Mixed oxides as multi-functional reaction media for chemical looping catalysis

Abstract

Over the past two decades, chemical looping combustion (CLC) has been extensively investigated as a promising means to produce electric power while generating a concentrated carbon dioxide stream for sequestration. We note that the chemical looping strategy can be extended well outside of combustion-based carbon capture. In fact, application of the chemical looping strategy in areas beyond combustion can result in somewhat unexpected energy and carbon dioxide savings without producing a concentrated CO₂ stream at all. Furthermore, it allows the looping-based technologies to tap into applications such as chemical production – a $4 trillion per year industrial sector with high energy and carbon intensities. The key resides in the design of effective oxygen carriers, also known as redox catalysts in the context of selective chemical conversion through chemical looping catalysis (CLCa). This contribution focuses on the design and applications of mixed oxides as multi-function reaction media in CLCa. Since typical mixed oxide oxygen carriers tend to be nonselective for hydrocarbon conversion, the first part of this article presents generalized design principles for surface modification of mixed oxides to improve their selectivity and catalytic activity. Applications of these redox catalysts in chemical looping – oxidative dehydrogenation (CL-ODH) of a variety of light alkanes and alkyl-benzenes are presented. This is followed with a discussion of computation assisted mixed oxide design based upon thermodynamic criteria. Finally, a few new directions for the chemical looping technologies are introduced.

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1. Introduction

The chemical industry contributes to nearly 1 gigaton of direct CO₂ emission each year.¹ In addition, nearly 50% of the oil and natural gas input in the chemical sector is consumed as the feedstock, which can lead to secondary emissions in the downstream industrial and/or consumer sectors.¹ Despite the projected strong growth, the chemical industry is expected to curb its emission within this decade according to the Net Zero by 2050 roadmap proposed by the International Energy Agency (IEA).² In fact, several leading chemical and petrochemical companies have pledged more aggressive emission reduction targets over the recent years. Not surprisingly, CO₂ emissions from state-of-the-art chemical production processes tie directly to the energy intensity of the manufacturing process as well as the annual production capacity. Fig. 1a summarizes the production capacity and emission levels from a few key commodity chemicals. Although the overall CO₂ emissions can be reduced, to some extent, by switching to low carbon intensity fuels and increasing the renewable contribution to the electrical grid, such incremental improvements to the existing technologies are unlikely to attain the net zero target by 2050. Rather, transformative chemical production technologies that are fundamentally different from state-of-the-art approaches are required.

Fig. 1 (a) The global annual production volumes and the associated CO₂ emissions of ethylene, propylene, butadiene, and styrene;^3–10 (b) top: comparison between the energy requirement of conventional dehydrogenation (DH) processes to that of the oxidative processes with in situ H₂ combustion (assuming an ideal process with 100% hydrogen byproduct combustion); bottom: comparison between the CO₂ emission from the existing processes to the CO₂ utilization potential from the energy resulting from in situ H₂ combustion (assuming an ideal scenario of utilizing the energy from H₂ combustion for CO₂ splitting to CO).^3–10

At present, the unsaturated hydrocarbons (i.e. Ethylene,³ Propylene,^4–8 Butadiene^3,7 and Styrene^9,10) listed in Fig. 1 are primarily produced by dehydrogenation or cracking processes, either thermally or in the presence of a heterogeneous catalyst. Common limitations for such processes include large heat requirement and equilibrium-limited single-pass yield due to the highly endothermic and endergonic nature of the reactions, shown in eqn (1). The limited product yield and selectivity in turn drive up the energy consumption and CO₂ emissions associated with product separation.

C_xH_y → C_xH_y−2n + nH₂(n ≥ 1) ΔH > 0, ΔG > 0 (unless at very high temperatures) (1)

It has long been recognized that (catalytic) oxidative conversion of the same feed molecules in the presence of gaseous oxygen would address the aforementioned limitations by changing the highly endothermic reactions into highly exothermic ones (Fig. 1b). The exergonic nature of H₂ oxidation also eliminates the equilibrium limitations, shown with eqn (2). Despite these advantages, commercial adoption of oxidative cracking (OC) or oxidative dehydrogenation (ODH) technologies have been scarce. In addition to the limited catalytic selectivity and stability, the key obstacles hindering the development of ODH technologies resides in the safety concerns over co-feeding oxygen with hydrocarbon feedstocks, and the high cost and energy consumptions for oxygen generation. In fact, state-of-the-art cryogenic air separation technology itself only achieves a 25% second law efficiency, or ∼0.78 GJ per tonne O₂ in terms of actual energy consumption,^11,12 making it a highly energy and emission intensive industrial process.

Resulting from three decades of research, a wealth of knowledge has been created for the chemical looping technology in terms of the oxygen carrier selection and performance, reactor design and operation, and the overall technological feasibility, primarily in the context of CO₂ capture from fossil fuel combustion via chemical looping combustion (CLC).^13–51 Shown in Fig. 2a, CLC technologies generally involves two cyclic steps: (1) the use of metal oxide(s) oxygen carrier as the oxidant to fully combust fossil fuels such as coal or methane; (2) the regeneration of the reduced metal oxide(s) in step 1 with air. We note that the core ideas of the CLC strategy reside in in situ oxygen separation from the air and indirect oxidation of carbonaceous fuels. Therefore, the use of chemical looping for efficient air separation (CLAS) represents a natural extension of CLC, as explored by many chemical looping researchers.^52–94 Furthermore, marrying the chemical looping strategy with oxidative catalysis offers a unique opportunity to intensify the production of a few important commodity chemicals with substantially decreased energy consumption and CO₂ emissions.^95–136 Given that separation processes consume ∼60% of the total energy usage in chemical and petroleum industries and heterogeneous catalysts are responsible for >80% of all chemical products worldwide, chemical looping catalysis (CLCa) in this article, has the potential to facilitate process intensification throughout the chemical manufacturing sector by combining catalytic reactions with separations.^{120,137–142} The abovementioned chemical looping process types are summarized in Table 1. In the context of CLCa, the oxygen carriers are denoted as redox catalysts to capture their dual functionality.

Fig. 2 Schematic illustration of: (a) a generic CLC process: a metal oxide redox catalyst combusts fuel into CO₂ in the oxidation step; the reduced metal oxide is reoxidized in the regeneration step with air; (b) a metal oxide redox catalyst oxidatively converts a light alkane into an olefin product in the oxidation step; the reduced metal oxide is reoxidized in the regeneration step with air.

