Comparison of Packed-Bed and Micro-Channel Reactors for Hydrogen Production via Thermochemical Cycles of Water Splitting in the Presence of Ceria-Based Catalysts

Abstract

Hydrogen production via two-step thermochemical cycles over fluorite-structure ceria (CeO₂) and ceria-zirconia (Ce_0.75Zr_0.25O₂) materials was studied in packed-bed and micro-channel reactors for comparison purposes. The H₂-temperature program reduction (H₂-TPR) results indicated that the addition of Zr⁴⁺ enhanced the material’s reducibility from 585 µmol/g to 1700 µmol/g, although the reduction temperature increased from 545 to 680 °C. Ce_0.75Zr_0.25O₂ was found to offer higher hydrogen productivity than CeO₂ regardless of the type of reactor. The micro-channel reactor showed better performance than the packed-bed reactor for this reaction.

1. Introduction

Hydrogen can be utilized in many modern-world applications. Its well-known challenges include production cost, transportation and storage. Hydrogen can be produced by various means, e.g., thermochemical processes, reforming processes, gasification, electrolysis, biological processes, and so on [1,2,3,4,5,6,7,8]. Conventional hydrogen production from either natural gas, coal, or biomass appears to be the most commercially available and affordable, although this unavoidably releases carbon emissions, radioactive elements, and air-borne pollutions into the atmosphere. Hydrogen production from water is a green technology in which water is split, producing high-purity gaseous hydrogen. Recent water splitting processes including, for instance, photo-catalytic, two-step thermochemical cycles, electrolysis and biological processes have been employed to generate high-purity hydrogen. The two-step thermochemical cycle possesses advantages over the others in terms of the product’s purity and yield [9]. The two-step thermochemical cycle reaction consists of (1) endothermic reduction, where the metal oxide material is reduced by thermal energy and/or chemical reducing agents, resulting in gaseous oxygen as a by-product and an active and non-stoichiometric reduced metal oxide, (2) exothermic oxidation, where the active metal oxides are oxidized by water, giving high-purity hydrogen as a product while the metal oxide is recycled back into its original stage [10,11,12]. The operating conditions—i.e., temperature, feed reactant and reaction time—of the two steps are different. Thus, the process can either be carried out in (1) two reactors between which the solid material is moved, or (2) one reactor in which the operating conditions are switched back and forth from the reduction to the oxidation step. The former is suitable for solid materials that have high mechanical strength, have changed phase, and have been fully reduced/oxidized based on their stoichiometry, such as ZnO/Zn [13], CdO/Cd [14], SnO₂/SnO [15], and GeO₂/GeO [16], although the process requires a sophisticated quenching and control system. The latter is commonly used with materials with no phase change. The materials could be either stoichiometric redox materials such as Fe₃O₄/FeO [17], MFe₂O₄ [18], and CoFe₂O₄/Al₂O₃ [19], or non-stoichiometric redox materials such as, for example, ceria and ceria-based materials [20,21,22], and perovskites [23]. Ceria (CeO₂) and ceria-zirconia (CeO₂/ZrO₂) were selected as the oxygen carrier in this work, as they offer (1) high oxygen storage capacity [24,25], (2) high thermal stability, and (3) the possibility of being fabricated as the whole reactor itself. Although the latter process does not require a quenching system and complicated fluidized bed operation, it needs a well-established control system that allows the operational conditions to switch from reduction to oxidation steps precisely. Regardless of the quality of the control system, the switching between conditions (temperature, feed) still causes low overall process efficiency (10 to 50%) [26,27,28,29]. Therefore, a reactor that offers rapid heat and mass transfer during the change is beneficial for this process. Additionally, the redox materials must be able to withstand the severe condition swing. This work studied the process under the same reduction and oxidation temperature, aiming to (1) optimize the process efficiency and product yield, (2) ease the control system, and (3) make it possible to use wider variety of the new catalysts. A micro-channel reactor was also applied to enable rapid mas/heat transfer, with shorter residence time.

