Tri-Reforming of Methane: Thermodynamics, Operating Conditions, Reactor Technology and Efficiency Evaluation—A Review

Abstract

The production of syngas with optimal energy usage, a minimal environmental impact, and an adjustable H₂/CO molar ratio is possible using tri-reforming of methane (TRM). Despite the number of studies dedicated to the TRM process, this process is still in its infancy, with many technical obstacles to overcome. Except for its kinetics and catalysts, which have been reviewed elsewhere, the TRM process is evaluated thoroughly in this work. First, feasibility studies of TRM and the TRM process are presented. Second, the impacts of various operating conditions on the rate of gas conversions, syngas production, and coke formation are discussed. Third, different reactor configurations are compared. This review then goes through the energy and energetic efficiency, economic, environmental, and safety aspects of the TRM process. Finally, a research path for the future is suggested.

1. Introduction

Year after year, the world’s energy consumption rises, and fossil fuels account for most of it [1]. This significantly contributes to the increased atmospheric level of CO₂, which leads to global warming and climate change [2]. As a result, reducing CO₂ levels in the atmosphere is critical. Furthermore, the challenges of implementing alternative energy sources on a broad scale, such as renewables, necessitate a thorough examination of more efficient and clean methods for utilizing fossil energy while reducing greenhouse gas emissions. According to Goeppert et al., converting CO₂ to fuels, particularly methanol, is a very successful technique for combating both climate change and fossil fuel depletion [3]. In this regard, an innovative process known as tri-reforming of methane (TRM) was firstly proposed by Song and Pan [4]. With this technique, CO₂ obtained from industrial flue gases is converted into syngas without pre-separation [5,6,7]. Therefore, it is possible to reduce the carbon footprint of industrial processes to combat global warming while maintaining economic strength compared to the classic steam reforming of methane (SRM) [8].

According to Equations (1)–(3), this process consists of three primary reactions of methane with carbon dioxide (dry reforming of methane (DRM)), steam (steam reforming of methane (SRM)), and oxygen (partial oxidation of methane (POX)), which occur simultaneously in a catalytic reactor [9,10]:

$${\mathrm{CH}}_4+{\mathrm{CO}}_2 \leftrightarrow { 2\mathrm{CO}+2\mathrm{H}}_2 { \Delta \mathrm{H}}_{298}^0=247.3 \frac{\mathrm{kJ}}{\mathrm{mol}} \mathrm{DRM},$$

(1)

$${\mathrm{CH}}_4+{\mathrm{H}}_2\mathrm{O} \leftrightarrow { \mathrm{CO}+3\mathrm{H}}_2 { \Delta \mathrm{H}}_{298 }^0=206.3 \frac{\mathrm{kJ}}{\mathrm{mol}} \mathrm{SRM},$$

(2)

$${\mathrm{CH}}_4+\frac{1}{2}{\mathrm{O}}_2 \leftrightarrow { \mathrm{CO}+2\mathrm{H}}_2 { \Delta \mathrm{H}}_{298}^0=-35.6 \frac{\mathrm{kJ}}{\mathrm{mol}} \mathrm{POX},$$

(3)

The advantages of this combined reforming include the high economic benefit of steam reforming, the high energy efficiency of partial oxidation, and the environmental benefit of carbon dioxide reforming [11,12,13]. Furthermore, owing to the presence of O₂ and H₂O, the carbon deposition on catalyst surfaces is drastically eliminated, resulting in an extended catalyst lifetime [8,14,15]. Another advantage of TRM is the ability to change the relative volume of H₂O, O₂, and CO₂ to effectively regulate the H₂/CO ratio of products and avoid the large energy consumption associated with CO₂ separation [8,10]. As a result, power plant flue gases containing CO₂, H₂O, and O₂ can be utilized without pre-separation to be converted into syngas [14], and then this syngas may be utilized to generate compounds such as methanol, dimethyl ether, and clean energy sources such as liquid hydrocarbons [16,17,18].

Because of the stability of the reactant molecules, TRM is usually carried out at high temperature (approx. 700–900 °C) and low pressure (atmospheric pressure), as will be explained in Section 5 [19,20].

In addition to Equations (1)–(3), in TRM processes a wide range of endothermic and exothermic side reactions, such as water–gas shift reaction (WGSR), reverse water–gas shift reaction (RWGSR), and methane cracking (Equations (4)–(6), respectively) [21] may occur.

$$\mathrm{CO}+{\mathrm{H}}_2\mathrm{O} \leftrightarrow {\mathrm{CO}}_2+{\mathrm{H}}_2{ \Delta \mathrm{H}}_{298}^0=-41 \frac{\mathrm{kJ}}{\mathrm{mol}} \mathrm{WGSR},$$

(4)

$${\mathrm{H}}_2+{\mathrm{CO}}_2 \leftrightarrow \mathrm{CO}+{\mathrm{H}}_2\mathrm{O}{ \Delta \mathrm{H}}_{298}^0=+41 \frac{\mathrm{kJ}}{\mathrm{mol}} \mathrm{RWGSR},$$

(5)

$${\mathrm{CH}}_4 \leftrightarrow { 2\mathrm{H}}_2{+\mathrm{C}\left(\mathrm{s}\right) \Delta \mathrm{H}}_{298}^0=+75 \frac{\mathrm{kJ}}{\mathrm{mol}} \mathrm{Methane} \mathrm{cracking},$$

(6)

Moreover, the TRM technique requires a catalyst to guide the reaction’s kinetics and selectivity [22].

2. Scope of the Current Review

The number of works dedicated to TRM has increased in the last decade. A search on Web of Science found around 130 publications from the year 2002 up to the year 2022, using the keywords “tri-reforming” and “methane”, excluding biogas reforming (Figure 1). However, the focus of this review is on the issues of tri-reforming, which have not been surveyed before. These include the feasibility of TRM, descriptions of the TRM process, the impact of different operating conditions, sensitivity analysis, evaluation of reactor technology, energy and exergetic approaches, and the economic, environmental, safety, and efficiency evaluations of the TRM process.

The work of Minh et al. [23] is dedicated to the reaction mechanisms and catalyst deactivation and regeneration. If the design of TRM catalysts is being examined, the review by Pham et al. [24] can be studied. Amin et al. also looked into TRM from the standpoint of its method, catalysts, and kinetic mechanisms [25]. Process intensification, modeling and simulation, synthesis of catalysts with high-temperature criteria and related process corrosive conditions, and CO₂ emission control were also reviewed by Arab Aboosadi and Farhadi Yadecoury [26]. Furthermore, Soloviev et al. provide a thorough analysis of structured nickel–alumina catalysts [27]. Furthermore, the use of biogas as a raw material for the production of synthesized gas through tri-reforming has already been investigated [26,28].

3. Feasibility Studies

In addition to some theoretical calculations [30,31], some experimental studies demonstrate that TRM is feasible [4,32]. Experimental investigations have been carried out using both pure gases and flue gases of power plants for CO₂ conversion. Furthermore, since 2008, Korea Gas Company (KOGAS) has incorporated TRM technology into a DME process that can produce syngas with proper control of the H₂/CO ratio with minimal coke formation [33]. Nevertheless, these studies are in their early stages and there are still several essential technical challenges to address such as configuration of effective catalysts for industrial applications, finding the proper processing conditions and feed compositions for a desirable ratio of products, inert gas (N₂) management, design of reactor and process, process integration into present plants, heat and energy management, and efficient use of heat dissipation. Moreover, all-encompassing economic assessments and consideration of chemicals and fuel mixtures that meet the demands of the market and that can potentially expand them should be considered. As a result, these aspects require thorough research and engineering assessments to develop and implement this innovative process concept [34,35].

4. TRM Process

As mentioned before, flue gas from industries or power plants can be utilized as a co-reactant in the TRM process to produce syngas [36]. Flue gases are mostly composed of CO₂, N₂, O₂, H₂O, CO, H₂, and trace amounts of NO_X and SO_X [8]. Because of the huge amount of N₂ in flue gas, the conversion will be reduced, resulting in more harmful NO_X output. As a result, to remove N₂ from the flue gas, an N₂ separation unit such as a nitrogen-selective membrane should be used [37,38] or the TRM process can be coupled to an upstream oxy-fuel combustion power plant to alleviate the negative impact of N₂ in flue gas from a traditional air–fuel combustion power plant [39].

Input natural gas must also undergo feed pre-treatment, which are desulfurization and pre-reforming processes. Desulfurization occurs when H₂ reacts with sulfur molecules in the feed to generate H₂S, which is then removed from the primary process. This avoids catalyst poisoning in the main reactor and extends the time between catalyst replacements [40,41]. However, if natural gas which meets the pipeline specifications is used [42], the sulfur contents are negligible [41]. In the pre-reforming section, C₂⁺ hydrocarbons are removed by steam to prevent soot generation and deposition in the main reactor [10,41]. In pre-reforming, the catalyst is mostly the same as in the TRM reactor [41].

As illustrated in Figure 2, to pre-reform the gas, the desulfurized natural gas is mixed with steam and heated up to 550 °C by a multi-pass fired-heater. The pre-reformed gas then mixes with CO₂ or flue gas and passes through the fired heater a second time to reach 650 °C before being transferred to the TRM reactor. The oxygen is likewise pre-heated to 650 °C before being delivered to the TRM reactor [10,41]. The TRM reactor’s temperature is kept between 950 and 1000 °C [10,43]. Typically, the TRM reactor is working in an adiabatic catalytic system and is packed with commercial Ni-catalysts (e.g., Ni/Al₂O₃) [41]. The outlet of the reformer, which contains synthetic gas, unreacted methane, carbon dioxide and steam, is cooled to 25 °C and compressed to 35 bar to separate water [8,44].

