**Preface to "Catalytic Conversion of Energy Resources into High Value-Added Products"**

Catalysis is involved in a large number of industrial processes for the production of chemicals, pharmaceuticals, fuels, and the cleaning or suppression of environmental pollutants. More than 90% of the chemical processes used in industry use catalysts, with heterogeneous catalysts being the most used (80%). Although the global market for catalysts is important per se, the greatest impact of catalysts comes from the value generated by the chemicals and fuels they produce.

The growing concern about the massive use of fossil fuels and its effect on climate change has encouraged research in the development of technologies related to a more efficient use of energy resources. Scientific advances in this area of knowledge are therefore essential to meet the challenges related to global warming and the finite nature of fossil fuels. Therefore, the development of active, selective and energy efficient heterogeneous catalytic processes is of paramount importance for the production of high-value-added products from energy resources in a more sustainable manner.

In this Special Issue of Energies, we collect some of the latest progress in the development of cleaner, more efficient processes for the conversion of these feedstocks into valuable fuels, chemicals and energy. A total of 8 high-quality papers focused on different catalytic systems are showcased. Most of the works are focused on the conversion of biomass related materials which clearly reflects the paramount importance that the biorefinery concept will play in the years to come.

The novelty and contributions of these papers are briefly summarized in the next paragraphs.

Ramirez-Reina et al address the upgrading of biogas, a gaseous mixture of methane and carbon dioxide produced by the anaerobic digestion of biodegradable matter, via dry reforming into high value syngas using nickel-based catalysts. The novelty of this work lies on the use of multicomponent Ni-Sn/CeO2-Al2O<sup>3</sup> catalyst and the optimization of the process conditions, namely temperature and biogas composition. When compared towards a benchmark catalyst, the multicomponent catalyst achieved similar conversion and benefited from greater coke resistance, confirming the promotion effect of both tin and ceria.

One of the most critical technical challenges in biomass gasification is the tars formation, which can cause serious risks to downstream equipment. Therefore, tars should be removed from the biomass gasification effluent stream. In the contribution by Millan et al, a commercial Ni/Al2O<sup>3</sup> catalyst was used to determine the optimum conditions of the steam reforming of toluene, used as tar model compound. Thus, the influence of reforming temperature, steam to carbon molar ratio (S/C) and gas hourly space velocity (GHSV) on the toluene reforming performance was thoroughly studied.

These process parameters are considered key to fine-tune the reaction and maximize the overall performance. A temperature of 800 ºC, GHSV of 61,200 h -<sup>1</sup> and S/C ratio of 3 provided the most suitable reaction conditions for toluene conversion and H<sup>2</sup> production in steam reforming of toluene. Under these optimized conditions, a steady state of toluene conversion over 94% and a H<sup>2</sup> production of 141.6 mol/mol toluene with no obvious deactivation observed in five-hour test was obtained.

Biofuels can be obtained through thermochemical routes based on the pyrolysis of biomass wastes and subsequent catalytic conditioning to reduce the oxygen content, a topic brilliantly addressed by Garc´ıa et al. Firstly, a two-step catalytic process (in-situ catalytic pyrolysis using CaO followed by a catalytic cracking of the vapours released using ZSM-5 zeolites) was used to directly obtain an upgraded bio-oil with an oxygen content of 16.4 wt.%. The resulting bio-oil was mixed with ethanol and gasoline in a 2/8/90 vol% ratio, and its behavior was studied in a stationary spark ignition engine, showing similar fuel consumption than pure gasoline at the same engine conditions. Additional benefits from gasoline blending with bio-oil and ethanol were a reduction in both the PAH and the carcinogenic equivalent concentrations, thus decreasing the environmental impact of the exhaust gases.

A different strategy for bio-oil upgrading based on the hydrodeoxygenation (HDO) reaction was used in the contribution of Suelves et al. In this work, a series of different carbon materials (namely, carbon nanofibers, carbon nanotubes, graphene oxide and activated carbons) were used as support for the preparation of Mo2C based catalyst via carbothermal hydrogen reduction (CHR). The differences in the catalyst characteristics and their catalytic behavior were addressed using guaiacol, an aromatic model compound of fast pyrolysis bio-oils. Mo2C catalyst supported on carbon nanofibers showed the best catalytic activity, with the highest selectivity to oxygen-free HDO products and guaiacol conversion as compared with other carbon-supported catalyst tested. This performance was related to its higher gasification resistance, good Mo2C dispersion and crystal formation in CHR, which makes it the most suitable carbon material support for Mo2C-based catalysts, paving the way for the deployment of HDO bio-oil upgrading technologies.

The use of novel catalyst to produce valuable chemicals from renewable resources is addressed in the three following contributions. Thus, Solsona et al studied the catalytic synthesis of γ-Valerolactone (GVL), a valuable chemical that can be used as a clean additive for automotive fuels, via dehydration and hydrogenation of levulinic acid, a biomass derived compound. A series of natural and low-cost clay-materials were used as Ni support. Different strategies aiming to avoid the addition of pressurized H<sup>2</sup> were pursued, such as the use of Zn in the aqueous media and the addition of an H<sup>2</sup> donor molecule (formic acid). Best results were obtained using Zn in the reaction media at 180º C and Ni supported on attapulgite or on a high surface area sepiolite as catalysts, leading to γ-valerolactone yields higher than 98%. The cellulose conversion into sorbitol via hydrolytic hydrogenation was addressed by Roman–Mart ´ ´ınez et al using mesoporous activated carbon supported Ru catalysts. Although a positive effect of a large amount of acidic oxygen surface groups was foreseen, such groups promoted the formation of by-products, and lower the sorbitol selectivity.

Therefore, highest sorbitol selectivity (91%) at 52% cellulose conversion was obtained under relatively mild reaction conditions with the pristine commercial mesoporous AC without further treatment. Another approach was used by Taylor et al to convert glycerol, a byproduct of biodiesel synthesis, into methanol and other useful chemicals, in a one-step low pressure process without the addition of hydrogen gas. A careful analysis of the process conditions was carried out and the influence of the surface area of CeO<sup>2</sup> as redox catalysts on their performance for glycerol conversion was addressed. The importance of the morphology of the catalyst and its impact on the of the reactivity of glycerol and its intermediates was proposed.

The last contribution of this special issue focused on the prospects of obtaining high quality hydrochar from red jujube branch, a by-product jujube industry, via hydrothermal carbonization. The hydrochar synthesis was optimized in terms of carbonization temperature and solid residence time, resulting in a material with very interesting properties as solid biofuel.

In conclusion, the papers collected in this Special Issue provide a very comprehensive snapshot about some of the most interesting trends in the catalysis area, such as the use of carbon supported catalyst, nickel as a replacement of noble metal and ceria as redox material. It is striking that most of the work here published are focused on the transformation of biomass derived molecules, clearly reflecting the enormous deal of attention that the technologies related to the development of the catalytic process in the context of biorefinery is gaining in the last years. This gives also a glimpse of the promising prospects of the catalysis in the area of biomass conversion technologies. We hope that all the reader enjoys these outstanding collections provided by our colleagues from some of the most prestigious research centers and universities wordwide, as we did as editing this special issue. Of course, all our gratitude for their outstanding work and implication. Special thanks to Ms. Vickie Zhang and Ms. Sally Xu for all their support during the edition of this special issue and its reprint as book format.

> **Jos ´e Luis Pinilla, Isabel Suelves, Tom´as Garc´ıa** *Editors*

### *Article* **Biogas Upgrading Via Dry Reforming Over a Ni-Sn/CeO2-Al2O3 Catalyst: Influence of the Biogas Source**

#### **Estelle le Saché, Sarah Johnson, Laura Pastor-Pérez, Bahman Amini Horri and Tomas R. Reina \***

Chemical & Process Engineering Department, University of Surrey, Guildford GU2 7XH, UK; e.lesache@surrey.ac.uk (E.l.S.); sj00174@surrey.ac.uk (S.J.); l.pastorperez@surrey.ac.uk (L.P.-P.); b.aminihorri@surrey.ac.uk (B.A.H.)