2. Methodology

2.1. Catalyst Preparation, Substrate Pretreatment and Catalyst Coating

2.1.1. Catalyst Preparation

Nitrate precursors of Ce and Zr were mixed with 0.1 M cetyl trimethyl ammonium bromide (CTAB) at room temperature while NaOH was added dropwise, keeping pH at 11 while stirring. Molar ratio of CTAB to total cations (Ce + Zr) was fixed at 0.8. The mixture was centrifuged at 4000 rpm for 10 min, and then washed with de-ionized water to remove any possible impurities. The precipitate was dried in the oven at 120 °C overnight and calcined at 700 °C for 3 h in air with 2 °C/min heating rate. For the micro-channel reactor, alumina solution was additionally synthesized by hydrolysis of aluminum tri-sec-butoxide (Aldrich, St. Louis, MO, USA, 97%) with a molar ratio of H₂O to Al = 100 to enhance the adhesion force between the catalyst power and the surface of the substrate. The solution was peptized by adding nitric acid (HNO₃/Al = 0.07) and refluxed at 85 °C for 20 h. After that, nitrate precursors of Ce and Zr were dissolved into the prepared alumina solution. The mixed solution was heated at 85 °C until a thixotropic solution was obtained. The gel was dried at 110 °C overnight and calcined at 800 °C for 6 h at a heating rate of 2 °C/min.

2.1.2. Stainless Steel Substrate Preparation

The lab-designed micro-structured stainless steel (316L) plates were commercially made (TGI, Chonburi, Thailand), as shown in Figure 1. Each substrate has 14 channels with 300 µm depth, 370 µm width and 25 mm length per channel. The plates were cleaned by etching with 20% citric acid in an ultrasonic bath for 30 min. They were subsequently annealed at 800 °C for 2 h in air to form a layer of mixed metal oxides for better adhesion strength.

2.1.3. Catalyst Coating

Polyvinyl alcohol was dissolved in deionized water while stirring at 400 rpm at 65 °C for 2 h. The solution was left overnight without stirring. The catalyst powder and acetic acid were added to the solution. The weight ratio of catalyst powder to water to binder to acid was 10 to 84 to 5 to 1. The resulting suspension was stirred at 65 °C for 2 h, cooled to room temperature, and subsequently stirred for 3 days. The non-coating area of the micro-structured stainless steel plates, such as the inlet and the outlet, were covered with the polymer film. The prepared suspension was wash-coated on the micro-channel substrate, then left to dry at room temperature for 6 h. After the removal of the polymer film, the substrate was dried in an oven at 120 °C and calcined at 500 °C for 3 h at a heating rate of 1 °C/min. The 2 micro-channel substrates were then laser-welded together. The gas inlet and outlet at the top and bottom were connected to 1/8-inch stainless steel tubing.

2.2. Characterization

The crystallinity of the synthesized catalysts was characterized using X-ray diffraction (XRD). The morphology of the stainless steel substrates, before and after the annealing process at different temperatures, was characterized using scanning electron microscopy (SEM). The reduction temperature of the samples was investigated using the H₂-temperature program reduction technique (H₂-TPR). During the H₂-TPR, 10% H₂/Ar was passed through the catalyst’s bed in the reactor. The temperature was increased from room temperature to 950 °C, at a heating rate of 5 °C/min, and held for 30 min. The gaseous products were analyzed using on-line mass spectrometer (MS, GSD 320 O1, OmniStar gas analysis) for all experiments.

2.3. Experimental Set-Up

The catalysts were pelletized and sieved to 180–212 µm for the packed-bed reactor, and <38 µm for the micro-channel reactor. The weight of the catalyst was 1 g and 8–15 mg for the packed-bed reactor and the micro-channel reactor, respectively. Each catalyst was placed inside the quartz tube packed-bed reactor (i.d. = 10 mm, length = 50 cm) between two layers of quartz wool. The packed-bed and micro-channel reactors were placed in the middle of an electrical furnace (Inconel, 20 cm heating zone). The schematic diagram of the experiments is shown in Figure 2. The system was purged by 300 mL/min of Ar using a mass flow controller (New Flow-TLFC-00-A-1-W-2, 10–500 mL/min) at room temperature for 1 h. Each catalyst was reduced before use in 10% H₂/Ar at 700 and 900 °C, in accordance with the TPR results, for 30 min. Steam was generated using a steam generator, whereby the amount of water was controlled by peristaltic pump (BT100M Model, 0.00067–65 mL/min). Steam was delivered to the reactor system through a 170 °C trace-heated line to prevent condensation. The total flow rate of 200 mL/min, which ensured the reaction control regime, was fixed for all of the experiments. The reaction was isothermally operated at 700 and 900 °C. The gaseous product stream was analyzed using an on-line mass spectrometer (Quadrupole, Omnistar, GSD 320 O1 Model).

