Enhancing CO₂ Capture Efficiency: Advanced Modifications of Solvent-Based Absorption Process—Pilot Plant Insights

Abstract

Since fossil fuels still dominate industry and electricity production, post-combustion carbon capture remains essential for decarbonizing these sectors. The most advanced technique for widespread application, particularly in hard-to-abate industries, is amine-based absorption. However, increasing energy efficiency is crucial for broader implementation. This study presents pilot-scale results from the Tauron Power Plant in Poland using a mobile CO₂ capture unit (1 TPD). Two innovative process modifications—Split Flow (SF) and Heat Integrated Stripper (HIS)—were experimentally investigated; they achieved a 10% reduction in reboiler heat duty, reaching 2.82 MJ/kg_CO2, along with a 36% decrease in overall heat losses and up to a 28% reduction in cross-flow heat exchanger duty. The analysis highlights both the advantages and challenges of these modifications. SF is easier to retrofit into existing plants, whereas the HIS requires more extensive modifications in the stripper section, thus making HIS more cost-effective for new installations. Moreover, as heat consumption constitutes the primary operational cost, even a moderate reduction in heat duty can lead to significant economic benefits. The HIS also offers substantial potential for thermal integration in industries with available waste heat streams. The pilot data underwent validation procedures to ensure reliability, which provides a robust foundation for process modeling, optimization, and scaling for industrial applications.

1. Introduction

Decarbonizing industry presents a major challenge, particularly in sectors where emissions are difficult to avoid, including cement manufacturing, metallurgy, fertilizer production, and refining. These industries, along with power generation, still rely heavily on fossil fuels, making post-combustion carbon capture (PCCC) a crucial technology for reducing CO₂ emissions [1,2,3]. Among the available methods, amine-based chemical absorption remains the most advanced and widely applicable, particularly for hard-to-abate sectors [4,5]. While initially developed for power plants, this approach has also proven effective in capturing CO₂ from biogenic sources (BECCS); it is currently the only available negative emissions technology [6,7]. Given the similar CO₂ concentrations found in industrial and power sector flue gases, amine-based absorption offers a scalable and versatile solution for large-scale emissions reduction. However, to enable widespread deployment across both industries, improving energy efficiency remains a key challenge that must be addressed [1,8].

One of the main challenges of amine-based CO₂ capture is the high energy demand associated with solvent regeneration, which typically requires at least 3.6–3.8 MJ/kg_CO2 in conventional systems [2]. Current research efforts aim to lower this energy consumption to below 2.0 MJ/kg_CO2 through advanced process modifications and novel solvent formulations [4]. Several modifications have been proposed to enhance process efficiency; however, those related to stripper modifications appear to be the least explored. In this context, Cold Solvent Split (CSS) and Stripper Interheating (SI) are the most commonly considered, while Split Flow (SF) and Heat Integrated Stripper (HIS) are investigated much less frequently [8,9,10].

The CSS, also referred to as Rich Split (RS), involves diverting a portion of the rich solvent stream and feeding it to the upper section of the stripper without prior heating. This configuration was recognized by Feron et al. in 2020 [11] as a new standard in amine-based CO₂ capture processes.

The SF configuration represents a more advanced process layout involving the separation of both lean and rich amine solvent streams. The key feature of this configuration lies in the strategic division of solvent flows for optimizing energy usage. Such a modification was initially proposed to reduce steam demand in the reboiler, which offered a more energy-efficient alternative to the conventional process arrangement. This concept was originally introduced in the early 20th century by Shoeld [12] for the removal of acid gases, such as H₂S, from high-pressure natural gas systems. Shoeld’s idea has since been further developed by several researchers [13,14,15]. Later studies confirmed its applicability for CO₂ capture from flue gases [16]. However, despite its potential, there are relatively few pilot-scale studies available that explore the practical implementation of such a configuration.

On the other hand, SI enhances heat distribution by extracting, heating, and reinjecting solvent at specific stripper sections while often utilizing the lean solvent as a heating medium [17,18,19]. Other potential heat sources include steam condensates, hot flue gases, or external low-temperature heat sources [16,20,21]. While SI has been extensively studied in simulations by Li et al. [22], Zhao et al. [23], and Muhammad et al. [24], experimental validation has been limited to studies by Stoffregen et al. [21] and Iijima et al. [17], who reported its application in MHI technology with the KS-1 solvent without detailed analysis. Recently, Danish researchers addressed this gap with pilot-scale experiments on Stripper Interheating at the Amager Bakke Waste-to-Energy plant in Copenhagen, utilizing a steam generator to facilitate the SI system, as presented by Vinjarapu et al. [25].

While other innovative stripper modifications, such as the Piperazine Advanced Stripper (PZAS) developed by the University of Texas at Austin [26] and the divided stripper configuration employed by the CAER 0.7 MWe pilot plant [27], have demonstrated promising results, challenges such as oxidative degradation of solvents persist. These studies highlight the need for continued exploration of alternative heat integration strategies. One such approach is the Heat Integrated Stripper (HIS), which was proposed by Leites et al. [28] and later extensively studied through modeling research by Oyenekan [29]. The HIS involves deep thermal integration by embedding heat exchangers along the significant height of the stripper, which are supplied with a lean solvent. As noted in reviews by Feron [8], Cousin et al. [16], and more recently by Salimi et al. [30] and Løge et al. [31], this concept has not yet been validated at the pilot scale.

A form of this solution has been implemented by the ITPE (Institute of Energy and Fuel Processing Technology) team in the Tauron pilot plant. Preliminary research findings were presented in [32,33], including an initial analysis of HIS efficiency and an economic evaluation based on pilot-scale experiments [34], which demonstrated its cost-effectiveness for industrial applications.

In this study, a combination of Split Flow (SF) and the Heat Integrated Stripper (HIS) is investigated using a mobile pilot plant with a 1 TPD CO₂ capture capacity (Figure 1). The experiments, conducted at the Tauron Power Plant in Poland, utilize a well-documented and widely tested MEA 30 wt.% solvent, making it a strong reference for evaluating advanced process configurations. The results for these modifications will be compared to baseline data from a previous study [35]. The findings demonstrate a significant reduction in reboiler heat duty, with energy consumption falling below 2.90 MJ/kg_CO2. Additionally, the HIS shows strong potential for integration with external waste heat sources, particularly in cement plants, steelworks, and waste incineration facilities. Beyond heat reduction benefits, the HIS could also facilitate CO₂ capture integration with exothermic catalytic processes, such as methanation, using green hydrogen [36,37].

