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Data will be made available on request.Fossil fuels are essential building blocks in the petrochemical industries for producing materials such as plastics, synthetic fibres, rubbers, lubricants, detergents, and solvents (Keim, 2010; Speight, 2011). However, their non-renewable nature poses a sustainability risk, prompting the search for sustainable alternatives based on renewable biomass sources (Ozturk et al., 2017). Oil produced from pyrolysis of lignocellulosic biomass, such as oil palm biomass, can be a sustainable alternative to fossil fuels, given its carbon-neutral properties with low sulfur and nitrogen content (Martínez et al., 2014; Palamanit et al., 2019). During pyrolysis, the thermal decomposition of oil palm biomass may produce more than 300 chemical compounds in the oil, which can be used as precursors for synthesizing petrochemical products (Keim, 2010; Machado et al., 2022). Oil palm biomass-derived pyrolysis oil mainly consists of oxygenated compounds due to its high oxygen content in raw biomass. Such oil requires modification to improve the hydrocarbon content (Palamanit et al., 2019). Co-pyrolysis of oil palm biomass with plastics like polypropylene (PP) is a promising method for increasing hydrocarbon content (Al-Maari et al., 2021). PP, rich in carbon and hydrogen, provides the hydrocarbon pool required for the deoxygenation reaction of oxygenated compounds from biomass to form hydrocarbons such as aliphatic and aromatic hydrocarbons in oil.Solid acidic catalysts can further promote the deoxygenation reactions (i.e., dehydration, decarbonylation, and decarboxylation) of pyrolytic volatiles to improve the hydrocarbon content in pyrolysis oil (Hassan et al., 2019; Shafaghat et al., 2019; Zhang et al., 2016). Due to its excellent catalytic performance for deoxygenation, which generates hydrocarbons such as olefins, aliphatic, and aromatic hydrocarbons, high-acidity zeolites have been widely used in several studies (Balasundram et al., 2018; Wang et al., 2020). In such a reaction, oxygen is typically removed by releasing by-products such as water, carbon dioxide, and carbon monoxide (Hassan et al., 2019). Zeolites, nonetheless, are microporous catalysts. Thus, micropore-related flow restriction can affect their deoxygenation catalytic performance, especially if relatively large molecules such as lignin-derived compounds are involved (Shafaghat et al., 2019). Such flow restriction can also cause coke formation, create pore blockage, catalyst deactivation, and catalyst poisoning, thereby reducing the performance of catalysts (Hassan et al., 2019; Shafaghat et al., 2019). To address this, mesoporous acidic catalysts such as titania (TiO2) and alumina (Al2O3) based catalysts with larger pore sizes were introduced, allowing large molecules to diffuse and reducing pore blockage and catalyst deactivation (Lu et al., 2010; Zhou et al., 2019). The high chemical and thermal stabilities of TiO2 and/or Al2O3-based catalysts have also sparked interest (Bagheri et al., 2014; Paranjpe, 2017). It has been proposed that doping of metals such as nickel, copper, molybdenum, cobalt, palladium, and cerium into TiO2 and/or Al2O3-based catalysts can improve deoxygenation (Bagheri et al., 2014; Lu et al., 2010; Zhou et al., 2019).Several works have evaluated TiO2 and/or Al2O3-based catalysts in oil upgrading through catalytic pyrolysis and co-pyrolysis in a nitrogen atmosphere. Dong et al. (2019) compared the catalytic performances of titania-based catalysts doped with different metals, including copper (10% Cu/TiO2), iron (10% Fe/TiO2), and molybdenum (10% Mo/TiO2) on the phenol conversion during the catalytic pyrolysis of corn straw lignin at 450 °C. They reported that the highest total phenol conversion was attained using 10% Mo/TiO2, followed by 10% Cu/TiO2, TiO2, and then 10% Fe/TiO2. Lu et al. (2010) studied the catalytic upgrading of oil from pyrolysis of poplar wood using the titania, zirconia, and titania-zirconia-based catalysts doped with cerium, ruthenium, and palladium at 600 °C. In general, all the catalysts reduced the sugars (i.e., levoglucosan) in the oil, while titania-zirconia-based catalysts yielded a high amount of hydrocarbons and ketones. TiO2-based catalysts promoted the formation of phenols. Mysore Prabhakara et al. (2021) investigated the catalytic performance of γ-Al2O3, dolomite, and hydrotalcite (HTC) MG70 with the addition of 20 wt.% Na2CO3 into the catalysts during the catalytic pyrolysis of beechwood at 500 °C. All these catalysts significantly reduced the oxygenated compounds and enhanced the formation of aliphatic, monoaromatic, and polyaromatic hydrocarbons. Zhou et al. (2019) investigated the utilization of NiO/γ-Al2O3 catalyst on the dehydration reaction mechanism during the pyrolysis of rice husks. Weak acid sites on Al2O3 were discovered to facilitate the dehydration reaction the most throughout the process. In addition, Imran et al. (2014) reported that the alumina-supported sodium carbonate (Na2CO3/γ-Al2O3) catalyst improved the quality of oil from the pyrolysis of wood fibers.No studies have used titania and alumina-based catalysts in the co-pyrolysis of OPT and PP to improve the targeted oil composition. The mesopores in these catalysts may facilitate the diffusion rate of large molecules (i.e., compounds derived from the thermal degradation of OPT and PP) through the pores of the catalysts and promote the conversion into hydrocarbons during catalytic co-pyrolysis. This study investigated the catalytic performance of a titania-based catalyst doped with nickel-molybdenum (Ni–Mo/TiO2) and an alumina-based catalyst with nickel (Ni/Al2O3) for the upgrade of oil generated from co-pyrolysis of OPT and PP. The effect of the catalysts on the oil composition was evaluated.OPT was collected from an oil palm plantation in Saratok, Sarawak. OPT was pre-dried in the oven at 105 °C for 24 h, ground (Fritsch rotary mill, PULVERISETTE 14), and sieved (Fritsch sieve shaker, ANALYSETTE 3 PRO) to obtain the samples with a particle size of 500 μm and below. Locally sourced PP food containers were cut into smaller sizes and sieved using a sieve shaker (Fritsch, ANALYSETTE 3 PRO) to obtain samples with a particle size of 500 μm and below. The sieved PP was stored under ambient conditions before use. Two catalysts used in this study, Ni–Mo/TiO2 and Ni/Al2O3, were synthesized based on the impregnation method reported by Aqsha et al. (2015).The catalysts’ specific surface area, average pore diameter, and pore volume were determined via nitrogen adsorption-desorption isotherm analysis (Brunauer–Emmett–Teller (BET) surface area and pore size analyzer, Quantachrome Nova 4200e). Before the analysis, the samples were degassed at 200 °C for 12 h to remove any surface-adsorbed residual moisture.The crystallinity of the catalysts was investigated using powder X-ray diffraction (XRD) (X-ray Diffractometer, Rigaku SmartLab). Cu-Kα radiation (λ = 0.154 nm) was used to measure the diffraction patterns in the range of 2θ from 5 to 100°.XRF was used to analyze the composition of the catalysts with an accelerating voltage of 15 kV and a current of 30 μA (Bruker S2 PUMA).The acidity of the catalysts was determined through NH3-temperature programmed desorption (TPD) analysis (Micromeritics Chemisorb 2750). The sample was pre-treated by heating it from room temperature to 200 °C in helium gas flow for 120 min. Adsorption of NH3 was carried out at 100 °C for 60 min (5% in He, v/v), followed by helium purging at the same temperature for another 60 min. Following that, NH3 desorption was carried out by heating from 50 to 800 °C at a ramping rate of 10 °C min-1 and holding at the final temperature of 800 °C for 15 min.The catalytic co-pyrolysis was carried out in a horizontal tube furnace (MTI, GSL-1100X) with a 400 mL min-1 nitrogen flow rate to form an inert condition in the tube furnace. 3 g of OPT and PP mixture sample (weight ratio of OPT: PP of 1:1) with 0.3 g of catalyst were loaded into the reactor and nitrogen purged for 5 min. The reactor was heated to the desired operating temperature (i.e., 500, 600, and 700 °C) at a heating rate of 10 °C min-1, with a holding time of 40 min. Afterwards, the reactor was cooled down to 200 °C while continuously purged with nitrogen gas. A cold trap in an ice bath (2–3 °C) was connected to the tube reactor outlet to collect the liquid product (oil) from the experiment. The collected oil was stored at 2–7 °C until further analysis. The non-condensable gases were released into the environment. The product yield obtained from the experiments was calculated using Equations (1)–(3) . (1) P y r o l y s i s o i l y i e l d ( w t . % ) = M a s s o f p y r o l y s i s o i l o b t a i n e d ( g ) M a s s o f s a m p l e ( g ) x 100 % (2) S o l i d y i e l d * ( w t . % ) = M a s s o f s o l i d o b t a i n e d ( g ) M a s s o f s a m p l e ( g ) x 100 % *Solid yield refers to all solid residues collected from the experiments, including feedstock residue, catalysts, and coke. (3) G a s y i e l d ( w t . % ) = 100 w t . % – p y r o l y s i s o i l y i e l d ( w t . % ) – s o l i d y i e l d ( w t . % ) The composition of pyrolysis oil was determined using a gas chromatography-mass spectrometer (GC-MS) with an HP-5MS column (Agilent, 30 m length x 0.25 mm inner diameter x 0.25 m film thickness) (Agilent, 6890 N). The column oven was programmed to operate at 40 °C for 3 min. Afterwards, it was heated from 40 to 200 °C at the rate of 8 °C min-1 with a holding time of 10 min. The temperature was then ramped from 200 to 220 °C at a rate of 10 °C min-1 and held for 10 min. The column was kept at a pressure of 7.04 psi and a flow rate of 1 mL min-1 of helium. The split ratio of 50:1 was used in the analysis. Before the analysis, 0.2 g of pyrolysis oil was diluted in 10 mL of acetone. A syringe filter was used to filter the diluted oil sample before it was transferred to the GC sample vial and injected into the equipment via auto-injection mode for analysis. Compounds were identified by comparing the NIST08 mass spectral data library entries. Table 1 presents the textural properties (i.e., specific surface area, pore volume, and average pore diameter) of Ni–Mo/TiO2 and Ni/Al2O3. The lower specific surface area of Ni–Mo/TiO2 compared to Ni/Al2O3 could be attributed to the accumulation of two types of metal particles on the catalyst's surface or within its pores (Kumar et al., 2019). Both catalysts are categorized as mesoporous since their average pore diameter sizes are between 2 and 50 nm (Thommes et al., 2015). The large pores allow large molecules, such as lignin-derived compounds, to flow in and out of the catalysts' pores for higher conversion of the compounds during the catalytic co-pyrolysis (Lu et al., 2010). Fig. 1 depicts the acidities of the catalysts analyzed with NH3-TPD, which reveals the acid site distribution. The temperature region where the ammonia desorption peak has located indicates the types of acid sites (i.e., weak, medium, and strong acid sites) on the surface of both catalysts. The weak acid sites correspond to the ammonia desorption peak at temperatures less than 250 °C. In comparison, the medium acid sites appear in the temperature region between 250 and 500 °C. The ammonia desorption peak, which appears at temperatures above 500 °C, represents strong acid sites (Phan et al., 2020). Strong acid sites with higher acid strength likely provide higher catalytic cracking activity for converting the compounds into desirable products through the catalyst (Li et al., 2020). Fig. 1(a) shows that most ammonia desorption peaks are between 250 and 500 °C, indicating the presence of medium acid sites for Ni–Mo/TiO2. On the other hand, weak, medium, and strong acid sites are present on Ni/Al2O3 catalyst surface as the ammonia desorption peaks are detected in all three temperature regions (Fig. 1(b)). A higher peak intensity value in Ni–Mo/TiO2 relative to that in Ni/Al2O3 contributes to the higher acidity in the former catalyst (Table 1). Fig. 2 shows the powder XRD patterns of Ni–Mo/TiO2 (upper) and Ni/Al2O3 (bottom) catalysts, respectively. Numerous peaks appear on the pattern of Ni–Mo/TiO2, indicating the presence of a mix of phases. Indexing reveals three major oxide phases, i.e., anatase (TiO2), molybdenum oxide (Mo9O26), and nickel oxide (NiO2). An intense peak at 2θ of 25.3° is detected for TiO2 phase, along with weak peaks at 2θ of 37.8°, 48.0°, 53.9°, 55.1°, and 62.7° (COD#96-720-6076). For Mo9O26 phase, intense peaks are observed at 2θ of 24.9° and 25.3°, while weak peaks are present at 2θ of 27.3°, 32.2°, and 33.0° (ICSD#98-002-7510). NiO2 has a weak characteristic peak at 2θ of 37.1° (ICSD#98-007-8698). Ni/Al2O3 has two-phase components, i.e., nickel oxide (NiO) and alumina (Al2O3). The intense peaks of NiO are observed at 2θ of 37.2° and 43.3° while the weak peak is detected at 2θ of 62.93° (ICDD#03-065-6920). On the other hand, the characteristic peaks of Al2O3 are observed at 2θ of 46.0° and 66.8° (ICDD#00-004-0858). The catalyst's composition from XRF analyses is presented in Table S1 in Supplementary Information. Fig. 3 depicts the product yield obtained from non-catalytic and catalytic co-pyrolysis of OPT and PP with Ni–Mo/TiO2 and Ni/Al2O3 at temperatures ranging from 500 to 700 °C. The solid yield in Fig. 3 refers to all solid residues collected from the experiments, including feedstock residue, catalysts, and coke. When the temperature rises from 500 to 700 °C, the solid yield decreases for non-catalytic and catalytic conditions due to the decomposition of char present in the solid fraction into the oil and gas with the rising temperature. According to Zhou et al. (2013), char formation is more favorable at a lower temperature (450 °C) due to the lower decomposition rate of the feedstocks. At temperatures above 450 °C, the decomposition of feedstocks into condensable volatiles and non-condensable gases improves while char formation decreases. The pyrolysis oil yield in non-catalytic co-pyrolysis is maintained at 16 wt.% from 500 to 600 °C and drops to 11.50 wt.% when the temperature rises to 700 °C. The pyrolysis oil yield in catalytic co-pyrolysis with Ni–Mo/TiO2 increases from 12.67 to 19.50 wt.% with the rise in temperature from 500 to 600 °C. Further increase of temperature to 700 °C reduces the oil yield to 17 wt.% due to the enhancement of the secondary reactions of the primary volatiles into the gaseous products at higher temperatures (>600 °C) (Fan et al., 2017; Zhou et al., 2013). The highest pyrolysis oil yield obtained from catalytic co-pyrolysis of Ni/Al2O3 is 17.17 wt.% at 500 °C, followed by a reduction to 12.33 wt.% at 600 °C. Such oil yield reduction is likely due to the increase of gas yield by 10 wt.% at this temperature. Ni/Al2O3 has been shown to improve the formation of gaseous hydrocarbons rather than liquid hydrocarbons during catalytic cracking of OPT and PP (Lin et al., 2020; Singh et al., 2019; Xue et al., 2017). This finding is consistent with the lower amount of liquid hydrocarbons obtained at 600 °C, as shown in Fig. 4 . On the other hand, the gas yield increases with rising temperatures from 500 to 700 °C for three cases (Fig. 3). The secondary reaction of primary volatiles into lighter compounds at higher temperatures results in the formation of non-condensable gases, increasing gas yield with temperature (Hassan et al., 2019). Fig. 4 shows the oil composition obtained from the non-catalytic and catalytic co-pyrolysis of OPT and PP with Ni–Mo/TiO2 and Ni/Al2O3 in the temperature range of 500–700 °C. The oil from non-catalytic co-pyrolysis consists mainly of oxygenated (39.74–52.10%) and phenolic compounds (34.01–41.85%). The oil contains a small amount of hydrocarbons (5.19–10.22%), as evidenced by the relatively low GC-MS relative area for these components. During non-catalytic co-pyrolysis, the oxygenated and phenolic compounds are generated from the thermal decomposition of OPT (i.e., hemicellulose, cellulose, and lignin) (Palamanit et al., 2019; Stefanidis et al., 2014). The thermal degradation of PP produces hydrocarbons via a series of reactions that include random chain scission, mid-chain β-scission, end chain β-scission, radical recombination, and hydrogen transfer reactions (Singh et al., 2019; Xue et al., 2017).When Ni–Mo/TiO2 and Ni/Al2O3 are used as the catalysts in the co-pyrolysis of OPT and PP, the hydrocarbons contained in the oil are significantly increased, as shown by an increase in the GC-MS relative area of up to 54.07–58.18% and 37.28–68.77%, respectively (Fig. 4). The amount of phenolic compounds is reduced, with the reduction in the GC-MS relative area for Ni–Mo/TiO2 (down to 8.46–20.16%) and Ni/Al2O3 (down to 2.93–14.56%). The presence of catalyst generally reduces the amount of oxygenated compounds, although no clear trend can be drawn concerning the parametric effect of temperature and catalyst type. Fig. 5 illustrates the proposed reaction mechanism for the hydrocarbon formation from the analyses based on relevant previous works (Dai et al., 2020; Lin et al., 2020; Singh et al., 2019; Xue et al., 2017). The increase of the hydrocarbon content in the catalytic co-pyrolysis is due to the catalytic cracking of PP and deoxygenation of oxygenated and phenolic compounds promoted by Ni–Mo/TiO2 and Ni/Al2O3 catalyst in addition to the thermal decomposition of PP (Fig. 5).The two catalysts used here rely on the presence of both metal (Ni and Ni–Mo) and acidic (TiO2 and Al2O3) sites to provide high deoxygenation ability and thus improve hydrocarbon production. The oxygenated and phenolic compounds undergo deoxygenation reactions via dehydration, decarbonylation, and decarboxylation to form hydrocarbons (Fig. 5) (Dai et al., 2020). The oxygen in the oil is removed during deoxygenation reactions with water, carbon dioxide, and carbon monoxide released as by-products. The acidic sites in the two catalysts, TiO2 and Al2O3, tend to the occurrence of dehydration reaction over decarbonylation and decarboxylation reactions, resulting in the removal of oxygen from the oil and its subsequent combination with hydrogen to form water as a by-product (Ding et al., 2020). This reaction pathway nonetheless consumes the hydrogen in the oil, which is required to produce hydrocarbon. The presence of metal sites, namely Ni and Ni–Mo, in the two catalysts is expected to partially counteract this pathway, resulting in a more dominant occurrence of decarbonylation and decarboxylation reactions in Ni–Mo/TiO2 and Ni/Al2O3-catalyzed co-pyrolysis of OPT and PP (Balasundram et al., 2018; Dai et al., 2020). Higher acidity of Ni–Mo/TiO2 relative to Ni/Al2O3 (Table 1) due to more abundant acidic sites and synergy between Ni and Mo leads to the formation of a higher amount of hydrocarbons from the catalytic co-pyrolysis of OPT and PP (Fig. 4).During the catalytic co-pyrolysis, the hemicellulose, cellulose, and lignin present in OPT undergo thermal decomposition to produce primary products or intermediates. Afterwards, these products and intermediates diffuse through the pores of Ni–Mo/TiO2 and Ni/Al2O3 and undergo catalytic cracking and deoxygenation reactions to produce secondary products (Balasundram et al., 2018; Lin et al., 2020). The thermal decomposition of hemicellulose primarily yields ketones, furans, and acids, which are then catalytically cracked into smaller oxygenates (i.e., acetic acid, acetone, and simple furans) and olefins on the acidic sites of the catalysts (Dai et al., 2020). Conversely, cellulose is degraded to form anhydrosugars as primary products (Lin et al., 2009). The acidic sites in the two catalysts aid in the dehydration of anhydrosugars to produce more furans. Likewise, the catalytic cracking and deoxygenation of furans form smaller oxygenates and olefins (Dai et al., 2020; Praveen Kumar and Srinivas, 2020). Table 2 shows the decrease of acids and furans in pyrolysis oil after adding Ni–Mo/TiO2 and Ni/Al2O3 catalysts. The result suggests their conversion into olefins, which are the important precursors for the formation of hydrocarbons (Peng et al., 2022). The presence of Ni and Mo in Ni–Mo/TiO2 promotes the decarbonylation and decarboxylation of oxygenated compounds (i.e., ketones, acids, and furans), producing olefins for the subsequent production of hydrocarbons (Balasundram et al., 2018; Xue et al., 2021). Despite this, the amount of ketones in the oil increases after adding these two catalysts (Table 2). This is likely due to catalyst-promoted radical interactions between OPT and PP (Lin et al., 2020).Compared to hemicellulose and cellulose, lignin has a more complex structure, thus producing larger molecules of oligomers during thermal decomposition (Jiang et al., 2010; Lu et al., 2010; Stefanidis et al., 2014). The mesoporous structure of Ni–Mo/TiO2 and Ni/Al2O3 catalysts with wide channels allow for higher diffusion of these lignin-derived oligomers, resulting in high conversion into simple phenols (Lu et al., 2010), which are then converted into olefins via deoxygenation (Hassan et al., 2019; Xue et al., 2017). During the thermal decomposition of PP, olefins can be produced via radical recombination and hydrogen transfer reactions of PP-derived radicals (Singh et al., 2019; Xue et al., 2017). These olefins would produce cyclic hydrocarbons via isomerization and oligomerization. The acidic sites in the catalysts have previously been reported to aid in the isomerization and oligomerization reactions resulting in the formation of cyclic hydrocarbons. Fig. 6 shows a higher amount of cyclic hydrocarbons in the oil derived from the catalytic co-pyrolysis than that from the non-catalytic co-pyrolysis (Peng et al., 2022).Aliphatic hydrocarbons, on the other hand, are produced during PP decomposition through random chain scission, β-scission, radical recombination, and hydrogen transfer reactions (Singh et al., 2019; Xue et al., 2017). Fig. 6 depicts an increase in aliphatic hydrocarbons in the oil produced by catalytic co-pyrolysis compared to non-catalytic co-pyrolysis. The metal sites (i.e., Ni and Ni–Mo) in the two catalysts promote the hydrogen transfer reactions (Peng et al., 2022). The presence of Mo in Ni–Mo/TiO2 promotes the transfer of electrons from Mo to Ni, which enhances the catalyst's electron density and thus improves the hydrogen transfer reaction (Maluf and Assaf, 2009). The lower relative amount of aliphatic hydrocarbons observed in Fig. 6 compared to cyclic hydrocarbons is consistent with the nature of aliphatic hydrocarbons as intermediates. Furthermore, some aliphatic hydrocarbons may go through additional isomerization and oligomerization reactions to become cyclic hydrocarbons, facilitated by the acidic sites of the catalysts (Xue et al., 2017). Table 3 compares the catalytic performances of the catalysts used in this work with other works (Imran et al., 2014; Lu et al., 2010; Mysore Prabhakara et al., 2021). Significantly higher content of hydrocarbons is obtained with the use of Ni–Mo/TiO2 and Ni/Al2O3 as compared to the other TiO2 and Al2O3-based catalysts. However, this is also contributed by the addition of PP as the co-feeding material that provides a sufficient hydrogen source. High oxygenated compounds in the oil reported in the other works are expected, mainly from the decomposition of the wood biomass in the presence of TiO2 and Al2O3-based catalysts.Ni–Mo/TiO2 and Ni/Al2O3 are mesoporous acidic catalysts based on nitrogen adsorption-desorption isotherm and NH3-TPD analyses. Between 500 and 700 °C, the pyrolysis oil yields from the catalytic co-pyrolysis of OPT and PP using Ni–Mo/TiO2 and Ni/Al2O3 were 12.67–19.50 wt.% and 12.33–17.17 wt.%, respectively. The acidic properties of both catalysts enhanced the production of hydrocarbon in oil by facilitating the deoxygenation of oxygenated and phenolic compounds and the catalytic cracking of PP. By adding transition metals (Ni and Mo) into the acidic TiO2 and Al2O3-based catalysts, the deoxygenation mechanism was shifted towards decarbonylation and decarboxylation, removing oxygen from oil as carbon dioxide and carbon monoxide gases, which can conserve hydrogen for hydrocarbon formation. Compared to the non-catalytic co-pyrolysis case, the high amount of cyclic hydrocarbons in oil from catalytic co-pyrolysis with Ni–Mo/TiO2 and Ni/Al2O3 catalysts indicates their high catalytic ability in promoting the isomerization and oligomerization reactions of olefins and aliphatic hydrocarbons.Liza Melia Terry: Methodology, Validation, Formal analysis, Investigation, Visualization, Writing – original draft. Melvin Xin Jie Wee: Methodology, Resources. Jiuan Jing Chew: Supervision, Resources, Writing – review & editing. Deni Shidqi Khaerudini: Resources, Writing – review & editing. Nono Darsono: Resources, Writing – review & editing. Aqsha Aqsha: Conceptualization, Resources, Funding acquisition, Writing – review & editing. Agus Saptoro: Resources, Writing – review & editing. Jaka Sunarso: Supervision, Resources, Writing – review & editing, Project administration, Funding acquisition.The authors declare that they have no known competing financial interests or personal relationships that could have appeared to influence the work reported in this paper.Liza Melia Terry gratefully acknowledges the Tun Taib Scholarship from Sarawak Foundation. The authors acknowledge the facilities, scientific, and technical support from Advanced Characterization Laboratories Serpong, National Research and Innovation Agency through E-Layanan Sains, Badan Riset dan Inovasi Nasional. The authors also acknowledge the facilities for GC-MS analysis and funding support from Curtin University Malaysia through Strategic Research Incentives (SRI).The following is the Supplementary data to this article: Multimedia component 1 Multimedia component 1 Supplementary data to this article can be found online at https://doi.org/10.1016/j.envres.2023.115550.