3. Result and Discussion

3.1. Characterization

Figure 3 shows diffractograms of all of the prepared catalysts, compared with pure alumina, shown as (a) CeO₂, (b) Ce_0.75Zr_0.25O₂, (c) 10%CeO₂/Al₂O₃, (d) 10%Ce_0.75Zr_0.25O₂/Al₂O₃, and (e) pure alumina.

CeO₂ and Ce_0.75Zr_0.25O₂ possessed a fluorite cubic structure with a face-centered cubic crystal system (FCC), which offers a high thermal stability and oxygen exchange kinetic rate [30,31,32]. CeO₂ and Ce_0.75Zr_0.25O₂ presented a lattice plane corresponding to the (111), (200), (220), (311), (222), (400) and (331) [33,34]. Their average crystal size, calculated using the Scherer equation, was 15.15 nm and 9.01 nm, respectively. Pure alumina was η-alumina and had a bayerite structure. Each peak corresponded to the (220), (311), (400), (333), and (440) lattice planes [35,36]. Both 10% CeO₂/Al₂O₃ and 10% Ce_0.75Zr_0.25O₂/Al₂O₃ showed η-alumina bayerite structures as a major crystalline phase, while CeO₂ and Ce_0.75Zr_0.25O₂ showed them as minor crystalline phase. The non-alumina catalysts were selected for the further study.

CeO₂ and Ce_0.75Zr_0.25O₂ showed two main reduction peaks at different temperatures, shown in Figure 4. Ceria and ceria-zirconia started to be reduced at the same temperatures, with the first peak at 300 °C (peaks α) and the second peak at 650 °C (peaks β). However, broader reduction peaks were observed in the ceria-zirconia due to the larger amount of oxygen release in ceria-zirconia. This led to a higher average reduction temperature for ceria-zirconia (680 and 950 °C), when compared to ceria (545 and 900°C), respectively. The results agreed with previous works reported by other researchers [37,38,39,40]. The first peaks of both catalysts were defined as surface reduction, evidenced by the steep reduction peaks, while the second peaks comprised bulk reductions, which were much broader compared to the first peaks due to the much slower solid-state oxygen diffusion within the materials. From the H₂-TPR profiles of both catalysts, it can be seen that ceria-zirconia had a higher reduction rate than ceria, and released a higher amount of oxygen within the studied temperature range. The addition of Zr into the ceria catalyst system has been suggested to affect the material’s cell volume, resulting in an increase in surface area [41,42,43,44]. The degree of the reduction, represented by the non-stoichiometric oxygen release (δ), is calculated and tabulated in

Table 1. H₂-TPR results of CeO₂ and Ce_0.75Zr_0.25O₂

Catalysts	Peak α			Peak β			Total OSC (µmol/g)	Total % Reduction (δ)
	Tred (°C)	OSC (µmol/g)	% Reduction (δ)	Tred (°C)	OSC (µmol/g)	% Reduction (δ)
CeO2-δ	545	585	9.34% (δ = 0.046)	900	915	14.64% (δ = 0.074)	1500	23.98% (δ = 0.12)
Ce0.75Zr0.25O2-δ	680	1700	547.40% (δ = 0.272)	950	1390	44346% (δ = 0.222)	3090	98.86% (δ = 0.49)

. For CeO₂, the surface oxygen was released at 545 °C, giving δ at 0.046 (theoretically maximum at 0.5) and percentage of reduction at 9.34%, where the material was reduced to CeO_1.95. The second peak of CeO₂ showed that the bulk oxygen was reduced at 900 °C, giving 0.074 of δ, which is equal to a reduction of 14.64%. At this stage, the CeO_1.88 became CeO_1.94. Similarly, ceria-zirconia was surface reduced and bulk reduced at 680 °C and 950 °C, respectively. The first reduction peak represented a δ of 0.277, with a reduction degree of 54.40%, while the second peak gave a δ of 0.222, with a reduction degree of 44.46%. The non-stoichiometric molecular formula of the ceria-zirconia after being reduced at 950 °C was Ce_0.75Zr_0.25O_1.51. Ce_0.75Zr_0.25O₂ conclusively showed better performance, compared to ceria, in terms of reduction rate and reducibility.

