Investigation on the Phase-Change Absorbent System MEA + Solvent

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Investigation on the phase-change absorbent system MEA + solvent A (SA)+H2O used for the CO2 capture from flue gas Kun Zhu, Houfang Lu, Changjun Liu, Kejing Wu, Wei Jiang, Jingxing Cheng, Siyang Tang, Hairong Yue, Yingying Liu, and Bin Liang Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.8b04985 • Publication Date (Web): 07 Feb 2019 Downloaded from http://pubs.acs.org on February 10, 2019

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Industrial & Engineering Chemistry Research

Investigation on the phase-change absorbent system MEA+ solvent A (SA)+H2O used for the CO2 capture from flue gas Kun Zhu †‡, Houfang Lu †‡, Changjun Liu †, Kejing Wu ‡, Wei Jiang †, Jingxing Cheng †, Siyang Tang †, Hairong Yue †, Yingying Liu ‡*, Bin Liang †‡.

† School of Chemical Engineering, Sichuan University, Chengdu, 610065, China; ‡ Institute of New Energy and Low Carbon Technology, Sichuan University, Chengdu, 610207, China ABSTRACT

Amine-based CO2 post-combustion capture is a commercialized technology for recovery CO2 from flue gas and the energy consumption is concentrated in the regenerative stage. The phase-change solvent system, monoethanolamine (MEA, absorbent) + solvent A (SA, solvent) + water (solvent) was investigated on their performances of CO2 absorption and desorption. CO2-rich solvent can be continuously separated for the feeding flow of regeneration tower; thereby, it significantly reduces the energy of regeneration. The results show that the solvent with suitable composition can easily separate CO2-rich phase from the solvent and the CO2-rich phase accounts for more than 98% of the total CO2. At 293–333 K, the temperature has no significant effect on the absorption

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Industrial & Engineering Chemistry Research 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

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rate and phase equilibrium. The estimation of energy consumption shows that the reboiler work can be reduced by 43.6% compared with that using 20 wt.% MEA aqueous solution.

1. Introduction CO2 emission by the combustion of fossil fuels has resulted in serious concerns for the global warming climate1. Although global CO2 emission has remained flat since 2014 to 2016, it reached a record-breaking 41 Gt in 20172. In order to control the global temperature rise under 1.4–2 °C, substantial reductions of CO2 emissions is required. More than 30% of CO2 comes from the flue gas of coal-fired power plants3. Currently, postcombustion CO2 capture is the most acceptable technology because it does not significantly change the equipment and process of power plants and it is the most effective large-scale capture technology in the short term4–6. Considering the high temperature, high flow rate, and low CO2 concentration of flue gas, chemical absorption with organic amines is maturer than physical adsorption, membrane separation, and low-temperature separation methods7. However, regeneration energy consumption with organic amines is high, which may reduce the net electricity efficiency by approximately 25% for the power plant equipped with a traditional MEA absorption CO2 post-capture device8,9. More and more endeavors are paid on the development of new absorbents to cut the regeneration energy consumption. Traditional amine absorption processes use 20–30 wt.% MEA as solvents and its regeneration energy consumption accounts for more than 60% of the total energy consumption10. The regeneration energy includes reaction heat, sensible heat and vaporization heat. New absorbents with low reaction heat, high boiling point and low heat capacity can reduce the energy consumption. There have been substantial researches on absorbents, such as methyldiethanolamine (MDEA), hindered-amine 2-amino-2-methyl-propyl alcohol (AMP)11, ammonia12,etc.. They have