**\*** Correspondence: t.ramirezreina@surrey.ac.uk; Tel.: +44-148-368-6597

Received: 15 February 2019; Accepted: 13 March 2019; Published: 15 March 2019

**Abstract:** Biogas is a renewable, as well as abundant, fuel source which can be utilised in the production of heat and electricity as an alternative to fossil fuels. Biogas can additionally be upgraded via the dry reforming reactions into high value syngas. Nickel-based catalysts are well studied for this purpose but have shown little resilience to deactivation caused by carbon deposition. The use of bi-metallic formulations, as well as the introduction of promoters, are hence required to improve catalytic performance. In this study, the effect of varying compositions of model biogas (CH4/CO2 mixtures) on a promising multicomponent Ni-Sn/CeO2-Al2O3 catalyst was investigated. For intermediate temperatures (650 ◦C), the catalyst displayed good levels of conversions in a surrogate sewage biogas (CH4/CO2 molar ratio of 1.5). Little deactivation was observed over a 20 h stability run, and greater coke resistance was achieved, related to a reference catalyst. Hence, this research confirms that biogas can suitably be used to generate H2-rich syngas at intermediate temperatures provided a suitable catalyst is employed in the reaction.

**Keywords:** biogas; syngas production; DRM; Ni catalyst; bi-metallic catalyst; ceria-alumina

#### **1. Introduction**

Rising greenhouse gas (GHG) emissions and the associated global warming threat are some of the largest challenges facing the world today. The energy sector alone accounts for two-thirds of total greenhouse gas emissions, and around 80% of carbon dioxide emissions [1]. The majority of CO2 emissions are produced from the combustion of fossil fuels for the generation of energy, and from other industrial processes. With global energy demand growing rapidly, the need for the decarbonisation of the energy industry has never been greater. Although CO2 is the primary greenhouse gas emitted, the detrimental impact of releasing methane, with a global warming potential (GWP) of 25 [2], into the atmosphere should not be underestimated.

In recent years, more efforts have been made by national governments to implement GHG mitigating processes and technologies, such as carbon capture and storage. However, there is still much to be done to meet the target agreed on by 'The Paris Agreement' of limiting the global average temperature to below 2 ◦C above pre-industrial levels. Therefore, it is crucial to redirect the global focus from fossil fuels to renewable energy sources in order to minimise the destructive effects of climate change.

Biogas is one possible source of renewable energy and is commonly referred to as a gaseous mixture of methane and carbon dioxide produced by the anaerobic digestion of biodegradable matter. Typical feedstocks for biogas are waste materials, such as municipal waste, sewage sludge and agricultural waste. Other components present in biogas include water vapour, nitrogen and hydrogen

sulphide. However, the actual composition, as well as the CH4/CO2 ratio, differs depending on the type of feedstock and the digestion process used [3]. Although biogas is the main renewable energy source contributing to the global energy supply [4], technical challenges still prevent it from being fully utilised for the generation of electricity.

Alternatively, biogas can be converted into high value syngas, a gaseous mixture of carbon monoxide and hydrogen, via the dry reforming of methane (DRM), as shown in Equation (1).

$$\text{CH}\_4 + \text{CO}\_2 \leftrightarrow 2\text{CO} + 2\text{H}\_2 \quad \Delta \text{H}\_{298\text{K}} = +247 \text{ kJ/mol} \tag{1}$$

Syngas is an extremely useful intermediate as it is used as a precursor to synthesise valuable fuels and chemicals such as methanol, as well as long chained hydrocarbons via the Fischer–Tropsch process [5,6]. Furthermore, syngas has been suggested as a feed gas to high temperature solid oxide fuel cells (SOFCs) with either internal or external reforming to generate electricity [7–9]. Figure 1 illustrates the opportunity to generate renewable energy from organic waste by combining anaerobic digestion with DRM and SOFCs. The high operating temperatures of SOFCs makes direct internal reforming feasible. However, the direct use of biogas in SOFCs is still associated with carbon deposition and sulphur poisoning in the fuel cell anode, which weakens cell durability and causes loss of cell performance [10].

**Figure 1.** Block flow diagram illustrating the potential of utilising organic waste to generate 'green energy'.

Due to the endothermic nature of the dry reforming reaction and the high stability of the reactants, high reaction temperatures and a stable catalyst are required to achieve high syngas yields [11]. However, high reaction temperatures often lead to the deactivation of the catalyst. Deactivation is caused by either solid carbon deposition forming on the catalyst, which physically blocks the active metal phase and is known as coking, or through sintering of the catalyst active phase at high temperatures. There are a number of side reactions which are responsible for the formation of carbon, and hence negatively affect the performance of the catalyst. In dry reforming, carbon is most commonly produced by the Boudouard reaction, as shown in Equation (2), methane decomposition, as shown in Equation (3) as well as carbon monoxide reduction, as shown in Equation (4).

$$2\text{CO} \leftrightarrow \text{C} + \text{CO}\_2 \quad \Delta \text{H}^\circ\_{298} = -171 \text{ kJ mol}^{-1} \tag{2}$$

CH4 ↔ C + 2H2 ΔH◦ <sup>298</sup> = 75 kJ mol−<sup>1</sup> (3)

$$\text{CO} + \text{H}\_2 \leftrightarrow \text{C} + \text{H}\_2\text{O} \quad \Delta \text{H}^\circ\_{\text{298}} = -131 \text{ kJ mol}^{-1} \tag{4}$$

Therefore, an effective dry reforming catalyst needs to be thermally stable in order to be resistant to coking and sintering, whilst yielding optimal conversions. Additionally, it must be economically viable so that it can be cost effectively scaled-up for industrial applications. All these factors are heavily dependent on the selected active metal, as well as the properties of the support and/or promoter [5].

Nickel-based materials are widely regarded as effective catalysts for the dry reforming of methane as they display good activity and respectable conversions [12]. However, the main drawback of Ni-based catalysts is that they suffer from quick deactivation as a result of carbon deposition and an inclination to sintering. Noble metal-based catalysts are frequently cited as a superior alternative due to greater stability and higher resilience to coking. Catalysts based on Ru and Rh in particular show higher activity compared to nickel [13]. Despite this, noble metal catalysts are far more expensive and have limited availability in comparison to Ni-based ones and are therefore unfeasible for use in industrial applications.

In order to develop a stable, high performing nickel-based catalyst, the material is often upgraded using an appropriate support or adding a promoter. Alumina (Al2O3)-based supports have been thoroughly investigated due to their high specific area which improves metal dispersion [14]. Yet implementing alumina as a support is still associated with catalyst deactivation, caused mainly by coke formation on the acid sites [14]. To combat this problem, basic promoters such as ceria (CeO2) can be used to stabilise the support. It has been shown that Ni catalysts supported on CeO2-Al2O3 systems showed far better performance than either CeO2 or Al2O3 supported nickel catalysts [15]. Indeed, ceria not only tunes the acid/base properties of the support but also provides excellent oxygen mobility (due to its high oxygen storage capacity), thus helping to prevent coking via oxidation of the coke precursors [16].

Moreover, bi-metallic systems can further enhance catalytic performance; the addition of a second metal has shown to improve the activity, as well as the stability of the catalyst [17–19]. It is believed that this is related to the metal–metal interactions which reduce the electron donor capacity of the active metal, limiting its tendency to form strong bonds with carbon precursors [20,21]. As recently demonstrated by Stroud et al., bi-metallic Ni-Sn supported Al2O3 and CeO2-Al2O3 catalysts can effectively generate good conversions for DRM whilst exhibiting high stability [16]. The positive effect of Sn as an added metal to Ni was further investigated by Guharoy et al. in a DFT study concluding that the beneficial effect of tin is credited to Sn atoms occupying C nucleation sites in the vicinity on Ni atoms, slowing coke formation (i.e., increasing the energy barrier for coke nucleation) [22].

Under these premises, the aim of this work is to investigate the performance of an optimised multicomponent Ni-Sn/CeO2-Al2O3 catalyst for the production of syngas from a biogas feed via the dry reforming of methane. The impact of the type of biogas and more precisely its origin (i.e., sewage, municipal waste, landfill and organic waste) is also a subject of this study which aims to showcase a suitable upgrading route for different types of biogases via reforming reactions.

