Hydrogen Purification through a Membrane–Cryogenic Integrated Process: A 3 E’s (Energy, Exergy, and Economic) Assessment

Abstract

Hydrogen (H₂) is known for its clean energy characteristics. Its separation and purification to produce high-purity H₂ is becoming essential to promoting a H₂ economy. There are several technologies, such as pressure swing adsorption, membrane, and cryogenic, which can be adopted to produce high-purity H₂; however, each standalone technology has its own pros and cons. Unlike standalone technology, the integration of technologies has shown significant potential for achieving high purity with a high recovery. In this study, a membrane–cryogenic process was integrated to separate H₂ via the desublimation of carbon dioxide. The proposed process was designed, simulated, and optimized in Aspen Hysys. The results showed that the H₂ was separated with a 99.99% purity. The energy analysis revealed a net-specific energy consumption of 2.37 kWh/kg. The exergy analysis showed that the membranes and multi-stream heat exchangers were major contributors to the exergy destruction. Furthermore, the calculated total capital investment of the proposed process was 816.2 m$. This proposed process could be beneficial for the development of a H₂ economy.

1. Introduction

Energy consumption is a critical issue in the global fight against climate change. Burning fossil fuels to meet the world’s growing energy demands has significantly increased greenhouse gas emissions. According to the International Energy Agency [1], the total global energy-related greenhouse gas emissions increased by 1.0% in 2022, reaching a record high of 41.3 Gt CO₂-eq., 89% of which is the direct CO₂ from combustion and other industrial processes. A more sustainable trend compared to 2021 has been observed, primarily because of the slow industrial growth in China and Europe [2] and increased implementation of clean technologies such as renewables, electric vehicles, and heat pumps.

Hydrogen (H₂) is considered to be one of the key players in the shift from a carbon economy to a more sustainable and clean-energy future. H₂ can be a game changer in shifting away from a carbon economy by providing a clean and versatile energy carrier for various sectors and applications [3]. H₂ can generate electricity with only water and heat as the byproducts in fuel cells. H₂ can also be used as a clean fuel for transportation, replacing fossil fuels and reducing the emissions of pollutants such as nitrogen oxides and particulate matter. H₂ can also enable the integration of renewable energy sources into the grid, enhance energy security and resilience, and create new business opportunities [4]. However, H₂ still faces many challenges, such as its high cost, low efficiency, infrastructure gaps, and safety issues [5]. The production and use of H₂ as a fuel have the potential to reduce greenhouse gas emissions and improve air quality significantly.

On the other hand, H₂ production can also generate emissions if it relies on fossil fuels [6]. Therefore, some countries are investing in developing emission-free H₂ production methods, such as using renewable energy or carbon capture and storage. H₂ can be produced from various low-carbon or carbon-free sources, such as renewable energy, nuclear power, and natural gases with carbon capture and storage technology [7]. Commercially, H₂ is produced from natural gas, coal, biomass, and water, resulting in impurities such as carbon monoxide, carbon dioxide, sulfur compounds, and water [8]. These impurities can negatively impact the performance of fuel cells and other H₂-based technologies. Therefore, H₂ purification is a crucial step in the production and use of H₂.

The importance of H₂ purification has grown significantly in recent years with the increasing interest in H₂ as a clean energy source. The use of H₂ as a fuel for transportation, power generation, and other applications requires high-purity H₂ to ensure the optimal performance and reliability [9]. The purity requirements of H₂ depend on its intended use. The purification standards for H₂ are the specifications that define the quality and purity of this H₂ for its different applications and use. For example, ISO 14687 is a standard that specifies the quality characteristics of H₂ fuel for its use in proton exchange membrane fuel cells [10]. For example, in fuel cell applications, the H₂ purity must be greater than 99.99% to prevent the contamination of the fuel cell electrodes and maintain their efficiency. ISO 19983 is a standard that specifies the safety requirements for H₂ separation and purification systems. The allowable concentration of carbon monoxide in H₂ for fuel cell applications is typically less than 10 parts per billion (ppb) [11].