After that, unconverted reactants are separated, recycled, mixed with the make-up stream, and fed back into the reactor at the product purification step [10,40]. The processed syngas is then compressed to a pressure of 50 bar and heated up to 220 °C if it enters a methanol reactor [44]. Given that most gas-to-liquid operations, including Fischer–Tropsch, methanol, and DME syntheses, are all carried out at pressures higher than 10 bar, it is more economical to implement TRM reactions at pressures consistent with the downstream processes [37]. Furthermore, water electrolysis [39,41,44,45] and cryogenic air separation units [45] are two common ways to supplement the TRM process with additional oxygen requirements.

5. Thermodynamic Analysis

This section discusses the effect of various operational parameters on feed gas conversions and H₂/CO ratios.

5.1. Effect of Temperature

Generally, the reaction kinetics improve as the temperature rises; therefore, the reactor volume required for the same conversion decreases [46]. Ren et al. found that when the temperature grew from 400 to 800 °C at 1 atm, the conversion of CH₄, CO₂, and H₂O increased significantly. However, a further increase in temperature led to a gradual increase in the conversion of the reactants, reaching close to 100% at 1000 °C [47,48]. When one compares Ren et al.’s experimental results with the results of ASPEN plus simulation (Figure 3) under the same condition, their findings are in conformity with the model.

Table 1. Simulation settings and conditions considered for Figure 3.

Process Conditions		Feed Gas Ratio CH4:CO2:H2O:O2	Minimization Method	Property Methods	ASPEN Technologies
Temperature	Pressure
400–1000 °C	1 bar	1:0.3:0.3:0.2	Gibbs free energy reactor	UNIFAC + STEAMNBS	ASPEN plus, V.12.0 (38.0.0.380), ASPEN Technology, Bedford, MA, USA

shows this simulation specification. The reason for higher conversions at higher temperatures is that the DRM and SRM prevail at higher reaction temperatures due to their endothermic character (Equations (1) and (2)) [4,47,49,50].

While the overall reaction enthalpy at lower temperatures up to 400 °C remains modestly exothermic, endothermicity grows rapidly as temperature rises [51]. According to Figure 4, methane oxidation reactions and WGSR, which are exothermic with negative Gibbs energy below 400 °C, are a reasonable source of the exothermicity of this process.

Depending on thermodynamics, a temperature of at least 650 °C is necessary for effective reactant conversion in methane reforming [19,49,51,52]. Moreover, at temperatures below 550 °C, the phenomena of negative CO₂ conversion can be observed due to WGSR (Equation (4) and Figure 5) [48], while at temperatures above 600 °C, RWGSR (Equation (5), Figure 5) accompanies the DRM reaction to consume CO₂ [51].

Regarding the products, increasing the temperature causes a rise in the CO mole fraction from zero at 400 °C to the maximum at about 850 °C. Meanwhile, the mole fraction of CO₂ starts to decrease and reaches zero at 850 °C. This suggests that CO₂ acts in the reactions as a limiting reagent. With increasing temperature, the molar fraction of H₂ continues to rise while the molar fraction of CH₄ continues to decrease, which confirms that there is a direct relationship between H₂ production and CH₄ consumption. Therefore, it is clear from the TRM reactions that the CH₄ intake is the main H₂ production source [53]. Following Zhang et al., the reason for H₂ production at temperatures below 400 °C is methane cracking (Equation (6)) [53]. However, this hypothesis is questionable because this reaction has positive Gibbs free energy at this temperature range (Figure 6). For the production of H₂, the POX reaction (Equation (3)), which has a negative Gibbs energy in all temperature ranges, appears to be the cause. WGSR is also likely to occur alongside the POX reaction to produce H₂ but with a lower probability (Figure 6).

Lower reaction temperatures have a significant impact on the H₂/CO molar ratios. The H₂/CO molar ratio might drop dramatically, for instance, from 11.8 to 2.1, when the temperature rises from 400 to 600 °C, regardless of the feed molar ratios. Further raising the reaction temperature to 1000 °C results in a slight decrease in H₂/CO molar ratio (for example, around 2.0), which is influenced by the composition of CH₄, CO₂, H₂O, and O₂ [4,47,52,54,55]. The reason is that above 650 °C, the RWGSR is predominant and consumes more H₂ and produces CO. In addition, the conversion of CO₂ improves; thus, it lowers the ratio of H₂/CO [19,50].

Based on the findings stated above, a high temperature along with low pressure is desirable to achieve high CO₂ conversion and H₂ yields. However, more energy demand is needed for a higher temperature reaction, which will inevitably increase the cost of the TRM operation, while at low pressures, temperatures above 850 °C do not provide significant benefits regarding the syngas ratio and conversions. Given the above consequences, the ideal reaction temperature is considered to be lower than 850 °C [51,56,57].

5.2. Effect of Pressure

Pressure is a crucial factor in setting the equilibrium state during TRM and is one of the primary variables that control the mole fractions of products [19]. The application of Le Chatelier’s principle to Equations (1)–(3) demonstrates that synthesis gas output is reduced at higher pressures [20]. Increasing the pressure up to 20 bar contributes to a drop in the H₂ mole fraction and an increase in the CH₄ mole fraction even at 850 °C. Regarding CO₂ and CO, the influence of pressure at temperatures below 450 °C is negligible, which means that DRM is not feasible in this temperature range. If the temperature rises above 450 °C, the rise in pressure leads to continuous growth in the concentration of CO₂ and a simultaneous decline in the concentration of CO [47]. Similarly, with increasing pressure, the molar fraction of H₂O increases constantly [19]. On the other hand, pressure has a minor influence on O₂ conversion [47]. The results of thermodynamic equilibrium concentrations and conversions simulated with ASPEN plus, which are presented in Figure 7 and Figure 8, respectively, also confirm these findings from the literature. The ASPEN plus simulation conditions and settings are presented in

Table 2. Simulation settings and conditions considered for Figure 7 and Figure 8.

Process Conditions		Feed Gas Ratio CH4:CO2:H2O:O2	Minimization Method	Property Methods	ASPEN Technologies
Temperature	Pressure
0–850 °C	1, 3, 5, 7, 10 bar	1:0.3:0.3:0.2	Gibbs free energy reactor	UNIFAC + STEAMNBS	ASPEN plus, V.12.0 (38.0.0.380), ASPEN Technology, Bedford, MA, USA

.

Based on the simulation results, and according to the minimum required temperature for O₂, CO₂, and H₂O conversions, one can conclude that the tendency of CH₄ to react with other reactants is as follows: O₂ > H₂O > CO₂, which is explainable by their Gibbs free energies (Figure 9).

Figure 10 shows the calculated equilibrium conversions depending on pressure for TRM. This calculation was performed by ASPEN plus and minimized the Gibbs free energy under conditions identical to the experimental conditions used by Ren et al. [47]. It is clear that the conversion of CH₄ is higher than that of CO₂ and H₂O at different pressure values [47]. In addition, in any pressure range, CO₂ conversion in TRM is much lower than in DRM, suggesting that the DRM reaction is hindered when O₂ and/or H₂O are present in the reaction environment [47,51].

When raising the pressure from 1 to 20 bar, the H₂/CO molar ratio increases with pressure up to 800 °C and then drops slightly to reach a steady value (i.e., 1.5), demonstrating that pressure has a more limited role in regulating H₂/CO molar ratios at higher temperatures [51]. This trend can be verified by ASPEN plus simulation as depicted in Figure 11 (the settings of simulation are according to Table 2). This observation is consistent with the study by Ren et al., which demonstrated that at 750 °C, the decrease in H₂/CO ratio was very small with increasing pressure (Figure 10) [47].

In conclusion, relatively low pressures are optimal for CO- and H₂-rich production and CO₂ conversion optimization [19,20,53,55]. However, the effect of pressure on conversions and product ratio is marginal when the temperature is higher than 1000 °C [47,51].

5.3. Effect of CH₄/Flue Gas Ratio

Since CH₄ reacts with all the active agents (CO₂, O₂, and H₂O), the CH₄/flue gas ratio has a noteworthy role in TRM reactions, especially in the formation of CO and H₂. However, the effect of the reaction temperature has to be considered in parallel. When the temperature is below 850 °C, higher CH₄/flue gas ratios increase the molar flow rate of H₂. For temperatures above 850 °C, H₂ molar flow follows a weak correlation of the CH₄/flue gas ratio, and the production of H₂ is also maximal under every CH₄/flue gas ratio. The CH₄/flue gas ratio has a similar effect on CO production as for H₂. However, the influence of the CH₄/flue gas ratio on the CO₂ molar flow is different. A rise in the ratio of CH₄/flue gas causes an increase in the molar flow of CO₂ at temperatures below 550 °C, while it causes a decrease in the molar flow of CO₂ at temperatures higher than 550 °C. When the reaction temperature exceeds 850 °C, the CO₂ content approaches zero, if the CH₄/flue gas ratio is between 0.4 and 1.0, indicating that CO₂ is the limiting reactant that inhibits additional syngas production. Therefore, when the operating temperature is higher than 850 °C, lower CH₄/flue gas ratios can be chosen to reduce CH₄ consumption [53].

5.4. Effect of O₂/CH₄ Ratio

By increasing the O₂/CH₄ ratio, CH₄ conversion becomes higher, mainly at temperatures lower than 850 °C, because CH₄ is the limiting reactant [19]. However, the rise in the O₂/CH₄ ratio results in a substantial drop in CO₂ conversion across the entire temperature spectrum [19,45,50,58,59] because the reaction between CH₄ and O₂ is thermodynamically preferred over the reaction between CO₂ and CH₄ [50,59,60] (Figure 9). The combustion of CH₄ (Equation (7)) overtakes the DRM reaction and thereby a net rise in the amount of CO₂ is created [50,61]. Similarly, with increasing O₂ content, H₂O conversion declines [62]. However, the H₂O conversion reduces more significantly compared to CO₂ conversion [31], which results in a lower H₂ yield, and consequently a lower H₂/CO ratio [19,31]. In other words, CH₄ is likely to be combusted to H₂O and CO₂ instead of H₂ and CO at a comparatively high O₂ concentration [19]. Under these conditions, the amount of CH₄ available for DRM and SRM is quite limited [61].

$${\mathrm{CH}}_4+{2\mathrm{O}}_2 \leftrightarrow { \mathrm{CO}}_2+{2\mathrm{H}}_2{\mathrm{O} \Delta \mathrm{H}}_{298}^0=-803 \frac{\mathrm{kJ}}{\mathrm{mol}} \mathrm{Full} \mathrm{combustion} \mathrm{of} \mathrm{methane},$$

(7)

When the O₂/CH₄ ratio is between 0.45 and 0.50 for varied H₂O/CH₄ ratios, the input CH₄ reacts completely [63]. The complete oxidation of methane releases large amounts of energy that provides heat for endothermic reactions and contributes to the thermal equilibrium of the process [50,60].

