Exergy analysis of a conceptual CO₂ capture process with an amine-based DES

Abbreviations

C _p

heat capacity (kJ‧kg⁻¹‧K⁻¹)

E

exergy (kJ‧kg⁻¹)

G

sour gas flow rate (kmol‧h⁻¹)

H,H ₀

enthalpy and reference enthalpy (kJ‧kg⁻¹)

L

solvent circulation rate (kmol‧h⁻¹)

m

mass flow rate (kg‧s⁻¹)

M _w

molecular weight (g‧mol⁻¹)

mHE2

heating fluid flow rate in heat exchanger 2 (kg‧kg CO₂ ⁻¹)

mHE3

cooling fluid flow rate in heat exchanger 3 (kg‧kg CO₂ ⁻¹)

mHE4

cooling fluid flow rate in heat exchanger 4 (kg‧kg CO₂ ⁻¹)

mCon

coolant flow rate in stripper condenser (kg‧kg CO₂ ⁻¹)

mReS

steam flow rate in reboiler (kg‧kg CO₂ ⁻¹)

P

pressure (atm)

P ₀

reference pressure (atm)

P _r

reduced pressure

P _a

absorber pressure (atm)

Q

heat load (kJ‧s⁻¹)

R

ideal gas constant

S, S ₀

entropy, reference entropy (kJ‧kg⁻¹‧K⁻¹)

T ₀

reference temperature (K)

T

temperature (K)

T _a

absorber temperature (K)

T _c

critical temperature (K)

T _n

normal boiling temperature (K)

T _r

reduced temperature

T _s

stripper temperature for DES process (K)

THE2_i, THE2

inlet, outlet temperature of heating fluid in HE2 (K)

THE3_i, THE3

inlet, outlet temperature of heating fluid in HE3 (K)

THE4_i, THE4

inlet, outlet temperature of heating fluid in HE4 (K)

T _st

stripper top temperature (K)

T _sd

stripper distillate temperature (K)

T _re

stripper reboiler temperature (K)

T _si

stripper inlet temperature (K)

T _sb

stripper bottom temperature (K)

W

work (kJ‧s⁻¹)

x _f, x _o

CO₂ mole fraction in a lean and rich solvent, respectively

1 Introduction

Globally, energy-intensive activities are the primary sources of anthropogenic carbon dioxide (CO₂) emissions, with fossil fuel combustion in transportation, power production, and industry accounting for around 90% of total emissions [1]. After 3 years of stagnant or little global emissions between 2014 and 2016 [2,3], CO₂ emissions rose by 1.6% in 2017 to 36.2 Gt (billion tons) and grew a further 2.7% in 2018 (range: 1.8–3.7%) to a record of 37.1 ± 2 Gt CO₂ [4]. According to the International Energy Agency (IEA), due to the increasing energy demand, global CO₂ emissions from the energy-producing sector increased by 1.7% in 2018, reaching a historic high of 33.1 Gt CO₂, the fastest pace of rise since 2013. Although the emissions from all fossil fuels increased, the power sector alone contributed nearly two-thirds of the emissions growth. Specifically, coal-based power plants alone accounted for more than 10 Gt CO₂. CO₂ is the primary contributor to global warming impacts along with other greenhouse gases [5,6]. CO₂ capture and storage (CCS) technologies could be a potential solution to mitigate the emission of CO₂ into the atmosphere by up to 20%. However, the debate regarding CCS becomes intense regarding the cost of energy production because it increases by approximately 30–80% when CCS is applied [6,7]. Therefore, for cleaner energy production, it is important to use cost-effective technology.