To accurately assess the impact of different modifications, comprehensive thermal balancing [38] and heat loss evaluations were performed [35]. These analyses are critical for pilot plants, where heat losses can constitute a significant portion of the energy balance, whereas, in industrial-scale applications, these losses represent a much smaller share. Additionally, the study examines the feasibility of combining multiple process modifications to maximize efficiency improvements.

Despite extensive simulation studies, the practical deployment of advanced stripper configurations remains underexplored. By providing pilot-scale validation, this study contributes to the optimization and industrial deployment of post-combustion CO₂ capture, reinforcing its role in decarbonization.

2. Materials and Methods

2.1. The Pilot Plant

The CO₂ capture unit investigated in this study builds upon extensive prior research conducted at Tauron Group power plants. While its design and operation parameters have been extensively described in previous studies [32,33], this paper specifically examines the impact of advanced process configurations, such as SF and the HIS, on energy efficiency and CO₂ capture performance.

Developed as a mobile unit, the pilot plant enables CO₂ capture at a rate of 1 ton per day via chemical absorption under different industrial conditions. It has been deployed three times in total, with key research conducted at the Jaworzno Power Plant [35], where it was connected to the flue gas stream from the hard coal-fired, fluidized bed boiler operating under normal industrial conditions. The present study includes additional results obtained during these campaigns.

Earlier studies at the Łaziska Power Plant played a crucial role in validating its performance and refining measurement methodologies [33]. Ultimately, after 2500 h of research, the plant was integrated into a research system for energy storage in the form of synthetic natural gas (SNG), demonstrating that chemical absorption can provide a CO₂ stream with parameters suitable for catalytic hydrogenation syntheses [36,37].

The modular pilot plant, typically handling flue gas streams of approximately 200 m³/h sourced directly from an industrial site, was equipped with a dedicated deep desulfurization section. During its deployment at the coal-fired power plant, this module played a critical role in flue gas purification prior to CO₂ capture studies. Before entering the desulfurization unit, the flue gas was first subjected to a washing and cooling stage, reducing its temperature to approximately 40 °C. The core process equipment was housed in the Technology Container while monitoring and analytical operations were conducted in the Supervision Container. The complete layout with the Storage Container is illustrated in Figure 1. The pilot plant was designed as a flexible testing platform, allowing for the evaluation of multiple process configurations during individual test campaigns, each typically lasting around 100 h. The most relevant of these configurations are discussed in detail in the present study.

2.2. Main Apparatus

Analyzing temperature profiles within the column is crucial when comparing different process configurations, as the tested SF and HIS setups affect both the absorber and the stripper.

Table 1. Characteristics of the sections in the columns of the pilot plant.

	Absorber			Stripper
Section No	Section Content	Packing Dimensions H × D [mm × mm]	Section No	Section Content	Packing Dimensions H × D [mm × mm]
S10	Sulzer Mellapak 750Y	810 × 273	S13	Sulzer Mellapak 750Y	830 × 273
S07	Sulzer Mellapak 500Y	1620 × 273	S11	Sulzer Mellapak 750Y	1040 × 273
S06	Sulzer Mellapak 500Y	1620 × 273	S09	Sulzer Mellapak 750Y	1050 × 273
S04	Sulzer Mellapak 350Y	2160 × 273	S08	Upper recuperator (REC-U) with Sulzer C ring 1”	2570 × 508
S03	Sulzer Mellapak 350Y	2160 × 273	S06	Lower recuperator (REC-L2) with Sulzer C ring 1”	1570 × 508
S02	Sulzer Mellapak 350Y	810 × 273	S04	Lower recuperator (REC-L1) with Sulzer C ring 1”	3050 × 508
			S02	Sulzer Mellapak 750Y	1370 × 323
			S01	Reboiler with el. heater, 63 kWel

presents the key data for these columns. Thermocouples were installed between individual packed sections and their positions corresponded to those shown in Figure 2.

The K-215 stripper was constructed from stainless steel and designed with 15 sections. Seven of these were dedicated to gas and liquid distribution, while the remaining sections, detailed in Table 1, were responsible for mass transfer and thermal integration.

The primary heat source was a reboiler situated in section S01, powered by a four-core electric heater manufactured by Selfa (Szczecin, Poland). To ensure precise process monitoring, six pressure measurement points were distributed along the stripper’s height, primarily for assessing packing performance and pressure variations. A pressure sensor and control valve (as shown in Figure A1 in Appendix A) were installed downstream of the F-223 separator to regulate system pressure.

To minimize thermal losses, the entire column was insulated with 80 mm thick mineral wool and enclosed in aluminum sheeting. Given the mobile nature of the plant, the stripper was designed for partial disassembly. Although the insulation was non-uniform at structural connection points, this had no significant impact on overall heat retention.

A key innovation in the stripper (total height of 15.3 m) was the incorporation of diaphragm-type, tube-in-tube heat exchangers, commonly referred to as recuperators. These exchangers harnessed the heat from the regenerated solvent, improving the temperature profile within the central part of the column and enhancing CO₂ desorption efficiency. To further optimize mass transfer, a 1” Sulzer C-ring packing was installed inside the recuperator tubes, while the remaining sections employed structured Sulzer Mellapak 750Y packing (Winterthur, Switzerland).

The system featured a dual-recuperator setup. The upper recuperator was positioned in section S08, while the lower recuperator was split into two parts, located in sections S04 and S06. This configuration enabled the implementation of the Split Flow (SF) modification by allowing the extraction of semi-lean solvent from section S05.

The technical data on the remaining apparatus and equipment were already presented in a previous publication [35], and they were not modified during the experiments described in this study.

2.3. Andvanced Proces Configurations

The process setup of the pilot plant was based on the conventional amine scrubbing system, which has been extensively described in the literature [39,40,41]. While grounded in this well-established framework, the installation was designed to offer considerable flexibility, enabling the implementation of several advanced process configurations. These included Double Absorber Feed (DAF), Cold Solvent Split (CSS), and the previously mentioned Heat Integrated Stripper (HIS) and Split Flow (SF).

Most of these configurations primarily differed in terms of the solvent dosing point within the column. However, the HIS configuration required modifications to the stripper by incorporating internal heat exchangers (recuperators), while the SF configuration involved splitting the lean solvent stream, necessitating the duplication of certain equipment such as pumps, heat exchangers, and coolers. Consequently, the process flowsheet presented in Figure 3 appears more complex than the conventional solvent-based CO₂ capture process.

