Methane Decomposition Kinetics over Fe - ACS Publications

Subscriber access provided by Kaohsiung Medical University

Kinetics, Catalysis, and Reaction Engineering

Methane decomposition kinetics over Fe2O3 catalyst in micro fluidized bed reaction analyzer (MFBRA) Sulong Geng, Zhennan Han, Yan Hu, Yanbin Cui, Junrong Yue, Jian Yu, and Guangwen Xu Ind. Eng. Chem. Res., Just Accepted Manuscript • DOI: 10.1021/acs.iecr.8b00662 • Publication Date (Web): 06 Jun 2018 Downloaded from http://pubs.acs.org on June 6, 2018

Just Accepted “Just Accepted” manuscripts have been peer-reviewed and accepted for publication. They are posted online prior to technical editing, formatting for publication and author proofing. The American Chemical Society provides “Just Accepted” as a service to the research community to expedite the dissemination of scientific material as soon as possible after acceptance. “Just Accepted” manuscripts appear in full in PDF format accompanied by an HTML abstract. “Just Accepted” manuscripts have been fully peer reviewed, but should not be considered the official version of record. They are citable by the Digital Object Identifier (DOI®). “Just Accepted” is an optional service offered to authors. Therefore, the “Just Accepted” Web site may not include all articles that will be published in the journal. After a manuscript is technically edited and formatted, it will be removed from the “Just Accepted” Web site and published as an ASAP article. Note that technical editing may introduce minor changes to the manuscript text and/or graphics which could affect content, and all legal disclaimers and ethical guidelines that apply to the journal pertain. ACS cannot be held responsible for errors or consequences arising from the use of information contained in these “Just Accepted” manuscripts.

is published by the American Chemical Society. 1155 Sixteenth Street N.W., Washington, DC 20036 Published by American Chemical Society. Copyright © American Chemical Society. However, no copyright claim is made to original U.S. Government works, or works produced by employees of any Commonwealth realm Crown government in the course of their duties.

Page 1 of 34 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Industrial & Engineering Chemistry Research

Methane decomposition kinetics over Fe2O3 catalyst in micro fluidized bed reaction analyzer (MFBRA) Sulong Geng1,2, Zhennan Han3, Yan Hu1,2,Yanbin Cui1, Junrong Yue1, Jian Yu1*,Guangwen Xu1,3* 1

State Key Laboratory of Multi-phase Complex Systems, Institute of Process Engineering, Chinese Academy of Sciences, Beijing 100190, China.

2

University of Chinese Academy of Sciences, Beijing 100049, China.

3

Institute of Industrial Chemistry and Energy Technology, Shenyang University of Chemical Technology, Shenyang 110142, China.

* Corresponding Authors, Tel: +86 10 82544886, Fax: +86 10 82629912. E-mails: [email protected] (J. Yu); [email protected].

ABSTRACT Micro fluidized bed (MFB) reactor offers a good way to determine the intrinsic kinetics of some reactions. For catalytic decomposition of methane (CMD), the hydrogen production rate appears much higher when catalyst is injected into an MFB reactor by pulse, in comparison with that via gas switching. The carbon nanotubes (CNTs) produced in an MFB were more dispersed than that in fixed bed. Therefore, the micro fluidized bed reaction analyzer (MFBRA) is proven to be an alternative tool for determining kinetics of methane decomposition and offering the values of higher reference for CMD in industrial fluidized beds. Kinetics of CMD and catalyst deactivation over Fe2O3 catalyst were studied using MFBRA at temperatures of 750 °C to 900 °C and partial pressures of methane among 0.22, 0.50, 0.75 and 1.0 atm. The reaction order and activation energy of CMD were estimated to be 2.27 and 50 kJ/mol, respectively.

Keywords: Kinetics; Micro fluidized bed reactor analyzer (MFBRA); Catalytic decomposition of methane (CMD); Carbon nanotubes (CNT).

1

ACS Paragon Plus Environment

Industrial & Engineering Chemistry Research 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

For Table of Contents Only

2


Page 2 of 34

Page 3 of 34 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


1. Introduction In recent years, as a kind of promising clean energy, hydrogen has attracted much attention.1,2 It can be transformed into electricity without pollution at high efficiency.3,4 Nowadays, hydrogen is mostly obtained by steam reforming and partial oxidation of natural gas.5-8 Inevitably, it can produce a lot of CO2, the major greenhouse gas. Having reported by many researchers, methane decomposition is a kind of way to produce hydrogen with high purity but without emission of greenhouse gas.9-11 However, methane decomposition without the help of catalyst need the temperature of approximately 1200 oC.12,13 Therefore, catalytic methane decomposition (CMD) has received a great attention due to relative lower reaction temperature and the single step process to produce pure hydrogen.14-17 This process also provides an important way to exploit biomethane effectively for easy carbon capture and storage.18-20 Another important advantage of CMD is the generation of highly valuable by-product, carbon nanotubes (CNTs) or carbon nanofibers (CNFs),16,21 with great application potentials in current nanoscience. Furthermore, CNTs have the exclusive electrical, chemical and mechanical properties, which make the production of CNTs from CMD process a bright prospect in electronics (batteries, optoelectronics integrated system, electro catalytic CO2 reduction, thermos-electric modules for energy conversion, etc), energy storage, nano composites and medical field.22-27 The popular catalysts used for the CMD is Ni, Fe, and Co, using different metal oxides (such as Al2O3, MgO and so on) as supporting materials. In particular, Ni-based catalysts are believed to be most effective for CMD because the Ni particles are highly active at moderate temperatures of 500-700 oC with a high growth rate of graphite carbon.3,14,15,28-30 However, the Ni- or Co- based catalysts show restricted applications for their heavy-metal toxicity influence on human beings and fast deactivation at higher temperatures (>873K), that are required for high conversion of methane3,28 and also for producing highly stable graphite carbon. Compared with Ni- or Co-catalysts, Fe-catalyst, as a kind of cheaper and environmental-friendly catalyst, operating at higher temperature (700-950 oC) with the same good activity, could be used for CMD to produce H2 and carbon materials, especially the formation of CNTs.24,31-34 The solid carbon as the main product of CMD often blocks of the reactor after a few hours of reaction in a fixed bed reactor.3,21,34,35 Therefore, the continuous collection of the produced carbon is a key factor for CMD process.36 In order to realize this purpose, the CMD reactions in different types of reactors have been investigated in the literatures,34,37-39 showing that the fluidized bed (FB) is a kind of suitable reactors for continuous industrial operation.40 As reported,34,41 there are three main advantages to use FBR for reaction of CMD, efficient mixing of 3