The conventional CCS approach is CO₂ absorption using liquid absorbents and is considered the most mature, practically operational, and successfully demonstrated technology [8,9]. Considering the techno-economic aspect of the CO₂ chemisorption process, the selection of solvents is decisive. Solvents with fast kinetics of CO₂ absorption, high CO₂ capture capacity, low vapor pressure, and high stability against thermal and oxidative degradation should be ideal for CO₂ capture [10]. Among the liquid solvents, the aqueous alkanolamine solutions such as monoethanolamine (MEA), diethanolamine (DEA), methyl diethanolamine (MDEA), triethanolamine, 2-amino-2-methyl-1-propanol (AMP), and 2-methylaminoethanol were the most commonly used solvents over the last 60 years [11,12]. New amine-based solvents such as dimethylamino-2-methyl-1-propanol [13] and aqueous blends of conventional amines [14] were tested as well. The aqueous blends of MEA, MDEA, and AMP investigated by Chen et al. [14] resulted in a desorption rate up to five times higher than the commercial aqueous 5 M MEA solution, combined with a significantly lower relative heat duty for regeneration and lower operating cost. However, despite the high affinity of amine solvents for CO₂ molecules, some serious technological drawbacks exist when using the conventional aqueous amine scrubbing process. Specifically, the high-energy requirement for absorbent regeneration and the significant equipment corrosion that occurs makes the process challenging and expensive because of the high-temperature difference between CO₂ absorption (50°C) and desorption (120–140°C) units. Therefore, a high amount of energy is required to maintain the thermal difference in this system.

Currently, the critical limitation preventing the implementation of large-scale CCS globally on the major sources of CO₂ emissions is approximately 3.2–4.0 GJ‧ton CO₂ ⁻¹. Through extensive fundamental studies and several industrial demonstrations (e.g., Boundary Dam, Canada, and Petra Nova, USA) [15], using amine-based technologies achieved a great reduction in energy requirements, with 2.3–2.4 GJ‧ton CO₂ ⁻¹ removed compared with 3.5 GJ‧ton CO₂ ⁻¹ removed with first-generation amine scrubbing [16]. However, the current energy penalty level of CO₂ chemisorption is still unbearable if a full-scale CO₂ removal process is to be implemented for a coal-fired power plant. Researchers focus on exploring refined techniques or finding new and more efficient amine solutions that could reduce the energy penalty during the regeneration part of the process [17,18,19]. It has been reported that water might not be the optimal solvent to use due to high heat capacity, the heat of vaporization, and vapor pressure [20], inclusive of high-energy requirements during the regeneration process [21]. Furthermore, the detrimental ability of water to induce the degradation/corrosion of the amine and process equipment is also a concern [22].

Nonaqueous solutions were proposed as possible alternatives, replacing water with room-temperature ionic liquids (ILs), CO₂-binding liquids, or organic solvents such as methanol or ethanol to reduce the energy penalty during the regeneration step [23,24,25,26,27]. The emergence of binary or ternary blended solvents of amines [28,29] or amines with organics [30], such as alcohols, ethers, and glycols, allowed for the design of advanced nonaqueous absorbents. Because of their lower specific heat than water and low desorption temperature, organic solvents could provide significant advantages for saving the regeneration energy and reducing corrosiveness and degradation. In a nonaqueous process using MEA/methanol as an absorbent, the removal of methanol below 373 K for solvent regeneration resulted in enormous energy savings [31]. A mixture of a nonaqueous absorbent of piperazine (PZ) with diethylene glycol to capture CO₂ in a rotating packed bed has been proposed elsewhere [32]. The results showed that using glycols as solvents significantly reduced energy consumption with almost negligible solvent evaporation and avoidance of thermal degradation. Another study demonstrated secondary alkanolamines/EG as cost-effective absorbents for rapid and reversible CO₂ capture [33]. Chen et al. studied the solubility of CO₂ in N-ethylmonoethanolamine solutions using the vapor–liquid equilibrium (VLE) and the absorption–desorption apparatus. Their results showed that the tertiary amine, N,N-diethylethanolamine, used as a nonaqueous solvent demonstrated more excellent performance than alcohols and glycols [34]. Recently, nonaqueous binary absorbents of conventional amines (MEA and MDEA) with low-viscosity ILs were investigated for CO₂ capture. The mixed absorbent MDEA + [BEIM]BF4 with low viscosity demonstrated a high recycling CO₂ capacity, energy-saving capacity, and good renewability [35]. The absorption–desorption performance of CO₂ into several blends of MEA or DEA with glycol ethers (2-methoxy ethanol (2ME) and 2-ethoxy ethanol (2EE)) as nonaqueous solvents has been reported at ambient pressure. It was found that the mixture of MEA and 2ME or 2EE demonstrated a significant reduction in energy consumption (55%) compared with the benchmark aqueous 5.0 M MEA system.