2.3.1. Standard Configuration with Cold Solvent Split (CSS)

In comparison to the conventional amine-based system, the applied setup included a process modification in which the top section of the stripper was supplied with a colder portion of the rich solvent (L₃₁), typically accounting for about 8% of the total rich solvent stream leaving the absorber. This adjustment constitutes the core of this process modification. It is now widely recognized as a standard feature in CO₂ capture systems and is commonly implemented [11,23,42,43].

The remaining portion of the rich solvent stream was split into two comparable parts and routed through cross-flow heat exchangers, where it exchanged heat with the lean solvent. These streams were subsequently introduced into the stripper as L₁₂ above section S09 and L₂₃ above section S08. This arrangement helped to partially cool the upper part of the stripper, thereby reducing water losses through the gas outlet stream.

During the presented campaign, this configuration was applied in all experiments and was not analyzed independently. For this reason, it is treated in this study as part of the standard system configuration (Figure 3).

Figure 3. The process flowsheet in the standard configuration of the TAURON pilot CO₂ capture plant.

Figure 3. The process flowsheet in the standard configuration of the TAURON pilot CO₂ capture plant.

2.3.2. Split Flow (SF)

This process modification involves drawing two solvent streams with different levels of CO₂ loading from the stripper and directing them back into the absorber at two distinct points (Figure 4). The first stream, deep regenerated (deep-lean) (L₄) and withdrawn from the bottom of the stripper, was introduced above the section (S07) of the absorber. The second, partially regenerated (semi-lean) (L₅) stream was extracted from an intermediate section of the stripper (S05) and fed into the middle part of the absorber above section S04.

This solution aimed to cool the middle section of the absorber and perform preliminary gas purification in its lower sections while enhancing the driving force for absorption in the upper stages. This approach allowed the system to maintain the required CO₂ removal efficiency without increasing the reboiler heat duty, as only a portion of the circulating solvent was subjected to deep regeneration in the stripper. Such a strategy is considered particularly useful for absorbents with a low reaction rate, such as tertiary amines [20].

2.3.3. Split Flow with Heat Integrated Stripper (SF-HIS)

In this configuration, diaphragm-based heat exchangers (recuperators) were integrated into the stripper, following a modified concept originally proposed by Leites [28]. In conventional strippers, an excessive driving force for desorption tends to concentrate in the reboiler and the top sections of the column. Leites suggested that optimal desorption performance can be achieved by distributing the driving force more evenly across the entire packed sections.

In the pilot plant, the lower recuperator (REC-L) consisted of two parts located in sections 4 and 6, separated by a distributor tray (section S05) that collected stream L₅₁ from the stripper. The REC-L facilitated heat recovery (without mass exchange) from the lean solvent stream L₄₁, which was drawn from the bottom of the stripper. The recovered heat was transferred then to the solvent within the recuperator’s internal tubes, intensifying desorption as the solvent flowed downward through the packing.

The upper recuperator (REC-U), located in section S08, operated similarly by recovering heat from the L₅₁ stream (collected in section S05) and transferring it to the solvent within the upper part of the desorption zone.

After passing through the recuperators, the partially cooled solvent streams (L₄₂ and L₅₂) were directed to cross-flow heat exchangers and subsequently cooled in solvent coolers before being returned to the absorber (Figure 5).

2.4. Main Proces Indicators

2.4.1. CO₂ Capture Efficiency

This parameter is a key performance indicator in solvent-based carbon capture processes, representing the proportion of CO₂ removed from the treated gas stream. It directly reflects the effectiveness of the absorption system and is crucial for assessing the viability of different process configurations, solvents, and operating conditions. A high η is desirable for maximizing CO₂ removal while minimizing solvent consumption and energy demand.

In this study, η was determined based on the mass flow rates of CO₂ in the inlet gas stream (G₂₁) and the outlet gas stream (G₂₂), both expressed in kg/h. The CO₂ capture efficiency was calculated using Equation (1):

$$\eta ={\displaystyle \frac{G_{21 C{O}_2}-G_{22 C{O}_2}}{G_{21 C{O}_2}}}·100\%$$

(1)

2.4.2. Reboiler Heat Duty

In solvent-based CO₂ capture, most of the energy consumption is associated with solvent regeneration, quantified as reboiler heat duty (q) in MJ/kg_CO2. This parameter is critical for assessing process efficiency, comparing solvent performance, and evaluating optimization strategies. Reducing reboiler heat duty is fundamental to improving the economic feasibility of CO₂ capture technologies. The heat required for solvent regeneration Q_reg is supplied to the stripper’s steam-powered reboiler, and the total specific energy demand can be expressed as in Equation (2) [40,44,45,46].

$$q={\displaystyle \frac{Q_{reg}}{m_{C{O}_2}}}={\displaystyle \frac{Q_{des}+Q_{vap}+Q_{sens}}{m_{C{O}_2}}}=q_{des}+q_{vap}+q_{sens}$$

(2)

where

• Heat of Desorption (Q_des)—Represents the minimum energy required to break the chemical bonds between CO₂ and the amine solvent, enabling solvent regeneration. In CO₂ capture studies [47,48,49,50], it is often assumed that the heat of desorption equals the heat of absorption, although this depends on the solvent composition. In this study, the heat of absorption value was adopted based on Bruder et al. [51].
• Heat of Vaporization (Q_vap)—Energy associated with solvent evaporation, primarily water. Some water exits the system with the desorbed CO₂, and the heat required for its vaporization is lost. The amount of water leaving the stripper in this manner depends on process parameters, including stripper operating pressure and temperature, solvent composition, and CO₂ loading. Higher desorption temperatures and lower solvent CO₂ loadings generally increase water vaporization losses.
• Sensible Heat (Q_sens)—Represents the energy required to raise the temperature of the solvent entering the desorption system. Its magnitude depends on the temperature difference between the inlet and outlet solvent streams and the specific heat capacity of the solution. In this study, the specific heat capacity of a MEA 30 wt.% solvent was determined using data from Hilliard [52].