#### **2. Materials and Methods**

#### *2.1. Catalyst Preparation*

The Ni-Sn/CeO2-Al2O3 catalyst was synthesised by sequential impregnation. First, the ceria promoted support was prepared by impregnation of the correct amount of Ce(NO3)2·6H2O (Sigma-Aldrich) on γ-Al2O3 powder (Sasol) in order to obtain a 20 wt % loading of CeO2. After 1 h of stirring, the solvent, acetone, was evaporated at reduced pressure in a rotary evaporator. The resulting powder was dried overnight at 80 ◦C and calcined at 800 ◦C for 4 h. The support was then impregnated in the same way with the necessary amount of Ni(NO3)2·6H2O (Sigma-Aldrich) dissolved in acetone, to achieve a 10 wt % metal loading. Lastly, the obtained solid was impregnated with SnCl2 (Sigma-Aldrich) following the same procedure to achieve a Sn/Ni molar ratio of 0.02. This ratio was chosen based on previous work [16]. For simplicity the catalyst will be referred to as Ni-Sn/CeAl from here on.

#### *2.2. Catalyst Characterisation*

N2-adsorption-desorption analysis was performed in a QuadraSorb Station 4 at liquid nitrogen temperature. Prior to the analysis, the catalyst was degassed at 250 ◦C for 4 h in vacuum. The surface area was calculated from the Brunauer–Emmett–Teller (BET) equation.

X-ray fluorescence (XRF) analysis of the catalyst was carried out on an EDAX Eagle III spectrophotometer using rhodium as the radiation source.

X-ray diffraction (XRD) analysis was conducted on fresh and spent samples using an X'Pert Powder instrument from PANalytical. The diffraction patterns were recorded over a 2θ range of 10 to 90◦. A step size of 0.05◦ was used with a time step of 450 s. Diffraction patterns were recorded at 30 mA and 40 kV, using Cu Kα radiation (λ = 0.154 nm).

Temperature-programmed reduction (TPR) in hydrogen was carried out on the calcined sample in a U-shaped quartz cell using a 5% H2/He gas flow of 50 mL·min−1, with a heating rate of 10 ◦<sup>C</sup> min−1. Prior to analysis the catalyst was treated with He at 150 ◦C for 1 h. Hydrogen consumption was measured by on-line mass spectrometry (Pfeiffer, OmniStar GSD 301).

Thermogravimetric analysis (DSC-TGA) was carried out on the samples post stability test in an SDT Q600 V8 from TA Instruments. The samples were ramped from room temperature to 900 ◦C at 10 ◦C min−<sup>1</sup> in air.

#### *2.3. Catalytic Activity*

The catalytic performance for the dry reforming of methane at varying compositions of model biogas (CH4 and CO2) was carried out in a continuous flow quartz tube reactor (10 mm ID) equipped with a thermocouple, at atmospheric pressure. Reactants and products were followed by an on-line gas analyser (ABB AO2020). In order to achieve a weight hourly space velocity (WHSVs) of 30 L g−<sup>1</sup> h<sup>−</sup>1, a mass of 0.2 g of catalyst was used for each screening. All samples were pre-reduced in situ using 100 mL·min−<sup>1</sup> of 10% H2 in N2 at 850 ◦C for 1 h.

Reactions were conducted using a total flow of 100 mL·min−<sup>1</sup> of varying biogas compositions for temperatures ranging from 550 to 850 ◦C in 50 ◦C increments. The catalytic performance was measured for CH4/CO2 molar ratios of 1, 1.25, 1.5 and 1.85 balanced in N2. These ratios were carefully chosen in order to model DRM of biogas produced from a range of residues [23]. Table 1 lists the types of biogas source, as well as the corresponding CH4 and CO2 content, being modelled by the specific CH4/CO2 molar ratio selected for each reaction.

**Table 1.** List of biogas sources and the corresponding methane and carbon dioxide content, and the resulting CH4/CO2 molar ratio. Adapted from Lau et al. [23].


Lastly, stability tests were carried out on Ni-Sn/CeAl and a reference nickel on alumina with the same Ni loading (Ni/Al) catalyst using CH4/CO2 molar ratio of 1 and 1.5 and a temperature of 650 ◦C for 20 h. The low temperature was chosen in order to simulate a syngas suitable for direct use in intermediate temperature SOFCs aiming for the hypothetical utilisation of biogas in an internal reforming SOFC.

Conversions of the reactants, as well as the H2/CO ratio and H2 yields, were calculated for each run to determine the effect of varying the biogas composition on the performance of the catalyst as follows:

$$\chi\_{\rm CH4}(\%) = 100 \ast \frac{[\rm CH\_4]\_{\rm In} - [\rm CH\_4]\_{\rm Out}}{[\rm CH\_4]\_{\rm In}} \tag{5}$$

$$\chi\_{\rm CO2}(\%) = 100 \ast \frac{[\rm CO2}{[\rm CO2}]\_{\rm In} - [\rm CO2}|\_{\rm Out} \tag{6}$$

$$\text{H}\_2\text{/CO ratio} = \frac{[\text{H}\_2]\_{\text{Out}}}{[\text{CO}]\_{\text{Out}}} \tag{7}$$

$$\text{Y}\_{\text{H2}}(\%) = 100 \ast \frac{[\text{H}\_2]\_{\text{Out}}}{2[\text{CH}\_4]\_{\text{In}}} \tag{8}$$

#### **3. Results and Discussion**

#### *3.1. Physicochemical Properties*

The N2 adsorption-desorption isotherm of the calcined catalyst is shown in Figure 2. The analysis generated a type IV isotherm with a characteristic hysteresis loop, which is associated with the presence of well-developed cylindrical mesopores. The steepness of the loop suggests that the mesopores were homogeneously distributed throughout the structure of the sample [24].

**Figure 2.** Nitrogen adsorption-desorption isotherm of Ni-Sn/CeAl sample.

The surface area, calculated using the BET method, the total pore volume and the average pore diameter of the calcined catalyst are listed in Table 2 along with its elemental composition determined by XRF. Although the textural properties were primarily governed by the γ-Al2O3 support, the surface area and pore volume measured were smaller than that of the support itself due to CeO2 and Ni particles present on the surface and covering the pores [25]. The metal content of the obtained catalyst was very close to the intended value.

**Table 2.** Textural properties of Ni-Sn/CeAl post calcination as determined by the Brunauer–Emmett– Teller (BET) equation and X-ray fluorescence (XRF) analysis.


#### *3.2. Reducibility: H2-TPR*

The temperature programmed reduction was conducted in order investigate the interactions between the support and metallic species, as well as the redox properties of the catalyst which are essential for DRM [5]. Figure 3 shows the H2-TPR profile of the calcined catalyst. The Ni-Sn/CeAl sample presents a main reduction peak at around 820 ◦C. This high reduction temperature is attributed to the reduction of Ni2+ to Ni, strongly interacting with the alumina support. The high reduction temperature may indicate the presence of NiAl2O4 spinel, which is difficult to reduce and is associated with high reduction temperatures of 600–870 ◦C [16,26]. It appears that a second reduction process occurs at 900 ◦C, possibly due to the reduction of CeO2 and Al2O3 to CeAlO3, in good agreement with the XRD results that will be discussed later in the paper [27,28].

**Figure 3.** H2-TPR profile of Ni-Sn/CeAl.

#### *3.3. XRD*

The structural properties of the Ni-Sn/CeAl catalyst after calcination (fresh) and after reduction under H2 at 850 ◦C for 1 hour were determined and the XRD profiles are shown in Figure 4. The fresh catalyst profile shows no presence of characteristic crystalline peaks associated to Sn or SnOx due to the low amount used, metallic nickel (Ni0) or oxidised nickel (NiO) species. The profile presents, however, peaks related to γ-Al2O3 (JCPDS# 00-048-0367) and NiAl2O4 (JCPDS# 00-010-0339). Although γ-Al2O3 and NiAl2O4 phases are hardly distinguishable due to their broad and overlapping diffraction peaks, the formation of nickel aluminate spinel (NiAl2O4) can be identified due to the slight shifts towards smaller angles of three strong diffraction peaks associated with γ-Al2O3: at 2θ 37.0◦, 45.5◦ and 66.3◦ [27]. Indeed, the nickel aluminate spinel can be formed at high calcination temperature (800 ◦C) via the reaction between NiO and Al2O3. Additionally, for the full formation of aluminate spinel, a Ni loading of around 33 wt % is necessary [29], therefore, since the Ni loading in this sample was much lower it is most likely that NiAl2O4 spinel co-exists with Ni nanoparticles supported on the γ-Al2O3 support.