Table 1 Generalized chemical looping reactions by process type (the reactions are not balanced. ODH was used as a CLCa example for illustration purpose)

Process type	Reactions
CLC	Fuels + MeOx → CO2 + H2O + MeOx−δ
	Air + MeOx−δ → MeOx + N2
CLAS	MeOx → O2 + MeOx−δ
	Air + MeOx−δ → MeOx + N2
CLCa	CxHy + MeOx → CxHy−2n + 2nH2O + MeOx−δ
	Air + MeOx−δ → MeOx + N2

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Taking ethane ODH as an example, the redox catalyst particles first convert ethane into ethylene and water using its active lattice oxygen. After completing this ODH step, the oxygen-depleted redox catalyst is exposed to air (and/or steam), to replenish the lattice oxygen.¹⁴³ Compared to conventional co-feed type ODH reactions, chemical looping process could partition the gaseous oxidant and the hydrocarbons, decreasing the selectivities towards unwanted CO_x and oxygenates. This CLCa process can be carried out either in circulating fluidized beds similar to a CFB combustor or parallel packed beds operated similar to the Houdry process.^3,137,144 The advantages of the redox catalysts and CLCa compared to conventional, heterogeneous ODH catalysts include: (i) integration of catalytic reaction with air separation (a simpler and safer process); (ii) potential to achieve high selectivity (absence of gaseous oxygen inhibits side reactions); (iii) potential to tailor heat of reactions, by varying the metal oxide's redox properties, for improved heat management in the redox steps.^137,145 Our recent studies indicated that up to 84% energy savings and emission reductions can be realized by CLCa.¹³⁷Fig. 2 illustrates a generalized schematic for CLCa in the context of oxidative dehydrogenation of hydrocarbons.

As will be discussed in the following sections, the potential of chemical looping extends well beyond oxidation catalysis.

2. Redox catalyst design in CLCa for olefin production

Complete oxidation of carbonaceous fuels (to CO₂ and H₂O) is expected for CLC because this will maximize the heat release, increase the power generation efficiency, and produce a near “sequestration ready” CO₂ stream.⁴⁶ In contrast, CLCa needs to avoid full oxidation since the target products would be value-add fuels and chemicals such as H₂, CO, or unsaturated organic molecules. With such expectations, typical CLC oxides alone would not be effective for CLCa because they tend to result in poor product selectivity. This is not at all surprising given that typical CLC oxygen carriers, particularly those composed of mixed oxides, are often specifically designed to have high equilibrium oxygen chemical potential (μ_O2), facile oxygen evolution kinetics, and (in many cases) highly basic surfaces.^{121,146–149} Therefore, the most critical aspect for CLCa resides in selectivity enhancement of redox-active metal oxides. Given that the (non-selective) reactions occur at the gas–oxide interfaces, surface modification of the oxides is almost always necessary in order to enhance the product selectivity.

Addition of dopants and/or modulating the metal–oxygen bonding strength on the surface have been shown be effective to enhance the selectivity towards light olefins for a number of redox oxides and reactions.^106,150 This section focuses on presenting a more “generalized strategy” for surface modification of mixed oxide in the context of oxidative dehydrogenation of hydrocarbons: instead of doping or impregnation of small amount of heteroatoms on the surface, we reasoned that complete or near-complete coverage of a non-selective oxide surface with a catalytically active layer would effectively suppress side reactions. This approach was first validated in our group for methane partial oxidation reactions,^122,151,152 and was recently extended to a series of ODH reactions using various core and shell materials, as illustrated in Fig. 3.^{119–121,138,143,144,147,153–161} We categorize this generalized ODH strategy into three types: Type 1, shown in Fig. 3a, involves covering the non-selective mixed oxides with an “inert” molten salt layer to simply block the non-selective sites on the surface, thereby inhibiting nonselective oxidation. In this case, the redox catalyst primarily function as an oxygen carrying agent for selective hydrogen combustion (SHC); Type 2, as illustrated in Fig. 3b, utilizes an “active” molten salt as the surface layer. With the assistance of the oxygen species supplied from the mixed oxide, this active surface layer can accommodate and/or generate active oxidants, e.g. electrophilic oxygen species or halogen atoms, to initiate C–H bond activation at the gas-molten salt interface. The reactions likely proceed through a surface-initiated gas phase radical reaction pathway, with concurrent hydrogen oxidation; Type 3, shown in Fig. 3c, involves covering the oxide substrate with a solid catalytic shell to catalyze the dehydrogenation reactions, and the H₂ byproduct is sequentially combusted by the oxygen supplied by the oxide core. The following sections further elaborate on these core–shell redox catalyst design strategies.

Fig. 3 General core–shell design frameworks for mixed oxide@molten salt catalysts and mixed oxide@catalytic shell catalysts: (a) Type 1: “Inert” molten salt promoted mixed oxides; (b) Type 2: “Active” molten salt promoted mixed oxide; (c) Type 3: catalytic material promoted mixed oxides. [O] stands for the oxygen species donated from the oxide surface, which is responsible for the oxidation reaction. Details of the various oxygen species and their evolution are discussed in Section 2.1.

2.1. Mixed oxide @ molten salt redox catalyst for light alkane conversion

2.1.1. Type 1: “Inert” molten salt modification. A Type 1 redox catalyst functions primarily as an oxygen separation and donation agent for selective hydrogen combustion. Therefore, the formation of unsaturated hydrocarbon products would rely on a separate reaction pathway such as thermal cracking or catalytic dehydrogenation. In this case, selectively combusting the hydrogen byproduct would eliminate the equilibrium limitations and supply the heat required by the endothermic dehydrogenation or cracking reactions. A typical application of Type 1 redox catalysts is ethylene production via thermal cracking of ethane. In this context, we developed redox catalysts based on selected molten salt and Mn-based oxides, e.g. Mg₆MnO₈,^121,147 (Fe/Mn)O_x,¹⁵³ (Mn/Si)O_x,¹⁵⁴ and CaMnO_3−δ,^144,158 with tungstates, vanadates and/or molybdates molten salts.^147,162,163

