<figure><img src="image_1.png" /></figure>

HAL
open science

# Catalytic methane combustion in plate-type microreactors with different channel configurations: An experimental study

Li He, Yilin Fan, Jérôme Bellettre, Jun Yue, Lingai Luo

► To cite this version:

Li He, Yilin Fan, Jérôme Bellettre, Jun Yue, Lingai Luo. Catalytic methane combustion in plate-type microreactors with different channel configurations: An experimental study. Chemical Engineering Science, 2021, 236, pp.116517. 10.1016/j.ces.2021.116517. hal-03155792

HAL Id: hal-03155792

https://hal.science/hal-03155792

Submitted on 2 Mar 2021

HAL is a multi-disciplinary open access
archive for the deposit and dissemination of sci-
entific research documents, whether they are pub-
lished or not. The documents may come from
teaching and research institutions in France or
abroad, or from public or private research centers.

L'archive ouverte pluridisciplinaire HAL, est
destinée au dépôt et à la diffusion de documents
scientifiques de niveau recherche, publiés ou non,
émanant des établissements d'enseignement et de
recherche français ou étrangers, des laboratoires
publics ou privés.
He, L., Fan, Y., Bellettre, J., Yue, J., & Luo L. (2021). Catalytic methane combustion in plate-type microreactors with different channel configurations: an experimental study. *Chemical Engineering Science, 236, 116517*. https://doi.org/10.1016/j.ces.2021.116517

4

5

# Catalytic methane combustion in plate-type microreactors with different channel configurations: an experimental study

Li He<sup>a, b</sup>, Yilin Fan<sup>a</sup>, Jérôme Bellettre<sup>a</sup>, Jun Yue<sup>b, *</sup>, Lingai Luo<sup>a, *</sup>

<sup>a</sup> Université de Nantes, CNRS, Laboratoire de thermique et énergie de Nantes, LTeN, UMR 6607, F-44000 Nantes, France

10

<sup>b</sup> Department of Chemical Engineering, Engineering and Technology Institute Groningen, University of Groningen, 9747 AG Groningen, The Netherlands

11

12

13

## ABSTRACT:

This paper presents an experimental study on the catalytic methane combustion (CMC) in plate-type microreactors with wall-coated Pt/γ-Al₂O₃ catalyst. Firstly, the influence of different operational conditions and coating properties on the CMC in the straight parallel-channel microreactor has been investigated. A specific catalyst loading of 57.6 g m⁻² was found to yield the highest methane conversion over 3.5 wt% Pt/γ-Al₂O₃. A higher or lower loading tended to decrease the methane conversion due to either the limited internal diffusion through the thicker coating layer or insufficient active sites in the thinner coating layer. Then, the above microreactor was compared with other five different geometries, including cavity, double serpentine microchannels, obstacled microchannels, meshed circuit and vascular network. The double serpentine microchannel geometry presented the highest methane conversion (especially at a

15

16

17

18

19

20

21

22

23

24

<footnote>* Corresponding authors.</footnote>

<footnote>Email addresses: yue.jun@rug.nl (J. Yue), lingai.luo@univ-nantes.fr (L.Luo)</footnote>
1 relatively low mixture flow rate) due to the appropriate control over the residence time and

2 catalyst coating surface area.

3

4 Keywords: Catalytic methane combustion; microreactor; washcoated catalyst; channel

5 configuration; methane conversion; flow distribution

6

7 Nomenclature

A Total cross-sectional area of the reaction microchannel in a microreactor, m²

$F_{CH_4,i}$ Inlet molar flow rate of CH$_4$, mol s$^{-1}$

$F_{CH_4,o}$ Outlet molar flow rate of CH$_4$, mol s$^{-1}$

$F_{CO_x,o}$ Outlet molar flow rate of CO or CO$_2$, mol s$^{-1}$

$F_{H_2,o}$ Outlet molar flow rate of H$_2$, mol s$^{-1}$

$h$ Channel height, m

$l_{j,tot}$ Total channel length between bifurcation indices $j$ and $j+1$, m

$l_{j,1}$ Length of horizontal channel at the index $j$ in the tree-like bifurcated structure, m

$Q_{tot}$ Total volumetric flow rate, m³ s$^{-1}$

$r$ Ratio of channel width of the downstream to that of the upstream in one bifurcation, -

$S$ Inner surface area of microchannels subjected to Pt/γ-Al$_2$O$_3$ coating, m²

$S_{H_2}$ Selectivity of H$_2$, %

$S_{CO_x}$ Selectivity of CO or CO$_2$, %

$U$ Mean velocity of methane-air mixture in the reaction microchannel, m s$^{-1}$
<table><tr><td>V<sub>tot</sub></td><td>Total microchannel volume for reaction in a microreactor, m<sup>3</sup></td></tr><tr><td>W<sub>cat</sub></td><td>Catalyst mass, g</td></tr><tr><td>w<sub>j</sub></td><td>Channel width at the index j in the tree-like bifurcated structure, m</td></tr><tr><td>X<sub>CH<sub>4</sub></sub></td><td>CH<sub>4</sub> conversion, %</td></tr></table>

## 1 Greek symbols

<table><tr><td>τ</td><td>Mean residence time, s</td></tr><tr><td>Φ</td><td>Inlet molar ratio of oxygen to methane, -</td></tr><tr><td>φ</td><td>Specific catalyst loading, g m<sup>-2</sup></td></tr><tr><td>μ</td><td>Dynamic viscosity, Pa s</td></tr><tr><td>ρ</td><td>Density, kg m<sup>-3</sup></td></tr></table>

## 2 Abbreviation

<table><tr><td>CMC</td><td>Catalytic methane combustion</td></tr><tr><td>GC</td><td>Gas chromatography</td></tr><tr><td>i.d.</td><td>Inner diameter</td></tr><tr><td>MW</td><td>Molecular weight</td></tr><tr><td>o.d.</td><td>Outer diameter</td></tr><tr><td>PVA</td><td>Polyvinyl alcohol</td></tr></table>

## 3

## 4 1. Introduction

5 Natural gas has been reported to have a largest increment in consumption in the past decade, 6 accounting for nearly half of the increase in global energy demand in year 2018 [1-3]. The 7 combustion of natural gas presents a particular advantage of higher energy content per CO<sub>2</sub> 8 emission (55.7 kJ g<sup>-1</sup> if fully based on methane as its main component) than coal (39.3 kJ g<sup>-1</sup>) 9 and petroleum (43.6 kJ g<sup>-1</sup>), a character which is essential to a low carbon future. In addition to 10 the abundant natural gas reserves proven worldwide, biomethane could be obtained from organic
wastes and residues by biomass gasification/digestion [4,5], providing a competitive energy
supply towards a more electrified world [3]. As a result, natural gas has been suggested as a
substitute for oil and coal as a future leading energy source for the next 20 years [3,6]. According
to the International Renewable Energy Agency (IRENA) [6], the utilization of petroleum and
coal would decline by 70% and 85%, respectively, by 2050, whereas natural gas would be the
largest source by then and achieve peak on its utilization at around 2027 [6]. Thus, a great number
of researches have been devoted to the development of energy-efficient natural gas combustion
systems in many aspects for their application in the industrial, transport and domestic areas, etc.
[7].

However, the gas industry faces some commercial and environmental challenges [5], due to the
releasing of pollutant emissions (e.g., NOx, CO and unburned hydrocarbon) by the conventional
flame combustion of methane typically occurring at above 1400 °C. The harmful impacts of these
emissions on the human health and environment have been well recognized [8] and the
applicable regulations over EU countries have become more and more stringent in recent years
[9]. In this context, the catalytic methane combustion (CMC) as a promising alternative has
received increasing attention [10]. A lower working temperature (e.g., < 600 °C) is needed for
the complete oxidation of methane in the presence of catalyst because of the reduced activation
energy (40-80 kJ mol⁻¹) compared to that for conventional combustion (100-200 kJ mol⁻¹).
Hence, the exhaust emissions (especially of NOx) can be remarkably abated [7,10-12]. Abundant
literature is available on the CMC, as summarized by recently published review papers [10,13-
15]. These existing CMC studies distinguish themselves by focusing on the catalyst development
[16-19] and mechanisms [20-24], different types of catalytic reactors [25-29] and (optimized)
reaction conditions [30-33], as well as the target applications [29,34,35].
Various types of catalysts have been studied for the CMC including noble metal catalysts and mixed oxides catalysts (perovskites and hexaaluminate) [10,36,37]. Among them, noble metal catalysts (e.g., Pt, Pd, Rh) are the most commonly used and well developed owing to their high specific surface area, high catalytic activity and low light-off temperature [30,38]. Bimetallic catalysts (e.g., Pd-Pt/Al₂O₃) often show higher catalytic activity and selectivity than their monometallic counterparts due to the synergistic effect by metal-metal interaction [39,40]. Additionally, different supports and additives (e.g., ZrO₂, CeO₂ [41,42]) have been introduced to improve the catalytic activity through the acceleration of oxygen exchange. In terms of the catalyst form, washcoated catalysts have undergone rapid development in recent years. The washcoated catalyst is commonly deposited as a thin layer on structured surfaces of monolithic reactors or (multichannel) microreactors, in order to avoid the high pressure drop and large temperature gradient likely existing in conventional fixed/packed bed reactors loaded with catalyst powders or pellets [27]. In this regard, the good adhesion and coverage of the coated catalyst layer on the channel walls are essential for the CMC performance of the reactor. The suspension method [43-46] and sol-gel method [47,48] are most commonly used for washcoating. It has been reported that the preparation method, binder nature, pH value and particle size could be the key factors that have to be adapted to obtain a high thermal stability and a uniform coverage of the coating [46].

Fixed/packed-bed reactor (catalysts typically in the form of fine powder) is conventionally employed for the research of the catalyst activity and reaction mechanism due to the easy operation and enhanced catalyst spatial density. Nevertheless, the low surface area, poor temperature uniformity and relatively high pressure drop are main disadvantages of the conventional fixed-bed reactor [15]. As an interesting alternative, the microreactors present a high surface-to-volume ratio, relatively low pressure drop, and high mass/heat transfer rate (especially when combining with washcoat catalyst), which is beneficial to prevent the catalyst
sintering in the long term due to the better temperature regulation. As more attention has been
paid on the research of microreactor, it could become a promising candidate for potential
industrial applications. In particular, plate-type microreactors containing a multitude of
microchannels arranged in parallel have attracted great interests [49,50]. In combination with the
well-adhered catalyst coating on the wall surface, the plate-type multichannel microreactors offer
numerous advantages, including the compact design, easy to manufacture and the scaling-up
potential to handle a large amount of reactive gas [51,52]. These features are especially beneficial
for handling the CMC or highly exothermic reactions in general by suppressing the presence of
temperature hot spots due to the local accumulation of reaction heat [53-55]. For instance, the
experimental study by O'Connell et al. [27] shows that a 100 % conversion could be obtained in
a multichannel microreactor (500 µm width × 250 µm depth, 14 channels in total) washcoated
with Pt-W or Pt-Mo/γ-Al₂O₃ catalyst at 600 °C and a total gas flow rate of 107 mL min⁻¹ (space
velocity of 74,000 h⁻¹) and an O₂/CH₄ molar ratio of 2.2. He et al. [46] experimentally performed
the CMC in a multichannel reactor (275 mm length × 1.5 width mm × 1 mm height, 32 channels
in total) over the washcoated Pt/γ-Al₂O₃ catalyst. A 95.75 % methane conversion was obtained
at 450 °C under a total gas flow rate of 110 mL min⁻¹ (space velocity: 6,557 h⁻¹) and an O₂/CH₄
molar ratio of 2. Moreover, the compact and modular feature of the plate-type multichannel
microreactor makes it possible for a better utilization of the reaction heat [56,57]. The
development of the so-called autothermal reactor that combines the CMC (exothermic reaction)
with another endothermic reaction (e.g., methane steam reforming to produce hydrogen or
syngas) within a single device has become a research hotspot over the past decade [58].
Mundhwa et al. [56,57] proposed a stacked autothermal reactor design which was composed of
two plates with parallel microchannels (5 cm length × 1 mm width, 15 channels, coated with ca.
20 µm thick Pt/Al₂O₃ catalyst) for the CMC and other two plates of cavity shape (5 cm length ×
5 cm width × 100 µm depth, with washcoated Ni/Al₂O₃ catalyst) for the methane steam reforming
reaction. Their numerical results show that this compact multiple plate design (especially under
1 the co-current mode between flows) could reduce the required amount of catalyst for the CMC
2 by 70% and maintain a uniform temperature profile for the methane steam reforming [56,57].
3 The scale-out strategy of microreactor stacks has been investigated numerically for coupling
4 such exothermic and endothermic processes [59]. The results illustrate that a certain number of
5 stacks (e.g., > 15) was needed so that the released reaction heat could compensate for the heat
6 loss and thus drive the reforming reaction effectively [59].