Pyrolysis oil from oil palm biomass can be a sustainable alternative to fossil fuels and the precursor for synthesizing petrochemical products due to its carbon-neutral properties and low sulfur and nitrogen content. This work investigated the effect of applying mesoporous acidic catalysts, Ni–Mo/TiO2 and Ni/Al2O3, in a catalytic co-pyrolysis of oil palm trunk (OPT) and polypropylene (PP) from 500 to 700 °C. The obtained oil yields varied between 12.67 and 19.50 wt.% and 12.33–17.17 wt.% for Ni–Mo/TiO2 and Ni/Al2O3, respectively. The hydrocarbon content in oil significantly increased up to 54.07–58.18% and 37.28–68.77% after adding Ni–Mo/TiO2 and Ni/Al2O3, respectively. The phenolic compounds content was substantially reduced to 8.46–20.16% for Ni–Mo/TiO2 and 2.93–14.56% for Ni/Al2O3. Minor reduction in oxygenated compounds was noticed from catalytic co-pyrolysis, though the parametric effects of temperature and catalyst type remain unclear. The enhanced deoxygenation and cracking of phenolic and oxygenated compounds and the PP decomposition resulted in increased hydrocarbon production in oil during catalytic co-pyrolysis. Catalyst addition also promoted the isomerization and oligomerization reactions, enhancing the formation of cyclic relative to aliphatic hydrocarbon.

Biomass represents a sustainable alternative carbon source compared to fossil resources like oil, gas and coal [1–3]. Considering the limited reserves of these fossil resources, growing research efforts are being devoted to the development of efficient catalytic systems for biomass valorisation into biofuels and biobased chemicals [1,3,4]. For the upgrading of biobased compounds into valuable chemicals, metallic catalysts are often required for one or more step(s) in a multi-step reaction that may involve hydrogenation, oxidation and/or hydrogenolysis [3,4]. On one hand, noble metal catalysts, such as Au, Pt, Pd and Ru nanoparticles, often exhibit excellent catalytic performance in specific reactions [3]; on the other hand, their high cost limits the extension of their application from lab-scale to the industry. Moreover, these catalysts often suffer from stability issues since the nanoparticles tend to aggregate and thus decrease their activity under hydrothermal reaction conditions [4,5]. As such, there is a strong need for developing noble-metal-free catalysts, which ideally should have comparable performance and better stability compared to those noble metal catalysts [3]. Among the biobased compounds that typically require the use of noble metal catalysts for their oxidation, glycerol is an attractive platform molecule [6,7]. It is produced in large amounts (above 1 million tons crude glycerol in 2016) as the major side product from the biodiesel industry by transesterification of vegetable oils with methanol [4,8]. This led to an oversupply of glycerol and, therefore, has prompted both academia and industry to develop efficient catalytic routes to convert it into several valuable chemical products [9–11]. Lactic acid and alkyl lactates can be produced from glycerol through a dehydrogenation-rearrangement pathway (Scheme 1 ) [8,12–14]. Lactic acid has a wide range of applications, including that as monomer of poly-lactic acid, a biodegradable bio-polymer with various applications in the food, pharmaceutical and packaging industry [12]. Currently, lactic acid is produced by fermentation of carbohydrates, which generates large amounts of salts in the product work-up section and has a relatively low volumetric production rate [15,16]. The chemocatalytic route involving the dehydrogenation of glycerol and consecutive rearrangement of the triose intermediates (Scheme 1) is considered a viable, sustainable alternative to the fermentation process [12]. This chemocatalytic route implies a nominal formation of H2 and in this sense can be correlated to the use of glycerol as feedstock for the sustainable production of H2 through aqueous-phase reforming (APR) [7,11,17]. Hydrogen is widely used in current chemical industry (e.g. ammonia synthesis, Fischer-Tropsch process, steel industry and various hydrogenation reactions) and in the power fuel cell systems as a clean power source [2,11,18]. Clearly, routes that allow producing H2 from a renewable source such as biomass represent a sustainable alternative to the current production through methane steam reforming, which is based on a fossil resource and requires extremely harsh conditions [2,19].The conversion of glycerol into lactic acid requires metallic sites for the first step, i.e. the dehydrogenative oxidation, and a base or a combination of Brønsted and Lewis acid sites for the second step (Scheme 1). Most studies used noble metal catalysts for the first step, such as Pt, Pd, Au and their alloys [12,20–22]. Pt/C was used for the hydrogenolysis of glycerol under He atmosphere and gave 55% selectivity to lactic acid at 95% conversion of glycerol [23,24]. Supported Au and its alloy catalysts (AuPt/TiO2) were firstly used with O2 as the oxidant, reaching 30% glycerol conversion and 86% selectivity to lactic acid at 90 °C [21]. The first report of a bifunctional catalyst for the conversion of glycerol into lactic acid without adding a base employed Pt supported on a zeolite (Sn-MFI) and achieved an excellent 81% selectivity towards lactic acid at 90% conversion of glycerol under O2 (6 bar) at a relatively mild temperature (90 °C) [14]. Catalysts based on non-noble transition metals, such as Ni, Co and Cu, were also found to be active in converting glycerol to lactic acid under inert atmosphere in the presence of a base [20,25–29]. A Ni/graphite catalyst tested at 250 °C for 2 h yielded 89% lactic acid at full glycerol conversion [20]. A series of 30%CuO/ZrO2 catalysts were also developed and reached 95% yield of lactic acid at 200 °C [29]. A recent study reported a 20%Co3O4/CeO2 catalyst that achieved 80% selectivity to lactic acid with 85% glycerol conversion at 250 °C for 8 h [27]. All these non-noble metal catalysts were employed in the presence of a homogeneous base (NaOH) and at relatively high reaction temperatures (200–250 °C), under which conditions the base alone would display a significant activity in the conversion of glycerol to lactic acid [30,31]. An additional drawback of the Ni, Cu and Co-based systems is the high metal-to-glycerol ratio that was needed to achieve acceptable reaction rates. Moreover, the Cu and Co-based catalysts suffered remarkable loss of activity upon reuse, probably due to leaching of metal species under the hydrothermal conditions [27,29]. If the conversion of glycerol to lactic acid (salt) is carried out under inert atmosphere, the initial dehydrogenative oxidation step (Scheme 1) nominally liberates one molecule of H2 per molecule of glycerol [14,25]. However, the hydrogen generated in such system is highly diluted by N2 in most cases and is thus difficult to collect. In this context, it is more attractive to utilise in-situ the hydrogen removed from glycerol in the reduction of relevant target compounds. Here, we report a bimetallic Ni-Co catalyst supported on CeO2 with remarkably high activity in the transfer hydrogenation between glycerol and several H2 acceptors, under relatively mild hydrothermal conditions (160 °C) and in the presence of NaOH as promotor. The choice of investigating a Ni-based catalyst was inspired by the above-mentioned activity of this metal in converting glycerol to lactic acid, combined with its well-known activity in catalysing hydrogenation reactions as significantly cheaper alternative to noble metals (e.g. Pt and Pd) [3,32]. The idea of using Ni in a bimetallic system was justified by previous reports that showed that the catalytic performance of Ni could be enhanced by incorporating another component, such as Co or Cu, which led to stronger metal-support interaction with consequent smaller metal particle size [3,4,33]. Namely, bimetallic Ni-based catalysts supported on ZrO2 showed much better performance in the dry reforming of methane (Ni-Co) or in the oxidative steam reforming of methanol (Ni-Cu) compared to their monometallic counterparts [34–36]. In this work, different oxides were tested as support for the Ni-based catalysts, with CeO2 leading to the highest activity in glycerol conversion. Our bimetallic Ni-Co catalytic system was also compared to its monometallic counterparts, showing higher activity and allowing to reach very high conversion of glycerol with excellent selectivity towards lactic acid, and to combine this reaction with the efficient hydrogenation of several unsaturated compounds in a one-pot process.Glycerol (99%), 1,3-dihydroxyacetone dimer (97%), glyceraldehyde (90%), glycolic acid (99%), lactic acid (98%), pyruvic aldehyde (40 wt% in H2O), cyclohexene (99%), cyclohexane (99.5%), sodium hydroxide (98%), benzene (99.9%), levulinic acid (99%), 4-hydroxypentanoic acid, γ-valerolactone (99%), nickel(II) nitrate hexahydrate (98.5%), cobalt(II) nitrate hexahydrate (98%), copper(II) nitrate hemi(pentahydrate) (98%), titanium oxide (P25), magnesium oxide (99%) cerium oxide (nanopowder, nominally < 25 nm, though some large particles were observed by TEM; this compound is denoted as CeO2 for the sake of simplicity, though it contains both CeIV and CeIII and it is thus actually CeO2-x), zirconium oxide (nanopowder, < 100 nm) were purchased from Sigma Aldrich. Glyceric acid (20 wt% in H2O), nitrobenzene (99.5%), aniline (98%), azobenzene (98%), azoxybenzene (98%) were purchased from TCI Chemicals. Active carbon Norit SX1G was purchased from Cabot. The H2O used in this work was always of MilliQ grade. All chemicals were used without further purification.A wet impregnation method was used for the preparation of catalysts based on Ni, Co, Cu, NiCo, NiCu supported on CeO2 and ZrO2. Typically, CeO2 (2 g) was mixed with an aqueous solution of Ni(NO3)2 or Co(NO3)2 or Cu(NO3)2 or the combination of two of them (2 M, with the volume of the solution being defined by the target loading of Ni, Co and Cu). The slurry was stirred at room temperature until the water evaporated. The solid mixture was then dried at 100 °C overnight. The resulting solids were milled to fine powder and then calcined at 550 °C in the oven under static air (heating rate 3 °C/min). The calcined catalysts were further reduced in a tube oven under H2 flow (99.9% and 200 mL/min) at 400 °C (heating rate 3 °C/min) for 2 h. The gas flow was switched to N2 for 1 h to wipe away the adsorbed H2 on the catalyst surface before taking the catalyst out from the tube oven. A typical reduced catalyst prepared by this method was named as 10NiCo/CeO2, in which 10 stands for the total loading of Ni and Co (wt%), in which the mass ratio between Ni and Co is always kept as 1:1. In addition, as a reference, the catalyst was also used directly after calcination at 550 °C without further reduction in H2, which was named as 10NiCo/CeO2-C.The catalytic experiments were carried out in a 100 mL Parr stainless steel autoclave reactor equipped with a Teflon liner and an overhead stirrer. In a typical test, a predetermined amount of the catalyst together with a mixture of aqueous solution of glycerol (0.5 M in 20 mL), NaOH (0.15 mol) and the selected hydrogen acceptor (0.2 mol, as organic phase) were loaded into the reactor. The reaction was performed under N2 (20 bar) for 4.5 h at 160 ᵒC (extra heating time 0.5 h) at a stirring speed of 800 rpm. Next, the reactor was depressurised and the reaction content (in two phases) was taken separately and filtered to remove the catalyst. The organic phase was analysed by gas chromatography using a Thermo Trace GC equipped with a Restek Stabilwax-DA column (30 m × 0.32 mm ×1 μm) and a FID detector. The aqueous phase was first neutralised and diluted by H2SO4 (1 M), then analysed by high performance liquid chromatography (HPLC, Agilent Technologies 1200 series, Bio-Rad AminexHPX-87H 300 × 7.8 mm column) at T = 60 °C, with 0.5 mM H2SO4 as eluent (flow rate: 0.55 mL/min) using a combination of refractive index detector and ultra-violet detector. For the analysis of nitrobenzene and its products, conversion and selectivity were determined by GC analysis using an Agilent Technologies 7980B GC equipped with an Agilent DB-5#6 (5%-Phenyl)-methylpolysiloxane column (15 m, 320 μm ID). The identification of the products was performed by GC-mass spectrometry (GC–MS) on an HP 6890 Series GC equipped with a Restek Rxi-5Si MS fused silica column (30 m, 250 μm ID) coupled to an HP 5973 Mass Selective Detector. Each component was calibrated using solutions of the individual components at 4 different concentrations.For the catalyst recycle tests, a small amount of the reaction mixture was collected for analysis, the remaining mixture was filtered and the catalyst was recovered. The catalyst was washed first with H2O (20 mL), then with ethanol (20 mL), and this procedure was repeated 3 times, after which the solid was dried overnight at 100 ᵒC. This solid was used for another batch experiment.Transmission electron microscopy (TEM), scanning transmission electron microscopy (STEM) and energy-dispersive X-ray spectroscopy (EDX) mapping measurements were performed on a FEI Tecnai T20 electron microscope operating at 200 keV with an Oxford Xmax 80 T detector. The samples were prepared by ultra-sonication in ethanol followed by drop-casting of the material on a copper grid.Nitrogen physisorption isotherms were measured at −196 °C using a Micromeritics ASAP 2420 apparatus. The Brunauer-Emmet-Teller (BET) method was used to calculate the specific surface area. The Barrett-Joyner-Halenda (BJH) method was used to calculate the pore volume.Inductively-coupled plasma optical emission spectrometry (ICP-OES) was performed using a Perkin Elmer Optima 7000 DV instrument in order to obtain the actual metal loadings on the supports.X-ray photoelectron spectroscopy (XPS) was measured by mounting the catalysts on a conductive tape adhered to the XPS sample holder. No further treatment was carried out prior to the XPS measurement. Then, the sample was loaded into the load lock and the pressure was reduced below 1·10−7 mbar. The XPS measurements were performed using a Surface Science SSX-100 ESCA instrument equipped with a monochromatic Al Kα X-ray source (hν =1486.6 eV). During the measurement, the pressure was kept below 2·10-9 mbar in the analysis chamber. For acquiring the data, a spot size with a 600 μm diameter was used. The neutraliser was on to avoid charging effects. All XPS spectra were analysed using the Winspec software package developed by LISE laboratory, University of Namur, Belgium, including Shirley background subtraction and peak deconvolution.Hydrogen-temperature programmed reduction (H2-TPR) measurements were performed on an Autochem II 2920 from Micromeritics. In a typical experiment, 80 mg of sample was pre-treated at 500 °C (heating rate 10 °C/min) for 1 h in a flow of He (30 mL/min). Subsequently, the sample was cooled down to 50 °C under the same flow of He. The reduction analysis was performed from 50 to 900 °C (10 °C/min) in a 30 mL/min flow of 5 vol.% H2 in He.X-ray diffraction (XRD) measurements were performed on a D8 Advance Bruker diffractometer with a CuKα 1 radiation (λ = 1.5418 Å). The XRD patterns were collected under 40 kV and 40 mA in the range of 10-80°.Definitions:The glycerol conversion (Conv./%) is defined by Eq. (1): (1) Conv . = C g , 0 - C ( g ) C ( g , 0 ) × 100 % in which C(g) is the molar concentration of glycerol after a certain reaction time and C(g,0) is the initial concentration of glycerol.Product selectivity for a compound P is defined by Eq. (2): (2) S p   = C ( p ) C ( g , 0 ) - C ( g ) × 100 % in which C(p) is the molar concentration of a product after a certain reaction time.The yield of transfer hydrogenation is defined by Eq. (3): (3) Y trans - H  = ∑ ( x * n p 1 ) n ( g , 0 ) - n ( g ) × 100 % in which x is the number of hydrogen atoms needed for the reduction of product 1, n(p1) is the molar amount of product 1, n(g) is the molar amount of glycerol after a certain reaction time and n(g,0) is the initial molar amount of glycerol.The term “lactic acid” is used in this article to describe the product obtained from the reaction mixture, which is actually sodium lactate (mixed with a small portion of lactic acid from hydrolysis).Our study of the conversion of glycerol into lactic acid coupled with the transfer hydrogenation to an unsaturated compound started with the investigation of the catalytic behaviour of Ni catalysts (10 wt%) as a function of the type of the material (activated carbon (AC) and various metal oxides) on which the metal particles were supported by wet impregnation. The five catalysts were tested at 160 °C in the presence of NaOH as promotor and using a model compound as cyclohexene as the H2 acceptor (Table 1 ). Ni supported on AC, MgO and TiO2 showed relatively low activity (entries 1–3, Table 1), whereas the activity was significantly higher when nanosized CeO2 and nanosized ZrO2 were used as support for Ni (glycerol conversion 53% and 63%, respectively; entry 4–5, Table 1), in line with previous reports on other (de)hydrogenation reactions [4,33]. In all cases, high selectivity towards lactic acid (> 91%) was observed. This is attributed to the presence of NaOH, which effectively promotes the deprotonation of glycerol and catalyses the successive isomerisation of the intermediates (glyceraldehyde and dihydroxyacetone) into the lactic acid salt (Scheme 1), thus granting very high selectivity towards the desired product [21,22,28,37]. Small amounts of glyceric acid, glycolic acid and propanediol were detected as side products (Table 1). Glyceric acid is formed through the further dehydrogenation of glyceraldehyde and glycolic acid probably originates from oxidative CC bond cleavage of glyceric acid [13]. Propanediol (as a mixture of 1,2- and 1,3-isomers) probably forms via the hydrogenolysis of glycerol [38–40]. In addition, for all reactions, very minor amounts of glyceraldehyde and propanoic acid were observed as side products, with selectivity below 0.2% for each of them.Based on this preliminary study, CeO2 and ZrO2 were selected as supports for further study of Ni-based catalysts. Then, we aimed at improving the catalytic performance by incorporating another metallic component, i.e. Co or Cu [3,4,33]. The activity of the bimetallic catalysts was compared to the monometallic counterparts (Table 2 ), while keeping the total loading of metal at 10 wt% (and with 1:1 mass ratio for the bimetallic systems). The incorporation of Co into the catalyst formulation was highly beneficial when CeO2 was used as support (10NiCo/CeO2), leading to 91% glycerol conversion (entry 1, Table 2) compared to 53% conversion obtained over 10Ni/CeO2 and 46% conversion over 10Co/CeO2 (entry 5, Table 2). Also the incorporation of Cu enhanced the activity compared to the monometallic counterparts, though the effect was less marked (compare entry 2 in Table 2 to entry 5 in Table 1 and entry 6 in Table 2). On the other hand, the 10NiCo/ZrO2 catalyst showed almost the same activity as the monometallic 10Ni/ZrO2, (compare entry 3 in Table 2 with entry 4 in Table 1), whereas the incorporation of Cu proved more beneficial when ZrO2 was the support, reaching 80% glycerol conversion (entry 4, Table 2). These results indicate a complex interplay between the type of metals and the supports. The benefit brought about by the bimetallic formulation will be elucidated further in the case of the optimum catalyst, i.e. 10NiCo/CeO2 (vide infra). In all these tests, the selectivity towards lactic acid remained very high (94–96%). Glyceric acid, glycolic acid and propanediol were detected as the main side products, with selectivity < 6% in total. Though the incorporation of Cu enhanced the activity of the Ni-based catalysts, leaching of metal species was observed in the basic medium under hydrothermal conditions, with significant amount of brown Cu-containing precipitate deposition on the stirring bar and reactor walls [28,33,41–43]. Therefore, 10NiCo/CeO2 was selected for further investigation aimed both at a deeper evaluation of the catalytic performance and at understanding the relationship between structure and catalytic behaviour.The catalysts presented in this work were prepared by wet impregnation, followed by calcination and finally reduction by H2. The actual loading of Ni and/or Co determined by ICP-OES (Table 3 ) was found to be very close to the nominal 10 wt% loading. In the bimetallic Ni-Co catalyst, the actual loading of Ni and Co is 5.6 wt% for both metals, which is slightly higher than the theoretical 5 wt%. The BET