Among these redox catalysts, Mg₆MnO₈@Na₂WO₄ was the most thoroughly studied.^121,147 As shown in the Fig. 4a, co-impregnating Na and W at any ratios on Mg₆MnO₈ would lead to significant increase in olefin selectivity and yield.¹⁴⁷ Of the various Na : W ratio investigated, the Mg₆MnO₈@Na₂WO₄ (Na : W = 2 : 1) redox catalyst exhibited one of the highest increase in C₂H₄ yield (∼45% increase, on an absolute scale) compared to the unpromoted Mg₆MnO₈ at 850 °C and 4500 h⁻¹. The yield increase was primarily resulted from the substantially decreased CO_x selectivity (>80%). Mg₆MnO₈@Na₂WO₄ also exhibited well defined crystalline phases at room temperature (Mg₆MnO₈ and Na₂WO₄), with only a minor Na₄Mg(WO₄)₃ phase (Fig. 4b). The core–shell structure of the catalyst was verified by low-energy ion scattering spectroscopy (LEIS) (Fig. 4c) and X-ray photoelectron spectroscopy (XPS) (Fig. 4d).¹²¹ Even after 10 cycles of sputtering, W and Na/Mg are still far more prevalent than Mn according to LEIS. It is noted that LEIS cannot accurately differentiate Na and Mg due to their similar atomic weights. XPS, which provides near surface elemental compositions (Fig. 4c), also confirmed the suppression of Mn by Na₂WO₄. In fact, impregnation of Na₂WO₄ onto many Mn-containing oxides has shown to create an oxide@Na₂WO₄ core shell structure. For instance, transmission electron microscope with energy dispersive x-ray spectroscopy (TEM-EDS) images published from our group by Hao et al. indicated a core shell structure of CaMn_0.75Fe_0.25O₃(CMFO)@Na₂WO₄ (Fig. 4e)¹⁵⁷ whereas LEIS and/or XPS data indicated surface enrichment of Na₂WO₄ on mixed (Fe/Mn)O_x,¹⁵³ (Mn/Si)O_x,¹⁵⁴ (Mn/Cu)O_x,^118,164 and a number of perovskite oxides. The formation of core–shell structure is likely to be due to the relatively low melting point of Na₂WO₄ (698 °C)¹⁶⁵ and low surface tension of the molten salt on typical oxide surfaces.

Fig. 4 (a) Performance data of the CL-ODH with Na, W promoted Mg₆MnO₈ at 850 °C and GHSV = 4500 h⁻¹;¹²¹ (b) X-ray diffraction characterization of Mg₆MnO₈@Na₂WO₄;¹²¹ (c) LEIS peak area over sputter cycles for Mg₆MnO₈@Na₂WO₄;¹²¹ (d) XPS near surface concentration of Na and Mn of the cycled sample;¹⁴⁷ (e) TEM-EDS of CMFO@Na₂WO₄.¹⁵⁷

Detailed mechanistic studies from our group were conducted by Yusuf et al. to understand the function and mechanism of the molten salt promoter for Type 1 redox catalyst.^121,147 Mg₆MnO₈@Na₂WO₄ was used as the model catalyst. Methanol TPSR experiment (Fig. 5a) showed that the presence of Na₂WO₄ led to significantly decreased CO₂ signal. This suggests that Na₂WO₄ would suppress the basic sites on the surface of Mg₆MnO₈.¹²¹ Mn 2p_3/2 region scan from XPS (Fig. 5b) showed that the peaks corresponding to Mn⁴⁺(HBE) and Mn⁴⁺ at 644.1 eV and 641.7 eV were suppressed in Mg₆MnO₈@Na₂WO₄ as compared to the unpromoted Mg₆MnO₈.¹²¹ This corresponds well to the decrease in the basic sites on the surface. Oxygen donation/reincorporation dynamics were also probed with ¹⁸O/¹⁶O isotope exchange experiments, shown in Fig. 5c and Table 2.¹²¹ The presence of molten Na₂WO₄ inhibits oxygen exchange, particularly in terms of limiting the incorporation rates of dissociatively adsorbed oxygen into the Na₂WO₄ layer. Gas switching experiments between H₂ and O₂ further indicated that H₂ solubility in Na₂WO₄ was negligible, thus indicating that the hydrogen combustion reaction would occur on the molten salt – gas interface. Therefore, it can be inferred that Na₂WO₄ decreases the CO_x formation by increasing the energy barrier for the release of the active lattice oxygen at the reaction interface. Electrochemical impedance spectroscopy (EIS) measurements confirmed that oxygen and electron can transport within the Na₂WO₄ molten layer to facilitate selective H₂ combustion (Fig. 5d).¹⁵⁸ It was further shown that Na₂WO₄ was redox active between W⁶⁺ and W⁵⁺ even though the bulk Na₂WO₄ phase was difficult to be fully reduced. As such, WO₄²⁻/WO₃⁻ redox pair would act as the intermediate to shuttle oxygen to and from the Na₂WO₄ core for selective hydrogen combustion following the following mechanism (Fig. 5e).

Fig. 5 (a) CO₂ signal Methanol TPSR of Mg₆MnO₈ and Mg₆MnO₈@Na₂WO₄;¹²¹ (b) XPS spectra of manganese 2p_3/2 peaks of cycled redox catalysts;¹⁴⁷ (c) ¹⁸O–¹⁶O isotope exchange experiment;¹²¹ (d) oxygen ion conductivity and electronic conductivity of Na₂WO₄;¹⁵⁸(e) proposed mechanism of Mg₆MnO₈@Na₂WO₄.¹²¹

Table 2 Arrhenius activation energies in the range of 800–850 °C for Mg₆MnO₈ and Mg₆MnO₈@Na₂WO₄¹²¹

Activation energy	Mg6MnO8 (kJ mol^−1)	Mg6MnO8@Na2WO4 (kJ mol^−1) (800∼850 °C)
Ro: overall oxygen exchange rate; Ra: dissociative oxygen adsorption rate; Ri: adsorbed oxygen incorporation rate.
E Ro	163.83	202.58
E Ra	227.18	205.04
E Ri	123.82	174.68

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The same strategy can be extended to many other Mn based oxides and molten salts.^{153,154,157,158,164} In a recent study, we demonstrated this strategy with Mg₆MnO₈ and (Cu/Mn)O_x as the redox core and W, V, Mo based alkali salts as the surface promoters.¹⁶⁴ All the abovementioned molten salts were shown to be effective to decrease the CO_x selectivities from >90% (unpromoted oxides) to <15% (Fig. 6a and b) at 850 °C. C₂H₄ selectivities were maintained >80% while achieving up to ∼90% H₂ conversion. As shown in Fig. 6a to ∼74% C₂H₄ yield at 80 vol% C₂H₆ feed and 850 °C can be achieved. Fig. 6a also demonstrated the effect of different molten promoters. The use of multiple alkali metal cations (e.g. Li and Na) and/or transition metals (e.g. Mo and W) for the salt anions were shown to improve the redox catalyst performance in several cases. Such combinative effect further increases the flexibility for designing the Type 1 catalyst.

