Prospects and Technical Challenges in Hydrogen Production through Dry Reforming of Methane

Abstract

Environmental issues related to greenhouse gases (GHG) emissions have pushed the development of new technologies that will allow the economic production of low-carbon energy vectors, such as hydrogen (H₂), methane (CH₄) and liquid fuels. Dry reforming of methane (DRM) has gained increased attention since it uses CH₄ and carbon dioxide (CO₂), which are two main greenhouse gases (GHG), as feedstock for the production of syngas, which is a mixture of H₂ and carbon monoxide (CO) and can be used as a building block for the production of fuels. Since H₂ has been identified as a key enabler of the energy transition, a lot of studies have aimed to benefit from the environmental advantages of DRM and to use it as a pathway for a sustainable H₂ production. However, there are several challenges related to this process and to its use for H₂ production, such as catalyst deactivation and the low H₂/CO ratio of the syngas produced, which is usually below 1.0. This paper presents the recent advances in the catalyst development for H₂ production via DRM, the processes that could be combined with DRM to overcome these challenges and the current industrial processes using DRM. The objective is to assess in which conditions DRM could be used for H₂ production and the gaps in literature data preventing better evaluation of the environmental and economic potential of this process.

1. Introduction

The energy transition is currently one of the major global challenges and will require the development of various new technologies allowing the economical production of low-carbon energy vectors. Hydrogen (H₂) is an energy vector that can be used as a fuel or as raw material in the chemical industry and in the production of synthetic fuels (ammonia, synthetic CH₄, synthetic liquid fuels and methanol) [1]. Hence, due to its large use in high-emitting sectors such as transport and industry, H₂ (when produced in a sustainable way) can be a key enabler of the energy transition.

H₂ is currently produced from fossil fuels, biomass or water. Currently, steam reforming of methane (SRM, Equation (1)) is the main process used for the dedicated H₂ production [1]. However, the production of H₂ from fossil resources, such as natural gas or coal, is responsible for significant CO₂ emissions. According to the International Energy Agency (IEA), 10 tons of CO₂ are generated per ton of H₂ when it is produced from natural gas.

In order for H₂ to have a significant impact on the energy transition, it has to be produced in a sustainable way. Water electrolysis using renewable electricity is the main technology currently investigated since it produces high-purity H₂ from renewable energy and water. However, there are still some drawbacks related to high costs, technology readiness level, renewable electricity supply and the coupling of electrolyzers with variable renewable energy depending on the type of electrolyzer used [2]. There are already a large number of reviews available in the open literature about H₂ production through water electrolysis [3,4]. Hence, this technology will not be discussed in the present paper. Another alternative technology that has been considered in the literature for H₂ production is dry reforming of methane (DRM, Equation (2)). This process allows the conversion of two major greenhouse gases (GHG) into a mixture of H₂ and carbon monoxide (CO) with a H₂/CO ratio of 1.

$$\mathrm{SRM}:\hspace{1em}{\mathrm{CH}}_4+{\mathrm{H}}_2\mathrm{O}\rightleftharpoons 3{\mathrm{H}}_2+\mathrm{CO}\hspace{1em}\hspace{1em}\Delta {\mathrm{H}}_{298\mathrm{K}}=206 \mathrm{kJ} \mathrm{mo}{\mathrm{l}}^{-1}$$

(1)

$$\mathrm{DRM}: {\mathrm{CH}}_4+{\mathrm{CO}}_2\rightleftharpoons 2{\mathrm{H}}_2+2\mathrm{CO}\hspace{1em}\hspace{1em}\Delta {\mathrm{H}}_{298\mathrm{K}}=247 \mathrm{kJ} \mathrm{mo}{\mathrm{l}}^{-1}$$

(2)

However, there are issues preventing the use of this process at larger scale, especially for H₂ production. DRM is a highly endothermic reaction and both CH₄ and CO₂ conversion are favored at high temperatures, especially above 600 °C. Figure 1 presents the equilibrium concentrations as a function of temperature for all the species when DRM is operated at atmospheric pressure [5]. Several side reactions (Equations (3)–(6)) producing solid carbon and water are also possible and they not only lead to catalyst deactivation, but they also decrease the final H₂/CO of the syngas produced, which is a major issue when DRM is considered for H₂ production. For temperatures above 800 °C, reverse water gas shift (RWGS) reaction (Equation (6)) is especially favored, which reduces the H₂ production [6].

$$\mathrm{Methane} \mathrm{cracking}:\hspace{1em}{\mathrm{CH}}_4\rightleftharpoons {\mathrm{C}}_{\left(\mathrm{s}\right)}{+2\mathrm{H}}_2{\hspace{1em}\hspace{1em}\Delta \mathrm{H}}_{298}=75 \mathrm{kJ}/\mathrm{mol}$$

(3)

$$\mathrm{Boudouard} \mathrm{reaction}: 2\mathrm{CO}\rightleftharpoons {\mathrm{CO}}_2{+\mathrm{C}}_{\left(\mathrm{s}\right)}{\hspace{1em}\hspace{1em}\Delta \mathrm{H}}_{298}=-172 \mathrm{kJ}/\mathrm{mol}$$

(4)

$$\mathrm{Steam} \mathrm{gasification} \mathrm{of} \mathrm{carbon}: {\mathrm{C}}_{\left(\mathrm{s}\right)}{+\mathrm{H}}_2\mathrm{O}\rightleftharpoons {\mathrm{CO}+\mathrm{H}}_2{\hspace{1em}\Delta \mathrm{H}}_{298}=131 \mathrm{kJ}/\mathrm{mol}$$

(5)

$$\mathrm{RWGS}:\hspace{1em}\hspace{1em}\hspace{1em}{\mathrm{CO}}_2{+\mathrm{H}}_2\rightleftharpoons {\mathrm{H}}_2\mathrm{O}+\mathrm{CO}{\hspace{1em}\Delta \mathrm{H}}_{298}=41 \mathrm{kJ}/\mathrm{mol}$$

(6)

Most of the recent DRM papers published in the literature are focused on catalyst development, especially on overcoming the constraints that lead to catalyst deactivation [7,8,9,10,11,12]. However, very few of them actually analyze the viability of this process for H₂ production [13]. The present paper aims to assess and present in which conditions DRM could be used as a viable pathway for H₂ production. The paper will first present the performance of the most recent DRM catalysts developed for producing syngas with high H₂/CO. Then, it will present other reforming technologies currently investigated that could eventually be combined with DRM to increase the syngas’ H₂/CO ratio. This part will be followed by pilot and commercial scale DRM process reported in the literature as well as the goals of these processes. Finally, based on these findings, an outlook will be presented highlighting the viability of this process for H₂ production based on the current findings and what it would take for its economic viability.

2. Recent Advances on Catalysts for Syngas Production

There are already numerous reports in the literature on the development of new catalysts for DRM to overcome catalyst deactivation via coking and sintering of support and active phase as well as to decrease the cost effectiveness of the overall catalytic system. Hence, these aspects will not be discussed in this work. The focus of the present section will be to identify the catalysts able to produce syngas with high H₂/CO, aiming at H₂ production as well as the process conditions required to achieve this ratio.

2.1. Noble Metal-Based Catalysts

The catalytic activity of noble metals from groups VIII to X (Ru, Rh, Ir, Pd, Pt) has been thoroughly investigated towards DRM when supported over different materials. These noble metals present high catalytic activity for CH₄ and CO₂ conversion (${\mathrm{X}}_{{\mathrm{CH}}_4}$ and ${\mathrm{X}}_{{\mathrm{CO}}_2}$, respectively) and distinct stability against thermal and oxidative sintering and, especially, coke deposition [14,15]. The observed activity of noble metals for DRM follow the order Rh > Ru > Ir > Pt > Pd [16]. Nonetheless, the use of noble metals for such purpose still present, as major drawbacks, the low abundance, high costs and usual toxicity [17]. Recently, research has focused not only on highly active catalysts using noble metals at low concentrations (Table 1), which has been proven effective, but also on the promoting effects of noble metals for transition metal-based catalysts, aiming to enhance their resistance against coke deposition [18,19].

Ruthenium is one of the cheapest noble metals (still almost 500 times more expensive than nickel) that presents a good catalytic performance for DRM [16]. Andraos et al. [20] explored the effect of different supports (Al₂O₃, MgAl₂O₄ and YSZ) on the DRM activity of mono- (Ru) and bimetallic (Ru, Ni) catalysts for hydrogen production. Despite the high catalyst activity of monometallic Ru/Al₂O₃ catalyst (T = 750 °C, ${\mathrm{X}}_{{\mathrm{CH}}_4}$ = 90%, ${\mathrm{X}}_{{\mathrm{CO}}_2}$ = 93%, H₂/CO = 0.94), the highest hydrogen production was obtained when 5 wt% Ni/Al₂O₃ catalyst was promoted with 1 wt% Ru (T = 750 °C, ${\mathrm{X}}_{{\mathrm{CH}}_4}$ = 94%, ${\mathrm{X}}_{{\mathrm{CO}}_2}$ = 97%, H₂/CO = 0.97). Ru functionalization also increased the overall catalyst stability by preventing catalyst sintering and the formation of poorly active phases, such as NiAl₂O₄ for the Ru-Ni/Al₂O₃ catalyst. On the other hand, poor metal dispersion after Ru promotion (maximum dispersion of 6%) was responsible for an increase in carbon deposition over the Ni-based catalyst (from 0.18% to 0.32% for Ni/Al₂O₃ and Ru-Ni/Al₂O₃, respectively) [20]. A high metal dispersion was found to promote a stronger interaction between Ru particles and the catalyst support and lead to lower carbon deposition and increased catalyst stability [21].

Anil et al. [22] investigated the performance of partially substituted perovskites LaAlO₃ by 2 wt% of noble metals Ru, Pt and Pd for the DRM reaction at 800 °C. The authors reported that Ru-substituted LaAl_0.98Ru_0.02O_3−δ presented the highest hydrogen production, reaching a H₂/CO ratio 0.85 with CH₄ and CO₂ conversions of 86% and 100%, respectively. Strong occurrence of RWGS reaction along with DRM was responsible for straying the H₂/CO from unity in all cases, especially on the Pt-substituted LaAl_0.98Pt_0.02O_3−δ (H₂/CO = 0.60). Moreover, on the stability study at 650 °C for 50 h, no significant efficiency reduction was observed on the LaAl_0.98Ru_0.02O_3−δ.

Rhodium has also been evaluated as an effective noble metal with distinct catalytic activity towards DRM [16]. In the study of Moreira et al. [23], the performance of the 0.1 wt% Rh/Al₂O₃ catalyst was compared with 10 wt% Ni/Al₂O₃. The authors reported that, when incipient wetness impregnation method was used for catalyst preparation, the Rh-based catalyst presented a higher S_BET (191.8 m² g⁻¹) and a much higher metal dispersion (45.1%) when compared to the 10 wt% Ni/Al₂O₃ catalyst (154.3 m² g⁻¹ and 1.9%, respectively). The catalytic tests at 700 °C showed that the hydrogen selectivity over the 0.1 wt% Rh/Al₂O₃ was slightly higher, leading to a H₂/CO ratio of 0.91, while the 10 wt% Ni/Al₂O₃ presented a H₂/CO ratio of 0.89 under the same test conditions. However, stability tests at 700 °C showed that the 0.1 wt% Rh/Al₂O₃ catalyst performed up to 184 h of time on stream (TOS) with H₂/CO ratio higher than 0.95, CH₄ conversion of 65% and CO₂ conversion of 58% with no signs of deactivation. On the other hand, the 10 wt% Ni/Al₂O₃ catalyst showed a strong deactivation trend due to sintering and coke deposition at 24–70 h, with significant decrease in the H₂/CO ratio to 0.80.

