Optimization of Blended Amines for CO2 Absorption in a Hollow-Fiber

Article pubs.acs.org/IECR

Optimization of Blended Amines for CO2 Absorption in a HollowFiber Membrane Contactor Zhen Wang,†,§ Mengxiang Fang,*,† Shuiping Yan,‡ Hai Yu,§ Chiao-Chien Wei,§ and Zhongyang Luo† †

State Key Laboratory of Clean Energy Utilization, Zhejiang University, Hangzhou, 310027, China College of Engineering, Huazhong Agricultural University, Wuhan, 430070, China § CSIRO Energy Centre, Mayfield West, 2304, Australia ‡

ABSTRACT: To understand and optimize CO2 absorption in binary amine systems, we experimentally and theoretically investigated CO2 absorption using typical amines and blended amines in a polypropylene hollow-fiber membrane contactor. The amines studied were monoethanolamine (MEA), diethanolamine (DEA), and N-methyldiethanolamine (MDEA), and their aqueous blends of MEA/MDEA, DEA/2-amino-2-methyl-1-propanol (AMP), and MDEA/piperazine (PZ). The predicted results, including overall mass transfer coefficients and CO2 removal ratio, agreed very well with those determined experimentally. For single amines, the optimal concentration was around 30 wt % for MEA and 20 wt % for DEA. MDEA concentration had little effect on the overall mass transfer coefficient. We optimized the formulation of blended amines using theoretical analysis. The optimal compositions in MEA/MDEA, DEA/AMP, and MDEA/PZ systems were respectively 30 wt % MEA, with MDEA in proportions from 0.1 to 0.3; 15 wt % DEA, with AMP in proportions from 0.5 to 0.8; and 20 wt % MDEA, with PZ in a proportion of 0.3. To further understand the CO2 membrane absorption process, we also analyzed individual mass transfer resistances as a function of additive concentration in blended amines and the effects of liquid velocity on the overall mass transfer coefficient. This shows that CO2 absorption is controlled by the liquid side for DEA/AMP blends and by combined liquid−gas phases for MEA/MDEA blends. For MDEA/PZ blends, control of CO2 absorption is characterized by a gradual transition from liquid side controlled to liquid−gas combined controlled as the concentration of PZ increases.

1. INTRODUCTION Carbon dioxide (CO2) is the major contributor to global warming, with 60% of CO2 emissions being produced by the power and industrial sectors.1 Consequently, controlling CO2 emissions from coal-fired power stations is a serious and pressing issue. Among various CO2 capture technologies, amine-based postcombustion capture (PCC) technologies, which typically use packed columns to achieve separation, are believed to be the most mature technologies for removing CO2 from flue gas. This is due to their successful commercial applications in many industries, their effectiveness in gas streams of low CO2 partial pressure, and their tail-end technology, which can be retrofitted to existing power plants.2 However, the liquid and gas streams cannot be controlled independently in amine-based PCC processes. This often results in operating problems such as foaming, flooding, and entrainment, thus limiting amine-based PCC applications.3 To overcome these drawbacks, an alternative technology CO2 membrane absorptionhas been proposed. Figure 1 shows a schematic diagram of CO2 membrane absorption in a hydrophobic membrane. In this process, a suitable membrane configuration, such as a hollow-fiber membrane contactor, is used to control gas and liquid fluids independently. CO2 molecules diffuse through the membrane pores and are absorbed by the liquid absorbent. Membrane absorption therefore presents more operational flexibility, and avoids the operating problems encountered in conventional contacting devices. A membrane contactor also potentially offers a much larger gas−liquid interfacial area than traditional reactors, with a reduction of 63−65% in absorber size;4 it can also be scaled up © 2013 American Chemical Society

Figure 1. Schematic diagram of CO2 membrane absorption in hydrophobic membrane.

linearly. Therefore, membrane contactors are a strong candidate for CO2 absorption compared with a conventional packed tower, and have great potential for PCC. Received: Revised: Accepted: Published: 12170

May 27, 2013 July 19, 2013 July 30, 2013 July 30, 2013 dx.doi.org/10.1021/ie401676t | Ind. Eng. Chem. Res. 2013, 52, 12170−12182

Industrial & Engineering Chemistry Research

Article

overall mass transfer coefficients and individual mass transfer resistances for different blended amines in the CO2 membrane absorption process. We used theoretical analysis to optimize the formulation of blended amine systems, to obtain the highest possible overall mass transfer coefficient between gas and liquid phases while minimizing the use of amines.

Unlike a traditional gas separation membrane, CO 2 selectivity during membrane absorption is provided by solvents, rather than by the membrane itself. The hydrophobic membrane, which prevents flooding problems in the membrane contactor, acts as a barrier between the liquid and gas phases to increase the gas−liquid contact area without dispersing one phase into another. Screening CO2 absorbents for membrane absorption is therefore of vital importance to improve CO2 absorption performance and reduce liquid-side mass transfer resistance. Amine solvents such as monoethanolamine (MEA), diethanolamine (DEA), and N-methyldiethanolamine (MDEA), which are preferred in conventional packed towers, are also extensively reported as absorbents in membrane absorption processes.5−7 Because these absorbents require regeneration after CO2 absorption, regeneration energy consumption is an important consideration. Until now, no single amine has been reported to achieve satisfactory performance in both absorption and desorption.8 Generally speaking, primary (e.g., MEA) or secondary (e.g., DEA) amines present fast reaction kinetics with CO2, but are difficult to regenerate. MDEA, which is a tertiary amine, has a very low CO2 reaction rate, but is easily regenerated. Hence, there has been rising interest in blending different amines to improve performance. Most studies of blended amine absorbents have focused on conventional gas−liquid reactors,9−15 with limited research into the use of amine blends in a hollow-fiber membrane contactor. Gong et al.16 found that CO2 absorption flux increases with MEA content in a blended MEA/MDEA solution. Lu et al.17 activated MDEA with piperazine (PZ) in a polypropylene (PP) hollow-fiber membrane and found that the average value of the overall mass transfer coefficient in MDEA/PZ is 2.25 times higher than in MEA. Paul et al.18 simulated the performances of different single and blended alkanolamine solvents in a hollowfiber membrane contactor, at a total amine concentration of 10 wt % in all cases. However, the combined experimental and theoretical analyses of CO2 absorption in different blends of amine solvents in the hollow-fiber membrane contactor were very limited. CO2 absorption can be improved by increasing the amine concentration, but it is impossible to improve the CO2 absorption performance to a very high level with increasing base or additive amine concentrations. This is because the viscosity changes with amine concentration and further reduces the mass transfer between the gas and liquid phases. Moreover, a higher concentration of absorbent can accelerate corrosion of facilities and amine volatile loss, and increase the operation and maintenance costs of PCC. Therefore, optimizing the formulation of blended amines is of practical significance for efficient, effective CO2 membrane absorption. To the best of our knowledge, the optimization of binary amine systems for CO2 removal in hollow-fiber membrane contactors has not yet been reported. In the present study, we investigated the performances of CO2 absorption from CO2/N2 gas mixtures in a hollow-fiber membrane contactor using three different single amines (MEA, DEA, and MDEA) and three blended amines (MEA/MDEA, DEA/2-amino-2-methyl-1-propanol (AMP), and MDEA/PZ). Polytetrafluoroethylene is recommended as the best material for membrane absorption due to its high hydrophobicity;19 however, in favor of commercial applications, we used a commercially available, low-cost PP membrane. The effects of amine concentration and liquid velocity were investigated and compared with our theoretical analysis. We also analyzed the