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relative low reaction heats, but their absorption rates are normally much lower compared with those of primary amine. One of the alternative ways to reduce regeneration energy is to cut the liquid volume, such as separating CO2 rich phase by phase-change solvents, which also reduces the sensible heat and vaporization heat. Phase-change absorbents automatically separate into CO2-rich and CO2-lean phases after CO2 absorption. It makes it possible to pump only the CO2-rich phase into the regeneration tower and recycle CO2-lean phase to the absorption tower. Hereby, it greatly reduces the regenerated absorbent volume and lowers the heat charge of the regeneration tower. Typical phase-change absorbent systems include the alkanolamines + lipophilic amine + H2O and alkanolamines + nonamine organic solvent + H2O systems13,14. Zhang et al.15–17 proposed a phase transition system containing alkanolamines + lipophilic amine, and investigated over 30 types of lipophilic amines, such as cycloheptylamine (CHPA), dibutylamine (BDA), and 1-ethylpiperidine (EPD). Absorbents containing primary amines or secondary amines are usually used as an absorbent activator, and tertiary amines are used as regeneration promoters because of their lower dielectric constants and thermal phase transition behavior. The DMXTM reagent developed by IFP is also this type of absorbent18–20, which can reduce the desorption reboiler work to 2.1 GJ/ton CO2. Furthermore, BDA-N,N-Diethylmonoethanolamine (DEEA)-H2O and 3-dimethyl aminopropyl amine (MAPA)-DEEA-H2O systems were studied by Xu et al.21,22 and Pinto et al.23, respectively. They found that two phases were formed after the absorption of CO2 because DEEA is insoluble in the absorption products, and the DEEA-rich phase is the upper phase due to its lower density. Recently, DEEA phase-change systems such as N,N-dimethylbutylamine (DMBA)-DEEA-H2O24， triethylenetetramine (TETA)-DEEA-H2O25 and 2-((2-aminoethyl) amino) ethanol (AEEA)DEEA-H2O26 show excellent absorption capacity, phase-change behavior at room temperature and


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significant potential in energy saving, and attract more and more attentions. However, lipophilic amines, predominantly tert-amine, can also absorb CO2 in aqueous solution. For instance, in the BDA-DEEA-H2O system, DEEA can react with CO2 and H2O and the product is soluble in the lower phase. As a result, DEEA in the upper phase gradually decreases and the volume of the lower phase increases. Although the alkanolamines + lipophilic amine aqueous systems show higher CO2 loading and lower desorption temperature, higher volume fraction of the rich phase makes the energy consumption reduced by phase-change less obvious, especially at high CO2 loading. Compared with the alkanolamines + lipophilic amine aqueous system, the alkanolamines + nonamine organic solvent + H2O system is stabler and has a lower volume percentage of the rich phase. Zhang et al. 27 developed an MEA-propanol-H2O system based on the salting-out effect. The solubility of propanol in the system reduced with increasing salt concentration, which led to separation of the solvent (propanol) and enrichment of the absorption products in the lower phase. In addition, diethylenetriamine (DETA)-sulfolane-H2O system was proposed27. Sulfolane is a good solvent for physical absorption of CO2 due to its high thermal stability, low corrosivity, and promotion effect on desorption, which can improve the circulation capacity by 35% compared with 30 wt.% MEA aqueous system. Moreover, the MEA/DEA-heptanol/octanol system and Nmethylmonomonoethanolamine (MMEA)-diethylene glycol dimethyl ether (DEGDEE) system were proposed by the Korea Institute of Energy Research28 and the Italy ICCOM Institute29, respectively. In the latter system, the cycling absorption efficiency can be 97.6%. Recently, Yang et al.30 developed MDEA/C4-C6 alcohol/H2O system by means of the cosolvent effect, which expanded the understanding of the phase-change mechanism. However, there are some disadvantages in the existing systems: the low boiling point of propanol, the high toxicity of