Different purification methods can remove specific impurities from H₂, including pressure swing adsorption (PSA), membrane separation, and catalytic processes [12]. PSA is a widely used method that uses adsorbent materials to separate H₂ from other gases by changing the pressure of the feed stream. PSA can achieve high purity levels (>99.9%) and high recovery rates (>90%), but it requires high pressure and large equipment [13]. Membrane-based separation is a promising method that uses selective membranes to allow H₂ to pass through while blocking other gases. Membrane-based separation can operate at a low pressure and temperature and has a low energy consumption and environmental impact. However, membrane-based separation may suffer from a low selectivity, stability, and high cost. Cryogenic distillation is a method that uses low temperatures to liquefy and separate H₂ from other gases. Cryogenic distillation can produce pure H₂ (>99.999%), but has a very high energy consumption and capital cost [14].

The choice of H₂ purification method depends on various factors, including the level of purity required, the impurities present in the feed gas, the operating conditions, the volume of H₂ to be purified, and the economic factors. Several hybrid H₂ purification techniques combine different methods to achieve higher levels of purity, address specific impurities, and achieve a better performance and efficiency. Amosova et al. [15] proposed an integrated system of membrane and PSA technologies for H₂ separation from a gas mixture of H₂/CO/CO₂/N₂/H₂S. The hybrid structure could achieve a 90–97% recovery for cases where the biosynthesis and petrochemical gases can have a 70% concentration of H₂ at the exit of the membrane section. The simulations of the hybrid PSA–membrane and membrane–PSA processes carried out by Li et al. [16] showed that, for the same feed gas (H₂ concentration 62.75%), the latter combination achieved a higher recovery (97.06%). Ohs et al. [17] applied a cost-minimization optimization of the membrane–PSA hybrid process for bioH₂ purification. The hybrid PSA–membrane model for an on-site ammonia-fed H₂ refueling station system presented by Lin et al. [18] demonstrated a 95% recovery rate. Ma et al. [19] used a similar combination for refinery off-gas to achieve net zero H₂ emissions. The CRYOCAP^TM process developed by Air Liquide integrates a cryogenic purifier with PSA to separate CO₂ from the off-gas. Studies [20,21] have shown a 98% recovery of H₂ and a 99.5% purity of CO₂, though H₂ purity has not been mentioned. Van Acht et al. [22] integrated PSA with an electrochemical H₂ purification and compression system to obtain a 99.999% purity using coke oven gas as the feed. The overall percentage of the recovery was not reported, so it is difficult to comment on the efficacy of this scheme.

Membrane–cryogenic hybrid processes are processes in which membrane technology and condensation equipment are combined to achieve a gas separation. This type of process is particularly efficient when at least one component is condensable [23]. Agrawal et al. [24] used a hybrid system combining membrane and cryogenic separation technologies, which resulted in a 96.5% H₂ purity and a 98% recovery from a fluid catalytic cracking unit off-gas. The power requirement of the hybrid approach is 35% less than that of the standalone cryogenic process and requires less capital investment, but this is at the cost of thermodynamic efficiency. The separation of propylene and nitrogen has been efficiently achieved by combining a hybrid process of cryogenic separation and membrane modules [25,26]. Lin et al. [27] used membrane and cryogenic technologies to separate H₂ from feed derived from a coal oxy-fired GE gasifier containing 56% H₂, achieving an optimum H₂ recovery of 98.7% and a purity of 90.3%. Liao et al. [28,29] focused on optimizing the economic performance of the CO₂ capture process using a hybrid membrane–cryogenic system. In the oilfields business, when employing the CO₂-enhanced oil recovery method, this technique can save around 10% of the consumed energy while achieving a 99% purity and 96% recovery [30]. This concept has also found its potential application in the air separation process, wherein pre-enriching oxygen can result in an 11% reduction in the flow rates to the cryogenic section [31]. Burdyny and Struchtrup [32] reported a 0.9% improvement in the efficiency of the air separation process. Another hybrid setup is a multistage membrane and distillation with an integrated heat pump for propylene/propane separation. Park et al. [33] reported purity and recovery percentages of >99%.