2. THEORY 2.1. CO2−Amine Reaction Mechanism. The reaction of CO2 with primary or secondary amines can be described by a two-step zwitterionic mechanism, which was first proposed by Caplow20 and reintroduced by Danckwerts.21 The first step of the reaction is to form a zwitterion as an intermediate, which can be expressed as k 2,amine

CO2 + R1R 2NH XooooooY R1R 2NH+COO− k −1

(1)

The zwitterion then undergoes deprotonation by a base B, as shown by kb

R1R 2NH+COO− + B → R1R 2NCOO− + BH+

(2)

−

where B could be an amine, OH , or H2O. The contribution of OH− can usually be ignored, due to its very low concentration in solution. With the assumption of quasi-steady-state condition for the zwitterion concentration, the change of zwitterion concentration is negligible. Then the CO2 reaction rate at quasi-steady state can be expressed by rCO2 − R1R 2NH =

k 2,R1R 2NH[CO2 ][R1R 2NH] 1+

k −1 k H2O[H 2O] + k R1R2NH[R1R 2NH]

(3)

For the reaction of CO2 with tertiary alkanolamines, such as MDEA, the reaction with CO2 is a base-catalyzed hydration of CO2,22 and can be described by k 2,amine

CO2 + R1R 2R3NH + H 2O ←⎯⎯⎯→ R1R 2R3NH+ + HCO3− (4)

Therefore, the reaction of CO2 with tertiary alkanolamines can be assumed to be of pseudofirst order with respect to both CO2 and amine. The reaction rate of CO2 can be expressed as rCO2 − R1R 2R3N = k 2,R1R 2R3N[CO2 ][R1R 2R3N]

(5)

For the reaction of CO2 with PZ, the reaction can be treated as a second-order reaction, as shown by23 rCO2 − PZ = k 2,PZ[CO2 ][PZ]

(6)

For absorption of CO2 into blended amine aqueous solutions, the overall CO2 reaction rate can be expressed as follows: robs = rCO2 − amine1 + rCO2 − amine2 + rCO2 − OH− + rCO2 − H2O (7)

In eq 7, the reaction of CO2 with H2O and OH− can be ignored, due to their small contributions.24 Therefore, the overall CO2 reaction rate in blended amines can be simplified as robs = rCO2 − amine1 + rCO2 − amine2 (8) In this work, for the reaction of CO2 with blended amine systems, the overall CO2 reaction rate can be expressed as the following equations, respectively. 12171

dx.doi.org/10.1021/ie401676t | Ind. Eng. Chem. Res. 2013, 52, 12170−12182


Article

k 2,AM2[CO2 ][AM2] k 2,AM1[CO2 ][AM1]

+ k 2,MDEACCO2CMDEA (9)

for blended DEA/AMP: k 2,AMPCCO2CAMP k −1 ∑ k bCb

(10)

for blended MDEA/PZ: robs = k 2,MDEACCO2CMDEA + k 2,PZCCO2C PZ

(11)

robs = kobs[CO2 ]

k 2,AM2[CO2 ][AM2]

Since the change of amine concentration is very small in the liquid phase, the overall reaction rate of CO2 with aqueous amines can be expressed as a pseudo-first-order reaction with respect to CO2 as follows: (12)

Gz =

uLd i 2 DCO2,L L

(15)

where uL is the liquid mean velocity and L is the effective length of the fiber membrane. 2.2.2. Membrane-Side Mass Transfer Coefficient. On the assumption that CO2 transport through the membrane is only caused by diffusion and the membrane is operated without a 12172

k 2,AM1[CO2 ][AM1] k 2,AM1[CO2 ][AM1] MDEA (AM1)

where Sh is the Sherwood number, DCO2,L is the diffusion coefficient of CO2 in solvent, and Gz is the Graetz number, expressed by

MEA (AM1), DEA (AM1)

(14)

single amines

kLd i = (3.673 + 1.623Gz)1/3 DCO2,L

Table 1. Expressions of CO2 Reaction Rates for Different Amine Systems

Sh =

rCO2AM1

where KO is the overall mass transfer coefficient; kL, km, and kg are the individual mass transfer coefficients in the liquid, membrane, and gas phases, respectively; E is the enhancement factor; m is the distribution coefficient between liquid and gas; di, dm, and do are the inner, log mean, and outer diameters of the hollow-fiber membrane; and dm = (do − di)/ln(do/di). In eq 13, the enhancement factor E is a function of pseudofirst-order reaction rate constant kobs; the detailed calculation of E is presented in section A.7. 2.2.1. Tube-Side Mass Transfer Coefficient. Generally, fluid flow in the tube side of a hollow-fiber membrane is considered to be a fully developed laminar flow. The Graetz−Leveque solution can be used to predict the mass transfer coefficient in the tube side. Kerulen et al.32 proposed a generalized solution of the Graetz−Leveque equation that is applicable for the complete range of Graetz numbers in the laminar flow:

k 2,AM1[CO2 ][AM1]

(13)

amines

1 1 1 di 1 di = + + KO mEkL k m dm kg do

1 + 1/((k H2O/k −1)[H 2O] + (kAM1/k −1)[AM1])

Table 1 presents the specific rate expressions for different amine systems. The values of kinetic constants for all the reactions of CO2 with amine systems at the temperature of 313 K are shown in Table 2. 2.2. Overall Mass Transfer Coefficient. For a hydrophobic hollow-fiber membrane with gas-filled pores and liquid absorbent in the lumen side, the overall gas phase mass transfer coefficient can be expressed by a resistance in series model:31

DEA (AM1)/AMP (AM2)

1+

1 + 1/((k H2O/k −1)[H 2O] + (kAM1/k −1)[AM1] + (kAM2/k −1)[AM2])

+

MEA (AM1)/MDEA (AM2)

1+

k −1 ∑ k bCb

blended amines

k 2,DEACCO2C DEA

rCO2AM2

robs =

MDEA (AM1)/PZ (AM2)

k 2,AM2[CO2 ][AM2]

k −1 ∑ k bCb

1+


k 2,MEACCO2CMEA

k 2,AM1[CO2 ][AM1]

robs =


for blended MEA/MDEA:



Article

Table 2. Kinetic Parameters of Single and Blended Amine Systems (313 K) aqueous amine systems MEA (AM1) + H2O MEA (AM1)/ MDEA (AM2) + H2O DEA (AM1) + H2O AMP (AM1) + H2O DEA (AM1)/AMP (AM2) + H2O MDEA (AM1) + H2O MDEA (AM1)/PZ (AM2) + H2O

k2,AM1 (m3 mol−1 s−1) 1.74 × 10

1

1.74 × 10

1

k2,AM1kH2O/k−1 (m6 mol−2 s−1) 1.675 × 10

−5

1.675 × 10

−5

k2,AM1kAM1/k−1 (m6 mol−2 s−1) 5.33 × 10

−3

5.33 × 10

−3

1.51 1.25

2.45 × 10−6 5.82 × 10−3

2.46 × 10−3 3.16

1.60

2.51 × 10−6

2.62 × 10−3

k2,AM1kAM2/k−1 (m6 mol−2 s−1)

k2,AM2 (m3 mol−1 s−1)

k2,AM2kH2O/k−1 (m6 mol−2 s−1)



ref 25

7.80 × 10

−4

1.29 × 10

−2

25, 26 27 28

7.17 × 10−3

1.20

3.29 × 10−5

1.00 × 10−3

2.80 × 10−2

1.29 × 10−2

27 26

1.29 × 10−2

1.51 × 102

29, 30

Table 3. Physicochemical Parameters of Single Amines Used for Prediction of Mass Transfer Coefficient at 313 K (Cin = 14 vol %, ω0 = 10 wt %) MEA

DEA

MDEA

ρL DCO2,L

param

kg m−3 m2 s−1

units

1000.03 2.38 × 10−9

1007.81 1.08 × 10−9

1002.35 9.88 × 10−10

DCO2,G

m2 s−1

1.87 × 10−5

1.87 × 10−5

1.87 × 10−5

−1

−9

Damine Dm ig Dkig HCO2,L

m s m2 s−1 m2 s−1 kPa m3 kmol−1

1.05 × 10 8.72 × 10−6 1.29 × 10−6 4065.84

3.71 × 10 8.72 × 10−6 1.29 × 10−6 4471.47

3.08 × 10−10 8.72 × 10−6 1.29 × 10−6 4156.79

ηL μG

Pa s Pa s

8.25 × 10−4 1.79 × 10−5

3.06 × 10−3 1.79 × 10−5

3.56 × 10−3 1.79 × 10−5

2

Sh = 5.85(1 − φ)(dh /L)Re 0.6Sc 0.33

pore-wetting problem, the mass transfer coefficient in the membrane can be expressed as33 1 τδ = e km Dig ς

(18)

where φ is the packing density, dh is the hydraulic diameter, and Re and Sc are the Reynolds and Schmidt numbers, respectively. Table 3 lists some of the main parameters of the single amines at a concentration of 10 wt % concentration that were used to calculate theoretical mass transfer coefficients. The detailed calculations of these physicochemical parameters, such as the Henry’s constant, enhancement factor, viscosity, and diffusion coefficient, can be found in the Appendix.

(16)

where ς is the membrane porosity, Dige is the effective membrane diffusion coefficient in pure gas-filled pores, δ is the membrane thickness, and τ is tortuosity; τ = 1/ς2. The calculation of Deig was determined by the mean pore diameter of the membrane pore. In this work, the pore diameter is in the range 1 × 10−7−1 × 10−5 m. Therefore, diffusion through the membrane is the combination of molecular diffusion and Knudsen diffusion, and Deig can be determined by

1 1 1 e = m + Dig Dig Digk

−10

3. EXPERIMENTAL SECTION 3.1. Materials. MEA (purity, >99%), DEA (purity, >98%), MDEA (purity, >98%), and PZ (purity, 99.5%) were purchased from Sinopharm Co. AMP (purity, >97%) was purchased from Fluka. All amines were used as received without further purification. The solutions were prepared to the desired concentrations by dissolving the amines in deionized water. High-purity (>99.9 mol %) CO2 and N2 gases were purchased from Hangzhou Jingong Gas Co. Ltd. The hydrophobic microporous PP hollow-fiber membrane modules were provided by Hangzhou Joint-future Membrane Co. The specifications of the membrane modules are listed in Table 4. 3.2. Experimental Apparatus and Procedure. A schematic diagram of the experimental apparatus for CO2 membrane absorption is shown in Figure 2. Aqueous amine absorbent in the lean solvent reservoir is first heated to the desired temperature and then pumped into the tube side of the hollow-fiber membrane contactor. The liquid flow rate is regulated by the flow controller, and the liquid-side pressure is adjusted at around 135 kPa by a back-pressure valve placed on

(17)

k in which Dm ig and Dig are the molecular and Knudsen diffusion coefficients of CO2, respectively. 2.2.3. Shell-Side Mass Transfer Coefficient. Many studies have been reported concerning shell-side mass transfer,34−36 but shell-side fluid flow is not yet fully understood. The fluid hydrodynamics of the shell side varies significantly and is affected by a range of parameters, such as the packing density, the irregularity of fiber spacing, the polydispersity of the fiber diameter, and the inlet and outlet effects. In this work, taking into consideration the operating conditions of shell-side flow, the following equation is used to predict the mass transfer coefficient at the given conditions of 0 < Re < 500, 0.04 < φ < 0.4:37

12173



Article

logarithmic mean driving force based on gas-phase concentration, and can be described as

Table 4. Specifications of Hollow Fiber and Membrane Module item

parameter

units

value

PP membrane fiber

fiber i.d. fiber o.d. pore size porosity effective length module i.d. no. of fibers contact areaa

mm mm μm % mm mm

0.344 0.424 0.2 × 0.02 40 200 20 500 0.108

membrane module

a

m2

ΔC =

* ) − (Cg,out − Cg,out * ) (Cg,in − Cg,in * )/(Cg,out − Cg,out * )] ln[(Cg,in − Cg,in

(20)

where Cg,in * and Cg,out * are the inlet and outlet gas phase concentrations in equilibrium with the corresponding liquid phase CO2 concentrations, respectively. In this study, C*g,in and Cg,out * can be assumed to be zero, due to very low CO2 loadings in the solution.

4. RESULTS AND DISCUSSION 4.1. CO2 Membrane Absorption in Single Amines. Figure 3 describes the experimental and theoretical CO2 overall

Contact area is calculated based on the inner diameter of the fiber.

the liquid outlet for preventing gas bubble formation in the liquid. Two gas stream inputs, CO2 and N2, are mixed by mass flow controllers at around ambient pressure (107 kPa) to produce a simulated flue gas with 14 vol % CO2. The gas mixture flows into the shell side of the hollow-fiber membrane contactor after being heated to the desired temperature. During the course of the experiment, the entire membrane contactor system is kept at a constant temperature of 40 °C. The CO2 inlet and outlet concentrations in the gas phase were measured using a gas analyzer (Rosemount NGA-2000), and CO2 loadings in solvents after absorption were measured using a titration method which was described in our previous work.38 3.3. Experimental Overall Mass Transfer Coefficient. The experimental overall mass transfer coefficient was used to assess the membrane’s CO2 absorption performance and compared with the theoretical mass transfer coefficient. It can be calculated by the following equation: K OEXP =

Vg,inCg,in − Vg,outCg,out

Figure 3. Influences of amine concentration on overall mass transfer coefficient for single amines (VG = 120 L/h, VL = 8 L/h, T = 40 °C, CCO2,in = 14 vol %, PL = 135 kPa, PG = 107 kPa, ΔPL = 6 kPa).

