. Remias, and Kunlei Liu* Center for Applied Energy Research, University of Kentucky, Lexington, Kentucky 40511, United States ABSTRACT: Mass transfer coefficients for CO2 absorption by aqueous 4 M, 5 M, and 6 M morpholine (MOR) solution at 30 °C under carbon loaded conditions are presented. Measurements were performed using a wetted wall column under simulated flue gas conditions for postcoal combustion CO2 capture. Similar to other major amines, it was found that the amine concentration has a positive impact on the overall mass transfer coefficient at lean loading while at rich loading the difference between each concentration is within experimental uncertainty. It was also observed that the viscosity increases more dramatically with amine concentration in MOR than monoethanolamine (MEA) within the studied carbon loading range due to different hydrogen-bond structures. The calculation of physical mass transfer resistance in the liquid film revealed that the reaction resistance dominates at lean loadings while the physical mass transfer takes a large portion at the rich loadings. The rate constant of the pseudo-first-order reaction of MOR and CO2 was calculated to be three times higher than MEA, which is consistent with the experimental mass transfer data from the wetted wall column experiment.

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INTRODUCTION Aqueous amine solutions have been identified as the most mature CO2 capture technology for postcombustion flue gas treatment. Among the alkanolamines, MEA has been used as the benchmark because of its fast reaction, high cyclic capacity, and low cost. Despite the advantages of MEA, the cost to capture CO2 from coal combustion flue gas using such solvents is expected to increase the energy consumption by 30%−40%.1 A sizable portion of the cost increase (approximately 20%) is due to the capital cost of the CO2 capture absorber.2 An advanced solvent that has a faster absorption rate of low concentration CO2 has the potential to decrease the absorber height and cost. In addition, the other drawbacks of MEA, including high heat of reaction, low resistance to thermal and oxidative degradation, and corrosion to equipment, make the solvent potentially less attractive in a utility application. Therefore, there is a high demand to develop new solvent systems that have a reasonable combination of high absorption rate, high working capacity, low energy consumption, good thermal and oxidative stability, and low cost. Morpholine (MOR) could be one of the promising amine candidates for CO2 capture because of its fast reaction rate and high thermal stability.3 Similar to piperazine (PZ), the MOR molecule has a six-ring structure with a highly sterically accessible amine site yielding fast reaction with a CO2 molecule. Because of the attractive advantages of MOR, studies have been done on the kinetics of CO2 capture by MOR. Little et al.4 reported MOR kinetic data at 303 K obtained by using a stirred cell reactor and interpreted their results with the zwitterion mechanism. Stop-flow-based method studies were employed for obtaining the kinetic data for the reaction of CO2 and MOR by Crooks et al.5 and Alper.6 Furthermore, Al-Juaied et al.7,8 proposed a rate-based model combined with the thermodynamics model to predict the rate of CO2 absorption for DGA, MOR, and DGA/MOR blends. Unfortunately, most of the studies were done at diluted amine concentrations and no CO2 loading, conditions which are prohibitive in the direct application of the CO2 absorption in utility. Therefore, the © 2013 American Chemical Society

mass transfer data for CO2 absorption by MOR at conditions consistent with postcombustion CO2 capture is needed and is studied in this work.

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EXPERIMENTAL METHODS Overall Mass Transfer Coefficient. Overall mass transfer coefficient was measured by wetted wall column (WWC) which is a well-known method for the determination of the fundamental gas−liquid absorption process (laminar film). The essential of the WWC is a gas−liquid contactor in which CO2 absorption by an aqueous solvent occurs on a known surface area. The schematic of the WWC used in this test is shown in Figure 1. In a typical experiment, liquid solvent was first loaded to the mol-C/mol-N level of approximately 0.1 (the

Figure 1. Schematic of wetted wall column apparatus. Received: Revised: Accepted: Published: 15932