the solid particles and mass transfer through the particles surface renewal via gaseous products on catalyst, an uniform temperature distribution in reactor and feasibility to realize operating continuously. Until now, there are many researches on the decomposition of methane in FB reactor.34,38,40,42-46 Pinilla and co-workers43-45 reported that CMD on metal catalysts can be implemented in a FB reactor without clogging at a suitable methane velocity. Torres et al.34 also reported the fluid dynamics and production of multiwall CNTs by CMD on Fe-based catalyst in a FB reactor. However, the reaction and deactivation kinetics of CMD in fluidized bed reactors are rarely reported, especially for Fe-based catalyst. Fukada et al.47 tested the kinetics of CMD over the Ni-based catalyst in a fixed bed reactor with the inner diameter of 6.3mm. Nasir Uddin et al.48 and Ashik et al.49 have also examined the methane decomposition kinetics over Ni-based catalyst in a fixed bed. Several works about the deactivation kinetic parameters of CMD on activated carbon catalyst were conducted using thermogravimetric analyzer (TGA).50-53 Up to now, there is no research about the kinetics of CMD on Fe-based catalysts in fluidized bed reactor. The micro fluidized bed reaction analyzer (MFBRA) was developed to realize the on-line pulse injecting of powder reactant or catalyst, rapid heating of fed particle, plug flow of gas to suppress the influence of diffusion in the reactor. These ensure the obtained kinetics to close the intrinsic values of the tested reaction. Through connecting with a process mass spectrometer (MS) to analyze the gaseous products, the MFBRA can also reveal the mechanism of reaction.54,55 Yu et al.54,56 have reported the kinetics and mechanism of some gas-solid reactions in MFBRA. Consequently, the kinetic study of CMD in MFBRA can possibly achieve the scientific basis for the development of CMD industrial-scale fluidized bed reactor in the future. This work for the first time carried out and compared three different methods of kinetic measurement in the MFBRA for the CMD reaction over Fe2O3 catalysts. The kinetics of methane decomposition and catalyst deactivation were obtained in the range of methane partial pressure from 0.22 to 1.0atm and at temperatures from 750 to 900 °C. The deposited carbon nanotubes were preserved and characterized using scanning electron microscopy (SEM). On all of these, the reaction mechanism of CMD was further analyzed.

2. Experimental methodology 2.1. Characterization of catalyst The catalysts containing 20 wt.% Fe2O3 are prepared by co-precipitation approach using Al2O3 and MgO as supporting materials. The nitrates Fe(NO3)3·9H2O, Al(NO3)3·9H2O and Mg(NO3)2·6H2O are used as precursors for Fe, Al and Mg, respectively. Amount-specified Fe(NO3)3·9H2O, Al(NO3)3·9H2O and Mg(NO3)2·6H2O were 4


Page 4 of 34

Page 5 of 34 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


dissolved in deionized water and further mixed with aqueous solution of ammonium carbonate (NH4)2CO3 to obtain a slurry mixture. The co-precipitated samples were carefully washed with distilled water and dried at 120 ◦C for 10 hours. The prepared catalyst materials were finally ground and sieved to make the samples used in tests. The related details of preparation approach can be found in a previous publication.57 The X-ray diffraction (XRD) pattern of fresh catalyst (Figure 1), was acquired in a PA Nalytical X’Pert PRO MPD instrument. According to the XRD patterns, the diffraction peaks at 2θ = 12, 23 and 50 principally represent the support Mg4Al2(OH)12CO3•3H2O (PDF#51-1528). There are several peaks at 2θ = 25, 33, 35, 41, 54, and 62-65, which are related to Fe2O3 (PDF#33-0664). The nano Fe2O3 particles are well dispersed on the support Mg4Al2(OH)12CO3•3H2O. The morphological characteristics of the fresh catalyst was analyzed through scanning electron microscopy (SEM) (Hi-tachi SU8000), given in Figure 2. A well-displayed virgulate structure can be identified and the size of catalyst particles is about 500nm. This indicates that the sample of Fe2O3 catalyst has higher crystallinity. Nitrogen adsorption-desorption measurements (BET method) were carried out in Micromeritics ASAP 2020 BET apparatus at -196 oC. The adsorption-desorption isotherms of Fe2O3 catalyst indicate the type III isotherm, signifying a nonporous or macroporous materials (Figure 3). The BET surface area, pore volume and pore diameter of the catalysts is 29.40 m2/g, 0.18cm3/g and 24.31 nm, respectively. The characterization shows that Fe2O3 is located dispersed on the surface of the support materials. This result is highly corresponding with the XRD data in Figure 1.