(19) AΔC where Vg,in and Vg,out are the inlet and outlet gas flow rates, respectively; Cg,in and Cg,out represent the gas phase CO2 concentrations in mol/m3 at the inlet and outlet, respectively; A denotes the contact area of membrane in m2. ΔC is the

mass transfer coefficients, KO, of single amines when the amine concentrations were varied from 5 to 40 wt %. A good

Figure 2. Schematic diagram of the experimental setup for CO2 absorption membrane absorption. 12174



Article

reaction zone may extend up to the axis of the fiber, and that an increase in MEA concentration has a positive influence on KO. For DEA, Gz increases almost 10-fold, from 25 to 286, as the DEA concentration increases from 10 to 90 wt %. This demonstrates that CO2 has difficulty diffusing into the liquid phase, in particular, in the high DEA concentration region. MDEA solution has an intermediate Gz number, which increases slightly as the concentration rises. Varying the MDEA concentration therefore has a relatively insignificant effect on KO. 4.2. Effects of Additive Concentration for Blended Amines. Figure 5 presents the simulated and experimental KO

agreement between the experimental and theoretical KO values was achieved within the range of amine concentrations investigated, which confirms the validity of the theoretical model. KO values in MEA solvent are much higher than those in DEA and MDEA. This can easily be explained by the difference in reaction rate constants for the reaction of CO2 with different amines. As listed in Table 2, MEA (as a primary amine) possesses the highest second-order forward reaction constant, k2, to form the zwitterions. This is 1 order of magnitude higher than that for DEA and 3 orders of magnitude higher than that for MDEA, respectively. Figure 3 also shows that, as the amine concentration increases, KO varies following different patterns in the three different amines. The KO in MEA increases greatly in the lowMEA-content region and then slowly increases at high MEA concentrations. Therefore, there is an optimal value for MEA concentration, beyond which KO does not increase noticeably with an increase in concentration. This transitional MEA concentration is around 30 wt %, as shown in Figure 3. Aqueous DEA also presents a critical concentration of 20 wt %, beyond which KO reduces with higher DEA concentrations. Similar results can also be found in Jamal’s work.39 In the case of MDEA, there is no obvious improvement of KO with rising MDEA concentration, due to its low reactivity and high viscosity. These phenomena were further analyzed by investigating the effects of amine concentration on the Graetz number (Gz); the results are presented in Figure 4. The dimensionless Gz

Figure 5. Effects of additive amine concentration on overall mass transfer coefficient (VG = 120 L/h, VL = 8 L/h, T = 40 °C, CCO2,in = 14 vol %, ω0 = 10 wt %, PL = 135 kPa, PG = 107 kPa, ΔPL = 6 kPa).

values for three different blended amine systems (MEA/ MDEA, DEA/AMP, and MDEA/PZ) as a function of additive (MDEA, AMP, and PZ) content. In this study, the base amine concentrations, ω0, remain constant at 10 wt % for all blended amines. As shown in Figure 5, in the MEA/MDEA blended amine, addition of MDEA weakly influences KO. A similar weak influence of additives was found in blended DEA/AMP solution, in which an increase in AMP content only slightly increases KO. Although MDEA and AMP are not good promoters of CO2 absorption, they can increase the CO2 capture capacity and are beneficial to CO2 desorption, due to their low CO2 reaction heats. The addition of PZ strongly increases KO in blended MDEA/PZ solution, which indicates that PZ is an effective absorbent to activate MDEA in CO2 membrane absorption. To further optimize blended amines, the effects of additive amine concentration on CO2 outlet concentration in blended amines were also investigated and are shown in Figure 6. For MDEA/PZ blends, the CO2 removal ratio improves significantly with an increase of ωPZ when ωPZ/ωMDEA is smaller than 0.4. At ωPZ/ωMDEA above 0.4, the room for an increase of the CO2 removal ratio is small, since the CO2 removal ratio is close to 0.9. In the blended MEA/MDEA solution, an increase of ωMDEA/ωMEA has little influence on the improvement of CO2 removal efficiency, which remains above 0.9 at all additive concentrations. In the DEA/AMP system, AMP promoted CO2 absorption, but the maximum CO2 removal ratio theoretically achieved in this study is only 75% as ωAMP/ωDEA increases to 1.

Figure 4. Change of Gz number with amine concentration (VG = 120 L/h, VL = 8 L/h, T = 40 °C, CCO2,in = 14 vol %, PL = 135 kPa, PG = 107 kPa, ΔPL = 6 kPa).

indicates the ratio of the penetration time of solute CO2 to reach the axis of the hollow fiber from the gas−liquid interface, to the average residence time of the liquid in the fiber. A higher Gz number means that the penetration depth of the CO2 diffusing from the gas−liquid interface is smaller than the fiber radius at a fixed gas−liquid contact time (i.e., constant liquid flow rate). Consequently, the free amine far from the interface is essentially undisturbed. Based on eq 15, Gz is inversely proportional to the CO2 diffusion coefficient DCO2,L, which is a strong function of the viscosity of amine solutions. In Figure 4, aqueous MEA has the lowest Gz number, which increases slowly as the concentration rises. This indicates that the CO2 12175



Article

In terms of KO, amine systems differ in their sensitivity to liquid velocity. Generally, KO is more sensitive to the liquid flow rate in amines with a higher CO2 absorption rate than in amines with lower CO2 absorption rate. For example, with increasing liquid velocity, the extent of increase of KO in MEA and MEA/MDEA is larger than that in DEA, DEA/AMP, and MDEA/PZ, and there is almost no increase of KO in MDEA solution. This is due to the difference in individual mass transfer coefficients in the liquid phase for amines. The liquid mass transfer coefficient is quite small in MDEA solution, and consequently, the size of the overall mass transfer coefficient is relatively limited by increasing liquid velocity. 4.4. Optimization of Blended Amine System. To achieve a greater CO2 absorption rate with the minimum use of amines, the profiles of the overall mass transfer coefficient as a function of base amine concentration and additive proportion, ωadditive/ω0, were plotted for three different blended amine systems. Figure 8 depicts the profile of KO with changing ωMEA from 5 to 50 wt % and ωMDEA/ωMEA from 0 to 1. In MEA/ MDEA blends, a small ωMDEA/ωMEA ratio improves the KO more effectively than a large ωMDEA/ωMEA ratio does. This is especially the case at ωMEA greater than 35 wt % and ωMDEA/ ωMEA greater than 0.15, when KO begins to decrease with rising MDEA content. This phenomenon can be attributed to the combined influences of MDEA’s high viscosity and low CO2 reaction activity. Although a higher ωMEA could increase the overall mass transfer coefficient, this effect becomes weaker at high ωMEA regions, due to a higher Gz number. A higher ωMEA could also raise the solution’s viscosity. Therefore, an optimal ωMEA is recommended to be approximately 30 wt %, with ωMDEA/ωMEA being 0.1−0.3 for blended MEA/MDEA. Figure 9 shows the profiles of KO in DEA/AMP blends as a function of DEA and AMP concentrations. Like MDEA in MEA/MDEA blends, adding more AMP will not always improve KO. When the value of ωDEA exceeds 35 wt %, KO declines with the increase of ωAMP/ωDEA, but when ωDEA is below 20 wt %, KO has an approximately linear dependence on ωAMP/ωDEA. Hence, a smaller ωAMP/ωDEA is recommended in

Figure 6. Effects of additive amine concentration on CO2 outlet concentration ratio in blended amines (VG = 120 L/h, VL = 8 L/h, T = 40 °C, CCO2,in = 14 vol %, ω0 = 10 wt %, PL = 135 kPa, PG = 107 kPa, ΔPL = 6 kPa).