From the literature review, it can be inferred that hybrid techniques offer advantages such as higher levels of purity, increased efficiency, and an improved selectivity for specific impurities. However, they can also be more complex and require additional equipment and infrastructure. Most of the prior works accessible in the open literature were related to CO₂ capture and separation. Only a few were focused on H₂ purification and efficiency and those too suffered from either a low thermodynamic efficiency and recovery or compromised on the final H₂ purity, which is an important factor for fuel cell applications. This leaves a big gap that needs immediate attention. Therefore, H₂ separation and purification have been targeted in the present study. For this purpose, membrane- and cryogenic (desublimation)-based technologies are integrated to produce high-purity H₂ with a high recovery. To the best of authors’ knowledge, this is a unique integration attempt for maximizing the throughputs. The emphasis is also on assessing and identifying the energy sinks and economic feasibility of the proposed membrane–cryogenic hybrid process.

2. Methods

A hybrid membrane–cryogenic process for the purification of H₂ is proposed in the present study. The following lines briefly introduce the details of the proposed process and the analysis approaches used to assess its efficacy and potential feasibility.

2.1. Process Description and Simulation

H₂ can be produced from various sources and its composition varies widely based on these sources. The present study considered the feed conditions and composition from the works of Xu et al. [34], as this mixture represents typical fuel conversion technologies. The feed gas enters the membrane section first, followed by the cryogenic separation to improve the purity level of the H₂ gas product. Figure 1 provides an overview of the proposed hybrid scheme, while the pertinent details of each section are mentioned in the following subsections.

2.1.1. Membrane Process

The choice of membranes and their module configuration heavily depends on the components involved in the gaseous mixture and their permeability or the selectivity of the material used. For mixtures like the one used in this study, the membranes can be selective towards H₂ or CO₂. Glassy membranes are selective towards H₂, but the low pressure on the permeate side (H₂) is a disadvantage [35]. Since a cryogenic process will be used for enhancing the purity, permeability has been preferred over selectivity, and a rubbery membrane type is used. A two-stage or a two-step approach is commonly used in membrane separation; however, to enhance this separation, a hybrid approach, as presented by Riaz et al. [36], has been considered in the present study. The process involves three spiral-wound membrane modules with H₂ on the permeate side. Two compressors (K-1 and K-2) are employed to raise the pressure of the CO₂-rich permeate streams to 10 bar. In this study, the membrane material is polyethylene oxide [37], which achieves a CO₂/H₂ selectivity of 8–15 or more [38]. The gas permeance values in the membrane simulation are CO₂ = 1580 GPU and H₂ = 195 GPU [37].

2.1.2. CO₂ Solidification Process

H₂ is the lightest of the known elements and has a very low boiling point (−252.8 °C), second only to Helium. However, CO₂ can be solidified via an anti-sublimation process, whereby CO₂ vapors are directly transformed into the solid state. This feat can be achieved within a specific temperature and pressure range, i.e., 1 bar and −81 °C to 6.3 bar and −56 °C [39]. The same technique has been applied to separating CH₄ from Biogas [40].

In the present study, the H₂-rich stream obtained from the retentate side of the membrane section is at 10 bar and 25 °C. The gas is expanded to 4 bar pressure via an expander (T-1) and passed through a multi-stream heat exchanger (CHX-1), wherein its temperature is further reduced to −58 °C. This cold gas stream then enters the specially designed cold box (CB-1), in which the temperature is further dropped to −61 °C. The small residual quantity of the CO₂ solidifies in this chamber and the pure H₂ is obtained from the top. The solidified CO₂ is periodically removed from the cold box. The suitable conditions to perform this solidification phenomenon are chosen based on the phase behavior of the CO₂, as shown in Figure 2.