2.2. Experimental set-up and operating procedure The CMD reaction was conducted in the MFBRA. The schematic diagram is shown in Figure 4, the MFBRA includes a micro fluidized bed (MFB) reactor with an inner diameter of 15 mm and a reaction zone of 60 mm in length, an sample injector, an electric furnace, and an online rapid mass spectrometer (MS, AMETEK, Dycor

system 2000 LC-D). As detecting system, MS is connected with the outlet of the MFB reactor. Some other details of the MFBRA can be found in the works reported before.58 All experiments were conducted at atmospheric pressure. The results of three typical experimental approaches were compared in this work. The first was the traditional method conducted in MFBRA (Method I). The properties of α-Al2O3 particles are shown in Table 1 and the minimum fluidization velocity Umf for the particles was determined by Eq. (1):59

5



2

33.7 + U mf =

0.0408d p3 ρ g ( ρ s − ρ g ) g

µ2

Page 6 of 34

− 33.7

d pρg / µ

(1)

After α-Al2O3 particles with 10 mm static bed height put inside the MFB reactor, the reactor was heated from 20 oC to the preset value (750-900 oC) with the heating rate of 20 oC /min by using pure Ar as the fluidized gas at a flow rate of 60 mL/min. The determined Umf is 0.0037 m/s, and the gas velocity in the reactor was 5-6Umf at temperatures of 750-900oC to maintain the expected stable bubble fluidization conditions. At the preset temperature, the CH4/Ar mixture in 60 mL/min was fed into the reactor to replace the Ar flow completely. The catalyst with about a mount of 15 mg was injected into the reactor instantaneously to mix with the fluidized α-Al2O3 particles in the MFB to initiate the CMD reaction. For the second and third approaches (Method II and Method III) all experimental conditions were kept the same as for the first method, but the mixture of catalyst and α-Al2O3 particles for the second approach and only catalyst particles for the third approach were put into the inside of the reactor in advance. In the second and third cases the CMD reaction started as the gas switched from pure Ar to CH4-Ar mixture at a preset temperature. The gas switching method was similar to that literature reported testes conducted in fixed bed reactors or thermo-gravimetric analyzers.48,50 Because the catalysts particles are cohesive and it was difficult to keep normal fluidization, the third set of experiments was actually conducted in a fixed bed with 15 mg pure and small-size catalyst particles. All the experiments were conducted at 750−900 °C and the produced gas product was detected on-line using MS, which had been calibrated by micro gas chromatography (GC, Agilent 3000A) to ensure the accuracy of hydrogen and methane concentrations measured by MS. Based on the compositions of gaseous product and kinetic data, the reaction characteristics were analyzed. In order to control the gas flow rates, mass flowmeter in 0-2 L/min were used. Methane-Argon mixture with different methane partial pressures PCH4 of 0.22 atm, 0.50 atm, 0.75 atm and 1.0 atm and pure Argon (99.999%) was purchased from DaTe Corp., China.

2.3. CMD thermodynamics According to stoichiometry of Eq. (2), the conversion of CH4 αCH4 can define by Eq. (3). The real-time H2% CH2 can be measured by MS based on the standard response curve of MS calibrated by micro gas chromatography. → C + 2

(2)

%

(%) =

%

(3)

Once αCH4 was calculated, the formation rate of hydrogen rH2 and decomposition rate of methane rCH4 6


Page 7 of 34 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


(mol/min·gcat-1) can be computed: =

× ×

(4)

=

(5)

where is the molar flow rate of methane in- mol/min, mcal is the mass of catalyst in g. The carbon deposited rate !"#$% (g/min·gcat-1) on catalyst is calculated by !"#$% =

× ×

(6)

The time dependence of carbon yield over the whole course of CMD is determined by ,

&'()* = +0 &'()* (,) ., =

×12 4

1&'2

,

+0 4 (,).,

(7)

Eq.(7) can estimate the accumulated carbon thermodynamically, although methane might be decomposed thermally.

3. Results and discussion 3.1. CMD kinetic from three methods Figure 5 reports the product-gas releasing characteristics of three different methods by experiments at methane with partial pressure of 0.75atm and temperature of 750 °C. Apparently, the pulse injection of catalysts (method I) resulted in a greater signal intensity (concentration) of H2 than that by gas switch (method II). For the first method, the peak of H2 appears (t1) earliest among all experiments. In order to figure out the initial reaction before t1, Figure 6 compares the result of blank test and experiments with catalysts addition. The formation of CO2 and H2O is followed by the injection of catalysts into the reactor (Figure 6a, Method I) or gas switching from Ar to Ar-CH4 mixture (Figure 6b, c). The reaction, 4Os+CH4→CO2+2H2O, occurs with the oxygen Os linked to iron on the oxide catalyst, which forms the stage of 0-t1 in Figure 5. However, as shown in Figure 7, it would take about 600s for Ar to be replaced by Ar/CH4 mixture completely in the reactor after the gas switch. As reported in the literatures,32,33 the reduction of Fe2O3 by CH4 goes through in three steps: Fe2O3—Fe3O4—FeO—Fe. Figure 8 expresses the reduction process of Fe2O3 over catalyst. Because of the difference of CH4 concentration in the reactor, the surrounded CO2 and H2O on the surface of catalyst for the method II lasts for a longer time than that of method I does, thus delaying the H2 production (t1) for the method II. In addition, the fluidization of catalysts and α-Al2O3 particles result in quickly renewal of gas species on the surface of catalyst. Therefore, the product gases CO2 and H2O in fixed bed (method III) are taken away slowly by CH4, causing a longer stage in Figure 8 (b) to (c). It is why the time t1 corresponding to H2 formation comes later in method III than that in method II. With the consumption of Os, the Fe2O3 in catalyst will be reduced completely by CH4. Meantime, as the active Fe0 accumulated to a certain 7