DEA/AMP may therefore not be a suitable blended amine for CO2 absorption due to its low KO and CO2 removal rate. The same conclusion was drawn by Paul et al.18 However, DEA/ AMP is still competitive, considering its relatively low regeneration energy demand. 4.3. Effects of Liquid Flow Velocity. Figure 7 shows the variation of KO with liquid velocity in single and blended amines. The ratio of additive amine to base amine, ωadditive/ω0, was fixed at 0.2 with ω0 being 5 wt % for all blended amine systems. KO rises with an increase in liquid velocity, which is expected, due to decreasing liquid-phase resistance. In practice, howeverconsidering the high power consumption and probability of membrane wetting as a result of high liquid pressureit may not be attractive to elevate the liquid velocity to a very high level (e.g., turbulent flow) to improve CO2 absorption in a membrane contactor with narrow fibers.

Figure 7. Effects of liquid velocity for single and blended amines (VG = 120 L/h, T = 40 °C, CCO2,in = 14 vol %, ω0= 5 wt %, ωadditive/ω0= 0.2, PL = 135 kPa, PG = 107 kPa, ΔPL = 6−10 kPa). 12176



Article

Figure 8. Profiles of overall mass transfer coefficient in MEA/MDEA blended amine as a function of MEA and MDEA concentrations (VG = 120 L/ h, VL = 8 L/h, T = 40 °C, CCO2,in = 14 vol %, PL = 135 kPa, PG = 107 kPa, ΔPL = 6 kPa).

Figure 9. Overall mass transfer coefficient profile of DEA/AMP blended amine as a function of AMP and DEA concentrations (VG = 120 L/h, VL = 8 L/h, T = 40 °C, CCO2,in = 14 vol %, PL = 135 kPa, PG = 107 kPa, ΔPL = 6 kPa).

the high ωDEA region, whereas a larger ωAMP/ωDEA is better for low ωDEA conditions. The optimal ωDEA is around 15 wt %, where the highest KO can be achieved. For DEA/AMP blends, the specific DEA concentration ωDEA is therefore recommended to be approximately 15 wt % with a ωAMP/ωDEA ratio of 0.5−0.8. The influence of ωMDEA and ωPZ/ωMDEA on KO for blended MDEA/PZ is shown in Figure 10. As ωMDEA increases beyond 10 wt %, the whole range of ωPZ/ωMDEA from 0 to 1 cannot be fully covered, due to the limited solubility of PZ in water at 40 °C. PZ clearly improves the CO2 absorption rate, although this influence becomes relatively weaker as ωPZ/ωMDEA increases. In addition, as a result of MDEA’s high viscosity, the extent of increase of KO reduces as the value of ωMDEA becomes larger. Thus, for an aqueous MDEA/PZ system, the most effective concentration of MDEA is nearly 20 wt %, and ωPZ/ωMDEA is optimal around 0.3.

4.5. Mass Transfer Resistance Analysis and Comparison. The analysis of individual mass transfer resistances for MEA/MDEA, DEA/AMP, and MDEA/PZ amine systems with the variation of ωadditive/ω0 is presented in Figure 11. The base amine concentration ω0 was kept constant at 10 wt %. For MEA/MDEA blends, the gas-side mass transfer resistance is comparable to the liquid-side resistance, but the membrane-side resistance is quite small. Therefore, the CO2 absorption process in MEA/MDEA blends is controlled by both gas-phase and liquid-phase mass transfer. In DEA/AMP blends, although increasing the AMP concentration can weaken the contribution of the liquid-side resistance, the liquid-phase resistance is still dominant; hence, CO2 membrane absorption is controlled by liquid-phase mass transfer. In MDEA/PZ blends, the liquid resistance declines sharply with an increase in PZ content, as a result of PZ’s high CO2 reaction activity; consequently, the CO2 absorption process gradually changes from liquid-phase 12177



Article

Figure 10. Overall mass transfer coefficient profile of MDEA/PZ blended amine as a function of MDEA and PZ concentrations (VG = 120 L/h, VL = 8L/h, T = 40 °C, CCO2,in = 14 vol %, PL = 135 kPa, PG = 107 kPa, ΔPL = 6 kPa).

highest KO and minimize amine use, the optimal concentration for DEA is 15 wt %, and the optimal concentration for MEA is around 30 wt %. The concentration of MDEA has little impact on KO. For all amines tested, increasing the liquid velocity leads to a slight improvement in KO due to the reduction of liquidside mass transfer resistance. For blended amines, solvent formulation optimization was studied intensively by theoretical analysis. The recommended optimal blended amine compositions are the following: for MEA/MDEA, 30 wt % MEA with a MDEA proportion of 0.1− 0.3; for DEA/AMP, 15 wt % DEA with an AMP proportion of 0.5−0.8; and for MDEA/PZ, 20 wt % MDEA with a PZ proportion of 0.3. The analysis of individual mass transfer resistances shows that CO2 absorption is controlled by the liquid side for DEA/AMP blends and by combined liquid−gas phases for MEA/MDEA blends. For MDEA/PZ blends, control of CO2 absorption is characterized by a gradual transition from liquid-side controlled to liquid−gas combined controlled as the concentration of PZ increases.

Figure 11. Individual mass transfer resistance comparison of different blended amines (VG = 120 L/h, VL = 8 L/h, T = 40 °C, CCO2,in = 14 vol %, ω0 = 10 wt %, PL = 135 kPa, PG = 107 kPa, ΔPL = 6 kPa).

■

controlled to combined liquid−gas controlled. In addition, the membrane-phase resistance in this study is relatively small compared to liquid-side resistance, and can usually be ignored if a nonwetted membrane is assumed. However, if the absorbents wet the membrane, CO2 absorption can deteriorate, due to dramatically increased membrane resistance.