7
8 It is noteworthy that most of the plate-type multichannel microreactors tested for the CMC or
9 simulated for its autothermal coupling have a simple configuration like parallel straight channels
10 [27,46,56-60]. In this respect, the influence of the operating parameters and the washcoated
11 catalyst properties (e.g., of the promising noble metal-based one) on the CMC performance still
12 needs more systematic investigations, as the current understanding is still far from being
13 sufficient. Furthermore, in order to provide an effective way to realize the miniaturization, the
14 design and optimization of microchannel geometries deserve particular attention. On one hand,
15 the improved the mixing and the enhanced the mass/heat transfer of appropriate microchannel
16 configurations (compared to parallel straight channel) will render better reactor performance so
17 that the (nearly) 100 % methane conversion could be achieved at a lower temperature or with a
18 lower needed amount of catalysts. The lifetime of noble metal catalysts could be prolonged for
19 a higher cost-effectiveness of the CMC technology. On the other hand, the better thermal
20 management of this strong exothermic reaction achieved by optimizing the reaction channel
21 configuration will facilitate the design of the associated cooling system, which is especially
22 beneficial for various aimed applications such as autothermal reactor and CMC boiler. However,
23 the testing and comparison of different microchannel geometries on the CMC performance have
24 been rarely reported in the literature. There is still much room for the reactor performance
25 improvement by optimizing the internal microchannel configurations that lead to a better process
26 control (e.g., narrowed residence time distribution and more uniform reactant flow distribution)
and catalyst utilization (e.g., high catalyst surface area and improved gas-solid mass transfer) in the CMC. Such microreactor performance optimization is equally important towards obtaining a better usage of the released reaction heat from the CMC in energy applications, such as the small-scale water boiler and the autothermal reaction device.

As a continuation of our previous study [46], this work has firstly presented an experimental investigation into the CMC performance over the microreactor with parallel straight channels coated with Pt/γ-Al₂O₃ catalyst. The influence of different operation conditions (e.g., temperature, flow rate and oxygen to methane molar ratio), catalyst loading and thickness, reaction microchannel length was studied. Then, the reaction performance of this type of microreactor was compared with those of other five type of microreactors with different channel configurations. The studied configurations cover some relatively simple channel geometries (cavity and double serpentine microchannels) and more complicated ones (obstacled microchannels, meshed circuit and vascular network). The methane conversion between different microreactor designs has been compared and explained based on the internal channel surface area, residence time and the specific catalyst loading. Moreover, the use of tree-like bifurcation structure for fluid distribution or collection in the microreactor was explored towards improved reaction performance. The findings of this work may help to achieve a flexible and compact design of microreactors that allows upscaling from laboratory to field-scale applications of the CMC with optimized performance.

21

## 2. Experimental

### 2.1. Experimental setup and procedures

The experimental test rig is shown in Fig. 1. Two mass flow controllers (MFC, Brooks SLA5850) were used to adjust the flow rates of methane and the synthetic air for the CMC experiment. The total gas flow rate was adjusted from 110 to 500 mL min⁻¹ (based on ca. 20 °C and 1 atm).

24

25

26
1

2 For each test, the oven (Nabertherm, LT 9/11/B170) was firstly heated up to the target

3 temperature at a ramp of ca. 15 °C min⁻¹ flowing with air to prevent any reaction happened during

4 heating up process. Once the oven temperature reached the target reaction temperature ranging

5 between 300 °C and 500 °C, the methane-air gas mixture (at an O₂/CH₄ molar ratio from 0.5 to

6) first passed through the inlet pre-heating coil (stainless steel; i.d.: 3.7 mm, o.d.: 6 mm, ca. 20

7 cm in length) in the oven, and was then introduced into the plate-type microreactor where the

8 catalytic reaction happened. The product gas first flew out through a condenser to remove water,

9 and was then analyzed by an online gas chromatography (GC).

<figure><img src="image_2.png" /></figure>

10

11 **Fig. 1.** Test rig for the CMC in the plate-type microreactor.

12

13 **2.2. Reactor design and fabrication**

14 The plate-type microreactor has an overall dimension of 190 mm in length, 130 mm in width and

15 10 mm in height. It was designed as a sandwich shape with a reaction plate (made of FeCrAlloy,

16 i.e., Kanthal A-1, 22 % Cr, 5.8 % Al and Fe for balance) in the middle, enclosed by two additional

17 blind plates (made of stainless steel) as the outside shell (Fig. 2). One blind plate has additional

18 inlet and outlet ports in order to interface with the reaction platelet (vide infra). A graphite gasket

19 (3 mm in thickness) was installed in between these plates to prevent the gas leakage and bolts

20 were used on the peripherals for further sealing. A liquid leak detector (Snoop) was applied on

21 the assembled microreactor in order to ensure the tightness before each series of the test.
1

2 The middle reaction plates have an overall dimension of 112.5 mm (length) × 50 mm (width) ×

3 3 mm (height). Each plate has a single inlet port and outlet port (i.d.: 6.5 mm) aligned with the

4 central line, the distance between their centers being 82.5 mm. Six shapes of the internal flow

5 network have been designed and machined on one side of the plate, all of them having a

6 symmetric geometry, as shown in Fig. 2. Each flow network consists of three parts: the inlet

7 distributor, reaction microchannel(s) subjected to the catalyst coating (on two sides and the

8 bottom surface of the microchannel wall) and outlet collector. A detailed description of the six

9 reactors (Reactors #1-6) and its flow network is presented as follows. Note that except for the

10 cavity case (Reactor #2), all the reaction microchannels have a rectangular cross-section of 1.5

11 mm in width and 1 mm in height.

12

13 Reactor #1 (denoted as the straight parallel channel microreactor) has a basic shape of 16

14 parallelized and straight reaction microchannels of 60 mm in length. The thickness of the

15 separating walls between neighbouring microchannels is 1.5 mm.

16

17 Reactor #2 (denoted as the cavity microreactor) has a simple cavity geometry for reaction

18 (dimension: 66.45 mm length × 46.5 mm width × 1.0 mm depth).

19

20 Reactor #3 (denoted as the double serpentine channel microreactor) consists of two serpentine

21 reaction microchannels arranged in axial symmetry. For each zigzag unit, the channel length of

22 the horizontal section is 21 mm and that of the vertical section is 4.5 mm, the angle of zigzags

23 being 90°. The thickness of the separating walls is 1.5 mm. The total length of each serpentine

24 microchannel is 547.02 mm.

25
1 Reactor #4 (denoted as the obstacled parallel channel microreactor) differs from the basic straight
2 parallel channel microreactor (Reactor #1) by the presence of solid obstacles (square shape, 1.5
3 mm × 1.5 mm) along the flow direction in the reaction microchannels. Each microchannel thus
4 assumes a split-and-recombine shape around successive obstacles (4.5 mm × 4.5 mm, 10 pieces
5 for each, 80 in total). The thickness of walls separating the microchannels is 1.5 mm. The
6 connection unit between repeating split-and-recombine units is 1.5 mm in length and 1.5 mm in
7 width.

8
9 Reactor #5 (denoted as the meshed circuit microreactor) presents a meshed flow circuit with
10 channel interlacing, inspired by the land crack architecture in nature [61]. The reaction
11 microchannels are intersected at a crossing angle of 90°. The solid diamond blocks formed
12 between the adjacent microchannels have a dimension of 2.74 mm × 2.74 mm with 143 pieces
13 in total.

14
15 Reactor #6 (denoted as the vascular microreactor) has two vascular blocks arranged in axial
16 symmetry. For each vascular block, there are two side vertical microchannels (69 mm length ×
17 1.5 mm width) for gas splitting whereas one central vertical microchannel (66 mm length × 1.5
18 mm width) for recombining. The side and central vertical microchannels are connected by means
19 of short horizontal microchannel sections (9 mm length × 1.5 mm width, 86 pieces in total) that
20 are arranged in a staggered pattern and separated by rectangular parallelepipeds (9 mm in length
21 and 1.5 mm in width).
<figure><img src="image_3.png" /></figure>

1

2 **Fig. 2.** Schematics of various plate-type microreactors. Reactors #1-6 denote, respectively, the

3 straight parallel channel microreactor, cavity microreactor, double serpentine channel

4 microreactor, obstacled parallel channel microreactor, meshed circuit microreactor and vascular

5 microreactor.

6

7 The inlet fluid distributor and outlet fluid collector fed with the reaction microchannel(s) also

8 differ to some extent in Reactors #1-6. The combination of a multi-scale bifurcated distributor

9 and a simple rectangular collector was employed for Reactors #1, 2, 4 and 5 (Fig. 2). Regarding

10 Reactors #3 and #6, the above structures are not very appropriate due to the geometrical
constraints, i.e., the reaction microchannel being either limited in number (only two in Reactor #3) or with a special shape (vascular in Reactor #6). Thus, a simple bilateral divergent or convergent channel network was used as the distributor or collector, respectively. The design, scaling relations and detailed dimensions for these distributors/collectors are found in Appendix A.

6

## 2.3. Catalyst preparation and coating procedures

The reaction platelets were first immersed with acetone at 45 °C for 30 min in the ultrasonic bath (PCE-UC 20) with the frequency being 40 kHz, in order to remove oil, grease and other dirt [62]. Thermal pretreatment was subsequently performed at 900 °C (ramp from room temperature: 20 °C min⁻¹; 10 h at the final temperature), in order to generate a thin alumina layer over the substrate which could be a strong bonding between the Pt/γ-Al₂O₃ coating layer and the substrate. The slurry suspension method was applied for the coating deposition in microreactors [63,64]. The materials used for preparing the catalytic coating were as follows: γ-Al₂O₃ (3 µm, 99.97 % on metals basis), PVA (98-99 % hydrolyzed), acetic acid and tetraammineplatinum (II) nitrate (Pt(NH₃)₄(NO₃)₂, 99.99 % on metals basis) purchased from Alfa Aesar. The γ-Al₂O₃ slurry was prepared by mixing γ-Al₂O₃ powder, PVA binder and acetic acid [43], based on the optimized composition identified in our previous study [46]: 5 wt% PVA (MW of 146,000 - 186,000), 20 wt% Al₂O₃ (3 µm) and 1 wt% acetic acid. The γ-Al₂O₃ slurry was then heated up to 65 °C for 2 h under 300 rpm stirring and stored at room temperature for at least 2 weeks to remove the bubbles inside the slurry before use.

22

The prepared slurry as the catalyst support precursor was first deposited on the walls (two sides and bottom) of the reaction microchannels for all Reactors #1 to 6 using syringe injection. The applied slurry weight was adjusted as per the required catalyst specific loading (vide infra). The excessive suspension outside the microchannel was immediately removed with a razor blade.

23

24

25

26
The platelets were then dried at room temperature overnight for at least 8 h, subsequently dehydrated at 120 °C for 8 h and finally calcined at 600 °C (ramp from room temperature: 2 °C min⁻¹; 2 h at the final temperature).

The incipient wetness impregnation was performed by manually dropping the Pt(NH₃)₄(NO₃)₂ solution with a certain Pt concentration (1.5 wt%, 3.5 wt% or 5 wt%) into the reaction microchannel of the platelet coated with γ-Al₂O₃. The impregnated coating was then dried at room temperature overnight for at least 8 h, and finally calcined at 500 °C (ramp from room temperature: 2 °C min⁻¹; 2 h at the final temperature). The adhesion of the Pt/γ-Al₂O₃ washcoat catalyst prepared following this procedure has been tested in our previous study by the ultrasonic test and no appreciable weight loss has been found [46]. Fig. 3 shows the prepared reaction plates with the coated catalyst. Some geometric parameters of these reactors and the involved Pt/γ-Al₂O₃ catalyst coating characteristics are presented in Table 1. Note that all the area and volume calculations are based on bare microchannels without considering the coating thickness. Based on SEM images from our previous study [46], the coating thickness range is ca. 18 µm (upper parts at side walls) to ca. 110 µm (corner parts), which is much smaller than the microchannel width/height. Given the negligible thickness of catalyst coating (64 µm average) compared with the reaction microchannel width and height, the pressure drop through the reaction microchannel remains almost unaffected by loading the catalyst.