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lowest limit deployed in the potential commercial CO2 capture process) with CO2 by sparging the solution reservoir with a concentrated CO2/N2 mixture for a predetermined amount of time. The loaded solution was then heated to the desired temperature by circulating through a preheater unit at 180 mL/ min. Once the solution was thermally stable, a CO2 gas feed mixed with N2 at 6.6 L/min supplied from AALBORG MFCs with precision of 0.5% was first bubbled through a water saturator which is maintained at the tested temperature and then allowed to contact with the liquid countercurrent on the surface of the column. The column is a stainless steel hollow rod with a height of 15.3 cm and OD of 1.26 cm, which gives the surface area of 60.56 cm2. Gas is introduced into the wetted wall chamber through 16 orifices uniformly distributed around the rod at the bottom of the chamber to minimize the gas side mass transfer resistance. The pressure at the wetted wall column is maintained at slightly higher than atmospheric pressure to simulate a scrubber tower. Absorption or desorption of CO2 occurred across the contacting area, which gave a CO2 concentration difference in the gas stream between the inlet and outlet of the column. CO2 concentration was measured by a CO2 analyzer (VIA-510, HORIBA, 0.5% precision) downstream of the WWC. Water vapor is removed using an ice bath and a Drierite bed before the CO2 analyzer. Losing CO2 by condensation with the water and excess volatile MOR is avoided by passing the effluent gas through a 10 M H2SO4 acid trap before the ice bath. Flux and driving force can be obtained from the CO2 pressure difference. Four CO2 concentrations between (2−14% CO2) in the gas stream were tested at each carbon loading. Liquid samples were collected during the test for carbon loading, viscosity, and density measurements. The above procedure was repeated for various carbon loadings with the high concentrated CO2 sparging cycle repeated in between each to increase the solution carbon loading. The overall mass transfer coefficient at the operating condition can be calculated from eq 1. KG =

ΔPCO2 =

i i PCO = yCO (Ptotal − Pwater) 2 2

⎛ 7207 Pwater = exp⎜72.55 − − 7.139 ln(T ) ⎝ T ⎞ + 4.046 × 10‐6T 2⎟ ⎠

2

out yCO

2

yNout 2

(4)

(5)

The equilibrium partial pressure of CO2, P*CO2, was calculated by making the flux NCO2 to be zero at zero driving force through a trial-and-error routine in MATLAB. The gas side mass transfer coefficient in this work was obtained by measuring the overall mass transfer coefficients of unloaded MEA solvent whose liquid side mass transfer properties are well established.9 The liquid film mass transfer coefficient (1/k′g) was calculated from the pseudo-first-order reaction model. By subtracting calculated k′g from measured KG, the gas mass transfer resistances and coefficients (kg) at different gas flow rates were obtained. Total Inorganic Carbon Loading. The total inorganic carbon loading was determined by an in-house carbon loading apparatus which has been described previously.10 Briefly, it consists of a gas sampling tube with stopcock for holding a strong acid. The inlet and outlet of the tube connect to a pure N2 gas cylinder and a CO2 analyzer (VIA-510, HORIBA, 0.5% precision) respectively. A liquid sample of known mass was injected to the sampling tube through an injection port located at the middle of the tube. The CO2 loading was determined by integrating the area under the curve for CO2 gas liberated from the sample using the CO2 analyzer. The calculated value is also referenced to a calibration curve from known standards of potassium carbonate. The HORIBA CO2 analyzer was calibrated with a certified CO2:N2 gas mixture (PurityPlus, 14.00% CO2) each day. The uncertainty was checked with three known analytical standard (1.5, 1.8, and 2.5 M K2CO3) prior and after each set of unknowns with the allowable discrepancy in CO2 measurement (|(expected − measured)/expected| × 100) set at less than ±3% absolute. The standard deviation for repeated measurements was ±2.7%. To make a fair comparison among the solutions with different MOR concentrations, the carbon loadings reported in this work are normalized by C/N which is the ratio of carbon loading in mol C/kg soln to amine concentration in mol N/kg soln. Alkalinity. Alkalinity is the capacity of solution to neutralize acid (CO2 in this case). Since MOR is the only basic amine to react with CO2 in this study, the alkalinity is mathematically equal to the MOR concentration. The alkalinity of a solution was measured by an automatic titrator (Metrohm Titrando 836). In each analysis, the diluted sample was titrated with 0.1 M sulfuric acid until the pH is below 2.5. The volume of the acid at the final equivalence point was used to calculate the

in which NCO2 is the flux of CO2, KG is the overall mass transfer coefficient, ΔPCO2 is the log mean of driving force, and A is the gas−liquid contacting surface area. The flux is calculated by the CO2 concentration difference at the inlet and outlet of the WWC as shown in eq 2.