Platinum is a specially studied noble metal for DRM [24,25,26]. Among the optimization goals on Pt-based catalysts, metal dispersion, size and, especially, phase of Pt particles (with ionic Pt²⁺ presenting higher DRM activity than metallic Pt⁰) have demonstrated a great deal of influence on the catalytic performance [27]. Carvalho et al. [28] evaluated the deposition of Pt over mesoporous alumina promoted with metal oxides (MgO, ZrO₂, CeO₂ and La₂O₃) and reported good stability under DRM conditions. On this study, the authors reported that the Pt/CeO₂-Al₂O₃ catalysts presented the highest activity for DRM at 700 °C with CH₄ and CO₂ conversions of 90% and 78%, respectively, and a high H₂/CO of 0.90. The authors hypothesize that a combination of CH₄ cracking and reverse Boudouard reaction may have compromised the H₂ production over the other oxide-modified catalysts. Moreover, this catalytic performance was maintained for over 24 h even though coke deposition had taken place.

The resistance of Pt against coke deposition as a main trigger for catalyst deactivation was also reported by Xie et al. [29]. The authors evaluated the addition of Pt to CeO₂-supported cobalt catalysts in order to describe a possible synergistic effect on the bimetallic catalyst responsible for an activity boost. Kinetic and density-functional theory (DFT) calculations revealed that the CO₂ is much more easily activated over the PtCo surface (69 kJ.mol⁻¹) when compared to the monometallic surfaces. In addition, the authors also reported that the enhanced presence of oxygen radicals on the support surface due to the oxygen storage capacity (OSC) of CeO₂ also favored the activation of CH₄ via oxygen-assisted activation pathway (${\mathrm{CH}}_4^{*}{+\mathrm{O}}^{*}\to { \mathrm{CH}}_3^{*}{+\mathrm{OH}}^{*}$). Furthermore, the authors reported that the Co/CeO₂ catalyst presented low activity on higher temperatures (>600 °C), which yielded the lowest H₂/CO ratio of 0.3. On the other hand, the Pt/CeO₂ and PtCo/CeO₂ catalysts presented around 95% conversion of both CH₄ and CO₂ at 750 °C with H₂/CO higher than 0.95.

Table 1. Summary of catalytic systems recently reported for H₂ production via DRM.

Catalyst	Preparation Method	Reaction Conditions	${\mathbf{X}}_{{\mathbf{CH}}_4}(\%)$	${\mathbf{X}}_{{\mathbf{CO}}_2}(\%)$	H2/CO Ratio	Coke Deposition	Ref.
Noble metal-based catalysts
0.1 wt% Rh- and 10 wt% Ni-based catalysts supported on γ-Al2O3	Incipient wetness impregnation	700–800 °C, CH4:CO2 = 1:1; GHSV = 12–108 h−1	92	98	>0.97	nr	[23]
2 wt% Ru-, Pt- and Pd-substituted LaAlO3	Combustion synthesis	400–800 °C, CH4:CO2:N2 = 0.1:0.1:0.8, GHSV = 48,000 h−1	>60	>80	0.6–0.85	nr	[22]
0.16 wt% Pt-CeZrO2 supported on carbon fibers	Co-precipitation	800 °C, CH4:CO2:Ar = 0.04:0.1:0.86, GHSV = 5000 h−1	40	52	0.30	nr	[30]
Core-shell 0.6 wt%Ru@SiO2	Reverse-phase microemulsion	700 °C, CH4:CO2:N2 = 1:1:1, WHSV = 34.8 L h−1 g−1	68	70	0.90	nr	[31]
Nickel-based catalysts
10 wt%Ni/MgAl2O4	Microwave-assisted combustion	800 °C, CH4:CO2:N2 = 0.1:0.1:0.8, WHSV = 72 L h−1 g−1	>75	>80	>0.90	2–6 wt%	[32]
1–5 wt% K-promoted 10 wt%Ni/Al2O4	Co-precipitation	650 °C, CH4:CO2:N2 = 0.2:0.2:0.6, WHSV = 30 L h−1 g−1	35–55	42–58	0.50–0.86	No coke for 5 wt% K promotion	[33]
1–8 wt% Mn-promoted 3–8 wt% Ni/Al2O3	Wet impregnation	750 °C, CH4:CO2:Ar = 1:1:1, WHSV = 36 L h−1 g−1	70–80	80–90	0.70–0.76	nr	[34]
Mg(Ni)Al2O4 spinel	Co-precipitation	850 °C, CH4:CO2:N2 = 30:80:90, WHSV = 120 L h−1 g−1	90	98	1.4	nr	[35]
10 wt% Ni/Al2O3	Microwave-assisted combustion	700 °C, CH4:CO2:N2 = 0.1:0.1:0.8, WHSV = 72 L h−1 g−1	72–87	72–87	1.09	13.7 wt%	[36]
1–2 wt% Fe promoted 5–15 wt% Ni/MgAl2O4 catalyst	Incipient wetness impregnation	650–850 °C, CH4:CO2 = 1:1, WHSV = 30,000 L g−1 h−1	57–100	62–95	0.82–1.01	2.1–34.2 wt%	[37]
10 wt% Mo promoted 10 wt% Ni/Al2O3 catalyst	Wet impregnation	550–850 °C, CH4:CO2 = 1:1, GHSV = 20,000 h−1	6–91	8–93	0.40–0.93	nr	[38]
5 wt% La, Ce and Zr promoted 25 wt% Ni/Al2O3	Ultrasound assisted co-precipitation	550–700 °C, CH4:CO2 = 1:1, WHSV = 12,000 mL h−1 g−1	0–70	40–80	0.6–1.1	nr	[39]
4.5 wt% Ni yolk-shell catalyst on hollow silica spheres (HSS)	Micro-emulsion impregnation	800 °C, CH4:CO2 = 1:1, WHSV = 36,000 mL h−1 g−1	94	98	0.95	No coke	[40]
6.5 wt% Ni core-shell catalyst on 3 wt% ZrO2/SiO2	Micro-emulsion impregnation	550–800 °C, CH4:CO2 = 1:1, Flow rate = 15 mL min−1	85–93	90–95	0.9–1.0	No coke	[41]
DMS-supported 7 wt% Ni nanoparticles	Ethylene glycol impregnation	700 °C, CH4:CO2 = 1:1, WHSV = 18 L h−1 g−1	76	83	0.9	nr	[42]
Sandwiched SiO2@8.9wt%Ni@ZrO2 catalyst	Organic wet synthesis	500–900 °C, CH4:CO2 = 1:1, WHSV = 180 L h−1 g−1	7–98	7–95	0.5–1.0	0–19 mg gcat−1 h−1	[43]
5 wt% Ni supported on CeO2-SiO2	Incipient wetness impregnation	600–800 °C, CH4:CO2 = 1:1, WHSV = 48 L h−1 g−1	66–97	51–92	0.77–0.94	nr	[44]
1–5 mol% Ce and Ca promoted Ni/MSC catalyst	Co-impregnation	650–900 °C, CH4:CO2 = 1:1, Flow rate = 120 mL min−1	45–97	54–97	0.8–1.0	nr	[45]
3 wt% Sc, Y, Ce and Pr promoted 15wt% NiMgAl catalysts	Co-precipitation	750 °C, CH4:CO2 = 1:1, Flow rate = 45 mL min-1, WHSV = 15 L g−1 h−1	86–88	95–96	0.97–0.98	nr	[46]
Cobalt-based catalysts
10–25 mol% Co/MgO catalyst	Sol-gel method	750 °C, CH4:CO2 = 1:1, WHSV = 24 L h−1 g−1	85–90	80–85	0.93–0.95	0–0.22 mg gcat−1 h−1	[47]
0.1–2 wt% Ce-promoted 10 wt%Co/Al2O3	Incipient wetness impregnation	650–750 °C, CH4:CO2:N2 = 0.2:0.2:0.6, WHSV = 36 L h−1 g−1	75–82	63–72	0.86–0.91	7.5–12.8 wt%	[48]
3 wt%, Y, La, Ce or Sm-promoted 10 wt%Co/MA	Incipient wetness impregnation	750 °C, CH4:CO2:N2 = 1:1:3.1, WHSV = 36 L h−1 g−1	71–84	73–90	0.86–0.96	7–28 wt%	[49]
3 wt% La-promoted 10 wt% Co/Al2O3	Hydrothermal synthesis	750 °C, CH4:CO2 = 1:1, WHSV = 36 L h−1 g−1	76–96	81–93	0.89–0.99	26–38 wt%	[50]
0.2–0.5 wt% Rd-promoted12 wt% Co/SBA-15	Two-solvent impregnation	550 °C, CH4:CO2 = 1:1, WHSV = 67 L h−1 g−1	15–49	18–43	0.5–1.1	nr	[51]
0.5–2 wt% Sc-promoted 5 wt% Co/TiO2	Incipient wetness impregnation	700 °C, CH4:CO2:N2 = 17:17:2, WHSV = 3.6 L h−1 g−1	82–84	82–85	0.94–0.96	6.8–32.3 wt%	[52]
Other transition metal-based catalysts
LaNi0.34Co0.33Mn0.33O3 perovskite catalyst	Microwave-assisted Pechini method	800 °C, CH4:CO2 = 1:1.05, WHSV = 12 L h−1 g−1	94	93	1.1–1.2	nr	[53]
LaNiO3 perovskite	One-step sol-gel method with chitosan	600–800 °C, CH4:CO2 = 1:1, WHSV = 6–24 L h−1 g−1	48–95	45–95	0.65–1.25	16–34 wt%	[54]
SmCoO3 perovskite	Sol-gel citrate method	700–800 °C, CH4:CO2 = 1:1, WHSV = 30 L h−1 g−1	88–93	89–93	0.8–1.1	nr	[55]
Bi-and trimetallic catalysts
3–15 wt% Ce promoted 1 wt%Pt-1 wt%Pd-1 wt%Ni/MgO trimetallic	Co-precipitation	700–900 °C, CH4:CO2 = 1:1 and 2:1, Flow rate = 30 mL min−1	60–83	39–98	0.4–1.1	2.4 wt%	[56]
0.1–0.5 wt% Rh-4.5–4.9 wt% Co/Al2O3	Wet co-impregnation	600–700 °C, CH4:CO2 = 1:1, GHSV = 1000–2000 h−1	87–94	86–91	0.99–1.0	No coke	[57]
10.5 wt%Ni-4.5 wt%Co supported on Al2O3-MgO catalysts	Sol-gel method	800 °C, CH4:CO2 = 1:1, Flow rate = 40 mL min−1	79	84	0.85–0.90	nr	[58]
Co2Ni2Mg2Al2 mixed oxide catalyst	Hydrotalcite route synthesis	500–800 °C, CH4:CO2 = 1:1, GHSV = 32,000 h−1	97	91–93	0.8	nr	[59]
12.5 wt% Ni and 12.5 wt%Ni-2 wt%Co supported on CeO2-ZnAl2O4	Co-precipitation	700 °C, CH4:CO2 = 1:1, WHSV = 180 L h−1 g−1	76	88	0.99	38 wt%	[60]
7 wt% Ni and 7 wt%Ni-1 wt%Ru supported on SiO2, ZrO2, Al2O3 and MgAl2O4	Wet impregnation and precipitation	800 °C, CH4:CO2 = 1:1, Flow rate = 400 mL min−1	39–100	48–85	0.5–1.0	0.1–0.42 wt%	[61]

nr—not reported; WHSV—weight hourly space velocity; GHSV—gas hourly space velocity; DMS—dendritic mesoporous silica; MSC—mesoporous silica-carbon; MA—mesoporous alumina; SBA-15—mesoporous silica.