34 5

amount in catalyst, decomposition reaction of CH4 occurs, 67 2 + . The amorphous carbon and H2 would produce simultaneously. Meanwhile, Fe interacts with amorphous carbon to form Fe3C, and amorphous carbon becomes graphite carbon under the help of Fe3C.32 Therefore, the period t1-t2 can be regarded as the stage to activate the catalyst. Methane decomposition kinetics is obtained from the highest point at t2, the initial hydrogen peak. After t2, the catalyst begins to deactivate because of deposition of carbon. For the method I, the methane concentration in the reactor is higher and more stable in the initial 600 s, leading to the higher and earlier appearance of its initial hydrogen peak (concentration) CH2,0. In the method II, the concentration of CH4 in the reactor varies from 0 to a stable value in about 600 s. When H2 begins to appear, Ar is yet not completely replaced, so that the concentration of CH4 in the reactor is lower than that for the method I. Furthermore, as shown in Figure 5, the period from t1 to t2 in the method II is longer than that in the method I. Carbon is formed in the stage of t1-t2, which influences the intrinsic kinetics of CMD. The longer the stage of t1-t2 is, the lower the initial hydrogen peak is. The lower CH4 concentration, along with the longer stage of t1-t2, leads to the lower initial hydrogen peak in the method II. Hydrogen begins to produce in fixed bed (method III) as the complete replacement of Ar by CH4 is reached. Therefore, the initial hydrogen peak in the method III is very similar with that in the method I. The kinetic curves obtained from the methods I and III are very similar. This shows the consistency in reaction kinetics, which should not have great differences for different methods. However, the kinetic data from the method I, especially the reaction time and the maximal reaction rate are more appropriate to represent the reaction characteristics taking place in the continuously fluidized bed reactor that actually simulates actual industrial processes. In addition, the reactant injection method makes MFBRA good adaptable for thermally instable reactants in obtaining their decomposition kinetics under isothermal conditions, for example, the reaction of fuel pyrolysis and carbonate decomposition.56,58,60,61 As reported in our previous work,56 the heating rate of the reactant particles can reach about 5*104 oC/s in a fluidized bed when the size of particle is smaller than 50 µm. According to this, heating up cold reactant particles to 900 oC in MFBRA only requires 0.018 s, which is much shorter than the reduction time of catalyst so that it is ignorable. Therefore, the reaction by the method I at temperatures of 700-950 oC should be under isothermal conditions to ensure its comparability with the other two methods. In summary, for having stable gas composition and temperature, the traditional method (pulse injection of catalyst) in MFBRA (Method I) is suitable for the kinetic analysis of CMD. This offers an alternative way to obtain 8


Page 8 of 34

Page 9 of 34 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


reaction kinetics closer the industrial of CMD process in fluidized bed, if comparing with the studies in fixed bed reactor or thermo-gravimetric analyzer. According to studies on CMD kinetics using thermogravimetric analysis,52 the formation rate of H2 at time t2 could be regarded as the initial formation rate, and the catalyst deactivation time is defined as t>t2. As shown in Figure 5, the decomposition rate of CH4 decreases with reaction time, representing the gradual deactivation of catalyst because of carbon deposition on active sites. Figure 9 shows the SEM images of the morphologies of carbon nanotubes from three different methods, (a) method I: pulse injection of powdery catalyst into MFBRA, (b) method II: gas switch in MFB, and (c) method III: gas switch in fixed bed reactor (catalyst bed without fluidized particles). Apparently, the product of CNTs from micro fluidized bed, Figure 9a and 9b, dispersed well, with a uniform average diameter of 40 nm. However, as shown in Figure 9c by arrows, we can see clearly particle agglomerates among CNTs. The different dispersions of CNTs prove actually the advantages of fluidized bed for CMD, which is consistent with other works. 34,62

3.2. Kinetics of catalytic methane decomposition 3.2.1. Reaction behavior under different conditions The reaction temperature T and methane partial pressure PCH4 significantly affect the hydrogen production rate during CDM over iron catalyst. In our tests, the conversion of methane was continuously recorded until the instantaneous hydrogen concentration was below 3%. Figure 10 (a-d) shows the time-series rate of hydrogen formation rH2 in CMD according to Eq. (4) under various PCH4 (0.22, 0.50, 0.75, and 1.0 atm) and at temperatures of 750, 800, 850 and 900 oC. The reduction time (0-t1) shortens with raising T and PCH4. The initial formation rate of hydrogen (peak height) rH2,0 at t=t2 also increased with the rise of temperature and PCH4. Similar varying tendency was also described by Ashik et al.49 over Ni/SiO2 catalysts. Due to the endothermic feature of CMD, rcarbon is largely related to temperature, especially in atmosphere with high CH4 partial pressures. The researches on CMD over activated carbon catalysts by Daud et al.51 revealed that increasing T and PCH4 resulted in the improvement of methane diffusion rate and in turn accelerated the decomposition reaction. Reported by Chen et al.,49,63 the driving force on carbon diffusion, determination of the hydrogen and carbon formation rate over catalytic sites, is up to PCH4. However, the higher temperature led to faster deactivation of the catalyst. Because of the higher carbon deposited rate rcarbon, the active sites of the catalyst would be quickly encapsulated. As reported, the carbon removal rate from the active sites by H2 is greatly lower than the rate of carbon deposition, thus deactivating the encapsulated active sites. Consequently, the deposited solid carbon encapsulates the active catalyst sites to hinder 9