However, regarding the effect of the O₂/CH₄ ratio on the syngas ratio (H₂/CO), there are two opposing findings. Some references mention that as the O₂/CH₄ ratio rises, the syngas ratio drops because the higher O₂/CH₄ ratio in addition to higher reaction temperatures not only increases DMR but also accelerates the RWGSR, boosting CO generation [55,64]. However, others report that oxygen has the same effect as steam on the H₂/CO ratio. Increasing the O₂/CH₄ ratio results in a higher H₂/CO molar ratio because the higher the oxygen content, the less methane is accessible in the reaction environment. As a result, SRM would replace DRM, resulting in a higher H₂/CO molar ratio [4,65]. These arguments appear to be inaccurate. In [4], authors changed the other feed gas ratios at the same time as oxygen, and in [64], in addition to using a membrane reactor to disperse O₂, the H₂/CO ratio was practically constant (1.57 and 1.58). Furthermore, according to Yang et al., the range of the O₂/CH₄ ratio is essential because when this ratio is between 0.1 and 0.4, the H₂ yield is increased by the O₂/CH₄ ratio, and when the O₂/CH₄ ratio is greater than 0.4, the H₂ yield decreases gradually [63]. The reason is that a high O₂/CH₄ ratio will result in a hydrogen combustion reaction [64,66]. Thus, the conclusions of references [55,65] contradict the conclusions of Yang et al., as [65] studied a lower range of O₂/CH₄ ratios (0.01–0.29) and [55] worked in a wider range of O₂/CH₄ ratios, from 0.2 to 0.75.

In addition, thermodynamic studies have revealed that to achieve high synthesis gas efficiency and avoid coke formation, adiabatic processes require a considerable oxygen input (about half of the methane feed) [64] but for non-adiabatic processes with stable methane reforming, a O₂/CH₄ ratio of 0.25 is required [67]. It is also evident that increasing the O₂ concentration reduces not only the coke formation but also the process’s energy consumption [51]. Moreover, it is also advised to utilize pure oxygen with a low pressure rather than air with a high pressure to decrease its impact on the reactor volume [46].

5.5. Effect of H₂O/CH₄ Ratio

With a rise in H₂O up to a specific concentration at low temperatures, CH₄ conversion falls [19,59,68], and it converts to a growing pattern after reaching an inflection point at relatively high temperatures. However, above 850 °C, a conversion of CH₄ of almost 100% is obtained, independent of the H₂O content [19].

There is clear evidence that the SRM outweighs DRM with increasing H₂O concentration [50]. The growth in H₂O content causes a substantial reduction in CO₂ conversion [59] because, for TRM, CO₂ and H₂O are both oxidizing agents [37,69], but CH₄ is thermodynamically more likely to react with H₂O [19,25,65]. Moreover, both SRM and WGSR are preferred at a higher molar ratio of H₂O/CH₄, which again lowers CO₂ conversion [16,46,59,70]. High H₂O/CH₄ ratios, on the other hand, promote methane conversion. The amount of methane engaged in SRM rises to more than 50% as the H₂O/CH₄ ratio rises, whereas the DRM falls to 7% [71].

The effect of H₂O content on H₂O conversion depends on temperature. When the temperature exceeds 650 °C, the H₂O conversion rate falls as the H₂O/CH₄ molar ratio rises. This is due to a lack of CH₄ available to react with H₂O. Furthermore, above 650 °C, RWGSR also dominates, leading to lower H₂O conversion rates and, consequently, lower H₂ yields and H₂/CO ratios [19]. Nevertheless, the conversion rate of H₂O rises with an increase in the H₂O/CH₄ molar ratio while the temperature is below 650 °C [19]. However, there is no adequate justification for this phenomenon in the literature. For instance, according to Zhang et al. and Singha et al., both WGSR and SRM reactions are equally important below 650 °C [19,59], consuming steam. However, this cannot be the case for SRM because it has a positive Gibbs free energy in this temperature range (Figure 12). Therefore, only WGSR can happen, and an increase in H₂O, which favors WGSR, drives this reaction to consume more H₂O.

Another point is that increases in feed gas content of H₂O can boost H₂ yield while decreasing CO yield [19,55,64,65]. Therefore, the H₂/CO molar ratio increases as the H₂O/CH₄ molar ratio increases [18,20,31,33]. However, based on the amount of H₂O brought in, the ratio between the DRM and SRM reactions can be influenced, and thus the H₂/CO ratio is highly adjustable [50,68]. Moreover, excessive steam might deactivate the catalyst by oxidizing its surface [16]. Furthermore, adding steam, unlike oxygen, increases the required energy of the process [51].

5.6. Effect of CO₂/CH₄ Ratio

CH₄ conversion improves by increasing the CO₂/CH₄ ratio, while CO₂ conversion declines [19,72]. However, the CO₂/CH₄ ratio has a far smaller impact on the methane conversion than the O₂/CH₄ ratio [63].

With different CO₂/CH₄ ratios at low temperatures, the effect of the CO₂/CH₄ ratio on CO and H₂ yields tends to be marginal. However, the H₂ yields decrease considerably when CO₂ is added at temperatures above 700 °C [19]. The foremost reason is that a higher CO₂/CH₄ ratio induces RWGSR to consume H₂ and produces more CO, leading to a reduction in the H₂/CO ratio [19,62,63,73]. Likewise, as the CO₂ level climbs, the amount of H₂O produced by RWGSR decreases. There is, in fact, a trade-off between maximal CO₂ conversion and syngas ratios of 2 or more. In other words, a lower synthesis gas ratio is available for higher CO₂ conversion [64].

The ideal CO₂/CH₄ ratio is 1.43; at lower ratios more methane is available, which cannot react efficiently [66].

5.7. Effect of O₂/CO₂ Ratio

When the O₂/CO₂ ratio increases from 0 to 1.5, while the CH₄ and H₂O molar ratios are constant, the average CO₂ conversion drops, for example, from 60 to 33% at atmospheric pressure and 750 °C [52]. The decline in CO₂ conversion is due to a higher amount of O₂ in the feed, which inhibits DRM and raises methane oxidation. Therefore, the higher O₂/CO₂ ratio produces syngas that is richer in H₂ as a consequence of the reduced CO₂ conversion, causing the higher H₂/CO ratios [50,52,74].

5.8. Effect of H₂O/CO₂ Ratio

Raising the H₂O/CO₂ ratio in the feed stream promotes SRM over DRM, resulting in a rise in the H₂/CO ratio, lowering CO₂ conversions. For instance, at atmospheric pressure and 750 °C, while the other gas molar ratios are constant, the CO₂ conversion falls from 73 to 30% [52,66]. When there is no rivalry between CO₂ and H₂O for adsorption sites in the absence of water vapor, the maximum CO₂ conversion is attained [49].

Table 3. Effect of various operational parameters on feed gas conversions and H₂/CO ratios.

	CO2	CH4	H2O	H2/CO	C Formation	Optimum Ratio	Comments
Temperature	D 1	D	D	I2	I	800 °C
Pressure	I	I	I	D	I or D (temperature-dependent)	1 bar	At low temperatures, carbon deposition reduces as pressure rises, but at temperatures above 600 °C, it increases considerably with pressure.
CH4/flue gas	I or D (temperature-dependent)					Between 0.4 and 1.0	When the temperature is lower than 550 °C, CO2 molar flow increases by increasing the CH4/flue gas ratio but after that it decreases.
O2/CH4	I	D	I	I or D (depending on O2/CH4 ratio)	I	Between 0.45 and 0.50	When this ratio is between 0.1 and 0.4, the H2 yield is increased by the O2/CH4 ratio, and when the O2/CH4 ratio is greater than 0.4, the H2 yield gradually decreases.
H2O/CH4	I	I or D (temperature-dependent)	I or D (temperature-dependent)	I or D (temperature-dependent)	I	1	When the temperature exceeds 650 °C, the H2O conversion and H2/CO ratio decrease but when the temperature drops below 650 °C, the opposite is true.
CO2/CH4	I	D	I	I		1.43
O2/CO2	I			D	I	0.17 for a high CO2 conversion
H2O/CO2	I			D	I

¹ D: Directly proportional. I: Inversely proportional.

provides a summary of the conclusions from these Section 5.1, Section 5.2, Section 5.3, Section 5.4, Section 5.5, Section 5.6, Section 5.7 and Section 5.8.

6. Optimal Operating Conditions and Feed Compositions

As mentioned previously, TRM is very sensitive to reaction conditions, specifically temperature and feed composition. In the TRM operation, the pressure, temperature, and feed ratios (H₂O/CH₄, O₂/CH₄, and CO₂/CH₄) are independent variables influencing the overall efficiency. These variables affect the thermodynamic equilibrium, the kinetics of the related reactions, and the resulting H₂/CO ratio [59]. However, the required energy of the system and CO₂ conversion is more sensitive to the feed composition variation rather than temperature. The effect of a 1% fluctuation in operating temperature (considering an industrial setting) on both energy consumption and CO₂ conversion is minimal, while a 1% change in feed composition resulted in a nearly 5% increase in both energy demands and CO₂ conversion [51].

Regarding the optimum feed composition, several feed ratios have been considered optimum in the literature, which are summarized in

Table 4. Optimum process conditions.