ILs and deep eutectic solvents (DESs) are among the most attractive substitutes that have been examined over the past few years because of their favorable properties. These properties include negligible vapor pressure, high thermal stability, nonflammability, and tuneability for the desired separation process by changing the combination of cations and anions [36]. For CO₂ capture, various ILs containing amine-functionalized cations/anions [35,37], amino acids [38], aprotic heterocyclic anions, super bases [39], and pyridinium anions [40] have been investigated. Nevertheless, the main problems associated with ILs regarding their application for CO₂ capture are their high susceptibility to contamination, poor biodegradability, high toxicity [8], increasing viscosity during CO₂ capture, and high cost for large-scale industrial application [41].

To overcome the problems associated with ILs while retaining their advantages, DESs appeared as a credible alternative [42,43]. DESs are sustainable solvents analogous to ILs but with several advantages over ILs such as biodegradability, high thermal and chemical stabilities, nonflammability, a wide liquid range, and high purity. Furthermore, DESs are less expensive and easier to synthesize by combining an organic halide salt (considered the hydrogen bond acceptor) with another hydrogen bond donor, such as an amine, amide, carboxylic acid, and alcohol in a certain ratio, resulting in the formation of a low melting-temperature solution by using hydrogen bonding interactions [44]. However, despite the significant efforts of researchers in developing novel solvents with outstanding performance regarding CO₂ absorption capacity and low-energy requirements for solvent regeneration, the implementation of these advanced solvents has not yet reached the commercial level, and more efforts are required to test their practical and economic viability on a large scale.

Furthermore, rigorous rate-based models are indispensable for the design, scale-up, optimization, and analysis of the CO₂ absorption/desorption processes inside packed columns [45]. Many researchers have used the ASPEN Plus simulator to analyze and optimize the CO₂ capture process [12,46,47,48]. Abu-Zahra et al. implemented the ASPEN PlusRadFrac subroutine model to simulate and optimize the energy required for lean solvent regeneration during the CO₂ capture process from a bituminous coal-fired plant. When using a 40 wt% MEA solution and a stripper operating pressure of 210 kPa, the obtained value was 3.0 GJ‧ton CO₂ ⁻¹, which is 23% lower than the base case of 3.9 GJ‧ton CO₂ ⁻¹ with MEA [46]. Li et al. recently used ASPEN Plus to develop a rate-based model for CO₂ capture by using the aqueous MEA absorption process from flue gas [47]. The model was systematically studied for process modification (absorber intercooling, rich-split, and stripper interheating) and the optimization of various parameters (MEA concentration, lean CO₂ loading, lean temperature, and stripper pressure). The minimum regeneration energy obtained from the model was 3.1 MJ‧kg CO₂ ⁻¹, whereas the highest reported value for the reboiler duty is 5.8 MJ‧kg CO₂ ⁻¹ [49]. Oh et al. [50] conducted a recent study based on minimizing the energy expenditure of the CO₂ capture process using aqueous MEA. The authors used the UniSim^® process simulator to perform four different types of structural process modifications, including flue gas splitting, multiple solvent feeding, and split-stream/semi-lean and absorber inter-cooling. They inferred that the minimum energy costs could be attained by applying a combination of all four types of structural modifications. However, significant reductions in energy cost could be achieved by implementing single-process modifications, specifically by splitting the flue gas and feeding it at two different locations in the column.

