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Modeling CO_{2}, H_{2}S, COS, and CH_{3}SH Simultaneous Removal Using Aqueous Sulfolane–MDEA Solution

^{1}

^{2}

^{*}

## Abstract

**:**

_{2}and H

_{2}S) and organic sulfur (COS and CH

_{3}SH) from natural gas with an aqueous sulfolane–MDEA solution. First, the accuracy of the thermodynamic model was validated by comparing the calculated partial pressure of CO

_{2}, H

_{2}S, and CH

_{3}SH with those of the experimental data reported in the literature. Then, the industrial test data were employed to validate the absorption model and the simulation results agreed well with the experimental data. The average relative errors of the removal rates of CO

_{2}, COS, and CH

_{3}SH are 3.3%, 3.0%, 4.1%, respectively. Based on the validated coupled model, the total mass transfer coefficient and mass transfer resistance of each solute component at different column positions were analyzed. The effects of the gas–liquid ratio, overflow weir height, and absorption pressure on the absorption performance of each component were studied, and the influence of the acid component concentration in the feed gas on the removal efficiency of methyl mercaptan (CH

_{3}SH) was also discussed. It is found that the improved absorption model can better characterize the absorption performance and be conducive to the optimal design of the absorber column.

## 1. Introduction

_{2}and H

_{2}S) and organic sulfur (COS and CH

_{3}SH), it needs further treatment before it can be used [4,5,6,7]. Conventional gas processing includes distillation, adsorption, membrane separation, and absorption [8]. Among these purification techniques, chemical absorption is the most commonly used method for acid-gas removal in the natural gas industries with its high efficiency and simultaneous strip of multiple acid gases [9].

_{2}and H

_{2}S content in raw gas. A common sulfone–amine solution is composed of N-methyl diethanolamine (MDEA), sulfolane, and water, which has an advantage in treating natural gas with high organic sulfur content [8,13]. Macgregor and Mather [14] reported the absorption of H

_{2}S and CO

_{2}with a mixed solvent consisting of MDEA (20.9 wt.%) + sulfolane (30.5 wt.%). Jou et al. [15] investigated acid-gas absorption with MDEA, methanethiol, and ethanethiol at 40 and 70 ℃ at thiol partial pressure within 0.1–15.8 kPa. Haghtalab et al. [16] measured the absorption of acid-gas in different mix solvents at 343 K and a total pressure of 0.1–0.21 kPa and reported variation in the solubility of CO

_{2}and H

_{2}S in different operating conditions and solvent compositions. However, although absorption with various mixed solvents has been reported, reliable thermodynamic modeling of the absorption process remains a big challenge for simulating and optimization of the acid-gas capture processes [17].

_{2}capture using a 2-amino-2-methyl-1-propanol (AMP) solution in a packed column. The simulation results of the absorber show that the rate-based model can better predict temperature and concentration curves than the equilibrium phase model. Al-Baghli et al. [22] simulated the process of removing CO

_{2}and H

_{2}S by MEA and DEA aqueous solutions using rate-based gas absorber model and obtained reliable simulation results. Pacheco and Rochelle [23] developed a framework to perform selective absorption of H

_{2}S using MDEA solution from a gas stream containing CO

_{2}. The Maxwell–Stefan and enhancement factor theories are used in the model. Mandal and Bandyopadhyay [24] conducted theoretical and experimental research on the simultaneous absorption of CO

_{2}and H

_{2}S into solutions containing MDEA and DEA. Moioli et al. [25] used the Eddy diffusivity theory in Aspen Plus and used an external subroutine to simulate the absorption of CO

_{2}and H

_{2}S from gas streams. In addition, they modified the parameters for vapor–liquid equilibrium (VLE) calculations and verified the simulation using data in the literature. Yang et al. [26] used an aqueous sulfone–MDEA solution to remove organic sulfur in natural gas and studied the factors that affect organic sulfur removal in natural gas purification devices. However, the process of the simultaneous removal of multiple impurities (i.e., CO

_{2}, H

_{2}S, COS, and CH

_{3}SH) is usually very complicated, and the influence of acidic gas on the absorption performance of organic sulfur during the removal process is still unclear.

_{2}and H

_{2}S) and organic sulfur (COS and CH

_{3}SH) from natural gas with an aqueous sulfolane–MDEA solution. An improved thermodynamic model was used considering the influence of acid component concentration on the removal efficiency of methyl mercaptan (CH

_{3}SH). The mass transfer characteristics at different column positions were analyzed. The influence of the gas–liquid ratio, overflow weir height, and absorption pressure on the removal rate of each component and the influence of acid component content on the removal rate of methyl mercaptan (CH

_{3}SH) were studied. Furthermore, with the modified rate-based absorption model, we successfully simulated the adsorption process for a wide range of feed gas compositions and operating conditions. This study is expected to further facilitate optimization of the operating conditions and device structure.