**Figure 4.** XRD profiles of the fresh and reduced Ni-Sn/CeAl catalyst.

The fresh sample presents, as well, the typical diffraction peaks of CeO2 fluorite cubic cells (JCPDS# 01-075-0390) at 28.5◦, 33.1◦, 47.8◦ and 56.3◦. However, once the sample was reduced, an additional CeAlO3 tetragonal phase (JCPDS# 01-081-1185) was detected. When reduced at temperatures above 600 ◦C, Ce2O3 and γ-Al2O3 reacted to produce CeAlO3 in good agreement with the results discussed in the TPR section [28]. However not all CeOx species were reacted to CeAlO3 since traces of CeO2 cubic phase can still be detected at 28.5◦. Finally, the XRD pattern of the sample reduced in hydrogen shows the presence of metallic Ni particles (JCPDS# 87-0712), which indicates that Ni<sup>0</sup> will be the predominant active phase for the dry reforming reaction. The shift towards the higher angles of the 37.0◦, 45.5◦ and 66.3◦ peaks confirms that NiAl2O4 has been reduced to Ni metallic and γ-Al2O3 upon reduction at 850 ◦C, confirming the TPR results.

#### *3.4. Catalyst Performance*

The catalytic performance for DRM was studied over a range of temperatures (550–850 ◦C) at CH4/CO2 molar ratios of 1, 1.25, 1.5 and 1.85. The effect of varying the biogas feed concentration, in terms of CH4 conversion as a function of temperature, is reported in Figure 5a. It was shown that with increasing methane concentration, the overall CH4 conversion decreases. Indeed, methane was introduced in excess for DRM. Additionally, CH4 conversion increased with temperature; this aligns with expectations, as the endothermic nature of the reaction requires higher temperatures to reach equilibrium conversion. Hence, superior activity of the catalyst was observed during the temperature screening at a 1:1 molar ratio of CH4/CO2, which reached a methane conversion of 93% at 850 ◦C.

**Figure 5.** Catalytic performance of Ni-Sn/CeAl at CH4/CO2 molar ratios of 1, 1.25, 1.5 and 1.85 as a function of temperature: (**a**) methane conversion; (**b**) carbon dioxide conversion.

Figure 5b shows the conversion of CO2 as a function of temperature for various biogas compositions. The temperature screening carried out at CH4/CO2 molar ratio of 1.5 displayed the highest conversion for most temperatures. For all ratios, the highest CO2 conversion was achieved at 850 ◦C, reaching around 98%. Overall, the conversion witnessed for CO2 was much higher than that of CH4; CO2 being the limiting reactant for ratios greater than 1. For the model biogas mixture (CH4/CO2 = 1) however, CH4 and CO2 conversions should be similar. The fact that CO2 conversion was slightly higher than CH4 conversion for temperatures lower than 850 ◦C was due to the occurrence of the reverse water gas shift (RWGS) reaction, consuming CO2 and H2 to form CO as previously reported elsewhere [5]. This can also be attributed to the high activation energy of methane [11,22].

The effect of varying the biogas feed concentration on the H2/CO ratio is illustrated in Figure 6a. As a general rule, the H2/CO ratio increased with higher temperatures, with the exception of the temperature screening carried out at a CH4/CO2 molar ratio of 1.85, where the H2/CO ratio initially decreased with the increment of the temperature and only increased once a temperature of 650 ◦C was reached. The high concentration of hydrogen in the products stream may be explained by the excess of methane in the feed gas. It is likely that this excess created favourable conditions for the methane decomposition reaction to take place, which contributed to a higher concentration (partial pressure) of H2 at lower temperatures. Furthermore, it was observed that at lower temperatures (550–600 ◦C), a larger CH4/CO2 ratio yielded a higher H2/CO ratio. Whereas, at higher temperatures (700–850 ◦C), the H2/CO ratio exhibited the largest value for a model biogas feed of 1:1 molar ratio of CH4/CO2.

The H2 yield obtained for each biogas composition as a function of temperature is shown in Figure 6b. As expected, yields increased along with temperature and CH4 conversion. As the CH4/CO2 molar ratio increased, H2 yields decreased due to excess CH4 in the feed and reduced CH4 conversion.

**Figure 6.** (**a**) H2/CO ratio and (**b**) H2 yield for Ni-Sn/CeAl at CH4/CO2 molar ratios of 1, 1.25, 1.5 and 1.85 as a function of temperature.

In order to establish a more general comparison with previously reported Ni-based catalysts for biogas reforming reactions, different reported catalysts and the conditions used in the biogas dry reforming are presented in Table 3. As shown in the table, our Ni-Sn/CeAl catalyst can be deemed to perform with superior behaviour in the compared conditions, especially in terms of hydrogen yields. Furthermore, our catalyst continues showing a good performance tested at lower temperatures and long-term tests, which reinforces the exceptional behaviour of the developed system.


**Table 3.** Catalysts employed, and conditions used in catalytic biogas dry reforming, using CH4/CO2 ratio: 1.5. WHSV: weight hourly space velocity.

#### *3.5. Stability Test*

Long-term stability tests were carried out at a temperature of 650 ◦C. The temperature was selected in order to simulate a syngas suitable for direct use i+n intermediate temperature SOFCs.

The results of the stability tests are displayed in Figure 7, which shows both CH4 and CO2 conversion as a function of time. Figure 7a compares the stability of Ni-Sn/CeAl under two biogas mixtures: 1 and 1.5 CH4/CO2 molar ratios. When tested with the model biogas mixture, the catalyst showed good performance with 79% CO2 conversion and 60% CH4 conversion. The performance was stable with a slight deactivation equivalent to 5% conversion loss. On the other hand, when the feed CH4/CO2 molar ratio was increased to 1.5, CH4 conversion was much lower, as previously observed, but the deactivation of the catalyst, although not drastic, was more pronounced than the one of the model mixture. The excess of methane in the feed stream seems to promote methane decomposition and therefore enhance the coking of the catalyst. This resulted in the loss of 10% and 15% conversion for CO2 and CH4, respectively.

**Figure 7.** Dry reforming of methane (DRM) stability test at 650 ◦C on (**a**) Ni-Sn/CeAl at CH4/CO2 molar ratios of 1 (model biogas) and 1.5 (sewage waste) and (**b**) Ni-Sn/CeAl and Ni/Al catalyst using CH4/CO2 molar ratios of 1.5: CO2 and CH4 conversions.

The long-term test using sewage waste biogas was compared to a reference Ni/Al2O3 catalyst on Figure 7b. The reference catalyst seems to perform slightly better than the promoted catalyst with slightly higher conversions. In fact, the introduction of Sn in the formulation of the catalyst was meant to prevent coke formation since Sn has a similar electronic configuration and was nucleating with Ni. Nevertheless, such an interaction may reduce the catalytic activity of the catalyst. Indeed, the deactivation of the reference catalysts was clearly more pronounced than that observed in the multicomponent catalyst, again highlighting the robustness of the Ni-Sn catalysts for long runs.

#### *3.6. Study of Carbonaceous Deposits*

Thermogravimetric analysis was conducted on the samples after reacting for 20 h to estimate the coke formation during reaction. Figure 8a shows the effect of biogas mixture on coke deposition. When subjected to the model biogas mixture, around 0.256 gC/gcat was formed whereas for sewage derived biogas, 0.268 gC/gcat was formed. The catalyst displayed faster deactivation under methane rich biogas, confirming that methane decomposition, as shown in Equation (3), must have been favoured, hence forming more coke on the catalyst. Figure 8b compares the carbon loading on two different catalysts after reacting with methane-rich biogas. The reference catalyst Ni/Al had similar catalytic performance to the Ni-Sn/CeAl catalyst, however, it was more prone to coking with 0.346 gC/gcat. The multicomponent catalyst benefits from the high oxygen storage capacity of ceria that facilitates carbon oxidation [5]. In addition, Sn atoms were occupying C nucleation sites in the vicinity of Ni, preventing carbon from poisoning the active sites [16]. Moreover, the gasification of the coke present on the Ni/Al catalyst occurred at higher temperatures, suggesting the formation of more graphitic carbon as previously observed in the literature [16].

**Figure 8.** Thermogravimetric analysis for the catalysts after (**a**) Ni-Sn/CeAl at CH4/CO2 molar ratios of 1 (model biogas) and 1.5 (sewage waste) and (**b**) Ni-Sn/CeAl and Ni/Al catalyst using CH4/CO2 molar ratios of 1.5.