Figure 2 shows that the product CO₂ conditions (−61 °C and 4 bar) lie in the solid region below the triple point, which ensures the CO₂ solidification conditions. Furthermore, this process is validated by a recent study from Yurata et al. [41]. In this study, an ethane- and propane-based mixed refrigerant (MR) in an external refrigerant cycle is used for cooling purposes in CHX-1. The MR is compressed to 14.4 bar and cooled to 23 °C in a series of compressors (K-3 and K-4) and coolers (E-4 and E-5). The MR is then passed through CHX-1 to reduce the temperature to −44.1 °C. After the temperature reduction, the pressure of MR is reduced to 1.17 bar. The low-pressure MR at 1.17 bar and −59 °C is again passed through CHX-1 to exchange its cold energy with the feed stream and MR stream. After exchanging heat, the MR stream (14) is recycled back to complete the refrigeration loop.

2.1.3. Process Simulation

The entire process is simulated in Aspen HYSYS^® V14, using Peng Robinson [42] as the fluid property package. The equation of state is the flagship thermodynamic model in the Aspen database, which has been endorsed by numerous scientists [41,43].

Table 1. Design parameters of the proposed H₂ separation and purification process [34].

Feed Conditions	Values
Vapor/Phase Fraction	1.00
Temperature (°C)	35.00
Pressure (bar)	5.00
Mass Flow (kg/s)	100.00
H2 (mol%)	0.20
Carbon Dioxide (mol%)	0.80

presents the design parameters used in the present study, whereas

Table 2. External refrigerant cycle operational parameters.

Refrigerants Conditions	Values
C2 (mol%)	25.33
C3 (mol%)	74.67
Suction Pressure (bar)	1.18
Discharge Pressure (bar)	14.40
Temperature (°C)	23.02

describes the refrigerants’ flows and cycle operating conditions.

It is pertinent to mention the assumptions made to simulate the proposed process successfully. This is crucial for the reproducibility of this work and in line with other works using similar or even more extreme conditions [44,45]. Following is the list:

• No experimental data are available on the amount of CO₂ that solidifies under the operating conditions; a 100% solidification rate is assumed.
• The compressors in the refrigeration cycle are operated with a compression ratio of ≤3.0.
• The pressure drop across all the coolers and heat exchangers is considered to be negligible.
• The efficiencies of the compressors, pumps, and expanders are maintained at 75%.
• The minimum internal approach temperature (MITA) for the multi-stream heat exchanger is maintained at 1.0–2.0 °C.
• The heat loss is considered to be negligible.

The composition data of the major process streams are presented in

Table 3. Composition data for the important process streams (mol%).

Stream ID	CO2	H2	C2	C3
Feed	0.8000	0.2000	0	0
CO2-m	0.9908	0.0092	0	0
H2-m	0.9996	0.0004	0	0
14	0	0	0.2533	0.7467
H2	0	1.000	0	0
sCO2	1.000	0	0	0

.

2.2. Process Optimization

In combination with the thermodynamic knowledge of the process, the built-in optimization environment of Aspen Hysys v14 is used to determine the optimal values of the design variables in the proposed process. In Aspen Hysys, many optimization schemes are available, such as Fletcher Reeves, Quasi-Newton, Box, sequential quadratic programming (SQP), and Mixed. In this study, a BOX optimization scheme is used to evaluate and determine the optimal values of the design variables against the minimal energy consumption. Box optimization is suitable for handling inequality constraints. Inside the cryogenic heat exchangers (CHX-1), the minimum temperature approach value of 1.0 °C is applied as a constraint during the optimization. The objective function of the optimization is to reduce the energy consumption.

2.3. Energy Analysis

The energy analysis of the proposed process includes an evaluation of the design variables and composite curves. The method for the energy analysis is given in the following subsections.