2

(3)

in which Ptotal is the total pressure and yCO2 is the mole fraction of CO2 measured by the CO2 analyzer at dry base. As the feed gas was saturated with water in the saturator, the partial pressure of water Pwater, the saturation pressure at the testing temperature T, can be written as

(1)

in out in NCO2 = NCO − NCO = yCO Ntin − yNin Ntin 2 2

in * ⎞ ⎛ PCO − PCO ln⎜ P out2 − P* 2 ⎟ ⎝ CO2 CO2 ⎠

The partial pressure of CO2, PCO2, is calculated by eq 4.

NCO2 AΔPCO2

in out − PCO PCO 2 2

(2)

in which the molar flow rates Ni of component i were calculated from ideal gas equation of states, and yi is the molar fraction of component i. Since the CO2 dynamically transfers from gas phase to liquid phase, the driving pressure of CO2 decreases along the WWC. To better represent the true average differential pressure of CO2 for absorption in the column, log mean of the driving forces was taken at the inlet and the outlet of the column as indicated in eq 3. 15933

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The liquid film thickness δ can be calculated by

alkalinity of the sample. The standard deviation for repeated measurements was ±0.7%. Density. The density of a solution was determined by measuring the weight of a known volume liquid on an analytical balance (PA-220, Symmetry). Samples were measured in quadruplicate. The volume of the liquid was defined by a pipet (1000 μL, Eppendorf). Viscosity. The kinematic viscosity was measured using a glass capillary viscometer (model 100, Cannon Instrument) in a temperature-controlled water bath. In each test, the viscometer was secured in a water bath whose temperature was adjusted to the tested temperature in the WWC experiment. Liquid was injected to the viscometer and stabilized for 10 min to reach thermal equilibrium. The time taken for the liquid level to pass between two marks on the viscometer was converted to kinematic viscosity using a vendor-supplied equation. Dynamic viscosity was obtained by multiplying the kinematic viscosity by the density of the liquid.

δ=

(11)

⎛ α α2 ⎞ B = exp⎜A + +C +D ⎟ Δ[CO2 ] T T T⎠ ⎝ * ΔPCO 2

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⎛ 1 α⎞ 1 × ⎜C + 2D ⎟ ⎝ T T ⎠ [MOR]

(12)

RESULTS AND DISCUSSION Liquid Side Mass Transfer Coefficient Comparison with Literature. To validate the WWC experimental method, the overall mass transfer coefficient for 30 wt % MEA at 40 °C was measured on WWC as proof-of-method due to its prevalence in the literature for comparison. The data collected from WWC experiments are overall mass transfer coefficients which include the resistances from both gas side and liquid side. The kinetic property of a solvent is decided by liquid side mass transfer coefficient, while the gas side mass transfer coefficient may be different in each apparatus and method. To make a fair comparison with the literature data, the liquid side mass transfer coefficient was separated from the overall mass transfer coefficient using eq 7. Liquid side mass transfer coefficients 30 wt % MEA at 40 °C were compared with the available literature data12,13 in Figure 2. It can be seen that the result from this

Resistance in eq 6 can also be written as the form of the reciprocal of mass transfer coefficients in eq 7.

(7)

in which KG is the overall mass transfer coefficient, kg is the gas side mass transfer coefficient, and kg′ is the liquid side mass transfer coefficient which can be further separated into the reaction resistance (1/k″g ) and the physical mass transfer resistance (1/k°l,prod (ΔP*CO2/Δα [MOR])). The reaction resistance term (1/kg″) accounts for the resistance from the chemical reaction of CO2 and MOR in the reaction film, which depends on the reaction rate constant and reactant concentration. The physical mass transfer resistance term describes the resistance due to the diffusion of reacted CO2 in the liquid diffusion film. To evaluate the physical mass transfer resistance, it is necessary to calculate the physical mass transfer coefficient k°l,prod and the equilibrium slope ΔP*/Δ[CO2]. A model proposed by Hobler11 in eq 8 is adopted to calculate the kl,prod ° . Q η 3 A π

(10)