#### *3.7. Post Reaction Characterisation*

Figure 9 shows the XRD profiles of the Ni-Sn/CeAl catalyst after reduction, after reacting from 550 to 850 ◦C and after the 20 h stability run at 650 ◦C in a biogas ratio of CH4/CO2 = 1.5. The reduced sample, as described earlier, presents the characteristic peaks of CeAlO3, Ni, CeO2 and Al2O3. After reacting at temperatures up to 850 ◦C, CeO2 seems to have completely reacted with Al2O3 to form CeAlO3 exclusively. The high temperatures and reducing atmosphere associated with the reaction conditions favoured the phase transition. After the stability test however, the CeO2 cubic phase displayed peaks of higher intensities than those of the reduced catalyst and the post temperature screening reaction catalyst. The low temperature of the stability test, 650 ◦C, seems to have favoured the reverse reaction and allowed CeAlO3 to re-oxidise back to CeO2. Indeed, the peaks associated with CeAlO3 decreased in intensity while the CeO2 diffraction peaks increased. Additionally, all post mortem samples presented a large peak at 26◦ associated to graphitic carbon. In terms of Ni crystallite size, Ni was estimated using the Scherrer equation at 17 nm, 20 nm and 11 nm in the sample reduced, after the temperature-dependant run and after the stability test, respectively. The temperature-dependant experiment was performed up to 850 ◦C, temperatures far above the Tammann temperature of Ni. Nickel sintering was prone to happen at such temperatures, explaining the increase in crystallite size. However, a reduction in the nickel crystallite size was observed after the 20 h stability test. First, the test took place at a much lower temperature (650 ◦C), which may have slowed down sintering. Second, the reaction atmosphere has an influence on the interactions amongst Ni, CeO2 and Al2O3. The re-oxidation of CeAlO3 to form CeO2 may have re-dispersed Ni crystallites on the catalyst surface. Zou et al. observed the same phenomenon on a Ni/ CeO2-Al2O3 catalyst reduced at 800 ◦C and after 105 h reaction at 350 ◦C. It appears Ni crystallites were re-constructed on the catalyst surface under the reaction atmosphere [27].

**Figure 9.** XRD characterisation spectra for Ni-Sn/CeAl reduced, post reaction and post stability at a CH4/CO2 molar ratio of 1.5.

#### **4. Conclusions**

In this work, a high performing Ni-Sn/CeO2-Al2O3 catalyst was developed for the conversion of biogas to syngas via the dry reforming of methane. The synthesised material was based on the DRM catalytic standard of Ni/Al2O3, which is an inexpensive alternative to high performing noble metal-based catalysts, and upgraded using CeO2 and Sn.

The performance of the multicomponent catalyst was investigated for a range of temperatures and model biogas compositions by determining the conversion of each reactant, as well as the H2/CO ratio of the syngas produced and the H2 yield. Additionally, multiple characterisation techniques were carried out on fresh, reduced and spent samples of catalyst, exhibiting the crystal structure changes and the reversible reducibility for the support.

Overall, the Ni-Sn/CeAl catalyst exhibited respectable conversions of both CO2 and CH4 ratios for all compositions of biogas. The effect of biogas composition on the stability of the catalyst was investigated, and although no extensive signs of deactivation were detected during 20 h, the presence of excess methane accelerated the deactivation of the catalyst due to additional methane decomposition. The stability of the catalyst was also compared with the standard Ni/Al2O3. Although both catalysts displayed similar catalytic activity, the multicomponent catalyst benefited from greater coke resistance, confirming the promotion effect of both tin and ceria.

In summary, our multicomponent catalyst is a suitable material—highly active and considerably stable—to be implemented in a biogas upgrading unit to generate renewable energy (i.e., using a SOFC) or added-value products using syngas as a platform chemical.

**Author Contributions:** Conceptualization, T.R.R. & B.A.H.; methodology, L.P.-P.; formal analysis, E.l.S.; investigation, S.J. & E.l.S.; writing—original draft preparation, S.J.; writing—review and editing, E.l.S. & T.R.R.; supervision, T.R.R. & B.A.H.; funding acquisition, T.R.R.

**Funding:** This research was funded by the Department of Chemical and Process Engineering at the University of Surrey and by the EPSRC, grant EP/R512904/1 as well as the Royal Society, Research Grant RSGR1180353. LPP acknowledge Comunitat Valenciana for her APOSTD2017 fellowship. This work was also partially sponsored by the CO2Chem through the EPSRC grant EP/P026435/1.

**Acknowledgments:** The authors acknowledge Sasol for supplying the alumina support.

**Conflicts of Interest:** The authors declare no conflict of interest. The funders had no role in the design of the study; in the collection, analyses, or interpretation of data; in the writing of the manuscript, or in the decision to publish the results.

#### **References**


© 2019 by the authors. Licensee MDPI, Basel, Switzerland. This article is an open access article distributed under the terms and conditions of the Creative Commons Attribution (CC BY) license (http://creativecommons.org/licenses/by/4.0/).

### *Article* **Catalytic Steam Reforming of Toluene: Understanding the Influence of the Main Reaction Parameters over a Reference Catalyst**

#### **Hua Lun Zhu, Laura Pastor-Pérez and Marcos Millan \***

Department of Chemical Engineering, Imperial College London, London SW7 2AZ, UK; h.zhu16@imperial.ac.uk (H.L.Z.); l.pastor-perez@imperial.ac.uk (L.P.-P.)

**\*** Correspondence: marcos.millan@imperial.ac.uk

Received: 31 December 2019; Accepted: 28 January 2020; Published: 13 February 2020

**Abstract:** Identifying the suitable reaction conditions is key to achieve high performance and economic efficiency in any catalytic process. In this study, the catalytic performance of a Ni/Al2O3 catalyst, a benchmark system—was investigated in steam reforming of toluene as a biomass gasification tar model compound to explore the effect of reforming temperature, steam to carbon (S/C) ratio and residence time on toluene conversion and gas products. An S/C molar ratio range from one to three and temperature range from 700 to 900 ◦C was selected according to thermodynamic equilibrium calculations, and gas hourly space velocity (GHSV) was varied from 30,600 to 122,400 h−<sup>1</sup> based on previous work. The results suggest that 800 ◦C, GHSV 61,200 h−<sup>1</sup> and S/C ratio 3 provide favourable operating conditions for steam reforming of toluene in order to get high toluene conversion and hydrogen productivity, achieving a toluene to gas conversion of 94% and H2 production of 13 mol/mol toluene.

**Keywords:** toluene; steam reforming; GHSV; S/C ratio; coke formation

#### **1. Introduction**

Greenhouse gas (GHG) emissions from fossil fuel combustion for power generation represent a major contribution to climate change. For this very reason, a switch from conventional to renewable power resources, i.e., solar, wind, hydroelectric energy and biomass is necessary [1].

Biomass can consistently provide energy and fuels and has an advantage over other renewable energies sources as it is more homogeneously distributed over the earth and is an abundant resource [2]. International Energy Outlook 2017 reported that biomass could provide over 14% of the world's primary energy consumption, which is the highest among renewable energy resource, and it will contribute a quarter or third of the global primary energy supply by 2050 [3].

For all the above and as a consequence of unstable oil prices and the alarming climate change, biomass gasification has increasingly received interest [4]. Indeed, this is a versatile and interesting way to re-use different wastes (e.g., agricultural and urban wastes, energy crops, food and industrial processing residues) to produce bio-syngas, which can be used for electrical power generation (fuel cells, gas turbine or engine), or as feedstock for the synthesis of liquid fuels and chemicals such as methanol [5]. Furthermore, the necessary technology for this process can be adapted from old coal gasification units [6]. However, one of the most critical technical challenges in biomass gasification is the formation of tars. Tar condensation can cause serious risks to downstream equipment. Therefore, tars should be removed from the effluent stream of biomass gasification [7].