Table 1. Summary of catalytic systems recently reported for H₂ production via DRM.

Catalyst	Preparation Method	Reaction Conditions	${\mathbf{X}}_{{\mathbf{CH}}_4}(\%)$	${\mathbf{X}}_{{\mathbf{CO}}_2}(\%)$	H2/CO Ratio	Coke Deposition	Ref.
Noble metal-based catalysts
0.1 wt% Rh- and 10 wt% Ni-based catalysts supported on γ-Al2O3	Incipient wetness impregnation	700–800 °C, CH4:CO2 = 1:1; GHSV = 12–108 h−1	92	98	>0.97	nr	[23]
2 wt% Ru-, Pt- and Pd-substituted LaAlO3	Combustion synthesis	400–800 °C, CH4:CO2:N2 = 0.1:0.1:0.8, GHSV = 48,000 h−1	>60	>80	0.6–0.85	nr	[22]
0.16 wt% Pt-CeZrO2 supported on carbon fibers	Co-precipitation	800 °C, CH4:CO2:Ar = 0.04:0.1:0.86, GHSV = 5000 h−1	40	52	0.30	nr	[30]
Core-shell 0.6 wt%Ru@SiO2	Reverse-phase microemulsion	700 °C, CH4:CO2:N2 = 1:1:1, WHSV = 34.8 L h−1 g−1	68	70	0.90	nr	[31]
Nickel-based catalysts
10 wt%Ni/MgAl2O4	Microwave-assisted combustion	800 °C, CH4:CO2:N2 = 0.1:0.1:0.8, WHSV = 72 L h−1 g−1	>75	>80	>0.90	2–6 wt%	[32]
1–5 wt% K-promoted 10 wt%Ni/Al2O4	Co-precipitation	650 °C, CH4:CO2:N2 = 0.2:0.2:0.6, WHSV = 30 L h−1 g−1	35–55	42–58	0.50–0.86	No coke for 5 wt% K promotion	[33]
1–8 wt% Mn-promoted 3–8 wt% Ni/Al2O3	Wet impregnation	750 °C, CH4:CO2:Ar = 1:1:1, WHSV = 36 L h−1 g−1	70–80	80–90	0.70–0.76	nr	[34]
Mg(Ni)Al2O4 spinel	Co-precipitation	850 °C, CH4:CO2:N2 = 30:80:90, WHSV = 120 L h−1 g−1	90	98	1.4	nr	[35]
10 wt% Ni/Al2O3	Microwave-assisted combustion	700 °C, CH4:CO2:N2 = 0.1:0.1:0.8, WHSV = 72 L h−1 g−1	72–87	72–87	1.09	13.7 wt%	[36]
1–2 wt% Fe promoted 5–15 wt% Ni/MgAl2O4 catalyst	Incipient wetness impregnation	650–850 °C, CH4:CO2 = 1:1, WHSV = 30,000 L g−1 h−1	57–100	62–95	0.82–1.01	2.1–34.2 wt%	[37]
10 wt% Mo promoted 10 wt% Ni/Al2O3 catalyst	Wet impregnation	550–850 °C, CH4:CO2 = 1:1, GHSV = 20,000 h−1	6–91	8–93	0.40–0.93	nr	[38]
5 wt% La, Ce and Zr promoted 25 wt% Ni/Al2O3	Ultrasound assisted co-precipitation	550–700 °C, CH4:CO2 = 1:1, WHSV = 12,000 mL h−1 g−1	0–70	40–80	0.6–1.1	nr	[39]
4.5 wt% Ni yolk-shell catalyst on hollow silica spheres (HSS)	Micro-emulsion impregnation	800 °C, CH4:CO2 = 1:1, WHSV = 36,000 mL h−1 g−1	94	98	0.95	No coke	[40]
6.5 wt% Ni core-shell catalyst on 3 wt% ZrO2/SiO2	Micro-emulsion impregnation	550–800 °C, CH4:CO2 = 1:1, Flow rate = 15 mL min−1	85–93	90–95	0.9–1.0	No coke	[41]
DMS-supported 7 wt% Ni nanoparticles	Ethylene glycol impregnation	700 °C, CH4:CO2 = 1:1, WHSV = 18 L h−1 g−1	76	83	0.9	nr	[42]
Sandwiched SiO2@8.9wt%Ni@ZrO2 catalyst	Organic wet synthesis	500–900 °C, CH4:CO2 = 1:1, WHSV = 180 L h−1 g−1	7–98	7–95	0.5–1.0	0–19 mg gcat−1 h−1	[43]
5 wt% Ni supported on CeO2-SiO2	Incipient wetness impregnation	600–800 °C, CH4:CO2 = 1:1, WHSV = 48 L h−1 g−1	66–97	51–92	0.77–0.94	nr	[44]
1–5 mol% Ce and Ca promoted Ni/MSC catalyst	Co-impregnation	650–900 °C, CH4:CO2 = 1:1, Flow rate = 120 mL min−1	45–97	54–97	0.8–1.0	nr	[45]
3 wt% Sc, Y, Ce and Pr promoted 15wt% NiMgAl catalysts	Co-precipitation	750 °C, CH4:CO2 = 1:1, Flow rate = 45 mL min-1, WHSV = 15 L g−1 h−1	86–88	95–96	0.97–0.98	nr	[46]
Cobalt-based catalysts
10–25 mol% Co/MgO catalyst	Sol-gel method	750 °C, CH4:CO2 = 1:1, WHSV = 24 L h−1 g−1	85–90	80–85	0.93–0.95	0–0.22 mg gcat−1 h−1	[47]
0.1–2 wt% Ce-promoted 10 wt%Co/Al2O3	Incipient wetness impregnation	650–750 °C, CH4:CO2:N2 = 0.2:0.2:0.6, WHSV = 36 L h−1 g−1	75–82	63–72	0.86–0.91	7.5–12.8 wt%	[48]
3 wt%, Y, La, Ce or Sm-promoted 10 wt%Co/MA	Incipient wetness impregnation	750 °C, CH4:CO2:N2 = 1:1:3.1, WHSV = 36 L h−1 g−1	71–84	73–90	0.86–0.96	7–28 wt%	[49]
3 wt% La-promoted 10 wt% Co/Al2O3	Hydrothermal synthesis	750 °C, CH4:CO2 = 1:1, WHSV = 36 L h−1 g−1	76–96	81–93	0.89–0.99	26–38 wt%	[50]
0.2–0.5 wt% Rd-promoted12 wt% Co/SBA-15	Two-solvent impregnation	550 °C, CH4:CO2 = 1:1, WHSV = 67 L h−1 g−1	15–49	18–43	0.5–1.1	nr	[51]
0.5–2 wt% Sc-promoted 5 wt% Co/TiO2	Incipient wetness impregnation	700 °C, CH4:CO2:N2 = 17:17:2, WHSV = 3.6 L h−1 g−1	82–84	82–85	0.94–0.96	6.8–32.3 wt%	[52]
Other transition metal-based catalysts
LaNi0.34Co0.33Mn0.33O3 perovskite catalyst	Microwave-assisted Pechini method	800 °C, CH4:CO2 = 1:1.05, WHSV = 12 L h−1 g−1	94	93	1.1–1.2	nr	[53]
LaNiO3 perovskite	One-step sol-gel method with chitosan	600–800 °C, CH4:CO2 = 1:1, WHSV = 6–24 L h−1 g−1	48–95	45–95	0.65–1.25	16–34 wt%	[54]
SmCoO3 perovskite	Sol-gel citrate method	700–800 °C, CH4:CO2 = 1:1, WHSV = 30 L h−1 g−1	88–93	89–93	0.8–1.1	nr	[55]
Bi-and trimetallic catalysts
3–15 wt% Ce promoted 1 wt%Pt-1 wt%Pd-1 wt%Ni/MgO trimetallic	Co-precipitation	700–900 °C, CH4:CO2 = 1:1 and 2:1, Flow rate = 30 mL min−1	60–83	39–98	0.4–1.1	2.4 wt%	[56]
0.1–0.5 wt% Rh-4.5–4.9 wt% Co/Al2O3	Wet co-impregnation	600–700 °C, CH4:CO2 = 1:1, GHSV = 1000–2000 h−1	87–94	86–91	0.99–1.0	No coke	[57]
10.5 wt%Ni-4.5 wt%Co supported on Al2O3-MgO catalysts	Sol-gel method	800 °C, CH4:CO2 = 1:1, Flow rate = 40 mL min−1	79	84	0.85–0.90	nr	[58]
Co2Ni2Mg2Al2 mixed oxide catalyst	Hydrotalcite route synthesis	500–800 °C, CH4:CO2 = 1:1, GHSV = 32,000 h−1	97	91–93	0.8	nr	[59]
12.5 wt% Ni and 12.5 wt%Ni-2 wt%Co supported on CeO2-ZnAl2O4	Co-precipitation	700 °C, CH4:CO2 = 1:1, WHSV = 180 L h−1 g−1	76	88	0.99	38 wt%	[60]
7 wt% Ni and 7 wt%Ni-1 wt%Ru supported on SiO2, ZrO2, Al2O3 and MgAl2O4	Wet impregnation and precipitation	800 °C, CH4:CO2 = 1:1, Flow rate = 400 mL min−1	39–100	48–85	0.5–1.0	0.1–0.42 wt%	[61]

nr—not reported; WHSV—weight hourly space velocity; GHSV—gas hourly space velocity; DMS—dendritic mesoporous silica; MSC—mesoporous silica-carbon; MA—mesoporous alumina; SBA-15—mesoporous silica.

2.2. Transition Metals-Based Catalysts

Transition metals from the groups VII to X, especially Co and Ni, have demonstrated good catalytic performance for DRM (Table 1). Due to their high catalytic performance and low cost when compared to noble metals, they have been considered as the way to go on improving DRM towards industrial hydrogen production [15,16]. However, the resistance of Ni- and Co-based catalysts against sintering and deactivation by coke deposition is much lower when compared to noble metals [18]. In this sense, several studies have been focused on improving synthesis methods, catalyst support and metal-support interactions in order to increase both the activity, hydrogen selectivity and the stability of transition metal-based catalysts [62].

2.2.1. Nickel Catalysts

Nickel-based catalysts have been extensively reported as highly active for DRM, with a catalytic activity comparable to iridium (a noble metal with higher DRM than Pt and Pd) [63]. Comprehensive studies have shown that the general bi-functional mechanism of DRM goes through a CH₄ dehydrogenation over non-valent metal sites, while basic sites on supports and promoters are responsible for CO₂ dissociation [64]. It has been reported that Ni active sites may present a favorable electron structure for favoring the activation and cleavage of the C–H bonds, which is an important mechanism step for CH₄ activation and H₂ production [65,66].