Fig. 6 (a) Performance data of the CL-ODH with Li, Na, K and W, V, Mo molten salt shell and Mg₆MnO₈ core at 850 °C at 4500 h⁻¹.¹⁶⁴ (b) Performance data of the CL-ODH with Na₂WO₄ molten salt shell and Cu/MnO_x core at 850 °C at 4500 h⁻¹.¹⁶⁴

Catalyst stability was also demonstrated through a long-term study using a more mechanically robust prototype redox catalyst designed based on the aforementioned principles.¹²⁹ As shown in Fig. 7a, stable product selectivity and ethane conversion were maintained, throughout 1400 redox cycles, in a bubbling fluidized bed reactor operated at 845 °C with space velocities varying between 2000 and 3250 h⁻¹. The pressure drop across the fluidized bed was also quite stable, indicating proper fluidization without significant attrition and particle entrainment.¹²⁹ The demonstrated higher conversion, along with in situ H₂ oxidation, leads to a lower energy requirement and potential for near an order of magnitude CO₂ emission reduction (Fig. 7b).

Fig. 7 (a) C₂H₆ conversion and selectivity profile of the prototype redox catalyst over 1400 cycles at 845 °C, 15–30 mol% ethane and GHSV = 2000–3250 h⁻¹; (b) CO₂ emissions of steam cracking and ODH processes under different scenarios for ODH 67/85/99, the numerical value indicates the ethane conversion value in percentage.¹²⁹

While Type 1 redox catalysts were most frequently used in the context of ethane thermal cracking, they can also be applied for oxidative cracking of naphtha. As shown in Fig. 8a and b, the promotion of Na₂WO₄ on BaFe₆Al₆O₁₉, CaMnO₃, or La_0.8Sr_0.2FeO₃ decreases the CO_x selectivity and increases the olefin yields from the cracking of naphtha model compounds such as n-hexane and cyclohexane.^158–160

Fig. 8 (a) Products yields of BaFe₆Al₆O₁₉ after promotion with different amounts of Na₂WO₄ in n-hexane cracking. Reaction condition: T = 700 °C; GHSV = 9000 h⁻¹ cyclohexane concentration ≈13 mol%¹⁶⁰ (b) selectivity profile of cyclohexane cracking with CaMnO₃@Na₂WO₄ and La_0.8Sr_0.2FeO₃@Na₂WO₄T = 750 °C; GHSV = 5400 h⁻¹; cyclohexane concentration ≈7 mol%.¹⁵⁸

A similar strategy has also been recently adopted by a few other researchers in the context of CL-ODH and methane oxydehydroaromatizations.^{96,97,118,166} In all cases, higher H₂ combustion selectivity were observed after promoting the oxide surface with the “inert” molten salt. This further demonstrates the feasibility of the Type 1 catalyst architecture. In fact, the selective hydrogen combustion function of Type 1 redox catalyst would work well when combined with other DH or cracking catalysts. This will be further detailed in Section 2.3.

To sum up, the design of Type 1 redox catalyst mainly aims to increase the selectivity towards hydrogen combustion, reducing the unwanted deep oxidation products such as CO_x. We apply an “inert” layer of molten salt, which is largely inactive towards C–H bond activation, while still being able to selectively oxidize hydrogen using the active lattice oxygen from the transition metal oxides. The term inert mainly refers to the inhibition of C–H bond activation since the molten salt promoters do actively participate in oxygen and electron transport as well as the hydrogen combustion reaction.

2.1.2. Type 2: active molten salt modification. Contrasting the “inert” nature of the Type 1 molten salt towards C–H bond activation, Type 2 redox catalysts utilize an “active” molten salt promoter to accelerate the ODH reactions. This is facilitated by the active species such as peroxides and halogen atoms generated from the interaction with the mixed oxide core. Our recent studies indicate that many mixed oxides, e.g. La_0.8Sr_0.2FeO_3−δ (LSF), when promoted by molten carbonate or halide salts, are quite effective for CL-ODH of C₂–C₄ alkanes.^167,168 Shown in Fig. 9a, LSF@Li₂CO₃ achieved >55% conversion and ∼90% C₂H₄ selectivity for CL-ODH of ethane at ∼750 °C.¹⁶⁷ The time-dependent gaseous products profile within the ODH half cycle (Fig. 9b) also demonstrated significantly higher ethane conversion than thermal cracking (denoted as blank). The molten Li₂CO₃ shell also significantly increased the ethylene selectivity for the LSF core. Compared to the minimal activity in thermal cracking, the significant increase in conversion with the Type 2 redox catalysts indicates that the molten salt layer catalyzes the ODH reactions besides its SHC function.

Fig. 9 (a) Ethylene selectivity and conversion profiles for LSF@10 wt% (or 20 wt%) Li₂CO₃via CL-ODH of ethane (our work) compared with literature results with various catalysts;¹⁶⁷ (b) ethane conversion profile vs. time-on-stream for: LSF@Li₂CO₃, thermal cracking (denoted as blank) and equilibrium conversion at 700 °C and 480 h⁻¹.