Process Conditions		Optimum Feed Composition CH4:CO2:H2O:O2	Conversions	H2/CO	Ref.
Temperature	Pressure
850 °C	1 atm	1:0.291:0.576:0.088	CO2 > 90%	2.0	[19]
850 °C	1 atm	1:0.282:0.574:0.1	CO2 > 90% CH4 > 90%	2.0	[36]
1000 °C	30 bar	1:0.2:0.35:0.48	CO2 = 50%	2.32	[75]
750 °C	1 bar	1:1:0.4:0.3	CO2 = 47.8%	-	[51]

. For example, Zhang et al. discovered that the best feed ratio was CH₄:CO₂:H₂O:O₂ = 1:0.291:0.576:0.088, which resulted in a 94.5% H₂ yield, with the H₂/CO ratio of 2.0 and the CO₂ conversion more than 90% [19]. Jarungthammachote observed the best feed ratio as CH₄:CO₂:H₂O:O₂ = 1:0.282:0.574:0.1, leading to the H₂ yield of 94.94%, H₂/CO ratio of 2.0, CO₂ and CH₄ conversions ≥ 90 [36]. These two groups of researchers used nearly identical feed ratios and obtained almost similar results at the reaction temperature and pressure of 850 °C and 1 atm. On the other hand, Rezaei et al. determined that the optimum feed composition was CH₄:CO₂:H₂O:O₂ = 1:0.2:0.35:0.48 to achieve the CO₂ conversion of 50% at 30 bar and 1000 °C. They also concluded that the feed should have a CO₂/CH₄ ratio of 0.1–0.2, an O₂/CH₄ ratio of 0.4–0.5, and H₂O/CH₄ ratios of 0.25–0.5 [75]. Moreover, Jarungthammachote observed that as the O₂ concentration of the co-reactant increased while the conversions and yields remained high, the optimal CO₂/CH₄ and H₂O/CH₄ ratios decreased. However, the maximum H₂ yield occurred beyond the optimum CO₂/CH₄ and H₂O/CH₄ ratios, where the CO₂ conversion was less than 90% [36]. Furthermore, Challiwala et al. discovered that the best working conditions are about 750 °C and 1 bar, with a feed mole ratio of CH₄:CO₂:H₂O:O₂ = 1:1:0.4:0.3. Under this condition, the process energy consumption is minimized and the CO₂ conversion is 47.84%, while carbon deposition is suppressed. This temperature was chosen as the ideal value since it benefits from both RWGSR and DRM reactions at the same time, whereas at high temperatures of about 900 °C, RWGSR has a detrimental influence on both the energy required for the process and the H₂/CO ratio of syngas [51]. The combined effect of low pressure and high temperature is favorable for the TRM process efficiency [55]. However, due to the catalyst strength limits and higher energy consumption, the temperature does not exceed the ceiling value. According to the thermodynamic equilibrium, the maximum temperature and pressure are 850 °C and 1 atm., respectively, to achieve more than 90% CO₂ conversion, a H₂/CO ratio of equal to 2.0, and to prevent solid carbon formation during reactions [19]. As a result, the appropriate feed composition depends on both the operating parameters of the process and the kind of catalysts utilized. Additionally, the desired condition (energy usage, coke avoidance, etc.) altered these best values.

7. Coke Formation Assessment

One of the significant problems faced by industrial reforming processes is the production of solid carbon (coke), which leads to catalyst deactivation [64]. In the reforming processes, coke forms through four reactions (Equations (6), and (8)–(10)). They have a comparatively low equilibrium constant, making them very sensitive to variations in reactant molar ratios. At low temperatures (<700 °C), the Boudouard and Beggs reactions (Equation (8) and Equation (9), respectively) and the hydrogenation of CO₂ (Equation (10)) dominate, but at high temperatures, CH₄ largely decomposes to form carbon deposits (Equation (6)) [51,67,76] (Figure 13).

$$2\mathrm{CO} \leftrightarrow { \mathrm{CO}}_2+\mathrm{C}(\mathrm{s}) {\Delta \mathrm{H}}_{298}^0=-172 \frac{\mathrm{kJ}}{\mathrm{mol}} \mathrm{Boudouard} \mathrm{reaction},$$

(8)

$$\mathrm{CO}+{\mathrm{H}}_2 \leftrightarrow \mathrm{C}(\mathrm{s})+{\mathrm{H}}_2\mathrm{O} {\Delta \mathrm{H}}_{298}^0=-131 \frac{\mathrm{kJ}}{\mathrm{mol}} \mathrm{Beggs} \mathrm{reaction},$$

(9)

$${\mathrm{CO}}_2+{2\mathrm{H}}_2 \leftrightarrow \mathrm{C}(\mathrm{s})+2{\mathrm{H}}_2\mathrm{O}{ \Delta \mathrm{H}}_{298}^0=-90 \frac{\mathrm{kJ}}{\mathrm{mol}}{ \mathrm{CO}}_2 \mathrm{hydrogenation} \mathrm{reaction},$$

(10)

Therefore, according to these reactions, the amount of synthesis gas that can be produced is negatively impacted by carbon formation. As a result, preventing the formation of carbon is highly preferred due to these detrimental impacts [77].

In order to comprehend how the development of coke is lessened, it is advisable to examine the C-H-O atomic phase diagram (Figure 14). A ternary C-H-O phase diagram depicts possible phases and their equilibria for a mixture of three components of C, H, and O under constant temperature and pressure conditions. With the help of this technique, an equilibrium line (boundary line) is drawn for any desired pair of temperature and pressure, which separates the graph into its upper and lower regions. The region above this line is the area where carbon formation is thermodynamically favored, and the region below it is the area where carbon formation is not favored. Moreover, each point on the graph indicates a mixture with the specific composition of atomic species of C, H, and O (e.g., CO is a mixture of 50% C and 50% O).

Thus, knowing the reaction temperature and pressure, as well as the ratio of C, H, and O atoms, allows one to determine whether or not carbon is formed for that specific mixture. Figure 14a shows the carbon deposition equilibrium lines for several temperatures at 1 bar, and Figure 14b shows the lines for two temperatures at three different pressure levels. Small circles on these images represent the main species, namely CO, CO₂, H₂O, CH₄, and the stoichiometry ratio of feed gases (St., CH₄:CO₂:H₂O:O₂ = 1:0.5:0.25:0.125).

As illustrated in Figure 14a, by raising the temperature while maintaining the same pressure, the carbon formation equilibrium lines shifted upward, expanding the coke-free zone. The reason is that as the reaction temperature rises, the chemical equilibrium of reactions of Equations (8)–(10) shifts toward consuming carbon to produce carbon monoxide or carbon dioxide (reverse direction of these reactions) [49].

In contrast, the equilibrium line shifts downward, and the coke-free zone shrinks when pressure rises at a constant temperature. For higher temperatures, this change is less pronounced. As shown in Figure 14b, the coke-free zone does not change significantly at 1000 °C since the equilibrium lines are adjacent to one another. Furthermore, as evident from these diagrams, carbon formation is always conceivable in a TRM process when the feed gas ratio is stoichiometric (point St.) and is not affected by temperature. The stoichiometric point must, therefore, move lower or to the right of the relevant equilibrium line in order to prevent carbon production. This suggests that the content of O₂, H₂O, and/or CO₂ should be increased. However, this results in a reduction of the heating value of the produced syngas [77,78]. Moreover, oxygen and steam can not only reduce the formation of coke, but also remove it through Equation (11) and the reverse reactions of Equations (9) and (10), which also explain why CO₂ conversion decreases by adding O₂ in the reactor feed [52].

$${\mathrm{C}+\mathrm{O}}_2 \leftrightarrow {\mathrm{CO}}_2{ \Delta \mathrm{H}}_{298}^0=-393.7 \frac{\mathrm{kJ}}{\mathrm{mol}},$$

(11)

Compared to increasing steam concentration, increasing oxygen concentration may be more successful in minimizing coke formation for industrial reforming processes because the conversion rate of the reaction between carbon and oxygen is faster and more efficient even at low temperatures (Figure 15), which makes it economically desirable [79]. However, due to the exothermic nature of the reaction with oxygen, overheating is a distinct possibility [80], making precise temperature control important. However, this can be achieved by controlling the partial pressure of oxygen.

What is more, in comparison to individual SRM, DRM, and POX processes, or a combination of SRM and DRM (bi-reforming of methane, BRM), the inclusion of three oxidants (CO₂, H₂O, and O₂) in the TRM process decreases carbon deposition considerably [51] as shown in Figure 16.

Another point is that the influence of reaction kinetics and the type of the catalyst on carbon formation should be paid attention to. In other words, in the carbon-free zone, carbon can develop due to kinetically restricted reactions or can be prevented by a proper catalyst type even when operating in the carbon formation zone [77].

8. Reactor Technology Evaluation

The most common reactor configuration used to investigate TRM in the literature is a fixed-bed reactor operating under adiabatic conditions. These reactors have smaller sizes and cheaper capital and operational expenses. Moreover, they surpass the non-adiabatic TRM processes at high pressure because, in adiabatic conditions when the pressure rises, the temperature also increases and can compensate for the effect of high pressure. However, these reactors demand high oxygen to maintain the temperatures, which increases the occurrence of hot spots [81].

Khajeh et al., on the other hand, demonstrated that employing a fluidized-bed tri-reformer instead of a fixed-bed reactor offers some advantages through a mathematical comparison. A fluidized bed aids in better mixing, improving heat transfer, and reducing the temperature differential in the catalyst bed, which results in a smooth and uniform distribution of heat. Endothermic reactions benefit from this proper heat distribution. Thus, the conversion of methane and CO₂ and the yield of H₂ in a fluidized bed reactor are greater. Furthermore, in a fluidized bed, the combustion of methane is faster than in a fixed bed, which leads to more methane consumption, while the amount of oxygen required to achieve maximum H₂ yield is lower and saves more energy. Furthermore, improved heat regulation throughout the reactor lowers the hot spot temperature by around 250 °C. In addition, fluidization reduces pressure drop and diffusion constraints. From this point of view, reactors with fluidized beds appear to be more promising [82]. However, catalyst attrition is a problem for these types of reactors. Therefore, the system requires a catalyst with sufficient mechanical stability to prevent particle elution and attrition [83,84,85].