20

<figure><img src="image_4.png" /></figure>

21
<caption>Fig. 3. Reaction platelets coated with Pt/γ-Al₂O₃ catalyst in Reactors #1-6.</caption>

**Table 1**

Geometric parameters of the tested microreactors and the involved Pt/γ-Al₂O₃ catalyst coating characteristics.

<table><thead><tr><th rowspan="2">Reactor</th><th rowspan="2">S (mm²)</th><th colspan="2">Bottom side <sup>a</sup></th><th colspan="4">W<sub>cat</sub></th></tr><tr><th>Wall side <sup>b</sup></th><th>Total</th><th>φ (g)</th><th>φ (g m<sup>-2</sup>)</th><th>V<sub>tot</sub> <sup>c</sup> (mm³)</th><th>τ <sup>d</sup> (s)</th></tr></thead><tbody><tr><td>#1</td><td>1476.00</td><td>1998.02</td><td>3474.02</td><td>0.27 <sup>e</sup></td><td>78.30 <sup>e</sup></td><td>1476.00</td><td>5.90</td></tr><tr><td>#2</td><td>3089.00</td><td>202.88</td><td>3291.88</td><td>0.25</td><td>76.40</td><td>3089.00</td><td>12.36</td></tr><tr><td>#3</td><td>1650.06</td><td>2200.08</td><td>3850.14</td><td>0.27</td><td>70.26</td><td>1650.06</td><td>6.60</td></tr><tr><td>#4</td><td>1620.00</td><td>1920.00</td><td>3540.00</td><td>0.24</td><td>67.51</td><td>1620.00</td><td>6.48</td></tr><tr><td>#5</td><td>1676.45</td><td>1815.26</td><td>3491.71</td><td>0.24</td><td>68.45</td><td>1676.45</td><td>6.71</td></tr><tr><td>#6</td><td>2166.00</td><td>2249.08</td><td>4415.08</td><td>0.26</td><td>58.28</td><td>2166.00</td><td>8.66</td></tr></tbody></table>

**2.4. Analytical procedure**

Gas product was analyzed by an online MicroGC (Agilent, R490 OBC) equipped with a thermal conductivity detector. A Molsieve 5Å column (MS5A, length: 20 m) and a pre-column (PPU, length: 10 m) were used. The concentrations of the reference gas used were 5.12 ± 0.1 mol% CH₄, 2.015 ± 0.04 mol% C₃H₈, 2.024 ± 0.061 mol% C₃H₆, 9.99 ± 0.2 mol% H₂, 9.98 ± 0.2 mol% CO, 1.994 ± 0.04 mol% C₂H₂, 1.999 ± 0.04 mol% n-C₄H₁₀, 2.017 ± 0.04 mol% C₂H₄, 9.92 ±

<footnote><sup>a</sup> The surface area of the bottom of the reaction microchannel. <sup>b</sup> The surface area of two sides walls of the reaction microchannel; <sup>c</sup> Volume of the reaction microchannel; <sup>d</sup> Residence time in the reaction microchannel, calculated based on a total gas flow rate of 150 mL min⁻¹. <sup>e</sup> Other catalyst weights and specific loadings were also tested. Here only the values for comparison with other type of microreactors are shown.</footnote>
1 0.02 mol% CO₂, and N₂ for the rest. Argon was used as the carrier gas. The MicroGC oven

2 temperature was heated from 40 °C up to 100 °C (ramp: 20 °C min⁻¹) and maintained for 7.5 min.

3 SOPRANE II software was used for the peak area integration and data analysis.

4

5 **2.5. Definitions**

6 The CH₄ conversion (X₄H₄), CO₂ (CO) selectivity (S₄COₓ) and H₂ selectivity (S₆H₂) are calculated

7 based on Eqs. (1) - (3), respectively.

$$
8 \quad X_{CH_4} = \frac{F_{CH_4,i} - F_{CH_4,o}}{F_{CH_4,i}} \times 100 \% \qquad (1)
$$

$$
9 \quad S_{CO_x} = \frac{F_{CO_x,o}}{F_{CH_4,i} - F_{CH_4,o}} \times 100 \% \qquad (2)
$$

$$
10 \quad S_{H_2} = \frac{F_{H_2,o}}{F_{CH_4,i} - F_{CH_4,o}} \times 100 \% \qquad (3)
$$

11 where *F* stands for the molar flow rate (based on ca. 20 °C and 1 atm). The subscripts *i* and *o*

12 indicate the inlet and outlet of the plate-type microreactor, respectively.

13

14 The mean residence time (τ) is defined as the total void volume of the coated microchannel(s)

15 (V<sub>tot</sub>) divided by the total volumetric flow rate of the gas mixture entering the reactor (Q<sub>tot</sub>).

$$
16 \quad \tau = \frac{V_{tot}}{Q_{tot}} \qquad (4)
$$

17 The catalyst specific loading φ (unit: g m⁻²) of the Pt/γ-Al₂O₃ washcoat catalyst deposited on the

18 reaction platelet is calculated as the catalyst mass (W<sub>cat</sub>) gained on the substrate (i.e., after

19 calcination) divided by the total surface area (S) of the reaction microchannel subject to the

20 coating.

$$
21 \quad \varphi = \frac{W_{cat}}{S} \qquad (5)
$$
1 The mean velocity of the gas mixture in the reaction microchannel is defined as the total
2 volumetric flow rate of the gas mixture divided by the total reaction microchannel cross-sectional
3 area (A).

4 $U = \frac{Q_{tot}}{A}$ (6)

5

6 **3. Results and discussion**

7 **3.1. CMC performance of the straight parallel channel microreactor**

8 In our previous study [46], the straight parallel channel microreactor of a larger dimension (317.5
9 mm length × 50 mm width × 3 mm height) with washcoated Pt/γ-Al₂O₃ catalyst has been
10 investigated for the CMC. Therein, the uniformity of γ-Al₂O₃ coating layer at optimized
11 synthesis conditions (pH = 3.5, 5 wt% PVA, 20 wt% γ-Al₂O₃ and particle size being around 3
12 µm) in the microchannels could be observed from SEM images. Moreover, the catalyst life time
13 test has been performed and the catalyst activity dropped from full conversion to 84.61% at 100
14 h [46], which could be a reference for the catalyst used in the current study given the same
15 protocol of synthesis. It has been reported that the introduction of promoters/additives (e.g.,
16 ZrO₂, CeO₂ [16,41,65]) into the catalyst could effectively prevent the catalyst deactivation by
17 increasing the temperature of thermal decomposition.

18

19 In order to gain a deeper understanding of the influence of among others the catalyst loading,
20 thickness, channel length and fluid distribution/collection structure on the methane conversion,
21 the similar type of microreactor, though with a relative smaller dimension (112.5 mm length ×
22 50 mm width × 3 mm height), has been prepared with Pt/γ-Al₂O₃ coating in this work and
23 examined under different reaction conditions.

24
### 1 3.1.1 Effect of the operating conditions

2 The methane conversion as a function of the reaction temperature (i.e., the oven temperature)
3 under different flow rates at an inlet oxygen to methane molar ratio ($\Phi$) of 2 is shown in Fig. 4.

4 For a given temperature, a higher methane conversion was obtained at a lower flow rate, due to
5 a longer residence time (c.f. Eq. (4)). At the beginning when the operating temperature is
6 relatively low (300 - 350 °C), the methane conversion has no significant difference between
7 different flow rates (150 - 500 mL min⁻¹). Under such circumstances, the reaction rate is slow
8 and (largely) controlled by kinetics (i.e., predominately determined by the reaction temperature
9 and catalyst mass). Thus, the methane conversion is at a very low level (< 3%) and the conversion
10 difference between each flow rate is not very discernible. As the temperature increased, the
11 methane conversion presented a significant increase and the light-off phenomenon occurred.

12 This indicates that the favorable coverage of the adsorbed methane and oxygen over the catalyst
13 surface facilitated by the temperature rise has been reached. Meanwhile, the significant increase
14 in the intrinsic kinetic rate and the catalytic activity during the light-off is likely attributed to the
15 rapid rise of the local temperature as well given the strongly exothermic reaction of methane
16 combustion. The light-off temperature corresponding to a 50% methane conversion (T₅₀ value)
17 relevant to conditions in Fig. 4a is shown in Table 2. When the operating temperature continued
18 to increase (e.g., > 450 °C), the reaction tended to be (more) governed by mass transfer, because
19 the kinetic rate rapidly increased (according to the Arrhenius equation), but the gas-catalyst mass
20 transfer rate was hardly to catch up with (given much less comparable increase in the species
21 diffusivity in the gas mixture and coating). It is worth noting that the decrement in the methane
22 conversion rendered a slowdown when comparing 150 mL min⁻¹ to 300 mL min⁻¹ and 300 mL
23 min⁻¹ to 500 mL min⁻¹ throughout the temperature from ca. 400 to 500 °C. This is in line with
24 T₅₀ value difference in Table 2.

25
Fig. 4b further compares the methane conversion as a function of the flow rate at a given temperature of 400 or 450 °C. The methane conversion presented an obvious decrease with the increasing total flow rate, as already revealed in Fig. 4a. It is worth noting that the decrease in the methane conversion at a sufficiently high flow rate tends to be insignificant. This could be due to the improved external mass transfer at high flow rates which compensate, to some extent, the conversion decrease (due to the shortened residence time). The conversion trends revealed here are in general similar to those for the straight parallel channel microreactor tested in our previous study [46]. These results confirm that the coating has been successfully applied onto the current microreactor and functioned well in the CMC.

<figure><img src="image_5.png" /></figure>

10
<figure><img src="image_6.png" /><figcaption>**Fig. 4.** Methane conversion as a function of (a) the temperature and (b) the total gas flow rate over the washcoated 3.5 wt% Pt/γ-Al₂O₃ catalyst in Reactor #1 (straight parallel channel microreactor). Conditions: T = 300 °C - 500 °C, Φ = 2, Qₜot = 110 - 500 mL min⁻¹.</figcaption></figure>

**Table 2**

T₅₀ value for the CMC experimental data presented in Fig. 4a.

<table><thead><tr><th>Total flow rate<br>(mL min<sup>-1</sup>)</th><th>T<sub>50</sub><sup>a, b</sup><br>(°C)</th></tr></thead><tbody><tr><td>150</td><td>451.8</td></tr><tr><td>300</td><td>487.6</td></tr><tr><td>500</td><td>516.0</td></tr></tbody></table>

Note: <sup>a</sup> T₅₀ indicates the temperature for reaching 50% methane conversion; Conditions are shown in Fig. 4a.
1 b Estimated from the polynomial regression line used for fitting the measured conversion curve
2 as a function of the temperature for different flow rates. The polynomial is in the form of $T_{50} = B_0 + B_1 X_{CH_4} + B_2 X_{CH_4}^2$, where $B_0, B_1$ and $B_2$ are the fitting constants.
3
4

### 3.1.2. Effect of catalyst loading and coating thickness

6 The reaction was further performed at various flow rates in Reactor #1 (straight parallel channel
7 microreactor) with different Pt loadings (1.5 wt%, 3.5 wt% and 5.0 wt% Pt/γ-Al₂O₃) by
8 maintaining almost a constant catalyst weight ($W_{cat}$= 0.25 ± 0.02 g). In more detail, during the
9 catalyst preparation only the concentration of Pt precursor solution was changed according to
10 the required Pt loading. The obtained results are presented in Fig. 5. At the same total gas flow
11 rate, a higher Pt loading (and with more Pt available in the catalyst) is accompanied by a higher
12 methane conversion. In other words, the overall reaction rate is significantly promoted by the
13 increased kinetic rate at higher Pt loadings. The methane conversion rendered an obvious drop
14 as increasing the total flow rate, due to the reduction of residence time. The extent of such
15 conversion decrease seems more pronounced especially at the highest catalyst loading studied
16 (5.0 wt% Pt/γ-Al₂O₃), indicating that the influence of the external mass transfer was more
17 pronounced therein given the highest kinetic rate involved. Furthermore, the external mass
18 transfer is related to many factors including the flow pattern, the diffusion rate, the channel
19 geometry, etc., by which the reactants have to be transported perpendicularly to the flow
20 direction from the center of microchannel to catalyst external surface [66]. A higher flow rate
21 (or larger Reynolds number) often leads to a higher external mass transfer rate [66]. However, a
22 higher flow rate also tends to incur a lower methane conversion due to the shorter residence time.
23 Thus, the decrement of methane conversion is generally dropped as shortened residence time as
24 presented in Fig. 5.
<figure><img src="image_7.png" /><figcaption>**Fig. 5.** Methane conversion as a function of the total gas flow rate over different Pt loadings in Reactor #1 (straight parallel channel microreactor). Conditions: T = 450 °C, Φ = 2, Q<sub>tot</sub> = 110 - 500 mL min<sup>-1</sup>, W<sub>cat</sub> = 0.25 ± 0.02 g.</figcaption></figure>