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sulfolane, the significant decline of absorption capacity in the MEA/DEA-heptanol/octanol system, and the relatively high price of MMEA. SA is a kind of ether. It is a non-proton polar solvent, which is miscible with water. Its boiling point is above 250 °C, and its low vapor pressure reduces the loss of solvent during the desorption process. Replacing part of the water with SA to develop a MEA-based phase-change absorbent is promising. Although the ether phase-change system of MMEA-DEGDEE has been reported, MEA is the most commonly used chemical absorbent and the development of MEA-based phase-change absorbent is easier to popularize in the existing process31. For capturing CO2 from flue gas with low energy consumption, the MEA-SA-H2O phase-change absorbent has advantages regarding the material price, low solvent loss, and decreased energy consumption due to reduced sensible heat, vaporization heat, and desorption reboiler work. In this study, a novel MEA-SA-H2O phase-change system was proposed. The absorption and separation performance of the MEA-SA-H2O system was studied for different solvent/water ratios, and the component ratio was optimized in view of the absorbing capacity, absorbing rate, viscosity, and CO2-rich phase volume. Then, the effect of temperature on the absorption and phase equilibrium was investigated and the mechanism of the phase-change was inferred. In addition, the energy consumption performance of the system was estimated. 2. Experimental 2.1 Materials The purities of MEA (purchased from Chengdu Chron Reagent) and SA were >98%. The deionized water has an electrical resistivity of 18.25 MΩ·cm at 295.1 K. The gases were purchased from Chengdu Dongfeng Gas Plant. The purities of CO2 and N2 were >99% and >99.99%, respectively.


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2.2 Absorption and desorption experiments The absorption and desorption experimental device is shown in Figure 1. In the absorption experiment, CO2 and N2 from the cylinders were mixed with the two-channel mass flow controller (MFC, CS200, Beijing Sevenstar Electronics Co. Ltd., China) to form the simulated flue gas with 12.5 vol.% CO2. The pressure was slightly higher than atmosphere. The mixture entered the flask and reacted with the solvent. An oil bath was used to maintain the temperature in the flask and the absorption temperature was 293–333 K. Using 600 r/min mechanical agitation to ensure full contact between the gas and liquid and the outlet gas passed through concentrated sulfuric acid and a cotton-filled drying tube before flowing into a CO2 infrared ray analyzer (FN316B, Shaanxi FeiEnTe Instrument Technology Co. Ltd., China) to detect the CO2 concentration. When the gas analyzer showed that the concentration reaches 12.5 vol.% and does not change for 30 min, the absorbent was considered saturated with CO2 and the experiment was stopped. After absorption, the product was taken out and separated into two phases. Only the CO2-rich phase was sent for desorption. Use 200 mL/min N2 as the purge gas and set the oil bath temperature at 383 K. The outlet gas also passed through the drying process into the CO2 analyzer. The cyclic absorption-desorption experiment was conducted by mixing the CO2-rich phase after desorption with the CO2-lean phase. The mixture was used to absorb CO2 again at 313 K then desorbed the CO2-rich phase by the same operation after phase separation. 2.3 Indexes and parameters involved Because N2 cannot be absorbed, the macroscopic rate of absorption (rabs) and desorption (rdes) can be calculated according to the Ideal Gas Law using the N2 flowrate, the inlet and outlet CO2 gas concentration as follows:

rabs  K

GN2





( 0  1 ) L 1  0 1  1

(1)


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rdes  K

GN2



( 1 ) L 1  1

(2)

where L represents the absorbent volume (L), ω0 and ω1 are the CO2 concentrations in the inlet gas and outlet gas, respectively. GN2 is the volume flow rate of N2 (L/min) at standard condition determined by MFC, and K is the inverse of the ideal gas molar volume (4.5×104 mmol/L). The time integration of the absorption rate is the absorption amount. The phase separation time is determined as follows: After the absorption process, stop the agitation, and place the absorbed solutions in a water bath at the corresponding absorption temperature. Keep observing the phase interface. When the phase interface is clear and no longer changes, it is considered that the phase separation is complete and the corresponding time is recorded as the phase separation time. The total CO2 concentration (cCO2) means the CO2 concentration in solution, including molecular, carbamate, carbonate, and bicarbonate CO2 forms; it is determined by BaCO3 precipitation titration method32 and the error is ±0.05 mol/kg. The total amine concentration (cA) refers to the MEA concentration in solution, including MEA molecular, carbamate, and proton MEA; it is determined by titration with hydrochloric acid22. The second titration jump point is taken as the endpoint and the error is ± 0.05 mol/kg. The water content (cw) is determined by KarlFischer method with an error of ± 0.02 kg/kg. The density is determined by the 5-mL pycnometer test method, with a precision of 0.0001 g/mL. The viscosity is determined by a cone-plate viscometer (DV2TRVCP, Brookfield, USA), with an error of ±2%. The temperature during density and viscosity measurement is controlled by a circulating water bath and the temperature change is ±0.2 °C. All analysis results are the average values of triplicate experiments.