in situ XRD was performed to study LSF@10 wt% Li₂CO₃ (Fig. 10a) and LSF (Fig. 10b).¹⁶⁷ The peak analysis indicates that the perovskite (LSF) phase largely remained stable in LSF@10 wt% Li₂CO₃ during the ODH step. In contrast, a significant portion of unpromoted LSF decomposed to La₂O₃, (La/Sr)₂FeO₄ and Fe phases. TGA studies showed that the oxygen capacity of LSF@10 wt% Li₂CO₃ was significantly lower than pure LSF (0.5 wt% vs. 12 wt%). These observations indicate that Li₂CO₃ inhibits the oxygen release from LSF and limits the reduction of Fe cation in the perovskite structure. Coupled with Mössbauer spectroscopy results, it was determined that for LSF@10 wt% Li₂CO₃, Fe⁴⁺ in LSF was reduced to Fe³⁺, whereas in unpromoted LSF, a notable fraction of Fe⁴⁺/Fe³⁺ were deeply reduced to form metallic Fe. The role of Li₂CO₃ was further characterized by electrochemical impedance spectroscopy (EIS, Fig. 10c). The results indicate that, at near melting point, Li₂CO₃ has increased oxygen ion conductivity but the electronic conductivity stayed near zero. This suggests that Li₂CO₃ could transport ionic oxygen species, but not electron. Therefore, such oxygen species must be in the oxidized form. Among all possible species, peroxide ion (O₂²⁻) is the most likely after excluding other potential species (molecular O₂, O²⁻, O₂⁻, and C₂O₄²⁻) based on O₂-TPD, ¹³C-NMR, and literature study.¹⁶⁹ This was also supported by the ethane and O₂ co-feed experiment with and without co-feeding CO₂ (Fig. 10d). It was shown that the cofeed of CO₂ would decrease the activity for ethylene formation, indicating that the peroxide species is likely suppressed by CO₂ (2O₂²⁻ + 2CO₂ → 2CO₃²⁻ + O₂). Density functional theory (DFT) calculation also indicates that the formation of the peroxide from the Fe⁴⁺ to Fe³⁺ is thermodynamically favorable (Fig. 10e). Based on these findings, the following reaction steps were proposed for the ODH reaction: (1) peroxide formation at the LSF/Li₂CO₃ interface through Fe⁴⁺ → Fe³⁺ transition; (2) dissolution and transport of the O₂²⁻ by Li₂CO₃; (3) H abstraction from C₂H₆ facilitated by O₂²⁻ at the molten-salt/gas interface. The subsequent reactions are likely to proceed through a surface initiated homogenous reaction pathway by releasing ethyl radicals into the gas phase.

Fig. 10 (a) In situ XRD of LSF@10 wt% Li₂CO₃ under CL-ODH cycles at 700 °C, the LSF phase remained stable, with a small amount of (La/Sr)₂FeO₄ formed during the ODH step; (b) in situ XRD of unpromoted LSF under CL-ODH cycles at 700 °C. Decomposition of the LSF phase to La₂O₃, (La/Sr)₂FeO₄, and Fe occurred during the ODH step; (c) electrochemical impedance spectroscopy (EIS) of Li₂CO₃; (d) ethane and O₂ co-feed experiment on pure Li₂CO₃ with and without CO₂ co-feed at 730 °C, 4000 h⁻¹; (e) DFT calculation of peroxide formation with Fe⁴⁺ to Fe³⁺ transition.¹⁶⁷

As one would expect, for different reactions and conditions, the required active species may change. For example, butane ODH to 1,3 butadiene conversion needs to be carried out at temperatures far below Li₂CO₃'s melting point (∼710 °C). As such, LSF@10 wt% Li₂CO₃ would not be effective. As can be seen from Fig. 11, LSF@10 wt% Li₂CO₃ does not have a high selectivity towards 1,3-butadiene.¹⁶⁸ For this reaction, LSF@AX (A = alkali, X = halogen) redox catalysts showed significant butadiene yields whereas the blank experiment indicated minimal thermal cracking.¹⁵⁵ The unmodified LSF exhibited 71% butane conversion, but with ∼93% selectivity towards CO_x products. In contrast, LSF@20 wt% LiBr showed 56.2% 1,3-butadiene selectivity with 75.6% n-butane conversion.¹⁶⁸

Fig. 11 1,3-Butadiene selectivity and conversion profiles for LSF@10 wt% LiBr via CL-ODH of n-butane (our work) compared with literature results with various catalysts at GHSV = 4500 h⁻¹ and 500 °C.¹⁶⁸

Compared to LiBr, other molten and active halogenated salt promoters such as NaBr, KBr and/or LiCl were less effective. This indicates that Br possibly act in conjunction with Li to facilitate selective dehydrogenation of butane. The role and interaction of Li and Br species were probed by ab initio molecular dynamics (AIMD) shown in Fig. 12a, the LiBr is likely oxidized by Li₂O₂ to form atomic Br, which then diffuse to the surface and facilitate H abstraction from butane at the gas/molten salt interface. DFT calculation shown in Fig. 12b indicates that the Br assisted C–H activation have energy barriers of 91.6 kJ mol⁻¹ (butane to butyl radical) and 34.7 kJ mol⁻¹ (1-butene to butenyl radical), a significant decrease compared to the C–H bond dissociation energy (∼400 kJ mol⁻¹). In addition, the desorption energies of the radicals are both ∼30 kJ mol⁻¹; it is therefore more likely that the radicals would be converted to the dehydrogenated counterpart on the surface of the molten salt layer. Additional gas-phase Reaction Mechanism Generator and Chemkin-Pro results indicate the 2-butene, instead of 1,3-butadiene, would be the primary product via gas phase radical reactions in the presence of butyl radical. The absence of 2-butene rejects the surface initiated radical reaction pathway. As such, the as-formed atomic Br would be primarily responsible for all the dehydrogenation steps. Based on these findings, the following reaction steps were proposed: (1) the as-formed peroxide on LSF oxidizes Li₂O to Li₂O₂; (2) Li₂O₂ oxidizes LiBr to atomic Br; (3) Atomic Br then abstracts H from C–H bond from n-butane and 1-butene, sequentially forming 1,3-butadiene and HBr; (4) HBr then react with Li₂O to form LiBr to complete the reaction cycle.

Fig. 12 (a) AIMD calculation of reaction pathway of 2LiBr + Li₂O₂ → 2Li₂O + 2Br; (b) AIMD calculation of reaction pathway of n-butane reacting with atomic Br-containing molten LiBr.¹⁶⁸

Generally speaking, a Type 2 redox catalyst functions through: (1) inhibiting CO₂ formation by covering the mixed oxide with the promoter, this aspect is similar to Type 1 redox catalysts;^167,168 (2) selectively catalyzing ODH reactions through the formation of active species from molten salt and oxide core interactions. It is also worth noting that different promoters in Type 2 redox catalysts can lead to different reaction pathways and mechanisms (Table 3). Selection of the oxide substrate is also important in terms compatibility with the promoter and ability to form desirable active oxygen species such as peroxide ions.