## 2. Model Theory

#### 2.1. Thermodynamic Framework

#### 2.1.1. Gas–Liquid Equilibrium

_{2}and H

_{2}S) and organic sulfur (COS and CH

_{3}SH) in the aqueous sulfolane–MDEA solution is calculated by Henry’s law:

_{i}is Henry’s law constant of component i in the mixed solvent of water, sulfolane, and MDEA, Pa. ${y}_{i}$ is the mole fraction of component i in the vapor phase, and ${x}_{i}$ is the equilibrium mole fraction of component i in the liquid phase. ${\phi}_{i}$ is the fugacity coefficient of component i in the vapor phase, which is calculated using the Peng–Robinson equation of state (EOS) [27]. ${\gamma}_{i}^{*}$ is the unsymmetric activity coefficient of component i in the mixed solvent solution of water, sulfolane, and MDEA. ${\gamma}_{i}^{*}$ is normalized to the mixed solvent infinite dilution reference state.

_{A}is the weighting factor, which can be calculated according to the method in the literature [13]. ${\gamma}_{iA}^{\infty}$ is the infinite dilution activity coefficient of component i in pure solvent A, ${H}_{iA}$ is Henry’s constant of component i in pure solvent A, and the values of ${H}_{iA}$ of solute i (CO

_{2}, H

_{2}S and CH

_{4}) can be obtained directly from the literature [10,11,12,13]. Some ${H}_{iA}$ values of CH

_{4}, CH

_{3}SH, and COS regressed according to the experimental data [29,30,31], which are summarized in Table 1.

_{3}SH) in the aqueous sulfolane–MDEA solution [26]. Equation (3) is used to quantitatively describe the influence of the acid component content on Henry’s constant for CH

_{3}SH.

^{+}in aqueous solutions. C is the effect factor, symbolizing the influence of the acid component on Henry’s law constant. ${H}_{C{H}_{3}SH}^{\prime}$ and ${H}_{C{H}_{3}SH}$ are Henry’s constants of CH

_{3}SH in the mixed solvent when the influence of acid component is not taken into account and is taken into consideration, respectively.

#### 2.1.2. Aqueous Phase Chemical Equilibrium

_{2}S and CO

_{2}react with MDEA, and the ionic equilibrium reactions are expressed as Equations (4)–(9) [35].

_{i}is the mole fraction of reactant component i in reaction j, ${x}_{{i}^{\prime}}$ is the mole fraction of product component $i\prime $ in reaction j, and ${\gamma}_{i}$ and ${\gamma}_{i\prime}$ are the unsymmetric activity coefficients of reactant component i and product component i′ in the aqueous solution. The unsymmetric activity coefficients are normalized to the aqueous phase infinite dilution reference state. ${v}_{i\prime}$ and ${v}_{i\prime}$ represent the stoichiometric coefficients of reactants and products, respectively.

#### 2.2. Rate-Based Model Assumptions

_{2}S and CO

_{2}) and organic sulfur (COS and CH

_{3}SH) from natural gas utilizing a tray column, a mathematical model was established using the two-film theory [37]. The basic assumptions are as follows:

#### 2.3. Material and Energy Balance

^{−1}. L and V represent the mole flow rates of liquid and vapor, respectively, kmol s

^{−1}. N is the mole transfer rate, kmol s

^{−1}. r is the reaction rate, which can be calculated from the component concentration. ${x}_{i}$ and ${y}_{i}$ are the mole fractions of component i in the liquid and vapor phases, respectively. Q is the heat input to a stage, J s

^{−1}; q is the heat transfer rate, J s

^{−1}; ${H}^{FL}$ and ${H}^{FV}$ are the enthalpies of the inflow liquid and inflow vapor, J kmol

^{−1}; and the thermodynamic properties including enthalpy and heat capacity utilized in heat transfer calculations can be obtained directly from the literature [10,11,12,13].

#### 2.4. Mass Transfer and Enhancement Factor

^{2}m

^{−2}; ${A}_{c}$ is the cross-sectional area of the column, m

^{2}; ${P}_{i}$ is the partial pressure of component i in the gas bulk; and ${P}_{i}^{*}$ is the partial pressure of component i in equilibrium with the liquid phase. The partial pressure of each component can be calculated by the established thermodynamic model.

_{i}is Henry’s law constant of component i; ${K}_{G,i}$ is the overall mass transfer of component i in gas phase; and ${k}_{G,i}$ and ${k}_{L,i}$ are the mass transfer coefficient of component i without reaction in the gas phase and the liquid phase, which are calculated according to the method used by Simon et al. [18] and Saimpert et al. [40]. The density, viscosity, and surface tension required for the calculation can be obtained directly from the Aspen database. ${E}_{i}$ is the enhancement factor of component i due to chemical reactions, which is the ratio of the mass transfer enhanced by a reaction over the mass transfer without the reaction [21]. The calculation of the enhancement factor in this work uses the formula developed by Danckwerts [41].

_{MDEAL}is the diffusivity of the MDEA in the liquid phase. ${C}_{i}^{In}$ is the concentration of absorption component i at the gas-liquid interface, and ${C}_{MDEA}^{Bulk}$ represents the concentration of MDEA in the bulk liquid. Pacheco and Rochelle [23] and Al-Ghawas et al. [32] measured the rate constants of reaction between CO

_{2}and MDEA and between COS and MDEA in the mixed solvent, respectively, as shown in Equations (22) and (23).