surface area was measured before and after loading Ni and Co, showing only a slight decrease (from 32 to 28 m2/g) compared to the fresh CeO2 support.To investigate the possible organisation of Ni, Co and Ni-Co species in crystalline phases on the CeO2 support, the catalysts were further characterised by XRD before and after reduction (Fig. 1 ). The materials before reduction (Fig. 1A) display the characteristic peaks of the CeO2 support together with the typical peaks of NiO (in 10Ni/CeO2-C) or Co3O4 (in 10Co/CeO2-C) [34,35,44]. The bimetallic 10NiCo/CeO2-C shows a broad peak at 36.7°, which is slightly shifted compared to the Co3O4 peak (37°) and has been attributed to the mixed oxide NiCo2O4 [34,45–47]. After reduction at 400 °C in H2 flow, besides the peaks of the CeO2 support, only one peak at 44.7° belonging to metallic Ni can be seen in the pattern of 10Ni/CeO2 (Fig. 1B). On the other hand, no signals stemming from Co and/or Ni phases were observed in 10Co/CeO2 and 10NiCo/CeO2. These results suggest that relatively large crystalline Ni particles formed upon reduction in 10Ni/CeO2, while the Co or Ni-Co species obtained after reduction were highly dispersed in the other two catalysts [45–48].To achieve deeper insight on the dispersion of Ni, Co and bimetallic Ni-Co catalysts supported on nanosized CeO2, TEM and STEM-EDX-mapping were used to investigate the average size of these metallic domains (Figure S1, 2 and 3). Since the atomic mass of cerium is much higher than that of nickel or cobalt, it is hard to determine the particle size of Ni, Co or Ni-Co alloy on the CeO2 support based on TEM pictures (Figure S1), as the darker zones are not necessarily corresponding to Ni or Co domains.Analysis by STEM coupled with EDX mapping was more informative as it allows identifying the elemental composition within the image (Fig. 2 ). The large green domains in Fig. 2A and 3B indicate the presence of Ni-containing nanoparticles on CeO2. Based on the XRD data (Fig. 1A), these domains are identified as large NiO nanoparticles (mainly around 100 nm, with some smaller particles, see Fig. 2A) in the sample before reduction (10Ni/CeO2-C), and to large domains of metallic Ni (around 75 nm, Fig. 2B) after the sample was reduced (10Ni/CeO2). For the monometallic material prepared by supporting Co on CeO2 and prior to reduction (10Co/CeO2-C), the Co3O4 identified by XRD (Fig. 1A) was found to be better dispersed on the CeO2 support (Fig. 2C) compared to NiO on CeO2. The 10Co/CeO2 material obtained upon reduction showed nearly homogeneously dispersed Co species (Fig. 2D), which indicates that the particle size of Co is lower than the detection limit of EDX-mapping (around 30 nm). The relatively small size of the Co nanoparticles is also in agreement with the absence of any signal due to metallic Co in the XRD pattern of 10Co/CeO2 (Fig. 1B), which suggests a strong metal-support interaction between Co and CeO2 [4,33,35,45,46,48].STEM and EDX-mapping of the Ni-Co bimetallic material prior to reduction (10NiCo/CeO2-C), showed that both Ni and Co are nearly homogeneously dispersed on the CeO2 surface (Fig. 4 A–D). This demonstrates that the presence of Co prevents the aggregation of Ni species, in contrast to the large domains observed in 10Ni/CeO2-C. After reduction at 400 °C under H2, Ni and Co still preserve very good dispersion, with no large metal particles (i.e. > 30 nm) being visible (Fig. 3 H). The strong interaction between Co and the CeO2 support, which promotes the observed high dispersion of both Co and Ni on the surface, has been shown to be related to the formation of a thin layer of reduced CeOx at the interface with the metallic Co [35]. Based on our results, we infer that this feature prevents Ni from forming large particles in the process of calcination and reduction [33,35,46].The reducibility of Ni, Co and Ni-Co supported on CeO2 was further investigated by H2-TPR (Fig. 4). The support, CeO2, exhibited two dominant peaks centred at 490 °C (from 300 to 550 °C) and 880 °C (from 700 to above 900 °C), which are attributed to the reduction of surface ceria and bulk ceria, respectively [35,49]. Besides the reduction peaks of CeO2 at 420 and 815 °C, which are slightly shifted to lower temperature, the monometallic 10Ni/CeO2-C displays two peaks at 213 °C (minor) and 320 °C (dominant), which are attributed to the reduction of adsorbed oxygen and NiO, respectively [35,50]. The monometallic 10Co/CeO2-C showed two main peaks at 260 and 315 °C, which are attributed to the two-step reduction Co3O4→CoO→Co [51,52]. The large and broad shoulder extending from 350 to 500 °C is probably due the reduction of surface CeO2. Compared to the monometallic Ni catalyst, the significant increase of the intensity of the reduction peak of surface CeO2 in the monometallic Co catalyst supports the existence of a strong metal-support interaction between Co and CeO2, which is in agreement with the formation of a thin layer of reduced support on the metallic Co surface reported in the literature [35,48]. The 10NiCo/CeO2-C material showed almost identical profile as the one of 10Co/CeO2-C, with all the peaks shifted by ca. 5 °C to lower temperature. This suggests that, in the bimetallic Ni-Co catalyst, the reduction behaviour is mainly dictated by the presence of Co, including the strong metal-support interaction indicated by the broad shoulder between 350 and 500 °C. This result explains the observed much better dispersion of the metal species in the bimetallic Ni-Co catalyst compared to the monometallic Ni catalyst (Figs. 2 and 4) [35].The characterisation by EDX-mapping and H2-TPR indicates a geometrical effect of the presence of Co on the dispersion of Ni on the CeO2 support. To investigate further the interaction between Co, Ni and the support, selected catalysts were analysed by XPS (Figure S2-4). The XPS signal of the Ni 2p3/2 core level region of the unreduced 5Ni/CeO2-C catalyst was deconvoluted into 3 main peaks: at 853.6 eV, assigned to NiO; at 855.6 eV, attributed to Ni(OH)2 and/or NiO(OH); and a satellite peak at 860.6 eV [53–56]. Similar peaks were identified by deconvoluting the Ni 2p3/2 signal of the unreduced 10NiCo/CeO2-C catalyst (Figure S2.A and B). After reduction (Figure S2.C and D), in addition to the 3 peaks mentioned above, the deconvolution allowed identifying a peak ascribed to Ni0 (at 852.3 eV) in catalysts 5Ni/CeO2 and 10 NiCo/CeO2 [54,55,57]. These data confirm the successful reduction to metallic Ni. The fact that the majority of the XPS signal stems from oxidised Ni species can be explained considering that XPS is a surface technique (information from the top 1–10 nm of the material) and that the surface of the particles is expected to tend to oxidise in contact with air and moisture [58,59]. The XPS signal of the Co 2p3/2 core level region of the unreduced 5Co/CeO2-C catalysts was deconvoluted into 3 main peaks: at 779.5 eV, assigned to cobalt oxides (CoO and/or Co3O4); at 781.5 eV, ascribed to Co(OH)2; and a satellite peak at 785.5 eV [55,60,61]. Analogous peaks were identified by deconvoluting the Co 2p3/2 signal of the unreduced 10NiCo/CeO2-C catalyst (Figure S3.A and B). After reduction (Figure S3.C and D), in addition to the 3 peaks mentioned above, the deconvolution showed a peak assigned to Co0 (at 778.0 eV) in the catalysts 5Co/CeO2 and 10 NiCo/CeO2 [55,62]. Similarly to what discussed in the case of the supported Ni particles, the presence of oxidised Co species in the reduced samples is attributed to the formation of a layer of oxides and hydroxides at the surface of the particles, generated by contact with air and moisture. The features of the XPS signal of the Ce 3d core level region support the anticipated strong interaction between Co and CeO2 (Figure S4). This is indicated by the surface reduction of Ce4+ and the increase in Ce3+ observed in the XPS spectra of the Co-containing catalysts (whereas this effect is absent in the spectra of the catalysts containing Ni but no Co). This matches well with the literature and with our H2-TPR results [35,52,54]. The XPS data are not conclusive on possible synergistic electronic effects between Ni and Co. Therefore, we infer that the main reason for the improved catalytic performance of bimetallic 10NiCo/CeO2 is the smaller size and better dispersion of the Ni-containing particles.Based on this characterisation study, the optimum activity observed with the bimetallic 10NiCo/CeO2 catalyst is attributed to presence of the more active Ni compared to the monometallic 10Co/CeO2, and to the better dispersion of the active metallic species compared to the monometallic 10Ni/CeO2 catalyst. To further confirm the nature of the active sites, unreduced Ni, Co and bimetallic Ni-Co catalysts were tested under the same conditions employed for the reduced catalysts (Table S1). In the unreduced materials, the metal oxides (NiO, Co3O4 and NiCo2O4) would be the catalytic sites rather than the metallic sites. All the unreduced catalysts had significantly lower activity compared to the reduced ones (Table 1 and 2), with the conversion of glycerol being < 16% in all cases. These results confirm that the metallic sites are the active site in this transfer hydrogenation reaction between glycerol and cyclohexene, in agreement with what shown in the literature [27–29].The Ni, Co and Ni-Co catalysts with different loading (2, 5 and 10 wt%) supported on CeO2 were tested to gain better understanding on the effect of the Ni and Co composition (Fig. 5 ). With the Ni/CeO2 catalysts, the conversion of glycerol increased with the metal loading up to 5 wt% Ni, at which it reached 55%, whereas it remained nearly constant upon further increase to 10 wt % of Ni. This trend is completely different from the one observed with the Co/CeO2 and NiCo/CeO2 catalysts, for which the glycerol conversion and the lactic acid yield exhibited an increasing trend with the increase in metal loading (Fig. 5A). The performance of these catalysts can be analysed also in terms of turnover number (TON) (Fig. 5B). These data show that the TON is nearly constant as a function of metal loading for the monometallic Co-catalysts, whereas an increasing loading of Ni causes a gradual decrease in TON, which is more marked for the monometallic Ni-catalysts compared to the bimetallic Ni-Co materials. These trends are in agreement with the tendency of Ni to form large particles at high loading (see Fig. 3.A–B), which implies that a smaller fraction of the metal is available to act as active site, thus leading to the observed lower TON. On the other hand, Co maintains small metallic domains on the CeO2 surface also at 10 wt% metal loading (Fig. 2.C–D), thus enabling to have a nearly constant TON as a function of metal loading. The highest TON was observed for 2Ni/CeO2 and 2NiCo/CeO2, whereas among the catalysts with 10 wt% metal loading, the highest TON was found for 10NiCo/CeO2, despite the decrease compared to the 2 wt% material. This confirms the higher intrinsic activity of Ni compared to Co in catalysing the dehydrogenative oxidation of glycerol. Non-noble metal catalysts are generally used with high loading to give high productivity. Indeed, when the catalytic performance is compared in terms of productivity (Fig. 5C) the highest value among the tested catalysts is obtained with the material with the highest TON among those with 10 wt% metal loading, i.e. 10NiCo/CeO2. This underlines the benefit of the presence of Co in combination with Ni on the catalytic performance [33–35,44,48].The 5NiCo/CeO2 catalyst, which achieved intermediate glycerol conversion at 160 °C, was selected for investigating the effect of the reaction temperature (in the range 140 to 200 °C, Figure S2). The conversion of glycerol increased with higher reaction temperature, from 11% (at 140 °C) to 99% (at 200 °C), while the selectivity to lactic acid remained > 98%. The selectivity towards the transfer hydrogenation was steady at around 25% in all range of temperatures. It should be noted that, when only NaOH was used in the reaction system, the conversion of glycerol was rather low, though it increased from 1.6 to 16% (from 140 to 200 °C, Figure S5). This demonstrates the need for a heterogeneous catalyst to carry out the dehydrogenation reaction in this range of relatively mild temperatures [30,31].To further investigate the effects of the catalyst amount on this reaction, different weights of catalyst (from 0.025 to 0.15 g) were used, while all other parameters were kept constant. The results show a gradual increase in the conversion of glycerol from 29% to > 99.9% upon increase of the loading of the 10NiCo/CeO2 catalyst (Figure S6).The role of NaOH was studied in more detail by varying the molar ratio between NaOH and glycerol (from 0 to 2, Figure S7). Without the addition of NaOH, both the conversion of glycerol and the selectivity to lactic acid were very low (conversion of glycerol = 3.5%). If the molar ratio between NaOH and glycerol was increased, the conversion of glycerol gradually increased reaching 91% with 85% yield of lactic acid salt at NaOH/glycerol = 1.5. However, a further increase in the NaOH/glycerol molar ratio to 2 caused a decrease in the conversion of glycerol to 81%, thus indicating that the employed ratio (1.5) is the optimum value. These results confirm that the presence of a base like NaOH in the reaction mixture is critical to induce the deprotonation of one of the hydroxyl groups of glycerol, thus promoting the dehydrogenation of glycerol [21,28]. Moreover, NaOH can catalyse the isomerisation of glyceraldehyde and dihydroxyacetone and lead to the formation of sodium lactate with very high selectivity.The reaction profile as a function of the reaction time was studied with the 10NiCo/CeO2 catalyst (Figure S8). The conversion of glycerol increased almost linearly within the first 4.5 h, corresponding to a productivity of lactic acid of 17.4 g(LA)/(g(metal)h). After 6.5 h of reaction, almost complete glycerol conversion (97%) was achieved, with 93% lactic acid (salt) yield. The selectivity towards lactic acid stayed always above 90% and the total selectivity towards by-products (glyceric acid, glycolic acid and propanediol) was around 4%. The selectivity towards the transfer hydrogenation slightly decreased with the reaction time, from 31% to 26%. These results suggest that under the employed reaction conditions the dehydrogenation of glycerol is the rate-determining step, and that once the dihydroxyacetone and/or glyceraldehyde formed, they would be transformed into lactic acid (salt) in a very fast and selective way.Catalyst 10NiCo/CeO2 was also selected for a reusability test (Fig. 6 ). The fresh catalyst showed 91% conversion of glycerol and 85% yield to lactic acid, while recycling after straightforward washing and drying led to a slight, gradual decrease in activity. After 5 runs, the conversion of glycerol decreased to 73%, while the selectivity towards lactic acid remained unaltered (> 94%). Meanwhile, the selectivity in the transfer hydrogenation gradually increased from 24 to 28% between the first and the fifth run. The gradual loss of activity is probably caused by the leaching of a small fraction of the active components in the alkaline hydrothermal reaction system, since the loading of Ni and Co decreased from 5.6 wt% (each) in the fresh catalyst to 4.4 wt% (each) after 5 runs (entry 5, Table 3).During the optimisation of the Ni-based catalyst presented above, cyclohexene was employed as hydrogen acceptor in the transfer hydrogenation reaction from glycerol. To expand the scope of applicability of the transfer hydrogenation, we tested a set of H2 acceptors with different features (a biobased compound as levulinic acid, an aromatic compound as benzene, a compound containing both an aromatic ring and another reducible group as nitrobenzene and a linear, terminal alkene as 1-decene). While cyclohexene and 1-decene were selected as model compounds, the hydrogenation of benzene, nitrobenzene and levulinic acid is of potential industrial relevance [63–69]. The tests were carried out with a 1:1 molar ratio between glycerol and the hydrogen acceptor, at 160 °C under N2 atmosphere (Scheme 2 and Table 4 ).When levulinic acid was employed as the H2 acceptor, two main products were observed: 4-hydroxypentanoic acid (27% yield), obtained by hydrogenation of the carbonyl group of levulinic acid, and γ-valerolactone (48% yield), obtained by subsequent dehydration (Scheme 2 and entry 1 in Table 4). γ-Valerolactone can be used as food additive, solvent and precursor for polymers [6,68,70,71]. This reaction also gave an 86% yield of lactic acid at 87% glycerol conversion with a very good 88% selectivity in the transfer hydrogenation.When benzene was tested as H2 acceptor, a very high selectivity (97%) in the transfer hydrogenation from glycerol was observed, with cyclohexane being the only product (corresponding to complete reduction of benzene). The reduction of benzene is the industrial route for the production of cyclohexane, which is employed as precursor in the synthesis of adipic acid used in the manufacturing of nylon [72,73]. The yield achieved here (25%) is promising considering that under the employed reaction conditions (1:1 molar ratio between glycerol and benzene), the maximum theoretical yield of cyclohexane is 33%. These results were coupled with 79% conversion of glycerol and 77% yield of lactic acid (entry 2, Table 4).When nitrobenzene was employed as hydrogen acceptor, the reduction of the nitro group is expected to be favoured over the reaction of the aromatic ring. Indeed, the observed products (azoxybenzene with 59% yield, azobenzene with 18% yield and aniline with 7.5% yield) all originate from the reduction of the nitro group (Scheme 2) [63,74–76]. These are all industrially valuable products, with azoxybenzene being utilised in dyes, reducing agents and polymerisation inhibitors; azobenzene being used in dyes, indicators and as additive in polymers; and aniline finding application in producing pesticides, dyes and as the precursor to polyurethane [77–79]. For this reaction, the selectivity in the transfer hydrogenation from glycerol was > 100%. This can be explained considering the strong oxidative ability of nitrobenzene, which led to the further oxidation of the triose intermediates to glyceric acid and glycolic acid (entry 3, Table 4), similarly to what is generally observed in the oxidation of glycerol in the presence of O2 [25,80–82]. Therefore, glyceric acid (52% yield) becomes the major product under these conditions, with lactic acid being obtained in much lower yield (23%).When 1-decene was selected as a linear H2 acceptor with a primary CC bond, 92% conversion of glycerol and 91% yield of lactic acid was achieved after reaction, while 85% of decene was hydrogenated to decane, corresponding to a remarkably high 94% selectivity in the transfer hydrogenation (entry 4, Table 4). This is much higher than what was found when using cyclohexene as the H2 acceptor (entry 5, Table 4). This result is probably due to the higher accessibility of the CC bond in a linear alkene with a terminal double bond as 1-decene compared to the more sterically-hindered cyclohexene.The study of substrate scope for the transfer hydrogenation reaction from glycerol demonstrated that our catalytic system based on 10NiCo/CeO2 is able to efficiently promote the conversion glycerol to lactic acid while exploiting the liberated hydrogen in the reduction of different unsaturated compounds to achieve the synthesis of useful target products without requiring an external H2 source.Bimetallic Ni-Co catalysts supported on CeO2 were prepared and tested for the transfer hydrogenation from glycerol to various unsaturated compounds, in which lactic acid and the corresponding hydrogenated products were obtained in a one-pot batch reaction. Introducing Co into the formulation of the Ni-based catalysts was crucial to prevent the aggregation of Ni into large particles. This was proven by the higher activity of the bimetallic 10NiCo/CeO2 catalyst compared the Ni- or Co-based counterparts, and by characterisation of the catalytic materials by EDX-mapping and H2-TPR, which demonstrated the high dispersion of Ni-Co sites on the CeO2 support. The bimetallic 10NiCo/CeO2 catalyst exhibited very high activity (91% glycerol conversion) and selectivity to lactic acid (94%) at 160 °C, 4.5 h under N2 atmosphere in the presence of NaOH as promoter. This result demonstrates that excellent conversion and selectivity can be achieved using a catalyst with a relatively low loading of Ni and Co and that operates at milder reaction conditions compared to other non-noble metal catalysts for glycerol dehydrogenation reactions [20,25–29]. Moreover, various H2 acceptors (levulinic acid, benzene, nitrobenzene, 1-decene, cyclohexene) were tested in the transfer hydrogenation from glycerol, exploiting in-situ the hydrogen liberated in the dehydrogenative oxidation of glycerol to generate several useful products.The authors declare that they have no known competing financial interests or personal relationships that could have appeared to influence the work reported in this paper.We would like to thank the financial support from the China Scholarship Council for the Ph.D. grant of Zhenchen Tang, the technical support from Leon Rohrbach, Jan Henk Marsman, Erwin Wilbers, Anne Appeldoorn and Marcel de Vries, the TEM-EDX support from Dr. Marc Stuart and the ICP-OES support from Johannes van der Velde. We also acknowledge Dr. Matteo Miola for useful scientific discussion of the XPS data.Supplementary material related to this article can be found, in the online version, at doi:https://doi.org/10.1016/j.apcatb.2019.118273.The following is Supplementary data to this article:

Bimetallic Ni-Co catalysts supported on nanosized CeO2 were prepared and investigated as heterogeneous catalysts for the transfer hydrogenation between glycerol and various H2 acceptors (levulinic acid, benzene, nitrobenzene, 1-decene, cyclohexene) to selectively produce lactic acid (salt) and the target hydrogenated compound. The bimetallic NiCo/CeO2 catalyst showed much higher activity than the monometallic Ni or Co counterparts (with equal total metal mass), thus indicating strong synergetic effects. The interaction between the metallic sites and the CeO2 support was thoroughly characterised by means of transmission electron microscopy (TEM), scanning transmission electron microscopy (STEM), energy-dispersive X-ray spectroscopy (EDX) mapping, X-ray photoelectron spectroscopy (XPS), hydrogen-temperature programmed reduction (H2-TPR) and X-ray diffraction (XRD). Combining characterisation and catalytic results proved that the Ni species are intrinsically more active than Co species, but that incorporating Co into the catalyst formulation prevented the formation of large Ni particles and led to highly dispersed metal nanoparticles on CeO2, thus leading to the observed enhanced activity for the bimetallic system. The highest yield of lactic acid (salt) achieved in this work was 93% at 97% glycerol conversion (160 °C, 6.5 h at 20 bar N2, NaOH: glycerol = 1.5). The NiCo/CeO2 catalyst also exhibited high activity and selectivity towards the target hydrogenated products in the transfer hydrogenation reactions between glycerol and various H2 acceptors. Batch recycle experiments showed good reusability, with retention of 80% of the original activity after 5 runs.

No data was used for the research described in the article.The efficient emission control of unburned methane in power plants and vehicles that use natural gas as a potential bridge fuel in the transition toward renewable energy is of vital importance [1,2], given that this pollutant is strongly involved in the greenhouse effect. Its potent greenhouse effect is around 25 times higher than that of CO2. Presently, the most adequate approach to minimize the negative impact of the release of residual methane to the atmosphere is catalytic oxidation, which allows the direct conversion of the hydrocarbon to carbon dioxide and water. Probably there is no doubt that noble metal-based catalysts, particularly Pd catalysts, are the systems with the highest intrinsic oxidation power for this abatement strategy [3–7]. However, although large efforts are continuously being made to increase its thermal and chemical stability under operating conditions, its wide use is fundamentally penalized by economic reasons [8]. Thus, the proposal of cheaper, highly efficient, alternative catalysts is a challenge of relevant interest. Most studies have been focused on the use of transition metal oxides, namely nickel [9,10], manganese [11,12], copper [13] or iron [14]. However, it is widely accepted that spinel cobalt oxide (Co3O4) is the most attractive oxide phase for the lean methane catalytic combustion owing to the presence of variable valance states (Co3+/Co2+), its lower bonding energy of Co-O bonds and the high mobility of active oxygen species capable of activating the C-H bond [15–17]. Nevertheless, bulk cobalt oxide, as well as other bulk transition oxides, usually exhibit very poor textural and structural properties, especially when synthesized by simple methodology routes [18,19]. Thus, their good behavior is mainly assigned to their high metallic content (>70 %wt.), thereby resulting in a markedly low intrinsic activity. For this reason, several strategies have been proposed in order to enhance the performance of Co3O4-based catalysts with the ultimate goal of maximizing the population of active sites. The selection of the support for depositing the active phase is the first obvious approach to take into consideration. Furthermore, advances in the optimized design of supported catalysts are highly relevant since the final configuration of a commercial catalytic unit will be a structured catalyst operating with large gas flows [20]. These catalysts will be surely prepared by washcoating a thin catalytic layer (metal oxide/support) onto the surface a monolithic/foam substrate.In this sense, it must be stated that, owing to the high affinity of cobalt for most typical inorganic supports (γ-Al2O3, SiO2 or MgO), a certain fixation of cobalt as less active CoAl2O4, CoSiO3 or Co-Mg mixed oxides must be assumed [21,22]. This unavoidably involves the use of relatively high Co loadings (20–40 %wt.) to compensate partially the useless presence of a fraction of deposited Co. The use of alternative supports such ceria or alpha-alumina prevented this strong undesired interaction but their relatively low intrinsic surface area do not usually lead to a substantially improvement in behavior of the resultant composite catalysts [23,24]. A complementary option to adjust the amount and/or reactivity of oxygen species is the addition of a promoter that could improve the reducibility of the resultant catalyst at low temperatures. Based on its comparable ionic radius and coordination and oxidation states to cobalt, nickel is the most preferred promoter. The incorporation of nickel is mainly justified by the notable activity shown by the NiCo2O4 spinel that can be formed from the interaction between cobalt and nickel. This approach is quite interesting for methane oxidative abatement [25–28], but requires a large amount of nickel (around 50 % of the Co content for a Ni/Co molar ratio of 0.5). In addition, the synthesis of stoichiometric nickel cobaltite is largely dependent on very well controlled synthesis conditions in terms of calcination temperature and selected preparation route, usually oriented to the synthesis of the mixed oxide in its bulk form. In other words, it would be of interest to explore alternatives for taking advantage of the known beneficial effects of nickel promoter, without the need of large amounts of this additive and using a relatively simple route for obtaining an active Ni-promoted cobalt catalyst.Therefore, the objective of this work is the study of Ni/Co-Al2O3 catalysts for the oxidation of methane under conditions similar to those found in the exhaust of vehicular natural gas engines (relatively low residence times, and presence of water and carbon dioxide). Thus, for a total metal loading of 30 % by weight, the effect of the addition of 5 % and 10 %wt.Ni on cobalt catalysts with a content of 25 % and 20 %wt.Co, respectively, was investigated. These samples were prepared by sequential precipitation of cobalt and nickel, with an intermediate calcination step. Along with these catalysts, monometallic cobalt (20 %, 25 % and 30 %wt.) and nickel (30 %wt.) catalysts with a content of 30 %wt. (30Co and 30Ni, respectively) were synthesized as well for comparative purposes.All oxide catalysts were synthesized following a precipitation route over a thermally-stabilized (calcined at 850 C for 8 h) γ-Al2O3 (Saint Gobain), which selected as the support. Three cobalt oxide catalysts, namely 20Co, 25Co and 30Co samples, were prepared by precipitation of aqueous solution of cobalt nitrate hexahydrate with an adjusted concentration to obtain the desired nominal Co loading (20, 25 and 30 wt. %, respectively), at 80 °C using an aqueous solution of sodium carbonate (1.2 M) until reaching a pH of 8.5. After precipitation, the precursors were dried at 110 °C overnight.Then, the catalyst precursors were calcined at 600 C for 8 h in static air. In the case of the reference nickel catalyst (30Ni, with a nominal Ni content of 30 %wt.), the starting salt was nickel nitrate hexahydrate. This sample was also submitted to the same aforementioned thermal treatment.Two bimetallic Ni-Co catalysts were obtained by sequential precipitation using the same metallic salts and precipitating conditions (pH = 8.5, 80 ºC). Thus, nickel was added to the previously prepared 20Co and 25Co samples with a nominal content of 10 % and 5 %wt., respectively. Thus, the total metallic loading of these samples was fixed at 30 %wt. Finally, the samples were again calcined at 600 C for additional 4 h. In this way, both monometallic and bimetallic catalysts were activated under identical thermal conditions. The resulting Ni-Co catalysts were designated as 10Ni/20Co and 5Ni/25Co.The supported catalysts were characterized by a wide number of analytical techniques, including scanning electron microscopy (SEM) coupled to energy dispersive X-ray spectroscopy (EDX), scanning transmission electron microscopy - high angle annular dark field (STEM-HAADF) coupled to EDX mapping, wavelength dispersive X-ray fluorescence (WDXRF), N2 physisorption, X-Ray diffraction (XRD), Raman spectroscopy, X-Ray photoelectron spectroscopy (XPS), temperature-programmed reduction with hydrogen (H2-TPR) and temperature programmed reaction with methane (CH4-TPRe). Experimental details on each of these techniques are included in the Supplementary Material.The activity of the synthesized catalysts for the oxidation of residual methane was determined in a fixed bed reactor (Microactivity by PID Eng&Tech S.L.) between 200 and 600 °C. The reaction products were quantified with an on-line gas chromatograph (Agilent Technologies 7890 N) equipped with a thermal conductivity detector. In each reaction test one gram of catalyst (particle size 0.25–0.30 mm) diluted with 1 g of inert quartz (particle size 0.5–0.8 mm) was used. A reaction mixture of composition 1 %CH4/10 %O2/89 %N2 was used with a total flow rate of 500 mL min−1, which represents an approximate space velocity of 60,000 h−1. To ensure that the mass and heat transfer effects were not affecting the kinetic results, the inter- and intraphase concentration and temperature gradients (Table S1, Supplementary Material) were verify to be negligible according to the criteria proposed by Eurokin [29]. The absence of mass and heat transfer limitations within the reactor was evaluated not only under differential conditions (X < 20 %) but also under the least favorable conditions (450–600 ºC). Additionally, the stability of the most promising catalyst with time on stream was evaluated at constant temperature (575 °C) for a total reaction interval of 150 h under alternate dry, humid (10 %) or CO2-rich (10 %) conditions while maintaining the O2/CH4 molar ratio at 10.Prior to the discussion of the characterization results of the prepared catalysts, it is highly relevant to remark that the variety of oxide phases that can be present in Co- and/or Ni-containing gamma-alumina supported catalysts thermally activated at moderate temperatures (600 °C) is wide. In addition to the expected Co3O4 and NiO oxides, and obviously the γ-Al2O3 support, the presence of mixed spinels such as CoAl2O4 and NiAl2O4 is normally unavoidable. These new metallic oxides are formed due to the strong interaction between the Co and Ni species and the support that results in the partial insertion of Co or Ni atoms into the lattice of the gamma alumina. Moreover, it is commonly accepted that the morphology of these spinels will be essentially amorphous since its transformation into a crystalline structure requires calcination temperatures as high as 800–850 °C [30,31]. Besides, the formation of Ni/Co mixed oxides can occur. Thus, based on these considerations both monometallic (20Co, 25Co, 30Co and 30Ni) and bimetallic (5Ni/25Co and 10Ni/20Co) catalysts were thoroughly investigated by a wide number of analytical techniques including N2-physisorption, SEM coupled to EDX, XRF, STEM-HAADF coupled to EELS or EDX, Raman spectroscopy, XPS, H2-TPR and CH4-TPRe. Table 1 include the textural properties of the metal oxide catalysts. The corresponding pore size distribution are included in Fig. S1, Supplementary Material. The thermally-stabilized (calcined at 850 °C for 8 h) blank alumina support showed a surface area of around 140 m2 g-1 and a pore volume of 0.56 cm3 g−1. Its pore size distribution was bimodal with maxima located at 110 and 150 Å. After the addition of increasing amounts of cobalt (20Co, 25Co and 30Co samples), the surface area appreciably decreased to 120–108 m2 g−1 due to pore blocking. Accordingly, their pore volume was notably affected since it decreased to 0.35–0.29 cm3 g−1. The resultant narrower average pore size was in the 94–98 Å range. It was then evident that cobalt species preferentially deposited on the larger pores of the support (150 Å). In the case of the nickel catalyst (30Ni sample), the addition of the metal affected the textural properties to lesser extent when compared with its cobalt counterpart with the same loading (30Co sample). Thus, a surface area close to 130 m2 g−1 was observed. This was probably connected to a trade-off effect between the pore blocking of the support by nickel and the newly formed NiAl2O4 phase with a high intrinsic surface area. This rationale was supported by the notable surface area (170 m2 g−1) of an as-prepared bulk NiAl2O4, which was prepared by precipitation and calcined at 600 °C.Regarding the bimetallic catalysts, the addition of nickel (5–10 %wt.) to the Co/Al2O3 samples produced a loss of specific surface area0 around 7–8 % with respect to the corresponding monometallic sample with the same Co content (25Co and 20Co samples). Furthermore, while the pore volume remained almost constant, a slight increase in the mean pore size (from 94 to 98–107 Å) was found. In view of these results, it could be concluded that the deposition of the promoter had no marked effect on the textural properties, as relatively similar surface areas, pore volumes, and pore diameters were obtained independently from the Ni/Co ratio of the bimetallic samples.The microstructural morphology of the four monometallic samples (20Co, 25Co, 30Co and 30Ni) and the two bimetallic samples (5Ni/25Co and 10Ni/20Co) was examined by SEM. Irrespective of the composition of the catalysts, the micrographs (Fig. S2, Supplementary Material) revealed a heterogeneous surface on which irregular particles with sizes ranging from 5 to 20 µm are arranged with an aggregated morphology. Elemental identification and quantitative compositional information could be obtained by an energy dispersive X-Ray analyzer. Thus, the average surface composition of various defined regions (40 ×40 µm with a sampling depth of about 1 µm) for each catalyst was determined. Table 2 compares the bulk and surface composition as analyzed by XRF and EDX. As for the monometallic samples, an expected surface enrichment was found as revealed by their comparatively higher metal (Co or Ni)/Al molar ratios in relation to the respective bulk molar ratios. Particularly, this ratio at the surface as determined by EDX increased by a factor of 1.6–2.1 in the case of the Co-containing catalysts, and a factor of 1.1 in the case of the 30Ni sample.The monometallic 30Co and 30 Ni samples were also examined by scanning transmission electron microscopy–high-angle annular dark field (STEM–HAADF). Additionally, EELS elemental maps (Fig. S3, Supplementary Material) were obtained for certain regions in each sample to examine the spatial distribution of these metals in the catalysts. It was revealed that both Co and Ni were homogeneously distributed over the surface and no large uncoated support regions were apparently observed. This suggested a relative good metallic coverage of the alumina surface. The samples were characterized by the presence of polycrystallites (in some cases formed by the apparent attachment of smaller crystallites) with sizes ranging from 10 to 40 nm. It is worth pointing out the detection of crystalline phases on the surface of the 30Ni catalyst was comparatively less frequent, thereby suggesting the deposited metallic species on this sample exhibited a more amorphous nature.As for the bimetallic Ni-Co catalysts, it must be pointed that, although the Ni/Co molar ratio at the surface was higher than the corresponding bulk ratio, this increase was not very marked, from 0.23 to 0.25 over the 5Ni/25Co sample and from 0.58 to 0.60 over the 10Ni/20Co sample. This suggested a partial Ni diffusion into the cobalt catalytic layer. Likewise, surface chemical mapping, in this case carried out by STEM-HAADF coupled to EDX, was carried out to study the distribution of both metals on the surface of the bimetallic catalysts. As seen in the compositional maps included in Figs. 1 and 2, both cobalt and nickel were relatively well dispersed over the surface, with no visible clustering or agglomeration of either metal. Seemingly, the mixing between cobalt and nickel seemed to be equally intimate for both Ni-Co catalysts.X-ray diffraction analysis was used to identify the crystalline phases present in each oxide catalyst. The corresponding patterns are included in Fig. 3. The monometallic cobalt catalysts (20Co, 25Co and 30Co) showed the characteristic signals of a cubic spinelic phase (2θ = 19.2, 31.4, 37.1, 45.1, 59.6 and 65.5°) that would be in agreement with the formation of Co3O4 (ICDD 00–042–1467) and/or CoAl2O4 (ICDD 00–044–0160). Certainly, as will be evidenced later by both Raman spectroscopy and H2-TPR analysis, these samples consisted of a mixture of these cobalt oxides. However, while assuming the present cobalt aluminate will be preferentially amorphous under mild calcination at 600 °C, the visible diffraction signals in these patterns could be exclusively assigned to highly crystalline Co3O4. On the other hand, the reference 30Ni catalyst evidenced the typical signals of a cubic phase at 2θ = 37.2; 43.1; 62.9 and 75.4° corresponding to the presence of nickel oxide NiO (ICDD 00–089–7131). It must pointed out that although the formation of nickel aluminate is highly likely, this could not be detected probably due to its poor crystallinity. Recall that no clear signals attributable to crystalline NiAl2O4 (ICDD 00–078–1601) were observed. However, and similar to the results found for the Co catalysts, the existence of this spinel will be verified by redox and structural studies. Finally, a weak signal attributable to γ-alumina (ICDD 01–074–2206) support was also observed at 2θ = 67.2° over these four monometallic samples. The diffractograms of the bimetallic Ni-Co catalysts did not reveal the presence of segregated NiO, which suggested that this oxide was finely dispersed on the Co/Al2O3 matrix. Thus, only diffraction signals related to Co3O4 were noted. The crystallite size of this oxide (Table 1) was determined from the full width half maximum of the characteristic signal at 37.1°, using the Bragg equation. It thus ranged from 19 to 21 nm for the 20Co and 25Co samples to 35 nm for the 30Co catalyst. Interestingly, the addition of nickel to the Co/Al2O3 samples did not significantly alter the crystallite size (17–19 nm). On the other hand, the crystallite size of the NiO phase in the 30Ni catalyst was 14 nm.The examination of the structure of the oxide catalysts was carried out by Raman spectroscopy ( Fig. 4). The Raman spectra of the 20Co, 25Co and 30Co samples displayed the five typical vibration modes of Co3O4 at 196, 480, 520, 619 and 687 cm−1 [32]. The presence of CoAl2O4 was also evidenced by the two shoulders located at 706 and 725 cm−1 [33]. Apparently the contribution of these two additional signals was more marked in the case of the 20Co and 25Co, thus suggesting that the formation of cobalt aluminate would be favored with lower loadings of cobalt. Therefore, the cobalt phases present in the studied Co/Al2O3 catalysts would