In the SF configuration, the lean solvent stream withdrawn from the stripper is divided into two separate streams. As a result, the specific sensible heat q_sens consists of two components. The value of q_sens-deep refers to the energy needed to heat the deep regenerated solvent stream (L₄), while q_sens-semi corresponds to the heating of the partially regenerated solvent stream (L₅). This relationship is shown in Equation (3).

$$q_{sens}=q_{\mathit{\text{sens-semi}}}+q_{\mathit{\text{sens-deep}}}={\displaystyle \frac{Q_{sens}}{m_{C{O}_2}}}={\displaystyle \frac{\sum (L{c}_p{t)}_{outlet}-\sum (L{c}_p{t)}_{inlet}}{m_{C{O}_2}}}$$

(3)

In practical applications, the actual heat supplied to the system is always higher than the theoretical reboiler heat duty due to inevitable heat losses. These losses vary depending on the type of installation, system scale, and process conditions. Proper quantification of heat losses must be considered when analyzing the thermal efficiency of the process. Heat losses can be categorized as follows:

• Stripper heat loss (Q_loss.S)—Heat dissipation from the stripper, influenced by factors such as design, insulation quality, number of connection points, instrumentation, and the surface-area-to-volume ratio of flowing media (scale effect).
• Exchanger heat loss (Q_loss.E)—Heat loss occurring in heat exchangers and other process equipment, including measurement devices and piping between the stripper and cross-flow heat exchangers. The extent of this loss depends on the complexity and the scale of the plant.
• Heat lost in coolers (Q_loss.C)—This occurs when the lean solvent is not sufficiently cooled due to inadequate heat exchange in cross-flow heat exchangers, leading to excess heat removal in coolers. Such a situation is more likely in research facilities where varying process configurations are tested, causing deviations from nominal design parameters. However, this type of heat loss does not occur in properly designed industrial systems operating near their design specifications. For this pilot-scale installation, a maximum acceptable (and not counted as a loss) temperature difference between the lean and rich solvent at the outlet of the cross-flow heat exchangers was set at 7 °C.
• The actual heat input required for solvent regeneration, considering all losses, is given by Equation (4):

$$q_h={\displaystyle \frac{Q_{reg}+Q_{loss.S}+Q_{loss.E}+Q_{loss.C}}{m_{C{O}_2}}}=q+q_{loss.S}+q_{loss.E}+q_{loss.C}$$

(4)

Heat losses can significantly impact pilot-scale measurements, with losses reaching up to 35%, depending on system design and insulation quality [53,54,55]. While such losses are acceptable in research settings, they can distort comparisons between different process configurations if not properly accounted for. The heat loss determination methodology for this pilot plant was thoroughly described by Tatarczuk et al. [35], using the process balance reconciliation method presented in [38,56].

To ensure accurate benchmarking, the use of q as a standardized parameter allows for a fair evaluation of process efficiency across various system configurations and operating conditions.

2.4.3. CO₂ Loading

The primary function of the stripper is to regenerate the solvent by reducing the concentration of dissolved CO₂ in the liquid phase, called the CO₂ loading of the solvent (α, mol_CO2/mol_A). This parameter represents the amount of CO₂ absorbed per mole of amine groups in the solvent and is crucial for assessing both absorption and regeneration efficiency.

In this study, the CO₂ loading was determined through precise density measurements of solvent samples collected during pilot-scale tests. The methodology relied on an experimentally derived correlation α = f(t,ρ), which is specific to the solvent used. For MEA 30 wt.% solutions, the procedure described by Spietz et al. [57] was applied.

3. Results

This study analyses the impact of advanced configurations on the chemical absorption process for CO₂ capture from flue gases. The comparison is based on pilot-scale test results, with global parameters presented in

Table 2. Summary of key raw parameters for pilot tests.

Parameter	Symbol	Value
Process configuration		S	SF	SF-HIS
Inlet gas (volumetric flow), m3n/h	G’21	223.3	214.7	213.5
CO2 capture efficiency, %	η	72.5	76.2	85.1
Liquid/gas flow ratio to the absorber, kg/kg	L/G	4.73	5.92	5.71
CO2 conc. in the inlet gas (dry basis), %vol.	c21_CO2	13.62	12.94	13.04
Reboiler heater power setting, kW	Qh	52.5	47.25	47.25
Total heat flow through recuperators, kW	QREC	0.0	0.0	25.89

. The SF and HIS configurations were evaluated following the scheme in Figure 4, using a previously published standard configuration test as a reference [35].

The mass balance results, including reconciled process stream data, are summarized in

Table 3. Mass flow rate of main components in process streams (data after balance reconciliation).

Symbol—Mass Flow Rate, kg/h	Test
Test Configuration	S	SF	SF-HIS
G21—Inlet flue gas to the absorber K-201	287.5	280.4	282.4
G21_CO2	55.5	51.4	52.3
G21_N2	208.3	202.5	204.7
G21_O2	18.1	19.9	19.6
G21_H2O	5.5	6.6	5.8
G22—Cleaned flue gas at the outlet of the absorber K-201	254.9	258.5	257.0
G22_CO2	15.2	12.2	7.8
G22_N2	208.3	202.5	204.7
G22_O2	18.1	19.9	19.6
G22_H2O	13.2	23.8	24.8
G33—Gas at the outlet of the F-223 separator	41.0	39.7	45.1
G33_CO2	40.3	39.1	44.5
G33_H2O	0.7	0.5	0.6
L11—Rich solvent at the inlet to the heat exchanger E-210	664.8	816.3	813.6
L11_CO2	75.0	88.1	92.9
L11_H2O	395.2	494.1	485.0
L11_A	194.6	234.1	235.7
L21—Rich solvent at the inlet to the heat exchanger E-213	632.9	744.4	706.7
L21_CO2	71.4	80.3	80.7
L21_H2O	376.2	450.6	421.3
L21_A	185.3	213.5	204.7
L31—Rich solvent to the top of the stripper	110.9	138.6	137.9
L31_CO2	12.5	15.0	15.7
L31_H2O	65.9	83.9	82.2
L31_A	32.5	39.8	40.0
L41—Lean solvent to the lower recuperator (w/o HIS to E-213)	637.6	796.1	750.3
L41_CO2	55.6	54.3	52.1
L41_H2O	388.7	503.2	469.7
L41_A	193.3	238.7	228.5
L51—Lean solvent to the upper recuperator (w/o HIS to E-214)	722.3	863.1	862.1
L51_CO2	63.0	89.9	92.7
L51_H2O	440.3	524.4	517.6
L51_A	219.0	248.7	251.9
L61—Condensate (H2O) to the top of the absorber K-201	7.7	0.5	0.7
L81—Make-up water to the upper part of the absorber K-201	8.4	17.8	19.6

. The stream notations align with Figure 4. Raw measurement data from the test campaigns are provided in Appendix B, with instrumentation notations matching the scheme in Figure A1.