Table 3 Proposed reaction pathway in the solid/gas and gas/liquid interfaces

C2H6 ODH with LSF@Li2CO3	C4H10 ODH with LSF@LiBr
At the LSF/molten Li2CO3 interface:	At the LSF-molten LiBr interface:
2Fe^4+ + 2O^2− → 2Fe^3+ + O2^2−	O2^2− + 2Li2O → O^2− + Li2O2
	Li2O2 + 2LiBr → 2Li2O + 2Br
At the gas-molten Li2CO3 interface:	At gas-molten LiBr interface:
Overall reaction:
C2H6 + O2^2− → C2H4 + O^2− + H2O	Li2O + 2HBr → 2LiBr + H2O

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2.2. Type 3: mixed oxide @ catalytic phase redox catalyst for hydrocarbon conversion

Type 3 redox catalysts are based on a design concept like Type 2 but differs in that a solid catalytic phase is used and a radical formation step is not likely to be involved. For instance, a well-studied K–Fe–O catalyst is highly active towards ethylbenzene DH to styrene.^170,171 By combining the K–Fe–O's catalytic function with transitional metals oxides’ SHC capability, equilibrium limitation can be circumvented, intensifying the overall process.

A notable example is the multi-functional (Ca/Mn)_1−xO@KFeO₂ redox catalyst developed from our group, which showed high activity towards ethylbenzene CL-ODH.¹⁶¹ As shown in Fig. 13a, up to 91% single-pass styrene yield can be achieved. This represents a 72% yield increase, on a relative basis, compared to commercial dehydrogenation. As can be seen in Fig. 13a, there is an unselective region at the beginning of the ODH reaction step, during which the ST selectivity was ∼54%. The selectivity of ST would increase to ∼95% as more ethylbenzene (EB) is injected. This unselective region can be avoided by limiting the air injection during the regeneration step, as shown in Fig. 13b to maintain a high selectivity (∼94.2%). Higher EB partial pressures closer to industrial conditions has also been tested at a lab scale (Fig. 13c), and the styrene yield surpassed the equilibrium limitation. In addition, the stability of this redox catalyst was tested for 100 cycles (Fig. 13d), albeit at a relatively low EB partial pressure.

Fig. 13 Performance of redox-ODH of ethylbenzene with (a) fully reoxidized and (b) partially reoxidized redox catalyst (EB partial pressure = ∼0.01 atm, temperature = 600 °C). (c) Redox-ODH performance comparing to DH equilibrium conversions in the range of 0.01–0.1 atm ethylbenzene feed partial pressure (balance Ar) using partially reoxidized redox catalyst. (d) Long-term cycle and product distributions in redox-ODH using fully reoxidized redox catalyst.¹⁶¹

Fig. 13a and b signify that there were dynamic changes in the redox catalyst. in situ XRD (Fig. 14a) and ex situ XPS (Fig. 14b and c) were used to characterize the changes in the shell and the core at different stages of the redox ODH reactions.¹⁶¹in situ XRD indicated that the unselective region shown in Fig. 13a was resulted from the solid-state reactions between the core and shell phases, forming a nonselective K_0.296Mn_0.926O₂ phase while decreasing the surface concentration of potassium. As the injection of EB continues, this nonselective phase disappeared and the catalytic layer KFeO₂ appeared throughout the rest of the EB injection step. Meanwhile, there was a continuous lattice oxygen release from the CaO–Mn^2+/3+O solid solution, which was responsible for selective hydrogen combustion. From Fig. 14b and c, the Mn 2p peaks shifted to lower binder energy level when in contact with EB (in both reduced and deeply reduced region), while the Fe 2p peaks only shifted after prolonged contact with the EB (deeply reduced region). Both results suggest that Fe³⁺ oxidation state, which is key to high DH activity, is reserved by sacrificing Mn^3+/4+ in the operating regime.

Fig. 14 (a) In situ XRD under cyclic ethylbenzene ODH and air reoxidation steps at 600 °C; (b) XPS of Mn 2p and (c) XPS of Fe 2p for redox catalysts at different reaction stages.¹⁶¹

The reaction mechanism was further studied by isotope experiments in our group. Similar to conventional catalytic DH, during the dehydrogenation reaction, ethylbenzene on the CLCa catalyst undergoes a two-step H abstraction on the KFeO₂ surface.^170–172 The KIE value (Fig. 15a) measured in the isotope experiment confirmed that H abstraction was the rate limiting step. This was further supported by the DFT results (Fig. 15b-1), where the α-H abstraction step was shown to be rate limiting. DFT results also indicated that the presence of water would facilitate the proton transfer in the water formation step on the KFeO₂ surface, leading to more than 4-fold decrease in activation energy (Fig. 15b-2). This was verified by dehydrogenation experiments on KFeO₂ both with and without cofeeding steam.

Fig. 15 (a) Kinetic isotope effect study with both temperature-programmed reduction and isothermal reduction of the (Ca/Mn)_1−xO@KFeO₂ using C₈H₁₀ and C₈D₁₀¹⁶¹; (b) computed energy potential profiles of (1) α and β H abstraction and (2) proton transfer and water formation with or without water assistance.¹⁶¹

Similar to Type 2 redox catalysts, design of Type 3 needs to consider the compatibility of the core and shell materials. The dynamic changes observed during the redox reactions signify the potential complexity of CLCa when compared to conventional heterogeneous catalysis systems.¹⁶¹

2.3. Synergistic effects between heterogeneous and redox catalysts

As briefly mentioned in Section 2.1, the abovementioned redox catalyst design, particularly Type 1, has the potential to function in concert with conventional heterogeneous catalysts. We recently demonstrated this approach towards catalytic cracking of cyclohexane by integrating cation exchanged ZSM-5 with a Type 1 redox catalyst, i.e. perovskite oxide@Na₂WO₄.¹⁵⁷ The primary function of the redox catalyst was to selectively oxidize hydrogen and to enhance the naphtha conversion. It is interesting to note that a significant increase in ethylene (C₂⁼) and propylene (C₃⁼) selectivity and C₃⁼/C₂⁼ ratio was observed by mixing the ZSM-5 based catalyst with the Type 1 redox catalyst (Fig. 16a and b). The presence of the perovskite oxide@Na₂WO₄ did not deactivate the zeolite catalyst. Rather, it changed the acidic properties of the zeolite by creating additional strong Brønsted acid sites, which is likely to be responsible for the enhanced olefin selectivity and yield. A similar concept was also shown to be effective for n-hexane conversion: Ba-exchanged ZSM5 mixed with CaMnO₃@Na₂WO₄ could exhibit up to 67% single-pass olefin yield from n-hexane with tunable propylene to ethylene ratio, with <5% CO₂ yield.¹⁷³ We note that in both cases, the nature of the interactions between ZSM-5 and the redox catalyst particles, as well as the fundamental reasons for the change in product selectivity have yet to be understood.¹⁵⁷ This represents an interesting area for in-depth studies. Synergistic utilization of heterogeneous catalysts with redox-active oxides can also significantly broaden the applicability for CLCa.