APPENDIX

A.1. Viscosities of Blended Amine and Mixture Gases

The viscosity of aqueous blended amines was estimated by the model of Grunberg and Nissan:40 ln ηm =

∑ yi ln ηi + ∑ ∑ yyi j Gij

where ηi is the viscosity of the ith pure fluid fraction of the ith component in the temperature-dependent parameters, which Mandal et al.41 For a mixed gas (binary system), the estimated by Wilke’s method:

5. CONCLUSION The CO2 absorption performance of three single amines (MEA, DEA, and MDEA) and three typical blended amines (MEA/ MDEA, DEA/AMP, and MDEA/PZ) was investigated experimentally and theoretically in a PP hollow-fiber membrane contactor. The predicted mass transfer coefficients and CO2 removal ratio results agree very well with those obtained experimentally. Amine concentration has diverse influences on the overall mass transfer coefficient, KO, of single amines. To achieve the

μG =

y1μ1 y1 + y2 M 2 /M1

+

(A.1)

and yi is the mole mixture; Gij are are provided by viscosity can be

y2 μ2 y2 + y1 M1/M 2

(A.2)

in which Mi is the mole weight of the ith component. 12178



Article

A.2. CO2 Diffusivity in Liquid Phase

ΩD =

The CO2 diffusivity in blended amine solutions can be determined by the analogy of N2O diffusivity in solution.42 It can be described as follows: ⎛ DCO ,W ⎞ 2 ⎟⎟ DCO2,L = DN2O,L ⎜⎜ ⎝ DN2O,W ⎠

Ωμ =

(A.3)

The CO2 diffusivity DCO2,W and N2O diffusivity DN2O,W in water can be expressed in the following equations, 42 respectively: DCO2,W = (2.35 × 10−6)e(−2119/ T )

(A.4)

DN2O,W = (5.07 × 10−6)e(−2371/ T )

(A.5)

T̅ is dimensionless temperature: T̅ = κT/ε, where κ is the Boltzmann constant and ε is the parameter of the Stockmayer potential, which can be estimated by the critical temperature. The Knudsen diffusion coefficient can be calculated by the following equation:46 Digk = 48.5d p

where ηL and ηW are the viscosities of the amine solution and water, respectively.

Henry’s constant, H, of CO2 in MEA solution also can be calculated by the N2O analogy:

The diffusion coefficients of mixed amines in water can be determined by the following correlation:43

HCO2,L = HN2O,L(HCO2,W /HN2O,W )

(A.7)

DB,F =

P(νF1/3 + νB1/3)2

+

(A.9)

(A.16)

(A.17)

n

X = ln HN2O,L −

∑ Φi ln HN O,i 2

i=1

(A.18)

where X is the excess Henry’s coefficient and Φi is the volume fraction of the ith solvent. In addition, the relationship of dimensionless Henry’s constant, H, and the distribution coefficient, m, can be determined by

A.5. Molecular and Knudsen Diffusion Coefficients of CO2

As introduced previously, the effective diffusion coefficient of CO2 in the membrane’s pores is the combination of molecular and Knudsen diffusion coefficients. The molecular self-diffusion coefficient of CO2 is calculated from the kinetic gas theory:45 RT Ωμ μ MP ΩD

HN2O,W = (8.55 × 106)e(−2284/ TL)

in which b1 and b2 are the parameters for calculating the solubility of N2O in pure amine. The solubility of N2O in aqueous mixed amines can be described by a semi-empirical model, by correlating the excess Henry’s coefficient presented as the following:48

1 MF

where TG is the gas phase temperature, MB and MF are the molecular weights of gases B and F (g/mol), P is the pressure in the gas phase (kPa), and νB and νF are the molecular volumes of gases B and F (cm3/mol), respectively.

Digm = 1200

(A.15)

HN2O − pureamine = b1 exp(b2 /T )

When gas B is diffused in gas F, the diffusivity of gas B can be calculated by the Maxwell−Gilliland equation:44 1 MB

HCO2,W = (2.82 × 106)e(−2044/ TL)

Wang et al.47 proposed the solubility of N2O in pure amine solvent as follows:

(A.8)

A.4. CO2 Diffusivity in Gas Phase

(4.36 × 10−5)TG 3/2

(A.14)

in which HN2O,L is the Henry’s constant of N2O in aqueous amine. Henry’s constants of CO2 and N2O in water, HCO2,W and HN2O,W, can be determined in the following equations provided by Versteeg et al.:42

in which M is the molar mass of the amine and ρ is the density of the amine. The diffusion coefficients were correlated for viscosity and temperature using the modified Stokes−Einstein relation:41 Damine,L

(A.13)

A.6. Henry’s Constant

A.3. Amine Diffusivity in Liquid Phase

0.6 T ⎛ ηW ⎞ ⎜⎜ ⎟⎟ = Damine,W 273 ⎝ ηL ⎠

TG M

where TG is the gas temperature, M is the molecular weight of the gas, and dp is the pore mean diameter.

(A.6)

⎛ M ⎞−0.54 Damine,W = 2.5 × 10−10⎜ ⎟ ⎝ρ⎠

1.16145 0.52487 2.16178 + + exp(0.7732T̅ ) exp(2.43787T̅ ) (T̅ )0.14875 (A.12)

The diffusion coefficients of N2O in the amine solution are calculated according to the modified Stokes−Einstein relation: DN2O,L ηL 0.6 = DN2O,W ηW 0.6

1.6036 0.193 1.03587 + + 0.1561 exp(0.47635T̅ ) exp(1.52996T̅ ) (T ̅ ) 1.76474 + exp(3.89411T̅ ) (A.11)

m= (A.10)

RT H

(A.19)

in which R is the ideal gas constant, 8.314 J/(mol K).

where Dm ig is the molecular self-diffusion coefficient, M is the gas molecular weight, and μ is the gas dynamic viscosity in Pa s. P is the gas pressure in kPa. Ωμ and ΩD are the viscosity collision integral and diffusion collision integral, respectively:

A.7. Enhancement Factor

The enhancement factor, E, in a fast reaction regime is given by DeCoursey:49 12179


Industrial & Engineering Chemistry Research E=

−Ha 2 + 2(E∞ − 1)

Article

kb = second-order forward reaction rate constant for base b, b = H2O or R1R2NH [m3 mol−1 s−1] kobs = pseudo-first-order reaction rate constant [s−1] KO = overall mass transfer coefficient [m s−1] kL = mass transfer coefficient in liquid phase [m s−1] km = mass transfer coefficient in membrane phase [m s−1] kg = mass transfer coefficient in gas phase [m s−1] L = fiber length of membrane module [m] m = distribution coefficient between liquid and gas M = mole weight [g mol−1] P = pressure [kPa] ΔPL = pressure drop in liquid side along the membrane contactor [kPa] r = reaction rate [mol s−1] robs = overall reaction rate [mol s−1] R = ideal gas constant [J mol−1 K−1] X = excess Henry’s coefficient T = temperature [K] T̅ = dimensionless temperature uL = liquid mean velocity [m s−1] yi = mole fraction of the ith component Gz = Graetz number Re = Reynolds number Sh = Sherwood number Sc = Schmidt number

E∞Ha 2 Ha 2 + +1 E∞ − 1 4(E∞ − 1)2 (A.20)

where Ha is the Hatta number and E∞ is the asymptotic infinite enhancement factor. For a system with a fast first-order irreversible chemical reaction, the Hatta number can be described as Ha =

kobsDCO2,L /kL

(A.21)

The asymptotic infinite enhancement factor, based on a penetration model, is described as 0.5 ⎛ CamineDamine ⎞⎛ DCO2,L ⎞ ⎟⎟⎜ E∞ = ⎜⎜1 + ⎟ βCCO2,LDCO2,L ⎠⎝ Damine ⎠ ⎝

(A.22)

where β is the stoichiometric coefficient of amine in the reaction.