3.1. Effect of the SF Configuration

3.1.1. Impact on the Absorber

The SF configuration affects both the absorber and the stripper, as evident from its influence on the temperature profiles in both columns.

Since chemical absorption is an exothermic process, temperature variations directly impact mass transfer between phases. As shown in Figure 6 (left), the main absorption phase occurred in sections S06–S07, where a significant temperature rise was observed. In all tests, lean solvent was introduced at section S08, which acted as a liquid distributor. It then flowed downward, contacting the flue gas, absorbing CO₂, releasing heat, and increasing the temperature in the sections below. In the SF tests, however, a partially regenerated solvent (L₅₅) at 40 °C was introduced above section S04 into the liquid distributor (section S05). This adjustment improved the process driving force by shifting operating conditions away from equilibrium and lowering the temperature in the middle section of the column [58]. This effect is comparable to intercooled absorption, where part of the solvent stream is withdrawn, cooled, and reintroduced into the column [16,59].

However, due to the higher CO₂ loading of this solvent, absorption in sections S02–S05 was less intensive. Ultimately, the SF configuration resulted in a rich solvent stream exiting the absorber at a temperature approximately 5 °C lower than in the standard test. Additionally, CO₂ capture efficiency improved by nearly 4 percentage points, while reboiler heater power setting was reduced by 10% (Table 2), enhancing overall process efficiency.

It should be noted that the tests for the S and SF configurations were not conducted under perfectly identical conditions, particularly with regard to the L/G ratio. However, as reported in an earlier study [32], increasing L/G beyond 5.0 in the standard configuration can negatively affect the capture efficiency, while the SF configuration continues to perform favorably under these conditions. Moreover, the S test was conducted at a slightly higher CO₂ inlet concentration and with higher reboiler heat input, both of which are conditions that typically enhance CO₂ capture efficiency. Nevertheless, the SF test achieved better performance. This implies that under identical conditions, the actual performance advantage of the SF configuration might be even greater. As is often the case in pilot-scale testing, a perfect alignment of conditions is not always feasible.

3.1.2. Impact on the Stripper

The temperature profiles along the stripper packing, presented in Figure 6(right), indicate that in all sections except the reboiler section (S01), the SF configuration resulted in the lowest temperatures compared to the standard test.

With the withdrawal of stream L₅₁ at section S05, the solvent flow through the lower sections of the stripper (S01–S04) was nearly halved compared to the standard case. At a lower Q_h, deeper solvent regeneration was achieved (Figure 7). Furthermore, this stream was heated to a slightly higher temperature in the reboiler. Simultaneously, the vapor flow generated in the reboiler (S01) was significantly lower, reducing heat transfer from the rising vapor to the descending solvent. As a result, temperatures in sections S02–S05 were noticeably lower. This trend was further amplified in sections S06–S09, where a reduced vapor flow contacted the entire rich solvent stream.

Additionally, the rich solvent introduced into sections S09–S12 had the lowest temperature in the SF configuration, leading to further cooling in the upper part of the stripper. An important factor influencing the temperature difference at the top of the stripper is the water vapor content in the gas phase leaving section S09. The gas temperature at this point was noticeably higher in the standard configuration compared to the SF test. Given that the operating pressure in the stripper remains similar across all cases (see

Table A1. Raw measurement data from pilot test with S, SF, and SF-HIS configurations.

	Configuration			S		SF		SF-HIS
Sym.	Measured Parameter	Unit	Meas. no	Value	u	Value	u	Value	u
G21	Absorber inlet flue gas flow	kg/h	FI-138	299.7	5.9	286.5	5.9	285.6	5.9
p21	Pressure of the absorber inlet flue gas	kPa(g)	PI-182	24.6	0.1	25.0	0.1	24.8	0.1
t21	Temperature of the absorber inlet flue gas	°C	TI-137	29.0	1.2	32.6	1.3	30.0	1.2
y21_CO2	The molar fraction of CO2 in the dry inlet flue gas	molCO2/ moldfg	AE-201	0.1366	0.0017	0.1299	0.0019	0.1305	0.0016
y21_O2	The molar fraction of O2 in the dry inlet flue gas	molO2/ moldfg	AE-201	0.0616	0.0017	0.0689	0.0020	0.0674	0.0016
G22	Absorber outlet flue gas flow	kg/h	FI-286	245.2	7.3	254.2	7.3	252.7	7.3
p22	Flue gas pressure at the top of the absorber	kPa(g)	PI-223	20.4	0.1	21.3	0.1	21.2	0.1
t22	Flue gas temperature at the top of the absorber	°C	TI-224	44.9	0.8	56.0	0.8	57.1	0.8
y22_CO2	The molar fraction of CO2 in the dry outlet flue gas	molCO2/ moldfg	AE-225	0.0412	0.0015	0.0336	0.0017	0.0217	0.0015
y22_O2	The molar fraction of O2 in the dry outlet flue gas	molO2/ moldfg	AE-225	0.0673	0.0018	0.0768	0.0021	0.0756	0.0017
G33	Separator S-223 outlet flue gas flow	kg/h	FI-285	40.9	0.4	39.6	0.4	45.1	0.4
p33	Outlet gas pressure from the Separator S-223	kPa(g)	PI-284	-3.0	0.1	-2.0	0.1	-1.8	0.4
t33	Outlet gas temperature from the Separator S-223	°C	TI-280	29.0	1.3	25.0	1.2	25.0	1.2
p31	Gas pressure at the top of the stripper	dm3/h	PI-272	30.0	0.1	29.9	0.1	30.1	0.1
L’11	E-210 HX-rich solvent inlet flow	dm3/h	FI-231	601.0	8.8	736.0	6.6	739.2	6.9
α11	CO2 loading of the rich solvent from the absorber	molCO2/ molA	ZA-305	0.5436	0.0095	0.5246	0.0091	0.5443	0.0094
t11	Temperature of the rich solvent from the absorber	°C	TI-212	49.0	0.8	43.0	0.8	43.0	0.8
t12	Temperature of the rich solvent to the stripper (sec. 10)	°C	TI-235	92.8	0.9	79.0	0.8	77.0	0.8
C%	Amine concentration in the CO2 unloaded solvent	%	ZA-304	33.8	1.7	32.4	1.7	32.4	1.7
L’21	E-213 HX-rich solvent inlet flow	dm3/h	FI-232	571.1	17.8	666.1	10.7	656.7	12.6
t22	Temperature of the rich solvent to the HX E-214	°C	TI-239	97.4	1.0	97.0	0.8	81.0	0.8
t23	Temperature of the rich solvent to the stripper (sec. 8)	°C	TI-240	100.5	0.8	95.0	0.8	81.0	0.8
L’31	Stripper (sec. 12) solvent inlet flow	dm3/h	FI-233	100.4	1.6	125.3	1.5	124.6	1.5
L’41	Stripper solvent outlet flow	dm3/h	FI-251	608.8	2.6	773.0	2.6	727.6	2.4
α41	CO2 loading of the lean solvent from the stripper	molCO2/ molA	ZA-304	0.3967	0.0067	0.3150	0.0055	0.3167	0.0055
t41	Temperature of the lean solvent from the stripper	°C	TI-260	107.0	0.8	108.0	0.8	109.0	0.8
t42	Temperature of the lean solvent to the HX E-213	°C	TI-241	107.0	0.9	108.0	0.8	87.0	0.8
t43	Temperature of the lean solvent to the cooler E-211	°C	TI-238	58.7	0.9	55.4	0.9	52.1	0.8
t44	Temperature of the lean solvent to the top absorber	°C	TI-237	40.0	0.8	40.0	0.8	40.0	0.8
L’51	Stripper (sec. 1/sec. 5) solvent outlet flow	dm3/h	FI-247/ FI-228	689.7	3.1	800.1	2.5	799.8	2.5
α51	CO2 loading of the lean solvent from the stripper (sec. 1/sec. 5)	molCO2/ molA	ZA-306	0.3967	0.0067	0.5007	0.0087	0.5118	0.0089
t51	Temperature of the lean solvent from the stripper (sec. 1/sec. 5)	°C	TI-260/ TI-295	107.0	0.8	85.0	0.8	93.0	0.8
t52	Temperature of the lean solvent to the HX E-214	°C	TI-242	107.0	0.8	85.0	0.8	83.4	0.9
t53	Temperature of the lean solvent to the HX E-210	°C	TI-236	99.2	0.9	86.0	0.8	83.0	0.8
t54	Temperature of the lean solvent to the cooler E-208	°C	TI-234	59.8	0.8	51.0	0.8	51.0	0.8
t55	Temperature of the lean solvent to the mid-section of the absorber	°C	TIC-230	40.0	0.8	40.0	0.8	40.0	0.8
L’61	Top absorber condensate inlet flow	dm3/h	FI-278	7.7	0.5	0.5	0.0	0.7	0.1
t61	Temperature of the condensate from the Separator S-223	°C	TI-280	29.0	0.9	25.0	0.7	25.0	0.7
L’81	Top absorber make-up water inlet flow	dm3/h	FI-207	11.8	0.7	22.1	1.8	20.1	1.6
t81	Temperature of the make-up water	°C	Ext.	12.6	0.7	17.0	0.7	17.0	0.7