Furthermore, there is another reactor layout that aids in lowering hot spot temperatures. In this way, the temperature profile of the reactor may be controlled by delivering oxygen into the reactor from multiple locations. Rezaei et al., for example, found that dosing oxygen at four evenly spaced points along the length of the reactor keeps the reactor’s maximum temperature below 1300 °C, whereas when only one injection point is available, the reactor temperature rises to 1879 °C near the inlet point. This rapid increase in reactor temperature around the injection points is due to the exothermic reaction of methane combustion. Then, the temperature steadily drops as the endothermic reforming reactions consume this heat content. However, when numerous injection points are employed, the temperature of the highest peak is suppressed because the distribution of oxygen across the catalytic section leads to the stepwise release of heat from the oxidation of methane. As a result, having more injection points results in fewer temperature peaks at the reactor entrance [75]. Farsi et al. also reported similar data and reached the same conclusion [86].

It is also feasible to boost synthetic gas production by distributing all feed components across more than one injection point. In this regard, a mathematical model developed by Fekri Larry et al. revealed that dividing a typically fixed bed reactor into three equal sections and distributing preheated steam and carbon dioxide among these sections results in more hydrogen production, less oxygen consumption, and an acceptable H₂/CO ratio in the final product. In this modified reactor, methane is given directly to the reactor, whereas oxygen is added to the first section, and carbon dioxide and steam are distributed between the sections. Injecting steam and carbon dioxide into the second and third beds enables the reversible reactions, which are at equilibrium conditions, to progress toward higher hydrogen production. When compared to traditional designs, this concept has the potential to increase hydrogen generation by 2.93% [40].

The idea of multiple injection points can lead one to think about membrane reactors as an alternative to conventional reformers. In this regard, Rahimpour et al. studied a permeable membrane reactor. They suggested that a multi-tubular fixed bed reactor would benefit from selective permeable oxygen and hydrogen membranes. This reactor was made up of two concentric pipes, similar to a tube-shell system (Figure 17). The inner tube walls were dense oxygen permeable membranes (perovskite ceramic membrane), and the outer porous stainless-steel tube supported a dense Pd-based membrane coating. Co-currently, the air is delivered into the oxygen membrane to permeate into the reaction side, and then the generated hydrogen penetrates through the Pd-based membrane on the opposite side of the reaction zone to be swept by a sweeping gas. The benefits of this reactor include the generation of pure hydrogen and the equivalence of the H₂/CO ratio for methanol production at lower inlet temperatures. In addition, the oxygen membrane omitted the air separation requirements, reducing the need for an external pure oxygen source and resulting in considerable cost savings. Moreover, the oxygen membrane distributes oxygen and, therefore, the heat along the reformer, resulting in an optimal temperature profile within the reactor [87].

Although perm-selective membranes increase control of the TRM process’s kinetics and temperature profile, substantial permeation fluxes and harsh working conditions favor the use of porous membranes over perm-selective membranes [64]. Jardim et al. investigated a porous membrane reactor to determine the optimal oxygen distribution for adiabatic operation. They created a model of a packed bed tube with a porous membrane wall that was inert. From the shell (exterior compartment), oxygen was delivered into the feed stream as well as the α-alumina membrane. They discovered that while the oxygen partition does not affect the overall performance of the reactor, it has a considerable impact on the thermal behavior. In addition, the percentage of oxygen delivered at the feed to that permeating via porous membrane should be 80% to avoid hotspots. The temperature profile showed no hot spot but rather a cold spot (725 °C) in the inlet region due to the rapid reforming processes, followed by a smooth rise in the equilibrium temperature (953 °C). As a result, the porous membrane reactor can execute TRM adiabatically and safely with a large amount of oxygen, but it imposes a non-uniform oxygen distribution that rises with reactor length. Furthermore, due to the low pressure of methane in the outlet zone, over-oxidation may occur. To prevent an undesirable flow of oxygen, the pressure drop, the transmembrane pressure, and the permeability of the membrane should all be adequately constructed and managed [64].

Further, Khademi et al. demonstrated that the feed gas distribution to the reactor influences the reactor sizing. They looked at three different types of micro-porous membrane reactors as well as conventional ones. These reactors are distinguished by their feed injection techniques. In each proposed membrane reactor, only one of the reactants (O₂, H₂O, or CO₂) is distributed along the reaction zone across the membrane. When O₂, H₂O, or CO₂ are distributed, the reactors are called O-TRM, C-TRM, or H-TRM, respectively, as shown in Figure 18b–d [88].

These authors also developed a mathematical model of the required minimum length of each reactor to produce syngas with desired H₂/CO ratio, which varies in the range of 1.7–2.3 in terms of acceptable CH₄ and CO₂ conversion. Their results indicate that the minimum required lengths are in this order: TRM < C-TRM < O-TRM < H-TRM. As a result, the conventional reactor has the shortest length, implying that the reforming reactions reach the equilibrium more quickly, whereas the membrane reactors require more length to achieve the desired behavior.

Table 5. Different types of micro-porous membrane reactors and the reason why they should be longer compared to conventional reactors [88].

Type of Membrane Reactor	Distributed Gas through the Membrane	Reason to Be Longer for Desired H2/CO Ratio
O-TRM	O2	To provide heat for reforming processes
C-TRM	CO2	To accommodate the consumption of CO2 diffused through the catalytic bed
H-TRM	H2O	To achieve a favorable H2/CO ratio

shows why they should be longer [88].

The literature also shows that for all designs, increasing the inlet gas temperature reduces the reactor length by speeding up the rate of conversions [46,88]. However, there is an optimum point for the O₂/CH₄ ratio. On the one hand, as the O₂/CH₄ ratio increases, the length of the reactor decreases since more heat is generated, which activates reforming reactions, and also more methane is consumed. Increasing the O₂/CH₄ ratio, on the other hand, results in more CO₂ production via methane combustion, requiring a longer reactor to convert the produced CO₂. The size of the reactor remains nearly constant at low H₂O/CH₄ ratios; however, at high H₂O/CH₄ ratios, more length of the reactor is required to achieve satisfactory CO₂ conversion due to a drop in the RWGSR rate [88].

Despite the benefits of membrane reactors, they cannot prevent coke formation because the oxygen distribution via membrane technology reduces the concentration of oxygen, whereas an O₂/CH₄ ratio of more than 45% is required to avoid coke formation and achieve high yields [81].

Another type of reactor is the autothermal plug flow reactor, which has received less attention in the literature. An autothermal reactor consists of two chambers: a combustion chamber (homogeneous section) and a catalytic chamber (heterogeneous section). The combustion chamber is a special refractory chamber that can resist the high temperatures generated by methane oxidation and prepare the heat required for the reforming processes. In the catalytic chamber, reforming reactions are performed in the presence of a catalyst. This concept was explored for the DME plant’s synthetic gases, in which lower syngas ratios are adequate [10]. When designing a plug flow reactor, the length of the homogeneous and heterogeneous sections, as well as operational variables, are critical. A major criterion for constant inlet temperature and composition is that the homogenous section is long enough to have the required residence time for combustion, leading to an increase in the temperature needed for the reforming reactions. If the oxygen conversion is greater than 99.9%, the homogeneous segment is regarded as long enough. The needed residence time is, however, greatly influenced by the inlet temperature and feed composition. By increasing the inlet temperature, the ignition is accelerated and the needed dwell time is reduced. The effect of the temperature is exponential and is more intense when there is no hydrogen in the input. Furthermore, the necessary residence time is strongly influenced by the flow rate of oxygen. With more oxygen, the ignition occurs more quickly, requiring a shorter residence period. Moreover, the flow of steam or carbon dioxide at the reactor inlet increases the required residence time linearly as they simply dilute the reactants and do not contribute to the combustion reaction. Additionally, the findings of verified mathematical models reveal that the dependency of hydrogen generation on the homogeneity to overall length ratio is stronger than the dependency of carbon monoxide generation. The ideal ratio for hydrogen generation is approximately 0.37, whereas the optimal ratio for carbon monoxide production is around 0.4, and the ratio of H₂/CO falls monotonically as the length ratio increases. These findings make it possible to adjust the reactor length according to the desired products ratio [10]. This design reduces catalyst degradation, but due to negative CO₂ conversion it performs poorly when generating synthetic gases with an H₂/CO = 2.0 ratio required [81].

The next less attractive TRM reactor design in the literature is the heated reactor. In a heated reactor, which resembles the industrial SRM design, the required heat is supplied externally in the arrangement of bundles of tubes within a furnace. The performance of this reactor is superior to the autothermal reactors at a high desired syngas ratio for methanol production and adiabatic reactors at low-pressure ranges (below 40 bar). However, considering the total CO₂ emission including furnace, its CO₂ conversion is lower [81].

In other work, Ahmed et al. compared the performance of methane reforming in a spherical reactor with a cylindrical one using a multi-objective optimization (MOO) model. The notable result was that the spherical reactor offers substantially lower reactor pressure loss and consequently about 70% less power dissipation for the same amounts of H₂ production and CO₂ emission. As a result, spherical reactors outperform tubular reactors in terms of energy savings [89].

9. Energetic and Exergetic Considerations

In this section, the benefits of integrating the TRM process with upstream and downstream units are first examined, and then the effects of various parameters on the exergy of the system are discussed. The main results are presented in Table 6 and Table 7.

In the TRM process, unlike other CO₂ mitigation techniques, owing to the possibility of simultaneously feeding CO₂ with other inlet gases, a CO₂ separation step before its recovery and utilization can be omitted. As a result, this reduces the necessary energy consumption by CO₂ capturing. For example, CO₂ capturing using chemical absorption with alkanolamines consumes 20% of the energy in a coal-fired power plant [90,91]. Additionally, a calculation-based comparative study revealed that TRM is preferred over other reforming methods (i.e., DRM and SRM) in terms of energy efficiency [4,63]. According to the standard enthalpies, the TRM reactions use around 29% less energy compared to the BRM reactions [67] as the heat generated during exothermic reactions (i.e., partial combustion of CH₄ and the WGSR) can supply the needed energy for reforming reactions [4,67,92].