Chein et al. [67] investigated hydrogen production over 10 wt% Ni/Al₂O₃ with CeO₂ modification to better understand the role of feed composition and CeO₂–Al₂O₃ interaction on the activity and stability of Ni-based catalysts. The authors reported around 75% conversion of CH₄ and CO₂ at 800 °C, over 10 wt% Ni/Al₂O₃ catalysts with a H₂/CO ratio of 0.90. Adding up to 5 wt% CeO₂ to the alumina support led to an increase in both CH₄ and CO₂ conversion to 85% and 90%, respectively. Moreover, H₂ selectivity also increased, leading to a syngas with a H₂/CO of 0.93. Increasing CeO₂ loading from 5 to 15 wt% led to the reduction in the surface area from 104.5 to 90.5 m² g⁻¹, increased the probability of Ni particles’ agglomeration and decreased the catalyst activity. Regarding feed composition, the authors reported that CH₄ conversion increased when feed composition changed from CH₄:CO₂ = 1:1 to 1:3. However, this increase was not followed by a higher H₂ production, and the overall best H₂/CO ratio (0.90) was achieved at CH₄:CO₂ = 1:1. This observation points to the favoring of parallel reactions on higher CH₄:CO₂ ratios.

Ceria (CeO₂) promotion was also found favorable for H₂ production in a recent study by Farooqi et al. [11]. The authors evaluated the effects of La₂O₃ and CeO₂ promotions on 10 wt% Ni/Al₂O₃ catalysts synthesized by the sol-gel method. The authors reported that La-promotion may present a dual role on DRM catalysts by increasing their resistance to coke as well as decreasing the acidic character of the support. However, CeO₂ promotion was more advantageous when compared to La₂O₃ promotion when H₂/CO was considered. CeO₂ promotion led to a syngas with a H₂/CO ratio of 0.95, while a H₂/CO ratio of 0.89 was obtained when La₂O₃ was used as a promoter in the same test conditions. The authors highlighted that the presence of the CeO₂ improved the active phase dispersion and reducibility, which led to higher H₂/CO on the produced syngas.

Cruz-Flores et al. [68] evaluated the impact of a direct synthesis method via Ni-phyllosilicates for the production of Ni-SiO₂ catalysts in comparison to the incipient wetness impregnation method on the catalytic activity and stability of Ni-based catalysts on DRM. The authors reported that the Ni-SiO₂ catalysts synthesized via Ni-phyllosilicates presented a higher metal dispersion even for high Ni loadings (15 wt% Ni), which led to a reduced carbon formation on short reaction times. However, for longer reaction times (TOS > 180 min), significant carbon formation was observed in the temperature range of 600–900 °C. The results indicate that for T < 700 °C, the Boudouard reaction was favored, leading to a high carbon deposition. For T = 800–900 °C, the high H₂ yields (above 90%) were followed by H₂/CO ratio > 1.0 and high carbon formation, indicating the occurrence of CH₄ cracking and CO disproportionation reactions.

Mahfouz et al. [69] used KIT-6 mesoporous silica as support for Ni catalysts synthesis via wet impregnation and reported a high surface area of S_BET = 422 m² g⁻¹ for the 15 wt% Ni/KIT-6 catalyst. The catalytic tests at T = 800 °C showed CH₄ and CO₂ conversions of 88% and 90%, respectively, and the production of syngas with a H₂/CO ratio of 0.92. Moreover, the authors evaluated that, although cerium promotion (ceria/KIT-6 ratio = 3/2) on the 15 wt% Ni/KIT-6 showed a slight increase on CH₄ and CO₂ conversion, the H₂/CO ratio at higher temperatures (>700 °C) was not altered. Furthermore, Mourhly et al. [70] have recently reported the synthesis of Ni-based catalysts over nano-sized mesoporous silica (MSN) as an alternative support. The MSN was prepared by alkaline leaching of silica-rich pumice rock. Different Ni loadings varying between 5 and 20 wt% were added to the MSN support by wet impregnation and tested in DRM at 550–750 °C. This study showed that increasing the amount of Ni on the catalysts did not improve CH₄ conversion due to the formation of Ni particles clusters with non-reactive aspects for hydrogen production. On the other hand, low Ni loading (5–10 wt%) led to high metal dispersion and strong metal-support interaction, favoring H₂ yields close to 70%.

One-pot microwave-assisted combustion synthesis was reported to be a fast and cost-effective method for producing high H₂-yielding 10 wt% Ni/Al₂O₃ catalysts by Medeiros et al. [71]. The authors showed that a low fuel/oxidant ratio for the combustion-based synthesis resulted in catalysts with high surface area (S_BET = 266 m² g⁻¹) and high reducibility (degree of reduction = 93%). Moreover, the microwave-synthesized 10 wt% Ni/Al₂O₃ catalysts presented an H₂ yield of 58% with low coke deposition (1.2 wt%) and H₂/CO ratio close to 1.0 over 20 h of TOS. The same research group has also reported microwave-assisted synthesis of alumina supports for DRM [36]. The synthesized alumina presented a high surface area (221.2 m²/g) and was used as support for the deposition of 10 wt% Ni via incipient wetness impregnation for DRM evaluation. The catalyst presented significant DRM activity, with over 87% CO₂ conversion (T = 700 °C, GHSV = 72 L g⁻¹ h⁻¹) at a 20 h test. The authors also report the production of a syngas with H₂/CO = 1.09 and 13.7 wt% carbon deposition.

Miri et al. [72] recently investigated the influence of several promoters (Fe, La, Zr, Ce and Ca) on the catalytic performance of 10 wt% Ni/MgAl₂O₄ for DRM and reported that La, Zr, Ce and Ca were able to improve Ni dispersion over the catalyst, which could increase catalytic activity. Furthermore, Ce was found to be the promoter which better increased CH₄ conversion (X_CH4 = 48%), at an optimal loading of 3 wt% Ce. Finally, the authors evaluated that the overall H₂/CO ratio could be tuned as a function of CH₄/CO₂ feed composition, and the highest H₂/CO = 1.4 was obtained at CH₄/CO₂ = 2.0.

2.2.2. Cobalt Catalysts

Along with nickel and its applications as DRM catalysts, cobalt has also been investigated as an alternative, inexpensive and earth-abundant metal with distinct catalytic activities for CH₄ reforming [73,74]. However, cobalt-based catalysts have similar deactivation issues as those observed with nickel catalysts. Studies have shown that coke deposition is also a major trigger for Co-based catalysts deactivation, along with active sites oxidation [47,75,76]. Mechanistic studies via DFT showed that the metal-support interaction, metal distribution (related to the metal particle size) and active site oxidation stage play a significant role on the reaction rates for H₂ production and on the deactivation of cobalt sites during DRM. [77].

The influence of cobalt loading on sol-gel synthesized Co/MgO catalysts was investigated by Sukri et al. [47] with focus on the carbon deposition over the cobalt sites. The authors evaluated Co-loadings from 10 wt% to 25 wt% and reported an increase in the number of active sites at higher Co loads (20–25 wt%). On the other hand, the active phase at higher Co loads presented a weaker interaction with the catalyst support which led to a higher rate of carbon deposition. The similar catalytic activity shown by all evaluated catalysts showed a saturation behavior on the loading increase in Co. Furthermore, the authors evaluated the stability of 10 wt% Co/MgO catalyst for 50 h at 750 °C and reported the production of syngas with a H₂/CO ratio of 0.94 along with high CH₄ and CO₂ conversions of 82% and 88%, respectively, and a low carbon formation rate (0.0072 g g_cat⁻¹ h⁻¹). Al Abdulghani et al. [78] investigated the promotion of ZrO₂-supported Co catalysts with boron (B) for increasing reactivity towards CH₄ and CO₂ conversion. The authors showed that B impurities interacted with Co particles on the catalyst and increased hydrogen production as a result of increased reactivity of the Co/ZrO₂ catalyst by three times towards CH₄ decomposition.

Ayodele et al. [79] investigated the influence of CH₄:CO₂ feed ratio from 0.1 to 1.0 on the DRM activity of CeO₂-supported Co catalysts synthesized with 20 wt% Co by incipient wetness impregnation. The authors also reported that the highest CH₄ (75–80%) and CO₂ (80–90%) conversions were obtained for a temperature of 750 °C. The final H₂/CO ratio reached was 0.98–1.0 when the feed ratio (CH₄:CO₂) varied between 0.8 and 0.9. Moreover, although no activity loss was observed during the tests, TEM images of spent catalysts showed that the particle size of active sites increased, which could, in turn, lead to sintering and irreversible deactivation.

Tran et al. [80] evaluated a novel filament-shaped mesoporous γ-Al₂O₃ as catalyst support for 10 wt%Co nanoparticles and reported an increase in H₂ production when compared to commercial γ-Al₂O₃ support. The authors reported that the highest CH₄ (76%) and CO₂ (82%) conversion was obtained at 800 °C and the H₂/CO ratio increase from 0.83 to 0.89 when comparing the commercial as synthesized γ-Al₂O₃ supports, respectively. On the other hand, the authors highlight that such H₂/CO increase was a result of CH₄ cracking reaction (Equation (3)), which resulted in a higher carbon deposition (47 wt%) for the synthesized alumina support.

Tungsten carbide (WC) has been referred to in literature as presenting activity comparable to noble metals and a “Pt-like” behavior for CO₂ conversion [81], which has motivated recent studies on its application [82,83]. Li et al. [84] have evaluated the DRM activity of 13 wt% Co-based catalysts supported on activated carbon-modified WC (WC-AC) and reported high stability of the prepared catalysts. According to the authors, at 800 °C, both CH₄ and CO₂ conversions were maintained over 90% for a 50 h test with minor activity drop, while H₂/CO ranged from 0.90 to 0.92 during the whole test. The authors evaluated that an optimal 6.5 wt% W loading on the support was responsible for decreasing the activation energy of the reactants and boosted the catalytic activity. Furthermore, Wang et al. [85] also evaluated the activity of the Co/WC-AC catalyst and reported the effects of La-promotion. Although no significant improvement on H₂/CO ratio was observed, the authors showed that a 10–20 wt% La addition to the catalyst could prevent catalyst sintering and the formation of poorly active CoWO₄ phase on the support, which could lead to catalyst deactivation.

2.2.3. Other Transition Metals

Ni and Co are the most investigated metals for their known and vastly reported catalytic activities towards DRM. However, recent studies have also investigated other transition metals and their role on the catalyst development process for a successful scaling up of DRM process. Dehimi et al. [86] evaluated the effect of temperature on the catalytic activity of Mo-based catalysts doped with Ni on DRM. The catalysts were prepared by incipient wetness impregnation with a nominal load of 20 wt% Mo and a variable load (2–10 wt%) of Ni. The authors reported that the increase in Ni loading to 10 wt% was responsible for a higher Mo dispersion over the catalyst support (Al₂O₃), which positively affected CH₄ and CO₂ conversion. However, the authors highlighted that although the high Mo loading may have been responsible for a low conversion at T = 800 °C (${\mathrm{X}}_{{\mathrm{CH}}_4}$ = 30%, ${\mathrm{X}}_{{\mathrm{CO}}_2}$ = 48%, H₂/CO = 0.40), the Mo-based catalysts presented no coke deposition after the catalytic tests, which may lead to a higher catalyst stability.