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CO2 loading (α) and the regeneration ratio (θ) indicate the saturation of CO2 in solution and the degree of desorption, respectively, as shown in equations (3) and (4).





cCO2 cMEA

(3)

 bdes   ades  bdes

(4)

where αbdes and αades represent the CO2 loadings before and after desorption (mol CO2/mol MEA). 2.4 13C NMR analysis A sample of 40 μL is placed in a nuclear magnetic tube and dissolved with a certain amount of D2O. 30 μL of 1,4-dioxane is added as an internal standard and oscillated until clear and transparent. One-dimensional (1D) 13C spectra are obtained using a Bruker AV-Ⅱ-400MHz NMR spectrometer. The testing temperature is 298 K, the acquisition time is 10 μs, and the delay time is 2 s for a total of 256 scans. The 13C NMR spectra are analyzed with MestReNova software. 2.5 Determinations of reaction heat and thermal capacity According to the literature33, the reaction heat of absorption (∆H) and constant pressure heat capacity (Cp) are tested by a simultaneous thermal analyzer (STA 449F3 Jupiter, Netzsch, Germany). STA combines thermogravimetric analysis (TG) with differential scanning calorimetry (DSC) to measure the mass and heat flux of the absorption process so that the CO2 absorption amount and absorption heat can be calculated by the following formula: q

Y2  Y1 44

H 

H * q

(5)

(6)


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where q is the amount of CO2 absorption (mol CO2/g absorbent), Y2 is the CO2 mass fraction after absorption (%), Y1 is the CO2 mass fraction before absorption (%), and ∆H* is the integral area of the heat flux (kJ/mol). The constant pressure heat capacity (Cp) is determined with a DSC curve and calculated using the following equation34,35:

Cp 

60(AS -A B ) msβ

(7)

where AS and AB are the heat flux in the sample test and blank test, respectively (mW), ms is the mass of the sample (mg), and β is the heating rate of 10 K/min. 2.6 Calculation method of dipole moment B3LYP functionals is chosen based on the density functional theory36. For C, H, O, and N atoms using the 6-311++g (d,p) whole electron base group, the potential energy surface is scanned and the stability structure of each molecule is stabilized at the minimum point. Then, the molecular configuration and polarity data are obtained based on the Natural Bond Orbital charge analysis. 3. Results and discussion 3.1 CO2 absorption and phase separation In the MEA-SA-H2O absorption system, MEA is the main absorbent of CO2, SA and water are mainly used as solvents. The effects of solvent ratios on CO2 absorption and phase-change are shown in Table 1. Before absorbing CO2, the mixtures of SA and water with different ratios are all transparent homogeneous liquids. After absorption, two situations exist with different SA ratios. When the fraction of SA is less than 40 vol.%, the solution is still homogeneous and clear after absorbing CO2; however, when the fraction of SA is higher than or equal to 40 vol.%, phasechange occurs after absorption, as shown in Figure 2. The solution turns into a yellow turbid liquid after absorbing CO2 and two clear layers appear after a period of settling.