■

AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. Tel.: 86-571-87952803. Fax: 86571-87951616. Notes

The authors declare no competing financial interest.

■

Greek Symbols

β = stoichiometric coefficient of amine in the reaction δ = membrane thickness [m] τ = tortuosity ς = membrane porosity φ = packing density η = viscosity of liquid [mPa s] μ = viscosity of gas [Pa s] ν = molecular volume of gas [cm3 mol−1] ρ = density [kg m−3] Ωμ = viscosity collision integral ΩD = diffusion collision integral Φi = volume fraction of ith solvent κ = Boltzmann constant [J K−1] ε = Stockmayer potential parameter [J] ω = mass concentration [wt %] ω0 = mass concentration of base amine

ACKNOWLEDGMENTS This work was financially supported by the National Natural Science Foundation of China (Nos. 51076139, 51276161) and Zhejiang Province Key Science Innovation Team Project (No. 2009R50048). Z.W. is grateful for financial support from the Australian government under the Australia−China Joint Coordination Group on Clean Coal Technology Partnership Fund, which allows him to carry out part of his work at CSIRO.

■

NOMENCLATURE A = liquid−gas contact area [m2] b = parameter for calculating N2O solubility in pure amine C = CO2 concentration [mol m−3] di = inner diameter of hollow-fiber membrane [m] dm = log mean diameter of hollow-fiber membrane [m] do = outer diameter of hollow-fiber membrane [m] dh = hydraulic diameter [m] dp = pore mean diameter [m] D = diffusion coefficient [m2 s−1] DCO2 = diffusion coefficient of CO2 [m2 s−1] DN2O = diffusion coefficient of N2O [m2 s−1] Deig = effective membrane diffusion coefficient [m2 s−1] 2 −1 Dm ig = molecular diffusion coefficient of CO2 [m s ] 2 −1 k Dig = Knudsen diffusion coefficient of CO2 [m s ] E = enhancement factor E∞ = asymptotic infinite enhancement factor Gij = temperature-dependent parameter for blended amine viscosity calculation H = Henry’s constant [kPa·m3 kmol−1] Ha = Hatta number k2,amine = reaction rate constant for amine [m3 mol−1 s−1] k2,R1R2NH = second-order forward reaction rate constant for primary or secondary amines [m3 mol−1 s−1] k2,PZ = second-order reaction rate constant for PZ [m3 mol−1 s−1] k−1 = reverse first-order reaction rate constant [s−1]

Subscripts

in = inlet of membrane contactor out = outlet of membrane contactor W = water L = liquid phase of membrane contactor G = gas phase of membrane contactor Superscripts

■

* = equilibrium EXP = experimental

REFERENCES

(1) Tan, C. S.; Chen, J. E. Absorption of carbon dioxide with piperazine and its mixtures in rotating packed bed. Sep. Purif. Technol. 2006, 49, 174−180. (2) Rochelle, G. T. Amine scrubbing for CO2 capture. Science 2009, 325, 1652−1654. (3) Gabelman, A.; Hwang, S. T. Hollow fiber membrane contactors. J. Membr. Sci. 1999, 159, 61−106. (4) Yeon, S.-H.; Lee, K.-S.; Sea, B.; Park, Y.-I.; Lee, K.-H. Application of pilot-scale membrane contactor hybrid system for removal of carbon dioxide from flue gas. J. Membr. Sci. 2005, 257, 156−160.

12180



Article

(5) Yan, S.; Fang, M.; Zhang, W.; Zhong, W.; Luo, Z.; Cen, K. Comparative analysis of CO2 separation from flue gas by membrane gas absorption technology and chemical absorption technology in China. Energy Convers. Manage. 2008, 49, 3188−3197. (6) Lu, J.; Zheng, Y.; Cheng, M. Wetting mechanism in mass transfer process of hydrophobic membrane gas absorption. J. Membr. Sci. 2008, 308, 180−190. (7) Zhang, H. Y.; Wang, R.; Liang, D. T.; Tay, J. H. Modeling and experimental study of CO2 absorption in a hollow fiber membrane contactor. J. Membr. Sci. 2006, 279, 301−310. (8) Zhu, D. C.; Fang, M. X.; Lv, Z.; Wang, Z.; Luo, Z. Y. Selection of blended solvents for CO2 absorption from coal-fired flue gas. Part 1: Monoethanolamine (MEA)-based solvents. Energy Fuels 2012, 26, 147−153. (9) Muraleedharan, R.; Mondal, A.; Mandal, B. Absorption of carbon dioxide into aqueous blends of 2-amino-2-hydroxymethyl-1,3-propanediol and monoethanolamine. Sep. Purif. Technol. 2012, 94, 92−96. (10) Dey, A.; Aroonwilas, A. CO2 absorption into MEA-AMP blend: mass transfer and absorber height index. Energy Procedia 2009, 1, 211− 215. (11) Aroonwilas, A.; Veawab, A. Integration of CO2 capture unit using blended MEA-AMP solution into coal-fired power plants. Energy Procedia 2009, 1, 4315−4321. (12) Vaidya, P.; Kenig, E. Y. Absorption of CO2 into aqueous blends of alkanolamines prepared from renewable resources. Chem. Eng. Sci. 2007, 62, 7344−7350. (13) Rodríguez, N.; Mussati, S.; Scenna, N. Optimization of postcombustion CO2 process using DEA-MDEA mixtures. Chem. Eng. Res. Des. 2011, 89, 1763−1773. (14) Choi, W. J.; Seo, J. B.; Jang, S. Y.; Jung, H. J.; Oh, K. J. Removal characteristics of CO2 using aqueous MEA/AMP solutions in the absorption and regeneration process. J. Environ. Sci. 2009, 21, 907− 913. (15) Bougie, F.; Iliuta, M. C. Analysis of regeneration of sterically hindered alkanolamines aqueous solutions with and without activator. Chem. Eng. Sci. 2010, 65, 4746−4750. (16) Gong, Y.; Wang, Z.; Wang, S. C. Experiments and simulation of CO2 removal by mixed amines in a hollow fiber membrane module. Chem. Eng. Process. 2006, 45, 652−660. (17) Lu, J.; Wang, L.; Sun, X.; Li, J.; Liu, X. Absorption of CO2 into aqueous solutions of methydiethanolamine and activated methyldiethanolamine from a gas mixture in a hollow fiber contactor. Ind. Eng. Chem. Res. 2005, 44, 9230−9238. (18) Paul, S.; Ghoshal, A. K.; Mandal, B. Removal of CO2 by single and blended aqueous alkanalamine solvents in hollow-fiber membrane contactor: modeling and simulation. Ind. Eng. Chem. Res. 2007, 46, 2576−2588. (19) Franco, J.; de Montigny, D.; Kentish, S.; Perera, J.; Stevens, G. A study of the mass transfer of CO2 through different membrane material in the membrane gas absorption process. Sep. Sci. Technol. 2008, 43, 225−244. (20) Caplow, M. Kinetics of carbamate formation and breakdown. J. Am. Chem. Soc. 1968, 90, 6795−6803. (21) Danckwerts, P. V. The reaction of CO2 with ethanolamines. Chem. Eng. Sci. 1979, 34, 443−446. (22) Little, R. J.; Van Swaaij, W. P. M.; Versteeg, G. F. Kinetics of carbon dioxide with tertiary amines in aqueous solution. AIChE J. 1990, 36, 1633−1640. (23) Zhang, X.; Zhang, C. F.; Qin, S. J.; Zheng, Z. S. A kinetics study on the absorption of carbon dioxide into a mixed aqueous solution of methyldiethanolamine and piperazine. Ind. Eng. Chem. Res. 2001, 40, 3785−3791. (24) Blauwhoff, P. M. M.; Versteeg, G. F.; Van Swaaij, W. P. M. A study on the reaction between CO2 and alkanolamines in aqueous solutions. Chem. Eng. Sci. 1984, 39, 207−225. (25) Liao, C. H.; Li, M. H. Kinetics of absorption of carbon dioxide into aqueous solutions of monoethanolamine+N-methyldiethanolamine. Chem. Eng. Sci. 2002, 5, 4569−4582.