The work of Halmannand and Stenfield showed that the TRM of the flue gases released from coal- and gas-fired burners of power stations containing N₂ leads to fuel saving in comparison to conventional methanol or H₂ production processes with SRM. Fuel-saving for the methanol and H₂ production via TRM would be of the order of 31% and 75%, respectively, for both flue gases. Moreover, TRM results in high exergy efficiency for methanol and H₂ production for these two flue gases of around 72% and 74%, respectively [93].

In another study, Minutillo and Perna investigated how a TRM process may be combined with a power plant and a methanol synthesis process to generate both electricity and methanol. They proposed using a heat recovery system to recover the thermal energy from the produced syngas and to use it to generate high-pressure steam that could be sent to the power island to boost electrical output. According to the results of a numerical approach, the thermal efficiency and the system efficiency can exceed 83% and 55%, respectively [94].

Compared to a traditional coal gasification plant for methanol production (CTM), the combination with a coke oven gas plant (COG) coupled with TRM can increase the carbon utilization and energy efficiency of the process by 4.3% and 11.4%, respectively. In addition, if the WGS unit is omitted in the integrated process, these efficiencies become 14.5% and 16.8% higher than those of the conventional CTM process, respectively [95].

Likewise, when a TRM process is coupled with coal to ethylene glycol technology (CTEG) in combination with a coke oven gas separation unit (COGS), its exergetic efficiency is the highest among conventional CTEG processes and the integration of the CTEG with other reforming processes such as SRM, DRM, and BRM. The order of exergy efficiencies is 10.3, 30.5, 34.9, 38.8, and 40.8% for CTEG, CTEG-SRM, CTEG-DRM, CTEG-BRM, and CTEG-TRM, respectively. The reason is that the reforming processes have higher efficiency, although they consume more utilities than the conventional CTEG method [63].

However, when employing pure oxygen, the energy necessary for an oxygen plant must also be taken into account. A cryogenic facility, for example, consumes 0.37 kWh/kgO₂ energy [92]. When air is utilized instead of pure oxygen, the nitrogen in the air stream reduces thermal efficiency since it does not react, raising the required compression and heating energy [96].

If the TRM process is combined with oxy-fuel combustion, the negative effects of nitrogen can be counteracted. In this configuration, a water electrolysis unit supplies the required oxygen for the oxy-fuel combustion unit and adds the generated hydrogen to the methanol production unit [39]. The net energy efficiency of this plant is 62%, which is significantly higher than the direct CO₂ hydrogenation method for methanol production [41]. Compared to air–fuel combustion, this configuration can boost the process’s CO₂ valorization potential by more than 3 times at the expense of its profit-generating potential because of the additional cost of water electrolysis [39,72]. Nonetheless, it is possible to enhance the profit of this concept by increasing the water electrolysis throughput and directing the extra produced oxygen to the TRM process to reduce TRM heat duty. This modified water electrolysis unit consumes less energy than a cryogenic unit [45].

A higher energy efficiency is achieved when the TRM process is integrated with downstream units [9], especially when using an optimal heat exchanger network [53,97], because the recovered heat from the reforming plant can provide the electric power required for the syngas compression [94].

Zhang et al. developed an optimal heat exchanger network for an integrated process of TRM and methanol production to minimize utility and capital costs. They used the general algebraic modeling system (GAMS) for their calculations and corroborated the results using ASPEN Energy Analyzer. According to their design, the gross energy saving is 34.3% when just utility cost reduction is included, but 32.2% when both utility and capital cost minimization are considered. This slight decrease in energy efficiency is because adequate heat exchange across the process flows reduces water and electricity costs while increasing capital costs since more heat exchangers are required. Moreover, when BRM is considered instead of TRM in this integration, BRM uses less specific energy (11.5 kWh/kgCO₂ vs. 19.0 kWh/kgCO₂) but produces twice less methanol than TRM [53].

These authors also conducted the same research to merge the TRM process with the dimethyl ether (DME) synthesis process, which yielded comparable results. The optimal heat exchanger network results in gross savings of 37% for the whole process and 33.3% for the DME process compared to the original case. Furthermore, the TRM process consumes a little more energy compared to the BRM for DME synthesis (0.7 kWh/kgCO₂ more energy) while generating slightly more DME (0.06 kg/kgCO₂). Therefore, the energy saving of the TRM method is 0.97 KW/kgCO₂ higher than BRM [97]. However, a more detailed economic examination considering catalysts and kinetics evaluation is required to completely show the economic viability of DME synthesis by the TRM process.

Table 6. Comparison of integrated TRM processes with conventional ones.

Process	Compared to	Saving Type: Amount %	Efficiency Type: Increased Amount %	Ref.
CO2 separation unit	Chemical absorption with alkanolamines	Energy of a coal-fired power plant: 20%		[90,91]
TRM reaction	BRM reactions	Energy: 29%		[67]
TRM of the flue gases released from coal- and gas-fired burners of power stations containing N2	Conventional methanol production with SRM	Fuel: 31%	Exergy: 72%	[93]
TRM of the flue gases released from coal- and gas-fired burners of power stations containing N2	Conventional H2 production with SRM	Fuel: 75%	Exergy: 74%	[93]
Combined TRM with COG and WGS unit	Conventional CTM		Carbon utilization: 4.3%	[95]
			Energy: 11.4%
Combined TRM with COG	Conventional CTM		Carbon utilization: 14.5%	[95]
			Energy: 16.8%
Combined TRM with CTEG and COGS	CTEG		Exergy: 30.5%	[63]
	Combined CTEG and SRM		Exergy: 10.3%
	Combined CTEG and DRM		Exergy: 5.9%
	Combined CTEG and BRM		Exergy: 2.0%
Combined TRM with oxy-fuel combustion equipped with water electrolysis unit	Direct CO2 hydrogenation method for methanol production		Net energy: 10%	[41]
Combined TRM with methanol production WITH heat integration	Combined TRM with methanol production WITHOUT heat integration	Gross energy: 34.3% (minimizing utility cost)		[53]
		32.2% (minimizing utility cost and capital cost)
Combined TRM with methanol production with heat integration	Combined BRM with methanol production with heat integration	Specific energy: 4.0 kWh/kgCO2		[53]
Combined TRM with DME production WITH optimal heat exchanger network	Combined TRM with DME production WITHOUT optimal heat exchanger network	Gross energy: 37% (whole process)		[97]
		33.3% (DME process)
Combined TRM with DME production with optimal heat exchanger network	Combined BRM with DME production with optimal heat exchanger network	Energy: 0.97 kWh/kgCO2		[97]

Table 6. Comparison of integrated TRM processes with conventional ones.

Process	Compared to	Saving Type: Amount %	Efficiency Type: Increased Amount %	Ref.
CO2 separation unit	Chemical absorption with alkanolamines	Energy of a coal-fired power plant: 20%		[90,91]
TRM reaction	BRM reactions	Energy: 29%		[67]
TRM of the flue gases released from coal- and gas-fired burners of power stations containing N2	Conventional methanol production with SRM	Fuel: 31%	Exergy: 72%	[93]
TRM of the flue gases released from coal- and gas-fired burners of power stations containing N2	Conventional H2 production with SRM	Fuel: 75%	Exergy: 74%	[93]
Combined TRM with COG and WGS unit	Conventional CTM		Carbon utilization: 4.3%	[95]
			Energy: 11.4%
Combined TRM with COG	Conventional CTM		Carbon utilization: 14.5%	[95]
			Energy: 16.8%
Combined TRM with CTEG and COGS	CTEG		Exergy: 30.5%	[63]
	Combined CTEG and SRM		Exergy: 10.3%
	Combined CTEG and DRM		Exergy: 5.9%
	Combined CTEG and BRM		Exergy: 2.0%
Combined TRM with oxy-fuel combustion equipped with water electrolysis unit	Direct CO2 hydrogenation method for methanol production		Net energy: 10%	[41]
Combined TRM with methanol production WITH heat integration	Combined TRM with methanol production WITHOUT heat integration	Gross energy: 34.3% (minimizing utility cost)		[53]
		32.2% (minimizing utility cost and capital cost)
Combined TRM with methanol production with heat integration	Combined BRM with methanol production with heat integration	Specific energy: 4.0 kWh/kgCO2		[53]
Combined TRM with DME production WITH optimal heat exchanger network	Combined TRM with DME production WITHOUT optimal heat exchanger network	Gross energy: 37% (whole process)		[97]
		33.3% (DME process)
Combined TRM with DME production with optimal heat exchanger network	Combined BRM with DME production with optimal heat exchanger network	Energy: 0.97 kWh/kgCO2		[97]

Table 7. Effect of different design factors on the energy and exergy efficiency of the system [98].

Design Factor	Effect on Energy Efficiency	Effect on Exergy Efficiency
Reactor inlet temperature	D 1	N 4
H2O concentration	D	I
Reactor pressure	I 2	N
CO2 concentration	I	I
Air concentration	O 3	O
Reactor length	O	I

¹ D: Directly proportional. ² I: Inversely proportional. ³ O: There is an optimum. ⁴ N: No substantial change.

Table 7. Effect of different design factors on the energy and exergy efficiency of the system [98].

Design Factor	Effect on Energy Efficiency	Effect on Exergy Efficiency
Reactor inlet temperature	D 1	N 4
H2O concentration	D	I
Reactor pressure	I 2	N
CO2 concentration	I	I
Air concentration	O 3	O
Reactor length	O	I

¹ D: Directly proportional. ² I: Inversely proportional. ³ O: There is an optimum. ⁴ N: No substantial change.