be a mixture of Co3O4 and CoAl2O4, with a higher relative abundance of the aluminate phase when the total Co content of the sample was lower. Lastly, the Raman spectra of the 30Ni catalyst was dominated by a wide signal located at 545 cm−1, which would be coherent with the presence of a mixture of NiO (its main Raman mode is located at 510 cm−1) and NiAl2O4 (its main Raman mode is located at 574 cm−1). The existence of the nickel spinel was further evidenced by the weaker signals at 746 cm−1 and 835 cm−1 [34]. On the other hand, the addition of nickel to the Co-Al2O3 catalysts did not substantially modify the spectra of the resulting samples (5Ni/25Co and 10Ni/20Co). Thus, the only observable Raman modes coincided with those corresponding to the parent cobalt catalyst, namely a mixed contribution of Co3O4 and CoAl2O4 phases. The marked presence of NiO and/or NiAl2O4 could be ruled out in these bimetallic catalysts.The surface composition of the samples and, more importantly, the distribution of the various metallic (cobalt and nickel) and oxygen species was investigated by analyzing the Co2p3/2 (777–792 eV), Ni2p3/2 (850–870 eV) and O1s (526–538 eV) XPS spectra of the samples, as shown in Fig. 5. Prior to the analysis in the XPS chamber, the as-calcined oxide catalysts were stored in airtight polyethylene containers in order to limit their exposure to ambient air. The Co2p3/2 spectra were deconvoluted into three main and two satellite contributions. The main contributions were located at 779.5, 780.7 and 782.4 eV, and were tentatively attributed to the presence of Co3+(Co3O4), Co2+(Co3O4 and/or CoAl2O4) and Co2+(CoO) species, respectively [35]. For all oxide catalysts, the relative abundance of the signal related to CoO was lower than 10 % of the total Co2p3/2 signal. This species was assumed to be formed by reduction under the vacuum conditions in the XPS chamber. The two signals located at 785.5 and 789.5 eV were assigned to the shake-up satellite peaks from Co2+ and Co3+ ions.Following a similar procedure, the Ni2p3/2 spectra were deconvoluted into five signals. The three main signals were centered at around 853.9, 855.4 and 856.9 eV and were associated with the presence of Ni2+(NiO), Ni2+(nickel belonging to a spinelic phase) and Ni3+(Ni2O3) species, respectively [36]. The satellite contribution of the spectra was dominated by an intense signal located at 861.0 eV, characteristic of the presence of Ni2+, and a small shoulder at 865.3 eV, which was a consequence of the relatively reduced presence of Ni3+ ions in these samples. Finally, the O1s spectra of the samples was characterized by three signals located at 529.3, 531.3 and 532.6 eV, which were attributed to oxygen species from the crystalline lattice (Olatt), superficially adsorbed oxygen species (Oads), and carbonate and hydroxyl species, respectively [37]. From the quantification of the aforementioned spectra, the elemental surface composition and the distribution of ionic species could be determined, as summarized in Table 2.As for the monometallic cobalt catalysts, it was observed that the Co3+/Co2+ molar ratio was in the 0.60–0.69 range, which was markedly lower than that expected for the exclusive presence of Co3O4. These moderate ratios suggested the presence of Co2+-rich oxides such as CoAl2O4, as previously pointed out by Raman spectroscopy. In fact, it could be inferred that cobalt aluminate was preferentially formed for low Co loadings, since the 20Co sample showed the lowest Co3+/Co2+ ratio (0.60). On the other hand, the Ni2p3/2 spectrum of the pure nickel catalyst (30Ni) clearly evidenced the presence of comparable amounts of nickel oxide and nickel aluminate. Hence, in addition to traces of Ni3+ species, the observed nickel was in the form of NiO (36 %) and Ni2AlO4 (49 %). The incorporation of nickel markedly affected the distribution of cobalt species on the 5Ni/25Co and 10Ni/20Co catalysts. Interestingly, the addition of this promoter favored the presence of Co3+ cations. This enrichment was a priori related to the partial insertion of Ni2+ ions into the structure of the Co3O4, which would imply the generation of the mixed NiCo2O4 spinel to some extent. In this sense, since the increased population of Co3+ ions was more noticeable for the 5Ni/25Co catalyst (with a Co3+/Co2+ molar ratio of 0.80) in comparison with the 10Ni/20Co sample (with a Co3+/Co2+ molar ratio of 0.64), a more extensive formation of nickel cobaltite was likely for low concentrations of the promoter. Accordingly, these samples showed a high population of Ni2+ species related a spinel-like phase (NiCo2O4) at the cost of Ni2+ as NiO. On the other hand, the presence at the surface of Ni3+ species (as Ni2O3) was also observed over the bimetallic Ni-Co and 30Ni samples, which was favored for low Ni loadings. Finally, it must be remarked that all these structural changes induced by nickel addition on the surface of the alumina supported cobalt catalysts led to an increase of lattice oxygen species. These are widely accepted to play a key role in methane oxidation [38]. Hence, the Olatt/Oads molar ratios were between 0.50 (10Ni/20Co) and 0.63 (5Ni/25Co), apparently higher than those of the respective Ni-free counterparts (0.25 for 20Co and 0.43 for 25Co).The analysis of the metallic catalysts by temperature-programmed reduction with hydrogen (H2-TPR) could be also helpful in identifying the nature of the oxide species present in each sample. The corresponding profiles are compared in Fig. 6. The redox behavior of the monometallic catalysts (20Co, 25Co, 30Co and 30Ni) was initially discussed in order to facilitate the subsequent interpretation of the results corresponding to the bimetallic Ni-Co samples. As for the Co-containing catalysts, two reduction events were clearly observable. Above 800 °C no measurable H2 consumption was noticed. Thus, the low-temperature uptake at 250–500 °C was assigned to the reduction of free Co3O4, according to the two-stage Co3+ → Co2+ → Co0 process [39]. The stoichiometric H2:Co molar ratio of this step is 1.33. The high-temperature consumption in the 550–750 °C corresponded to the reduction of the present cobalt aluminate [40]. The H2:Co stoichiometry for full reduction of this oxide is 1. Note that the presence of this highly stable oxide was in agreement with the results derived from both Raman and XPS spectroscopies. Table 3 includes the total H2 uptake of each monometallic sample, which increased from 3.5 to 5.2 mmol H2 g−1. The comparison of these values with the theoretical consumption expected when assuming that all cobalt was exclusively present as Co3O4 (which would vary from 4.2 to 6.1 mmol H2 g−1) resulted in reducibility degrees around 84–85 %. From these values, the relative distribution of Co atoms as Co3O4 or CoAl2O4 could be estimated. Hence, the abundance of cobalt as cobalt oxide gradually increased with the Co loading from 35 % to 37 % and 39 % over the 20Co, 25Co and 30Co samples, respectively.On the other hand, the fixation of the deposited metal as aluminate due to the strong metal-support interaction was observed for the 30Ni nickel catalyst as well. Therefore, its reduction trace also revealed two distinct H2 uptakes at moderate (400 °C) and high (700 °C) temperatures, which were associated with the presence of NiO and NiAl2O4, respectively [41], in consonance with the Raman and XPS results. It is worth pointing out that the stability of nickel aluminate was significantly higher than that of cobalt aluminate since its full reduction needed temperatures higher than 800 °C. Both oxides present a H2:Ni stoichiometry of 1. A relative good agreement was found between the experimental (4.5 mmol H2 g−1) and theoretical (4.6 mmol H2 g−1) consumptions. Consequently, a reducibility close to 100 % was evidenced. An estimation of the relative contribution of each nickel species suggested a roughly similar population of both oxide phases (43 %Ni as NiO and 57 %Ni as NiAl2O4).The incorporation of nickel to the 20Co and 25Co samples did not significantly altered the shape of the corresponding redox patterns since these also showed two reduction uptakes at low (250–475 °C) and high temperatures (550–750 °C). In view of the reduction pattern of the 30Ni sample, it was reasonable to expect that the reduction of the Ni2+ species present in the bimetallic samples (preferentially as free NiO) would occur mainly at the low temperature window, thus simultaneously coinciding with the reduction of Co3O4 species. Likewise, a small uptake at around 170 °C, which was not observed in the monometallic Co-Al2O3 counterparts, was visible. This consumption was assigned to the reduction of finely dispersed NiO nickel species [42,43], and was comparatively more noticeable for the 5Ni/25Co catalyst. In addition, the Ni-Co samples exhibited an appreciable shoulder at around 800 °C that was related to the reduction of nickel aluminate, probably due the strong interaction of added nickel with trace amounts of uncovered alumina.As shown in Table 3, owing to their higher total metallic loading the quantitative analysis of the reduction profiles expectedly evidenced a higher H2 uptake for the bimetallic samples (5Ni/25Co and 10Ni/20Co) in comparison with the respective Ni-free cobalt counterparts (25Co and 20Co, respectively). Thus, the overall reducibility increased from 4.4 to 5.2 mmol H2 g−1 in the case of 25Co and 5Ni/25Co catalysts, and from 3.5 to 5.2 mmol H2 g−1 in the case of the 20Co and 10Ni/20Co catalysts. Note that the total uptake of the Ni/Co samples (5.2 mmol H2 g−1) was virtually identical to that of the 30Co catalyst (5.2 mmol H2 g−1). Also relevant was the fact the reducibility, within the experimental error, of the bimetallic samples was promoted after the addition of nickel. Thus, it increased from 84 % over the 25Co sample to 90 % over the 5Ni/25Co sample, and from 84 % over the 20Co sample to 92 % over the 10Ni/20Co sample. This suggested that the incorporation of nickel promoted the presence of Co3+ cations with a higher H2 consumption per Co (1.5). As revealed by XPS, the simultaneous presence of nickel and cobalt could result in the formation of NiCo2O4-like spinel that ultimately increased the catalyst overall reducibility. Moreover, keeping in mind that the catalytic activity in the methane oxidation is expected to be mainly controlled by oxygen species consumed in the low-temperature range, it was found that the introduction of nickel was efficient for achieving this purpose. Hence, this uptake increased from 1.2 (20Co) to 1.6 mmol H2 g−1 (10Ni/20Co), and from 1.7 (25Co) to 2.0 mmol H2 g−1 (5Ni/25Co). In this latter case, a comparable uptake was found with respect to the 30Co catalyst.The reactivity of the available oxygen species present in the synthesized catalysts was complementary investigated by monitoring the conversion of methane in the absence of oxygen at increasing temperature (CH4-TPRe). The explored temperature range was 50–600 °C with a heating ramp of 10 °C min−1. The samples were then kept at 600 °C for 15 min. The composition of the product stream was followed by mass spectrometry (m/z = 44 (CO2), 28 (CO) and 2 (H2) signals). The resulting profiles of the bimetallic Ni-Co and monometallic (30Co and 30Ni) catalysts are shown in Fig. 7. Theoretically, methane is expected to be oxidized to carbon oxides at relatively low temperatures by active oxygen species at the catalyst surface. This will result in a progressive reduction of the metallic oxides, and a concomitant high-temperature conversion of methane into reforming products including CO, H2 and CO2, and/or cracking products (H2 and carbonaceous deposits) that will be catalyzed by partially reduced or metallic cobalt and/or nickel. Following this rationale, which is schematically depicted in Fig. S4 (Supplementary Material), the most relevant findings derived by this characterization technique were essentially those corresponding to the low temperature range, at which the complete oxidation of methane would be favorably occurring.The CH4-TPRe profiles revealed the formation of substantial amounts of CO2 at two relatively well-discernible temperature windows. On the one hand, the signal detected at lower temperatures (400–450 °C) was attributed to the gradual complete oxidation of methane by oxygen species. Note that no CO or H2 were detected in this temperature range. On the other hand, when the total oxidation process was no longer possible, the progressive reduction of the catalyst by methane then activated the transformation of the feed into CO, CO2 and H2, as can be evidenced by the co-existence of these three products at higher temperatures (500–550 °C). Moreover, the XRD analysis of the spent samples evidenced the presence of metallic cobalt (ICDD 00–015–0806) and nickel (ICDD 00–001–1258), and crystalline coke (ICDD 01–075–1621) (Fig. S5, Supplementary Material).As aforementioned, only the oxygen species involved in the low-temperature CO2 formation signal will be assumed to be highly active in the catalytic combustion reaction. After a proper quantification of the amount of formed CO2, the corresponding amount of consumed oxygen species could be estimated. In this sense, the 5Ni/25Co bimetallic catalyst showed the largest consumption (0.16 mmol O2 g−1) followed by the 10Ni/20Co sample (0.09 mmol O2 g−1) and the Co and Ni monometallic catalysts (0.08 and 0.04 mmol O2 g−1, respectively). In addition, it is worth pointing out that the 5Ni/25Co the oxidation reaction also started at significantly lower temperatures (200 °C) in comparison with the other samples (300–500 °C).The efficiency in the oxidation of methane into carbon dioxide of the four samples having the same nominal metallic content (30 %wt. %), namely 5Ni/25Co, 10Ni/20Co, 30Co and 30Ni catalysts, was analyzed operating at 300 mL CH4 g−1 h−1 between 200 and 600 °C. Three consecutive light-off tests were conducted over each catalyst. After the first test, which could be understood as an equilibration step of the catalyst under reaction conditions, a certain decrease in conversion was observed. Interestingly, no significant differences in conversion were found between the second and third tests resulting in a virtually identical light-off curve. Thus, the conversion profiles shown in Fig. 8 correspond to the third catalytic reaction run. All Co-based samples exhibited 100 % CO2 selectivity in the whole temperature range. Nevertheless, substantial amounts of carbon monoxide were formed over the 30Ni sample, leading to CO2 selectivity of only 90 % even at the highest reaction temperatures (600 ºC). It was observed that bimetallic catalysts exhibited a considerably better performance compared with the monometallic samples. The T50 values, listed in Table 4, were similar for the two monometallic catalysts (550 °C) and higher than those shown by the bimetallic counterparts (535 °C for 10Ni/20Co and 525 °C for 5Ni/25Co). Table 5.The specific reaction rates, calculated using the differential method (for conversions less than 20 %) at 450 °C, revealed a higher intrinsic activity of the 5Ni/25Co catalyst (0.80 mmol CH4 h−1 g−1), compared with the monometallic 30Co and 30Ni samples (Table 4). The other investigated bimetallic sample (10Ni/20Co) showed an intermediate behavior (0.63 mmol CH4 h−1 g−1). When referred to the total metallic loading, the best intrinsic activity of the 5Ni/25Co sample was also evidenced. From the correlations depicted in Fig. 9, the observed catalytic activity trend was coherent with the abundance of Co3+ species in the samples. Thus, the 5Ni/25Co catalyst presented the highest Co3+/Co2+ molar ratio due to the more efficient insertion of Ni2+ ions in the structure of the Co3O4 spinel leading to the generation of the nickel cobaltite-like species. As shown in Fig. S6 (Supplementary Material), this dependence was also valid when referred to the reaction rate normalized per gram of metal. The excellent behavior of this mixed oxide as oxidation catalyst for a variety of hydrocarbons [44,45], carbon monoxide [46], carbonaceous particulate matter [47] and methane [48] as well has been previously reported. On the other hand, it must be pointed out that both NiAl2O4 and CoAl2O4 spinel are not particularly active for the complete oxidation of methane [49,50], owing to their relatively low reducibility and highly stable oxygen species that penalized methane oxidative conversion by the Mars–van Krevelen mechanism. Besides, their formation could be detrimental for the generation of NiCo2O4 due to the decrease in the amount of available Co3O4 and Ni for their mutual interaction. In our study, the formation of this highly active mixed spinel was apparently enhanced with adding small amounts of nickel, since a Ni content as high as 10 %wt. did not lead to a better efficiency than the 30Co catalyst. This was probably owing to the fact that the incorporated Ni was more efficiently dispersed over the 5Ni/25Co catalyst in comparison with the 10Ni/20Co counterpart, as evidenced by its lower NiO/Ni molar ratio. This favored the interaction between Co3O4 and the deposited Ni to form NiCo2O4 to a larger extent as suggested by its large amount of Co3+. This increased presence of easily reducible Co3+ was accompanied by a concomitant higher presence of active lattice oxygen species that were able to activate the oxidation of methane at relatively low temperatures. This was also evidenced by the strong dependence of the intrinsic activity with the Olatt/Oads molar ratio and the amount of consumed oxygen at low temperatures in the CH4-TPRe runs (Fig. 9). On the other hand, it was found that the intrinsic activity of Olatt species present in the 30Ni catalyst was significantly lower than that exhibited by the Olatt species in the Co-containing catalysts.The Mars-van Krevelen mechanism, also known as the redox mechanism, has been widely used for kinetics modeling of methane oxidation over metal oxides. This is based on the assumption of a constant oxygen surface concentration on the catalyst, with reaction occurring by interaction between a molecule of reactant and an oxidized portion of the catalyst. Thus, the model assumes that the oxidation of the hydrocarbon occurs in two steps. In the first step, the compound react with the lattice oxygen resulting in its reduction and the corresponding formation of oxygen vacant site. In the second step, the reduced metal oxide is reoxidized by the gas phase oxygen present in the feed. In the steady state, the rates of the reduction and oxidation steps must be equal. Then, the kinetic equation (Eq. 1) can be expressed as: (1) ( − r ) = k red k ox P CH 4 P O 2 k ox P O 2 + γ k red P CH 4 where kred is the rate constant of the oxidation of the hydrocarbon by the lattice oxygen, kox the rate constant of the lattice re-oxidation and γ is the overall stoichiometry of the reaction. For conditions with oxygen excess (in our case, a PO2/PCH4 ratio of 10 at the inlet of the reactor), the term koxPO2 is considerably larger than γkredPCH4. Consequently, the rate equation simplifies to a power rate law equation (Eq. 2). (2) ( − r ) ≅ k red P CH 4 Accordingly, the integral method was applied to estimate the apparent activation energy when assuming a first pseudo-order for methane and a zeroth pseudo-order for oxygen [38,51]. Conversions between 10 % and 90 % were fit to the following linearized equation for the integral reactor (Eq. 3) where X is the fractional conversion of methane, k0 is the pre-exponential factor of the Arrhenius equation and FCH40/W is the weight hourly space velocity. The goodness of the numerical fit is depicted in Fig. S7 (Supplementary Material). It was observed that the apparent activation energy of the 30Ni catalyst (128 kJ mol−1) was markedly higher than that of the cobalt catalysts, in line with the lower activity of this catalyst for complete oxidation. The bimetallic catalysts and the 30Co catalyst showed a relatively similar value between 90 and 103 kJ mol−1. It is worth pointing out that this range of values was appreciably higher than that found for this reaction catalyzed by bulk Co3O4 (70–75 kJ mol−1) [52–54], thereby suggesting that the intrinsic activity of the examined cobalt catalysts was negatively affected by the presence of cobalt aluminate. (3) ln − ln 1 − X = ln k 0 C CH 4 0 W F CH 4 0 − E a RT Finally, given the presence of notable amounts of water vapor and carbon dioxide in the real exhaust gases of a natural gas engine, an attempt to evaluate the stability of the most efficient catalyst, namely the 5Ni/25Co sample, with time on stream was made. Thus, the evolution of conversion at 575 °C was examined when the composition