2.3.1. Design Variable Analysis

The proposed integrated process is divided into two main subsections, i.e., the membrane section and the desublimation section. The design variable analysis is conducted based on each section’s SEC. The design variables include the membrane pressure, refrigeration cycle flowrates, suction, and discharge pressure of each cycle. The values of these design variables are presented in Table 5.

2.3.2. Composite Curve Analysis

A composite curve (CC) analysis of the proposed study is performed to analyze the heat flow within the hot and cold streams of the multi-stream heat exchanger (CHX-1). A variation in the heat flow concerning the temperature along the length of the CHX-1 is observed in terms of the temperature heat flow CCs (THCCs). The variation in the temperature approach concerning the temperature is analyzed in temperature difference CCs (TDCCs). The minimum temperature approach in TDCC is expressed as the minimum internal temperature approach (MITA). The investigation of the MITA value is essential, as it validates the feasibility of the heat exchanger. The value of the MITA in the proposed study is set as 1 °C [39].

2.4. Exergy Analysis

The exergy analysis of the process provides insight into its inefficiencies. The total exergy of the process system consists of the potential, kinetic, physical, and chemical exergies. The proposed process is assumed to be in a steady state, at ground level, and involve no chemical reaction; therefore, the potential, kinetic, and chemical exergies are neglected for the studied system. The values of the physical exergy are taken from the Aspen Hysys v14. The analysis is performed by calculating the exergy destruction across each piece of equipment. The formulas for the physical exergy (Equation (1)) are shown below:

$${ex}_{ph,i}=\left(h_i-h_o\right)-T_o(s_i-s_o)$$

(1)

whereas ${ex}_{ph,i}$ is the physical exergy, $h_i$ is the enthalpy of the component $i$, $h_o$ is the enthalpy at the standard state, $S_i$ is the entropy of the component $i$, $S_o$ is the entropy at the standard state, and $T_o$ is the temperature at the standard state. The standard state is 1 atm and 298.15 K. The exergy destruction of each piece of equipment is calculated using the equations given in

Table 4. Equations for the calculations of exergy destruction [39].

Equipment	Exergy Destruction Equations
Compressors	$\mathit{E}\mathit{x}=\dot{\mathit{m}}\left({\mathit{e}\mathit{x}}_{\mathit{i}\mathit{n}}-{\mathit{e}\mathit{x}}_{\mathit{o}\mathit{u}\mathit{t}}\right)+\dot{\mathit{W}}$
Expanders	$\mathit{E}\mathit{x}=\dot{\mathit{m}}\left({\mathit{e}\mathit{x}}_{\mathit{i}\mathit{n}}-{\mathit{e}\mathit{x}}_{\mathit{o}\mathit{u}\mathit{t}}\right)-\dot{\mathit{W}}$
Pumps	$\mathit{E}\mathit{x}=\dot{\mathit{m}}\left({\mathit{e}\mathit{x}}_{\mathit{i}\mathit{n}}-{\mathit{e}\mathit{x}}_{\mathit{o}\mathit{u}\mathit{t}}\right)+\dot{\mathit{W}}$
Phase separators	$\mathit{E}\mathit{x}=\dot{{\mathit{m}}_1}{\mathit{e}\mathit{x}}_1-(\dot{{\mathit{m}}_2}{\mathit{e}\mathit{x}}_2+\dot{{\mathit{m}}_3}{\mathit{e}\mathit{x}}_3)$
Heat exchangers	$\mathit{E}\mathit{x}=\dot{\mathit{m}}\left({\mathit{e}\mathit{x}}_{\mathit{i}\mathit{n}}-{\mathit{e}\mathit{x}}_{\mathit{o}\mathit{u}\mathit{t}}\right)$
Cryogenic heat exchanger	$\mathit{E}\mathit{x}={\displaystyle \sum _{\mathit{i}=1}^{\mathit{n}}}\dot{{\mathit{m}}_{\mathit{i}}}\left({\mathit{e}\mathit{x}}_{\mathit{i}\mathit{n}}-{\mathit{e}\mathit{x}}_{\mathit{o}\mathit{u}\mathit{t}}\right)$
Cold chamber	$\mathit{E}\mathit{x}=\left(\dot{{\mathit{m}}_1}{\mathit{e}\mathit{x}}_1-\dot{{\mathit{m}}_2}{\mathit{e}\mathit{x}}_2-\dot{{\mathit{m}}_3}{\mathit{e}\mathit{x}}_3\right)+\mathit{Q}$
Membranes	$\mathit{E}\mathit{x}=\dot{{\mathit{m}}_1}{\mathit{e}\mathit{x}}_1-(\dot{{\mathit{m}}_2}{\mathit{e}\mathit{x}}_2+\dot{{\mathit{m}}_3}{\mathit{e}\mathit{x}}_3)$