Sadeghi et al. conducted a thorough investigation of the effects of several design factors on system performance. Given that the main products are hydrogen and carbon monoxide, the overall energy efficiency of the system and the reactor are as follows:

$${\mathsf{\eta }}_{\mathrm{reactor}}=\frac{{\mathrm{n}}_{{\mathrm{H}}_{2,\mathrm{outlet}}}{\times \mathrm{HHV}}_{{\mathrm{H}}_2}{+\mathrm{n}}_{\mathrm{CO},\mathrm{outlet}}{ \times \mathrm{HHV}}_{\mathrm{CO}}}{{\mathrm{n}}_{{\mathrm{CH}}_4,\mathrm{inlet} }\times { \mathrm{HHV}}_{{\mathrm{CH}}_4}} \times 100,$$

(12)

$${\mathsf{\eta }}_{\mathrm{overall}}=\frac{{\mathrm{n}}_{{\mathrm{H}}_{2,\mathrm{outlet}}}{ \times \mathrm{HHV}}_{{\mathrm{H}}_2 }{+\mathrm{n}}_{\mathrm{CO},\mathrm{outlet}}{ \times \mathrm{HHV}}_{\mathrm{CO}}}{{\mathrm{n}}_{{\mathrm{CH}}_{4,\mathrm{inlet}}}{ \times \mathrm{HHV}}_{{\mathrm{CH}}_4}{+\mathrm{W}}_{{\mathrm{CH}}_4,\mathrm{COMP}}{+\mathrm{W}}_{{\mathrm{CO}}_2,\mathrm{COMP}}{+\mathrm{W}}_{\mathrm{AIR},\mathrm{COMP} }{+\mathrm{W}}_{\mathrm{PUMP}}} \times 100,$$

(13)

Since H₂ is the most valuable product and has the highest HHV (higher heating value), increasing any parameter that results in a higher H₂ concentration in the product will increase energy efficiencies under Equations (12) and (13). Therefore, increasing the reactor inlet temperature or H₂O concentrations or decreasing the reactor pressure or CO₂ concentration increases the system efficiency. Additionally, if air is utilized instead of O₂, there is an optimum air concentration because the N₂ molar flow rate also increases, increasing the compressor workload. There is even an optimal reactor length since there is no chemical potential in the mixture beyond a specific length of the reactor to produce high-quality syngas [98]. Table 7 summarizes these findings.

The same authors similarly assessed the exergy of the system due to the irreversibility that occurs during the process, which is not considered by energy analysis alone. They argued that increasing the volume of H₂O and CO₂ as well as the length of the reactor reduces the exergy efficiency of the system. However, when the inlet temperature exceeds 900 °C and the reactor pressure increases, there is no substantial change in the system’s exergy efficiency [96,98]. Rather, there is a value for air (say, 3.35 mol/s) at which the overall energy efficiency of the system is at its lowest. The reason is similar to that previously stated for energy efficiency due to the presence of N₂ [98]. Table 7 highlights the effect of various parameters on the system’s exergy efficiency.

10. Economic Assessment

This section describes how specific factors and the TRM process, when combined with other facilities, affect system costs. Additionally, the economic advantages of using the TRM method in comparison to alternative processes are covered.

Research shows that when the synthesis gas produced by TRM and BRM processes is delivered to a methanol plant, the price of methanol can be competitive with the traditional method of methanol production. TRM’s specific capital cost (CAPEX/tonMeOH) and specific operating cost (OPEX/tonMeOH) are also somewhat lower than BRM’s (3.7% and 1.3%, respectively) [99]. Compared to BRM, TRM has a lower specific CAPEX at any capacity level because it uses oxygen, which allows the process to become more efficient and thus reduces equipment size, and this behavior becomes more visible at large production levels. However, for all scenarios, the specific CAPEX decreases as methanol production increases [8,92]. More than 40% of the total CAPEX for these technologies comes from the cost of purchasing equipment [8]. The dominant equipment cost of the BRM and TRM belongs to the reformer (a box furnace) and the air separation unit (ASU), respectively [92]. Nonetheless, BRM has a relatively lower OPEX because an air separation unit is not required [25].

A comparison of the TRM and BRM techniques with a hydrogenation-based process, in which CO₂ reacts with H₂ to produce methanol, reveals that the investment costs of a hydrogenation-based process are approximately 2.5-times lower than those of the TRM and BRM processes for any capacity because a hydrogenation-based process does not require additional processing units such as pre-reformers or reformers. However, only a small fraction of the total cost is capital costs (about 2% and 10% for hydrogenation and reforming processes, respectively) and the rest of the costs are the operating costs. Since the cost of hydrogen from renewable sources such as wind, sun, biomass, hydropower, and specifically the solar systems with high-temperature electrolysis is high, the reforming process is advantageous over the hydrogenation approach and it has superior results by lowering the yearly total cost of methanol (37% with BRM and 39% with TRM). Furthermore, when the methanol plant size increases, the operating costs decrease, and the gap between the hydrogenation process and BRM and TRM becomes slightly bigger. Therefore, maybe in the future, the availability of free hydrogen from industrial by-products or cheaper hydrogen from renewable sources would actively lower the cost of methanol synthesis by hydrogenation. Until then, however, BRM and TRM can be seen as a transition technology from a traditional to an alternative methanol plant as they not only have a higher investment appeal than a traditional methanol factory, but they also have environmental advantages [8].

Another study also shows that when water electrolysis utilizing carbon-free energy to be converted into O₂ and H₂ is coupled with TRM, it outperforms the direct CO₂ hydrogenation process for producing methanol in terms of economic feasibility [41]. A significant portion of CAPEX and OPEX (34% and 51%, respectively) is accounted for in the water electrolysis process. Furthermore, a discounted cash flow (DCF) model along with a sensitivity study revealed that a break-even point could be achieved with a methanol price of USD 491/ton, a net payout time (NPT) of 19.8 years, and a net present value (NPV) of USD 11.4 million after 20 years of operation. As a result, for their facility to be commercially feasible, a high methanol sale price of around USD 500/ton is necessary [41].

Likewise, if a TRM process is combined with methanol production, it will be a cost-effective method of turning waste carbons into usable fuels and chemicals [53]. For instance, Borreguero et al. demonstrated that the investment in methanol production utilizing TRM with natural gas, including a pre-reactor to convert ethane and propane to syngas, could be recovered in 7 years, and the process would be economical even with a 22% drop in methanol production mass flow. However, methanol prices lower than EUR 0.28/kg lead to negative NPV values [46]. Heat integration, when electricity is utilized for chilling below 40 °C, lowers the utility expenses by 78.4% and capital costs by 43.8%. [53].

Similarly, the combination of a TRM process with CTEG technology and a COGS unit reduces the capital investment by 18.82% compared to a typical CTEG process. The reason is that this linked process does not require WGS units and consequently has lower equipment costs. In addition, since the TRM process uses less coal and steam than other reforming processes, this combination results in the lowest production costs. Furthermore, the TRM method has the highest internal rate of return (18.85%). Therefore, the CTEG process in combination with a TRM system offers the best economic results compared with the conventional CTEG process or a combination of CTEG with SRM, DRM, and BRM processes [63].

In addition to the combination of these units, process parameters and design factors can also alter process costs. For example, Sadeghi et al. reports on the effect of feed composition and process parameters on the cost of the TRM process. They found that an increase in the concentrations of CO₂ or H₂O, reactor pressure, or inlet temperature, results in higher total capital investment, but for different reasons: this is mostly due to the necessity for larger heat exchangers in the case of higher H₂O concentrations and inlet temperatures, the higher cost of CO₂ compressor for more CO₂ flow rates, and a more expensive compressor and a reactor in the case of high pressure. However, there is an ideal value for the air concentration and the length of the reactor since two opposing effects are occurring at the same time. On the one hand, the capital cost of an air compressor grows as air concentration rises. Increased air concentration, on the other hand, results in a greater reactor exit temperature, a smaller heat exchanger size, and thus a lower capital cost. A longer reactor increases the capital cost, but also increases the concentration of H₂ in the produced syngas, outstripping the total investment cost. However, there is an optimum length (e.g., 1.9 m) where the unit cost of synthesized gas produced is minimal. When the length of the reactor is greater than this value, increasing the length of the reactor has little effect on the amount of hydrogen production. Thus, the yield of hydrogen remains relatively constant, while the total cost of the system increases. However, increasing the air molar flow rate, reactor pressure, and either CO₂ or H₂O concentrations causes a continuous rise in the total product unit cost, whereas increasing the reactor inlet temperature reduces the total product unit cost of the system at a specific value (e.g., 1200 K), in which the H₂ concentration increase overcomes the increase in the total product unit cost of the system. As a result, the unit cost of produced syngas is computed as USD 4.48/GJ when the TRM process is optimized [98].

Furthermore, other arguments exist to support the requirement for TRM to operate at higher pressures. According to Dwivedi et al., TRM reactors operating with large volumes of feed gas at high temperatures and atmospheric pressures (850 °C, 1 atm) are not feasible in terms of capital costs due to the reactor’s large size [20,100]. Ren et al. asserted, for example, that running the TRM at higher pressures is more cost-effective due to downstream units where the manufacture of chemicals and fuels from syngas is kinetically favored at higher pressures, such as methanol production at 50 bar. Additionally, methane as a feedstock is often stored and supplied at higher pressures. Furthermore, using the tri-reformer at a greater pressure would be industrially more efficient [47]. There is also evidence that process conditions with high pressures are the norm for SRM, DRM, and catalytic POX, which operate at pressure ranges of 20–30 bar, 10 bar, and 15–40 bar, respectively [101]. As a result, the recommendation is to employ a TRM reactor that operates at high pressures of 30 bar to limit the size of the reactor and therefore lower the process’s capital expenses [75].

Another case in point is the use of fluidized bed membrane reactors to reduce operational expenses. They can cut capital costs by increasing hydrogen production rates, decreasing heating energy needs, and removing or downsizing downstream gas treatment and purification facilities [102,103].