), the equilibrium vapor pressure of water in the SF case is significantly lower, resulting in reduced water vapor content in the gas phase. Consequently, the gas phase in the S case contains a higher fraction of water vapor, which is more prone to condensation upon contact with the colder L₃₁ stream. This leads to a reduced cooling effect of the gas in the standard configuration, whereas in the SF case, the gas containing less water vapor undergoes a more substantial temperature drop due to more effective sensible heat exchange with the L₃₁ stream.

However, the lower temperature in the packed sections of the stripper during the SF test (with an average reduction of 8.6 °C compared to the standard configuration) led to a less intensive mass transfer. This is evidenced by the graphical representation of operational stream parameters (L₁₂, L₂₃, L₃₁—inlets; L₄₁, L₅₁—outlets) relative to the equilibrium curve p_CO2 = f(t,α), as shown in Figure 8.

The equilibrium curves for the average and maximum temperatures, determined using Bruder et al.’s equation [51], indicate that the lower average temperature in the stripper’s packed section during the SF test resulted in lower equilibrium CO₂ partial pressure. The operating stream parameters for the SF test were closer to equilibrium conditions, reducing the mass transfer driving force compared to the standard test and resulting in a lower stream of removed CO₂ (Figure 7). The above conditions contributed to the fact that the total heat required for solvent heating (q_sens) was highest for the SF test and amounted to 0.98 MJ/kg_CO2.

When the remaining components of the q_h from Figure 9 are compared, it is evident that in the SF test, a significant reduction in heat losses in the stripper (q_loss.S) was observed. This is most likely associated with a lower solvent flow rate reaching the stripper’s reboiler due to the removal of the L₅₁ stream in section S05. The reduction in q_loss.S can be further confirmed by analyzing the temperature profile in the stripper for the SF test, presented in Figure 6. In contrast to test S, where the temperature decreased by 1.0 °C, an increase in the temperature of the solvent flowing downward along the packing in section S04 was observed.

On the other hand, the increased solvent flow outside the stripper contributed to a slight increase in heat losses within the heat exchanger island (q_loss.E) and the coolers (q_loss.C). Ultimately, the total heat losses were reduced by 12% compared to the standard test.

3.2. Impact of HIS Implementation on Process

To evaluate the potential for improving process efficiency in the SF configuration, an additional test was conducted, incorporating a heat-integrated stripper (HIS). The HIS configuration was activated by adjusting valve positions, enabling the lean solvent to flow through recuperators. Both tests were conducted consecutively under comparable operating conditions to minimize external influences.

The implementation of the HIS had no significant impact on the absorption process, as temperature profiles in the absorber remained highly similar. A slight temperature increase was observed, which may indicate a higher intensity of CO₂ absorption from the gas phase. However, the key benefit of the HIS was the substantial improvement in solvent regeneration, leading to an almost 9 percentage point increase in CO₂ capture efficiency.

During the SF test, the L₄₁ solvent stream at 108.0 °C was extracted from section S01 of the stripper and directed to the cross-flow heat exchanger E-213. Simultaneously, the L₅₁ stream at 85.0 °C was withdrawn from section S05 and routed to E-214. In contrast, during the SF-HIS test, the L₄₁ stream at 109.0 °C was directed to the lower recuperator (REC-L) in sections S04 and S06, where it transferred Q_REC-L = 17.50 kW of heat to the stripper. It then exited as L₄₂ at 87.0 °C and was routed to E-213. The semi-lean solvent stream L₅₁, at 93.0 °C, was extracted from S05 and introduced into S08, where it transferred Q_REC-U = 8.39 kW of heat in the upper recuperator before leaving the stripper. The L₅₂ stream, at 83.4 °C, was then directed to E-214 and subsequently to the cross-flow heat exchanger E-210.