Existing techniques for tar removal after a gasifier include separation either by physical (mechanical) methods, using ceramic candle filters or wet scrubbers, or thermochemical conversion methods using high temperature thermal or catalytic cracking to convert tar into syngas [8]. Physical separation methods would cause secondary pollution since they only remove tar from gas products, resulting in a waste stream that needs treatment. Conversely, thermal cracking has received increasing attention because tar can be converted into useful gas products and increases the overall efficiency of the gasification process [9]. Thermal cracking without catalysts operates at high temperature (>1000 ◦C) to decompose the tars in smaller non-condensable molecules. The high energy consumption makes this process less interesting. By contrast, catalytic cracking of tars can be carried out at lower temperatures converting tars into useful gases in a more efficient manner and is being widely studied as a principal method for tar removal [10]. As a steam reforming reaction, the proposed reaction could remove tar by a catalytic process and produce fuel H2 and CO at relatively low temperatures. At the same time, tar steam reforming poses some challenges that must be addressed related to the reaction conditions. Reforming catalysts can lose activity over time due to carbon deposition and sintering over the active phases [5]. These problems can be minimised with optimal operating conditions in the presence of the right catalyst. Real tar composition is highly complex and most studies use model tar compounds such as toluene, benzene or naphthalene to ascertain the catalytic mechanism [11–13]. However, previous studies have shown that these compounds represent a worst-case scenario in the tendency of the system to form carbon deposits [14,15].

Noble metals and Ni are most widely active phases used in reforming catalysts. Noble metal catalysts including Pt, Rh and Ru are known for their exceptionally good activity and stability in tar steam reforming. However, these catalysts have had limited use due to their high costs [16].

Nowadays, the aim is to develop an economically viable material, ideally not containing noble metals, which produces the same high levels of conversion and reaction performance as the noble ones. Ni is an attractive choice as steam reforming catalytic metal thanks to its good performance in the conversion of different types of hydrocarbons [16], being, for instance, the most popular active phase in methane steam reforming [17,18]. In particular, Ni/Al2O3 catalysts are considered as the state-of-the-art materials for steam reforming processes. Furthermore, the high surface area of alumina and its mechanical properties result in an excellent choice as support for nickel nanoparticles [19].

Under these premises, this work showcases the application of a Ni/Al2O3 catalyst in the steam reforming of toluene (C7H8) as a tar model compound in a fixed bed reactor. Until now, the catalytic performance in the steam reforming of toluene has been mainly evaluated as a function of catalyst design variables, such as the nature of the support [20–22] and metal [23,24], but little attention has been paid to the reaction conditions, especially for this benchmark catalytic formulation [25]. Identification and optimisation of the reaction parameters in the presence of a commercial-like catalyst (Ni/Al2O3) are vital to achieving the best catalytic performance. However, few studies involving parameter screening have been carried out [8].

Herein we analyse the influence of reforming temperature, steam to carbon molar ratio (S/C) and gas hourly space velocity (GHSV) on the toluene reforming performance. These are considered the key parameters to fine-tune the reaction and maximise the overall performance.

Parallel reactions in this system can be numerous. The main reactions that can occur during toluene steam reforming are represented as follows:

Toluene steam reforming

$$\rm{C}\_{7}H\_{8} + 14 \, H\_{2}O \to 7CO\_{2} + 18 \, H\_{2} \tag{1}$$

$$\text{C}\_7\text{H}\_8 + 7\text{ H}\_2\text{O} \to 7\text{CO} + 11\text{ H}\_2\tag{2}$$

Water–gas shift

$$\rm{CO} + \rm{H}\_2\rm{O} \leftrightarrow \rm{CO}\_2 + \rm{H}\_2\tag{3}$$

Hydrodealkylation

$$\rm C\_7H\_8 + H\_2 \to C\_6H\_6 + CH\_4 \tag{4}$$

Methane steam reforming

$$\text{CH}\_4 + \text{H}\_2\text{O} \leftrightarrow \text{CO} + 3\text{H}\_2\tag{5}$$

Boudouard reaction

$$2\text{CO} \leftrightarrow \text{CO}\_2 + \text{C} \tag{6}$$

Steam reforming of toluene is irreversible and Reactions (1) and (2) are dependent on the S/C ratio used. CH4 is produced from hydroalkylation (Reaction (4)) and pyrolysis of toluene. Methane reforming followed by a water–gas shift reaction converts the produced CO with steam to H2 and CO2. A Boudouard reaction is an exothermic reaction that produces C from CO. All of these reactions are heavily conditioned by the temperature, space velocity and reactants ratio [26–28]. Hence, a careful assessment of the impact of these parameters will allow us to identify the optimum conditions towards the generation of added value products.

For all the above, this work fills an essential gap in tar reforming literature via the systematic study of key reaction parameters using a state-of-the-art catalyst to reveal the optimum process conditions.

#### **2. Experimental**

#### *2.1. Catalyst Preparation*

The nickel-based catalyst was prepared following a wet impregnation method. The necessary amount of Ni(NO3)2·6H2O (≥97.0%, Sigma-Aldrich) to obtain 20 wt.% NiO was dissolved in an excess of acetone (≥99.8%, Sigma Aldrich). Then, the support γ- Al2O3 (≥98.0% purity, Sasol) was added into the solution and, after stirring for 2 h, the solvent was removed under vacuum at 60 ◦C by using a rotating evaporator. The remaining mixture was dried overnight at 110 ◦C. Finally, the solid was calcined at 600 ◦C with a ramping rate of 2 ◦C·min−<sup>1</sup> for 4 h. It has been reported that Ni/Al2O3 catalysts are stable under reaction conditions despite the calcination temperature being lower than those of the experiments [29,30]. Lower calcination temperatures have been shown to lead to better catalytic performance in steam reforming [31]. The obtained sample was labelled Ni/Al2O3. The Ni content assuming complete reduction from NiO to Ni is 16.4 wt.%.

#### *2.2. Characterisation*

Thermogravimetric analysis (TGA) was carried out to investigate the coke deposition on the catalyst in a Pyris 1 thermogravimetric analyser from PerkinElmer (Waltham, MA, USA). The samples were ramped from room temperature to 900 ◦C at a rate of 10 ◦C·min−<sup>1</sup> in air.

N2-adsorption-desorption analysis was conducted in a TriStar 3000 V6.07 A analyser from Micromeritics (Norcross, GA, USA). Before the analysis, the catalyst was degassed at 150 ◦C for 4 h in a vacuum. The Brunauer–Emmett–Teller (BET) method was used to calculate the surface area of the catalyst.

#### *2.3. Catalytic Toluene Steam Reforming Tests*

Toluene steam reforming was carried out in a fixed bed reactor used in previous bio-oil reforming work [24]. Before the reaction, the reactor was purged with N2 to remove air. The catalyst was reduced under 50 mL·min−<sup>1</sup> of H2 up to 700 ◦C for 1 h before each test.

Figure 1 shows the schematic diagram of the experimental reaction set-up, and the reaction zone is shown in Figure 2. The reactor was heated up by two copper electrodes; toluene and steam were injected by two syringe pumps from the top of the reactor and preheated at 200 ◦C to the vapour phase in a preheating chamber. Toluene was carried by N2 with a fixed concentration of 100 g Nm<sup>−</sup>3. Then the reactant stream entered an incoloy alloy 625 tube (12 mm i.d., 2 mm thick, 253 mm long), equipped with an inner quartz tube (9 mm i.d., 1 mm thick, 300 mm long) to prevent any contact between the reactant gas stream and the incoloy internal surface. 500 mg of Ni/Al2O3 catalyst with a

particle size in the range of 250–500 μm was placed right in the middle of the quartz tube. A K-type thermocouple was used to determine the catalytic bed temperature.

**Figure 1.** Schematic diagram of the catalytic toluene steam reforming system.

The product gases after reaction pass through two condensers in series to collect any liquid product as well as unreacted toluene and water. Ice and dry ice were used as coolant in the two condensers, respectively. The products identified in the gas phase were H2, CH4, CO2 and CO. Two on-line gas analysers were used to determine the product gas compositions: an MGA3000 Multi-Gas infrared analyser (ADC Gas Analysis, Herts, UK) for CO2, CH4 and CO, followed by a K1550 thermal conductivity H2 analyser (Hitech Instruments, Luton, UK).