Page 10 of 34

its (their) chemisorption of methane molecules further and shortens the active lifetime of catalyst.49 For t>t2, the hydrogen formation rate rH2 decreases with reaction time at all experimental temperatures but there is a big slop at high temperatures and high PCH4. The results are consistent with many previous studies.49,51 In addition, the higher carbon accumulation at higher T and PCH4 resulted in rapider filling of catalyst pore mouths. This then accelerates the encapsulation rate of active metal phases and prevents methane molecules from access into the internal pores of the catalyst. This indicates another cause of catalyst deactivation.49,64 The reported initial hydrogen concentration CH2,0 and formation rate rH2,0 in Table 2 increases with increasing temperature and PCH4. The initial hydrogen production was 8.02 vol.% at PCH4 of 0.50 atm and temperature of 750oC, which increased to 23.6 vol.% at 900 oC. The similar trend was observed also at other experimental PCH4 including 0.22, 0.75, and 1.0 atm. Hence, the chemical conversion of methane over the tested catalyst improved when increasing temperature. 3.2.2. CMD kinetics over Fe2O3 catalyst For the endothermic feature of CMD, the initial hydrogen formation rate rH2,0 increases as raise T and PCH4, which means the initial methane decomposition rate rCH4,0 increases with raising T and PCH4. Figure 11(a) exhibits the relationship between initial decomposition rate of CH4 (Ln(rCH4,0)) and partial pressure of CH4 (Ln(PCH4)) at different experimental temperatures. According to the reaction rate equation of CMD: rCH4=kpPCH4n Ln(rCH4)=Ln(kp)+nLnPCH4

(8) (9)

The slopes from Figure 11(a) represent the order of reaction, which are 2.36, 2.19, 2.20 and 2.32 at 750, 800, 850 and 900 oC respectively. Ln(kp) can be determined from the intercepts of Figure 11(a). Their regression coefficients are high and ranged in 0.98-1.0. The average value of the four slopes obtained from Figure 11 (a), 2.27, is regarded as the order of CMD reaction over Fe2O3 catalyst. Then, the maximal reaction rate, also the initial methane decomposition rate, is expressed: Rmax = rCH4,0=kpPCH42.27.

(10)

Such a reaction order represents the sensitivity of maximal reaction rate to partial pressure of CH4 over Fe2O3 catalysts. For metal-supported and carbon catalysts the literature reported values are 0.5 to 2.65.14,32,48,49,65 For example, the reaction order for CMD over Ni-based catalyst was reported to be 2.65 and 1.4 by Nasir Uddin and co-workers.48,49. As a matter of fact, the true mechanism of CMD was quite complicated and the kinetic laws of each intermediate steps are still unclear.49 Zhou et al.32 reported that Fe3C played the major role for the production of graphite carbon over Fe-based catalyst but there were not studies on kinetics. 10


Page 11 of 34 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


According to the Arrhenius law, the rate constant kp is expressed as: kp=Ae-Ea/RT

(11)

in which, A is the pre-exponential factor (mmol/(gcat·min·atm2.27)), Ea is the activation energy (kJ·mol-1) and R is the gas constant (8.314J·K-1·mol-1). Typically, taking the logarithm of the Eq. (11): Ln(kp)=Ln(A)-Ea/R·1/T

(12)

The slope determined by Ln(kp) vs. 1/T, as in Figure 11(b), gives the activation energy (-Ea/R) for the CMD over Fe2O3 catalyst. The value of Ea and pre-exponential factor from Figure 13(b) are 50kJ·mol-1and 10605.2mmol/(gcat·min·atm2.27), respectively. Hence, the initial methane decomposition rate is expressed as rCH4,0=Ae-Ea/RT·PCH42.27=10605.2e-60266/T·PCH42.27

(13)

The activation energy values reported in other studies are most about Ni-based catalysts48,49 and carbon catalysts51,53, which are 50-90 kJ·mol-1, and about 200kJ·mol-1, respectively. As reported, the wide range of activation energy about metal catalysts should first due to the variance of catalytic traits, for example, multiple forms of transition metals in single or bimetallic conditions, different crystalline faces and types of supports.49 The different characteristics of catalysts cause different interactions among catalyst nanoparticles. Another factor is the difficulty in obtaining the intrinsic kinetic data from the recent dynamic reaction.66 In this work, catalysts in MFBRA could be well dispersed by fluidized particles and have sufficient opportunity to contact methane. Also, the concentration of CH4 in fluidized bed is more stable and higher than that in fixed bed. Consequently, the interfacial and intra-particle diffusion of gas can be avoided at the possibly highest extent.

3.3. Catalytic activity and deactivation kinetics During the process of CMD, rH2 reduces gradually due to the deactivation of catalyst. The highest initial hydrogen formation rate at 900 oC and PCH4 of 1.0 atm is 1.86 (mmol/(gcat·s)), declining to 0.032 (mmol/gcat·s) after 10 min of reaction. At 750, 800 and 850 ℃the initial hydrogen formation rates under PCH4 = 1.0 atm are 0.97,1.67 and 1.76 (mmol/gcat·s), which abridges to 0.11, 0.059 and 0.047 (mmol/gcat·s) by 10 minutes of reaction, respectively. The results show that the decreasing trend of methane decomposition intensity is less at lower reaction temperature, or the activity loss calculated by Eq. (14) increases with raising temperature and partial pressure PCH4, which reached 90% in a short time of 10 min at 800-950 oC. 8&,9:9,; 2)== 8>% = 100 −

(@$A.%) !C >% (@$A.%)!C >%

× 100%

(14)

The catalyst had lower activity loss rate under PCH4 of 0.22 and 0.50 atm at 750 oC, revealing the phenomena of deactivation at higher T is more evident under higher PCH4. The activity of the catalyst, a(t), is defined 11



Page 12 of 34

as:48,50,52,53 a(t) = rc(t)/rc,0

(15)

In Eq. (15), rc(t) and rc,0 are known as the carbon deposited rate at time t and the initial carbon deposited rate, respectively. They are functions of PCH4 and reaction temperature.50 For → C + 2 , the molar carbon formation rate is half of hydrogen formation rate. In Figure 10, the hydrogen formation rate decreases because of the carbon deposition on the exterior surface of catalyst. The decline of rH2 also indicates the deactivation of the catalyst. Surely, the deactivation rate of catalyst is the function of T, PCH4, and state of catalyst.50,52 We thus have (on the gas phase concentration): -da/dt=kdPCH4mad=kd0e-Ed/RTPCH4mad