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Table 1 shows that the absorption capacity decreases with the decrease of water in the solvent. Moreover, the capacity of the system with a volume ratio of 2:8:0 is 0.37 mol/kg less than that of the system with a volume ratio of 2:0:8. In the aqueous MEA solution, there are three main reactions:

HOCH 2 CH 2 NH 2 +CO 2   HOCH 2 CH 2 NH +2 COO-

(8)

HOCH 2 CH 2 NH +2 COO - +HOCH 2 CH 2 NH 2   HOCH 2 CH 2 NHCOO- +HOCH 2 CH 2 NH 3+ (9)

ˆ ˆˆ†ˆ HCO3- +NH 2 CH 2 CH 2 OH (10) HOCH 2 CH 2 NHCOO - +H 2 O ‡ In the aqueous solution, in addition to reactions (8) and (9), a hydrolysis reaction (10) of the absorption product carbamate occurs, forming bicarbonate and MEA37, and the MEA generated from hydrolysis can continue to absorb CO2 to increase the absorption capacity. For systems with a volume ratio increase from 2:4:4 to 2:7:1, the absorption capacity decreases by 2.5% with a reduction of water. However, for the 2:8:0 system (without water), its absorption capacity is 13% less than that of the system with a ratio of 2:4:4. The absorption rate of each absorption system is shown in Figure 3. The overall absorption rate decreases slowly at first, then quickly, and finally slowly again until it reaches a balance. With the increase of SA, the initial absorption rate increases; for the 2:8:0 system, it is 16.2% higher than that of the 2:0:8 system. Thus, the solvent effect of SA on the system significantly increases the absorption rate. The effect of SA on the absorption rate may be similar to that of methanol reported in the literature38 because the relatively better solubility of CO2 benefits CO2 diffusion in the solution and decreases the liquid film resistance. Complete phase-change to form CO2-poor and CO2-rich phases is favorable for cycling absorption. The physicochemical properties and components of the two phases will affect the phase behavior of the CO2-absorbed system, such as density and viscosity. An obvious density


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difference of the two phases is more favorable for separation, while high viscosity inhibits the mass transfer process, lowering the phase-change rate. The density and viscosity of each phase are shown in Figure 4. For the system with a volume ratio of 2:4:4 for MEA, SA, and water, both the upper and lower phases have similar properties. However, for the other systems, the densities and viscosities are quite different between the upper and lower phases. With the increase of the SA volume fraction, the density and viscosity decrease in the upper phase and increase in the lower phase. The viscosities of the lower phases for systems with volume ratios of 2:7:1 and 2:8:0 are higher than other systems, resulting in high liquid-liquid mass transfer resistance and slow phase separation rates. As for the 2:4:4 system, the difference of density and viscosity between the two phases is small; thus, the phase separation rate is slower than that of 2:6:2 and 2:5:3 systems. Table S1 shows the total CO2 concentration, total amine concentration, and water content in the upper and lower phases after phase separation. The concentration of each component has the same regularity with density and viscosity. Except for the 2:4:4 system, the component concentration is very different in the upper and lower phases. With an increase of the initial SA volume fraction, the total amine concentration and total carbon concentration in the lower phase increase gradually. Based on 100 g of absorbent, the mass distribution of components in absorption product (Figure 5) is calculated using the concentration data in Table S1. Most of the MEA, CO2, and water are separated into the lower phase, and SA is mostly in the upper phase; only a very small part of SA enters the lower phase. The mass ratio of water in the upper phase and lower phase is close to 1:3. In all phase-change systems, the amount of CO2 in the lower phase is more than 85% of the total, which indicates that the system enriches CO2 into one phase. In theory, a smaller volume of the lower phase indicates low energy consumption. The lower phase volume fraction of the 2:8:0 system is the lowest, but the viscosity of the lower phase is too high, which inhibits fluid flow in