_{2}and COS, the enhancement factor is needed for the mass transfer rate calculations, which are calculated using Equation (18). As the reaction between H

_{2}S and MDEA is very fast, the enhancement factor can be calculated using the infinite fast reaction rate, as shown in Equation (19). In contrast, the absorption of CH

_{3}SH is mainly a physical effect, so there is also no need to consider the influence of chemical reactions.

#### 2.5. Computational Implementation

## 3. Modeling Results

_{2}and H

_{2}S) and organic sulfur (COS and CH

_{3}SH) from natural gas with an aqueous sulfolane–MDEA solution.

_{3}SH) is also discussed.

#### 3.1. Thermodynamics Model Validation

_{2}partial pressure data for the CO

_{2}–H

_{2}O–MDEA system and H

_{2}S partial pressure data for the H

_{2}S–H

_{2}O–MDEA system at different compositions and different temperatures. As shown, the calculated CO

_{2}/H

_{2}S partial pressure values are in good agreement with the experimental data reported by the literature [14,43,44,45], thus validating the proposed thermodynamic model. From Figure 2, it is found that the CO

_{2}partial pressure is more sensitive to temperature. With the temperature increasing from 313.15 K to 338.75 K, the CO

_{2}partial pressure experiences a significant increase. However, the MDEA mass fraction seems to have little effect on the CO

_{2}partial pressure when the temperature is 313.15 K. Figure 3 shows the comparison of experimental data and calculated data in terms of H

_{2}S partial pressure for the H

_{2}S–H

_{2}O–MDEA system. As shown, the proposed thermodynamic model is also effective and robust. This means that the improved thermodynamic model has good applicability.

_{2}–H

_{2}O–sulfolane–MDEA system and the H

_{2}S–H

_{2}O–sulfolane–MDEA system, are used to validate the thermodynamic model. The comparison results of the calculated values and experimental data in terms of CO

_{2}/H

_{2}S partial pressure are given in Figure 4 and Figure 5. As shown, the calculated results agree well with the experimental data in terms of CO

_{2}/H

_{2}S partial pressure. Generally, the addition of sulfolane is beneficial to enhancing the absorption of acidic gases, and therefore, high acidic gas loading is observed at low partial pressure. When compared with Figure 4 and Figure 5, it can be found that H

_{2}S has a higher solubility than CO

_{2}in the H

_{2}O–sulfolane–MDEA solution.

_{2}and H

_{2}S partial pressure data calculated in this paper are compared with the values calculated with reaction parameters proposed by Austgen et al. [36]. The average relative deviations of the partial pressure data of CO

_{2}and H

_{2}S calculated in this paper are 10.7% and 24.4%, respectively, while the average relative deviations of the partial pressure data calculated with reaction parameters proposed by Austgen et al. [36] are 38.1% and 36.8%, respectively. The calculated results using the improved thermodynamic model are not only in good agreement with the experimental data but also better than the calculated values of the existing models in the literature (see Figure 6 and Figure 7). This is because we refitted the exponents of the equilibrium equations of reactions (4) and (5) and calculated the reaction equilibrium constants with the new parameters in Table 2.

_{3}SH in the system, the calculated value of the model is compared with the value measured by Jou et al. [15] in Figure 8. It can be seen that the predicted CH

_{3}SH partial pressure data are in good agreement with the experimental value. This shows that the improved thermodynamic model not only has higher prediction accuracy but also has a larger application range.

#### 3.2. Rate-Based Absorption Model Validation

_{2}O, 40 wt% sulfolane, and 40 wt% MDEA. The tower was operated under a different number of plates (22, 26, and 30) and overflow weir heights (0.1 and 0.15 m). The Inlet gas–liquid ratio (vol./vol.) was adjusted from 490.2 to 720.0. CO

_{2}and H

_{2}S loading were varied from 4.11vol% to 4.73 vol% and 1.34 vol% to 1.57 vol%, respectively. The concentration of CH

_{3}HS and COS were changed from 14.12 to 25.41 mg/m

^{3}and 11.18 to 34.58 mg/m

^{3}, respectively. As shown in Table 3, the simulation results of removal efficiency for absorption components are in good agreement with the data from the industrial absorber. For CO

_{2}, COS, and CH

_{3}SH, the average relative errors are only 3.3%, 3.0%, and 4.1%, respectively. Therefore, the rationality of the absorption model can be confirmed.

#### 3.3. Rate-Based Absorption Model Calculation

_{2}O, sulfolane, and MDEA in the mixed solvent are 0.20, 0.40, and 0.40, respectively. The parameters of the absorber are obtained from (the desulfurizing column of) PetroChina Southwest Oil and Gasfield Company, and the operating parameters are obtained by analyzing the operating parameters of typical industrial absorbers and by making appropriate extensions.