It is believed that the second-generation amine system performs better in energy savings than the traditional MEA chemical absorption process. Zhang et al. presented a systemic study on the modeling and process analysis of the CO₂ capture process with an AMP + PZ aqueous solvent [51]. Their results revealed that by lowering the CO₂ removal rate, increasing the ratio of AMP in the solvent, and increasing the stripper pressure, the reboiler duty could be reduced. The authors applied various process configuration modifications, including intercooled absorber (ICA), lean vapor compressor (LVC), and rich solvent split (RSS), reducing the energy consumption by 6.7%, 2.7%, and 8.5%, respectively. However, the combination of ICA + RSS + LVC decreased the energy demand by 15.2%. Another group of researchers performed a scale experiment via the process development unit for CO₂ removal from a gas mixture by applying MEA and AMP promoted with PZ. Their findings revealed that using AMP/PZ enables the reduction of the solvent heat duty.

Our group investigated the solubility of CO₂ in different DESs at a moderate pressure of 10 atm [52]. The study showed promising results and promoted using methyltriphenylphosphonium bromide (MTPB)-MEA DES for CO₂ capture from power plant exhaust gases. The model developed based on these results could determine the required solvent circulation rates, CO₂ loadings, and flashing temperatures for specific CO₂ recovery rates and given flue gas conditions [53]. The main objective of this work is to quantify the energy requirements of a DES-based CO₂ capture process and determine the source of the energy losses by performing an exergy analysis. The results are benchmarked against the well-established MEA process.

2 Process description and analysis

A schematic flow sheet of the CO₂ capture process is shown in Figure 1. The process is a combination of an absorber and a stripper column. In the absorber, a lean sorbent is used to remove the CO₂ associated with the sour gas in a 40-stage column. The spent sorbent is regenerated in the stripper, where the exact amount of captured CO₂ is released with the top product. The regenerated solvent is recycled back into the absorber. In this work, the stripper is treated as an atmospheric flash drum where the amount of CO₂ to be released occurs at specific VLE conditions. The baseline operating conditions for this process are listed in Table 1. These data are selected based on our previous work [53] and are in the range used in literature [49]. We assumed that only CO₂ is miscible in the solvent (MTPB-MEA DES), the liquid species transferred to the gas phase are insignificant, the CO₂ absorption is only because of physical interactions, and no chemical reactions occur.

Figure 1

Schematic of CO₂ capture process using DES.

Indeed, as shown in our previous work [52], samples of the DES before and after the absorption of the CO₂ were analyzed using high-performance liquid chromatography (HPLC) and the results showed that only 10% of the amine reacted with the CO₂, while for the aqueous solution of MEA, all of the amines reacted with the CO₂. On the other hand, the detailed design equations and solution methodology of the process are given in our previous work [53]. The numerical solution for the baseline conditions is shown in Figure 1.

In the classical MEA process, the reboiler temperature is limited to 120°C because of the amine thermal instability above this temperature [54]. However, to evaluate the thermal stability of the DES used in this work, we performed a thermogravimetric analysis (TGA) under a nitrogen atmosphere using a Mettler Toledo TGA/DSC instrument. The results obtained are provided as supporting material (Figure A1 in Appendix). This analysis indicates two decomposition temperatures: one corresponding to the decomposition of MEA (at about 150°C) and the other corresponding to the decomposition of MTPB at a much higher temperature (at about 400°C). Therefore, the CO₂ capturing process that runs with this solvent could operate at a higher temperature than the classical MEA aqueous process.