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the pipeline and transfer of heat and mass. From the 2:7:1 to 2:4:4 system, the absorption capacity does not change substantially, and the viscosity of the lower phase is not too high. Although the 2:6:2 system has more advantages in terms of the phase separation rate, the 2:7:1 system has the lowest phase volume fraction, which exhibits the greatest potential advantages for energy saving. As for lower phase separation rate caused by relative high viscosity, some solutions could be used to overcome the disadvantage, such as ultrasonic method39 or appropriate increased temperature. As a result, the 2:7:1 system is selected to study the effect of temperature on CO2 capture. 3.2 Effect of temperature on the absorption performance Theoretically, high temperature leads to a rapid reaction rate. However, the absorption capacity will decrease with an increase of temperature. The absorption performance of the 2:7:1 system and its absorption rate are studied at different temperatures. The results are shown in Table 2 and Figure 6 separately. From Figure 6, temperature has no significant effect on the change rule of absorption rate from 293 K to 333 K. Increasing the absorption temperature can increase intrinsic reaction rate, while it will reduce the equilibrium solubility of CO2 in solution and thus decrease the driving force. The combination of these two factors results in a similar absorption rate within 293-333 K. With an increase of temperature, the absorption capacity and the volume fraction of the lower phase decrease. Although the amplitude of the decrease is small, the absorption capacity is reduced by 11.5% from 293 K to 333 K. Figure 7 shows the density and viscosity of the upper and lower phases under different temperatures. Both the density and viscosity decrease with an increase of temperature, and the viscosity of the lower phase is more sensitive to temperature. The viscosity at 333 K decreases by


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89.8% compared with that at 293 K; this large viscosity change causes an obvious difference in the phase separation time which is shown in Table 2. Table S2 and Figure 8 show the total CO2 concentration, total amine concentration, water content of the upper and lower phase, and the distribution of components in the two phases after absorption at different temperatures. Temperature has no significant effect on the distribution of components after absorption, and the amount of CO2 in the lower phase is more than 98 % of the total. Therefore, within the experimental temperature range, the absorption performance is not highly sensitive to temperature. Considering the viscosity of the lower phase and temperature of the flue gas, 323 K is a suitable temperature for absorption. 3.3 Analysis for phase-change behavior A theoretical understanding of phase-change behavior is meaningful for developing a new highperformance phase-change absorbent. As can be seen from Table S1, the α value of the lower phase is approximately 0.50, which is the theoretical absorption capacity of MEA. In this situation, MEA is mainly in the form of MEA carbamate, i.e., a protonated MEA and MEA carbamate40. Before absorption, MEA is homogeneously dissolved in the solvent. After absorption, MEA exists as MEA carbamate, forming the CO2-rich phase together with water and a small amount of SA. Most SA is separated from the system, forming the CO2-poor phase with some water. The salt effect can explain this behavior41–43. Adding salt to a non-electrolyte aqueous solution will change the activity coefficient of the non-electrolyte, which will increase or decrease the solubility of the non-electrolyte in the system. In the MEA-SA-H2O system, SA is a nonelectrolyte, while MEA carbamate is a salt. Zhang et al.31 also explained the phase-change phenomenon with the salt-out effect; the formation of carbamate decreased the solubility of


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propanol in the salt solution, resulting in two phases. The salt effect has also been observed in different phase-change systems14,22,23,27,31,44. According to van der Waals theory45,46, the salt effect depends on the electrostatic force and dispersion force between ions and molecules. If electrostatic forces play the dominant role, the electrostatic forces between ions and water will be stronger than those between ions and the nonelectrolyte. As a result, water will gather around the ions, reducing the solubility of the nonelectrolyte, leading to salt-out. If the ions and non-electrolyte molecules are similar in structure, the dispersion force will be strong, and it often displays salt-in. The dipole moment and the molecular size of the main components in the MEA-SA-H2O system are calculated, as shown in Table 3. On the one hand, the polarity of SA is weak, and the polarity of carbamate is very high; thus, the electrostatic effect should not be strong between them. On the other hand, the size of carbamate is smaller than SA, and the role of dispersion force is not significant, so salt-out effect may occur. Therefore, the dipole moment difference and the molecular diameter between the salt formed after absorbing CO2 and the non-electrolyte solvent are in favor of the phase-change behavior31. 3.4 Absorption-desorption circulating The effective circulation of absorbent is the guarantee to realize CO2 capture. Figure 9 shows the results of three cycles for the 2:7:1 absorption system. The three CO2 desorption processes are consistent, and the absorption rate of the first cycle is higher than the subsequent two cycles, which is due to incomplete desorption. Samples of the lower phase before the first desorption and after the third desorption were analyzed by 13C NMR to investigate the change in each component (Figure 10). Before desorption, CO2 exists mainly in the form of MEA carbamate, as well as a small amount of carbonate or bicarbonate. After desorption, the carbonate or bicarbonate