The internal mass transfer is mainly correlated to catalyst property including the catalyst thickness, tortuosity, porosity and reactant diffusivity, etc. [66]. To further investigate the influence of internal diffusion limitation (if present) on the microreactor performance, the coating thickness of 3.5 wt% Pt/γ-Al₂O₃ catalyst was varied in the experiment by applying different catalyst mass (W<sub>cat</sub>) onto the reaction microchannel walls in Reactor #1. That is, the higher the W<sub>cat</sub>, the thicker the coating layer and subsequently the higher the specific catalyst loading φ (Eq (5)). Microreactors with different values of φ varying from 39.45 to 90.39 g m<sup>-2</sup> in Reactor #1 have been prepared, and the catalyst layer visually presented a uniform and well-dispersed coating throughout the parallelized reaction microchannels. Fig. 6 illustrates that the methane conversion presents a maximum at about φ = 57.57 g m<sup>-2</sup> for a given temperature (more clearly seen at above 350 °C) and total flow rate. This phenomenon is probably due to a more or
less optimized match between the internal diffusion rate and the intrinsic kinetic rate at $\varphi = 57.57$ g m$^{-2}$. For lower $\varphi$ values (e.g., 39.45 g m$^{-2}$), this might be explained by the fact that although the thinner coating layer herein facilitated the internal diffusion due to the shorter path of diffusion, and the accompanied insufficient amount of active sites for the reaction (given the Pt weight proportionally decrease in the catalyst) has reduced the methane conversion. For higher $\varphi$ values (e.g., 78.30 and 90.39 g m$^{-2}$), the more significant internal diffusion resistance (longer path of diffusion) is likely present due to the increased thickness of the catalyst layer, resulting in a lower methane conversion (despite the higher intrinsic reaction rate due to Pt weight increase in the catalyst based on Arrhenius equation). To be more specific, the higher Thiele modulus and lower effectiveness factor could be obtained at the thicker catalyst coating, suggesting that the effect of internal mass transfer at thicker coating is more pronounced on the methane conversion. Thus the $\varphi = 57.57$ g m$^{-2}$ is expected to be the optimized compromise between the (internal) mass transfer and instinct kinetics. For a more quantitative prediction of the conversion trend as shown in Fig. 6, a good knowledge on the reaction kinetics and reactant diffusion within the coating layer (which depends primarily on the pore structure and morphology of the coating) need to be further acquired.
<figure><img src="image_8.png" /><figcaption>1

2 **Fig. 6.** Methane conversion as a function of the reaction temperature under different specific
catalyst loading (φ) in Reactor #1 (straight parallel channel microreactor). Conditions: T = 300
3
4 - 500 °C, Φ = 2, Q<sub>tot</sub> = 300 mL min<sup>-1</sup>, 3.5 wt% Pt/γ-Al<sub>2</sub>O<sub>3</sub> catalyst.</figcaption></figure>

5
6
### 3.1.3. Effect of reaction microchannel length

7
8
9
10
11
12
13
14
15
To gain further insights into the influence of external mass transfer on the reaction rate, the
methane conversion in experiments with the straight parallel channel microreactors of different
reaction microchannel lengths (50-275 mm) is compared in Fig. 7, where the microchannel width
& height is the same. The specific catalyst loading is kept more or less identical (φ = 82.5 ± 5 g
m⁻²). As displayed in Fig. 7, the methane conversion increased with increasing microchannel
length at 400 °C for a given mean velocity of the methane-air mixture (U; Eq. (6)). This is logical
given the increased residence time in the longer microchannel. A significant increase in the
microchannel length (from 60 to 275 mm) did not result in a conversion increase of similar
extent. This shows that the reaction is (at least) limited by the external mass transfer, given that
1 under internal diffusion limitation or kinetic control, the overall reaction rate should increase

2 linearly with the catalyst weight applied on the microchannel wall.

<figure><img src="image_9.png" /></figure>

3

**Fig. 7.** Methane conversion as a function of the reaction microchannel length at different mixture velocities (U). Conditions: T = 400 °C, Φ = 2, 3.5 wt% Pt/γ-Al₂O₃ catalyst, φ = 81.33, 78.30, and 87.37 g m⁻² for the microchannel lengths of 50, 60 and 275 mm, respectively, Data for 60 mm microchannel length is based on experiments with Reactor #1 (straight parallel channel microreactor). Data for 50 mm and 275 microchannel lengths are from the respective experiments with another microreactor of otherwise the same geometry and the microreactor used in our previous study [46] (but with different numbers of microchannels of the same width/height).

12

13
14
15
16
17
As already revealed in Fig. 4b, the methane conversion at a given temperature decreased with the increasing flow rate in the microreactor (Reactor #1), which is insignificant at sufficiently high flow rates (e.g., Qtot = 400 mL min⁻¹; corresponding to a U of 0.28 m/s). The results of Fig. 7 further suggest that at such high Qtot or U values, the methane conversion can be still improved by increasing the microchannel length (translated into a longer residence time). This also implies
1 that if the reaction is operated at a much higher $Q_{tot}$ (than 400 mL min⁻¹) in Fig. 4b, a significant
2 drop in the methane conversion is expected (i.e., the conversion does not level off at such high
3 flow rates).

4

### 3.1.4. Effect of bifurcated tree-like structure as distributor or collector

6 The flow distribution behaviour within the reaction microchannel network may have a great
7 impact on the microreactor performance in the CMC. In the experiments performed above in
8 Reactor #1, the bifurcated tree-like structure was used as the inlet fluid distributor whereas a
9 simple rectangular collector at the microchannel outlet (Fig. 2). Regarding the multi-scale tree-
10 like component, its relevant position with respect to the reaction microchannel will influence
11 flow distribution properties, as shown in the previous study [67]. In order to experimentally
12 verify the influence of such tree-like component on the catalytic performance, additional
13 experiments were performed with Reactor #1, by switching the microreactor outlet and inlet (i.e.,
14 by feeding the gas mixture at the microreactor outlet). Thus, the tree-like structure functioned
15 here as the product collector and the rectangular chamber as the fluid distributor.

16

17 Fig. 8 presents an interesting observation: for Reactor #1, the tree-like structure used as the outlet
18 product collector exhibited a higher methane conversion than that used as the inlet gas
19 distributor. This indicates a somewhat significant difference in the pressure (drop) between the
20 distributor case and collector case [67]. When the tree-like structure was used as the gas collector,
21 the pressure drop in the collector side is higher than the case of using the simple rectangular
22 chamber as the collector, and the pressure drop in the distributor side is lower. This means that
23 in the former case, the local pressures at the inlet and outlet of the reaction microchannel are
24 higher than those in the latter case, which is beneficial for a more uniform fluid distribution and
25 thus a better methane conversion. In other words, the lower pressure drop on the distributor side
26 tends to improve the uniform delivery of fluid into each reaction microchannel, and the higher
pressure at the reaction microchannel outlet would make the flow therein less affected by the
flow behavior within the outlet collector structure. These are also in line with the literature
reporting that less vortex was produced and less energy was dissipated when using such tree-like
structure as the collector compared with the distributor case [67].

Fig. 8 further shows that the conversion improvement in the former case is more pronounced at higher total flow rates. This is possibly because the bifurcated tree-like distributor is shown to be capable of guaranteeing a relatively uniform flow distribution among parallel channels/tubes at small flow rates [68,69]. Nevertheless, the inevitable flow distribution non-uniformity at high flow rates due to the increasing impact of inertial forces is an intrinsic character of such structure [68]. Thus, the tree-like component used as the outlet collector is able to provide a more uniform and stable flow distribution among parallel straight microchannels even under high flow rate conditions. The residence time difference across different reaction microchannels is thereby smaller, leading to a better usage of the catalysts and consequently higher methane conversion.

<figure><img src="image_10.png" /></figure>

15
**Fig. 8.** Influence of the location of the tree-like bifurcated component (i.e., used as the inlet gas distributor or outlet product collector) on the methane conversion in Reactor #1 (parallel straight channel microreactor). Conditions: T = 450 °C, Φ = 2, Q<sub>tot</sub> = 110 - 500 mL min<sup>-1</sup>, 3.5 wt% Pt/γ-Al<sub>2</sub>O<sub>3</sub>.

5

## 3.2. Comparison of the CMC performance in microreactors with different internal channel configurations

In the following sub-sections, the CMC performance of Reactor #1 (parallel straight channel microreactor) has been compared with other five plate-type microreactors with different internal channel configurations as shown in Fig. 2, with regard to the effect of the operating temperature, total gas flow rate and molar ratio of O<sub>2</sub> to CH<sub>4</sub> (Φ). In all reactors, the washcoated 3.5 wt% Pt/γ-Al<sub>2</sub>O<sub>3</sub> catalyst has the almost identical weight (0.25 ± 0.02 g; Table 1).

13

### 3.2.1. Effect of temperature on the performance of different microreactors

The methane conversion as a function of the reaction temperature in different microreactors is shown in Fig. 9a-c for different total flow rates (150 - 500 mL min<sup>-1</sup>) and Φ = 2. The same trend as presented in Fig. 4a was observed. At the lower temperature of 300 - 350 °C, the methane conversion in all Reactors #1-6 is very low (< ca. 5%) and has no obvious difference at different flow rates. This is because the reaction is mainly controlled by kinetics and mass transfer is not the main factor determining the reaction rate. As the temperature increased, the methane conversion presented a significant increase (especially at a lower flow rate), and the light-off phenomenon occurred. An obvious difference in the methane conversion was found among the different channel configurations of Reactors #1 - 6 at higher temperatures (> ca. 400 °C) when the (external) mass transfer gradually became the limiting factor, especially at a relatively low gas flow rate (e.g., 150 mL min<sup>-1</sup>; Fig. 9a). At a relatively high flow rate (e.g., 500 mL min<sup>-1</sup>;

15

16

17

18

19

20

21

22

23

24

25
1. Fig. 9c), the gap in the methane conversion between different reactors is narrowed, due to the

2. (largely) improved external mass transfer rate.

<figure><img src="image_11.png" /></figure>

3
<figure><img src="image_12.png" /></figure>

<figure><img src="image_13.png" /><figcaption>**Fig. 9.** Methane conversion as a function of the reaction temperature at the total flow rate of (a) 150 mL min⁻¹, (b) 300 mL min⁻¹ and (c) 500 mL min⁻¹ in plate-type microreactors with different channel configurations (Reactors #1 - 6). Conditions: T = 300 °C - 500 °C, Φ = 2, 3.5 wt% Pt/γ-Al₂O₃ catalyst. Other reactor details and catalyst properties are shown in Table 1.</figcaption></figure>

With regard to the different channel configurations in microreactors, the methane conversion ranked with the following order: double serpentine channels (Reactor #3) > obstacled straight
channels (Reactor #4) > vascular network (Reactor #6) > straight parallel channels (Reactor #1)
> cavity (Reactor #2) > meshed circuit (Reactor #5). This trend is more pronounced at a lower
total flow rate (e.g., 150 mL min⁻¹; Fig. 9a) and becomes insignificant at higher flow rate (e.g.,
500 mL min⁻¹; Fig. 9c). The straight parallel channel microreactor (Reactor #1) presented a
higher methane conversion than the cavity microreactor (Reactor #2). The relatively lower
methane conversion in Reactor #2 might be due to the negative effect of the non-uniform velocity
profile of the gas since only 1 flat cavity was involved (Fig. 2), even though it has the longest
mean residence time (c.f. Table 1). In other words, the catalyst coating around the cavity
experienced different contact times with the reactant mixture, leading to an overall conversion
decrease. In contrast, the gas flow distribution should be (much) improved in Reactor #1, given
the flow was distributed into 16 parallel straight microchannels contributing to a more uniform
contact time between the reactants and catalyst. The obstacle reactor (Reactor #4) is actually an
improved version of Reactor #1, with a slightly higher coating surface area and longer mean
residence time (c.f. Table 1). Besides, the structure of successive obstacles with the split-and-
recombine shape enabled to improve the gas mixing and further enhance the external mass
transfer compared to Reactor #1, especially at the mass transfer limited regime (at relatively
higher temperatures or low flow rates). It thus offered a slight increase in the methane conversion
under the same working conditions compared with Reactor #1. Among all the tested reactors,
the highest methane conversion has been achieved by using the double serpentine channel
microreactor (Reactor #3). This could be attributed primarily to a relatively higher coating
surface area (than all the others except Reactors #6) and relatively longer residence time (than
Reactors #1 and 4). In addition, the double serpentine structure could effectively promote the
gaseous mixing due to the formation of secondary flow on the channel cross-section [70], thus
improving their transfer to the coating external surface and contributing to a better conversion.
The vascular reactor (Reactor #6) has the highest coating surface area as well as a longer
residence time than the other reactors except Reactor #2 (Table 1). Despite these advantages, it
presented a methane conversion level in between those of Reactors #4 and 1 (Figs. 9a and b).