Vroulias et al. [87] has recently reported the investigation on the effect of W loading on Ni/Al₂O₃ catalysts for DRM. The catalysts prepared by wet co-impregnation had a W-loading ranging from 0 to 26.2 wt% of W and were tested for DRM at 700 °C. The highest concentrations of W on the catalysts (19.8 wt% and 26.2 wt%) inhibited DRM activity (${\mathrm{X}}_{{\mathrm{CH}}_4}$ < 20%) and the best results were achieved with intermediate W loadings (6.3 wt% and 11.9 wt%; T = 700 °C, ${\mathrm{X}}_{{\mathrm{CH}}_4}$ = 55–60%, ${\mathrm{X}}_{{\mathrm{CO}}_2}$ = 60–70%). These loadings, however, led to the production of syngas with a low H₂/CO ratio of 0.5–0.8. Finally, the authors reported that the presence of W was responsible for suppressing coke formation due to a double mechanism: the restrained carbon diffusion on the lattice of Ni-W alloys, and the carbon formation-gasification mechanism that took place on the surface of the catalysts and yielded the formation of tungsten carbides.

2.2.4. Bi- and Trimetallic Catalysts

Bi- and trimetallic catalysts have been widely explored for DRM (Table 1). Studies have shown that the presence of more than one metal on the catalyst may result in a synergistic effect, with all metals acting as active phases or, by a promotion effect, when the active metal has its intrinsic activity enhanced by the presence of a second/third doping metal. Such synergistic effects could be explained by the electronic modifications on the catalyst surface as a result of the presence of a second/third metal, which could translate as increasing reactants chemisorption, for example [88]. Those effects can range from the stabilization of metal particles to improvements in metal-support interactions, metal dispersion and reducibility [89,90]. Increased metal dispersion and strong metal-support interactions, for example, have already been linked to performance improvements on DRM, such as higher H₂ yields and suppressing of coke formation [91]. Furthermore, recent DFT modeling and calculations have considered that the appropriate fine-tuning active phase composition can lead to the suppressing of DRM side reactions, such as RWGS, which would contribute to higher H₂/CO ratio on the final syngas [92].

Al-Fatesh et al. [93] evaluated the effect of calcination temperature and reduction protocols on the DRM activity of Ni-Co bimetallic catalysts supported on Al₂O₃-ZrO₂ mixed oxide. The catalysts were prepared by co-precipitation and subjected to different in situ and ex situ calcination and reduction protocols before being tested on DRM reaction at 500–800 °C. The authors demonstrated that the catalysts calcined at 800 °C and reduced in situ prior to catalytic evaluation at 700 °C yielded the best performance, with CH₄ and CO₂ conversions of 67% and 76%, respectively, and a H₂/CO ratio of 0.91. The authors reported that the calcination at a high temperature triggered sintering and particle growth to a degree that increased both catalytic activity and stability. Bimetallic Ni-Co catalysts were also investigated by Movasati et al. [60]. The catalysts were prepared by co-impregnation using CeO₂-ZnAl₂O₄ mixed oxide as support. In this study, the authors reported that the presence of cobalt increased the catalyst’s surface area when compared to Ni monometallic catalysts from 39.87 m² g⁻¹ to 46.27 m² g⁻¹. The number of basic sites also increased from 180.3 µmol g⁻¹ to 190.1 µmol g⁻¹ when Co was added as a promoter. The metal-support interaction was also favored and prevented the sintering of the active phase. Finally, Ni-Co/CeO₂-ZnAl₂O₄ catalyst yielded a CH₄ and CO₂ conversion of 76% and 88% at T = 700 °C, respectively, with a high H₂/CO ratio of 0.99.

Turap et al. [94] investigated different Co/Ni ratios on the preparation of CeO₂-supported Co-Ni catalysts for DRM application. The authors prepared a series of bimetallic catalysts by co-impregnation with Co/Ni ratios of 0.3–1.0 with a fixed 10 wt% Ni content and evaluated the catalytic activity for DRM between 600 and 850 °C. Reducibility studies showed that at the ratio of 0.8 Co/Ni, the Co addition improved catalyst reduction capacity, increased the metal-support interaction, when compared to Co- and Ni-based monometallic catalysts, and changed the configuration of bulk oxygen from CeO₂ support. In addition, the 0.8Co–Ni/CeO₂ catalyst presented the best catalytic performance at 850 °C, reaching CH₄ and CO₂ conversions of 99% and 100%, respectively, and a H₂/CO ratio of 0.98.

Trimetallic NiFeCu alloys with a Ni/Fe = 3.0 and different amounts of Cu were prepared by co-precipitation and investigated as a catalyst for hydrogen production via DRM in a recent work by Jin et al. [95]. The catalytic performance tests showed that the excess addition of Cu (Cu/Fe = 1.5) results in a decreased initial activity for CH₄ and CO₂ conversion and severe deactivation by carbon deposition (12.7 wt%), due to possible coverage of Ni active sites by Cu. Overall, the Cu addition in the ratio of Cu/Fe = 0.5 delivered the highest H₂/CO = 0.60 at T = 750 °C with a CH₄ conversion of 50%. H₂ production of the trimetallic catalyst was comparable to the bimetallic NiFe catalysts, but the presence of Cu increased the amount of oxygen species on the catalyst surface and decreased carbon deposition by 50%.

2.2.5. Role of Catalyst Support

Although studies have stated that the catalyst support itself would not take part in the DRM reaction [16], it has been demonstrated that several support’s features, such as the surface basicity, oxygen mobility and storage capacity, as well as the ability of interacting with active phases, influence catalytic performance [64,90].

The support’s surface acidic-basic equilibrium may influence the reaction performance in several different aspects. While acidic sites favor CH₄ cracking, which leads to carbon formation [96], the increase in the support’s basicity is related to an enhanced removal of carbon deposits, which may lead to higher catalyst stability [97]. Gao et al. [98] indicate that such acidic-basic equilibrium on the catalyst support may also influence the electronic environment of the active sites and further favor/inhibit side reactions such as CH₄ cracking and CO disproportionation.

The addition of MgO to Al₂O₃ support is a well-described method of increasing support basicity for DRM. Song et al. [99] have described the role of MgO as two-fold. First, the basic properties of MgO increase the adsorption of CO₂, which leads to a higher release of oxygen on the support’s surface and, in turn, increases the coke resistance via the gasification of carbon deposits (reverse Boudouard reaction). Second, the presence of MgO leads to the formation of a layered double hydroxide, in which the Mg²⁺ and Al³⁺ ions are uniformly dispersed and help decrease the agglomeration of active metal particles and, in turn, increase sintering resistance. On the other hand, excessive basicity may also lead to carbon deposition, decrease metal dispersion, lead to lower H₂/CO ratios and compromise the overall catalyst performance [98,100]. Bagabas et al. [97] investigated the optimization of the MgO content on Al2O3-supported Ni catalysts for DRM and reported that the highest H₂/CO = 0.95 (T = 800 °C, ${\mathrm{X}}_{{\mathrm{CH}}_4}$ = 86%, ${\mathrm{X}}_{{\mathrm{CO}}_2}$ = 91%) was obtained with 2.0 wt% MgO, while higher contents were actually detrimental for the process.

Al-Fatesh et al. [101] evaluated hydrogen production over a series of transition metal-modified (Ti, Mo, Si, W) alumina supports with Ni active phase and reported that adding Si and W increased catalyst thermal stability and yielded 60% and 70% H₂ production, respectively, at T = 700 °C. In this study, the authors highlighted the importance of metal-support interaction on the catalyst activity and showed that harmful interactions between the active phase and the oxide support may significantly compromise the catalyst performance. The formation of poorly active NiTiO₃ and NiMoO species, for example, on Ti- and Mo-modified alumina was responsible for a drop in H₂ yields to 30% and 45%, respectively, at T = 700 °C.

Ceria and ceria-modified supports have been investigated for their oxygen storage abilities due to the capacity of cerium to easily change oxidation states from Ce⁴⁺ to Ce³⁺. Studies have shown that high oxygen retention accounts for a higher catalyst reducibility, which may lead to increased activity, as well as higher oxygen mobility and carbon gasification effectiveness [102,103]. The oxygen mobility on the catalyst support bulk structure may influence not only the reaction mechanisms, for example, under DRM conditions, but also the metal-support interaction and the overall catalyst stability [104]. On DRM process, the role of oxygen vacancies on the reaction mechanism is two-fold. First, the oxygen vacancies in the lattice of the support act as activation sites for the CO₂ to promote the cleavage of the C–O bonds, which increases the overall amount of oxygen in the support surface. Second, the presence of oxygen vacancies is responsible for oxygen mobility on the bulk catalyst support, which can react with carbon deposits to produce CO and avoid catalyst deactivation by coke deposition [14].

The effect of CeO₂ modification on catalyst supports was investigated by Faria et al. [105] using a ceria-zirconia mixed oxide promotion on alumina-supported Ni catalysts. The use of ZrO₂ in a solid solution with CeO₂ has been proved to increase and stabilize the oxygen vacancies, and, consequently, the oxygen storage capacity of the mixed oxide, via the atomic mobility inside the lattice to accommodate the two components with different atomic radii [106]. The ceria-zirconia promotion over the alumina support enhanced the oxygen mobility of the catalyst and was found to also increase metal-support interaction, but the probable occurrence of RWGS reaction was responsible for a syngas composition below 1.0 (0.75–0.85) [105].

Overall, catalyst development for hydrogen production via DRM is a major challenge, since the reaction also produces CO in the same amount of H₂ (H₂/CO = 1) and since side reactions (Equations (3)–(6)) further decrease the H₂/CO ratio. Low hydrogen production resulting from competing reactions, catalyst deactivation due to carbon build-up and particle agglomeration can be overcome by the incorporation of promoters and advances on catalyst synthesis methods. However, there is a lack of literature data on the development of catalysts with high H₂ selectivity in DRM, even in papers dedicated to H₂ via DRM. Further mechanism and kinetic studies are required to better understand how to increase the H₂/CO in this process while maintaining the catalytic performance of the catalysts.

Furthermore, different operating strategies, such as combining DRM with other reforming technologies, can also be used to increase the H₂/CO ratio of the syngas produced. The catalytic aspects of different reforming processes and their combinations are discussed in the next sessions.

3. Alternative Processes for Increased H₂/CO Ratio

CH₄ is a major chemical feedstock and the global primary source for hydrogen, accounting for more than 90% of the annual hydrogen production, via different catalytic reforming processes [107]. While SRM is the most mature and most used technology, several other processes, such as DRM, chemical looping methane dry reforming (CL-DRM), methane partial oxidation (POM), combined steam and dry-reforming (CSDRM), for example, have been investigated with a focus on reducing the costs and overcoming the environmental and technical challenges of SRM [12]. The H₂-production potential of each of those technologies can be evaluated by their resulting H₂/CO ratio of produced syngas, which is a function of different catalytic and reaction conditions. Different processes can produce syngas with different H₂/CO ratios [108] and, in the case of DRM, this ratio is close to the unity [109], which is suitable for hydrocarbons and methanol production [110,111]. When it comes to hydrogen production, the importance of securing a H₂/CO ratio as close as possible to the unity is even greater [112], for obvious reasons, but when it comes to syngas production, this is also essential for the direct utilization of the syngas with no need for further hydrogen addition.