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disappear, with CO2 only remaining in carbamate form, but the 4’ peak and 5’ peak indicate a decrease in carbamate concentration. The 1’ peak and 2’ peak shift to low frequencies after desorption, mainly because the protonated MEA translates to MEA during desorption, so that the cations decrease its shielding effect on α and β carbon. A little of non-numbered peaks exist before and after desorption and remain basically unchanged, which belong to the solvent SA. The results of NMR indicate that the desorption reaction is a process of CO2 release and transformation from carbamate, and protonated MEA to MEA, which is the reverse of the absorption reaction process. In addition, the mass difference after each desorption is no more than 2%, indicating that the solvent and MEA mass loss are very small during the desorption process. A set of 10 cycles of experiments was used to verify the stability of the reversible cycle and cyclic capacity for a long-time. As can be seen in Figure 11, the absorption capacity is maintained at 0.48-0.53 mol CO2/mol MEA, and the average of the cyclic capacity is 0.35±0.02 mol CO2/mol MEA. 3.5 Estimation of regeneration energy consumption for MEA + SA + H2O system with a volume ratio of 2:7:1 This section aims to estimate the regeneration energy consumption of the MEA+SA+H2O phasechange system and compare the results with that of the 20 wt.% MEA aqueous solution to evaluate the potential energy saving of the phase-change system. This estimation only focuses on the heating load of the reboiler in the desorption process. In the desorption process of the phase-change system, the CO2-rich phase needs to be separated by the decanter and heated to a certain temperature through the heat exchanger before entering the desorption tower. Then, the CO2-lean phase after desorption is pumped back into the absorption tower, in contrast to the traditional


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aqueous MEA solution desorption process. Figure 12 shows the absorption-desorption process for the phase-change system. According to the energy consumption calculation method suggested by Leites10, when desorbing a ton of CO2, the energy consumption, Qstr, can be divided into three parts: the heat of reaction (QR), the sensible heat for rising temperature (Qs), and the latent heat for vaporization of rich liquid components (Qv). The calculation equations are: QS 

QR 

QV 

C p mL T mCO2 nCO2 H mCO2



nw H H2O mCO2

C p T



(11)

cCO2 MrCO2 H MrCO2



(12)

nw  H H2O nCO2  MrCO2

(13)

where ∆T is the temperature difference between the top and bottom of the desorption tower (K), MrCO2 is the molecular weight of CO2 (44 g/mol), and nw/nCO2 is the mole ratio in the top of the tower, which equals to pH2O/pCO2. Regeneration energy consumption is affected greatly by the operation conditions; therefore, more gas-liquid equilibrium data are required for a detailed energy calculation. Here, it is compared with 20 wt.% MEA aqueous solution under the same assumed desorption conditions. We assume that the desorption conditions are as follows: 1. Full absorption is achieved in the absorption tower, the absorbent is separated to heat the CO2rich phase to 363 K, and the rich phase is taken into the desorption tower. 2. The desorption temperature is 383 K, the desorption pressure is 101.3 kPa, and the regeneration ratio is set to 50 %.