This is likely due to the non-uniform fluid distribution among the horizontal microchannel sections and thus an insufficient utilization of the coated catalyst. Some microchannels could be overloaded with gas mixture in the vascular network while others were underloaded, which might be related to the vascular structure causing a less uniform distribution of gas mixture. An improved design by carefully adjusting the resistance in the vertical diverging/converging channels may help solve this problem, based on the study of Tondeur et al. [61]. The lowest methane conversion was found in the meshed circuit microreactor (Reactor #5) under the same tested conditions. It has been reported that the flow pattern in the meshed circuit was similar to a continuous stirred tank reactor, and likely tended to be the divergent flow pattern even at low Reynolds numbers [71]. This indicates that the gas has the chance to flow towards different directions due to divergent flow appeared in the meshed circuit. This probably resulted in the non-uniform flow distribution and/or somewhat broad residence time distribution, and thus a lowest methane conversion.

Among the tested microreactors, the Reactor #3 (double serpentine microchannels) has the highest pressure drop due to the longer channel length and the higher velocity. But its total pressure drop was estimated to be smaller than 0.01 bar under the current experimental conditions (500 mL min⁻¹), lower than that present in fixed/packed bed reactors (e.g., mostly above 0.1 bar [72]) especially when fine catalyst powders are used.

In summary, guidelines regarding the optimal channel configurations in plate-type microreactors for the CMC should comprehensively consider the coating surface area, residence time (distribution), fluid distribution uniformity, and the efficient use of catalyst in order to achieve a desirable methane conversion. Further detailed investigation on the local flow and temperature
distribution characteristics by using simulation or optical visualization techniques would help give a clearer view and better understanding in this aspect.

Fig. 10 depicts the methane conversion as a function of the operating temperature at $\Phi = 6$ and a total flow rate of 150 mL min$^{-1}$ in Reactors #1 - 6. Similar conversion trends and the same microreactor performance ranking were observed as compared to the case of $\Phi = 2$ (Fig. 9a). A lower methane conversion was obtained at $\Phi = 6$ than $\Phi = 2$. A comparison of $T_{50}$ values at $\Phi = 2$ and 6 for different channel configurations is also given in Table 3. It is found that the $T_{50}$ value at $\Phi = 2$ is slightly lower than that at $\Phi = 6$ for each channel configuration tested. In other words, the methane conversion decreased a bit when $\Phi$ increased from 2 (stoichiometric ratio) to 6 (oxygen-rich). The competitive adsorption between methane and oxygen over the catalyst surface could be one of the main reasons, given the methane adsorption energy is higher than that of oxygen [73]. The adsorbed oxygen is likely to inhibit the weakly adsorbed methane in these light-off experiments, leaving less chances for methane being further oxidized [74]. Thus, the favorable fractional coverage of adsorbed reactants over the active sites is essential for achieving the desired methane conversion. However, caution must be taken when interpreting the trend because the difference in the $T_{50}$ value under $\Phi = 2$ and $\Phi = 6$ is only several degree Celsius (c.f. Table 3). An opposite trend may be observed when the operating procedures change as will be further discussed in the following sub-section.
<figure><img src="image_14.png" /><figcaption>**Fig. 10.** Methane conversion as a function of the temperature in plate-type microreactors with different channel configurations (Reactors #1 - 6). Conditions: T = 300 - 500 °C, Φ = 6, Q<sub>tot</sub> = 150 mL min<sup>-1</sup>, 3.5 wt% Pt/γ-Al<sub>2</sub>O<sub>3</sub> catalyst. Other reactor parameters and catalyst properties are shown in Table 1.</figcaption></figure>

6

7 Table 3

8 T<sub>50</sub> value for the CMC experimental data presented in Figs. 9a and 10.

<table><thead><tr><th rowspan="2">Reactor</th><th colspan="2">T<sub>50</sub> (°C)<sup>a, b</sup></th></tr><tr><th>Φ = 2</th><th>Φ = 6</th></tr></thead><tbody><tr><td>#1</td><td>460.0</td><td>462.8</td></tr><tr><td>#2</td><td>451.8</td><td>456.4</td></tr><tr><td>#3</td><td>437.5</td><td>443.5</td></tr><tr><td>#4</td><td>446.0</td><td>448.4</td></tr><tr><td>#5</td><td>470.6</td><td>472.2</td></tr></tbody></table>
1. $T_{50}$ indicates the temperature for reaching 50% methane conversion; Conditions are shown in Figs. 9a and 10.

3. Estimated from the polynomial regression line (cf. other details in Table 2).

4.

The results of Fig. 9 were obtained in Reactors #1 - 6 with almost identical catalyst mass. Due to the surface area difference, the specific catalyst loading ($\varphi$) differs to some extent per reactor (Table 1). This implies a coating thickness difference, which could lead to a different influence of internal diffusion on the overall reaction rate in each microreactor. To further shed light on this, some additional experiments were conducted in Reactors #1, 3 and 6, but now with almost the same specific catalyst loading ($\varphi = 90.5 \pm 0.5 \text{ g m}^{-2}$) by adjusting the coated catalyst mass accordingly. As shown in Fig. 11, the same conversion trend as a function of the temperature (at a total flow rate of 300 mL min$^{-1}$ and $\Phi = 2$) was observed in Reactors #1, #3, #6 as presented in Fig. 9b. Also, the methane conversion decrease follows the same order of Reactor #3 > Reactor #6 > Reactor #1, especially at relative higher temperatures (e.g., > 400 °C). Since in these reactors the coating thickness can be considered identical, the difference of methane conversion between each reactor is attributed to different channel geometries resulting in the different coating surface area, fluid mixing behaviour and residence time property therein.
<figure><img src="image_15.png" /><figcaption>**Fig. 11.** Methane conversion as a function of the temperature in plate-type microreactors with different channel configurations (Reactor #1, 3 and 6). Conditions: T = 300 - 500 °C, Φ = 2, Q<sub>tot</sub> = 300 mL min<sup>-1</sup>, 3.5 wt% Pt/γ-Al<sub>2</sub>O<sub>3</sub> catalyst, φ = 90.5 ± 0.5 g m<sup>-2</sup> (realized by adjusting the coated catalyst mass). Other reactor parameters and catalyst properties are shown in Table 1.</figcaption></figure>

### 3.2.2. Effect of molar ratio of O<sub>2</sub> to CH<sub>4</sub>

To investigate the effect of the inlet molar ratio of O₂ to CH₄ (Φ), the CMC experiments were further conducted in Reactors #1 - 6 at 450 °C and 150 mL min⁻¹. The results of different tested reactors are shown in Fig. 12, which reveal the same trend in terms of the microreactor performance rank as that shown in Fig. 9a. As Φ was increased from 0.5 up to 1.5, the methane conversion was improved gradually in each microreactor, which is due to the (severely) insufficient supply of oxygen limiting the adsorbed methane to be further converted. A relatively smaller difference in methane conversion presented in this region between different microreactors compared to that of Φ ≥ 1.5, and the (external) mass transfer shown a more pronounced effect. The highest methane conversion was found at ca. Φ = 1.5 (instead of 2 as the
stoichiometric ratio for the methane combustion) for each tested channel configuration. Oh et al. [75] also reported that operating at methane-rich conditions provided the best catalyst performance and increasing the oxygen concentration resulted in a sharp decrease in the catalyst activity. It could be thus due to firstly at $\Phi = 1.5$, there is an increased oxygen supply together with the consumption of methane in not only the combustion reaction, but also side reactions (such as methane partial oxidation and steam reforming) that require less amount of oxygen [76]. Moreover, a higher light-off temperature is required for the above-mentioned side reactions than for the methane combustion [77]. Such side reactions are more favoured to occur at higher temperature levels (e.g., at 450 °C relevant to Fig. 12). The existence of a variety of complex reactions under oxygen-lean conditions can be confirmed by the selectivities of CO₂, CO and H₂ as displayed in Fig. 13. For $\Phi \ge 2$, only CO₂ could be found as product ($S_{CO_2} = 100 \%$) while no CO or H₂ was detected in the product gas. For $\Phi < 2$, more amounts of CO and H₂ were produced, i.e., the selectivity of CO is ca. 2% - 4% and H₂ ca. 4% - 16% while the rest product is CO₂. Moreover, it seems that a balance between the adsorbed oxygen and methane has been achieved on the catalyst surface at ca. $\Phi = 1.5$ for the best methane conversion, which largely explains the significant conversion drop at $\Phi = 2$ or above. A way to further increase the methane conversion (e.g., to above 90%) at this stoichiometric ratio is by raising the reaction temperature (e.g., see Fig. 9a).

In our experiments, the proportion of CO₂ in the product at ca. $\Phi = 1.5$ is too high to achieve a desirable synthesis gas ratio of H₂/CO. It might be due to the reactant ratio and the catalyst used in this study that result in the overoxidation of products. The effects of the reactant ratio, pressure, temperature and mechanisms on the product selectivity have been evaluated in many earlier studies [78-80], but additional works remain to be performed in this respect for the purpose of syngas production by methane partial catalytic oxidation.
With $\Phi$ increasing from 2 to 6, the hysteresis effect plays an important role in determining the reaction conversion. The concentration hysteresis phenomenon has been demonstrated via switching the molar ratio of O₂ to CH₄ [81,82]. That is, a (relatively) high conversion state has been previously obtained at $\Phi$ = 0.5 - 1.5 could be maintained to some extent when further increasing the $\Phi$ value afterwards. Under such circumstances, the catalyst surface was primarily dominated by the adsorbed oxygen over the catalyst surface, and the oxygen concentration on the catalyst could be considered as constant in the present molar ratio experiment due to high conversion state. Given the less methane present in the feed and competing for a constant number of active sites, relatively more methane had the possibility to be adsorbed and converted, leading to a slight conversion increase at $\Phi$ > 2.

To further elucidate the observed conversion trend especially at $\Phi$ > 2, Table 4 compares the methane conversion in the light-off experiment (Figs. 9a and 10) and that in the molar ratio experiment here (Fig. 12) under the same working conditions. The results shown that the methane conversion displayed a slightly decrease as increasing O₂/CH₄ molar ratio, and on the contrary, the methane conversion turned out to be consistently higher in the molar ratio experiment. In the light-off experiment, as the experiment has been performed by increasing the temperature from 300 °C to 500 °C at a constant O₂/CH₄ molar ratio, the concentration hysteresis phenomenon is thus absent when switching from the low temperature (low conversion state) to high temperature (high conversion state). The favorable coverage between the adsorbed methane and oxygen has been achieved at $\Phi$ = 2, and the lower methane conversion at oxygen-rich environment ($\Phi$ = 6) could be reasonably explained by the competitive adsorption. The adsorbed oxygen could prevent the further oxidation of adsorbed methane over the catalyst surface due to the lower oxygen adsorption energy [73]. However, in the molar ratio experiment, a higher methane conversion obtained at $\Phi$ = 6 can be explained by the concentration hysteresis phenomenon present (procedurally/prehistorically related), as has been explained above.
1

2 Furthermore, based on the study of Pakharukov et al. [82], it could be further explained by the

3 difference in Pt electronic state of Pt/γ-Al₂O₃ catalyst, which is attributed to variation in O/Pt

4 surface ratio caused by changing the oxygen concentration. It has been demonstrated [83,84] that

5 the C-H bond activation over the oxygen-depleted Pt surface (bare Pt or O*-Pt pair sites),

6 compared to the oxygen-covered Pt surface (O*-O* pair sites), will lead to the remarkable

7 increase in the catalyst activity for the methane conversion.

8

9 The hysteresis effect has been supposed to first activate catalyst by varying catalyst prehistory

10 (e.g., O₂ concentration, reaction temperature) in order to achieve the ignition, and subsequently

11 to maintain the high conversion rate when switching back to the low conversion regime, [81]. In

12 addition to the concentration hysteresis, the temperature hysteresis effect has also been observed

13 in our experiments. Table 5 compares the CMC performance of Reactor #1 between the

14 ascending temperature (ignition process) and the descending temperature procedure (extinction

15 process). Methane conversion being higher during the extinction process than that during ignition

16 process has been observed. Furthermore, the higher conversion state has been first triggered at

17 the higher temperature (e.g., above 450 °C) and can still be maintained even after switching to

18 the low temperature (e.g., 300 °C), clearly showing that the catalytic activity is not identical

19 during the ignition or extinction process. This observation is in line with the published literature

20 [81,85-87], providing an alternative approach to take advantage of heat released from the

21 reaction.