.

The mass flow rate of the stream under observation is given as $\dot{m}$. Additionally, the exergies of the input streams are given as ${ex}_{in}$ and as ${ex}_{out}$ for the output streams, $\dot{W}$ is the work required to provide refrigeration or heating, and $Q$ is the heat supplied.

2.5. Economic Analysis

The economic analysis of the proposed process includes calculations of the total capital investment (TCI), total operating cost (TOC), and total annualized cost (TAC). In the proposed process, the membrane section includes membranes, compressors, and coolers. In the desublimation section, compressors, coolers, expanders, a heat exchanger, and the cold box are the major equipment. The cost of the membrane is calculated using relations from the literature [46]. The cost relations for each piece of equipment are taken from Turton et al. [47]. However, the cost relation for the compressors (Equation (2)) is taken from a study by Ahmad et al. [48].

$${\mathrm{C}}_{\mathrm{C}\mathrm{o}\mathrm{m}\mathrm{p}}=8650\times {\left(\frac{{\mathrm{W}}_{\mathrm{C}\mathrm{o}\mathrm{m}\mathrm{p}}}{{\mathsf{\eta }}_{\mathrm{C}\mathrm{o}\mathrm{m}\mathrm{p}}}\right)}^{0.82}$$

(2)

where ${\mathrm{C}}_{\mathrm{C}\mathrm{o}\mathrm{m}\mathrm{p}}$ = the cost of the compressor, ${\mathrm{W}}_{\mathrm{C}\mathrm{o}\mathrm{m}\mathrm{p}}$ = the power required by the compressor (hp), and ${\mathsf{\eta }}_{\mathrm{C}\mathrm{o}\mathrm{m}\mathrm{p}}$= the efficiency of the compressor.

The relation for estimating the purchasing cost is presented in Equation (3). The operating costs include the electricity, cooling, labor, maintenance, and other costs.

$$log{E}_{pc}=C_1+C_{2}logA+C_3{\left(logA\right)}^2$$

(3)

whereas $E_{pc}$ = purchased cost of the equipment, $C$ = the equipment cost constants, and $A$ = the equipment capacity.

3. Results and Discussion: Process Analysis

3.1. Energy Analysis

In the energy analysis, the design variables and CC analysis are the main two parameters upon which the analysis of the proposed integrated system is performed.

3.1.1. Design Variables Analysis

In the design variable analysis, the effect on the system is observed by varying the values of the design variables until the optimal value is gained. The design variables’ values are presented in

Table 5. Design variables analysis results of the proposed integrated system.

Section	Variables	Values
Membrane Separation	High pressure (bar)	10.0
	Low pressure (bar)	0.10
	Mem-1 area (m2)	9000
	Mem-2 area (m2)	500
	Mem-3 area (m2)	1000
	CO2 purity (%)	99.08
	CO2 recovery (%)	98.13
De-sublimation cycle	C2 (kg/s)	2.305
	C3 (kg/s)	9.962
	Discharge Pressure (bar)	14.40
	Suction Pressure (bar)	1.177
	Desublimation temperature (°C)	−61
	Desublimation pressure (bar)	4.00
	CHX-1 MITA (°C)	1.00
	H2 purity (%)	99.96
	H2 recovery (%)	95.9

.