3.1. Steam Reforming of Methane (SRM)

Steam reforming of methane (SRM) is the most mature and currently applied technology for hydrogen production. This highly endothermic process (Equation (1)) is operated at high temperature (700–1000 °C) at around 3–40 bar to produce a hydrogen-rich syngas with H₂/CO = 3.0 [113]. Figure 2 presents the thermodynamic equilibrium of the SRM process when operated at two different H₂O/CH₄ ratios. When an equimolar H₂O/CH₄ is considered (Figure 2A), H₂ is produced above 180 °C, and this production increases with the temperature increase. Solid carbon is usually obtained in the range between 450 and 800 °C, which can compromise the operation and lead to catalyst deactivation. For H₂O/CH₄ = 3.0 (Figure 2B), the carbon deposits are potentially eliminated and CH₄ can be completely consumed above 700 °C, which justifies the usual operation with excess steam. For hydrogen production purposes, the SRM is usually combined with water-gas shift reaction (WGS; Equation (7)) operated simultaneously to reduce the CO content and boost the hydrogen concentration [114].

Alumina-supported Ni-based catalysts are also commonly used in SRM. For this reaction, similar problems with coke deposition and catalyst sintering are also observed, with the aggravating factor that the carbon deposits can not only reduce the catalytic performance but contribute to an increased pressure drop on the reforming reactors. The operation of SRM under excess steam conditions (H₂O:CH₄ = 3–5) has been known to help in controlling the carbon deposition over the catalysts [116].

CO + H₂O → CO₂ + H₂  ΔH_298K = −41 kJ.mol⁻¹

(7)

Noh et al. [117] evaluated the performance of Ni-based catalysts supported over calcium aluminate modified SiC for SRM. In this study, the authors aimed to stabilize the H₂ production via SRM by taking advantage of the Lewis basicity of calcium aluminate support to avoid coke deposition. Results revealed that the structured catalysts produced delivered a CH₄ conversion above 80% during 20 h long tests, with H₂ yields above 95% and negligible carbon deposition (T = 850 °C, H₂O:CH₄ = 3). Controlling the coke deposition as a means of enhancing hydrogen production was also the goal of Cho et al. [118] when investigating eggshell-type Ni/MgAl₂O₄ pellets as a catalyst for SRM. The authors showed that homo-type catalyst pellets were more prone to carbon deposition and particle agglomeration due to low metal dispersion. On the other hand, superior metal distribution on eggshell-type catalysts led to lower carbon accumulation (2.5% and 39.5%, for eggshell- and homo-type catalyst pellets, respectively) and more stable SRM operation. In this case, CH₄ conversion remained around 50% for TOS = 6 h with high H₂ selectivity (>85%) and overall H₂/CO = 0.80–0.85.

Membrane-type reactors and catalysts have recently been investigated for SRM applications towards hydrogen production. In membrane-based configurations, the produced hydrogen can be removed from the reformed gas while the reaction is still taking place. This sort of purification process can produce a high purity H₂ stream without the requirements of a second downstream stage and alter the reaction equilibrium to promote higher CH₄ conversions. Kim et al. [119] investigated the use of a Pd composite membrane reactor packed with Ru/Al₂O₃ catalyst for CH₄ reforming at mild conditions (T = 500 °C, H₂O:CH₄ = 3). Despite the low operating temperature, CH₄ conversion reached 79.5% when high transmembrane pressure difference (5 bar) was applied to the system. Moreover, the purity of the H₂ permeate stream remained above 97% on 145 h long tests and accounted for 99% hydrogen recovery. Wang et al. [120] reported using Ni-based hollow fiber membrane catalysts for SRM and reaching 99% CH₄ conversion at 800 °C with hydrogen purity above 95%. In both cases, the authors stated that the transmembrane pressure difference proved to be a very important parameter influencing CH₄ conversion for promoting solution and diffusion mechanisms of H₂ through the membrane structure.

3.2. Chemical Looping Dry Reforming of Methane (CL-DRM)

Chemical looping (CL) is a relatively recent process that has been investigated mainly for CO₂ capture and utilization (CCU) [121,122]. In the CL process, a reaction is “broken down” to a series of multiple sub-reactions that take place separately. In this process, materials with a high oxygen storage capacity, here called oxygen carriers (OCs), are used as reaction intermediates that can be recycled along the process and regenerated by oxygen and energy transfer [123,124]. Several CH₄ processing strategies such as partial oxidation and dry reforming can benefit from the stepwise CL strategy. For DRM, chemical looping dry reforming of methane (CL-DRM; Figure 3) consists of two steps (Equations (8) and (9)). First, CH₄ is reacted with an OC and is oxidized to produce a hydrogen-rich syngas mixture (H₂/CO = 2.0). This step is usually called the “reduction step” since the OCs are reduced in the process of syngas production. Second, carbon dioxide is used as an oxidant to recycle the OCs in the usually called “oxidant step”, producing CO in the process, which can be recovered separately or mixed with the syngas from the first step [125].

CH₄ + OC_oxidized → 2H₂ + CO + OC_reduced

(8)

CO₂ + OC_reduced → CO + OC_oxidized

(9)

The CL-DRM operation offers numerous advantages when compared to the classical co-feed strategy. Overall, the global process corresponds to a classically operated DRM reaction and still presents an endothermic character. However, the equilibrium of the reaction is no longer dependent solely on the thermodynamics and the reactivity of the reactants and the catalysts, but rather on the redox properties of the OCs, which are the major focus of research and optimization [126]. Chein and Hsu [126] evaluated the thermodynamic equilibrium of CL-DRM as a function of the OC:CH₄ ratio, using Fe₂O₃ as OC. For OC:CH₄ < 1.0, POM is the main observed reaction (Equation (10)), yielding H₂ and CO as the main products (H₂/CO ≈ 2.0). For OC:CH₄ > 1.0, H₂ and CO production decrease as temperature increases, and CH₄ combustion is favored (yielding CO₂ and H₂O) as the OC:CH₄ also increases, reaching total methane combustion for OC:CH₄ > 12.0. Overall, the authors evaluate that an optimal OC:CH₄ = 1.0 could be used to obtain syngas with H₂/CO = 1–2.5 according to the CH₄/CO₂ feed ratio.

Furthermore, two problems observed on DRM can be avoided by the stepwise operation of CL-DRM. The accumulation of carbon on the catalysts (which is now an OC) due to CH₄ dissociation on the reduction step is overcome by the re-oxidation cycle. In this scenario, the deposited coke is converted to CO during the oxidation step. Moreover, since the two reactants, CH₄ and CO₂, are not in contact with CL-DRM operation, the produced H₂ on the reduction step is never in contact with the CO₂ injected on the oxidation step. Hence, the occurrence of RWGS reaction (Equation (6)), which decreases the H₂ selectivity, is also avoided [127,128].

Guerrero-Caballero et al. [129] investigated the performance of a series of CeO₂-based oxygen carriers with Ni, Co and Fe for CL-DRM application. The authors evaluated the influence of preparation method and metal active phase on the catalytic activity of the obtained powders. The catalytic activity of all samples was tested over 12 cycles of CL and presented a relatively stable conversion of both CH₄ and CO₂ over the reduction/oxidation cycles. Ni/CeO₂ presented the best catalytic performance (T = 800 °C, ${\mathrm{X}}_{{\mathrm{CH}}_4}$ = 90%, ${\mathrm{X}}_{{\mathrm{CO}}_2}$ = 85%) and maintained a H₂/CO ratio slightly above 2.0, which is the ratio expected for CL-DRM (Equations (8) and (9)). Despite a lower reactant conversion observed for CeNi_0.3O_y (T = 800 °C, ${\mathrm{X}}_{{\mathrm{CH}}_4}$= 62%, ${\mathrm{X}}_{{\mathrm{CO}}_2}$ = 62%), the catalysts prepared by co-precipitation presented no coke formation, a steady H₂/CO = 2.0 over the 12 CL cycles and no particle sintering, which was related to a better dispersion of the Ni particles on the CeO₂ structure. The influence of the active metal (Ni, Co and Fe) was investigated over wet impregnation-prepared catalysts. The CL-DRM experiments showed that Co/CeO₂ (T = 800 °C, ${\mathrm{X}}_{{\mathrm{CH}}_4}$= 80%, ${\mathrm{X}}_{{\mathrm{CO}}_2}$= 85%) catalysts presented a slightly lower CH₄ conversion when compared to Ni/CeO₂, while maintaining a similar H₂/CO between 2.0 and 2.1. Moreover, similar results were obtained at long-term stability tests that comprised 60 CL cycles, with stable H₂/CO ratio between 2.0 and 2.1. As for the Fe/CeO₂ catalyst, very low CH₄ conversion of less than 5% and rapid complete deactivation (after 6 CL cycles) were obtained. These results were related to the formation of poorly active CeFeO₃ mixed oxide. Interestingly, the analysis of spent catalysts showed that Ni and Co were not able to follow the oxidation/reduction cycles, i.e., once they were reduced on CL-DRM conditions, the metallic species were not oxidized again, which highlighted the importance of the oxygen storage capacity of CeO₂ on the operation of CL-DRM.

Kim et al. [130] described an efficient CL-DRM process using Ni-Fe-Al mixed oxides as oxygen carriers. In this study, the group showed the superior activity of the trimetallic mixed oxide over binary metallic mixtures, which were submitted to significant deactivation over 12 CL cycles, either by activity loss towards CH₄ decomposition or inefficient coke removal on the oxidation step. Overall, Ni-Fe-Al mixed oxides were able to maintain the catalysts’ activity over 20 CL cycles, with no signs of deactivation, to avoid particle agglomeration and activity suppression and to deliver a stable H₂/CO of 2.1. Interestingly, the authors also demonstrated the potential of Ni-Fe-Al mixed oxide for classical co-fed DRM. In this case, results showed that a H₂/CO ratio slightly below 1.0 was achieved for CH₄:CO₂ = 1:1 feed at T = 650 °C, with over 80% conversion of both reactants.

Sastre et al. [131] evaluated the perovskite-type La_0.9Sr_0.1FeO₃ supported on yttria-stabilized-zirconia (YSZ) for DRM and CL-DRM and reported a high H₂ production (H₂/CO = 6.5) during the reforming step for the looping operation. The authors evaluated the DRM process at 850 °C for 24 h and obtained an H₂/CO between 0.95 and 0.98 during the tests. For the CL operation, 20 min looping cycles were performed up to 7 h. The authors reported that for the reforming step (CH₄ injection), a predominant CH₄ cracking (Equation (3)) reaction took place and yielded high H₂/CO ratios (up to 6.5) with high coke formation. On the other hand, coke was successfully removed during the oxidation step (CO₂ injection) and the catalysts did not present significant deactivation during CL operation.

3.3. Combined Steam and Dry Reforming of Methane (CSDRM)

A combination of steam and dry reforming of methane (CSDRM) has been investigated as an alternative for modulating the H₂/CO ratio with a focus on increasing H₂ production and allowing for a higher versatility on the chemical and fuels production steps [132]. The use of steam on the feed for DRM may account for a reduced coke formation by oxidation of adsorbed CH_X that could otherwise lead to carbon deposits [133]. In addition, studies have shown that the convenient adjustment of the CH₄/H₂O ratio on feed composition leads to the desired modulation of H₂/CO ratio [132,134].

Figure 4 presents the thermodynamic equilibrium of CSDRM at different pressures, considering a CH₄:CO₂:H₂O feed ratio of 3:1:2. As both SRM and DRM have an endothermic aspect, the conversion of all reagents is favored at higher temperatures. Lower pressures are related to a decrease in both CH₄ and CO₂ conversion and lower H₂ selectivity. H₂/CO ratio decreases with the temperature increase and stabilizes around 2.0 in the range of 800–1000 °C [135].