was alternated following this sequence: 1 %CH4/10 %O2/N2, 1 %CH4/10 %CO2 /10 %O2/N2, 1 %CH4/10 %O2/N2, 1%CH4/10 % H2O 10 %O2/N2, 1 %CH4/10 %O2/N2, and 1 %CH4/10 % H2O/10 %CO2/10 %O2/N2. For each composition, a reaction time interval of 25 h was analyzed, with an accumulated time on stream of 150 h ( Fig. 10). During the first 15–20 h under base conditions (absence of water and CO2) a slight decrease in conversion from 80 % to 70 % was noticed. Then this conversion was stable, and was not affected by the addition of carbon dioxide for additional 25 h. Therefore, after an initial equilibration of the catalyst under reaction conditions, a relatively good thermal stability and resistance to the presence of CO2 was evidenced (75 h time on stream). However, after the admission of water into the reactor during additional 25 h, conversion dropped to a stable value of 40 % due to water adsorption on the surface [55]. Interestingly, when water was subsequently cut off, the methane conversion was almost fully recovered (65 %) upon returning to dry conditions. Thus, it was evidenced that this temporary inhibiting effect of water did not result in a remarkable irreversible deactivation of the sample. The catalyst was submitted to a further analysis under humid conditions but combined with the addition of carbon dioxide as well (25 h). Again, a decrease in conversion to 35 % was appreciated due to competitive effects caused by water.The state of the used catalyst in this long-term run was carried out by N2 physisorption, XRD and CH4-TPRe. The textural analysis revealed a slight decrease in surface area to 101 m2 g−1 (107 m2 g−1 for the fresh counterpart), thus suggesting the sintering of the active Co3O4 phase as in parallel confirmed by XRD. It is worth highlighting that irreversible poisoning was ruled out in view of the composition of the gas flow at the reactor inlet (CH4/O2/H2O/CO2). Besides, the formation of carbonaceous deposits (coke) was not observed given the net oxidizing character of the feedstream (PO2/PCH4=10 at the inlet of the reactor) that inhibited the eventual decomposition/cracking of methane. Hence, an enlargement of the crystallite size (25 nm, 19 nm for the fresh sample) was verified. These structural changes led in turn to a poorer oxidation ability at low temperatures judging from the results by CH4-TPRe analysis (Fig. S8, Supplementary Material). A shift of around 10 °C was noted for the peak oxidation temperature, from 410 °C (fresh sample) to 420 °C (used catalyst). However, it must be pointed out that the total amount of active oxygen species was not substantially modified (0.16 mmol O2 g−1).From a structural point of view, the monometallic samples consisted of a mixture of crystalline Co3O4 and amorphous cobalt aluminate in the case of the Co-containing catalysts (20Co, 25Co and 30Co), and a mixture of crystalline NiO and nickel aluminate in the case of the 30Ni sample. The formation of these undesired Al-based spinels due to the unavoidable strong interaction between the transition metal and gamma alumina was appreciable since around 40–65 % of the deposited metal was fixed as a metal-Al mixed oxide. It is worth pointing out that the generation of these aluminates was unfavored with the metallic loading.As revealed by STEM-HAADF coupled to chemical mapping the added nickel was homogeneously deposited on the surface of the corresponding cobalt catalyst as no clusters or visible agglomerates were distinguished. Thus, a relative good dispersion of the promoter could be inferred. As a result, the overall redox properties of the bimetallic catalysts were enhanced, which was essentially attributed to the formation of a new NiCo2O4-like spinel that increased the relative population of Co3+ species in the resulting Ni-Co samples. Hence, these structural changes induced by nickel led to an increase in the amount and mobility/reactivity of lattice oxygen species at lower temperatures with respect with the reference pure Co counterparts, which eventually resulted in a higher intrinsic activity and lower ignition temperatures for methane abatement. The optimal catalyst composition, which globally enhanced the abundance of Co3+ by a proper combination of highly active Co3O4 and NiCo2O4 phases, was that of the 5Ni/25Co sample. The 10Ni/20Co and the 30Co catalysts exhibited a similar efficiency. Therefore, this study demonstrated that the synergistic effect between the two metal sites is an efficient strategy to activate lattice oxygen species, which can affect the catalytic oxidation activity significantly. Andoni Choya: Investigation, Writing - original draft. Beatriz de Rivas: Methodology, Formal analysis, Validation. Jose Ignacio Gutiérrez-Ortiz: Methodology, Formal analysis, Funding acquisition. Rubén López-Fonseca: Conceptualization, Writing - review & editing, Supervision, Funding acquisition, Project administration.The authors declare that they have no known competing financial interests or personal relationships that could have appeared to influence the work reported in this paper.This research was funded by the Spanish Ministry of Science and Innovation (PID2019-107105RB-I00 AEI/FEDER, UE), Basque Government (IT1509-22) and the University of The Basque Country UPV/EHU (DOCREC21/23). The authors wish to thank the technical and human support provided by SGIker (UPV/EHU). In addition, authors acknowledge the use of instrumentation as well as the technical advice provided by the National Facility ELECMI ICTS, node ‘Advanced Microscopy Laboratory’ at University of Zaragoza.Supplementary data associated with this article can be found in the online version at doi:10.1016/j.jece.2022.108816. Supplementary material. .

In this work bimetallic Ni catalysts supported over Co-Al2O3 and monometallic Co-Al2O3 and Ni-Al2O3 catalysts were examined for the complete oxidation of methane. With a 30 % total metallic loading, the samples were synthesized by a sequential precipitation route. All samples were characterized by nitrogen physisorption, X-ray fluorescence, X-ray diffraction, Raman spectroscopy, scanning electron microscopy, scanning-transmission electron microscopy, X-Ray photoelectron spectroscopy, and temperature-programmed reduction with hydrogen and methane. Their catalytic performance was investigated in the temperature range of 200–600 °C with a space velocity of 60.000 h−1. The bimetallic catalysts showed a better behavior in the oxidation reaction than the monometallic counterparts, mainly due to the good dispersion of Ni on the surface of the Co-Al2O3 samples. This has enabled the insertion of Ni2+ ions into the cobalt spinel lattice, which in turn provoked an increase in the amount of Co3+ species, and a subsequent enhanced mobility of oxygen species in the spinel. In this sense, the 5Ni/25Co catalyst showed the best performance, thus reducing the value of the T50 by 25 °C with respect to the monometallic catalysts.

Brunauer–Emmett–TellerBarret–Joyner–HalendaCold gas efficiencyCassava rhizomeEnergy dispersive spectroscopyEquivalence ratioHigher heating valueInternational Union of Pure and Applied ChemistryLower heating valueMobil Composition of Matter No. 41Relative pressureScanning electron microscopeSimulated flue gasX-ray powder diffractionX-ray FluorescenceBiomass is one of the potential renewable energy resources [1]. Gasification is a promising thermochemical conversion, the main driver for converting biomass composition into useful gases and chemicals. Gaseous fuels from biomass gasification can be sources of producer gas (CO, H 2 , CH4, CO2), and syngas (CO, H 2 ). However, the inherent drawbacks of biomass are low energy density, hydrophilic materials, bulky volume and short time storage. Torrefaction is a pre-treatment process at a temperature of 200–300 °C in an inert atmosphere to increase the volumetric energy density, which can enhance the biomass conversion efficiency [2]. During torrefaction, the original component in biomass such as volatile compounds, lignocellulosic materials, inter and intra-molecular hydrogen and CO, CH bonds are destructed at different temperatures [3–5]. Generally, conventional torrefaction is carried out in a nitrogen atmosphere, which leads to a higher operating cost, stemming from the requirement of separation of N 2 from air. Oxidative torrefaction is another special torrefaction process, in which biomass is torrefied in an oxidative environment (containing 3–10 vol.% O 2 ). The study of Wang et al. [6] indicated that torrefied sawdust’s properties and its pellets in oxidative exposure, such as density and higher heating value, were close to those in inert atmospheres. From the work of Chen et al. [7], it was reported that higher heating value of liquid product derived from the torrefied palm oil fiber pellets in inert and oxidative exposure was in range of 10.10–13.20MJ/kg, which could increase to 23.20–28.70MJ/kg after dehydration. From research by Li et al. [8] which torrefied pine and poplar under CO2 carrier gas at temperature ranging from 220–340 °C. They concluded the higher temperature plays the reaction of deacetylation and dehydration while the mainly reaction decarboxylation occurs at the low temperature torrefaction. It can thus be concluded that using combustion flue gases as carrier gases for the torrefaction of biomass is feasible.Cassava rhizome (CR) biomass, agricultural residues in Thailand, can be converted into a gaseous product by gasification or pyrolysis. Many studies have evaluated CR behavior in thermo-chemical processes such as combustion, gasification, and pyrolysis. Previous studies have investigated cassava residues during later process such as fast pyrolysis of stalk and rhizome of cassava plants by a pyrolysis GC/MS [9]. There were some slow pyrolysis researches of palm kernel cake and cassava pulp residue in a fixed-bed reactor [10]. Homchat et al. [11] conducted slow pyrolysis of fresh and dried CR in a large scale metal kiln which resulted in less charcoal than fresh CR, due to the effect of the moisture content. Most of previous researches investigated the type of reactor; however, high oxygen component in CR (38-57 wt%) caused low heating value and oxygenated compound emission in the bio-oil product. In a few recent studies, it has been reported that torrefied biomass can significantly affect the efficiency of biomass gasification. Phanphanich and Mani [12] investigated the fuel characteristics and grindability of pine chips and logging residues torrefied at temperatures ranging from 225 °C to 300 °C and 30 min residence time. They found that high hemicellulose and lignin in the biomass produce more tar during the gasification. Tremel et al. [13] found that the overall gasification efficiency and carbon conversion efficiency of the entrained flow gasifier was observed to be superior for the smaller ( 160 μ m ) particles torrefied biomass compared to that of the larger ( 250 μ m ) particles.Various zeolite catalysts such as dolomite, olivine, and metal oxide have been introduced in biomass gasification or pyrolysis in order to improve the quality of the product. Several catalysts have been tested, either for coal or biomass gasification i.e., dolomite, fluid catalytic cracking catalysts (FCC) [14], and metal based catalysts [15]. Particularly, zeolites are widely applied in more than 90% of petrochemical and refining industries. During the thermo-chemical process, zeolite catalyzes to upgrade biomass (i.e., cellulose, cellobiose, D-glucose and xylitol) at moderate temperatures of 400–600 °C and enhances the yields of aromatic and aliphatic hydrocarbons. MCM-41 zeolite properties are a regular array of uniform and one-dimension mesopores. The extremely high surface area of ca. 900–1000 m2/g makes these materials promising candidates as catalysts or as catalysts support. Generally, metal such as nickel (Ni), shows excellent catalytic activity. Previous works have recorded improvement of catalytic activity and stability in steam gasification of biomass through Ni/MCM-41 [16], partial oxidation of CH4 [17], and CO 2 reforming of CH4 [18,19]. Many supporters of Ni catalyst, such as MgO, Al2O3, ZrO2, and CeO2 were tested in the activities. Moreover, porous structure of MCM-41 interaction with Ni metal are important for catalytic process during steam reforming of hydrocarbons into light products ( C 1 C 5 ) or the gasoline ( C 5 C 12 ) [16,20].MCM-41 can be synthesized from waste, such as cold fly ash or rice husk. Because abundance of silicon source composed in the solid waste. Research by Li et al. [8] found that the BET surface area and average pore diameter of MCM-41 synthesized from coal fly ash, were 1347 m2/g and 3.80 nm, respectively. Illite is a raw material in many industrial applications particularly in ceramics and refractories. Almost all the illite clay waste in Thailand was disposed of in the mining area after mining and dressing illite clay, which caused an environmental problem. To date there have been no systematic studies of the recovery of illite waste for MCM-41 synthesization. Illite waste can be major a silica (Si) source for MCM-41 zeolite synthesis. Illite waste was treated with an alkaline solution or silica, which make alumina to be the first extracted from clay with hot alkaline solution and consequently this process resulted in the supernatants. Then, the supernatants were applied as a starting material for MCM-41 zeolite synthesis by hydrothermal processing.According to, new trend of the renewable energy and zero waste and circular economy, the utilization of illite waste as the raw material for zeolite synthesis was focused on this work. The objective of this synergy study is to propose the value-added pathway on solving of the illite mining waste, flue gas emission and drawback of CR fuel. The transition metal Ni can be loaded on MCM-41 by impregnation or post-synthesis which is generally a desirable method. The focus of this work is to study MCM-41 synthesized by illite waste (Ni/MCM-41) for catalytic gasification of torrefied CR at 700 °C for 30 min for generation of high-quality gas products.Torrefaction of CR particle size of 0.425–0.850 mm and 0.850–2 mm was conducted at temperature of 260 °C for 60 min in nitrogen gas atmosphere and simulated flue gas (SFG) and used as a raw material for gasification. In case of SFG mixed, CO2 (15 vol.%) and O 2 (5 vol.%) in N 2 balance was applied in this work. The picture of the CR samples is displayed in Fig. 1. The element of initial CR before torrefaction such as carbon, hydrogen, nitrogen, and oxygen were 37.60, 5.41, 0.37, and 55.93, respectively. The properties of CR sample are listed in Table 1. The chemical compositions and phase analysis of illite waste were characterized by X-ray Fluorescence (XRF) and X-ray powder diffraction (XRD) techniques. The elements that are found in the highest quantities are O, Si, Al, Fe, K, and Na. These are also the major elements found in illite waste. The chemical composition of illite sample mainly consisted of 73.01wt.% SiO2, 16.52 wt% Al2O3, 5.28 wt% K 2 O and 2.38wt.% Fe2O3 and low contents of MgO, TiO2, Na2O, and SO3. Phase analysis of illite powders was determined by XRD (PANalytical, model X’ Pert Pro) with 40 kV, Cu K α radiation. The scanning ranges from 10-60° with a step size of 0.02 are shown in Fig. 2(a). Microstructure of illite sample was measured by scanning electron microscope (SEM) (Hitachi, model SU-5000) as shown in Fig. 2(b). Illite was fused with NaOH with the weight ratio of NaOH-Illite at 1.2:1. Calcined temperature of sample was 550 °C for 1 h. The illite fusion was slowly dissolved in 22.50 ml of deionized water for 24 h. The supernatant was obtained after filtration of suspension. 3.45 g of cetyltrimethylammonium bromide (99%, Aldrich Chem Co) was dissolved in 45 ml of deionized water (D.I.), mixed with 5.4 ml of ammonium hydroxide and continuously stirred at 25 °C for 30 min. After addition of 8.33–13 ml of tetraethyl orthosilicate (98%, Aldrich Chem Co) stirring continued until a homogeneous mixture emerged, with pH = 10.5–11.5 adjustment by acetic acid. The mixed liquid was transferred to a Teflon-lined stainless-steel autoclave and heated at 110 °C for 72 h. The precipitated powder of MCM-41 was filtered and washed with deionized water. MCM-41 was dried at 105 °C in oven and then calcined in air at 540 °C for 4 h. 5Ni/MCM-41 catalyst was prepared by impregnation and evaporation. A certain amount nickel (II) nitrate hexahydrate (Ni (NO3)2 ⋅ 6H2O, 98.5%, Aldrich Chem Co) loading (5 wt%) was dissolved in ethanol. MCM-41 powder was added to the mixture and stirred for 3 h followed by evaporation of the mixture at 50 °C. The solids obtained were calcined in a muffle furnace at 550 °C for 4 h with a heating rate of 1 °C/min in the presence of air.A downflow gasifier system consists of five main parts: (1) biomass feeder (2) carrier gas unit, (3) stainless steel reactor and catalyst holder, (4) condenser, and (5) gas filters and collection unit. Gasification zone and catalytic section temperatures were set at 700 °C, and 500 °C, respectively. The 5Ni/MCM-41 catalyst was mixed with silicon carbide (SiC) in the ratio of 1:55 and placed in a top holder section. Three thermocouples were fitted at the top and bottom part of the reactor, and at catalyst holder. The reactor was purged with N 2 to avoid combustion before operating. The carrier gas ( N 2 ) entered the reactor along with the gasifying agent ( O 2 ) which was fed into the bottom of the reactor. The ratio of N 2 and O 2 was adjusted to a target equivalence ratio (ER) of 0.4. The CR was continuously fed into the system of 1.0 g/min for a 30 min. The condensate products, such as tar were retained in condensers and gas washers. Gas produced was measured by means of volumetric gas meter after separation of condensate before conveyed into the main gas line by a vacuum pump at a flow rate of 0.5 L/h. The gaseous products such as CO2, CO, H2, CH4 and other hydrocarbons as C x H y were measured by Gasboard-3100p instrument. The solid portion was later collected for further analysis.X-ray diffraction pattern of synthesized catalyst was derived by XRD (Rigaku TTRAX III) with low angle range of 0.5–5° and XRD (PANalytical, X’ Pert Pro) for wide range of 20–70°. Ni metal dispersion of catalysts was analyzed by Energy dispersive X-ray spectroscope (EDS). Surface area, average pore size, and total pore volume of the fresh catalysts were determined by N 2 adsorption and desorption Brunauer–Emmett–Teller (BET) isotherms. Pore size, pore size distribution and pore volume were obtained by the Barret–Joyner–Halenda (BJH) pore analysis.As expected, at any torrefaction conditions, the oxygen component was decreased in the sample while more carbon was retained in the torrefied CR. For torrefaction in N 2 atmosphere, the carbon content for CR size 0.425–0.850 mm and 0.850–2 mm were 53.40 wt% and 51.49 wt%, respectively. Carbon contents in torrefied CR under SFG atmosphere drastically improved when compared to the original CR as noticed in Table 1. This effect of SFG is prominent for carbon and oxygen components in all torrefied CR particle sizes. The decrease of hydrogen and oxygen in torrefaction of CR process because of the breakage of OC and CC bonds [20]. After undergoing SFG torrefaction, OC atomic ratio (O/C) and HC atomic ratio (H/C) were reduced to become less than those under N 2 carrier gas. Proximate analysis of these samples was shown in Table 1, the loss of some of the organics affected the loss of the volatiles while ash content increased. Additionally, the heating value of CR was originally 15.6–15.9 MJ/kg and after torrefaction with N 2 and SFG, heating value was increased to 20.20–22.07 MJ/kg and 22.07–24.37 MJ/kg, respectively. These results can be expressed by the low-energy bond of HC and OC reduction and high energy bond of CC. Nitrogen and ash content increased in torrefied CR. This is simply attributed to the fact that all the components containing nitrogen and other minerals (in ash) retain in the biomass solid phase, whereas C, H, and O leave the solid.The XRD pattern of MCM-41 after calcination is illustrated in Fig. 3(a). The MCM-41 catalyst samples display the high intensity peak in 2 θ of 2.16°, 3.74°, 4.30°, and 5.72° and sharp diffraction peak ( d 100 ), ( d 110 ), ( d 200 ) and ( d 210 ), respectively. These diffraction peaks indicate a long-range ordered hexagonal mesoporous structure of MCM-41 synthesis from illite waste. The hexagonal diameter and pore wall thickness can be calculated by equation of a 0 = 2d 100 ∕ 3  [21]. The a 0 of MCM-41 was 45.71 Å.The XRD pattern of fresh 5Ni/MCM-41 is exhibited in Fig. 3(b). The diffraction peaks of Ni phase at 42° and 50° occurred at the NiO crystalline phase at 37°, 43° and 51°. The other crystalline such as Ni2SiO4 (nickel silicate) and Al2SiO5 (aluminum silicate) were observed at 26° and 61°, respectively. It can be concluded that Ni atoms displayed good dispersion in the support porous. The nitrogen adsorption–desorption isotherm and pore size distributions of catalyst sample are illustrated in Fig. 4(a) and 4(b). All the samples are isotherm type IV according to IUPAC classification. In Fig. 4, the immediate increase in the region of 0.25 < P/P0 < 0.40 is related to the capillary condensation inside the mesoporous wall [17,19]. In general, the long range at higher relative pressure suggested that the adsorption continued on the surfaces of MCM-41 sample at P/P0 > 0.45 due to an increase in pore size. The isotherm of MCM-41 sample presented mesoporous filling steps with pore size larger than 40 Å. In Fig. 4b, the isotherm shows an identical shape, although the adsorption capacity decreased with Ni loading on MCM-41 support because the particles of nickel oxide finely dispersed inside the MCM-41 supported porous by impregnation with ethanol. The textural properties of synthesized catalysts are presented in Table 2. The surface area of 5Ni/MCM-41 catalysts was slightly decreased from 804.03 m2/g to 737.88 m2/g. The 5Ni/MCM-41 pore diameter was close to MCM-41 supported sample, because of the dispersion of Ni metal particles in MCM-41 and less blockage of MCM-41 supported pore by the impregnation method [22]. It can be noted that the porosity of the catalyst is not significantly changed in this work. SEM technique was applied for analysis of the surface topology and to assess the dispersion of Ni components covered on supporter. SEM image of illite waste, MCM-41 support and 5Ni/MCM-41 can be seen in Fig. 5. All the samples contained irregular shapes and grain sizes [23]. The surface morphology of MCM-41 synthesized at pH 10.5–11.5 gave well-order with diameter size around 100 nm. The small amount of Ni particles can be observed on MCM-41 surface, whereas some Ni particles were found inside the MCM-41 pores [24,25]. Ni metal dispersion was obtained for 3.6 and 5.6 ± 0.2 wt% by EDS for analysis. In this work, the torrefied CR was used in catalytic gasification. The presence of nickel enhanced the gas fraction of product yield. Liquid yield from CR size 0.425–0.850 mm was lower than the larger size in gasification with a catalyst. Carbon and hydrogen conversion were calculated as the molar of CO and H 2 produced. The syngas composition from no catalyst gasification experiment consisted of CO (3.32–7.90 vol.