Table 5 shows that the value of the high pressure in the membranes is 10 bar. This pressure value is opted for to maximize the separation and minimize the membrane areas. The values of the optimum membrane areas are 9000, 500, and 1000 m², respectively. In the desublimation cycle, the values of the refrigerant flowrate, i.e., C₂ and C₃, are 2.305 kg/s and 9.962 kg/s, respectively. The discharge and suction pressure values are 14.4 bar and 1.17 bar, respectively. The desublimation temperature and pressure conditions are adopted considering the phase behavior of the CO₂ (see Figure 2). These values are optimally selected, keeping the MITA at 1.0 °C. The final purity and recovery of the H₂ obtained are 99.99% and 95.9%, respectively. The goal of the optimization is to minimize the SEC, achieving a high purity and recovery of H₂ while keeping the MITA within its limits. The values of the membrane section SEC, desublimation section SEC, and net process SEC are presented in

Table 6. Net overall specific energy consumption (SEC) of the proposed integrated process.

Membrane section SEC (kWh/kg)	2.26
Desublimation section SEC (kWh/kg)	0.11
Net SEC (kWh/kg)	2.37

.

The values of the energy consumption and energy generation and the MITA values of the turbines, expanders, and CHX-1, respectively, are significantly impacted by varying the values of the design variables. If the flow rate of the refrigerant increases, so does the energy consumption, and vice versa. It is also analyzed that the energy consumption also increases by increasing the discharge pressure of a cycle. Similarly, by increasing the suction pressure, the net energy consumption will also increase. The effect of the design variables is proportional to the net energy consumption.

3.1.2. Composite Curve (CC) Analysis

A CC analysis of the proposed process is performed. Figure 3 presents the THCC and TDCC of the CHX-1.

Figure 3a shows the significant gap between the hot and cold CCs at the hot end of the CHX-1. At the hot end, the temperature difference between the hot and cold CCs is 70 °C. This high temperature is because of the temperatures of streams 26 and 14. The temperature of stream 26 is fixed because of the cooler outlet temperature; however, stream 14’s temperature can be adjusted by adjusting the refrigerant compositions, flowrate, or pressures. By adjusting these stream conditions, the gap at the hot end can be reduced in order to minimize the entropy generation or exergy destruction. However, adjusting the temperature of stream 14 affects the MITA value at both the hot and cold ends, which leads to a negative MITA value. Similarly, Figure 4b shows the change in the approach temperature (or MITA) value along the length of the CHX-1. The higher the peak of CC, the higher the MITA value will be. This peak should be low (~1 °C). At the cold end, the value of the approach temperature is 1 °C, which shows that the cold end is efficient in terms of generating less entropy.

3.2. Exergy Analysis

The values of the exergy destruction (in kW) of the equipment are illustrated in

Table 7. Exergy destruction values of different equipment are segregated based on unit operations.

Section	Equipment	Exergy Destruction (kW)
Membrane section	Compressors	77,274.1
	Coolers	368,514.5
	Membranes	388,083.5
Desublimation section	Compressors	287.4
	Expanders	353.5
	Pumps	5.3
	Vessels	277.1
	Coolers	12.8
	CHX-1	693.2
	Cold box	0.5

, which shows the results of the performed exergy destruction analysis.