22
<figure><img src="image_16.png" /></figure>

1

2 **Fig. 12.** Methane conversion as a function of Φ in various plate-type microreactors.

3 Conditions: T = 450 °C, Φ = 0.5 - 6 (sequentially changed from small to large values in the experiments), Q<sub>tot</sub> =150 mL min<sup>-1</sup>. 3.5 wt% Pt/γ-Al<sub>2</sub>O<sub>3</sub>, Other reactor parameters and catalyst properties are shown in Table 1.

<figure><img src="image_17.png" /></figure>

6

7 **Fig. 13.** Selectivities of H<sub>2</sub>, CO, CO<sub>2</sub> as a function of Φ in the straight parallel channel microreactor (Reactor #1). Other conditions are the same as in Fig. 12.

8
1

Table 4

Comparison of the methane conversion in different reactors between the light-off experiment and the molar ratio experiment.

4

<table><thead><tr><th rowspan="2">Reactor</th><th colspan="2">X<sub>CH<sub>4</sub></sub> in light-off experiment <sup>a</sup></th><th colspan="2">X<sub>CH<sub>4</sub></sub> in molar ratio experiment <sup>b</sup></th></tr><tr><th>Φ = 2</th><th>Φ = 6</th><th>Φ = 2</th><th>Φ = 6</th></tr></thead><tbody><tr><td>#1</td><td>38.6</td><td>36.3</td><td>45.6</td><td>48.7</td></tr><tr><td>#2</td><td>48.4</td><td>46.6</td><td>51.6</td><td>54.9</td></tr><tr><td>#3</td><td>60.4</td><td>57.3</td><td>62.3</td><td>65.8</td></tr><tr><td>#4</td><td>54.8</td><td>52.6</td><td>56.2</td><td>58.4</td></tr><tr><td>#5</td><td>28.6</td><td>26.2</td><td>37.7</td><td>40.0</td></tr><tr><td>#6</td><td>52.1</td><td>48.4</td><td>53.3</td><td>56.2</td></tr></tbody></table>

7

Table 5

Comparison of ascending temperature and descending temperature procedures in CMC.

9

<table><thead><tr><th></th><th>X<sub>CH<sub>4</sub></sub> in ascending temperature <sup>a</sup><br>(ignition process)</th><th>X<sub>CH<sub>4</sub></sub> in descending temperature <sup>b</sup><br>(extinction process)</th></tr></thead><tbody><tr><td>300 °C</td><td>0.00</td><td>14.86</td></tr><tr><td>400 °C</td><td>22.04</td><td>40.23</td></tr></tbody></table>

Note: <sup>a</sup> Data from Fig. 9, ascending temperature operation, reactor #1, Q<sub>tot</sub> = 150 mL min<sup>-1</sup>, Φ = 2, 3.5 wt% Pt/γ-Al<sub>2</sub>O<sub>3</sub> catalyst. <sup>b</sup> Descending temperature operation, naturally cooling down, other conditions the same as ascending temperature operation.

13

<footnote><sup>a</sup> Conditions are shown in Figs. 9a and 10, T = 450 °C;</footnote>

<footnote><sup>b</sup> Conditions are shown in Fig. 12.</footnote>
1 **3.2.3. Effect of total flow rate**

2 The methane conversion as a function of the total flow rate in various plate-type microreactors
3 (Reactors #1 - 6) is further displayed in Fig. 14. Generally, the methane conversion remarkably
4 decreased when the total flow rate increased, mainly due to the decreased mean residence time
5 ($\tau$; cf. Eq. (4)) in every tested reactor. With the increased flow rate, the methane conversion
6 decrease tended to slow down and became insignificant especially at sufficiently high flow rates
7 (ca. 300 mL min$^{-1}$ for Reactors #1 and 5; ca. 450 mL min$^{-1}$ for Reactors #2 - 4 and 6). This could
8 be due to the (much) improved external mass transfer at such high flow rates that contributed
9 positively to the methane conversion. To improve the conversion at such high flow rates, the
10 catalyst loading/mass and/or the operating temperature should be increased, so that the kinetic
11 rate increases and contributes more to the overall reaction rate.

<figure><img src="image_18.png" /><figcaption>**Fig. 14.** Methane conversion as a function of the total gas flow rate in various plate-type microreactors. Conditions: T = 450 °C, Φ = 2, 3.5 wt% Pt/γ-Al₂O₃. Other reactor parameters and catalyst properties are shown in Table 1.</figcaption></figure>

12

13

14

15

16
# 4. Conclusions and prospects

An experimental investigation has been firstly performed for the CMC in the straight parallel channel microreactor (made of FeCrAlloy) with washcoated Pt/γ-Al₂O₃ catalyst. At increasing operating temperature, a remarkable increase in the methane conversion was observed. An obvious decrease in the methane conversion was found upon increasing the total flow rate due to the reduction of the residence time, and such conversion decrease is insignificant at sufficiently high flow rates due to the improved external mass transfer. The existence of an optimal catalyst specific loading (φ = 57.6 g m⁻²) rendering the highest methane conversion has been observed. A higher specific loading (e.g., φ = 90.39 g m⁻²) hinders the mass transfer due to a more significant internal diffusion limitation in the thicker coating layer, whereas a lower specific loading (e.g., φ = 39.45 g m⁻²) could not provide sufficient active sites, both leading to a lower methane conversion. In this type of microreactor, the bifurcated tree-like structural component placed at the outlet as the gas collector has presented a higher methane conversion than that using it as the inlet gas distributor. This could be attributed to the presence of a more uniform gas distribution in the former case.

Other five type of microreactors with different channel configurations (comprising cavity, double serpentine microchannels, obstacled microchannels, meshed circuit and vascular network) have also been tested. Their performance in the CMC has been compared with that of the straight parallel channel microreactor, with regard to the influence of various factors including the reaction temperature, inlet molar ratio of O₂ to CH₄ and total flow rate. The results suggest that the double serpentine channel microreactor exhibited the highest methane conversion, whereas the meshed circuit microreactor showed the lowest conversion. The effective usage of the coating surface area, a more uniform flow distribution and the sufficient residence time are essential factors that contribute to the higher methane conversion. A higher methane conversion was obtained in the molar ratio experiment than that in the light-off
experiment under otherwise identical conditions, mainly due to the concentration hysteresis phenomenon. The highest methane conversion can be achieved in each microreactor at $\Phi = 1.5$, because of the additional involvement of other complex side reactions under oxygen-lean conditions. The methane conversion for each tested reactor has appeared to decrease with the total gas flow rate increase due to the reduced residence time, despite the gradually improved external mass transfer.

Moreover, the partial methane oxidation under molar ratios of $O_2:CH_4 < 2$ has been briefly mentioned in this paper, over the same $Pt/\gamma-Al_2O_3$ catalyst. The carbon deposition over the catalyst mainly due to thermal decomposition of $CH_4$ and CO at high temperatures will easily result in the catalyst deactivation. Special attention is needed on the influence of different $CH_4/O_2$ ratios on the product selectivity distribution for downstream processes, such that an $H_2/CO$ ratio of 2 is favorable for further methanol synthesis.

In addition, a close investigation into the fluid distribution and temperature characteristics in different channel configurations of the plate-type microreactor using CFD simulation or optical visualization techniques could be the next research step to further elucidate the reaction performance. Better understanding of the conjugated transport phenomena and more experimental data available would be beneficial to develop new mathematic models predicting the CMC catalytic performance in microreactors.

This study is expected to provide useful guidelines regarding the optimal channel configurations in plate-type microreactors for the CMC, in view of future scaling-up in practice. Some difficulties remain to be overcome for the industrial application of CMC by using microreactors. Firstly, the catalyst stability (without detachment) in the microchannel and the selection of substrate material are the requirements related to specific industrial application. Secondly,
extending the catalyst lifetimes (especially for noble metal as active component) is particularly
essential to improve the cost-effectiveness of the CMC technology. More deep investigations for
catalyst regeneration and replacement methods are therefore needed. Finally, more efficient
reaction heat recovery is subject to further studies in order to improve the global energy
efficiency of the CMC system, for instance, the CMC coupling with endothermic reactions.

6

7 Acknowledgements

This work was supported by the Region Pays de la Loire (Chaire Connect Talent ODE) from
French side; and the Ubbo-Emmius Fund 2015 from University of Groningen and the start-up
package in the area of green chemistry and technology (for Jun Yue) from the Netherlands' side.

9

10

The authors thank the technical support from Gwenaël Biotteau and Nicolas Lefevre at the
University of Nantes, Erwin Wilbers, Marcel de Vries, Léon Rohrbach and Daniël Vreugdenhil
at the University of Groningen.

14

15 Appendix A

16 *Design and sizing of the inlet fluid distributor and outlet fluid collector structures*

17 The design of tree-like bifurcated structure is based on the literature [88,89], and the detailed
18 configuration is presented in Fig. A.1a. The gas mixture flows through the cascade bifurcations,
19 allowing a (uniform) splitting into the downstream channels. The sharp fillet structures were
20 designed, but it tends to be the concentric fillet structure in actual plate-type multichannel
21 reactors due to the manufacture limitation. It is reported that flow separation and recirculation
22 could occur in the sharp fillet structure, and the concentric fillet structure without backflow is
23 recommended [90]. The detailed geometric parameters of the tree-like bifurcation are presented
24 in Table A.1. The design of channel width at the index j ($w_j$) is based on the following equation:

18

19

20

21

22

23

24

25

$$w_j = w_n/r^{n-j} \tag{A.1}$$
1 where $r$ is equal to 0.77 and $w_n$ (width of the channel at the last index $n$) is equal to 1.50 mm

2 [89]). The calculation of $l_{j,tot}$ (channel length between indices $j$ and $j+1$) is based on the equation:

$$3 \qquad l_{j,tot}/w_j = 0.05Re_j \qquad (A.2)$$

4 where $Re_j$ is the Reynolds number at each index $j$ and calculated as

$$5 \qquad Re_j = \frac{\rho \times \frac{Q_{tot}/2^j}{w_j h} \times \frac{2hw_j}{(h+w_j)}}{\mu} \qquad (A.3)$$

6 Thus, there is

$$7 \qquad Re_{j+1}/Re_j = \frac{\frac{h+w_n}{r^{n-j}}}{2(h+w_n/r^{n-j-1})} \qquad (A.4)$$

8 In order to save more space for the reaction microchannel, a simple rectangular structure was

9 utilized as the gas collector for Reactors #1, 2, 4 and 5, with its dimensions presented in Fig.

10 A.1c. Simple bilateral divergent/convergent channels (Fig. A.1b) were used for both the inlet

11 and outlet structures of Reactors #3 and 6.

<figure><img src="image_19.png" /></figure>

<figure><img src="image_20.png" /></figure>

<figure><img src="image_21.png" /></figure>
1 **Fig. A.1.** Geometry and dimensions of the inlet/outlet structures: (a) the inlet distributor for
2 Reactors #1, 2, 4, 5; (b) the inlet distributor and outlet collector for Reactors #3 and 6; (c) the
3 outlet collector for Reactors #1, 2, 4 and 5.

4

5 Table A.1

6 Geometric parameters of tree-like bifurcated structure.

<table><thead><tr><th>Index number j</th><th>w<sub>j</sub><br>(mm)</th><th>l<sub>j,tot</sub><br>(mm)</th><th>w<sub>j</sub>/w<sub>j-1</sub><br>(-)</th><th>l<sub>j,tot</sub>/w<sub>j</sub><br>(-)</th><th>l<sub>j,1</sub><br>(mm)</th><th>Re<sub>j</sub><sup>a</sup><br>(-)</th></tr></thead><tbody><tr><td>Index 1</td><td>3.48</td><td>3.25</td><td>0.77</td><td>0.93</td><td>27.48</td><td>18.64</td></tr><tr><td>Index 2</td><td>2.68</td><td>1.52</td><td>0.77</td><td>0.57</td><td>14.68</td><td>11.35</td></tr><tr><td>Index 3</td><td>2.06</td><td>0.70</td><td>0.77</td><td>0.34</td><td>8.06</td><td>6.82</td></tr><tr><td>Index 4</td><td>1.50</td><td>0.30</td><td>0.73</td><td>0.20</td><td>4.50</td><td>4.03</td></tr></tbody></table>

7 Note: <sup>a</sup> Re<sub>j</sub> value is based on a total flow rate of 300 mL min<sup>-1</sup>.

8

9

10 References

11 [1] C. Zou, Q. Zhao, G. Zhang, B. Xiong, Energy revolution: from a fossil energy era to a new energy era, Nat. Gas Ind. B. 3 (2016) 1-11.
12 https://doi.org/10.1016/j.ngib.2016.02.001.