_{2}S is fast, and the concentration of H

_{2}S in the feed gas drops rapidly after entering the absorption column; it is completely removed in the middle of the column. This is the result of the synergistic effect of physical absorption and chemical reaction. However, the absorption rate of CH

_{3}SH is low. After the feed gas enters the column, the concentration of CH

_{3}SH decreases slowly and the final removal rate is low. The main reason for this is that the physical adsorption and absorption processes are controlled by equilibrium.

_{2}S and CO

_{2}. It can be seen that the gas phase and the liquid phase exchange heat in the column, and as the concentration of H

_{2}S and CO

_{2}increases, the heat release during the absorption process increases, resulting in an increase in the outlet temperature of the liquid phase at the top of the column. When the feed composition is the same, the H

_{2}S is completely absorbed in the middle section of the column, and then the heat release is rapidly reduced, resulting in a rapid drop in the gas phase temperature at the top of the column.

_{2}and COS, the gas mass transfer resistance can almost be ignored. In the whole absorption column, the gas mass transfer coefficient is less than 0.01 kmol/m

^{2}/kPa. For H

_{2}S, however, the gas mass transfer coefficient is almost ten times greater than that of CO

_{2}and COS. Additionally, the liquid mass transfer coefficient of H

_{2}S is also much higher than that of CO

_{2}and COS, accounting for 15–25% of the total mass transfer resistance, so the gas-side mass transfer resistance cannot be ignored. The gas-side mass transfer resistance of CH

_{3}SH accounts for 5–10% of the total mass transfer resistance, which also needs to be considered in the calculation process. This is consistent with the calculation method of the absorption model.

#### 3.4. The Influence of the Operating Parameters

_{3}SH because CH

_{3}SH is physically absorbed and controlled by equilibrium. When the gas–liquid ratio increases from 400 to 800, the removal efficiency of CH

_{3}SH decreases from 0.8 to 0.4. Obviously, a more liquid phase is conducive to the physical absorption process. Since the absorption rate of CO

_{2}and COS is determined by mass transfer rate, it is less affected by the gas–liquid ratio.

_{2}in the absorption solvent is very fast, the absorption rate is mainly controlled by the mass transfer rate of CO

_{2}. As the overflow weir height increases, the liquid thickness over the tray increases, and, thus, the mass transfer area increases, resulting in an increase in the removal rates of CO

_{2}and COS. For CH

_{3}SH, the overflow weir height has little effect on its removal rate. One reason is that the absorption rate of CH

_{3}SH is controlled by its physical solubility in the mixed solvent. The second reason is that the increase in the overflow weir height strengthens the absorption of CO

_{2}and H

_{2}S, which reduces the solubility of CH

_{3}SH in the liquid phase. Figure 15 shows the effect of acid component content in the feed gas on the removal rate of CH

_{3}SH. As the acid component content increases, the absorption rate of CH

_{3}SH decreases significantly. The reason for this is that, as the acid component content in the liquid phase increases, the Henry constant of CH

_{3}SH absorption increases significantly.

_{3}SH achieves a significant increase when compared with the removal efficiency of CO

_{2}and COS. Due to the chemical absorption being the rate-limited step for the removal of CO

_{2}and COS, the absorption pressure on their removal efficiency is insignificant. As shown in Figure 14, the removal efficiency of COS increases by less than 10% when the absorption pressure increases from 4 MPa to 8 MPa.

## 4. Conclusions

_{2}and H

_{2}S) and organic sulfur (COS and CH

_{3}SH) from natural gas with an aqueous sulfolane–MDEA solution. The thermodynamic model was improved by improving the calculation method of the chemical equilibrium constant and by incorporating the influence of acid components for COS removal, which made the gas–liquid equilibrium (GLE) data more accurate. The absorption model was validated by the experimental data obtained from an industrial device, and the average relative errors of the removal rates of CO

_{2}, COS, and CH

_{3}SH obtained by experiment and calculation are 3.3%, 3.0%, and 4.1%, respectively. The validated coupled model indicated that the gas mass transfer resistance can almost be ignored for CO

_{2}and COS, but both gas and liquid mass transfer resistance for H

_{2}S and CH

_{3}SH should be considered. Additionally, the analysis of influencing factors shows that the gas–liquid ratio has a greater impact on the physical absorption process controlled by equilibrium but has a limited effect on the absorption process determined by a chemical reaction. An increased overflow weir height increases the gas–liquid contact time and mass transfer area, thus resulting in an increase in the removal rates of CO

_{2}and COS, which, however, inhibits the absorption of CH

_{3}SH due to the increased Henry coefficient. Furthermore, the physical absorption process rather than the chemical absorption process is more sensitive to the tower pressure.