3.2.1. Upper Sections of the Stripper Performance

The temperature profile for the SF-HIS test (Figure 6) indicates that the use of solvent flow through the upper recuperator did not increase the temperature in sections S08 and S09. In this test, the upper recuperator was supplied with stream L₅₁ from section S05. The temperature of this stream was only 3 °C higher than t_avg, resulting in a low heat exchange driving force. Therefore, the heat flow transferred in the upper recuperator was twice as low as in the lower recuperator, and the temperature profiles for these stripper sections in the SF and SF-HIS tests were similar but clearly lower than in the standard configuration. Consequently, the q_vap remained consistently low, with a value of 0.07 MJ/kg_CO2, indicating that heat integration in the SF configuration had no significant impact on this parameter.

The desorption heat q_des remained unchanged across all tests despite having the largest contribution to the reboiler heat duty. Its value is independent of process configuration [47,60], and therefore, its role is not significant in comparing tests that used the same solvent. Given that MEA 30 wt.% was the solvent used in all tests, q_des was assumed to be 1.92 MJ/kg_CO2 (Figure 9) based on data from Bruder et al. [51].

3.2.2. Lower Sections of the Stripper Performance

The presence of recuperators in the Split Flow configuration resulted in a 4.4 °C increase in the average temperature (t_avg) within sections S02–S09 of the stripper during the HIS test. However, the maximum temperature (t_max) in section S01, where L₄₁ was extracted, remained the highest, despite the lowest heat duty (Q_h). The higher t_avg in the packed section of the stripper during the HIS test resulted in a correspondingly higher equilibrium CO₂ partial pressure (Figure 8), enhancing the mass transfer driving force compared to the SF test conditions.

Intense mass transfer in the HIS test is further evidenced by the positioning of the L₅₁ stream parameters. As shown in Figure 7, the CO₂ loading of L₅₁ (α₅₁) decreased by 0.0361 mol_CO2/mol_A relative to α₁₁, corresponding to an additional 12.5 kg/h of CO₂ transfer to the gas phase, which accounted for 28% of the total captured CO₂ in this test. Conversely, in the SF test, α₅₁ decreased by only 0.0201 mol_CO2/mol_A, leading to 7.1 kg/h of desorbed CO₂, or 18% of the total captured CO₂. Additionally, the regeneration of the L₄₁ stream was more intensive in the HIS test, resulting in a greater Δα_11–L41 value.

3.2.3. Impact of HIS on Heat Losses

The increase in heat exchange intensity within the stripper after implementing the HIS resulted in a noticeable rise in q_loss.S compared to the SF test, reaching a level similar to that observed in the S test. However, a significant decrease in other heat loss components was recorded. The reduction in the amount of heat directed to cross-flow heat exchangers lowered q_loss.E by 0.40 MJ/kg_CO2, while the existing heat exchange surface was sufficient to enhance heat transfer between the lean and rich solvent. This improvement led to a reduction in cooler heat losses (q_loss.C) by 0.29 MJ/kg_CO2.

As a result, the total heat losses for the test with the HIS were 27% lower compared to the SF test and 36% lower compared to the S test. Nevertheless, they still accounted for over 26% of q_h, which is an acceptable level. Pilot facilities typically experience higher heat losses, reaching up to 35% depending on design and construction [53,54,55].

4. Discussion

4.1. Reduction in Process Heat Demand

Pilot plant studies using a well-known MEA 30 wt.% solvent demonstrated the beneficial impact of innovative process modifications. The implementation of the SF and HIS modifications resulted in a 10% reduction in heat demand for the process (Figure 10), which, in industrial-scale applications, could lead to significant cost savings and an overall improvement in the economic efficiency of decarbonization processes.

Ultimately, the reboiler heat duty for the SF-HIS test was 2.82 MJ/kg_CO2, a value comparable to commercial technologies such as the DMX™ process, developed by Axens and IFPEN, which requires regeneration energy between 2.3 and 2.9 MJ/kg_CO2 [61]; the technology developed by Linde and BASF using the OASE^® blue solvent, with a heat demand of 2.7 MJ/kg_CO2 [62], and MHI’s KS-1 solvent, which ranges from 2.44 to 2.53 MJ/kg_CO2 [17].

The key findings of this study, highlighting the benefits of process modifications, are presented in Figure 10. Additionally, previously published results of the HIS implementation in the standard configuration [35] were included to identify common trends.

4.2. Reduction in Heat of Vaporization

The application of the SF configuration significantly reduced the temperature at the top of the stripper, creating conditions that limited water losses through vapor emissions from the system. This reduction directly lowered the energy required for water evaporation in the stripper, thereby decreasing q_vap.

The comparison of tests presented in Figure 11 confirms that a lower temperature at the top of the stripper results in a reduced q_vap. In the test using the SF configuration, a significant decrease in q_vap was observed compared to the standard test. While the HIS configuration also contributed to a notable temperature reduction at the top of the stripper [35], the SF modification lowered it to such an extent that the introduction of the HIS had no further impact. This suggests that when combined with the split-flow configuration, heat recuperation in the lower section of the stripper is sufficient.

This indicates that combining multiple configurations may not provide a synergistic benefit, as they may influence the process in a similar manner, affecting the same aspect of Equation (2). Previously published studies have shown that implementing the HIS in the standard configuration can reduce heat demand by up to 7% (Figure 10). However, in the case of integration with SF, a lower reduction was achieved, partly due to the limited impact on q_vap.

An additional advantage of these two modifications is the reduction in cooling requirements for the captured gas. The amount of heat that must be removed to condense moisture is largely equivalent to q_vap, as minimizing water content before further processing (including compression) is typically necessary. Therefore, the application of the SF and HIS configurations reduces cooling water demand, leading to lower operational costs. Additionally, the required surface area of the gas cooler and the size of the condensate separator are reduced, contributing to lower capital expenditures (CAPEX).

4.3. Reduction in Sensible Heat

A characteristic decrease in stripper temperature during the SF test, particularly in terms of t_avg, had a negative effect on q_sens, which increased significantly compared to the standard test. In both HIS tests shown in Figure 11, an increase in t_avg was observed relative to tests without the HIS, intensifying the solvent regeneration process. The higher CO₂ capture rate in the SF-HIS test at an identical Q_h led to a lower overall q_sens for the lean and semi-lean solvents, which was 0.15 MJ/kg_CO2 lower compared to the SF test (Figure 9). Since heat integration did not influence q_vap, the reduction in q_sens was the primary contributor to the overall benefit of using the HIS in the split-flow configuration.