The performance of catalysts was evaluated by the conversion into gaseous products (based on a carbon balance between the inlet and the outlet stream of the reactor), selectivity to main products (where "i" is CO2, CO and CH4 in moles) and hydrogen yield, which were defined as follows:

$$\% \text{ Carbon Concentration} = \frac{\text{C in the gas product}}{\text{C fed into reactor}} \text{} \text{100} \tag{7}$$

% "i" selectivity <sup>=</sup> "i" produced in moles C atoms in the gas products <sup>∗</sup> 100 (8)

$$\text{Hydrogen yield} = \frac{\text{total Hydrogen production in moles}}{\text{toluene fed into reactor in moles}} \tag{9}$$

The experimental error in toluene conversion, gas selectivity and gas yield is ±2%. Toluene conversion and H2 production could be influenced by experiment conditions and parameters. Reforming temperature, S/C ratio and residence time are reported to be the key factors that would affect the total conversion and H2 yield. In this paper S/C ratios of 1, 2 and 3; temperatures of 700, 800 and 900 ◦C, and GHSV of 30,600, 61,200, 91,800 and 122,400 h−<sup>1</sup> were investigated.

**Figure 2.** Reactor diagram of catalytic toluene steam reforming.

#### *2.4. Thermodynamic Simulation*

The ASPEN software package (AspenTech, Bedford, MA, USA) was used to determine the thermodynamic equilibrium of the toluene reforming reactions over the different reaction conditions. An ideal property method, an RGIBBS reactor (based on Gibbs free energy minimisation) was selected to investigate the thermodynamic equilibrium. Material flows into the reactor are identical to those from the corresponding experiment. The influence and effects of experimental parameters, including reforming the temperature and S/C ratio on the toluene conversion, the yield of main light gases and the carbon deposition was investigated.

#### **3. Results and Discussion**

#### *3.1. Textural Properties of the Synthetised Catalyst*

The N2 adsorption-desorption isotherm of the calcined Ni/Al2O3 catalyst is shown in Figure 3, which shows a type IV isotherm with a characteristic hysteresis loop for mesoporous materials. The adsorption average pore width was 9.1 nm, the total pore volume was 0.35 cm3·g−<sup>1</sup> and BET Surface Area was 153 m2·g<sup>−</sup>1. The BET surface area, pore volume and pore width of the original Al2O3 was 230 m2·g<sup>−</sup>1, 0.5 cm3·g−<sup>1</sup> and 10 nm respectively. Textural properties are actually governed by the primary support gamma-alumina, which provides mechanical and thermal stability as well as high surface area.

**Figure 3.** Toluene conversion to C-containing gases and CO/CO2 selectivity at different temperatures (S/C:3, gas hourly space velocity (GHSV):61,200 h<sup>−</sup>1).

#### *3.2. Influence of Temperature on Toluene Steam Reforming*

The steam reforming temperature was concluded to have a significant influence on toluene conversion since higher temperatures could increase syngas production and conversion from toluene to gas products [32]. When reforming is envisaged as a downstream tar upgrading unit after gasification, the most interesting temperature interval for atmospheric reforming is between 600 and 900 ◦C, since the gasification effluent temperature will normally be lower than 900 ◦C [33]. Research also suggested that high temperature might reduce H2 yield as the reverse water–gas shift reaction is favoured due to its endothermic nature [34–36]. The experiments were conducted at different temperatures to investigate the most suitable conditions for toluene conversion and H2 production. Catalytic performance tests were performed at 700, 800 and 900 ◦C, with a fixed S/C ratio of 3 and a GHSV of 61,200 h−<sup>1</sup> (corresponding to a N2 flow rate of 300 mL·min<sup>−</sup>1) for 5 h.

Figure 3 shows the toluene conversion and CO2/CO selectivity (based on conversion to C-containing gases) at the three different temperatures in a steady state within the five hours of reaction and compares with equilibrium selectivity. CO and CO2 were the main gases. The experimental selectivity of CH4 at all temperatures was under 0.1%. Conversion and selectivity of the gases approached that of thermodynamic equilibrium as temperature increased. Thermodynamic equilibrium predicts total toluene conversion at 700 ◦C or higher temperatures. At 700 ◦C, the thermodynamic equilibrium indicated a CH4 selectivity of 1.1%, and for higher temperatures only CO and CO2 were predicted as C-containing gas products. Experimental toluene conversions ranged from 84% to 96% and approached equilibrium with the temperature increase from 700 to 900 ◦C. Coke formation only accounted for less than 1% of toluene conversion, the rest being to gas phase products. Unreacted toluene was condensed in the cooling trap, which enabled closing the mass balance within 98%. The high conversions of toluene achieved for the three temperatures highlight once again the good performance of this commercial-like catalyst (Ni/Al2O3). Both experimental and equilibrium selectivity showed that the CO/CO2 ratio increases as temperature increases. This is likely dominated by a reverse water–gas shift reaction in the high-temperature area (Reaction (3), as the increasing temperature would favour the endothermic direction, and nickel contents would also promote the reaction at high temperatures [11].

Table 1 shows the gaseous product yields including CO, CO2, H2 and CH4. The yields of CO and CO2 have similar trends to selectivity. However, for H2 yield the maximum was achieved at 800 ◦C with a H2 production of 13.0 mol/mol toluene, while the highest H2 production in equilibrium conditions was predicted at 700 ◦C. This can be due to, in the experimental test at 700 ◦C, the lowest toluene conversion into gases was obtained. Toluene conversion to gas stayed over 94% at 800 ◦C or above, while a higher temperature would lead to a slight decrease in the content of H2 in gaseous products due to the presence of the reverse water–gas shift reaction as discussed above. It is interesting to remark that only at 700 ◦C the undesirable CH4 side product was obtained and it was only in a small amount. The absence of CH4 at higher temperatures can be due to the methane reforming to CO and other parallel processes consuming methane as suggested by the equilibrium results.

**Table 1.** Product yields for the gaseous products at the three different temperatures (S/C:3, GHSV: 61,200 h<sup>−</sup>1).


Thermogravimetric analysis was conducted on the catalysts after a five-hour reaction to estimate the coke formation during the three different temperatures reaction tests. It has been reported that coke deposition on Ni/Al2O3 catalysts is the main cause for deactivation and high nickel contents could also favour coke formation. Although coke deposition is thermodynamically unfavourable at high temperatures (>600 ◦C), methane decomposition in the high temperature range could lead to the production of solid carbon [37]. Table 2 shows the carbon conversion from toluene to coke and the fraction of coke on the catalyst as a function of temperature. As the temperature increased, the conversion to carbon deposits was slightly higher. It is likely that carbon at the higher temperatures was formed from the reforming of CH4 observed at lower temperatures, as CH4 is known to favour coke formation on Ni-based catalysts [27]. Notwithstanding coke formation at these temperatures, the amount of coke is very low, and the catalyst remains stable during the whole experiment.

**Table 2.** Toluene conversion to coke and fraction of coke deposited on the catalyst at different reforming temperatures (S/C:3, GHSV:61,200 h<sup>−</sup>1, 5-h test).


Based on the results presented above, 800 ◦C was considered the most suitable reforming temperature due to the highest H2 production, an excellent overall conversion (94%) and acceptable coke deposition levels. On the contrary, despite the 900 ◦C experiment showing a slightly higher toluene conversion, lower H2 production, greater coke deposition and lower energy efficiency made this temperature not preferable with the 800 ◦C test.

Gas analysis as a function of time on stream at the chosen temperature is shown in Figure 4 for a five-hour test. There was no CH4 detected throughout the test and CO, CO2 and H2 concentration stayed stable *ca* 11%, 23% and 66% (figures corrected from N2 dilution), respectively, during the whole experiment. Furthermore, no obvious change in the CO/CO2 ratio, or drop in H2 yields and catalyst deactivation were observed in this test.

**Figure 4.** Gas product concentration at 800 ◦C, S/C:3, GHSV:61,200 h−<sup>1</sup> in a 5-h test.

#### *3.3. Influence of GHSV in Toluene Steam Reforming*

In view of the results obtained in the temperature screening, a temperature of 800 ◦C was chosen to study the effect of GHSV to elucidate the suitable residence time for a complete conversion of toluene. High GHSV, or lower residence time, might inhibit toluene conversion and H2 production. Literature suggests that suitable GHSVs for steam reforming were mostly between 10,000 and 70,000 h−<sup>1</sup> [38,39] over different catalysts. For our benchmark catalyst, tests were carried out with 500 mg of catalyst, and a fixed temperature (800 ◦C) and S/C ratio of three during 5 h of reaction. The feeding rates of toluene, steam and N2 were changed according to the desired GHSV. The GHSV studied were 30,600, 61,200, 91,800 and 122,400 h<sup>−</sup>1. Some of the space velocities selected in this study are indeed over the standard ranges mentioned above. In fact, high space velocities normally involve lower reactor volumes, thus decreasing the capital cost of the reformer.