## Author Contributions

## Funding

## Institutional Review Board Statement

## Informed Consent Statement

## Data Availability Statement

## Conflicts of Interest

## Nomenclature

P | system pressure | Pa |

y_{i} | mole fraction of component i in the vapor phase | - |

x_{i} | mole fraction of component i in the liquid phase | - |

H_{i} | Henry’s law constant of component i | Pa m^{3}/mol |

φ_{i} | fugacity coefficient of component i in the vapor phase | - |

γ_{i}* | unsymmetric activity coefficient in the mixed solvent solution | - |

w_{A} | weighting factor | - |

γ^{∞} | infinite dilution activity coefficient | - |

H_{iA} | Henry’s constant of component i in pure solvent A | Pa m^{3}/mol |

H_{CH3SH} | Henry’s constants of CH_{3}SH considering the influence of acid gas | Pa m^{3}/mol |

H_{CH3SH’} | Henry’s constants of CH_{3}SH ignoring the influence of acid gas | Pa m^{3}/mol |

f | unremoved rate of MDEA | |

c | effect factor | |

K | chemical equilibrium constant | |

F | mole flow rate of feed | kmol/s |

L | mole flow rate of liquid | kmol/s |

V | mole flow rate of liquid | kmol/s |

N | mole transfer rate | kmol/s |

Q | heat input | J/s |

q | heat transfer rate | J/s |

H^{F} | enthalpy of feed | J/kmol |

H^{V} | enthalpy of the vapor | J/kmol |

H^{L} | enthalpy of the liquid | J/kmol |

K_{G} | overall mass transfer coefficient | kmol/m^{2} s kPa |

P_{i}^{*} | partial pressure of component i in equilibrium with the liquid phase | - |

Pi | partial pressure of component i in the gas bulk | - |

a_{p} | effective mass transfer area of the column per unit area of the tray | m^{2}/m^{2} |

A_{c} | cross-sectional area of the column | m^{2} |

k_{G} | mass transfer coefficient without reaction in the gas phase | kmol/m_{2} s kPa |

k_{L} | mass transfer coefficient without reaction in the liquid phase | m/s |

E | enhancement factor | - |

Ha | Hatta number | - |

E_{∞} | enhancement number for infinite fast reactions | - |

D_{L} | diffusivity in an aqueous sulfolane–MDEA solution | m^{2}/s |

C^{In} | concentration at the gas–liquid interface | kmol/m^{3} |

C^{Bulk} | concentration in the bulk liquid | kmol/m^{3} |

k_{2t} | rate constant | m^{3}/kmol s |

T | temperature | K |

Subscripts | ||

i | component i | |

i’ | product i’ | |

j | stage number | |

Superscripts | ||

L | liquid phase | |

V | vapor phase |

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**Figure 2.**Comparison of experimental data and calculated data of the CO

_{2}–H

_{2}O–MDEA system under different conditions. ■, measured by Jou et al. [43], T = 313.15 K, MDEA mass fraction = 0.35; ●, measured by Macgregor and Mather [14], T = 313.15 K, MDEA mass fraction = 0.209; ▲, measured by Qian et al. [44], T = 310.95 K, MDEA mass fraction = 0.20; ▼, measured by Qian et al. [44], T = 338.75 K, MDEA mass fraction = 0.20; and ◆, measured by Sidi-Boumedine et al. [45], T = 313.15 K, MDEA mass fraction = 0.257. The solid line represents the calculated data in this work.

**Figure 3.**Comparison of experimental data and calculated data of the H

_{2}S–H

_{2}O–MDEA system under different conditions. ■, measured by Jou et al. [43], T = 313.15 K, MDEA mass fraction = 0.35; ●, measured by Jou et al. [43], T = 313.15 K, MDEA mass fraction = 0.50; ▲, measured by Macgregor and Mather [14], T = 310.95 K, MDEA mass fraction = 0.209; ▼, measured by Qian et al. [44], T = 338.75 K, MDEA mass fraction = 0.20; and ◆, measured by Qian et al. [44], T = 338.75 K, MDEA mass fraction = 0.20. The solid line represents the calculated data in this work.

**Figure 4.**Comparison of the experimental data and calculated data for the CO

_{2}–H

_{2}O–sulfolane–MDEA system under different conditions. ■, measured by Jou et al. [43], T = 313.15 K, MDEA mass fraction = 0.305, sulfolane mass fraction = 0.209; ●, measured by Jou et al. [43], T = 373.15 K, MDEA mass fraction = 0.305, sulfolane mass fraction = 0.209. The solid line represents the calculated data in this work.

**Figure 5.**Comparison of the experimental data and calculated data of the H

_{2}S–H

_{2}O–sulfolane–MDEA system under different conditions. ■, measured by Jou et al. [43], T = 313.15 K, MDEA mass fraction = 0.305, sulfolane mass fraction = 0.209; ●, measured by Jou et al. [43], T = 373.15 K, MDEA mass fraction = 0.305, sulfolane mass fraction = 0.209. The solid line represents the calculated data in this work.

**Figure 8.**Parity plot for CH

_{3}SH partial pressure, experimental data vs. model predictions. CH

_{3}SH–H

_{2}O–MDEA system: ■, T = 313.15 K, MDEA mass fraction = 0.50; ▲, T = 343.15 K, MDEA mass fraction = 0.50. CO

_{2}–H

_{2}S–CH

_{3}SH–H

_{2}O–MDEA system: ●, T = 313.15 K, MDEA mass fraction = 0.50; ▼, T = 343.15 K, MDEA mass fraction = 0.50.