The implementation of heat integration through recuperators in the form of heat exchangers inside the stripper reduces the required external heat exchange surface. During the SF test, cross-flow heat exchangers and coolers transferred a total of 77.41 kW (Q_cross + Q_loss.C) from the lean solvent. The introduction of flow through the recuperators in the subsequent test significantly reduced the amount of heat exchanged outside the stripper (Figure 12). Since the mass flow rate of released CO₂ also increased, a fair comparison was made by normalizing Q_cross and Q_loss.C to the mass of separated CO₂. The SF-HIS test demonstrated a 36% reduction in external heat exchange compared to the SF test and a 28% reduction compared to the standard test, with the majority of this improvement attributed to a decrease in Q_cross due to the relatively small contribution of Q_loss.C. This reduction contributes to lower capital expenditures (CAPEX) related to the heat exchanger island.

It should also be noted that introducing process configurations into the system will inevitably involve additional investment costs.

For the SF configuration, the required modifications include the following:

• A dual pumping system for lean and semi-lean solvents;
• Modification of the stripper to allow semi-lean solvent extraction;
• Redesign of the absorber to accommodate changed solvent flows;
• Increased heat exchange surface in cross heat exchangers.
• For the HIS configuration, the modifications include the following:
• Significant modification of the stripper, particularly due to the integration of internal recuperators;
• A slightly increased pumping load for the lean solvent.

Even if the SF configuration requires more changes, these adjustments are generally easier to implement in existing facilities. In contrast, the HIS requires more extensive modifications, particularly in the stripper section, making it a more cost-effective option for new installations. Nevertheless, heat consumption remains the primary operational cost, and a 5–10% reduction in heat duty can generate substantial economic benefits, helping to offset investment costs [34].

4.4. Comparison with Literature Data

The available literature does not report practical applications of thermal integration in the stripper via built-in heat exchangers (recuperators), but several studies have explored this concept through simulations. The most frequently analyzed and similar configuration is Stripper Interheating (SI), where solvent heating is performed externally before re-entering the stripper.

Zhao et al. [23] conducted a simulation of CO₂ capture using MDEA/Pz (30/20%) with a stripper pressure of 2.1 bar_(abs). Their results showed that SI reduced q from 2.74 to 2.56 MJ/kg_CO2, representing a 7% decrease. They also identified reductions in q_vap and q_sens, which aligned with the findings from this study. However, differences in operating pressure influenced the magnitude of these effects.

Similarly, Li et al. [22] developed an optimized CO₂ capture model, achieving a heat duty reduction from 3.60 to 3.38 MJ/kg_CO2 when implementing SI, corresponding to an over 6% improvement. Their findings also indicated reductions in q_vap and q_sens, with q_des remaining unchanged.

Recent pilot-scale experiments on Stripper Interheating with MEA 30 wt.% solvent at the Amager Bakke Waste-to-Energy plant demonstrated a specific reboiler duty reduction from 3.71 to 2.71 MJ/kg_CO2 while maintaining an 83% capture efficiency. Further modeling suggested a reduction to 2.25 MJ/kg_CO2 at higher capture efficiencies [25]. An external heat exchanger powered by a separate steam source was used, which, in the long term, could enable the utilization of waste heat sources in a similar manner.

In this regard, the Heat Integrated Stripper also offers significant potential for thermal integration. Its design allows recuperators to be supplied with external waste heat sources available in various industrial sectors, such as cement plants, steelworks, and waste incineration facilities. Moreover, the HIS enables the integration of solvent-based CO₂ capture with exothermic catalytic processes, such as CO₂ hydrogenation for synthetic natural gas (SNG) production. Excess heat from this process could be used through recuperators to improve the stripper’s temperature profile and enhance solvent regeneration efficiency, as proposed in Tauron’s patent PL238451B1 [63]. Such integration strategies represent a promising direction for further research.

Until now, the only reported industrial studies on the application of the Heat Integrated Stripper for solvent-based CO₂ capture have been conducted within the Tauron pilot plant, as presented in this work. However, previous publications have compared the performance of the HIS with the standard configuration [34,35,38].

5. Conclusions

This study demonstrates that energy efficiency improvements in solvent-based post-combustion CO₂ capture can be achieved not only through advanced solvents but also via innovative process modifications. The integration of Split Flow and Heat Integrated Stripper with a widely studied MEA 30 wt.% solvent resulted in a reboiler heat duty reduction to 2.82 MJ/kg_CO2, which is comparable to commercial-scale technologies, thus indicating a potential for significant operational cost reductions. Further improvements could be expected when applying these modifications in combination with next-generation solvents.

Comparison with the standard process configuration revealed that SF alone reduced the reboiler heat duty by 5%, while the addition of the HIS further decreased it by another 5%. The combination of these configurations also resulted in an over 12 percentage point increase in CO₂ capture efficiency. The SF configuration improved absorption conditions and significantly lowered the stripper’s top temperature, reducing water vapor emissions and associated heat losses. As a result, implementing this solution is expected to reduce cooling water demand and the capital costs associated with the stripper condensation system.

The integration of the HIS with SF further enhanced solvent regeneration by increasing the average temperature within the stripper, particularly in the zones affected by the implanted heat exchangers. This improved mass transfer conditions and had a secondary effect of enhancing CO₂ capture efficiency. Additionally, heat integration led to a 36% reduction in overall heat losses and a decrease in external heat exchanger duty by up to 28%. These benefits will directly translate into lower investment and operating costs compared to the standard solvent-based CO₂ capture process.

While these modifications present clear economic and process advantages, it is important to consider potential challenges, particularly in adapting absorber and stripper column designs to accommodate these configurations. However, the results from over 2500 h of industrial-scale (at Tauron facilities) testing indicate no operational issues related to the implanted heat exchangers, thus providing strong support for their further scale-up and commercialization.

The findings presented in this study provide a valuable basis for the engineering-scale design of CO₂ capture facilities. The implementation of SF and the HIS in mature solvent-based processes can significantly enhance efficiency and accelerate decarbonization efforts, particularly in sectors where chemical absorption remains the most viable option until emerging carbon capture technologies reach commercial maturity.