(26) Ko, J. J.; Li, M. H. Kinetics of absorption of carbon dioxide into solution of N-methyldiethanolamine+water. Chem. Eng. Sci. 2000, 55, 4139−4147. (27) Ali, S. H. Kinetics of the reaction of carbon dioxide with blends of amines in aqueous media using the stopped-flow technique. Int. J. Chem. Kinet. 2005, 37, 391−405. (28) Xu, S.; Wang, Y. W.; Otto, F. D.; Mather, A. E. Kinetics of the reaction of carbon dioxide with 2-amino-2-methyl-1-propanol solutions. Chem. Eng. Sci. 1996, 51, 841−850. (29) Mandal, B. P.; Biswas, A. K.; Bandyopadhayay, S. S. Absorption of carbon dioxide into aqueous blends of 2-amino-2-methyl-1propanol and diethanolamine. Chem. Eng. Sci. 2003, 58, 4137−4144. (30) Rinker, E. B.; Ashour, S. S.; Sandall, O. C. Kinetics and modeling of carbon dioxide absorption into aqueous solutions of Nmethyldiethanolamine. Chem. Eng. Sci. 1995, 50, 755−768. (31) Kreulen, H.; Smolders, C. A.; Versteeg, G. F.; van Swaaij, W. P. M. Microporous hollow fiber membranes as gas-liquid contactors. Part 2. Mass transfer with chemical reaction. J. Membr. Sci. 1993, 78, 217− 238. (32) Kreulen, H.; Smolders, C. A.; Versteeg, G. F.; van Swaaij, W. P. M. Microporous hollow fibre membrane modules as gas-liquid contactors. Part 1. Physical mass transfer processes. A specific application: mass transfer in highly viscous liquids. J. Membr. Sci. 1993, 78, 197−216. (33) Mavroudi, M.; Kaldis, S. P.; Sakellaropoulos, G. P. A study of mass transfer resistance in membrane gas-liquid contacting processes. J. Membr. Sci. 2006, 272, 103−115. (34) Zheng, J. M.; Xu, Y. Y.; Xu, Z. K. Shell side mass transfer characteristics in a parallel flow hollow fiber membrane module. Sep. Sci. Technol. 2003, 38, 1247−1267. (35) Wu, J.; Chen, V. Shell-side mass transfer performance of randomly packed hollow fiber modules. J. Membr. Sci. 2000, 172, 59− 74. (36) Wickramasinghe, S. R.; Semmens, M. J.; Cussler, E. L. Mass transfer in various hollow fiber geometries. J. Membr. Sci. 1992, 69, 235−250. (37) Prasad, P.; Sirkar, K. K. Dispersion-free extraction with microporous hollow fiber module. AIChE J. 1988, 34, 177−188. (38) Fang, M.; Wang, Z.; Yan, S.; Cen, Q.; Luo, Z. CO2 desorption from rich alkanolamine solution by using membrane vacuum regeneration technology. Int. J. Greenhouse Gas Conrol 2012, 9, 507−521. (39) Jamal, A. Absorption and Desorption of CO2 and CO in Alkanolamine Systems. Dissertation, University of British Columbia, Vancouver, 2002. (40) Grunberg, L.; Nissan, A. H. Mixture law for viscosity. Nature 1949, 164, 799−800. (41) Mandal, B. P.; Kundu, M.; Bandyopadhyay, S. S. Density and viscosity of aqueous solutions of (N-methyldiethanolamine + monoethanolamine), (N-methyldiethanolamine + diethanolamine), (2-amino-2-methyl-1-propanol + monoethanolamine), and (2-amino2-methyl-1-propanol + diethanolamine). J. Chem. Eng. Data 2003, 48, 703−707. (42) Versteeg, G. F.; van Swaaij, W. P. M. Solubility and diffusivity of acid gases (CO2 and N2O) in aqueous alkanolamine solution. J. Chem. Eng. Data 1988, 33, 29−34. (43) Glasscock, D. A. Modeling and Experimental Study of Carbon Dioxide Absorption in Aqueous Alkanolamines. Dissertation, University of Texas at Austin, Austin, 1990. (44) Gilliland, E. R. Diffusion coefficients in gaseous systems. Ind. Eng. Chem. 1934, 26, 681−685. (45) Reid, R. C.; Prausnitz, J. M.; Sherwood, T. K. The Properties of Gases and Liquids, 3rd ed.; McGraw-Hill: New York, 1977. (46) Cussler, E. L. Diffusion: Mass Transfer in Fluid Systems, 3rd ed.; Cambridge University Press: New York, 2009. (47) Wang, Y. W.; Xu, S.; Otto, F. D.; Mather, A. E. Solubility of N2O in alkanolamines and in mixed solvents. Chem. Eng. J. 1992, 48, 31−40. 12181



Article

(48) Bensetiti, Z.; Iliuta, I.; Larachi, F.; Grandjean, B. P. A. Solubility of nitrous oxide in amine solutions. Ind. Eng. Chem. Res. 1999, 38, 328−332. (49) DeCoursey, W. J. Absorption with chemical reaction: development of a new reaction for the Danckwerts model. Chem. Eng. Sci. 1974, 29, 1867−1872.

12182


Optimization of Blended Amines for CO2 Absorption in a Hollow-Fiber

Recommend Documents