Where OSC stands for oxygen storage capacity, calculated from H₂ consumption. The percentage of reduction degree is calculated using Equations (1)–(3) below:

$$\%X_{red}=\left(\frac{n_{[O]real}}{n_{[O]\max }}\right)\times 100$$

(1)

$$n_{[O]real}=\frac{n_{H_2,consumed}}{m_{solid}}$$

(2)

$$n_{[O]\max }=\left(\frac{m_{solid}}{M{W}_{solid}}\right)\times {\delta }_{\max }$$

(3)

where $n_{{\left[O\right]}_{max}}$ is the maximum amount of O₂ release/uptake (mol/g) as a function of δ, δ is the stoichiometric coefficient of O in lattice (for this material, δ is 0.5 for maximum O release/uptake), $n_{{\left[O\right]}_{real}}$ is the number mole of O released per gram of the catalyst, which is equal to an integration of the area under the H₂ consumption curve, m_solid is the weight of the solid reactant, and MW_solid is the molecular weight of the catalyst [45].

Figure 5 shows the surface morphography of the prepared substrates (a) before and after annealing process at (b) 600, (c) 700, and (d) 800 °C. From the results, it can be seen that the oxides of the stainless steel’s surface were formed by annealing, occurring best at the highest temperature: 800 °C.

Energy-dispersive X-ray spectroscopy technique (EDX) was applied to identify the oxides which were formed with metals consisting of the stainless steel.

Table 2. Percentage of each element on the surface of the substrate.

Condition	O	Cr	Mn	Fe	Ni	Others	Total
Before annealing	7.69	28.61	−	47.87	5.65	10.18	100
Annealed at 600 °C	24.84	29.53	2.72	35.95	4.22	2.75	100
Annealed at 700 °C	30.04	26.34	−	30.44	3.75	9.43	100
Annealed at 800 °C	41.78	27.90	4.82	18.70	−	6.79	100

presents the percentage of each element. The results suggested that the formation of oxides increased with an increase in annealing temperature, as evidenced by the higher percentage of oxygen. The results are in agreement with the SEM results, as reported in the previous section.

3.2. Catalytic Performance Experiments

3.2.1. Effect of ZrO₂ Addition

Catalytic performances of CeO₂ and Ce_0.75Zr_0.25O₂ towards the water splitting were tested in a packed bed reactor. The oxidation and reduction temperatures were paired at the same temperature to avoid temperature switching, aiming to maximize the overall process efficiency and minimize the thermal stress of the reactor. The selected temperatures of reduction and oxidation (Tred/Tox) were 700/700 and 900/900 °C.

From Figure 6, it can be seen that Ce_0.75Zr_0.25O₂ had better performance towards this reaction than CeO₂ for both selected temperatures. The effect was more obvious at the lower temperature (700 °C) than the higher temperature (900 °C). At 700 °C, the H₂ production of Ce_0.75Zr_0.25O₂ was 88.71% more than that produced via CeO₂ (increased from 483.42 to 912.26 μmol/g), although it was only 8.40% higher (increased from 1563.21 to 1694.55 μmol/g) at 900 °C. This result shared the same trend as previous work reported by Z. Zhao et al. [46], although both catalysts in this research offered around two times higher of H₂ productivity for both temperatures, which could be due to the benefit of a surfactant-assisted method which allows smaller fine particles and larger specific surface area [47].