2⎞ ⎛ * = exp⎜A + B + C α + D α ⎟ PCO 2 T T T⎠ ⎝

THEORETICAL BASIS The overall mass transfer resistance of CO2 in a gas−liquid absorption process can be written as the contribution from three aspects as shown in eq 6. The first term in eq 6 represents the CO2 diffusion resistance in the gas phase. The second term refers to the reaction resistance of CO2 and aqueous amine. The last term is the physical mass transfer resistance in the liquid phase. R overall = R gas + R reaction + R physical (6)

kl°, prod =

3μQ ρgW

in which μ is the viscosity of solution, ρ is the density of solution, and W is the circumference of the WWC. The equilibrium slope (ΔP*CO2/Δα [MOR]) is determined by the thermodynamics of the system and is a function of temperature and carbon loading. To simplify the problem, an empirical equation that describes the relationship of partial pressure of CO2, temperature and carbon loading in eq 11 is adopted in this work. The equilibrium slope is calculated by taking the derivate of eq 11 with respect to carbon loading α and multiplying the amine molarity as shown in eq 12. The coefficients in eqs 11 and 12 are regressed using thermodynamic data collected previously.3

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* ΔPCO 1 1 1 1 1 1 2 = + = + + KG kg kg′ kg kg″ kl°,prod Δα ·[MOR]

3

(8)

in which Q is the liquid flow rate, A is the surface area of WWC, and η is the dimensionless penetration distance that can be written as

η=

Dprod h δ 2us

(9) Figure 2. Comparison of liquid side mass transfer coefficients of CO2 in 30 wt % MEA at 40 °C from this work (blue ●), Aboudheir, Tontiwachwuthikul et al. (2003, cited in Dugas, 2009)13 (purple ◆), Dang and Rochelle 200312 (green ▲), and Dugas 200913 (red ■).

where Dprod is the diffusion coefficient of product adopted from literature,8 h is the height of WWC, and us is the liquid film velocity. 15934

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work matches well with the literature data, especially at the rich loading. At lean loading, some discrepancies among the literature data are observed. Since the liquid mass transfer coefficient is obtained from the overall mass transfer coefficient, the relatively high gas side mass transfer resistance at lean loading in experiments at various gas velocities may result in a higher error in the kg′ calculation. Despite of the discrepancies in the literature data, the data from this work is in the range of the reported data. Thermodynamic Correlation. The thermodynamics properties of the MOR solution were studied in an earlier work.14 The CO2 partial vapor pressure for carbon loaded 34.8 and 43.5 wt % MOR solution at 40 °C, 60 °C, 100 °C, and 120 °C was measured and fit into an commercial electrolyte-NRTL simulation model. Although the rigorous model accurately represents the thermodynamic properties of the MOR−H2O− CO2 system, it is mathematically difficult to extract the simple relationship of CO2 partial pressure and carbon loading. Naser et al.14 used an empirical relationship proposed by Xu and Rochelle15 to fit the vapor−liquid equilibrium experimental data and calculated the heat of absorption. Although the large number of coefficients (six) in Xu and Rochelle’s empirical equation helps represent the true correlation of CO2 partial pressure and carbon loading. They will require a large data set to minimize the relatively large error in the coefficient regression which will pose a significant cost for a solvent evaluation. Therefore, to reduce the level of effort and minimize the regression error for a small set of data, the number of variables in the empirical equation was controlled to four through a sensitivity study in this work. The final expression of the empirical equation is shown in eq 11. However, the potential problem for such multivariable nonlinear regression is that the optimization search would be trapped in local minima. To solve the local minima problem, a pattern search which does not rely on the gradient of the objective function was selected here to ensure the global minimization. The regression was performed in MATLAB 2012b Global Minimization Toolbox. Standard error was also calculated from the 95% confidence interval. The best-fit coefficients in eq 11 for the MOR system as well as their standard deviation are listed in Table 1. Even

Figure 3. CO2 partial pressure above aqueous MOR solvent at 40, 60, 100, and 120 °C14 and the corresponding empirical correlation.

water, is selected in this work. As it can be seen in Figure 4, the overall mass transfer coefficients for all three studied

Figure 4. Overall mass transfer coefficients of CO2 in 4M, 5M, and 6 M MOR and 5 M MEA with respect of carbon loading at 30 °C.