**Figure 9.**The concentration profile of each component in the absorption column (overflow weir height: 0.15 m; absorption pressure: 6 MPa; inlet gas flow rate: 5.0 × 10

^{6}Nm

^{3}/d; inlet gas–liquid ratio: 600; inlet gas loading: CO

_{2}, 4.5 vol.%; H

_{2}S, 1.5 vol.%; COS, 15 mg/m

^{3}; and CH

_{3}SH, 15 mg/m

^{3}).

**Figure 10.**The gas–liquid temperature profile of each component in the absorption column (overflow weir height: 0.15 m; absorption pressure: 6 MPa; inlet gas flow rate: 5.0 × 10

^{6}Nm

^{3}/d; inlet gas–liquid ratio: 600).

**Figure 11.**Total mass transfer coefficient and mass transfer resistance of different column positions (overflow weir height: 0.15 m; absorption pressure: 6 Mpa; inlet gas flow rate: 5.0 × 10

^{6}Nm

^{3}/d; inlet gas–liquid ratio: 600; inlet gas loading: CO

_{2}, 4.5 vol.%; H

_{2}S, 1.5 vol.%; COS, 15 mg/m

^{3}; CH

_{3}SH, 15 mg/m

^{3}).

**Figure 12.**The effect of gas–liquid ratio on the removal efficiency and the acid components (CO

_{2}and H

_{2}S)/MDEA ratio (number of plates: 25; overflow weir height: 0.15 m; absorption pressure: 6 Mpa; inlet gas flow rate: 5.0 × 10

^{6}Nm

^{3}/d; inlet gas loading: CO

_{2}, 4.5 vol.%; H

_{2}S, 1.5 vol.%; COS, 15 mg/m

^{3}; CH

_{3}SH, 15 mg/m

^{3}).

**Figure 13.**The effect of overflow weir height on removal efficiency (number of plates: 25; absorption pressure: 6 Mpa; inlet gas flow rate: 5.0 × 10

^{6}Nm

^{3}/d; inlet gas–liquid ratio: 600; inlet gas loading: CO

_{2}, 4.5 vol.%; H

_{2}S, 1.5 vol.%; COS, 15 mg/m

^{3}; CH

_{3}SH, 15 mg/m

^{3}).

**Figure 14.**The effect of absorption pressure on removal efficiency (number of plates: 25; overflow weir height: 0.15 m; inlet gas flow rate: 5.0 × 10

^{6}Nm

^{3}/d; inlet gas–liquid ratio: 600; inlet gas loading: CO

_{2}, 4.5 vol.%; H

_{2}S, 1.5 vol.%; COS, 15 mg/m

^{3}; CH

_{3}SH, 15 mg/m

^{3}).

**Figure 15.**The effect of acid component content in feed gas on the removal efficiency of CH

_{3}SH (number of plates: 25; overflow weir height: 0.15 m; absorption pressure: 6 Mpa; inlet gas flow rate: 5.0 × 10

^{6}Nm

^{3}/d; and inlet gas–liquid ratio: 600).

Solute i | Solvent A | A | B | C | D | Data Source |
---|---|---|---|---|---|---|

CH_{4} | sulfolane | 26.68 | −1538.38 | 0 | 0.02 | Jou et al. [31] |

CH_{3}SH | H_{2}O | 21.128 | −1299.310 | 0 | 0 | Bedell and Miller [29] |

CH_{3}SH | sulfolane | 12.987 | 0 | 0 | 0 | Bedell and Miller [29] |

COS | H_{2}O | 27.402 | −2407.192 | 0 | 0 | Al-Ghawas et al. [32] |

COS | MDEA | 19.323 | −603.363 | 0 | 0 | Al-Ghawas et al. [32] |

COS | sulfolane | 11.004 | 0.170 | 0 | 0.015 | Shokouhi et al. [30] |

Reaction | A | B/T | C | D/K^{−1} | T Range/K | Source |
---|---|---|---|---|---|---|

4 | 819.8 | −37655.9 | −124.5 | 0 | 273–498 | In this Work ^{a} |

5 | −553.4 | 28412.7 | 77.7 | 0 | 273–423 | In this Work ^{a} |

6 | −9.4165 | −4234.98 | 0 | 0 | 298–333 | Austgen et al. [36] |

7 | −32.0 | −3338.0 | 0 | 0 | 287–343 | Austgen et al. [36] |

8 | 216.049 | −12431.7 | −35.4819 | 0 | 273–498 | Austgen et al. [36] |

9 | 132.899 | −13445.9 | −22.4773 | 0 | 273–498 | Austgen et al. [36] |

^{a}The equilibrium reaction order of CO

_{2}and H

_{2}S are obtained by fitting the experimental data. The reaction orders of CO

_{2}and HCO

_{3}

^{−}are 1.25, the reaction orders of H

_{2}S and HS

^{−}are 1.35, and the reaction orders of the remaining reactions are 1.

