**1. Introduction**

The high energy efficiency of fuel cells has drawn considerable attention toward the development of hydrogen production technologies. Hydrogen can be produced from fossil fuels either via hydrocarbon pyrolysis or hydrocarbon reforming processes including steam reforming, partial oxidation or autothermal reforming [1–5]. Biomass processes consisting of biological (bio-photolysis, dark fermentation, photo fermentation) and thermochemical (pyrolysis, gasification, combustion, liquefaction) methods as well as water splitting processes such as electrolysis, thermolysis or photolysis can be alternatively applied for the production of H2 from renewable energy sources [2,6]. However, the latter approaches are facing some major obstacles mostly related to high cost and low H2 yields. Currently, steam reforming of light hydrocarbons, including natural gas, ethane, propane, butane and liquified petroleum gas (LPG), are considered among the most promising and economical routes for hydrogen production [7]. Propane, which is the main component of LPG, has many advantages such as high energy density, compressibility to a transportable liquid at normal temperature and well-developed infrastructure which enable its use worldwide [8–10]. Moreover, propane can be stored and transferred as LPG through a wide distribution network or in high pressure cylinders in order to be supplied in remote places (e.g., agricultural, inaccessible or camping areas) or for domestic uses (e.g., households) [3,7].

Under propane steam reforming conditions, the water-gas shift reaction occurs simultaneously at low temperatures contributing to H2 and CO2 production, whereas CO/CO2 methanation may also run in parallel yielding CH4 and H2O. Methane can be also formed via hydrogenation of CH*x* species derived by the dissociative adsorption of propane on the catalyst surface or through propane decomposition accompanied by ethylene production. In certain cases, the C2H4, CH4 and CO thus produced are further decomposed

**Citation:** Kokka, A.; Petala, A.; Panagiotopoulou, P. Support Effects on the Activity of Ni Catalysts for the Propane Steam Reforming Reaction. *Nanomaterials* **2021**, *11*, 1948. https:// doi.org/10.3390/nano11081948

Academic Editors: Nikolaos Dimitratos and Alberto Villa

Received: 30 June 2021 Accepted: 27 July 2021 Published: 28 July 2021

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**Copyright:** © 2021 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 (https:// creativecommons.org/licenses/by/ 4.0/).

leading to the formation of coke on the catalyst surface and consequently, to its progressive deactivation [11,12].

Supported noble metal (Rh, Ru, Pt, Pd) catalysts have been proposed to efficiently catalyze the production of H2 via propane steam reforming exhibiting high resistance to coke formation [3,11,13,14]. However, the high cost and low availability of noble metals are major drawbacks restricting their use in practical applications [15–17]. Noble metals can be replaced by nickel, which is less expensive and able to convert propane with high H2 yields. However, Ni is susceptible to coke formation and particles sintering considered to be responsible for catalyst deactivation [5,10,17]. The lifetime of Ni-based catalysts can be improved by optimization of the reaction conditions, catalyst promotion, improvement of the catalyst synthesis method as well as by the proper selection of the support [3,9,10,12,16,18,19].

Regarding the support material, it has been proposed that metal oxides characterized by high oxygen storage capacity, such as CeO2, YSZ, TiO2, ZrO2 or CeO2-ZrO2, are able to suppress carbon deposition through the participation of their lattice oxygen in carbon removal [3,16]. Moreover, Ni catalysts supported on metal oxides or promoted metal oxides, which favor steam adsorption and mobility of surface hydroxyls, have been found to facilitate coke gasification [17,20].

Metal-support interactions have been also reported to impose a dramatic effect on both carbon deposition and metal particles sintering under conditions of hydrocarbons reforming [5,21,22]. It was demonstrated that stronger interactions between Ni and MgAl2O4 support resulted in high Ni dispersion and inhibition of the formation of large Ni clusters [22], whereas weak interactions between Ni and SiO2 carrier were found to accelerate sintering and coke formation [21].

Therefore, the economic viability and practical applicability of H2 production via propane steam reforming may be facilitated by the development of efficient Ni catalysts supported on suitable metal oxides, which will be able to possess both high activity and resistance against carbon deposition and particles sintering to realize economically viable reforming processes. In the present study the effect of the nature of the support (TiO2, CeO2, Al2O3, YSZ, ZrO2) on the activity and selectivity of Ni-based catalysts for the propane steam reforming reaction was investigated. Mechanistic aspects related to the support influence on the reaction pathway were also studied and are discussed.

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

#### *2.1. Catalyst Preparation and Characterization*

The wet impregnation method was applied to prepare Ni (5 wt.%) catalysts supported on commercial metal oxide powders by using an aqueous solution of Ni (Ni(NO3)2·6H2O) as the metal precursor salt. The commercial metal oxide carriers were used as received and were (a) activated aluminum oxide (Al2O3), catalyst support, 99% (metals basis) (Alfa Aesar, Kandel, Germany), (b) AEROXIDE® TiO2 P25 (TiO2) (Evonik Industries AG, Essen, Germany), (c) cerium (IV) oxide (CeO2), nanopowder, 99.5% min (Alfa Aesar, Kandel, Germany), (d) yttria-stabilized zirconia (YSZ) (8Y-SZ, Tosoh, Amsterdam, The Netherlands) and (e) zirconium (IV) oxide (ZrO2) 99% (metal basis) (Alfa Aesar, Kandel, Germany). The resulting materials were dried at 110 ◦C for 24 h followed by reduction at 400 ◦C under H2 flow for 2 h. The selection of reducing Ni catalysts at 400 ◦C for 2 h was based on previous studies which indicated that under these reducing conditions nickel oxide species are able to be completely converted to metallic nickel [23–25].