(16)

In Eq. (16), kd is the rate constant of catalyst deactivation in (atmm·s)-1, d is the order of deactivation, m is the methane gas-phase concentration dependency, Ed is the activation energy of deactivation, and kd0 is the pre-exponential factor in (atmm·s)-1. Integrating of Eq. (16) in terms of t at given PCH4 and T leads to Eq. (17): a1-d=1-(1-d)kdPCH4mt

(17)

By assuming different values of d, plots the line of a1-d versus t, with the slope s: s=-(1-d)kdPCH4m, S=s/(d-1)=kdPCH4m

(18)

The value of d with the highest linearity of a1-d vs. t for the Eq. (17) is the deactivation order. The selected d ought to suitable for data at all test PCH4 and T during the total deactivation progress. In this work, as shown in Figure 12, d is equal to 1.8 with the regression coefficients varying from 0.97 to 1.0. The resulting values of S for each PCH4 and T can be taken to calculate the other kinetic parameters for deactivation. When S was determined, kd and m can be easily got through plotting Ln(S) against Ln(PCH4) at different T. Ln(S)=Ln(kd)+mLn(PCH4)

(19)

As shown in Figure 13(a), the slope of four temperatures determines the value of m=0.95, and the intercept determines the different values of kd at various T. According to Arrhenius equation: kd=kd0e-Ea/RT

(20)

Plotting Ln(kd) vs. 1/T, shown in Figure 13(b), comes into being the straight line with a slope of –Ea/RT. The intercept on Y axis is value of kd0. The value of Ea for the deactivation of CMD over Fe2O3 catalyst is 121.37 kJ/mol, and kd0 is 87903.95 (min·atmm)-1. Therefore, the Arrhenius equation for the deactivation of CMD over Fe2O3 catalyst is kd=87903.95exp(-16905/T) 12


(21)

Page 13 of 34 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


Some works have reported that the Ea for the deactivation of CMD over Ni-based catalysts in fixed bed is in 75-97kJ/mol.48,67 In fact, different catalysts, reactor and experimental approaches can lead to different results. The activation energy of a reaction reflects the sensitivity of the reaction to temperature. In this work, the activation energy required for the CMD over Fe2O3 catalysts (50 kJ/mol) is lower than that for catalytic deactivation (121.37 kJ/mol). This shows that the catalytic deactivation is more sensitive than that the reaction of CMD. With raising temperature, the rate of catalyst deactivation varies rapider than the methane decomposition rate does.

4. Conclusion For catalytic methane decomposition (CDM) over iron catalyst, the hydrogen production rate appeared to be higher when catalyst was injected into a micro fluidized bed reactor by pulse, if comparing with the gas switch from Ar to CH4 into a catalyst bed. Nanotubes deposited in fluidized bed were more dispersed than in fixed bed according to evaluation by SEM. Therefore, the micro fluidized bed reaction analyzer (MFBRA) was used to obtain the kinetics of CMD, and the kinetics would be closer to the actual reactions in industrial fluidized bed reactors. A series of methane decomposition experiments were conducted over a self-made Fe2O3 catalyst in MFBRA at temperatures of 750-900°C and four different partial pressures of methane of 0.22, 0.50, 0.75 and 1.0 atm. The calculated kinetics based on the specific molar hydrogen formation rate clarified that the reaction order of CMD is 2.27 with an Ea of 50 kJ/mol. The catalyst deactivation kinetics was also determined, and its deactivation order, methane concentration dependency and activation energy are 1.8, 0.95, and 121.37 kJ/mol, respectively.

13



Acknowledgements The authors are grateful to the financial supports of the National Basic Research Program of China (2014CB744303) and the National Natural Science Foundation of China (91534125)

14


Page 14 of 34

Page 15 of 34 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


Symbols

αCH4 —— Conversion of CH4, % rH2—— The formation rate of hydrogen, mol/min·gcat-1 rCH4—— The decomposition rate of methane, mol/min·gcat-1 fCH4—— The molar flow rate of methane, mol/min CH2——The concentration of H2, % CH2,0——The initial concentration of H2, % mcal—— The mass of catalyst, g rcarbon—— The deposited carbon rate, g/min·gcat-1 rCH4,0—— The initial methane decomposition rate, mol/min·gcat-1 rH2,0—— The initial formation rate of hydrogen, mol/min·gcat-1 rcarbon,0—— The initial deposited carbon rate, g/min·gcat-1 PCH4—— The partial pressure of methane, atm T —— The temperature, K

kp

—— The rate constant

A

—— The pre-exponential factor of Arrhenius law

Ea —— The activation energy, kJ·mol-1 a(t) —— The normalized activity of catalyst, % kd —— The rate constant of catalyst deactivation, (atmm·s)-1 m —— The methane gas-phase concentration dependency. Ed —— The activation energy of deactivation, kJ·mol-1 d —— The order of deactivation kd,0 —— The pre-exponential factor, (atmm·s)-1