Fig. 16 Cyclohexane conversion and product distributions over (a) HZSM5 and different metal cation (Sr, Ca, Cr or Fe) exchanged ZSM-5; (b) composite composed of ZSM-5s and CaMn_0.75Fe_0.25O_3−δ@Na₂WO₄ redox catalyst.¹⁵⁷

When integrating CLCa with a heterogeneous catalyst, an important aspect to consider is whether the heterogeneous catalyst would function well under significant oxygen partial pressure swings. In addition, steam generated from hydrogen combustion by the CLCa redox catalyst can have a negative impact on the active sites of the catalyst and can lead to side reactions (e.g. steam reforming). For example, MoC_x, active for dehydroaromatization (DHA) of methane,^174–181 can react with steam produced from SHC to form MoO_x, which would deactivate the catalyst. Therefore, care must be taken in selecting robust and compatible hybrid catalyst systems involving both chemical looping and heterogeneous catalysts.

2.4. Computational assisted oxide selection

Although the previous section focused on surface promoters for the redox-active oxides, the selection of the oxide also represents a crucial and enabling step to CL-ODH. In fact, a sizable amount of chemical looping research focused on discoveries and the selections of oxide-based oxygen carriers. To date, over 2000 research articles have covered different aspects of oxygen carrier compositions, performance, and underlying mechanism. Despite the extensive research and development efforts, a rationalized strategy for the effective design and selection of oxygen carriers is still lacking. In fact, the high complexity of redox reactions, which extends beyond the interactions of gaseous molecules with poorly defined oxide surfaces, makes it near impossible for ab initio design of redox-oxides. On the other hand, the emergence of oxygen carriers composed of redox-active mixed oxides, e.g. perovskite oxides, has greatly expanded the design space for chemical looping material selection, making oxygen carrier optimization an even more daunting task.

We note that while it is impractical to expect an ab initio oxygen carrier selection model to comprehensively consider all aspects of oxygen carrier design, some relatively simple selection criteria can nevertheless be adopted, particularly when capitalizing on modern computational and data science tools, to greatly reduce the experimental efforts in oxygen carrier development. Take CLCa of light alkanes as an example, the light olefin yields tie directly with the equilibrium oxygen partial pressure (P_O2,eq) or equivalently oxygen chemical potential (μ_O2,eq) of the redox oxide from a thermodynamic standpoint (Fig. 17a). In fact, the thermodynamic selection rule of the oxides in CLCa of alkane ODH largely overlaps with that of chemical looping combustion and chemical looping air separation (CLAS). This is not at all surprising given that a primary function of the oxygen carrier in CLCa is in situ air separation, besides acting as a heterogeneous catalyst. Given that the thermodynamic requirement represents a prerequisite for oxygen carrier selection, we recently used such a criterion to select oxides with a general formulation of Sr_xA_1−xFe_yB_1−yO_3−δ (A = Ca, K, Y, Ba, La, Sm; B = Ti, Ni, Mn, Mg, Cu, Co) for applications such as CLAS and chemical looping dry reforming of methane. Fig. 17b illustrates the generalized simulation and experimental validation strategies as well as representative results.¹⁸² Despite adopting simplified assumptions to facilitate density functional theory (DFT) based high throughput computations, the simulation results were shown to be quite effective in predicting various perovskite oxides with interesting properties, that could be adopted in the both CLAS or CLCa.¹⁸² These redox active oxides materials have excellent potential to be adopted for CLCa. We also note that the proposed model is not without limitations. For instance, the model was based exclusively on thermodynamic criteria without considering kinetic effects. To ensure manageable computational intensity, we assumed random distribution of the dopants and did not consider potential defect clustering. The U parameters selected to correct the strong on-site coulomb interaction of the d-electrons may need to be further tuned for the perovskite oxides. Potential formation of alternative phases and structures besides the base SrFeO₃ structure were not considered. These represents potential areas of improvements for future modelling efforts. To date, only a few other groups have also use high-throughput method for materials screening.^183,184 Fan et al. screened over 1500 M₁–M₂–O/N pairs and 170 M₁–M₂–N_rich/N_poor pairs to identify optimized bicationic materials for CL-ammonia synthesis from different route.¹⁸³In the context of CLC of methane, Singstock et al. screened over 1300 redox pairs (i.e. ABO_x/AO_y + B, A(SO₄)_x/AS_x) and predicted 152 pair to demonstrate >99% methane conversion with limited byproducts.¹⁸⁴

Fig. 17 (a) Relationship between the equilibrium ethane conversion in CL-ODH and the equilibrium P_O2 of the redox oxide at 600 and 800 °C (only the ODH reaction is considered in this calculation without considering the dehydrogenation reaction); (b) a general strategy for computationally guided oxide selection and the key results related to CLAS material selection published by Wang et al.¹⁸²

3. Summary and outlook

The pressing need for decarbonizing the modern society calls for transformative technologies to intensify state-of-the-art manufacturing processes. The chemical looping strategy, with an inherent capability for chemical transformation combined with feedstock and/or product separations, offers unique opportunities to significantly simply the chemical manufacturing processes, improve energy efficiency, and reduce CO₂ emissions.

This review covered a generalized redox catalyst design strategy, involving a redox-active mixed metal oxide core and a surface promoter shell, for the CL-ODH of hydrocarbons to olefins. An interesting aspect of this design strategy is the ability to independently tune the properties of the oxide core and the promoter shell provided that the two materials are compatible under the operating conditions. For example, for ethane conversion with Type 1 catalyst, we recently predicted that C₂₊ yield could achieve 78.2–86.7% based on the existing data from molten salt promoted Mg₆MnO₈ and Cu/MnO_x. These analyses suggest that further increase in single-pass olefin yield is highly possible by optimizing the core and the shell compositions and ratios. For Type 2 redox catalysts, the heat of reactions of ethane CL-ODH can be tuned to be mildly exothermic in both the ethane ODH and the re-oxidation steps, allowing more efficient heat management. On the other hand, the selection of the core would also affect the activity and selectivity of the redox catalysts. In the case of LSF, Fe⁴⁺ in the perovskite structure was proposed to be responsible for the formation of the peroxide in molten phase. As an example, further adjusting the La:Sr ratio could tune the propensity to form peroxide and therefore optimize the ODH performance. Type 3 redox catalysts, and the potential to integrate CLCa redox catalyst with a heterogeneous catalyst in tandem, also offer exciting opportunities.