%), H 2 (2.48–3.02 vol.%), CH4 (2.23–3.29 vol.%), C x H y (0.07–0.16 vol.%) and CO2 (9.30–16.52 vol.%). Torrefied CR had lower O/C ratio, and when it was gasified, the torrefied CR produced lower CO2. Gasification with 5Ni/MCM-41 showed higher H 2 and CO for torrefied CR with lower CO 2 concentration, in tune with the findings of Tapasavi et al. [26]. Tar was effectively removed by 5Ni/MCM-41 catalyst. Typically, Ni based catalysts exhibit high tar cracking. This metal property along with the ongoing Boudouard and water gas shift reaction activities allow favorable composition adjustment of H 2 and CO more than that of CO2 in the product gas. The CO and H 2 concentration obviously increased with the use of a catalyst, as can be seen in Table 3. In the downflow reactor, the gasification of torrefied CR reactions can be explained the following main equations (Eqs. (1)–(6)). Devolatilization in gasification of torrefied CR occurs more than others since volatiles react with themselves in gas–gas phase. In addition, the water-gas shift reaction can increase the CO2 amount in syngas. These are explained in gas-phase phenomena of volatile gases released during gasification. In gas phase reactions, the water gas shift reaction is important for increasing the H 2 in syngas, while the methanation influences the CH4 product. The Boudouard reaction converts CO2 into CO. Because temperatures are below 1000 °C, this reaction is in equilibrium and the CO remains in the synthesized gas. (1) Boudouard reaction  C + CO 2 → 2 CO (2) Water-gas shift reaction  CO 2 + H 2 ↔ CO + H 2 O (3) Methanation reaction  C + 2 H 2 → CH 4 (4) C + 3 H 2 → CH 4 + H 2 O The nitrogen compounds contained in materials enhanced the direct release of isocyanic acid as part of volatile matter during degradation of CR at low temperature in downflow reactor. HNCO can react with steam from water gas shift reaction, yielding ammonia (NH3) and CO2. The moisture in CR could participate in the reaction involving hydrogen cyanide (HCN) and give rise to NH3 and CO as per equation below [27]. This results in a comparatively steady amount of CO. (5) HNCO + H 2 O → NH 3 + CO 2 (6) HCN + H 2 O → NH 3 + CO The effect of 5Ni/MCM-41 catalyst on torrefied CR gasification was studied. The catalyst holder was placed on the top of reactor. The operating temperature of catalyst was set at 500 °C, which is generally the range for tar cracking in the gasification, similar to the use of Ni catalyst in gasification temperature of 500 °C and 600 °C of cedar wood and sunflower stalk, respectively [16,25]. Surface acidity active sites of 5Ni/MCM-41 can improved the cracking or reforming reaction. In this work, the gas production and carbon conversion slightly increased from 66 to 73 wt% and 75 to 80%, respectively. H conversion obviously increased from 18.47 to 27.39% as illustrated in Fig. 6. Having surface acidity active sites on 5Ni/MCM-41 can improve the cracking or reforming reactions. All researches reported over 60% conversion of reforming biomass tar over Ni-based catalysts [28–30]. Moreover, synthesized 5Ni/MCM-41 can crack the vapor fraction in the system which increase the product gas volume due to the decomposition of gaseous components of the synthesized gas obtained during cracking of liquid products [31]. 5Ni/MCM-41 will drive Sabatier reaction forward and yield more CH4 as described in equation below [32]. (7) CO 2 + 4 H 2 → CH 4 + 2 H 2 O Combined with the characterization results, the trend of the catalytic performance of the active Ni surface areas and dispersion can be clearly observed. Ni/MCM-41 enhanced the higher H2 production from CR gasification, suggesting that some small Ni particle sized of 3 nm maybe inside the MCM-41 porous which promoted water gas shift and reforming reactions of C x H y and CH4. This is because of the reactants’ longer residence time inside of the MCM-41 pores [20].Cold gas efficiency (CGE) is the fraction of energy output over the energy input. The heating value of the gas products obtained from the gasification of torrefied CR is presented in Fig. 7. A lower O/C ratio in torrefied CR indicate the increasing amount of C and the calorific value which implies higher gasification efficiency. The ranges of CGE varies with heating value of gas products. The minimum of gas heating value (lower heating value) was 7.78–8.37 MJ/kg without catalyst gasification, while the maximum range was 9.38–10.03 MJ/kg. These results indicated a high CO, H 2 and hydrocarbon gaseous mixture in the gas product. Gasification efficiency in gasification without catalyst has some issues worth mentioning as per following. Utilizing CR size 0.425–0.850 mm will yield slightly more CGE than larger CR while smaller CR obtaining from torrefied under SFG will yield considerably 10% less than larger size. CGE is fundamentally calculated from heating value of produced gas and CR biomass sample, yet there is another parameter namely gas-volumetric which induced CGE to become disproportional and not corresponding with other results. For example, gasification of CR size 0.850–2 mm dropped for 3%, which is statistically insignificant. It can be concluded that addition of Ni/MCM-41 enhanced overall efficiency. Catalytic gasification of CR size 0.85–2 mm obtaining from torrefied under SFG showed outstanding result in term of CH conversion and H 2 /CO data.Gasification of torrefied CR with 5Ni/MCM-41 catalyst was investigated in this work under partial oxidative atmosphere such as simulated flue gas from typical power plant. Results from the experiments confirmed the benefits of torrefaction of CR prior to gasification under described atmosphere. Illite waste can be utilized as a precursor to MCM-41 synthesis and added as a catalyst in gasification of torrefied CR. This work has revealed the synergy of utilizing torrefaction and catalyst from waste for energy generation considering from its efficiency in various factors such as gasification of torrefied CR under SFG atmosphere and N 2 with catalyst would yield 10% less unfavorable liquid. MCM-41 supported Ni metal was prepared by ethanol assisted impregnation method. NiO dispersed throughout on MCM-41 supported and partially formed into Ni2SiO4 which showed high surface area and pore volume. 5Ni/MCM-41 plays an important role on total gas yields and CO and H 2 conversions with its excellent properties in gasification and tar cracking. In summary, applying SFG for CR torrefaction is a promising technique to produce quality green fuel at economical cost for further utilization as renewable energy via catalytic gasification process employing illite waste as precursor for synthesizing effective catalyst.The authors declare that they have no known competing financial interests or personal relationships that could have appeared to influence the work reported in this paper.This work was supported by Royal Golden Jubilee Ph.D. Programme, Thailand [grant no. PHD/0212/2557]. Additional research scholarships were provided by Overseas Academic Presentation Scholarship for Graduate Students, Thailand, and the 90th Anniversary of Chulalongkorn University Fund (Ratchadaphiseksomphot Endowment Fund), Thailand . The authors would like to thank the National Metal and Materials Technology Center (MTEC) and Interdisciplinary Program in Environmental Science, Graduate School, Chulalongkorn University.

In this work, torrefaction of cassava rhizome (CR) under nitrogen gas ( N 2 ) and a simulated flue gas (CO2 (15 vol.%) and O 2 (5 vol.%) in N 2 balance) atmosphere was examined in a downflow reactor at 260ºC for a residence time of 60 min to produce a superior solid fuel for subsequent 5Ni/MCM-41 catalytic gasification of CR utilization. Mesoporous molecular sieves (MCM-41 zeolite) was synthesized from illite waste as a silica source. The MCM-41 synthesis was carried out by hydrothermal and post-synthesis for Ni loading. Various characterization techniques, such as XRD, SEM, and BET were employed to thoroughly characterize catalyst. High surface area (737.88 m2/g) and a typical type IV pattern of hysteresis loop (0.25 –0.40) obtained 5Ni/MCM-41 catalyst is calculated by N 2 adsorption–desorption technique. Catalyst characterization and discussion of results are presented in this work. 5Ni/MCM-41 catalyst strongly enhances the H 2 and CO production from gasification of torrefied CR at a temperature of 700ºC. Carbon and hydrogen conversions were 80.17% and 27.39%, respectively while liquid yield was lower than 10 wt%. The syngas from the conversion maintained H 2 /CO ratio of 0.55 with the highest gaseous efficiency of 49.35%. Obviously, synergy of synthesized 5Ni/MCM-41 catalyst and torrefied CR with gasification is valuable useful as potential renewable energy generation process.

Utilization of CO2 is currently a hot topic in catalysis due to the chance to decrease anthropogenic CO2 emissions on the one hand and to recycle it as a C1 source in exchange to fossil fuels on the other hand. So called power-to-gas (PtG) and power-to-liquid (PtL) technologies enable chemical storage of surplus energy from regenerative sources by reaction of renewable H2 with CO2 to energy carrier such as methane (PtG) or liquid fuels (PtL) [1–3]. Especially the PtG technology has high potential as a chemical energy storage technology since infrastructure for fast energy generation as well as a natural gas grid based on fossil natural gas is already well established and a state-of-the-art-technology. Hence, renewably produced CH4 via PtG can be easily feed into the existing gas grid and in a future perspective completely replace fossil natural gas.Ni is the state of the art catalyst for CO2 methanation (Eq. (1)) already since its discovery by Paul Sabatier in 1902 [4] and has been center of several studies on various supports, whereat reviews can be found elsewhere [1–3,5–14]. (1) CO2 + 4 H2 → CH4 + 2 H2O Besides Ni, also other metals are active in CO2 methanation [15,16]. Mills and Steffgen classified the important metals for methanation catalysts by its activity (Ru > Fe > Ni > Co > Mo) and selectivity to methane (Ni > Co > Fe > Ru) [17].Ni shows high activity with a very good selectivity to CH4. Nevertheless, traditional Ni-catalysts suffer from deactivation by sintering of the Ni particles upon heat evolution from the highly exothermic methanation reaction [18]. Deposition of coke and formation of volatile nickel carbonyls contribute to additional catalyst deactivation [19,20]. Besides, Ni is of toxicological concern. The sequences of Mills and Steffgen point out, that Fe has a very high activity for CO2 activation but suffers from low selectivity. In contrast to Ni, iron is not toxic, is much more abundant and hence around 180 times cheaper than nickel.Surprisingly, only a few studies focus on optimization of Fe based catalysts for CO2 methanation. Kirchner et al. investigated bare iron oxide samples in the CO2 methanation and obtained best activity for nano-sized γ-Fe2O3 with maximum CH4 yield of 60 % at 400 °C and ambient pressure [21]. In addition, pure α-Fe2O3 based catalysts can be promoted with 2 wt % Mg in order to increase the basicity and hence interaction of CO2 with the catalysts. This promotion leads to improved CH4 yield up to 32 % at 8 bar and a GHSV of 10,000 h−1 [22]. The results emphasize that the methanation takes place predominantly on surface carbon and iron carbide species on promoted bulk Fe2O3 catalysts [22]. In general, the high activity of Fe for CO2 activation results from the high reverse-water-gas-shift (RWGS) activity (Eq. (2)) and especially at elevated pressure its further capability of CO hydrogenation via Fischer–Tropsch-Reaction (FTR) (Eq. (3)). (2) CO2 + H2 → CO + H2O (3) nCO + m/2n + 1 H2 → 1/n CnHm + n H2O Lee et al. investigated the CO2 hydrogenation via FTR on Fe catalysts at 1–25 bar and in various H2/CO2 ratios [23]. They found that metallic Fe transforms into mixtures of magnetite and carbides under reaction conditions. Especially in the pressure range of 1–10 bar the increase of pressure leads to an increase of the chain length and higher temperature increases the CO2 conversion as well as CO and CH4 yield. In contrast, the produced H2O from FTR contributes to the equilibrium of the RWGS-reaction that limits the CO2 conversion [23]. In line, on K and S promoted Fe-based catalysts it was shown that the CO2 methanation activity is strongly influenced by the H2/H2O ratio effluent from the reactor [24,25]. It was claimed that conversions increase with increasing H2/CO2 ratio and cannot be further improved than their maximum CO2 conversion of 44 % obtained at 20 bar and a H2/CO2 ratio of 8 [24].With the aim of tailoring Fe-based materials as CO2 methanation catalysts, studies on increasing the C2–C4 fraction in CO-FTR, with CH4 as an undesirable product, provide information on the direction of necessary properties for high CH4 yields: In general, iron carbides are considered as the active phase in FTR and active carbon sites contribute to the chain growth mechanism [26]. In addition, the activity and selectivity is closely related to the particle size of the Fe-based catalysts [26]. Smaller Fe nanoparticles (<7–9 nm) lead to higher CH4 selectivity [27–29]. It was concluded, that low coordinatively unsaturated corner and edge sites are important for CH4 formation, while terrace sites of the bigger Fe particles are responsible for olefin generation [29,30]. Hence, the selectivity of Fe based catalysts for CO2 methanation could be improved by decreasing the Fe particle size. This stands in contrast to the particle size dependency of Ni based CO2 methanation catalysts, which decrease in selectivity if the particle size decrease below 2 nm [31,32].Supporting Fe on zeolites enables a way to produce stable and highly dispersed Fe species. This has been proven by their use as highly stable selective catalytic reduction (SCR) catalysts [33–35]. Despite the high Fe dispersion, zeolites offer additional tailoring possibilities and have shown to positively influence the CO2 methanation performance of Ni-based catalysts [36]. Namely by their compensating cation [37], Si/Al ratio [38] and zeolite framework type such as FAU, BEA, MFI and MOR [39]. Due to the high affinity of zeolites to adsorb water they allow further improvement of catalytic activity by applying so called sorption enhanced conditions whereat H2O is adsorbed by the zeolite and in that way pulled away from the reaction center [40–44]. To the best or our knowledge it was not investigated yet how the combination of Fe supported on zeolites perform in CO2 methanation. In the present study a series of differently loaded Fe on zeolite catalysts are investigated at ambient and elevated pressure up to 15 bar with the aim to increase the CO2 methanation performance. In order to avoid restricting the product spectrum resulting from pore size effects within the zeolite, 13X was selected as zeolite support. On the one hand due to its relatively large and three dimensional pore structure where molecules up to a kinetic diameter of 7.35 Å can form and diffuse freely along all axis including CO, CH4, CH3OH as well as C–C coupled products up to at least C6 compounds. On the other hand a large range of Fe loadings can be theoretically ion exchanged due to the high aluminium content of 13X. As main focus the trends in activity and selectivity with increasing pressure as well as iron loading are carefully analysed and correlated with the properties of the catalysts. This leads to a justification if CO2 methanation on Fe-based catalyst will become feasible as an attractive alternative.1, 5 and 10 wt % Fe/13X catalysts were synthesised by wet impregnation with a 0.05 M Fe(NO3)3 · 9H2O (99 % Sigma-Aldrich) solution in ethanol on commercial Na-13X zeolite (ZEOCHEM, Si/Al = 2.5; Faujasite structure). After ion exchange for 24 h at room temperature under intense stirring ethanol was evaporated in a rotary evaporator. The resulting solids were dried at 80 °C for 12 h and calcined at 400 °C (heating ramp = 5 K/min) for 5 h in a continuous flow of air. The 5 wt % catalyst with collapsed zeolite structure was synthesised by wet impregnation for 30 min with a 0.05 M Fe(NO3)3 ∙ 9H2O aqueous solution on commercial Na-13X. Water was evaporated in a rotary evaporator, the resulting solid dried at 80 °C for 12 h and calcined at 500 °C (heating ramp 5 K/min) for 5 h in a flow of air.The weight loadings of iron of all samples were analysed via Inductively Coupled Plasma Optical Emission Spectroscopy (ICP-OES) on an Agilent 720 ES. The X-ray powder diffraction pattern were measured on a Bruker D8 Advance diffractometer with Ni filtered Cu Kα radiation (λ = 1.5406 Å) and a step size of 0.2° from 2°θ = 20–90. Crystallite sizes of the Fe-particles were calculated according the Debeye–Scherrer equation using the half width of the reflex at 44.7°. UV/vis spectra were collected on a UVVISNIR Lambda 950 spectrometer from Perkin Elmer equipped with a 150 mm integration sphere to analyse the diffuse reflectance of the Fe-zeolites. The spectra where recorded in reflexion mode in a wavelength region of 800–200 nm and a step size of 5 nm. Specific surface area, pore diameter and pore diameter dispersion were analysed by N2 physisorption at 77 K in a Quantachrome Autosorb IQ TPX. All samples were degassed for 12 h in vacuum at 200 °C. The pore diameter and dispersion were analysed according the BJH method from the desorption branch and specific surface area (SSA) by using the BET method. The pressure range for analysis was defined by rouquerol analysis in order to stay in the linear regime of the BET analysis [45]. The microporous surface area was distinguished from the external and mesoporous surface area by the t-plot method. Temperature controlled analysis were performed in the same Quantachrome Autosorb IQ TPX in dynamic mode and with a thermal conductivity detector. For temperature programmed reduction (H2-TPR) all samples were degassed at 400 °C in a flow of N2 