Dan et al. [136] investigated the performance of 10 wt% Ni/Al₂O₃ catalyst, which is a classical reforming catalyst, on CSDRM conditions. The study has shown that a higher hydrogen production can be achieved by the CSDRM operation at 700 °C with 100% conversion of CH₄ and a high H₂O/CO₂ ratio on the inlet feed. For H₂O/CO₂ ratios of 2.5, 7.3 and 12.7, the H₂/CO ratios achieved on this study were 2.2, 5.0 and 8.0, respectively. In addition, the stability of the 10 wt% Ni/Al₂O₃ catalyst was evaluated for 24 h at CSDRM conditions and 100% CH₄ conversion was maintained for the whole test when H₂O/CO₂ = 7.3 and 12.7 were used at 700 °C. Moreover, analysis of spent catalyst showed no oxidation or size changes of Ni sites, indicating that deactivation by catalyst sintering or oxidation was not triggered on the reaction conditions, and no carbon deposits were detected.

Batebi et al. [137] reported the influence of feed composition on the CSDRM activity and final hydrogen yields over Ni- and Pd-based catalysts supported on Al₂O₃. Mono- and bimetallic catalysts were prepared by sol-gel method and tested for CSDRM at 500–1000 °C with various (CO₂ + H₂O)/CH₄ ratios (1–3) and CO₂/H₂O ratios (1–3) under atmospheric pressure. A central composite experimental design was used to evaluate the statistical significance of the influence of each reaction condition on the final CH₄ conversion and H₂ yield. Based on the ANOVA results, the authors showed that higher temperatures (T > 700 °C) were crucial to achieve higher CH₄ conversions (>75%) due to the endothermic character of both SRM and DRM reactions. As for the influence of feed composition, the authors confirmed that the CO₂/H₂O ratio had no significant impact on the final CH₄ conversion and H₂ yields. However, when the amount of CO₂ was higher in the inlet feed, lower H₂/CO ratios were obtained. Finally, higher H₂ yields were favored by lower (CO₂ + H₂O)/CH₄ ratios at temperatures lower than 1000 °C, when competitive RWGS reaction takes place and reduces the H₂ final production.

3.4. Partial Oxidation of Methane (POM)

The catalytic partial oxidation of methane (POM) has attracted attention for being a more energy-efficient process when compared to commercial SRM. The fast kinetics of POM lead to a good response time and system compactness, while its exothermic character is less energy demanding and avoids the requirements of high temperature steam [138]. In the partial oxidation process, CH₄ (or other hydrocarbons) is reacted with stoichiometric amounts of pure oxygen and yields a H₂/CO ratio of 2.0 for the produced syngas (Equation (10)), which is suitable for Fischer-Tropsch synthesis [139]. However, controlling the reaction selectivity to avoid total combustion can be challenging [140]. Figure 5 presents the thermodynamic equilibrium for POM as a function of temperature for different operating pressures for a stoichiometric feed (CH₄:O₂ = 2.0, Equation (10)). Enger et al. [141] highlight that, at higher pressures, higher temperatures are required for obtaining a high selectivity towards H₂ and CO (Figure 5b,c). In addition, the authors also comment that high H₂ selectivity (>80%) can be obtained for temperatures as low as 550 °C if CH₄:O₂ ranges between 2.0 and 5.0.

Classical noble- and transition-metal-based catalysts are the most reported for POM, and carbon deposition is also a major trigger for catalyst deactivation. Moreover, the temperature gradients resulting from the exothermic POM reaction increase the probability of catalyst deactivation due to particle sintering, which leads to more challenges on catalyst development [139].

CH₄ + ½O₂ → 2H₂ + CO   ΔH_298K = −36 kJ.mol⁻¹

(10)

Alvarez-Galvan et al. [139] evaluated the performance of different Ni-based catalysts for hydrogen production via POM. Catalysts containing 5 wt% of Ni doped on Al₂O₃, CeO₂, La₂O₃, MgO and ZrO₂ supports were tested for POM activity under stoichiometric feed conditions at T = 750 °C and atmospheric pressure. Among the prepared catalysts, 5 wt% Ni/MgO and 5 wt% Ni/La₂O₃ presented low activity and rapid deactivation, due to the formation of non-catalytic phases between the metal and the oxide support. The alumina-supported 5 wt% Ni/Al₂O₃ revealed to be the most efficient and most stable catalytic system, with >80% CH₄ conversion and H₂ yields throughout the catalytic tests. In addition, the authors showed that Rh-promotion over the Ni/Al₂O₃ catalyst was responsible for a slight increase in CH₄ conversion and H₂ yield up to 90%. Finally, the H₂/CO ratio for the obtained syngas was stable around 2.5 for the Ni/Al₂O₃ catalyst with or without Rh-promotion. Ding et al. [142] also investigated the hydrogen production via POM over Ni-based catalysts and reported the use of mesoporous oxides La₂O₃, Yb₂O₃, ZrO₂ and CeO₂ for the modification of SiO₂ support. The catalysts were prepared with a 10 wt% Ni loading and tested for their POM activity under stoichiometric feed conditions, atmospheric pressure and T = 800 °C. The authors demonstrated that the modifications with mesoporous solid oxides were able to increase Ni dispersion on the support, which increases catalytic activity and CH₄ conversion > 75% (reaching 90% for 10 wt% Ni/ZrO₂-SiO₂). Moreover, the authors reported that hydrogen production via POM can profit from oxygen mobility of oxide supports such as ZrO₂, which promotes the combustion-reforming mechanism. In this mechanism, CH₄ and O₂ first react to form CO₂ and H₂O, followed by a combined steam and dry reforming promoted by the combustion products and the excess CH₄ [139]. In this case, H₂ selectivity increased almost 20% and the final H₂/CO ratio of produced syngas stabilized at 3.0.

3.5. Methane Tri-Reforming (TRM)

The methane tri-reforming (TRM) process refers to a combination of the SRM, DRM and POM processes, in which methane is reacted with a mixture of steam, CO₂ and oxygen to produce syngas [143]. According to its thermodynamic aspects (Figure 6), TRM is favored at higher temperatures (T ≥ 700 °C) due to the already discussed endothermic aspects of the individual reforming reactions (SRM and DRM) and the stability of the reactants. Moreover, lower pressures (1 bar) also favor the TRM equilibrium, although typical reforming reactors usually operate at higher pressures (3–30 bar) to decrease reactor size and meet high-pressure requirements for syngas downstream processes [115,144]. Under stoichiometric conditions (2 mol of CH₄, 1 mol of H₂O, 1 mol of CO₂, 0.1 mol of O₂), CH₄ is almost completely converted (>98%) above 850 °C, oxygen is completely consumed via POM and high syngas selectivity of 1.65 is reached above 800 °C [115]. However, the main advantage of the TRM process is the adjustable H₂/CO ratio of the produced syngas according to the feed composition, which makes TRM a versatile CH₄ reforming process that can be controlled for various downstream applications, such as methanol or Fischer-Tropsch synthesis [145].

As a combination of different catalytic reforming processes, TRM also profits from the low cost and known catalytic activity of Ni-based catalysts for reaction with C–H bonds. Maciel et al. [146] evaluated the performance of 5.75 wt% Ni/Al₂O₃ catalyst for TRM under 1 bar and feed composition CH₄:CO₂:H₂O:O₂ = 1.0:0.49:0.3:0.04. The authors reported that the highest hydrogen yield of 37% was achieved at temperatures higher than 850 °C. The competing RWGS reaction was predominant for T < 850 °C and decreased the H₂ yield to 4.4%. Majewski and Wood [147] prepared core-shell-type 11 wt% Ni/SiO₂ catalysts for TRM and reported high hydrogen production and low coke deposition for T = 750 °C. At this temperature, SRM and DRM reactions were favored, which increased H₂ production, while for T ≥ 800 °C, RWGS took place and decreased the overall H₂/CO ratio to 1.5–1.7. In this study, the authors reached a H₂/CO ratio of 2.6 with 73% and 56% of CH₄ and CO₂ conversion, respectively, and low carbon formation (5.0 mg g⁻¹ catalyst) for a feed composition of CH₄:CO₂:H₂O:O₂:He = 1.0:0.5:0.5:0.1:0.4.

Pino et al. [148] investigated hydrogen production by TRM over 1.76 wt% Ni/CeO₂ catalyst at T = 800 °C and described the impact of La-doping and different feed compositions on the final H₂/CO ratio. The authors reported that La modifications on the catalysts were able to increase the presence of oxygen vacancies on the ceria support and enhanced Ni dispersion, which account for high CH₄ (96%) and CO₂ (86%) conversions. However, the La-doping did not show any significant impact on hydrogen production. The modifications on feed composition, on the other hand, were able to modulate the final H₂/CO ratio of the syngas. The authors demonstrated that the H₂/CO ratio could drop from 2.8 to 1.3 when the inlet molar ratio of CO₂/H₂O varied from 0.0 (in the absence of CO₂) to 2.6. According to Pham Minh et al. [145], the modulation on the H₂/CO ratio as a function of the feed composition is due to competitive CH₄ reforming reactions that take place as a result of the presence of multiple oxidants (CO₂, H₂O and O₂) on the TRM process.

García-Vargas et al. [149] investigated the influence of different support materials for 5 wt% Ni-based catalysts. The authors synthesized four different catalysts supported over Al₂O₃, yttria-stabilized-zirconia (YSZ), CeO₂ and silicon carbide (β-SiC). The high thermal conductivity, mechanical strength and chemical inertness of β-SiC proved to be excellent catalytic features for TRM. The 5 wt% Ni/β-SiC presented the highest H₂/CO ratio of 2.0 on the catalytic tests. Moreover, both 5 wt% Ni/β-SiC and 5 wt% Ni/CeO₂ catalysts showed high CH₄ conversion and the highest reaction rates (11.70 × 10⁴ mol s⁻¹ g_Ni⁻¹ and 11.13 × 10⁴ mol s⁻¹ g_Ni⁻¹, for 5 wt% Ni/CeO₂ and 5 wt% Ni/β-SiC, respectively).

4. Large-Scale Applications of DRM

The potential benefits of industrial implementation of DRM have driven several research groups into investigating not only the catalyst optimization aspects but also the process scaling up to commercial level [150]. Over the years, different groups have evaluated several aspects of the DRM process on a semi-pilot and pilot scale, aiming at bringing it into commercial application. Kahle et al. [151] have demonstrated the use of Pt-based catalysts for DRM at 850–950 °C and 20 bar and have considered different process aspects such as catalyst position, temperature gradient and feed composition aiming at reducing coking formation at pilot-scale conditions. The authors reported reaching a H₂/CO = 0.78 at T = 850 °C, 20 bar, CH₄:CO₂ = 1:1. In addition, studies have also addressed other aspects, such as the type of reactor [152] and catalyst structure (in the case of different shapes of pelletized catalysts) [150], aiming at maximizing DRM activity and reducing deactivation rates at pilot scale.

While CH₄ reforming technologies, such as autothermal reforming and steam reforming, are already used in industrial scale, the commercial application of DRM has been limited as a part of some industrial processes, such as the CALCOR™, SPARG™ and MIDREX™ processes [153]. Although such processes do not target H₂ production, understanding the commercial-level applications of DRM is a valuable way of assessing the advantages and limitations of the technology.