Experimental No. | Number of Plates | Overflow Weir Height (m) | Inlet Gas Flow Rate (10^{4} Nm^{3}/d) | Inlet Gas-Liquid Ratio (vol/vol) | Pressure (Mpa) | Inlet Gas Loading H_{2}S (vol.%) | Inlet Gas Loading CO_{2} (vol.%) | CH_{3}SH Concentration (mg/m^{3}) | COS Concentration (mg/m^{3}) | Experimental Removal Efficiency (%) | Simulated Removal Efficiency (%) | ||||
---|---|---|---|---|---|---|---|---|---|---|---|---|---|---|---|

CO_{2} | COS | CH_{3}SH | CO_{2} | COS | CH_{3}SH | ||||||||||

1 | 26 | 0.15 | 384.0 | 490.2 | 6.14 | 1.49 | 4.26 | 15.53 | 12.33 | 74.1 | 82.9 | 76.9 | 76.4 | 80.2 | 70.7 |

2 | 26 | 0.15 | 417.0 | 568.4 | 6.14 | 1.44 | 4.17 | 15.62 | 14.64 | 70.4 | 79.2 | 61.9 | 72.6 | 78.3 | 62.8 |

3 | 26 | 0.15 | 446.0 | 608.6 | 6.13 | 1.41 | 4.32 | 15.53 | 11.18 | 73.3 | 78.0 | 56.3 | 73.1 | 78.4 | 58.1 |

4 | 30 | 0.15 | 535.2 | 614.3 | 6.20 | 1.55 | 4.62 | 22.42 | 32.53 | 73.4 | 79.4 | 57.7 | 73.4 | 80.3 | 58.6 |

5 | 26 | 0.15 | 458.0 | 622.8 | 6.13 | 1.41 | 4.31 | 15.75 | 13.78 | 76.6 | 76.6 | 53.9 | 72.6 | 78.1 | 56.8 |

6 | 26 | 0.15 | 553.0 | 645.4 | 6.11 | 1.49 | 4.73 | 20.20 | 41.18 | 69.7 | 81.1 | 56.8 | 70.3 | 76.5 | 54.0 |

7 | 30 | 0.15 | 553.4 | 655.1 | 6.44 | 1.57 | 4.58 | 25.41 | 34.58 | 74.8 | 79.7 | 52.1 | 76.6 | 82.0 | 56.1 |

8 | 26 | 0.15 | 570.0 | 659.7 | 6.18 | 1.50 | 4.71 | 19.40 | 29.94 | 68.8 | 79.1 | 51.2 | 70.9 | 76.9 | 53.2 |

9 | 26 | 0.15 | 442.0 | 661.1 | 6.18 | 1.40 | 4.28 | 16.09 | 23.57 | 72.0 | 79.1 | 51.5 | 72.1 | 78.2 | 54.0 |

10 | 30 | 0.15 | 600.2 | 661.8 | 6.11 | 1.52 | 4.26 | 24.39 | 28.08 | 70.4 | 76.7 | 50.5 | 71.8 | 79.0 | 53.7 |

11 | 26 | 0.15 | 448.0 | 662.4 | 6.18 | 1.41 | 4.17 | 16.12 | 24.29 | 73.6 | 79.1 | 50.4 | 71.7 | 77.9 | 54.1 |

12 | 26 | 0.15 | 596.0 | 665.8 | 6.11 | 1.51 | 4.23 | 23.49 | 19.78 | 67.8 | 75.9 | 54.1 | 69.3 | 75.3 | 53.2 |

13 | 22 | 0.15 | 600.0 | 666.0 | 6.11 | 1.52 | 3.69 | 20.54 | 20.81 | 57.1 | 67.2 | 53.2 | 59.5 | 67.6 | 54.5 |

14 | 22 | 0.15 | 602.3 | 673.0 | 6.10 | 1.51 | 4.49 | 24.33 | 25.01 | 61.0 | 71.0 | 54.0 | 59.5 | 67.9 | 52.7 |

15 | 30 | 0.15 | 612.0 | 674.3 | 6.19 | 1.51 | 4.31 | 24.31 | 31.01 | 66.8 | 74.3 | 56.6 | 71.7 | 79.0 | 53.3 |

16 | 26 | 0.15 | 589.0 | 681.2 | 6.16 | 1.50 | 4.51 | 14.85 | 21.59 | 66.7 | 70.9 | 52.9 | 68.9 | 75.6 | 51.9 |

17 | 26 | 0.15 | 614.7 | 720.0 | 6.26 | 1.51 | 4.48 | 17.42 | 17.45 | 66.7 | 75.9 | 47.7 | 69.3 | 76.0 | 49.5 |

18 | 26 | 0.1 | 475.0 | 547.6 | 6.19 | 1.35 | 4.34 | 15.49 | 17.66 | 67.0 | 61.2 | 66.3 | 65.9 | 63.0 | 65.0 |