concentrations decrease gradually as the solvent takes up CO2. As CO2 is absorbed in the solvent, the amine is depleted making the reaction rate drop dramatically, which decreases the mass transfer coefficient. In addition, the formation of carbamate and protonated amine because of the reaction of CO2 and MOR increases the ionic strength, which reduces CO2 solubility (Henry’s constant) and causes the viscosity increase accordingly at rich loading as shown in Figure 5. The increased viscosity results in a high CO2 physical diffusion resistance as indicated in eqs 8, 9, and 10. The combination of the increasing

Table 1. Best-fit Parameters for Empirical Correlation of CO2 Partial Pressure and Carbon Loading coefficients A B C D

best-fit value 30.98 −8440 3657 212

± ± ± ±

0.55 258 725 974

though fewer variables are included in the equation, they can adequately represent the partial vapor pressure of CO2 above the MOR solution as shown in Figure 3 with a reasonably narrow error. The empirical equation with the regressed coefficients in Table 1 will be used in the kinetics discussion later. Overall Mass Transfer Coefficient of MOR. The overall mass transfer coefficients of CO2 were measured in 4 M, 5 M, and 6 M aqueous MOR solution at 30 °C using the WWC. Because of the high volatility of MOR, such solvent most likely will be operated at a relatively low temperature to minimize the amine loss in the absorber. Therefore, a temperature of 30 °C, a reasonably low temperature that can be achieved by cooling

Figure 5. Dynamic viscosity of carbon loaded 5 M aqueous MOR and MEA at 30 °C. 15935

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in reaction resistance and diffusion resistance leads to the drop in overall mass transfer coefficient with carbon loading. Effect of MOR Concentration. High amine concentration is usually preferred in a CO2 capture process as it may lead to a high cyclic carbon loading per unit mass of solvent. But it is limited by the elevated corrosion rate and viscosity in practice. As shown in Figure 5, the viscosity increases more dramatically with the increasing of amine concentration for MOR than MEA at similar concentration. This is due to the structure difference between the two amine molecules. Both amines can form a hydrogen bond through the amine site and the hydrophilic sites (−OH on MEA and −O− on MOR). MOR can form an intermolecular hydrogen bond due to the rigid six-member ring structure, while the hydrogen bond in MEA solution can be intramolecular because of the linear molecular structure. The intermolecular forces in MOR solution make the solution more resistant to flow than MEA, which results in a higher sensitivity of amine concentration on solution viscosity. A high viscosity is usually unwanted in an aqueous amine packing scrubbing system as it negatively impacts the physical diffusion rate, heat transfer, and pumping energy. Of course, as expected, amine concentration also has direct and indirect (through diffusivity, viscosity, and Henry’s constant) impacts on the overall mass transfer coefficient. As shown in Figure 4, the overall mass transfer coefficients increases with the increasing of MOR concentration at lean loadings below carbon loading of around 0.2 C/N. At rich loading, however, the difference between the curves of the three studied concentrations is within the experimental uncertainty, which indicates that the impact of MOR concentration is not as significant at rich loading as that at lean loading. This could be due to the chemical reaction resistance is more significant at lean loading while physical diffusion of CO2 dominates at rich loading. At lean loading where the overall mass transfer rate is expected to be primarily determined by the pseudo-first-order reaction, higher MOR concentration has a positive effect on the mass transfer coefficient. Such impact is negated to some degree by higher viscosity and lower CO2 solubility at higher MOR concentration. At rich loading where the physical diffusion resistance could dominate, 1/k°l,prod increases with an increasing MOR concentration due to the dramatic viscosity increase as discussed above. However, it is balanced out by the decreased equilibrium slope in responding to the MOR concentration change as indicated in eq 7. Such counteractive effects on the mass transfer coefficients were also observed in the kinetic study for MEA.16 To gain an insight into the contribution from reaction and diffusion resistances, the physical mass transfer resistance was calculated using the model described in the Theoretical Basis section for 4 M MOR at 30 °C. The contribution from physical mass transfer resistance in the total liquid side mass transfer resistance is then calculated and shown in Figure 6. The percentage of physical diffusion of CO2 in overall liquid phase mass transfer resistance increases exponentially with the increasing carbon loading. The primary cause for the increasing physical diffusion resistance is due to the increasing equilibrium slope (ΔPCO * 2 /Δα [MOR]) determined by the thermodynamic correlation. In addition, the increased viscosity at rich loading as shown in Figure 5 also leads to the increasing in k°l,prod according to eqs 8, 9, and 10. At carbon loading above a C/N of 0.3, the physical diffusion resistance takes more than 50% of the total resistance. The dominance of physical diffusion resistance weakens the