for 30 min. Subsequent to the cooling down procedure to 40 °C, TPR was started in a flow of 5 vol% H2 in N2, with a total flow rate of 25 mL/min and a heating ramp of 5 K/min up to 850 °C and isothermally treated at the end temperature for additional 30 min. NH3 was used in order to analyse the acidic properties of the zeolite in the temperature controlled desorption (TPD) experiments. Prior to the analysis all samples were reduced in a flow of 50 % H2 in N2 at 400 °C for 30 min, accordingly to the pre-treatment of the catalytic tests. Residual adsorbed hydrogen was flushed-off from the samples by additional 2 h treatment in N2 at 400 °C. Subsequently, adsorption of 10 % NH3 in N2 was performed at 100 °C and physisorbed NH3 was purged in a flow of N2 at 100 °C for 30 min. TPD was performed in a flow of 25 mL/min N2 and a heating ramp of 10 K/min up to 800 °C. Scanning electron microscopy (SEM) analysis was performed in a Thermo Scientific Phenom XL equipped with a back scattered detector. Concurrent elemental mapping was carried out by using the integrated EDX detector.Methanation tests were performed in fixed bed flow reactor system with an inner diameter of 6 mm at ambient and elevated pressure (5, 10, 15 bar) at a GHSV = 4186 h−1. Prior to the catalytic tests all catalysts were reduced within the reactor in a flow of 50 % H2 in N2 for 30 min at 400 °C and ambient pressure. In a typical run 25 mL/min CO2, 100 mL/min H2 and 12 mL/min N2 as internal standard were supplied by mass flow controller (Bronkhorst, El Flow). During the methanation tests the temperature was raised from 200 to 400 °C in steps of 50 °C and kept constant at reaction temperature for 30 min. The composition of effluent gases from the reactor was monitored by online raman spectroscopy (Kaiser Raman RXN2 spectrometer equipped with AirHead probes). The conversion X, selectivity S and reaction rate of CO2 conversion (r(CO2)) were calculated according Eqs. (4)–(6): (4) X C O 2 = n ˙ C O 2   i n -   n ˙ ( C O 2   o u t ) n   ˙ ( C O 2 i n ) (5) S C H 4 =   n ˙ ( C H 4   o u t ) ∑ n ˙   ( p r o d u c t s ) (6) r C O 2 =   X C O 2   ×   n ˙ ( C O 2 )     n F e c a t   With n ˙ i as the molar flow of component i, and n(Fecat) as the molar amount of Fe in the catalyst bed within the reactor.Catalysts with three different weight loadings (1, 5, 10 wt %) of Fe on 13X were prepared via impregnation. Elemental analysis via ICP-OES confirms the presence of Fe on 13X close to the aimed amounts of Fe on the samples (Table 1 ).Since the zeolite framework is prone to destruction by iron, the integrity of the structure was validated via XRD analysis.The impregnation procedure and calcination temperature strongly influences the stability of the iron impregnated zeolites. Hence, a synthesis optimization was conducted: the zeolite structure stays intact only by avoiding H2O as a solvent and using ethanol as well as decreasing the calcination temperature to 400 °C (Fig. 1 ). Nevertheless, with higher Fe-loading the decrease of intensity of reflexes shows the incipient destruction of the framework even by applying the optimized procedure. Compared to the pure 13X, 1 wt % Fe/13X shows nearly no changes in intensity and all catalysts show reasonable stability. In contrast, the zeolite structure of 5 wt % Fe/13X impregnated in H2O and calcined at 500 °C completely vanishes. For this reason, all catalysts were prepared in ethanol and by calcination at 400 °C and it was avoided to exceed this temperature at any time. As a pretreatment in the catalytic test a reduction of the catalysts in 50 % H2 in N2 at 400 °C for 30 min was performed. Comparison of XRD of ex situ reduced and as calcined catalysts (Fig. 1) ensure that the zeolite structure stays intact for all Fe loadings during the pre-reduction and confirm the Fe-reduction by the raise of the specific reflex of metallic Fe at 2°θ of 44.7° (inset in Fig. 1). According to the Debeye–Scherrer equation extracted Fe crystallite sizes from this reflex are 33 and 23 nm for 5 wt % and 10 wt % Fe/13X, respectively. Solely the reduced 1 wt % Fe/13X does not show this specific reflex. This could be due to two reasons, or a combination thereof: Either the Fe loading is too low for the sensitivity of XRD or the Fe species are highly dispersed within the framework of the zeolite.N2 physisorption analysis confirms the presence of microporosity of all Fe/13X catalysts calcined at 400 °C. Nevertheless, the specific surface area decreases from 612 to 161 m2/g with increasing Fe loading. In line with the decreasing reflex intensity of the zeolite lattice from XRD analysis the micropore area extracted from t-plot analysis decreases from 573 down to 32 m2/g (see Table 1).The dispersion of Fe within the zeolite framework after calcination was analysed with UV/vis spectroscopy. The line shape of the spectra arising from O → Fe3+ charge transfer are rather similar (Fig. 2 ). In all spectra, four distinct peaks are separated by deconvolution (Figs. S1–S3). Two strong bands are found below 300 nm that are assigned to isolated Fe3+ ions. Whereat the band centered at 205 nm attributes to charge transfer from tetrahedral coordinated Fe3+ and the band at 250 nm relates to Fe3+ in higher coordination [34]. The two bands above 300 nm arise from agglomerated Fe-species. Whereby the band from octahedral Fe3+ species in small oligomeric FexOy cluster appears at 350 nm and from large Fe oxide particles as a very broad band at 436 nm. Quantitative analysis of the deconvoluted bands shows that all samples have the same relative amount of Fe3+ in tetrahedral sites. Contrary to this, 1 wt % Fe 13X shows with 55 % of all Fe3+ ions relatively more Fe ions in dispersed and oligomeric octahedral sites. Solely 30 % of the Fe ions agglomerate to particles. In comparison to this, the two higher loaded samples have comparable factional amounts of Fe in all sites and more than 55 % of Fe agglomerate into particles.The reducibility of the Fe/13X catalysts was investigated by H2TPR experiments (Fig. 3 ). In line with the Fe loading of catalyst the intensity of the signals increases and the features of 5 and 10 wt % Fe/13X samples are rather similar. These two samples show a very intense and broad signal between 200–550 °C with a peak maximum that shifts to lower temperatures from 442 to 405 °C with increasing Fe loading from 5 to 10 wt %. In agreement with literature these signals correspond to the reduction of Fe of the agglomerated FeOx particles and dominate the TPR [46]. In addition, two more signals appear at temperatures higher than 550 °C that go in line with the collapse of the zeolite structure.In the TPR of 1 wt % Fe/13X three distinct peaks appear in the temperature region of zeolite’s thermal stability with peak maxima at 375, 424 and 498 °C. According to literature, reduction of Fe3+ within the zeolite structures as well as reduction of Fe2O3 to Fe3O4 from oligomeric and small cluster takes place at lower temperature [47]. The visibility of the fine structure of reduction under the same measurement conditions shows on the one hand that agglomerated FeOx-species are not the main species, and on the other hand, that Fe species coordinated on different sites of the zeolite framework are present in this sample.Temperature programmed desorption of NH3 was performed in order to analyse the influence of the Fe loading on the zeolites acidity (Fig. S4). In line with the decrease of reflex intensity of the zeolite framework in the XRD with increasing Fe loading the total number of acid sites decreases. The main signal in the TPD appear at the same temperature region. Hence, even though the number of acid sites decreases with increasing Fe load, the acid strength as well as nature of acid sites remain constant in all samples. Therefore, it can be excluded to significantly influence the selectivity of the catalysts.All prepared materials were investigated in a temperature region of 200–400 °C and at pressures of 1, 5, 10 and 15 bar.In a first step the two 5 wt % Fe/13X with a collapsed (prepared in H2O and calcined at 500 °C, broken lines in Fig. 4 ) and intact zeolite structure (calcined at 400 °C & exclusion of H2O from the synthesis, solid lines in Fig. 4) were compared by their catalytic performances. In case of the catalysts with a collapsed framework after synthesis a rather low CO2 conversion of 10 % was observed by increasing the temperature up to 400 °C, even at 15 bar. For this reason, the temperature range of the catalytic test was expanded to 550 °C for this catalyst. At 1 bar no significant CO2 conversion was observed up to 550 °C. Likewise all other investigated catalysts, the CO2 conversion increases with temperature and increasing pressure of the catalytic tests. With 5 wt % Fe on collapsed 13X reasonable CO2 conversion was achieved up to 74 % at 550 °C and 15 bar. On this catalyst, CO is the main product at low pressure. With increasing pressure, selectivity towards CH4 increases up to 85 % at 15 bar and 550 °C. No Fischer–Tropsch products were observed under any conditions.On the contrary, 5 wt % Fe/13X with an intact zeolite framework shows already reasonable CO2 conversion of 33 % at 1 bar and 400 °C. CO2 conversion increases with temperature and pressure up to 88 % at 400 °C and 15 bar. Incipient activity is already obtained at 250 °C. Hence, the comparison of these two catalysts clearly demonstrates that an intact zeolite framework is essential to obtain and support high catalytic performances at reasonable temperatures.In order to stay in the kinetic regime the Fe-normalized reaction rates at 300 °C are used to compare the activity of produced catalysts with intact zeolite structure and different Fe loading at pressures from 1 to 15 bar (Fig. 5 ). The catalyst masses included in the reactor and corresponding Fe-content from ICP analysis were used to calculate the Fe molar reaction rates. The two catalysts with 5 and 10 wt % Fe/13X show similar and increasing reaction rates at increasing pressure of up to 12 and 8 mmol(CO2)/(mol(Fe)∙s), respectively. This points out that the main active sites are the same in these two catalysts. In opposition to these results, 1 wt % Fe/13X shows much higher reaction rates at all investigated pressures up to 42 mmol(CO2)/(mol(Fe)∙s) at 10 bar. These trends show in correlation to the UV/vis analysis that finely dispersed Fe-species, which are the main species in 1 wt % Fe/13X, have a much higher catalytic activity than the agglomerated Fe-species, that are the main species of the 5 and 10 wt % loaded Fe/13X catalysts. Nevertheless, the reaction rate decreases upon further increase of the pressure from 10 to 15 bar against the principle of Le Chatelier. This might be either due to hampered desorption of one of the products from the catalyst surface or reconstruction of the Fe-species or zeolite framework at elevated pressure.The catalytic performances at 350 °C are used to compare the variation of product selectivity of different Fe loading and at varying pressures (Fig. 6 ).At 1 bar the 10 wt % Fe/13X catalyst shows a very high selectivity towards CO (S(CO) = 97 %) and minor selectivity to CH4 (Fig. 6a). With increasing pressure to 15 bar the selectivity towards C–C coupled products and CH4 increases monotonously. CH4 becomes the main product at 10 bar and reaches its maximum selectivity of 61 % at 15 bar. The selectivity to C–C-coupled products increases up to 22 %, while the selectivity towards CO decreases to 15 %.Likewise, 5 wt % Fe/13X catalyst shows the same trend with increasing pressure (Fig. 6b). At high pressures it increases its selectivity towards the desired product CH4 up to 68 %, while the selectivity towards CO (S(CO) = 14 %) and CC-coupled products (S(CC) = 17 %) stays relatively low.In contrast to the behavior of the two higher loaded Fe catalysts 1 wt % Fe/13X shows at 1 bar already significant CH4 selectivity of 22 % (Fig. 6c). The CO selectivity of 21 % at 350 °C and 1 bar is relatively low and selectivity towards C–C coupled products is at 56 % and therefore surprisingly high in that sequence. In opposition to the trend of 5 and 10 wt % Fe/13X as well as literature [26] on Fe-based Fischer–Tropsch catalysts, the selectivity towards C–C-coupled products decreases with increasing pressure on 1 wt % Fe/13X. The main product is CH4 from 5 to 15 bar with a selectivity up to 76 % at 10 bar and 350 °C. Comparable product selectivites of S(CO) = 11 % and S(CC) = 14 % are observed at 10 and 15 bar. This trend of decreasing selectivity towards C–C coupled products and increasing CH4 selectivity with increasing pressure is opposed to the general trend of Fe-based Fisher–Tropsch catalysts reported in literature [26], in which Fe3C is regarded as active species. But it stands in line, that more coordinative unsaturated Fe species have a higher tendency to produce CH4 [29,30].XRD analysis of the catalyst after the catalytic tests shows a decrease of the reflexes from the zeolite framework for all samples (Fig. 7 ). Nevertheless, 1 wt % Fe/13X shows considerably high intensities. Hence, the integrity of the framework is still given in the mayor fraction of the sample even though a small fraction of the zeolite framework collapses. In the XRD of this sample no other reflexes from other phases than the 13X framework are visible. In contrast to the results of 1 wt % Fe catalysts 5 and 10 wt % Fe catalyst show significant decrease and a full depletion of the reflexes from the zeolite. Additionally, reflexes from a Fe3C phase appear in the diffractograms of both catalysts after the methanation experiments. Given by the sharp shape of metallic Fe reflex in the 5 and 10 wt % Fe/13X catalysts the crystallite size of Fe increases during the catalysis. On the low loaded 1 wt % Fe/13X still no reflexes origin from metallic Fe or Fe3C, respectively, under reaction conditions. Hence, the included Fe is very stable within the framework of the zeolite.The collapse of the zeolite with high Fe loading is visible in the SEM micrographs as well. The spherical shape of the zeolite crystallites is still visible in the used 5 wt % Fe/13X catalysts (Fig. S5). This shape nearly vanishes completely on the 10 wt % Fe/13X catalyst after operation. Larger fragments with different morphology, consisting of Al and Si, become obvious instead (Fig. 8 bottom, middle & right). In addition to this, the formation of larger Fe particles is visible, too. In comparison to SEM images of the reduced catalysts prior to the catalytic testing it seems that Fe migrates out of the zeolite particles and forms, together with deposited carbon, an outer shell around the support (Fig. 8 bottom). In the case of the used 10 wt % Fe/13X catalyst the EDX mapping indicates that residual zeolite particles do not contain Fe at all, while the concentration of Fe is comparably high on the amorphous fragments (Fig. 8). As opposed to this the spherical shape of the 1 wt % Fe/13X zeolite catalyst particles appears to be unchanged after the reaction in the SEM micrographs (Fig. 8 top). In addition, EDX analysis shows a homogeneous dispersion of Fe over the sample and no larger particle agglomerations of Fe or creation of a common Fe–C shell/layer is visible. Hence, the SEM micrographs confirm together with XRD analysis the destruction of the higher loaded zeolite during the catalytic run under formation of a Fe3C shell at the outer layer of the catalysts.The hydrogenation of CO2 towards CH4 on differently loaded Fe/13X catalysts was investigated at ambient and elevated pressure (5–15 bar). Comparison of the catalytic performances with a catalyst with collapsed zeolite framework shows, that an intact zeolite structure and hence high dispersion of Fe within the catalyst is essential for high CO2 conversion at temperatures below 400 °C and all investigated pressures.Catalytic tests on 10, 5 and 1 wt % Fe/13X catalysts with intact zeolite structure revealed a different reactivity of the two higher loaded catalysts compared to the 1 wt % Fe/13X catalyst. Higher Fe-loading leads to relatively low reaction rates of up to 12 mmol(CO2)/(mol(Fe)∙s) at 15 bar. CO is the main reaction product at low pressures of 1 and 5 bar. With increasing pressure the selectivity towards CH4 as well as C–C-coupled products increases. The low Fe-loading of 1 wt % leads to a significant increase of the molar reaction rate at all investigated pressures up to 42 mmol(CO2)/(mol(Fe)∙s) at 300 °C and 10 bar. In contrast to both higher Fe-loadings, the lower Fe-loading leads to high selectivity for C–C-coupled products of 56 % at 1 bar. The selectivity towards desired CH4 increases up to 76 % with increasing pressure at the expenses of the formation of CO and C–C-coupled products.Physico-chemical characterization before and after the catalytic run show on the one hand that in 5 and 10 wt % Fe catalysts, Fe is mainly present as agglomerated particles. This leads to a destabilization of the zeolite and further agglomeration of Fe under reaction conditions with simultaneous formation of Fe particles embedded in a Fe3C-phase as an outer shell layer. On the other hand, in 1 wt % Fe/13X, Fe is mainly present as octahedrally coordinated dispersed and oligomeric species. This leads to a higher hydrothermal stability of the catalysts and neither formation of larger Fe agglomerates nor Fe3C-phase formation under operation. The high dispersion of Fe within the material suppresses CC coupling reactions at higher pressure due to confined neighboring Fe sites and this in turn supports the hydrogenation of CO2 to methane. At 15 bar the selectivity towards CO is limited down to 8 %. Even though the performance is not yet fully optimized, the presented results show, that the utilization of Fe-based catalysts as alternative to more expensive and especially hazardous Ni-catalysts for e.g. biogas upgrading and feed into the natural gas grid becomes considerable and provides essential prerequisites for the direction of further catalyst optimization.The authors declare that they have no known competing financial interests or personal relationships that could have appeared to influence the work reported in this paper. Tanja Franken: Conceptualization, Methodology, Validation, Formal analysis, Investigation, Writing - original draft, Writing - review & editing, Visualization, Project administration, Funding acquisition. Andre Heel: Conceptualization, Methodology, Writing - review & editing, Supervision, Project administration, Funding acquisition.The authors kindly acknowledge funding of the SmartHiFe Project by the Swiss Federal Office of Energy (grant number: SI/501754-01). Additionally the authors kindly thank Michal Gorbar and Dr. Roman Kontic for their support during the SEM and XRD analyses.Supplementary material related to this article can be found, in the online version, at doi:https://doi.org/10.1016/j.jcou.2020.101175.The following are Supplementary data to this article:

The raise of regenerative but unsteadily produced energy demands a highly flexible way to store the energy for time periods when less energy is produced than consumed. In the current study, it is investigated if catalysts based on environmentally more attractive and less hazardous to health Fe might be able to be considered as an alternative to Ni catalysts in the CO2 methanation at elevated pressure. For this a set of catalysts with 1–10 wt % Fe supported on the zeolite 13X is analysed in CO2 methanation at 1–15 bar. The trends of activity as well as selectivity with varying Fe loading and pressure are presented. Correlation with thorough characterization of the materials shows that a very high dispersion of Fe in octahedral sites within the zeolite is necessary to generate CH4 as the main reaction product and suppress the Fischer–Tropsch activity towards CC coupling reactions at elevated pressure. Especially with low Fe loading such as 1 wt % high reaction rates of 42 mmol(CO2)/(mol(Fe)∙s) with a CH4 selectivity of 76 % at 300 °C and 10 bar are obtained. In contrast to that, highly Fe loaded catalysts tend to form increasing amounts of Fischer–Tropsch products at increasing pressure. In addition, highly Fe-loaded catalysts are much more susceptible to destruction of the zeolite under reaction conditions. At the same time, highly loaded catalysts form a Fe3C shell around the remaining support. Hence, avoiding the formation of a Fe3C phase is crucial for high CH4 selectivity. The results presented here therefore show that catalysts with a very high Fe-dispersion in particular can gain considerably in importance as alternatives to Ni-methanation catalysts at elevated pressure.

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