Table 2 lists the process performance at different selected operating pressures for the absorber column. The required liquid sorbent, the lean and rich solvent loadings, and the necessary stripping temperature are estimated at each operating pressure. The results in Table 2 form the baseline for the intended exergy analysis. Figure 2 illustrates the process performance with variable absorber pressures. The required solvent amount increases dramatically at low absorber vessel pressures because, at low pressure, the solubility of CO₂ in the DES diminishes. Therefore, an additional quantity of solvent is needed to achieve an exact recovery rate of 90%. Similarly, Figure 2 demonstrates how the stripping temperature necessary for separating the captured CO₂ from the spent sorbent lowers with pressure. The stripping temperature increases because of the amplification in the solvent circulation rate. Consequently, the rich loading decreases, and a much deeper lean is necessary, demanding additional energy to release the captured CO₂.

Figure 2

CO₂ capture process performance at selected values for P _a using DES: squares for 90% recovery and circles for varied recovery.

Our simulation indicates that 90% of the CO₂ recovery cannot be achieved by using DES at an absorber pressure of less than 2.2 atm. Furthermore, the simulation results revealed that the process operation below 7 atm is not physically acceptable because the required flashing temperature is beyond the boiling point of the applied DES. Therefore, the simulation was repeated, incorporating a flashing temperature constraint. Table 3 and Figure 2b show the results. The simulation results when operating at a low absorber pressure and a flashing temperature less than the DES boiling point of 155°C are shown in Table 3. At P _a = 5 atm, the recovery must be lowered to 80% or less to ensure a low stripping temperature. Similarly, at P _a = 3 atm, 60% is the maximum allowable recovery that guarantees a stripping temperature below 155°C.

Furthermore, it is found that operating the absorber vessel at a pressure lower than 3 atm does not allow for a flashing temperature less than the boiling point, even at exceptionally low recovery rates.

An absorber column operating at 3 atm is not economically appealing from an energy consumption standpoint because it requires a high stripping temperature. Operating at this low pressure (less than 7 atm) could result in additional costs, such as the cost of materials and pumping. The results in Tables 2 and 3 indicate that the solvent flow rate could reach up to 13 times that required at P _a = 10 atm. Furthermore, the process is operating at a lower recovery rate. Notably, although the process efficiency at low pressure is unappealing, it will be included in the following exergy analysis for demonstration purposes.

3 Exergy principles

Exergy is defined as the maximum possible reversible work that can be obtained from a system or a process. It is useful to measure and compare the energy efficiency of different processes or a process with different configurations [55,56]. Exergy analysis could identify areas of improvement and indicates the priorities toward efficiency enhancement and optimization, which are essential for the creation of new engineering decisions toward sustainable development.

The general exergy (Ex) expression is given by Eq. 1:

were H ₀ and S ₀ represent the enthalpy and entropy values, respectively, at the environmental temperature T ₀, which is usually equal to 298 K. H and S represent the enthalpy and entropy, respectively, at a specified temperature, T (K).

All the processes involve the conversion and consumption of exergy; thus, high efficiency is critical, implying that the exergy use is managed and effective tools are applied. Although the total energy is always conserved, real process exergy is never balanced because of being irreversible, i.e., exergy destruction always occurs. The exergy destruction, ΔE, indicates possible process improvements. In engineering, different flow diagrams are often used to describe the different exergy flows through a process, enabling the selection of the process presenting the least exergy destruction. Generally, the exergy destruction or loss can be written as follows:

where m _i and m ₀ represent the mass input and output to the system, Q is the heat flow to the system, and W is the workflow to the system.

If the system is adiabatic (Q = 0) and does not require any work (W = 0), then Eq. 2 becomes Eq. 3:

The standard conditions considered in this work are T ₀ = 298 K and P ₀ = 1 atm.

4 Enthalpy and entropy change calculations

The estimation of exergy requires calculating the enthalpy and entropy of the fluids involved in the process. For the fluids used in the system, such as water and the liquid solvent, the enthalpy and entropy are functions of the temperature only, which could be calculated using the following equations for constant heat capacity:

For compressible gases at the elevated pressure, the enthalpy can be calculated as follows:

where the residual property can be calculated from the following correlation [49]:

Similarly, the gas entropy can be found using Eq. 8:

Eqs. 6–9 allow the calculation of enthalpy and entropy changes for any temperature and pressure variation.