Figure 6. Percentage of physical mass transfer resistance of CO2 in total liquid phase mass transfer resistance for 4 M MOR at 30 °C.

contribution from reactive resistance benefited from high MOR concentration, which makes the impact of MOR concentration much less important at rich loading. Comparison with Benchmark Primary Amine MEA. For reference and comparison purposes, the overall mass transfer coefficient for 30 wt % MEA at 30 °C is also plotted in Figure 4. It can be seen that the overall mass transfer coefficient for MOR is about 50% higher than that for MEA at lean loading. It decreases dramatically with the increasing of carbon loading and drops below the MEA curve at carbon loading greater than 0.4 C/N. Such a comparison is sometimes misleading because each solvent system may have its own optimal working range which is highly related to the concentration driving force ΔP (PCO2 − P*CO2). To better evaluate the performance of such solvent in the CO2 capture process, carbon loadings were converted to the equilibrium CO2 partial pressure (P*CO2) using the empirical equation in eq 11 with the regressed coefficients in Table 1. Therefore, the overall mass transfer coefficients was plotted against PCO * 2 in Figure 7. It can be seen that, within the typical working range

Figure 7. Overall mass transfer coefficients of CO2 in 5 M MOR and MEA with respect of equilibrium partial pressure at 30 °C.

(P*CO2 between 0.5 and 5 kPa),17 the overall mass transfer coefficients of MOR are about three times higher than those for the MEA while the cyclic CO2 capacity (mol-CO2/mol-N) is 26% higher than that of MEA. The different observation in the comparison with MEA in Figure 4 and Figure 7 is because the equilibrium partial pressure of CO2 for MOR is higher than that of MEA which shifts the MOR curves toward lean loading when there is a conversion from carbon loading to equilibrium partial pressure. The high mass transfer coefficient within the 15936

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CONCLUSION Aqueous MOR solution overall mass transfer coefficients for CO2 absorption at 30 °C were collected at carbon loadings of 0.1 to 0.5 mol-C/mol-N. From the comparison of 4 M, 5 M, and 6 M MOR, the MOR concentration was found to have a positive impact on the overall mass transfer coefficients at carbon loadings less than 0.2 mol-C/mol-N. The difference of overall mass transfer coefficients at rich loading is within the experimental uncertainty. Calculation revealed that the difference of MOR concentration impact at lean and rich loadings is due to the dominance of reaction resistance at the lean loading and physical mass transfer resistance at rich loading. At leaner loadings there is predominantly a balance between the higher amine concentration being negated by higher viscosity and lower CO2 solubility. A comparison of MOR and MEA indicates that the overall mass transfer coefficient for MOR is higher than that for MEA at mol-C/mol-N < 0.4 on the basis of carbon loading (mol-C/mol-N), while MOR is three times faster than MEA across the working carbon loading range on the equilibrium partial pressure basis. The rate constant of the CO2 and MOR pseudo-first-order reaction obtained from experimental mass transfer data is three times higher than that of MEA due to the cyclic molecular structure and lower than that of PZ due to the low basicity.

working range can greatly help lower the solvent circulating rate or decrease the size of absorber which takes about 30% of the capital cost for CO2 capture in a reference MEA case.18 Rate Constant. At lean CO2 loading condition where physical mass transfer resistance of reaction products is negligible, the mass transfer of CO2 in the liquid film can be represented by the pseudo-first-order reaction as shown in eq 13.4 1 H = k′g k 2[MOR]DCO (13)

2

in which k2 is the rate constant of pseudo-first-order reaction, [MOR] is the MOR concentration, and H is the Henry’s law constant of CO2 in solution. The diffusivity of CO2 in amine solution, DCO2, can be calculated from a modified Stokes−Einstein equation in eq 14.19 (DCO2)amine soln