In conclusion, due to the concurrent effects of uncertainties in key economic factors such as natural gas prices and CO₂ taxes on the plant’s NPV, simulation with the Monte-Carlo method revealed that when TRM is used for methanol production, there is an 84% probability of a methanol plant being feasible. If the natural gas and CO₂ tax costs increased to USD 0.1569/stdm³ and USD 140/1000 kgCO₂, respectively, this plant would not remain operational since it would lead to a negative NPV. NPV is most sensitive to the price of natural gas, followed by changes in the price of O₂ and then CO₂. However, an 80% increase in the price of O₂ or CO₂ still results in a positive NPV. The proposed plant included a pressure swing adsorption (PSA) unit that recovered hydrogen from unreacted gases to improve the syngas H₂ ratio and generated 2000 tons/day of methanol. This facility has a NPV of USD 161 million during a 15-year economic life based on a methanol selling price of USD 390/ton [75]. Moreover, the profitability of the TRM process is highly dependent on the price of the final commercialized products (e.g., methanol) [46]. Borreguero et al. highlighted that the TRM process is economical at a methanol price of 0.320 USD/kg. Negative NPV parameters are determined at lower prices. This is consistent with the break-even point, implying that the process is economically viable [46]. Shi et al. also found that when water electrolysis is coupled with TRM, the break-even point is USD 0.491/kg of methanol [41].

11. Environmental Assessment

With a share of 47% of global CO₂ emissions, fossil fuel power plants are the major contributors to anthropogenic CO₂ emissions [104]. According to calculations, flue-gas treatment with TRM can reduce CO₂ emissions by 59% and conserve fossil fuels by 75% [26,94,105]. Therefore, reforming-based pathways, except for SRM, can be seen as a promising option for CO₂ treatment during the transition from a carbon-based to a carbon-free production [8,99]. Specifically, the TRM process leads to the greatest avoidance of CO₂ emissions when methanol production is considered instead of other products [45]. For instance, for producing methanol, the CO₂ emissions when using TRM and BRM technology are 1.1- to 1.4-times lower than conventional technologies (SMR-based methanol plants), which release 1.49–1.9 tonCO₂/tonMeOH into the air. Between the two TRM and BRM processes, the TRM process emits lower CO₂ (1.39 vs. 1.41 tonCO₂/tonMeOH) [8,75,106,107,108,109].

When a TRM-based methanol plant uses a PSA hydrogen separation unit and a cryogenic air distillation unit, its net CO₂ emissions are 0.91 tonCO₂/tonMeOH, which is half of the CO₂ emissions of a conventional plant and 35% lower than a TRM-based methanol plant, which maintains oxygen from water electrolysis [75] with net emissions of 1.39 tonCO₂/tonMeOH [8]. Methanol, on the other hand, may be generated with net negative carbon emissions of −0.14 tonCO₂/tonMeOH when utilizing a TRM-based methanol plant paired with carbon-free water electrolysis, which results in a net CO₂ reduction of 570,000 tons per year, when the plant’s production rate equals 2095 tons of methanol per day [41].

However, the product of reforming technology is syngas, which can be used as a raw material to make low-emission synthetic fuels such as hydrogen, methanol, synthetic gasoline and DME, or it can be used directly to generate electrical energy and heat in high-temperature fuel cells [110]. Therefore, reforming technology is a more sustainable alternative for producing clean fuels. For example, compared to the conventional CTM process, combining COG and coal gasification with a TRM process for methanol production can reduce CO₂ emissions by 44% [95], and the combination for ethylene glycol synthesis decreases overall CO₂ emissions by 59% [63]. Integrating fossil-fuel-fired power stations with the TRM process can alleviate CO₂ emissions up to 83–90% [9,71]. However, because this unique method still uses steam and energy, the CO₂ emission reduction is mostly attributable to the direct avoidance of CO₂ emissions rather than indirect CO₂ emissions [63].

Furthermore, controlling operating conditions favors the desired environmental benefit. For example, increasing the inlet gas temperature above a certain value promotes the consumption of CO₂ [111]. Alternatively, raising the O₂/CH₄ ratio from 0.1 to 0.2 reduces net CO₂ emissions by roughly 22% [36] due to the lower energy requirement for heat delivery as a result of introducing oxygen to the reforming plant [8].

12. Safety Assessment

When hydrocarbons are combined with oxygen at temperatures above the auto-ignition point, an explosion can occur. As a result, suitable precautions must be taken in large-scale industrial operations where CH₄ and oxygen come into contact at the upper flammability limit and high temperatures [52,112,113,114]. Therefore, regardless of the concentration of hydrocarbons, lowering the oxygen concentration below its minimum flammability concentration limit prevents an explosion [115]. For safety reasons, Khadem et al. considered a lower and upper range for O₂/CH₄ ratios between 0.35 and 0.65 for a TRM system including CH₄, CO₂, O₂, and H₂O [88], and Khajeh et al. proved that at a O₂/CH₄ ratio of 0.6, methane conversion is 97% [116]. In addition, due to the exothermic nature of the reaction between CH₄ and O₂, heat management in the TRM process is difficult, and some points inside the reactor are exposed to hot spots [54]. These hot spots, at best, result in catalyst loss, decreased conversions, and subpar product quality, and, at worst, cause fires and explosions due to the higher flammability potential and the reaction runaway conditions, in which there is a progressive increase in the rate of heat generation, temperature, and pressure [115,117,118].

Based on these findings, some methods can improve the safety of the TRM process. One possibility is to reduce the amount of oxygen in the reactor at any point. This can be accomplished through a variety of methods, including boosting the methane mole fraction above the upper explosion limit by injecting oxygen at different points along the reactor length (side-feeding policy) [111,119,120], using an inert gas (i.e., N₂ or He) to reduce the oxygen concentration to a safe level, for example, a molar ratio of O₂:CO₂:H₂O:CH₄:He = 1:1:2.1:5:18 [54,112], or utilizing air rather than pure oxygen [121]. However, the usage of inert gases contributes to the dilution of the product gas, thereby reducing the efficiency of the process and adds the cost of inert gas separation to the final product [122,123]. For oxygen feed distribution, two distinct schemes are available. In the first design, oxygen is injected into a sequence of fixed beds, with oxygen injection occurring between each bed [40]. In the second arrangement, each reactor tube has inside itself a distributor tube with the necessary number of holes. As the number of holes grows, the designs approximate those of a membrane reactor [119].

The second method is using chemical-looping reforming of CH₄ to avoid direct contact between O₂ and CH₄. This method involves two steps: reducing a metal oxide, known as an oxygen carrier, in contact with CH₄ in one reactor and then replenishing the metal with oxygen by an oxidizing agent, usually air, in another reactor. Thus, CH₄ is physically separated from O₂ [114,124].

A third alternative could be to adopt a fluidized bed reactor by integrating it with membranes. In this configuration, the excellent oxygen dosage capabilities of a membrane reactor with the high heat transfer efficiency of a fluidized bed reactor can be combined to make the operation safer and more efficient even in highly exothermic reactions. A membrane reactor, as shown earlier, helps maintain the local oxygen concentration at a level low enough to prevent the formation of hot spots as well as an explosive mixture. The fluidized bed reactor improves the turbulence and results in lower temperature gradients. However, this configuration still has several drawbacks, such as challenges related to the membrane sealing at the reactor wall and its durability and stability under fluidization conditions [115].

Finally, membrane reactors have shown to be a feasible solution for delivering a high volume of oxygen in a safe manner [64]. Alipour-Dehkordi and Khademi proved that the use of a microporous membrane with a side-feeding approach improves the safety of the process by avoiding direct contact with the explosive gases and limiting the hot spot temperatures [111,120]. As a result, including a membrane into a catalytic reactor while regulating the right axial oxygen concentration can assist with temperature management and the safety of the system [64,87,115]. Nonetheless, since the TRM process involves both exothermic and endothermic reactions, the heat generated by the POX reaction helps to reduce the energy requirements of the SRM and DRM processes. As a result, this acts as a buffer on the reaction temperature, preventing it from rising above a specific degree. This is crucial for avoiding temperature escape and/or the development of hot spots in the reactor. Furthermore, these reactions can be rendered slightly endothermic, almost thermoneutral, or slightly exothermic by adjusting the process conditions (i.e., temperature and CH₄/O₂ ratio in feed). In this situation, the TRM method is not only energy efficient but also safe to use [32,113,125]. Even though the current experimental inquiry into TRM [16,22,52,58,59,65,68,77,85,126,127,128,129] shows that tri-reforming of natural gas looks to be safe, its industrial application will require more engineering evaluations and studies.

13. Conclusions and Outlook

Tri-reforming of methane can convert waste flue gases containing carbon dioxide to syngas on its own and, subsequently, can produce green fuels and chemicals in conjunction with other synthesis processes such as a methanol plant. The findings suggest that combining the TRM unit with other plants can result in a design that is more efficient, cost-effective, and environmentally beneficial compared to conventional production processes. As a result, it may be a more sustainable solution to ensure clean fuel and chemical production. However, compared to other methane-reforming processes, the execution of the TRM process has more technical challenges and needs a thorough understanding of the impact of various parameters individually and in interaction with each other during this process. Thus, future research should concentrate on the following aspects to improve the layout of the TRM process and to develop this innovative process concept for large-scale plant capacity:

• Optimization studies have been independently undertaken focusing only on one or two specific aspects of TRM technology; thus, an integrated optimization of a TRM process should be performed to find the optimal design while taking safety, environmental, and economic aspects into account.
• Determination of the ideal operating conditions from economic and efficiency points of view.
• A comprehensive sensitivity analysis on operational parameters of the process.
• Determination of the optimal O₂ concentration considering both in situ coke removal, energy balance and safety of the process.
• Pilot studies on different reactor configurations to find the best design, since most studies on reactors are of a theoretical nature.
• Investigation on the economic feasibility and commercial viability of the proposed reactor and process configurations.
• Commercialization feasibility study of using membrane technology to separate nitrogen from flue gases for industrial levels.