15



References (1) Pudukudy M, Yaakob Z, Mohammad M, Narayanan B, Sopian K. Renewable hydrogen economy in Asia – Opportunities and challenges: An overview [J]. Renew. Sust. Energ. Rev., 2014, 30(2): 743-757. (2) Pudukudy M, Yaakob Z, Takriff M S. Methane decomposition into COx free hydrogen and multiwalled carbon nanotubes over ceria, zirconia and lanthana supported nickel catalysts prepared via a facile solid state citrate fusion method [J]. Energy Convers. Manage., 2016, 126: 302-315. (3) Wang G, Jin Y, Liu G, Li Y. Production of Hydrogen and Nanocarbon from Catalytic Decomposition of Methane over a Ni–Fe/Al2O3Catalyst [J]. Energy Fuels, 2013, 27(8): 4448-4456. (4) Lubitz W, Tumas W. Hydrogen: An Overview [J]. Chem. Rev., 2007, 107(10): 3900-3. (5) Shen W, Wang Y, Shi X, Shah N, Huggins F, Bollineni S, Seehra M, Huffman G. Catalytic Nonoxidative Dehydrogenation of Ethane over Fe–Ni and Ni Catalysts Supported on Mg(Al)O to Produce Hydrogen and Easily Purified Carbon Nanotubes [J]. Energy Fuels, 2007, 21(6): 3520-3529. (6) González A R, Asencios Y J O, Assaf E M, Assaf J M. Dry reforming of methane on Ni–Mg–Al nano-spheroid oxide catalysts prepared by the sol–gel method from hydrotalcite-like precursors [J]. Appl. Surf. Sci., 2013, 280(280): 876-887. (7) Izquierdo U, Barrio V L, Cambra J F, Requies J, Güemez M B, Arias P L, Kolb G, Zapf R, Gutiérrez A M, Arraibi J R. Hydrogen production from methane and natural gas steam reforming in conventional and microreactor reaction systems [J]. Int. J. Hydrogen Energy, 2012, 37(8): 7026-7033. (8) Sun N, Wen X, Wang F, Peng W, Zhao N, Xiao F, Wei W, Sun Y, Kang J. Catalytic performance and characterization of Ni–CaO–ZrO 2 catalysts for dry reforming of methane [J]. Appl. Surf. Sci., 2011, 257(21): 9169-9176. (9) Abbas H F, Wan Daud W M A. Hydrogen production by methane decomposition: A review [J]. Int. J. Hydrogen Energy, 2010, 35(3): 1160-1190. (10) Li Y, Li D, Wang G. Methane decomposition to CO x -free hydrogen and nano-carbon material on group 8–10 base metal catalysts: A review [J]. Catal. Today, 2011, 162(162): 1–48. (11) Pudukudy M, Yaakob Z, Takriff M S. Methane decomposition over Pd promoted Ni/MgAl2O4 catalysts for the production of COx free hydrogen and multiwalled carbon nanotubes [J]. Appl. Surf. Sci., 2015, 356: 1320-1326. (12) Muradov N, Smith F, Huang C. Autothermal catalytic pyrolysis of methane as a new route to hydrogen production with reduced CO2 emissions [J]. Catal. Today, 2006, 116(3): 281-288. (13) Muradov N Z, Veziroǧlu T N. From hydrocarbon to hydrogen–carbon to hydrogen economy [J]. Int. J. Hydrogen Energy, 2005, 30(3): 225-237. (14) Abbas H F, Daud W M A W. Hydrogen production by methane decomposition: A review [J]. Int J Hydrogen Energ, 2010, 35(3): 1160-1190. (15) Awadallah A E, Aboul-Enein A A, Aboul-Gheit A K. Effect of progressive Co loading on commercial Co–Mo/Al 2 O 3 catalyst for natural gas decomposition to CO x -free hydrogen production and carbon nanotubes [J]. Energy Convers. Manage., 2014, 77(1): 143-151. (16) Li Y, Chen J, Yongning Qin A, Liu C. Simultaneous Production of Hydrogen and Nanocarbon from Decomposition of Methane on a Nickel-Based Catalyst [J]. Energy Fuels, 2012, 14(6): 1188-1194. (17) Pudukudy M, Yaakob Z, Akmal Z S. Direct decomposition of methane over SBA-15 supported Ni, Co and Fe based bimetallic catalysts [J]. Appl. Surf. Sci., 2015, 330: 418-430. (18) Liu C, Lu X. Carbon reduction pattern in China: Comparison of CCS and biomethane route [J]. Ciesc Journal, 2013, 64(1): 7-10. (19) Moghaddam E A, Ahlgren S, Nordberg Å. Assessment of Novel Routes of Biomethane Utilization in a Life Cycle Perspective [J]. Front Bioeng Biotechnol, 2016, 4: 89. (20) Koornneef J, Breevoort P V, Noothout P, Hendriks C, Luning U, Camps A. Global Potential for Biomethane 16