13 [2] G. Karavalakis, T.D. Durbin, M. Villela, J.W. Miller, Air pollutant emissions of light-duty vehicles operating on various natural gas compositions, J Nat. Gas Sci. Eng., 4 (2012) 8-16. https://doi.org/10.1016/j.jngse.2011.08.005.

14 [3] International Energy Agency (IEA). Natural Gas Information. https://www.iea.org/reports/natural-gas-information-2019. [Acessed 20 April 2020].

15 [4] M. Mehrpooya, M. Khalili, M.M.M. Sharifzadeh, Model development and energy and exergy analysis of the biomass gasification process (Based on the various biomass sources), Renew Sust Energ Rev. 91 (2018) 869-887. https://doi.org/10.1016/j.rser.2018.04.076.

16 [5] F. Ahmad, E.L. Silva, M.B.A. Varesche, Hydrothermal processing of biomass for anaerobic digestion-A review, Renew Sust Energ Rev. 98 (2018) 108-124. https://doi.org/10.1016/j.rser.2018.09.008.

17 [6] D. Gielen, Global energy transformation - a roadmap to 2050, International Renewable Energy Agency (IRENA), 2018.
[7] A.W. Petrov, D. Ferri, F. Krumeich, M. Nachtegaal, J.A. van Bokhoven, O. Kröcher, Stable complete methane oxidation over palladium based zeolite catalysts, Nat. Commun. 9 (2018) 2545-2553. http://dx.doi.org/10.1038/s41467-018-04748-x.

[8] I. Manisalidis, E. Stavropoulou, A. Stavropoulos, E. Bezirtzoglou, Environmental and health impacts of air pollution: a review, Frontiers in Public Health. 8 (2020).

[9] Official Journal of the European Union. Commission regulation (EU) No. 813/2013, http://data.europa.eu/eli/reg/2013/813/oj; [Accessed 2 August 2013].

[10] J. Chen, H. Arandiyan, X. Gao, J. Li, Recent advances in catalysts for methane combustion, Catal. Surv. Asia. 19 (2015) 140-171. https://doi.org/10.1007/s10563-015-9191-5.

[11] J.H. Lee, D.L. Trimm, Catalytic combustion of methane, Fuel Process. Technol. 42 (1995) 339-359. http://doi.org/10.1016/0378-3820(94)00091-7.

[12] R.J. Farrauto, Low-temperature oxidation of methane, Science. 337 (2012) 659-660. https://doi.org/10.1126/science.1226310.

[13] J. Yang, Y. Guo, Nanostructured perovskite oxides as promising substitutes of noble metals catalysts for catalytic combustion of methane, Chin. Chem. Lett. 29 (2018) 252-260. https://doi.org/10.1016/j.cclet.2017.09.013.

[14] A. Cruellas, T. Melchiori, F. Gallucci, M. van Sint Annaland, Advanced reactor concepts for oxidative coupling of methane, Catal. Rev. 59 (2017) 234-294. https://doi.org/10.1080/01614940.2017.1348085.

[15] L. He, Y. Fan, J. Bellettre, J. Yue, L. Luo, A review on catalytic methane combustion at low temperatures: Catalysts, mechanisms, reaction conditions and reactor designs, Renew Sust Energ Rev. (2019) 109589.

[16] R.J. Farrauto, M. Hobson, T. Kennelly, E. Waterman, Catalytic chemistry of supported palladium for combustion of methane, Appl. Catal. A. 81 (1992) 227-237. https://doi.org/10.1016/0926-860X(92)80095-T.

[17] D. Wang, J. Gong, J. Luo, J. Li, K. Kamasamudram, N. Currier, A. Yezerets, Distinct reaction pathways of methane oxidation on different oxidation states over Pd-based three-way catalyst (TWC), Appl. Catal. A. 572 (2019) 44-50. https://doi.org/10.1016/j.apcata.2018.12.022.

[18] I.E. Beck, V.I. Bukhtiyarov, I.Y. Pakharukov, V.I. Zaikovsky, V.V. Kriventsov, V.N. Parmon, Platinum nanoparticles on Al₂O₃: Correlation between the particle size and activity in total methane oxidation, J. Catal. 268 (2009) 60-67. https://doi.org/10.1016/j.jcat.2009.09.001.

[19] C.-P. Hwang, C.-T. Yeh, Platinum-oxide species formed on progressive oxidation of platinum crystallites supported on silica and silica-alumina, J. Catal. 182 (1999) 48-55. http://dx.doi.org/10.1006/jcat.1998.2304.

[20] E. Garbowski, C. Feumi-Jantou, N. Mouaddib, M. Primet, Catalytic combustion of methane over palladium supported on alumina catalysts: Evidence for reconstruction of particles, Appl. Catal. A. 109 (1994) 277-291. https://doi.org/10.1016/0926-860X(94)80124-X.

[21] C.A. Müller, M. Maciejewski, R.A. Koeppel, A. Baiker, Combustion of methane over palladium/zirconia: effect of Pd-particle size and role of lattice oxygen, Catal. Today. 47 (1999) 245-252. https://doi.org/10.1016/S0920-5861(98)00305-8.

[22] D. Ciuparu, E. Altman, L. Pfefferle, Contributions of lattice oxygen in methane combustion over PdO-Based catalysts, J. Catal. 203 (2001) 64-74. http://doi.org/10.1006/jcat.2001.3331.

[23] C.A. Müller, M. Maciejewski, R.A. Koeppel, R. Tschan, A. Baiker, Role of lattice oxygen in the combustion of methane over PdO/ZrO₂: combined pulse TG/DTA and MS study with ¹⁸O-labeled catalyst, The Journal of Physical Chemistry. 100 (1996) 20006-20014. https://doi.org/10.1021/jp961903a.
[24] S. Seimanides, M. Stoukides, Catalytic oxidation of methane on polycrystalline palladium supported on stabilized zirconia, J. Catal. 98 (1986) 540-549.
https://doi.org/10.1016/0021-9517(86)90342-8.

[25] P. Jodłowski, R. Jędrzejczyk, D. Chlebda, A. Dziedzicka, Ł. Kuterasiński, A. Gancarczyk, M. Sitarz, Non-noble metal oxide catalysts for methane catalytic combustion: sonochemical synthesis and characterisation, Nanomaterials. 7 (2017) 174.
https://doi.org/10.3390/nano7070174.

[26] Y.S. Seo, S.P. Yu, S.J. Cho, K.S. Song, The catalytic heat exchanger using catalytic fin tubes, Chem. Eng. Sci. 58 (2003) 43-53.
http://dx.doi.org/10.1016/S0009-2509(02)00475-X.

[27] M. O'Connell, G. Kolb, R. Zapf, Y. Men, V. Hessel, Bimetallic catalysts for the catalytic combustion of methane using microreactor technology, Catal. Today. 144 (2009) 306-311.
https://doi.org/10.1016/j.cattod.2008.10.053.

[28] Q. Dong, S. Zhang, Z. Duan, Q. Zhou, An energy analysis of the catalytic combustion burner, Heat. Technol. Energy Eff. (2006).
http://hdl.handle.net/1969.1/5530.

[29] D. X.Chun, Applied studies on methane catalytic combustion development of a natural gas premixed catalytic burner and boiler for household applications. 2005; Sichuan University. Thesis in chinese.

[30] R. Burch, P.K. Loader, Investigation of Pt/Al₂O₃ and Pd/Al₂O₃ catalysts for the combustion of methane at low concentrations, Appl. Catal., B. 5 (1994) 149-164.
http://dx.doi.org/10.1016/0926-3373(94)00037-9.

[31] M. Lyubovsky, L.L. Smith, M. Castaldi, H. Karim, B. Nentwick, S. Etemad, R. LaPierre, W.C. Pfefferle, Catalytic combustion over platinum group catalysts: fuel-lean versus fuel-rich operation, Catal. Today. 83 (2003) 71-84.
http://dx.doi.org/10.1016/s0920-5861(03)00217-7.

[32] Z. Wang, J. Deng, Y. Liu, H. Yang, S. Xie, Z. Wu, H. Dai, Three-dimensionally ordered macroporous CoCr₂O₄-supported Au-Pd alloy nanoparticles: Highly active catalysts for methane combustion, Catal. Today. 281, Part 3 (2017) 467-476.
https://doi.org/10.1016/j.cattod.2016.05.035.

[33] M. Halabi, M. De Croon, J. Van Der Schaaf, P. Cobden, J. Schouten, Intrinsic kinetics of low temperature catalytic methane-steam reforming and water-gas shift over Rh/CeαZr₁-αO₂ catalyst, Appl. Catal. A. 389 (2010) 80-91.
https://doi.org/10.1016/j.apcata.2010.09.005.

[34] H. Hao, Z. Liu, F. Zhao, W. Li, Natural gas as vehicle fuel in China: A review, Renew Sust Energ Rev. 62 (2016) 521-533.
https://doi.org/10.1016/j.rser.2016.05.015.

[35] S. Su, X. Yu, A 25 kWe low concentration methane catalytic combustion gas turbine prototype unit, Energy. 79 (2015) 428-438.
http://dx.doi.org/10.1016/j.energy.2014.11.031.

[36] M. Bhagiyalakshmi, R. Anuradha, S.D. Park, T.S. Park, W.S. Cha, H.T. Jang, Effect of bimetallic Pt-Rh and trimetallic Pt-Pd-Rh catalysts for low temperature catalytic combustion of methane, Bull. Korean Chem. Soc. 31 (2010) 120-124.
https://doi.org/10.5012/bkcs.2010.31.01.120.

[37] P. Gélin, M. Primet, Complete oxidation of methane at low temperature over noble metal based catalysts: a review, Appl. Catal., B. 39 (2002) 1-37.
http://dx.doi.org/10.1016/S0926-3373(02)00076-0.

[38] S.B. T.V. Choudhary, V.R. Choudhary, Catalysts for combustion of methane and lower alkanes review, Appl. Catal. A. 234 (2002) 1-23.
https://doi.org/10.1016/S0926-860X(02)00231-4.

[39] K. Persson, K. Jansson, S. Jaras, Characterisation and microstructure of Pd and bimetallic Pd-Pt catalysts during methane oxidation, J. Catal. 245 (2007) 401-414.
https://doi.org/10.1016/j.jcat.2006.10.029.
[40] K. Persson, A. Ersson, A.M. Carrera, J. Jayasuriya, R. Fakhrai, T. Fransson, S. Järås, Supported palladium-platinum catalyst for methane combustion at high pressure, Catal. Today. 100 (2005) 479-483. https://doi.org/10.1016/j.cattod.2004.08.018.

[41] G. Pecchi, P. Reyes, T. López, R. Gómez, Pd-CeO₂ and Pd-La₂O₃/alumina-supported catalysts: their effect on the catalytic combustion of methane, J. Non-Cryst. Solids. 345 (2004) 624-627. https://doi.org/10.1016/j.jnoncrysol.2004.08.110.

[42] S. Specchia, E. Finocchio, G. Busca, G. Saracco, V. Specchia, Effect of S-compounds on Pd over LaMnO₃·2ZrO₂ and CeO₂·2ZrO₂ catalysts for CH₄ combustion, Catal. Today. 143 (2009) 86-93. https://doi.org/10.1016/j.cattod.2008.10.035.

[43] R. Zapf, C. Becker-Willinger, K. Berresheim, H. Bolz, H. Gnaser, V. Hessel, G. Kolb, P. Löb, A.K. Pannwitt, A. Ziogas, Detailed characterization of various porous alumina-based catalyst coatings within microchannels and their testing for methanol steam reforming, Chem. Eng. Res. Des. 81 (2003) 721-729. https://doi.org/10.1205/026387603322302887.

[44] N.R. Peela, A. Mubayi, D. Kunzru, Washcoating of γ-alumina on stainless steel microchannels, Catal. Today. 147 (2009) S17-S23. https://doi.org/10.1016/j.cattod.2009.07.026.

[45] M. Liauw, M. Baerns, R. Broucek, O. Buyevskaya, J. Commenge, J. Corriou, L. Falk, K. Gebauer, H. Hefter, O. Langer, Microreaction Technology: Industrial Prospects. Springer, 2000, pp. 224-234.