19 | 26 | 0.1 | 399.0 | 558.9 | 6.15 | 1.39 | 4.37 | 15.23 | 11.90 | 71.6 | 64.8 | 66.5 | 68.5 | 65.9 | 63.2 |

20 | 30 | 0.1 | 412.0 | 575.8 | 6.14 | 1.38 | 4.18 | 14.63 | 11.82 | 68.4 | 66.1 | 63.2 | 71.2 | 68.2 | 62.4 |

21 | 26 | 0.1 | 407.0 | 576.0 | 6.15 | 1.39 | 4.30 | 14.95 | 11.57 | 70.0 | 61.7 | 63.0 | 67.9 | 65.4 | 61.6 |

22 | 30 | 0.1 | 414.0 | 582.3 | 6.15 | 1.38 | 4.37 | 14.12 | 12.23 | 67.7 | 68.9 | 64.2 | 71.6 | 68.9 | 61.4 |

23 | 30 | 0.1 | 448.0 | 630.8 | 6.14 | 1.38 | 4.15 | 14.26 | 13.48 | 68.4 | 67.7 | 53.2 | 72.7 | 71.0 | 56.4 |

24 | 26 | 0.1 | 599.0 | 653.3 | 6.27 | 1.45 | 4.23 | 16.70 | 26.95 | 60.8 | 57.3 | 52.9 | 61.8 | 59.6 | 56.2 |

25 | 26 | 0.1 | 470.0 | 653.8 | 6.22 | 1.36 | 4.11 | 15.48 | 9.35 | 65.9 | 61.3 | 56.6 | 65.2 | 63.0 | 55.6 |

26 | 30 | 0.1 | 595.0 | 656.2 | 6.42 | 1.35 | 4.53 | 17.76 | 25.27 | 65.6 | 63.3 | 55.1 | 68.3 | 66.5 | 56.5 |

27 | 22 | 0.1 | 605.0 | 660.6 | 6.26 | 1.34 | 4.53 | 17.78 | 24.15 | 61.1 | 55.7 | 53.6 | 55.5 | 54.0 | 55.1 |

28 | 26 | 0.1 | 602.0 | 661.0 | 6.51 | 1.46 | 4.58 | 18.48 | 30.04 | 62.7 | 59.5 | 57.1 | 62.6 | 60.7 | 57.0 |

29 | 30 | 0.1 | 603.0 | 663.6 | 6.45 | 1.45 | 4.41 | 17.35 | 27.17 | 66.3 | 63.1 | 56.9 | 68.8 | 67.2 | 56.2 |

30 | 26 | 0.1 | 597.0 | 664.9 | 6.25 | 1.41 | 4.45 | 17.15 | 28.45 | 61.8 | 56.8 | 59.5 | 61.5 | 59.7 | 54.7 |

31 | 22 | 0.1 | 599.0 | 671.3 | 6.31 | 1.35 | 4.50 | 17.21 | 28.49 | 61.1 | 55.1 | 57.0 | 55.3 | 53.7 | 54.8 |

32 | 22 | 0.1 | 612.0 | 681.6 | 6.35 | 1.35 | 4.60 | 17.72 | 29.77 | 58.9 | 53.6 | 54.7 | 55.2 | 53.8 | 54.2 |

Parameters | Data |
---|---|

Column diameter (m) | 3.4 |

Number of plates | 25 |

Overflow weir height (m) | 0.08–0.2 |

Inlet gas flow rate (10^{4} Nm^{3}/d) | 500 |

Inlet gas temperature (°C) | 40 |

Inlet liquid temperature (°C) | 20 |

Inlet gas-liquid ratio | 400–800 |

Absorption pressure (MPa) | 4.0–8.0 |

Inlet gas loading CO_{2} (vol.%) | 0–6.0 |

Inlet gas loading H_{2}S (vol.%) | 0–6.0 |

CH_{3}SH concentration (mg/m^{3}) | 15.0 |

COS concentration (mg/m^{3}) | 15.0 |

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**MDPI and ACS Style**

Liu, K.; Chang, H.; Xiong, G.; He, J.; Liu, Q.; Li, J.
Modeling CO_{2}, H_{2}S, COS, and CH_{3}SH Simultaneous Removal Using Aqueous Sulfolane–MDEA Solution. *Processes* **2021**, *9*, 1954.
https://doi.org/10.3390/pr9111954

**AMA Style**

Liu K, Chang H, Xiong G, He J, Liu Q, Li J.
Modeling CO_{2}, H_{2}S, COS, and CH_{3}SH Simultaneous Removal Using Aqueous Sulfolane–MDEA Solution. *Processes*. 2021; 9(11):1954.
https://doi.org/10.3390/pr9111954

**Chicago/Turabian Style**

Liu, Ke, Honggang Chang, Gang Xiong, Jinlong He, Qisong Liu, and Jinjin Li.
2021. "Modeling CO_{2}, H_{2}S, COS, and CH_{3}SH Simultaneous Removal Using Aqueous Sulfolane–MDEA Solution" *Processes* 9, no. 11: 1954.
https://doi.org/10.3390/pr9111954