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3. In the estimation of latent heat, as the boiling points of MEA and SA are much higher than the desorption temperature, their latent heat of vaporization are ignored. Moreover, only water is required to reflux. 4. pH2O/pCO2, is predominantly determined by the water content in the liquid and regeneration temperature47.

ptotal  pCO2  pH2O (14) pH2O  xw pwvap

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pH2O on the top of the tower is determined by Raoult’s law 48. The temperature on the top is 363 vap

K, and the corresponding saturated vapor pressure ( pw ) is 70.1 kPa. For the phase-change system, the mole per cent of water is approximately 40%. And the pressure ratio, pH2O/pCO2, could be set to 0.38. For the 20 wt.% MEA aqueous solution, the mole per cent of water is approximately 88% and the pressure ratio could be 1.55. Firstly, the thermodynamic parameters for estimating energy consumption need to be measured, i.e., the desorption reaction heat and constant pressure heat capacity (Cp) of the rich phase. The desorption reaction heat is approximately equal to the reaction heat during the absorption process (∆H). The curves are shown in Figure S1 and Figure S2. Calculations indicate that the reaction heat of the system is 88.2 kJ/mol; this result matches the reaction heat of MEA aqueous solution in literature33. All parameters required for estimation are listed in Table S3 and Table S4. As shown in Table 4, the phase-change system reduces the total energy consumption by reducing the sensible heat and latent heat in the desorption process. The regeneration energy consumption of the phase-


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change system is 43.6% lower than that of the 20 wt.% MEA aqueous solution for per ton of CO2 released. It should be noticed here that the energy consumption estimate in this work is the reboiler duty at the assumed conditions, and the optimization of operating parameters and the heat recovery are not taken into account. Thus, the calculation result of the 20 wt.% MEA aqueous solution in this paper is much higher than those reported in literature49. If the heat recovery ratio of Qs and Qv is considered at 50%26, the energy consumption of the 20 wt.% MEA aqueous solution will be reduced to 3.74 GJ/ton CO2, while the energy consumption of the phase-change absorbent will be reduced to 2.55 GJ/ton CO2. Compared with the energy consumption of other phase-change systems reported in Table 5, this energy consumption is slightly higher, with lack of process optimization. 4. Conclusion In this study, a phase-change CO2 absorption system of MEA+SA+H2O is proposed. The system can automatically separate into two phases after absorption of CO2 where the lower phase is CO2rich phase, in which CO2 occupies more than 85% of the total. MEA-SA-H2O solution with a volume ratio of 2:7:1 exhibits stable CO2 absorption capacity and phase-change behavior from 293–333 K, and good performance in multiple absorption-desorption cycles. The estimated regeneration energy consumption indicates that the system reduces the heat load of the reboiler by 43.6% more than that of 20 wt.% MEA aqueous solution. SA can increase the CO2 absorption rate. The salt effect can be used to explain the phase-change behavior. Although this preliminary study on the MEA-SA-H2O system shows good absorption and phasechange performance and excellent reduction of energy consumption, it should be noted that the viscosity of the enriched phase is high and its high concentration of carbamate could cause


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degradation and corrosion problems under high desorption temperatures. These issues require further exploration.

ASSOCIATED CONTENT Supporting Information. The methods and data used in absorption and desorption (Part 1), the parameters and data used in estimation of regeneration energy consumption (Part 2). (PDF)

AUTHOR INFORMATION Corresponding Author * Contact email: [email protected] ACKNOWLEDGMENT The authors gratefully thank the National Natural Science Foundation of China (No.21878190) for financial support. REFERENCES (1)

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TABLES. Table 1. Absorption capacity and phase-change performance of the MEA-SA-H2O system Absorption VMEA:VSA:VH2O (vol. capacity ratio) (mol/kg)

Number of PhaseVolume percentage of phases after separation the lower phase % absorption time

2:0:8

1.78

1

——

——

2:1:7

1.74

1

——

——

2:2:6

1.70

1

——

——

2:3:5

1.63

1

——

——

2:4:4

1.62

2

83.1

~1 h

2:5:3

1.62

2

49.7

Investigation on the Phase-Change Absorbent System MEA + Solvent

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