⎛ η ⎞0.8 water ⎟⎟ = (DCO2)water ⎜⎜ ⎝ ηamine soln ⎠

(14)

in which ηwater and ηamine soln are the viscosity of water and amine solution, respectively. The rate constant in the pseudo-first-order reaction was calculated using the experimentally determined kg′ at lean CO2 loading through eq 13. Both water and MOR are considered as solvent in this study, therefore the Henry’s law constant of CO2 in aqueous MOR solution is solvent composition dependent. The value of Henry’s law constants were reported previously14 and adopted in this work. The calculated rate constant at 30 °C is listed in Table 2 and compared with several reported rate constants for MOR,

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rate constant (L/mol s)

temperature (K)

MOR MOR MOR MOR

26900 19952 12400 22259

303 298 298 298

MOR MEA PZ

20520 8008 53700

298 303 298

AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. Tel: (859) 257-0293. Fax: (859) 257-0292. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS The authors acknowledge the financial support provided by the Carbon Management Research Group (CMRG) with the members of Duke Energy, East Kentucky Power Cooperative (EKPC), Electric Power Research Institute (EPRI), Kentucky Department of Energy Development and Independence (KYDEDI), Kentucky Power (AEP), and LG&E and KU Energy.

Table 2. Rate Constants of CO2 Reaction with Aqueous Amines amine

Article

resources This work Sharma (1965)20 Littel et al. (1992)4 Al-Juaied and Rochelle (2006)7 Alper (1990)6 Versteeg et al. (1996)21 Bishnoi and Rochelle (2000)22

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REFERENCES

(1) DOE Fossil Energy RD&D: Reducing the Cost of CCUS for Coal Power Plants; DOE/NETL: Pittsburgh, PA, 2012; p 10. (2) Woods, M.; Pinkerton, L. L.; Varghese, E. Updated Costs (June 2011 Basis) for Selected Bituminous Baseline Cases; NETL/DOE-341/ 082312; Department of Energy: Washington DC, 2012; p 61. (3) Freeman, S. A.; Rochelle, G. T. Thermal Degradation of Aqueous Piperazine for CO2 Capture. 1. Effect of Process Conditions and Comparison of Thermal Stability of CO2 Capture Amines. Ind. Eng. Chem. Res. 2012, 51 (22), 7719−7725. (4) Littel, R. J.; Versteeg, G. F.; Van Swaaij, W. P. M. Kinetics of CO2 with primary and secondary amines in aqueous solutionsII. Influence of temperature on zwitterion formation and deprotonation rates. Chem. Eng. Sci. 1992, 47 (8), 2037−2045. (5) Crooks, J. E.; Donnellan, J. P. Kinetics and mechanism of the reaction between carbon dioxide and amines in aqueous solution. J. Chem. Soc., Perkin Trans. 2 1989, 0 (4), 331−333. (6) Alper, E. Kinetics of reactions of carbon dioxide with diglycolamine and morpholine. Chem. Eng. J. 1990, 44 (2), 107−111. (7) Al-Juaied, M.; Rochelle, G. T. Thermodynamics and Equilibrium Solubility of Carbon Dioxide in Diglycolamine/Morpholine/Water. J. Chem. Eng. Data 2006, 51 (2), 708−717. (8) Al-Juaied, M. Carbon dioxide removal from natural gas by membranes in the presence of heavy hydrocarbons and by aqueous

primary amine MEA, and cyclic amine PZ. It can be seen that the rate constant for MOR from this work at 30 °C is higher than the other reported values at 25 °C. Using the energy of activation determined by Al-Juaied,7 the value here can be calculated at the 298 K temperature used by others to be 24700 L/mol·s. In addition, the rate constant of MOR is three times higher than the primary amine MEA, which is consistent with the experimental results from the WWC experiments. The higher rate constant of MOR is due to the low steric hindrance of the six-membered, heterocyclic molecular structure which helps the reaction with the CO2 molecules. On the other hand, the rate constant of MOR is lower than PZ which has a similar molecular structure as MOR. This could be due to the lower basicity of MOR (pKa of 8.5) than that of PZ (pKa of 9.8).14 15937

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dx.doi.org/10.1021/ie402570u | Ind. Eng. Chem. Res. 2013, 52, 15932−15938