Table 2. Summary of large-scale DRM applications.

Technology	Catalyst	Feed Composition	Operation Conditions	H2/CO Ratio	TRL	Process Application	Ref.
DRM	10 wt% Ni/Al2O3-MgO	CO2:CH4:N2 = 2:1:0.5	600 °C, 1 bar	>1.0	Semi-pilot scale	Syngas production	[150]
DRM	Pt-based	CH4:CO2:Ar = 1:1:0.05	850–950 °C, 20 bar	0.7–0.8	Pilot-scale	Syngas production	[151]
CALCOR™ (Caloric GmbH)	Ni-based	CH4, LPG, CO2. Feed composition not reported.	Low pressure, high temperature	0.42–1.0	Commercial	CO production	[154,155]
SPARG™ (Haldor-Topsoe)	Partially sulfur-poisoned Ni	CH4, H2O, CO2, H2S (pre-reforming, for catalyst passivation). CO2/CH4 = 0.07–2.5; H2O/CH4: 0.1–0.26	875–945 °C	0.55–3.2	Commercial	Syngas production	[156,157,158]
MIDREX™ (Midrex Technologies)	Ni-based, supported on α-Al2O3 and/or MgO	CH4:CO2:H2O = 1:1:1	850–1000 °C, 2–3 bar	1.7–2.5	Commercial	Metallic Fe production	[159,160,161]
DRYREF™ (Linde and BASF)	Ni-based oxide-supported tablets (4-hole quadrilobes)	CH4, H2O, CO2 (S/C < 1.8)	800–950 °C, 20–40 bar	1.0–3.0	Commercial demonstration (since 2019)	CO and H2 production	[162,163]

TRL—Technology readiness level; LPG—Liquified petroleum gas.

presents a summary of the main aspects of the pilot- and commercial-scale applications of DRM.

The CALCOR™ process (Caloric GmbH) aims at the production of highly CO-concentrated syngas (Figure 7). In the first stage of this process, syngas with a low H₂/CO ratio (0.42) is produced from the dry reforming of liquefied petroleum gas (LPG) or natural gas (NG) with excess CO₂, over a Ni-based catalyst [164]. In this process, the coke deposition is minimized by using a catalyst bed with different stages of catalyst reactivity and different catalyst geometries [165]. As DRM is a highly endothermic process [166], part of the fuel is burnt to provide heat for the reaction. The process outlet stream is a mixture of CO, CO₂, H₂, H₂O and traces of unreacted CH₄. The stream is cooled for H₂O removal, then it is scrubbed through packed towers with monoethanolamine (MEA) or methyldiethanolamine (MDEA) for CO₂ removal and, finally, is purified by a pressure swing adsorption unit (PSA) for H₂ separation that would operate according to the CO purification requirements (can reach up to 99.95% CO) [155]. The unreacted CO₂ is recycled to the reformer and the H₂ produced during the reaction is either burnt as fuel or sold as a separate product. Despite the toxicity, CO is an important chemical building block for the production of several carbon-based chemicals such as acetic acid and phosgene [155,164,167].

The SPARG™ (sulfur passivated reforming) process, which was initially designed by Haldor-Topsoe and commercialized by Sterling Chemical Inc. in 1987, is another example of an industrial application of DRM for the production of syngas with lower H₂/CO ratios. This process takes advantage of the already existing infrastructure for SRM, combining characteristics of both SRM and DRM, and it is able to deliver syngas with a H₂/CO ratio ranging from 1.8 to 3.0 [158,165,168]. The process can operate with different CO₂ and H₂O mixtures and can be considered a type of CSDRM [165]. In this process (Figure 8), sulfur is selectively used to poison the catalyst sites that are most active for coke formation in order to minimize it during the reaction, while keeping part of the reforming activity [168]. The SPARG™ process is operated between 915 °C and 945 °C, and the coke formation over the Ni catalyst is minimized by the sulfur pretreatment that passivates the main coke-active sites on the catalyst [165].

Both the CALCOR™ and the SPARG™ processes are designed to prevent the formation of coke under DRM conditions. Unfortunately, these processes require specific conditions for their full applicability. Therefore, in order to reduce these barriers and increase the range of applicability of these processes, a better understanding of the mechanisms of coking and measures related to the development of catalysts are necessary, as well as the choice of the reactor material [158].

The MIDREX™ process (Figure 9) was developed in the 1960s and remains, to this day, a leading process in the field of the direct reduction technology [169]. This process takes advantage of a combined DRM and SRM process to produce syngas which is later used as a reducing agent for the production of metallic iron [161]. In the MIDREX™ process, the reformer is fed with a mixture of CO₂:CH₄:H₂O = 1:1:1. The MIDREX reforming process is performed over a Ni-based catalyst supported on α-alumina, MgO or a combination thereof, and operated between 850 and 1000 °C at 1.9–3.0 bar. Basic promoters such as Zr, Ce and La are also frequently used on commercial catalysts for this process, which can increase the overall basicity of the catalysts, prevent carbon deposition and stabilize Ni particles [160,161]. Although detailed information on the syngas produced via the MIDREX™ process is not available, since the syngas is not actually the final product, modeling has indicated that the H₂/CO ratio of the syngas should be in the range between 2.0 and 2.5 [160]. The MIDREX™ process operates commercially since the 1970s and has currently over 100 plants installed in 21 countries, having decommissioned only the pilot- and demonstration-scale plants built pre-1975 [159].

Linde and BASF have recently collaborated to the advancement of DRM industrial implementation. The DRYREF™ technology (Figure 10) has been developed by Linde with the aims of not only contributing to the efforts towards anthropogenic CO₂ utilization, but also for reducing the capital (CAPEX) and operational (OPEX) costs of CH₄ reforming technologies [170]. The process is based on an innovative catalyst developed by BASF, the SYNSPIRE™ G1-110, which is a Ni-based oxide-supported catalyst, specially designed to prevent coke formation and to operate at pressurized reforming conditions (20–40 bar) [162]. The DRYREF™ technology is an application of CSDRM, in which the use of CO₂ is designed to reduce the energy requirements of the full SRM process, due to the energy-demanding steam production step. In the process developed by Linde, a lower steam-to-carbon ratio (S/C < 1.8) is combined with CO₂ capture and recycling in order to reduce the overall carbon footprint of the process and the high steam requirements of SRM (S/C ≥ 2.5). The result is an overall 5% and 3% reduction in OPEX and CAPEX, respectively, and a variable syngas H₂/CO ratio (dependent on the initial S/C ratio) in the range of 1.0–3.0, which can be suitable for several downstream processes, such as aldehydes, methanol, acetic and formic acid synthesis [163]. The project started in 2010 and is currently on commercial demonstration scale since 2019, after the commissioning of the Linde Pilot Reformer, in Pullach, Germany [170].

The highly endothermic aspect of DRM requires high operating temperatures for high syngas yields, which promotes undesired parallel reactions and deactivation of conventional catalysts by sintering mechanisms. In addition, the limited stability of conventional reactor materials at the required high temperatures further hinders the application of this process in industrial scale. Technological advancements are therefore required to obtain both reactors and catalysts that can withstand reaction temperatures above 1000 °C [171].

None of the processes presented above target the production of H₂ as the final product. Hence, large-scale data on DRM aiming at a high H₂/CO ratio, as well as detailed information on sizing and overall energy efficiency of the process and specific facilities, are not available in the open literature. Besides catalytic development and process optimization at laboratory scale, life-cycle assessment and techno-economic analysis of this process are also required to evaluate the real environmental advantages of producing H₂ through this pathway and to assess its economic viability.

5. Conclusions and Perspectives

DRM has been extensively reported in the recent literature as a possible alternative pathway for sustainable H₂ production since it uses CH₄ and CO₂ (two main GHGs) as feedstock. However, this process still suffers from catalyst deactivation in the harsh reaction conditions, mainly due to coking and sintering of the active phase. Moreover, economic issues must also be taken into account, especially when H₂ production is considered, since this reaction also produces 1 mol of CO per mol of H₂ and side reactions (Equations (3)–(6)) may also occur, further decreasing the H₂/CO ratio of the syngas produced. However, as presented in Section 2, the focus of catalyst development has been on improving the stability and activity of catalysts and the H₂/CO ratio of the syngas produced is often secondary, even in papers that target DRM as a pathway for H₂ production. Some of the works cited in Section 2 suggested that low metal loadings could lead to small metal particles that would be more available to the reactants and could increase the H₂ production. High surface area of catalysts and high reducibility of the active phase would increase its activity and increase H₂ production. Finally, tuning process conditions, such as CH₄/CO₂ feed ratio, can also improve CH₄ decomposition and increase the H₂/CO ratio of the syngas produced. However, there is still a lack of detailed data on the literature dedicated to improving the catalyst design to favor higher H₂/CO ratios via this reaction.

Some of the technical and economic challenges in DRM for H₂ production could be partially overcome by either using a chemical looping process or by combining DRM with other technologies, such as steam reforming and partial oxidation. DRM operated in a chemical looping is an interesting process, since only CH₄ and CO₂ are used as feedstock, but the process produces syngas with H₂/CO generally around 2.0, which is higher than the H₂/CO obtained in classical DRM reactions. This is achieved by using oxidation and reduction cycles so the H₂ produced from CH₄ oxidation with the catalyst is not in contact with the CO₂, which prevents the RWGS reaction and assures a high H₂/CO ratio. However, catalyst development is still required to improve their oxygen storage capacity before utilization at larger scales. Combined steam and dry reforming of methane (CSDRM) have also been considered as an alternative process to the classical DRM to improve the H₂/CO ratio. Steam is added to the feed allowing the production of syngas with a flexible H₂/CO ratio up to 8.0 depending on the amount of steam used. The addition of steam also contributes to preventing carbon depositions and catalyst deactivation. However, issues related to the endothermicity of this reaction must also be considered. Methane tri-reforming, which adds steams and oxygen to the feed mixture, has also been investigated, since it can produce syngas with a higher H₂/CO of around 3.0. This process can also cope for the endothermicity of the reactions, since part of the CH₄ is combusted to generate heat for the reactions. However, there are issues related to catalyst deactivation via sintering and vapor-solid reactions that occur on the surface of catalysts.

Due to these issues related to catalyst performance and thermodynamics, only very few industrial processes use DRM at large scale. However, none of them aim at the production of H₂ or of syngas with a high H₂/CO ratio. The CALCOR™ process produces a CO-rich syngas while the SPARG™ process produces syngas with a H₂/CO ratio lower than 1. Finally, the MIDREX™ process does not target H₂ or syngas as the final product. The syngas produced in this process is used as a reducing agent for metallic iron production.

DRM is definitely a very promising process for the production of energy vectors, especially considering that GHGs are used as feedstock. However, the role that it will play in the future large scale of H₂ production is still debatable. DRM still faces technical and economic issues related to catalyst deactivation and the endothermicity of the reaction. Moreover, there are still a lot of data missing on how to improve the performance of the catalysts and of the process to increase the selectivity to H₂. Mechanistic and kinetic studies for better understanding the performance of the catalysts for H₂ production and optimizing reactor design and process conditions are still required before the use of this process at larger scale. Finally, there is also a lack of data on life-cycle and techno-economic analysis of DRM for H₂ production. Estimating the costs and the carbon intensity of H₂ produced via this process are crucial to better understand the environmental and economic potential of using DRM for H₂ production.