The X-ray diffraction patterns of the catalysts were recorded using an X-ray powder diffractometer (A Brucker D8 Advance instrument, Bruker, Karlsruhe, Germany) using Cu *Ka* radiation ( *λ* = 0.15406 nm, 40 kV, 40 mA) and a scan rate of 0.025◦/s over a range of 2*θ* between 20 and 80◦. The diffraction pattern was identified by comparison with those supplied from the JCPDS data base, whereas the primary crystallite size of MxOy (dMxOy) was estimated according to Scherrer's equation:

$$\mathbf{d}\_{\mathrm{M\_xO\_\gamma}} = \frac{0.9 \cdot \lambda}{\mathrm{B} \cdot \cos \theta} \tag{1}$$

where θ is the angle of diffraction corresponding to the peak broadening, B is the fullwidth at half maximum intensity (in radians) and λ = 0.15406 nm is the X-ray wavelength corresponding to Cu *Ka* radiation.

The specific surface area (SSA) of the supported Ni catalysts were measured by N2 adsorption at 77 K (B.E.T. technique) using a Gemini III 2375 instrument (Micromeritics, Norcross, GA, USA). Carbon monoxide chemisorption measurements at 25 ◦C were applied for the determination of Ni dispersion and mean particle size using a modified Sorptomatic 1900 apparatus (Fisons Instruments, Glaskow, UK) and assuming a CO:Me stoichiometry of 1:1, an atomic surface area of 6.5 Å2 and spherical particles. CO chemisorption measurements were used instead of H2 chemisorption in order to avoid overestimation of Ni dispersion due to hydrogen spillover effects, which have been previously found to occur over supported Ni catalysts [26,27]. Nickel particle size was calculated according to the following equation:

$$\mathbf{d}\_{\rm Ni} = \frac{60000}{\rho\_{\rm Ni} \cdot \mathbf{S}\_{\rm Ni}} \,\mathrm{[\,\dot{\mathbf{A}}\,]}\tag{2}$$

where dNi is the mean crystallite diameter, ρNi (= 8.9 <sup>g</sup>·cm<sup>−</sup>3) is the density of Ni and SNi [m2/gNi] is the surface area per gram of Ni.

Transmission electron microscopy (TEM) images were obtained with a JEM-2100 system (JEOL, Akishima, Tokyo, Japan) operated at 200 kV (point resolution 0.23 nm) using an Erlangshen CCD Camera (Model 782 ES500W, Gatan Inc., Pleasanton, CA, USA). Samples were dispersed in water and spread onto a carbon-coated copper grid (200 mesh). Details related to the equipment and procedures used for catalyst characterization have been described in detail elsewhere [28].

#### *2.2. Catalytic Performance Tests and Kinetic Measurements*

The catalytic performance of the synthesized materials was studied in a tubular fixed-bed quartz reactor under atmospheric pressure using an apparatus which has been described in detail elsewhere [11]. The reaction conditions were as follows: temperature range 400–750 ◦C, H2O*/*C = 3.25, and gas hourly space velocity (GHSV) = 55,900 h−1. The reactor was loaded with 150 mg of catalyst (particle diameter: 0.15 < dp < 0.25 mm) and placed in an electric furnace, where it was reduced in situ at 300 ◦C for 1 h under 50%H2/He flow (60 cm<sup>3</sup> min−1) to ensure that the Ni exists in its metallic phase prior to catalytic performance tests. Catalyst reduction was followed by heating at 750 ◦C under He and subsequent switch of the flow to the feed stream consisted of 4.5%C3H8 + 0.15%Ar + 44%H2O (He balance). Argon was used as internal standard in order to account for the volume change. Water was fed through an HPLC pump (LD Class Pump, TELEDYNE SSI, PA, USA) into a vaporizer maintained at 180 ◦C and mixed with the gas stream coming from mass-flow controllers. A condenser immersed in an ice bath was placed at the exit of the reactor to condensate water prior to introduction of the gas stream to the analysis system. Reaction gases (He, 30%C3H8−1%Ar/He, H2) are supplied from high-pressure gas cylinders (Buse Gas, Bad Hönningen, Germany) and are of ultrahigh purity. Measurements of reactants' and products' concentrations were obtained by stepwise decreasing temperature up to 400 ◦C. The effluent from the reactor was analyzed using two gas chromatographs (Shimadzu, Kyoto, Japan) which were connected in parallel. The procedure used for gas phase analysis was described in our previous study [11]. The conversion of propane (XC3H8 ) was calculated using the following expression:

$$\chi\_{\text{C}\_{3}\text{H}\_{8}} = \frac{[\text{Carbon}]\_{\text{total,out}}}{[\text{Carbon}]\_{\text{total,out}} + [\text{C}\_{3}\text{H}\_{8}]\_{\text{out}}} \times 100\tag{3}$$

where [Carbon]total,out is the sum of the concentrations of all carbon containing products:

$$\left[ \text{Carbon} \right]\_{\text{total,out}} = \frac{\left[ \text{CO} \right] + \left[ \text{CO}\_2 \right] + \left[ \text{CH}\_4 \right]}{3} + 2 \times \frac{\left[ \text{C}\_2 \text{H}\_4 \right] + \left[ \text{C}\_2 \text{H}\_6 \right]}{3} \tag{4}$$

Selectivity toward reaction products containing carbon was defined using Equation (5). The factor *n* corresponds to the number of carbon atoms in the corresponding molecule (e.g., for CO is 1, for C2H4 is 2 etc.):

$$\mathbf{S\_{Cn}} = \frac{[\mathbf{C\_n}] \times \mathbf{n}}{3 \times [\mathbf{C\_{carbon}}]\_{\text{total,out}}} \times 100\tag{5}$$

Selectivity toward hydrogen production was defined as the concentration of hydrogen produced divided with the concentration of all products containing hydrogen according to Equation (6). The factor m represents the number of hydrogen atoms in the corresponding molecule (e.g., for CH4 and C2H4 is 4).

$$\text{S}\_{\text{H}\_2}(\%) = \frac{[\text{H}\_2]}{[\text{H}\_2] + \text{m}/2 \times [\text{C}\_\text{n}\text{H}\_\text{m}]} \times 100\tag{6}$$

The intrinsic reaction rates for propane steam reforming reaction were measured for low propane conversions (XC3H8< 10%) by varying W/F using the following expression:

$$\mathcal{R}\_{\mathbb{C}\_3\text{H}\_8} = \frac{[\mathbb{C}\_3\text{H}\_8]\_{\text{in}} \cdot \text{F}\_{\text{in}} - [\mathbb{C}\_3\text{H}\_8]\_{\text{out}} \cdot \text{F}\_{\text{out}}}{\mathcal{W}} \times 100\tag{7}$$

where RC3H8 is the molar rate of C3H8 consumption (mol s<sup>−</sup><sup>1</sup> gcat<sup>−</sup>1), [C3H8]in, [C3H8]out, are the inlet and outlet concentrations (*v*/*v*) of C3H8, respectively, Fin and Fout are the total flow rates in the inlet and outlet of the reactor (mols−1), respectively, and W is the mass of catalyst (gcat).

Turnover frequencies (TOFs) of propane conversion were estimated following Equation (8) taking into account the measurements of both the reaction rates and nickel dispersions:

$$\text{TOF} = \frac{\mathbb{R}\_{\text{C}\_{\text{S}}\text{H}\_{8}} \cdot \text{AW}\_{\text{Ni}}}{\text{D}\_{\text{Ni}} \cdot \text{X}\_{\text{Ni}}} \tag{8}$$

where AWNi is the atomic weight of nickel (gNi/molNi), XNi is the nickel loading (gNi/gcat) and DNi is the dispersion of nickel.

#### *2.3. In Situ FTIR Spectroscopy*

In situ Fourier transform infrared (FTIR) experiments were carried out using an iS20 FTIR spectrometer (Nicolet, Thermo Fischer Scientific, Waltham, MA, USA) equipped with an MCT detector, a KBr beam splitter and a diffuse reflectance (DRIFT) sampling system (Specac, Orpington, UK) accompanied by an environmental chamber suitable for the study of diffusely reflecting solid samples in a controlled atmosphere. A flow system equipped with mass flow controllers, a steam saturator and a set of valves used for controlling the gas stream interacted with the catalyst surface, was directly connected to the gas inlet of the environmental chamber.

In a typical experiment, the catalyst powder was placed in the sampling system and heated at 500 ◦C in flowing helium for 10 min and then reduced under hydrogen flow at 300 ◦C for 30 min. The flow was then switched to He and the temperature was increased at 500 ◦C. After remaining 10 min at this temperature the sample was cooled at 100 ◦C. While cooling, the background spectra were recorded at the desired temperatures. Finally, the flow was switched to the reaction mixture, which consisted of 0.5%C3H8 +5%H2O (in He). Steam was introduced to the system via an independent He line passing through a saturator containing water maintained at 60 ◦C. The resulting gas mixture was fed to the DRIFT cell through stainless steel tubing maintained at 60 ◦C by means of heating tapes. A spectrum was collected at 100 ◦C after 15 min-on-stream followed by a stepwise increase of temperature up to 500 ◦C. During heating, spectra were recorded at selected temperatures after an equilibration for 15 min. In all experiments, the total flow through the DRIFT cell was 30 cm<sup>3</sup> min−1. Reaction gases (He, 2%C3H8/He, H2) are supplied from high-pressure gas cylinders (Buse Gas, Bad Hönningen, Germany) and are of ultrahigh purity.