3.2.2. Comparison of Micro-Channel Reactor and Packed-Bed Reactor

Ce_0.75Zr_0.25O₂ was selected for this experiment. H₂ production via both packed-bed and micro-channel reactors were compared. Uncoated micro-channel reactor was also introduced to the reactant stream and performed as a blank test. H₂ productivity in all experiments was calculated per weight of catalyst used. H₂ productivity from the blank micro-channel reactor, uncoated reactor, was subtracted from the catalyst-coated micro-channel reactor. The blank test was performed only for the micro-channel reactor, because the formation of metal oxides on the substrate surface after annealing, such as on Cr₂O₃, Mn₂O₃ and Fe₂O₃, could possibly be involved in the catalytic reaction. From Figure 7, the 5-cycle average amount of H₂ production using the packed-bed reactor and the micro-channel reactor was 912.26 and 14,308.32 μmol/g. Thus, micro-channel reactor showed roughly 16 times better performance than the packed-bed reactor, in terms of H₂ production. This was suggested to be the effect of its high surface-to-volume ratio, leading to an intrinsic reaction occurring at the molecular level [48,49,50,51]. It can be noticed that the H₂ productivity of the packed-bed reactor decreased while that of the micro-channel reactor increased when the number of cycles increased. The decrease in H₂ productivity in the packed-bed reactor was presumably due to the catalyst’s coagulation when repetitively used at such temperatures. On the other hand, the increase in H₂ productivity in the micro-channel reactor was possibly the result of the reactive oxides, in which they were formed by the reaction of metals in the stainless steel and the oxygen in the system. Thus, the more cycles the reaction was run, the more H₂ productivity was achieved. However, these H₂ productivities in both type of reactor were supposed to be constant after a certain number of cycles.

In addition, the decline of the H₂ production rate in the micro-channel reactor was noticed to be much slower than that in the packed-bed reactor, leading to a longer reaction time. This was due to (1) the micro-channel reactor having a higher active surface area to volume, allowing better access of the reactants to the catalyst’s surface, and (2) channeling and/or mass transfer limitation occurring in the packed-bed reactor, and thus, the reaction time being shorter than it should be.

3.2.3. The Influence of Alumina Addition into the Catalyst System in the Micro-Channel Reactor

Alumina was added to the Ce_0.75Zr_0.25O₂ catalyst system and used in this experiment, as it was believed to increase the adhesion force between the active catalyst powders and the surface of stainless steel substrates [52,53]. From Figure 8, it can be seen that the 5-cycle average H₂ amount, produced in the presence of the bare reactor, Ce_0.75Zr_0.25O₂, and Ce_0.75Zr_0.25O₂/Al₂O₃, was estimated at 130.83, 245.61 and 108.45 μmol, respectively. This concluded that the addition of alumina had a negative effect on H₂ productivity. In addition, H₂ production when using Ce_0.75Zr_0.25O₂/Al₂O₃ was lower than when using the blank reactor. This means that the addition of alumina inhibited the access of water to react with Ce_0.75Zr_0.25O₂. Meanwhile, the surface of the blank reactor was obviously a catalyst itself, as it was formed by the oxidation of metals in stainless steel, such as Cr, Mn, and Fe. These metal oxides are known as redox catalysts, and could therefore catalyze this reaction.

4. Conclusions

Ce_0.75Zr_0.25O₂ was found to have better catalytic performance towards the two-step thermochemical cycles of water splitting, compared to CeO₂. The higher oxygen storage capacity was suggested to be the cause. Oxygen mobility in the lattice depended on the effective radius of the cations. Thus, the partial substitution of Ce⁴⁺ (higher ionic radius of 0.97 Å) with Zr⁴⁺ (smaller ionic radius of 0.84 Å) could create a smaller unit cell volume and larger channel radius in the lattice while the desired fluorite structure of ceria was remained. The infinitesimal cell volume required less energy for the hopping of oxygen ions; therefore, the active oxygen could be easily migrated from one vacancy to the others through the channel radius in the lattice. In a packed-bed reactor, it gave 1694 µmol/g of H₂ productivity at 900 °C. Because the micro-channel reactor was fabricated from stainless steel, the highest operating temperature of the micro-channel reactor was limited to 700 °C. The micro-channel reactor was proved to show 16 times higher H₂ productivity compared with the packed-bed reactor when operated at 700 °C. This was the result of the high surface-to-volume ratio of the micro-channel reactor, which allows better access of the gaseous reactant to react with the catalyst.