From a practical standpoint, molten salt promoted oxides may not be easily fluidizable. However, it is possible to adjust the surface promoter compositions to ensure good fluidization properties and long-term stability.¹²⁹ Molten salt promoted redox catalysts have also been operated in packed bed reactors in our recent studies, showing satisfactory long-term stability up to 1000 hours.^119,120 Other issues to consider for future studies and practical applications include the potential vaporization of molten salts at high temperatures and reactor material selection to minimize potential corrosion. A recent long term study on Li₂CO₃ promoted LSF showed signs of slight increase in CO_x selectivity.¹¹⁹ This may be attributed to the loss of Li₂CO₃. Therefore, replenishing of Li₂CO₃ either ex situ or during the reactor operation may be considered for future studies.

The application of chemical looping concept is not limited to olefin productions. Table 4 summarizes recent publications related to CLCa. An extensively investigated topics is methane to syngas/H₂ conversion.^{95,109–117,123,125–127,130,136,185,193–208} A few other studies explored a phase transition sorbent concept where a single mixed oxide particle combines the function of a redox oxide (for oxygen separation), a carbonate sorbent (for CO₂ separation), and a catalyst (to catalyze chemical reactions).^{95,123,130,197,198,209–211} For instance, a recent study adopted the chemical looping strategy to selectively oxidize ammonia with excellent NO selectivity for nitric acid production.¹⁹² A few other studies explored the feasibility of using nitrides/imides/hydrides^{131–134,189–191,212–214} as intermediates for ammonia synthesis and using alkali molten salt as the medium for combined CO₂ capture and utilization.²¹⁵ Engineering the chemical potential of redox oxides has also shown to be able to circumvent reaction equilibrium limitations in conventional processes while maintaining high selectivity, making it possible to achieve previously unattainable yields for water gas shift reactions.^126,194

Table 4 Selected example of recent progress in chemical looping catalysis

Target product	Carrier	Operating condition	Key performance result	Ref.
Syngas/CO	Fe2O3@SBA-15	750–935 °C, 17.8–37.5 mL per h per mgFe2O3	∼100% CO selectivity	108
	La0.5Ce0.5FeO3	850 °C, flow rate: 80 mL min^−1	X methane: 82%	185
		0.8 g catalyst	S Syngas: 93%
		Reduction: 10 mol% CH4/N2
		Oxidation: 10 mol% CO2/N2
	BaFe1−xSnxO3−δ	900 °C, 0.1 g catalyst	>99% syngas selectivity	114
		Reduction: 5% CH4/He (20 mL min^−1)	Y syngas = 19.2 mmol g^−1
		Oxidation: 5% O2/He (20 mL min^−1)
	MgO supported Ca2Fe2O5	1000 °C, 2 g catalyst 1500 h^−1	X methane > 99%	117
		Reduction: 10% CH4/He (25 mL min^−1)	S Syngas: 98.1%, H2/CO ∼2
		Oxidation: 40% CO2/He (25 mL min^−1)
	Ni-doped Fe2O3/Al2O3	900 °C, 2 g catalyst 1500 h^−1	X methane > 95%	186
		Reduction: 20% CH4 (75 mL min^−1)	S Syngas: ∼96%, H2/CO ∼2.3
		Oxidation: 20% O2 (75 mL min^−1)
Ethylene (via CL-OCM)	(Li,W)–Mg6MnO8	850 °C, 2400 h^−1	X methane: 50%	106
		Reduction: pure% CH4 (200 mL min^−1)	C2+ Yield: 28.6%
Ethylene (via CL-ODH)	0.2Ce/SrFeO3	700∼750 °C 3000∼6000 h^−1	Ethylene Yield: 23.8%	100
		Reduction: ∼12.5 mol % C2H6	X CO2 69.4%
	Sr0.8Ca0.2FeO3−δ@Na2CO3 + MoVNbTeOx	500 °C, 600 h^−1 GHSV	S ethylene: 91%	96
		Reduction: 6 mol% C2H6, balance N2	X ethane: 42%
	Ce-incorporated FeTiOx	600 °C, reduction: 20 vol% C2H6/N2, 30 mL min^−1	S ethylene: 84.1%	187
		Oxidation 20 vol% CO2/N2, 30 mL min^−1	X ethane: 14%
Propylene (via CL-ODH)	Mo–V–O (V/Mo = 6)^188	500 °C, 2500 h^−1 GHSV	X propane: 36%	188
		Reduction: pure% CH4 (200 mL min^−1)	S propylene: 89%
Ammonia	Li–Pd/Li2NH	300 °C and 1 bar	r NH3 = 693 μmolNH3 g^−1 h^−1	189
		100% N2 or H2 at 30 mL min^−1
	Si modified γ-Al2O3 ZrO2 supported	1 bar	X AlN = 60.4%	190
		N-sorption: 1200 °C 100 mL min^−1 N2
		NH3-desorption: 1200 °C 80 vol% H2O 500 mL min^−1
	20 wt%TiN/MgO	Nitridation: 1 atm, 808 nm laser at 30 W cm^−2 (∼550 °C)	r NH3 = 1.67 μmolNH3 g^−1 h^−1	191
NO (NH3 oxidation)	V2O5	300–650 °C	NO selectivity: 99.8%	192
		Reduction: 5 mol % NH3/Ar, 30 mL min^−1	NH3 conversion: 97%
		Oxidation: 30 mol % O2/Ar, 30 mL min^−1

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Most of the processes summarized above harness the oxidizing potential (e.g. SHC) of the metal oxides at high temperatures during the half cycle for metal oxide reduction. The reduced metal oxides are regenerated in an O₂, CO₂, or steam rich environment, allowing integrated air separation, CO₂-splitting, or hydrogen production. It can be fully anticipated that well-designed metal oxides, after reduction, could be used for producing complex, value-added chemicals via hydrogenation and/or carbonylation during the regeneration half cycle. This offers the potential to produce two different value-added products in a two-step redox process or even multi-step chemical loops. Therefore, many new and highly exciting areas of fundamental research and practical applications can be anticipated through the unique combination of chemical looping with heterogeneous catalysis.

Conflicts of interest

There are no conflicts to declare.

Acknowledgements

This work was supported by the U.S. National Science Foundation (Award No. CBET-2116724 and CBET-1923468), US Department of Energy (Award No. FE0031869, EE0008809, FE0031918), and the Kenan Institute for Engineering, Technology and Science at NC State University.

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