Figure 5 shows the toluene conversion and CO2/CO selectivity (based on conversion to C-containing gases) at the four different GHSVs in a steady state within the five hours of reaction and compares them with equilibrium selectivity. Toluene conversion decreased from 96% to 86% when GHSVs increased from 30,600 to 122,400 h<sup>−</sup>1. When GHSV was 91,800 h−<sup>1</sup> or lower, toluene conversion to gas remained higher than 90%, without notorious differences for the two lowest GHSV studied (30,600 h−<sup>1</sup> or 61,200 h<sup>−</sup>1), which showed a stabilised conversion at around 95%. The selectivity of CO and CO2 approached the equilibrium results slowly when GHSV decreased. It could be seen that toluene conversion and CO2 selectivity declined with the increasing of GHSV, as a result of shorter residence time and reaction period. It is interesting to note that the selectivity is less affected by the residence time than the conversion. Negligible or no methane was detected in the product gas, which was in line with literature that suggested that at temperatures higher than 750 ◦C very small amounts of CH4 are produced [11].

**Figure 5.** Toluene conversion to C-containing gases and CO/CO2 selectivity at different GHSVs (S/C:3, 800 ◦C).

The gas product CO2, CO and H2 yields are shown in Table 3. As expected, GHSV 30,600 h−<sup>1</sup> led to the highest H2 production (13.2 mol/mol toluene). As a trend, CO, CO2 and H2 yields dropped when GHSV increased. This decrease in yields for all gas products at high GHSV is linked with the drop in conversion as the residence time became shorter. An exception occurred at 61,200 h−<sup>1</sup> GHSV as the reverse water–gas shift reaction could cause a small increase in CO. Previous studies suggest that some tar model compounds (naphthalene) had no apparent trends for the hydrogen yields or selectivity, as the product was affected by the equilibrium of the water–gas shift reaction and other side reactions [40]. In this work, hydrogen yield and total conversion showed a slightly increasing trend in conversion and yields with the decrease of residence time, but this trend was more remarkable in the conversion of toluene than in selectivities.

**Table 3.** Product yields for the gaseous products at different GHSV (S/C:3, 800 ◦C).


Table 4 shows the carbon conversion from toluene to coke with different GHSVs. This data led us to a better understanding of the coke deposition. The reactant toluene and steam feeding rate increased three times when GHSV increased from 30,600 to 122,400 h<sup>−</sup>1; however, the coke conversion reduced from 0.38% to 0.22%. This means that the higher GHSV used resulted in a higher coke amount deposited but lower conversion rate. This result is the balance between the higher feed of toluene used and the decrease in conversion due to the high GHSV used.

**Table 4.** Toluene conversion to coke deposited on the catalyst with different GHSV (S/C:3, 800 ◦C, five-hour test).


Although the lower space velocity leads to higher conversion, greater selectivity and the lower net amount of carbon deposits, for practical applications (i.e. manufacturing cost savings) higher space velocities are desired. In this sense, 61,200 h−<sup>1</sup> yields very similar conversion and selectivity levels, low net coking and lower conversion to coke. Hence, this space velocity was selected to optimise the next reaction parameter.

#### *3.4. Influence of S*/*C Ratio in Toluene Steam Reforming*

Steam, as a principal reactant in catalytic steam reforming, is widely recognised to have a strong influence on H2 production. Steam is involved in most of the relevant reactions in the toluene reforming and therefore the steam to carbon ratio was chosen as a key variable to study. A high steam partial pressure improves gasification reactions and moves the water–gas shift equilibrium towards hydrogen production, while the partial pressure of toluene in the gas stream is lower due to dilution as the S/C ratio rises [37]. Suitable S/C ratios to investigate the catalyst performance were mostly between one and four in the literature [16,41]. Although steam is the main reactant in a reforming process, a large excess of water in these experiments could condense into ice in the cooling trap and block the system. It is also reported that the saturation of the catalyst surface by steam at a high S/C ratio would not favour the conversion of toluene or H2 production [5].

In this study, the catalytic performance tests were performed at three different S/C ratios of 1, 2, 3, at 800 ◦C, 500 mg of samples and a GHSV of 61,200 h−<sup>1</sup> for five hours.

Figure 6 shows the influence of the S/C ratio on the selectivity of carbon containing products. Both experimental and equilibrium results showed that an increment in S/C ratio results in a large improvement of CO2 selectivity and toluene conversion. In the experimental tests, toluene to gas conversion increased from 53% to 94% when the S/C ratio increased from one to three. Equilibrium results indicated that a higher S/C ratio always increases the H2 production, as the excess steam would promote the water–gas shift reaction. A higher S/C ratio also increased toluene conversion to gas. Some researchers suggested that the most suitable S/C ratio was mostly between 2.5 and 3.5, because more excess steam would not increase toluene conversion or H2 production, causing a drop in toluene partial pressure [26].

Table 5 compares the product yields of CO, CO2 and H2 with different S/C ratios. Equilibrium simulation and experiments both showed large increases in CO2 and H2 production with the S/C ratio. Experimental H2 yield increased from 5.1 to 13.0 mol/mol toluene when the S/C ratio raised from one to three. The opposite equilibrium trend is expected for CO, but this did not happen in the experiments. Instead, the CO yield was observed to increase with the amount of steam as the trend was dominated by the higher conversion achieved at higher S/C ratios. Despite the experimental CO yield showing a slight increase, in general the CO/CO2 ratio decreased with the increase of the S/C ratio in line with the equilibrium results.

**Figure 6.** Toluene conversion to C-containing gases and CO/CO2 selectivity at different S/C ratios (GHSV:61,200 h<sup>−</sup>1, 800 ◦C).



The coke weight on the spent catalyst with S/C ratio 1, 2, 3 is 0.289, 0.222, 0.167 gC/gcat, respectively. As expected, the higher S/C ratio promoted carbon gasification avoiding coke formation [42]. Table 6 shows the carbon conversion from toluene to coke for the three different S/C ratio studied. As discussed above, it can be seen that the excess of steam inhibited, but did not suppress, the formation of coke. For all the above, the better choice is to use a S/C ratios of three since it permits the highest H2 production and toluene conversion and the lowest coke amount.

**Table 6.** Toluene conversion to coke deposited on the catalyst at different S/C ratios.


#### **4. Conclusions**

To remove tar produced from biomass gasification, catalytic steam reforming was conducted for toluene as a model tar compound. Simulations of a thermodynamic equilibrium based on Gibbs free energy minimisation and experiments in a fixed bed reactor using a Ni/Al2O3 catalyst were carried out. The effect of reforming temperature, S/C ratio and GHSV on toluene conversion and product distribution was studied.

Increasing the temperature from 700 to 900 ◦C increased total conversion, with a potential risk of higher coke deposition. A temperature of 800 ◦C observed the highest H2 production, high toluene conversion (>94%) and relatively lower coke deposition.

The Ni/Al2O3 catalyst only requires a very short residence time (GHSV < 91,800 h<sup>−</sup>1) for toluene reforming with this catalyst and effectively removes low density toluene in a mixed gas stream. H2 yield and toluene conversion increased slightly and approached simulated thermodynamic equilibrium results when GHSV decreased. Coke deposition increased at a lower rate as GHSV increased.

A high S/C ratio would greatly increase total conversion, hydrogen production and reduce the coke formation on the catalyst. The presence of excess steam could shift the equilibrium of the water–gas shift reaction to produce more H2.

A temperature of 800 ◦C, GHSV of 61,200 h−<sup>1</sup> and S/C ratio of three provided the most suitable reaction conditions for toluene conversion and H2 production in steam reforming of toluene, obtained a steady state of a toluene to gas conversion over 94%, a H2 production of 141.6 mol/mol toluene in a five-hour test, with no obvious deactivation observed in five hours. Based on these results, this condition would be suitable for tar model compound removal.

**Author Contributions:** Conceptualization, M.M. and H.L.Z.; experiments, H.L.Z.; data curation and analysis, H.L.Z. and L.P.-P.; writing—original draft preparation, H.L.Z.; writing—review and editing, M.M. and L.P.-P.; supervision, M.M. and L.P.-P.; All authors have read and agreed to the published version of the manuscript.

**Funding:** This research received no external funding.

**Conflicts of Interest:** The authors declare no conflict of interest.

#### **References**


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