The exergy destruction in the membrane section is higher than that in the desublimation section. In this section, the coolers and membranes show the highest exergy destruction among the other unit operations in the whole process. The compressors, on the other hand, show the second-highest exergy destruction. The main reason for this high exergy destruction in the membrane section is due to the feed flow rate. The feed flow rate is high in the membrane section compared to the desublimation section because of the separated CO₂ from the membrane section. In the desublimation section, the input feed flow rate is less, which leads to a low exergy destruction. In this section, the CHX-1 is the major source of the exergy destruction, i.e., 693 kW, followed by the expanders (353.5 kW) and compressors (287 kW). The major reason for this high exergy destruction in the CHX-1 is because of the large temperature difference between the streams. Figure 4 shows the percentage of the exergy destruction of the equipment in the membrane and desublimation sections.

According to Figure 4, the membrane chamber’s exergy destruction share is 46.54%, followed by the coolers share of 44.19 %. The remaining share of 9.27% is occupied by the compressors. Similarly, in the desublimation section, the CHX-1 occupies the highest share (42.5%), followed by the expanders (21.6%) and compressors (17.6%). The exergy destruction in the CHX-1 can be further reduced by reducing the temperature difference between the streams.

3.3. Economic Analysis

Figure 5 shows the cost breakdown of the proposed process.

Figure 5 shows that the TCI of the process is 816.2 m$, whereas the TAC and TOC of the process are 460.8 m$/y and 420.7 m$/y, respectively. The TCI comprises all the equipment costs related to the membrane and desublimation sections. The share of the equipment costs in both sections is presented in Figure 6.

According to Figure 6, the major share of the TCI in the membrane section comprises the compressors (95.7%), followed by the membranes (3.05%) and coolers (1.19%). The major reason for the high TCI of the compressors is because of the high energy consumption of the compressors. In the desublimation section, the major share of the TCI comes from the compressors (67.78 %), followed by the expanders (21.03%) and CHX-1 (6.86%). The high energy consumption in the compressors is the main reason for this high TCI in the compressors.

3.4. Comparison with Conceptual Studies

In

Table 8. Presents a brief comparison of the proposed process with other conceptual studies.

Technology	Product Conditions	Energy Requirements	Cost	Ref.
Membrane–PSA	H2 purity: 99.98 % H2 recovery: 97%	N/A	N/A	[16]
Membrane–PSA	H2 purity: 99.0 % H2 recovery: 98%	Energy efficiency: 77.5%	H2 production cost: 0.6 €/kg (0.65 $/kg) *	[17]
Cryogenic–Membrane	H2 purity: 90.4% H2 recovery: 98%	Energy consumption: 36.6 MW	N/A	[27]
Cryogenic–Membrane	H2 purity: 90% H2 recovery: 98%	SEC: 0.546 GJ/ton CO2	TAC: 40.9 m$/y	[28]
This study (Membrane–Cryogenic)	H2 purity: 99% H2 recovery: 96%	SEC: 2.37 kWh/kg	TAC: 460.8 m$/y

* 1 EUR = 1.08 USD [49].

, a brief comparison of the proposed study and previous studies is presented. The proposed study is efficient in terms of a high H₂ purity, whereas the H₂ recovery is competitive with other processes. However, it is important to note that Table 8 provides a brief overview of the studies’ comparison. A detailed comparison is not possible because of the different feed and product conditions, design parameters, and process configurations.

4. Conclusions and Future Work

With the growing attention towards the H₂ economy, the production of high-purity H₂ is becoming essential. High-purity H₂ is used in many applications, such as H₂ fuel cells and the production of liquid H₂. The production of high-purity H₂ with a high recovery and low energy consumption is a challenging issue. Unlike standalone technologies, integrated processes show commitment towards high-purity H₂. The proposed integrated membrane–cryogenic process produced H₂ with a 99.99% purity and 95.9 % recovery. The net SEC of the proposed process was 2.37 kWh/kg. In the integrated process, the membrane and multi-stream heat exchanger were the major sources of the exergy destruction. The total annualized cost of the integrated process was 460.8 m$/y. This proposed integration could be advantageous for boosting a H₂ economy. The future work on the proposed process may include a detailed evaluation from an advanced exergy analysis.