[46] L. He, Y. Fan, L. Luo, J. Bellettre, J. Yue, Preparation of Pt/γ-Al₂O₃ catalyst coating in microreactors for catalytic methane combustion, Chem. Eng. J. 380 (2020) 122424. https://doi.org/10.1016/j.cej.2019.122424.

[47] X. Xu, H. Vonk, A. Cybulski, J.A. Moulijn, Alumina washcoating and metal deposition of ceramic monoliths, Stud. Surf. Sci. Catal., Elsevier, 1995, pp. 1069-1078.

[48] K. Haas-Santo, M. Fichtner, K. Schubert, Preparation of microstructure compatible porous supports by sol-gel synthesis for catalyst coatings, Appl. Catal. A. 220 (2001) 79-92. https://doi.org/10.1016/S0926-860X(01)00714-1.

[49] B. Schmidt, M. Liauw, Mass transfer limitations in microchannel reactors, Catal. Today. 110 (2005) 15-25. https://doi.org/10.1016/j.cattod.2005.09.019.

[50] P.L. Mills, D.J. Quiram, J.F. Ryley, Microreactor technology and process miniaturization for catalytic reactions-A perspective on recent developments and emerging technologies, Chem. Eng. Sci. 62 (2007) 6992-7010. https://doi.org/10.1016/j.ces.2007.09.021.

[51] J. Ganley, K. Riechmann, E.G. Seebauer, R. Masel, Porous anodic alumina optimized as a catalyst support for microreactors, J. Catal. 227 (2004) 26-32. https://doi.org/10.1016/j.jcat.2004.06.016.

[52] J.C. Ganley, E.G. Seebauer, R.I. Masel, Porous anodic alumina microreactors for production of hydrogen from ammonia, AlChE J. 50 (2004) 829-834. https://doi.org/10.1002/aic.10078.

[53] C.M. Miesse, R.I. Masel, C.D. Jensen, M.A. Shannon, M. Short, Submillimeter-scale combustion, AlChE J. 50 (2004) 3206-3214. https://doi.org/10.1002/aic.10271.

[54] K.F. Jensen, Microreaction engineering-is small better?, Chem. Eng. Sci. 56 (2001) 293-303. https://doi.org/10.1016/S0009-2509(00)00230-X.

[55] M.H. Exchangers, Microreactors: new technology for modern chemistry, Weinheim: Wiley/VCH, 2000.

[56] M. Mundhwa, R.D. Parmar, C.P. Thurgood, A comparative parametric study of a catalytic plate methane reformer coated with segmented and continuous layers of combustion catalyst for hydrogen production, J. Power Sources. 344 (2017) 85-102. http://dx.doi.org/10.1016/j.jpowsour.2017.01.082.

[57] M. Mundhwa, C.P. Thurgood, Numerical study of methane steam reforming and methane combustion over the segmented and continuously coated layers of catalysts in a plate
1 reactor, Fuel Process. Technol. 158 (2017) 57-72.

2 http://doi.org/10.1016/j.fuproc.2016.12.002.

3 [58] R.C. Ramaswamy, P.A. Ramachandran, M.P. Duduković, Coupling exothermic and

4 endothermic reactions in adiabatic reactors, Chem. Eng. Sci. 63 (2008) 1654-1667.

5 http://dx.doi.org/10.1016/j.ces.2007.11.010.

6 [59] M.S. Mettler, G.D. Stefanidis, D.G. Vlachos, Scale-out of microreactor stacks for

7 portable and distributed processing: coupling of exothermic and endothermic processes

8 for syngas production, Ind. Eng. Chem. Res. 49 (2010) 10942-10955.

9 https://doi.org/10.1021/ie100459b.

10 [60] H. Mei, C. Li, S. Ji, H. Liu, Modeling of a metal monolith catalytic reactor for methane

11 steam reforming-combustion coupling, Chem. Eng. Sci. 62 (2007) 4294-4303.

12 http://dx.doi.org/10.1016/j.ces.2007.05.011.

13 [61] D. Tondeur, Y. Fan, J.-M. Commenge, L. Luo, Uniform flows in rectangular lattice

14 networks, Chem. Eng. Sci. 66 (2011) 5301-5312.

15 https://doi.org/10.1016/j.ces.2011.07.027.

16 [62] V. Meille, S. Pallier, G. Santacruzbustamante, M. Roumanie, J. Reymond, Deposition of

17 γ-Al₂O₃ layers on structured supports for the design of new catalytic reactors, Appl. Catal.

18 A. 286 (2005) 232-238. https://doi.org/10.1016/j.apcata.2005.03.028.

19 [63] L. Giani, C. Cristiani, G. Groppi, E. Tronconi, Washcoating method for Pd/γ-Al₂O₃

20 deposition on metallic foams, Appl. Catal., B. 62 (2006) 121-131.

21 https://doi.org/10.1016/j.apcatb.2005.07.003.

22 [64] P.A.R. Cebollada, E. Garcia Bordejé, Optimisation of physical properties of γ-alumina

23 coating microreactors used for the growth of a carbon nanofiber layer, Chem. Eng. J. 149

24 (2009) 447-454. https://doi.org/10.1016/j.cej.2009.02.016.

25 [65] R.J. Farrauto, T. Kennelly, E.M. Waterman, M.C. Hobson Jr, Process conditions for

26 operation of ignition catalyst for natural gas combustion. 1993; Google Patents.

27 [66] H.S. Fogler, Elements of chemical reaction engineering, Pearson Higher Education &

28 Professional Group, 2013.

29 [67] Y. Fan, R. Boichot, T. Goldin, L. Luo, Flow distribution property of the constructal

30 distributor and heat transfer intensification in a mini heat exchanger, AlChE J. 54 (2008)

31 2796-2808. https://doi.org/10.1002/aic.11597.

32 [68] P. Zhou, D. Tarlet, Y. Fan, X. Hu, L. Luo, Water-in-oil emulsification in a bifurcated

33 tree-like network: Flow distribution properties and their impact on the emulsion

34 polydispersity, Chem. Eng. Res. Des. 134 (2018) 420-433.

35 https://doi.org/10.1016/j.cherd.2018.04.031.

36 [69] X. Guo, Y. Fan, L. Luo, Multi-channel heat exchanger-reactor using arborescent

37 distributors: A characterization study of fluid distribution, heat exchange performance

38 and exothermic reaction, Energy. 69 (2014) 728-741.

39 https://doi.org/10.1016/j.energy.2014.03.069.

40 [70] R.H. Liu, M.A. Stremler, K.V. Sharp, M.G. Olsen, J.G. Santiago, R.J. Adrian, H. Aref,

41 D.J. Beebe, Passive mixing in a three-dimensional serpentine microchannel, Journal of

42 microelectromechanical systems. 9 (2000) 190-197. https://doi.org/10.1109/84.846699.

43 [71] Z. Ni, E. Seebauer, R.I. Masel, Effects of microreactor geometry on performance:

44 differences between posted reactors and channel reactors, Ind. Eng. Chem. Res. 44 (2005)

45 4267-4271. https://doi.org/10.1021/ie048956w.

46 [72] G. Eigenberger, Fixed Bed Reactors, Universitätsbibliothek der Universität Stuttgart,

47 2009.

48 [73] O. Deutschmann, R. Schmidt, F. Behrendt, J. Warnat, Numerical modeling of catalytic

49 ignition, Symposium (International) on Combustion. 26 (1996) 1747-1754.

50 http://dx.doi.org/10.1016/S0082-0784(96)80400-0.
[74] S. Oh, P. Mitchell, R. Siewert, Methane oxidation over noble metal catalysts as related to controlling natural gas vehicle exhaust emissions, Catalytic control of air pollution. ACS Publication, Washington, DC, 1992.

[75] S.H. Oh, P.J. Mitchell, R.M. Siewert, Methane oxidation over alumina-supported noble metal catalysts with and without cerium additives, J. Catal. 132 (1991) 287-301. https://doi.org/10.1016/0021-9517(91)90149-X.

[76] G.S. Bugosh, V.G. Easterling, I.A. Rusakova, M.P. Harold, Anomalous steady-state and spatio-temporal features of methane oxidation on Pt/Pd/Al₂O₃ monolith spanning lean and rich conditions, Appl. Catal., B. 165 (2015) 68-78. https://doi.org/10.1016/j.apcatb.2014.09.058.

[77] K.H. Delgado, L. Maier, S. Tischer, A. Zellner, H. Stotz, O. Deutschmann, Surface reaction kinetics of steam-and CO₂-reforming as well as oxidation of methane over nickel-based catalysts, Catalysts. 5 (2015) 871-904. https://doi.org/10.3390/catal5020871.

[78] A.P. York, T. Xiao, M.L. Green, Brief overview of the partial oxidation of methane to synthesis gas, Top. Catal. 22 (2003) 345-358. https://doi.org/10.1023/A:1023552709642.

[79] C. Diehm, O. Deutschmann, Hydrogen production by catalytic partial oxidation of methane over staged Pd/Rh coated monoliths: Spatially resolved concentration and temperature profiles, Int. J. Hydrogen Energy. 39 (2014) 17998-18004. https://doi.org/10.1016/j.ijhydene.2014.06.094.

[80] R. Schwiedernoch, S. Tischer, C. Correa, O. Deutschmann, Experimental and numerical study on the transient behavior of partial oxidation of methane in a catalytic monolith, Chem. Eng. Sci. 58 (2003) 633-642. https://doi.org/10.1016/S0009-2509(02)00589-4.

[81] A. Amin, A. Abedi, R. Hayes, M. Votsmeier, W. Epling, Methane oxidation hysteresis over Pt/Al₂O₃, Appl. Catal. A. 478 (2014) 91-97. https://doi.org/10.1016/j.apcata.2014.03.032.

[82] I.Y. Pakharukov, A.Y. Stakheev, I.E. Beck, Y.V. Zubavichus, V.Y. Murzin, V.N. Parmon, V.I. Bukhtiyarov, Concentration hysteresis in the oxidation of methane over Pt/γ-Al₂O₃: X-ray absorption spectroscopy and kinetic study, ACS Catal. 5 (2015) 2795-2804. https://doi.org/10.1021/cs501964z.

[83] Y.-H. Chin, C. Buda, M. Neurock, E. Iglesia, Reactivity of chemisorbed oxygen atoms and their catalytic consequences during CH₄-O₂ catalysis on supported Pt clusters, J. Am. Chem. Soc. 133 (2011) 15958-15978. https://doi.org/10.1021/ja202411v.

[84] M. García-Diéguez, Y.-H.C. Chin, E. Iglesia, Catalytic reactions of dioxygen with ethane and methane on platinum clusters: Mechanistic connections, site requirements, and consequences of chemisorbed oxygen, J. Catal. 285 (2012) 260-272. https://doi.org/10.1016/j.jcat.2011.09.036.

[85] L. Chen, J.P. McCann, S.L. Tait, A re-examination of the catalyst activation and temperature hysteresis in methane combustion on Pt/Al₂O₃, Appl. Catal. A. 549 (2018) 19-30. https://doi.org/10.1016/j.apcata.2017.09.008.

[86] A. Abedi, R. Hayes, M. Votsmeier, W.S. Epling, Inverse hysteresis phenomena during CO and C₃H₆ oxidation over a Pt/Al₂O₃ catalyst, Catal. Lett. 142 (2012) 930-935. https://doi.org/10.1007/s10562-012-0861-x.

[87] R. Raj, M.P. Harold, V. Balakotaiah, Steady-state and dynamic hysteresis effects during lean co-oxidation of CO and C₃H₆ over Pt/Al₂O₃ monolithic catalyst, Chem. Eng. J. 281 (2015) 322-333. https://doi.org/10.1016/j.cej.2015.06.057.

[88] P. Li, D. Coopamah, N. Dhar. Analysis and optimization of flow distribution channels for uniform flow in fuel cells. in ASME 2008 Fluids Engineering Division Summer Meeting collocated with the Heat Transfer, Energy Sustainability, and 3rd Energy Nanotechnology Conferences. 2008. American Society of Mechanical Engineers Digital Collection.
1 [89] H. Liu, P. Li, Even distribution/dividing of single-phase fluids by symmetric bifurcation of flow channels, Int. J. Heat Fluid Flow. 40 (2013) 165-179.
3 https://doi.org/10.1016/j.ijheatfluidflow.2013.01.011.

4 [90] H. Liu, P. Li, K. Wang, The flow downstream of a bifurcation of a flow channel for uniform flow distribution via cascade flow channel bifurcations, Appl. Therm. Eng. 81 (2015) 114-127. http://dx.doi.org/10.1016/j.applthermaleng.2015.02.028.

7