The energy required for fluid pumping is calculated using Eq. 10:

In Eq. 10, Q is the volumetric flow rate in cubic meters per second, ΔP is the pressure difference in pascals, and η _p is the pump efficiency. Similarly, the energy required for gas compression can be calculated using Eq. 11:

where m is the mass flow rate of the gas, H is the specific enthalpy for the inlet and outlet streams, and η _c is the compressor efficiency.

5 Results and discussion

5.2 Exergy analysis of DES-based CO₂ capture process

After determining all flow rates and operating parameters with DES as the solvent (Table A2), we performed an exergy analysis for the overall process denoted by the dotted boundary in Figure 1. In this case, the exergy analysis included all streams that cross the boundary (streams 1–4). Therefore, Eq. 2 incorporates the contribution of heat exchangers and compressors to overall exergy loss in which case Table 4 and Figure 3 show the results.

Table 4

Exergy flow and overall exergy loss of CO₂ capture process

Solvent	P a	T s	Exergy flow in streams and units										Overall
			S1	S2	S3	S4	HE2	HE3	HE4	C01	Reb	Cond	Exin	Exout	∆EX
	(atm)	(°C)	(MW)	(MW)	(MW)	(MW)	(MW)	(MW)	(MW)	(MW)	(MW)	(MW)	(MW)	(MW)	(MW)
DES	3	147.8	0.00	8.82	0.26	0.00	6.70	−0.38	−0.35	10.64	—	—	17.33	9.80	7.53
	5	125.2	0.00	12.78	0.15	0.00	1.74	−0.14	−0.55	16.77	—	—	18.51	13.62	4.89
	7	125.6	0.00	15.36	0.18	0.00	0.75	−0.08	−0.70	21.28	—	—	22.03	16.32	5.71
	10	65.6	0.00	18.17	0.02	0.00	0.23	−0.04	−0.87	26.51	—	—	26.75	19.11	7.64
	10 + ERD	65.6	0.00	0.00	0.02	0.00	0.23	−0.04	−0.88	5.25	—	—	5.48	0.94	4.54
MEA	1	40	0.04	0.04	0.018	4 × 10−5	—	−0.211	—	—	6.26	−0.64	6.5	0.8	5.7

Figure 3

Exergy destruction of the CO₂ capture process using DES at selected pressure values.

Figure 3 shows that the system’s exergy destruction to exergy input decreases with operating pressure but increases again at P _a = 10 atm. To achieve a high recovery rate and the lowest exergy losses, the absorber should be operated at P _a = 7 atm. Examining the precise exergy flow in Table 4 can help you understand the exergy trend with pressure. Because the stripping temperature reaches 148°C at P _a = 3 atm, the thermal energy consumption for heating the rich sorbent in HE2 is the highest. As a result, the equivalent gross exergy loss is the greatest. However, for P _a = 10 atm, the work for compressing the flue gas takes precedence over the stripping energy. The additional compression work raises the net exergy destruction once again. As a result, the compression and heat energies must be balanced. As previously stated, running at low absorber pressures is undesirable due to the increased solvent circulation rate and low CO₂ removal efficiency. Table 4 depicts the exergy flow linked with the process streams. Because of its high pressure, stream 2, the sweetened gas stream, has the largest exergy flow of the other streams. As the pressure decreases, so does the exergy.

It is useful to exploit the large exergy leaving the system by stream 2. The energy consumption when the process is operating at 10 atm (or higher) could be reduced by using the exergy associated with stream 2. This could be achieved by using a dual-work exchanger energy recovery system. These low-cost energy recovery devices (ERDs) provide greater energy-saving payback and greater operational flexibility for liquids [57,58].

A similar idea is utilized for gases, consisting of a gas turbine (expander) followed by a gas compressor. The power created by the gas turbine is utilized to power the following compressor, which pressurizes another low-pressure gas stream in this device. This device is also known as the gas pressure exchanger or booster, adopting the fundamentals of turbochargers in combustion engines. The so-called ERD for gases is proven useful in industry, for example, in ammonia factories. The purpose of this equipment is to recover energy from the pressurized fluid stream leaving the absorber column. The modified process is shown in Figure 4. Usually, such devices are limited by thermodynamic efficiency; therefore, the pressure of the exiting stream would be less than the inlet pressure of the pressurized stream. For an ERD efficiency of 60%, the pressure of the gas stream exiting the ERD can be 6.3 atm and the corresponding temperature 217°C. In this case, a compressor is still needed to complement the ERD. The main advantage of the modified process is that the purified gas stream exits the process at 1 atm. Consequently, the calculated overall exergy of the modified process is 6.36 MW.

Figure 4

Schematic of CO₂ capture process using DES with the aid of ERD.

By using the pressure of the exit gas stream with an ERD efficiency of 60%, the process changed from an exergy deficit of 7.64 to 6.36 MW, which is equivalent to a 17% enhancement because the energy recovery in the form of the pressure head within the whole system enhances the conservation of exergy (decrease in exergy loss). If the ERD efficiency can be increased further, the performance of the DES process regarding exergy losses could be enhanced substantially (Table 5). This situation makes the DES-driven CO₂ capture process, supported by the energy recovery system, more energy efficient. Usually, the sour gas treated for CO₂ removal is the exhaust gas of a power plant. The flue gas leaves the combustion chamber of a power plant at an elevated temperature and pressure (1,494°C and 30 bar, respectively) [59,60]. The immense potential energy of the flue gas is usually used in power plants to produce high-pressure steam. Therefore, the exergy loss due to the compression work can be avoided by integrating the DES-driven CO₂ capture process with the power plant.

Table 5

Performance of DES + ERD at 10 atm and different ERD efficiency

Efficiency (%)	∆EX (MW)	En (MJ‧kg CO2 −1)
50	8.370	3.041
60	6.357	2.060
70	4.542	1.400
80	2.884	0.935

6 Conclusions

The energy demand of CO₂ removal from flue gas using amine-based DESs in the absorption–desorption process was investigated in this work. High absorption column pressures are required to provide effective CO₂ physical absorption and, as a result, improved process performance and design requirements. Depending on the processing capacity and operating pressure, the raised pressure increases the exergy deficit to 4–8 MW and the energy consumption to roughly 6–38 MJ‧kg CO₂ ⁻¹. We discovered that the main source of exergy shortage is the compression energy of the flue gas. However, when compared to the MEA-based process, it was found that the DES-based process has a lot better exergy destruction because the exergy inserted is much superior to the exergy lost in the DES process. The energy demand of the MEA process, on the other hand, was discovered to be nearly 50% lower than that of the DES process, where the energy consumption is mostly related to the compression work to the required high pressure. Adding an ERD unit can lower the DES process’s energy need by up to 80%. To improve the process economically, we recommend lowering the operating pressure, which reduces exergy loss at the expense of performance. Alternatively, ERDs could be incorporated into the process to reduce compression demand and, as a result, exergy loss. Nonetheless, when applied to contaminated exhaust gases at high pressure, the DES-driven CO₂ capture process could be energy efficient. The study’s major goal is to identify a less energy-intensive CO₂ capture technique than the standard MEA solvent. DES solvents are used to alleviate this. However, the suggested DES process necessitates high-pressure operation to maintain the same CO₂ recovery and solvent-to-gas ratios as the standard MEA method, resulting in increased energy consumption owing to flue gas compression. Yet, by employing an ERD, the DES process can outperform the conventional technique. These encouraging findings pave the way for additional research into alternatives to standard MEA solvents.