Page 16 of 34

Page 17 of 34 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


Production with Carbon Capture, Transport and Storage up to 2050 ☆ [J]. Energy Procedia, 2013, 37(37): 6043-6052. (21) Awadallah A E, Aboul-Enein A A, El-Desouki D S, Aboul-Gheit A K. Catalytic thermal decomposition of methane to COx-free hydrogen and carbon nanotubes over MgO supported bimetallic group VIII catalysts [J]. Appl. Surf. Sci., 2014, 296: 100-107. (22) Wolfbeisser A, Klötzer B, Mayr L, Rameshan R, Zemlyanov D, Bernardi J, Föttinger K, Rupprechter G. Surface modification processes during methane decomposition on Cu-promoted Ni–ZrO2 catalysts [J]. CATAL SCI TECHNOL, 2015, 5(2): 967-978. (23) Awadallah A E, Gad F K, Aboul-Enein A A, Labib M R, Aboul-Gheit A K. Direct conversion of natural gas into CO x -free hydrogen and MWCNTs over commercial Ni–Mo/Al 2 O 3 catalyst: Effect of reaction parameters [J]. Egyptian Journal of Petroleum, 2013, 22(1): 27-34. (24) Calafat Á, Sánchez N. Production of carbon nanotubes through combination of catalyst reduction and methane decomposition over Fe–Ni/ZrO2 catalysts prepared by the citrate method [J]. Appl. Catal. A: Gen., 2016, 528: 14-23. (25) Liu Y, Wang S, Liu H, Peng L M. Carbon nanotube-based three-dimensional monolithic optoelectronic integrated system [J]. Nat Commun, 2017, 8: 15649. (26) Zhang X, Wu Z, Zhang X, Li L, Li Y, Xu H, Li X, Yu X, Zhang Z, Liang Y, Wang H. Highly selective and active CO2 reduction electrocatalysts based on cobalt phthalocyanine/carbon nanotube hybrid structures [J]. Nat Commun, 2017, 8: 14675. (27) Zhou W, Fan Q, Zhang Q, Cai L, Li K, Gu X, Yang F, Zhang N, Wang Y, Liu H, Zhou W, Xie S. High-performance and compact-designed flexible thermoelectric modules enabled by a reticulate carbon nanotube architecture [J]. Nat Commun, 2017, 8: 14886. (28) Wang W, Wang H, Yang Y, Jiang S. Ni–SiO2 and Ni–Fe–SiO2 catalysts for methane decomposition to prepare hydrogen and carbon filaments [J]. Int. J. Hydrogen Energy, 2012, 37(11): 9058-9066. (29) Chen J, Li Y, Li Z, Zhang X. Production of COx-free hydrogen and nanocarbon by direct decomposition of undiluted methane on Ni–Cu–alumina catalysts [J]. Appl. Catal. A: Gen., 2004, 269(1): 179-186. (30) Awadallah A E, Aboul-Enein A A, Aboul-Gheit A K. Various nickel doping in commercial Ni–Mo/Al 2 O 3 as catalysts for natural gas decomposition to CO x -free hydrogen production [J]. Renew Energy, 2013, 57(3): 671-678. (31) Avdeeva L B, Reshetenko T V, Ismagilov Z R, Likholobov V A. Iron-containing catalysts of methane decomposition: accumulation of filamentous carbon [J]. Appl. Catal. A: Gen., 2002, 228(1): 53-63. (32) Zhou L, Enakonda L R, Harb M, Saih Y, Aguilar-Tapia A, Ould-Chikh S, Hazemann J L, Li J, Wei N, Gary D, Del-Gallo P, Basset J M. Fe catalysts for methane decomposition to produce hydrogen and carbon nano materials [J]. Appl. Catal. B-Environ., 2017, 208: 44-59. (33) Reddy Enakonda L, Zhou L, Saih Y, Ould-Chikh S, Lopatin S, Gary D, Del-Gallo P, Basset J M. Methane-induced Activation Mechanism of Fused Ferric Oxide-Alumina Catalysts during Methane Decomposition [J]. ChemSusChem, 2016, 9(15): 1911-5. (34) Torres D, de Llobet S, Pinilla J L, Lázaro M J, Suelves I, Moliner R. Hydrogen production by catalytic decomposition of methane using a Fe-based catalyst in a fluidized bed reactor [J]. J NAT GAS CHEM, 2012, 21(4): 367-373. (35) Pinilla J L, Utrilla R, Karn R K, Suelves I, Lázaro M J, Moliner R, García A B, Rouzaud J N. High temperature iron-based catalysts for hydrogen and nanostructured carbon production by methane decomposition [J]. Int. J. Hydrogen Energy, 2011, 36(13): 7832-7843. (36) Zhou L, Guo Y, Hideo K. Unsupported nickel catalysts for methane catalytic decomposition into pure hydrogen [J]. AlChE J., 2014, 60(8): 2907-2917. (37) Muradov N. Hydrogen via methane decomposition: an application for decarbonization of fossil fuels [J]. Int. J. Hydrogen Energy, 2001, 26(11): 1165-1175. 17



(38) Shah N, Ma S, Wang Y, Huffman G P. Semi-continuous hydrogen production from catalytic methane decomposition using a fluidized-bed reactor [J]. Int. J. Hydrogen Energy, 2007, 32(15): 3315-3319. (39) Amin A M, Croiset E, Malaibari Z, Epling W. Hydrogen production by methane cracking using Ni-supported catalysts in a fluidized bed [J]. Int. J. Hydrogen Energy, 2012, 37(14): 10690-10701. (40) Ammendola P, Chirone R, Ruoppolo G, Russo G, Solimene R. Some issues in modelling methane catalytic decomposition in fluidized bed reactors [J]. Int. J. Hydrogen Energy, 2008, 33(11): 2679-2694. (41) Vahlas C, Caussat B, Serp P, Angelopoulos G N. Principles and applications of CVD powder technology [J]. Mater. Sci. Eng., R, 2006, 53(1): 1-72. (42) Ammendola P, Chirone R, Ruoppolo G, Russo G. Production of hydrogen from thermo-catalytic decomposition of methane in a fluidized bed reactor [J]. Chem. Eng. J., 2009, 154(1-3): 287-294. (43) Pinilla J, Moliner R, Suelves I, Lazaro M, Echegoyen Y, Palacios J. Production of hydrogen and carbon nanofibers by thermal decomposition of methane using metal catalysts in a fluidized bed reactor [J]. Int. J. Hydrogen Energy, 2007, 32(18): 4821-4829. (44) Pinilla J L, Suelves I, Lázaro M J, Moliner R, Palacios J M. Parametric study of the decomposition of methane using a NiCu/Al2O3 catalyst in a fluidized bed reactor [J]. Int. J. Hydrogen Energy, 2010, 35(18): 9801-9809. (45) Pinilla J L, Lázaro M J, Suelves I, Moliner R, Palacios J M. Characterization of nanofibrous carbon produced at pilot-scale in a fluidized bed reactor by methane decomposition [J]. Chem. Eng. J., 2010, 156(1): 170-176. (46) Dunker A, Kumar S, Mulawa P. Production of hydrogen by thermal decomposition of methane in a fluidized-bed reactor—Effects of catalyst, temperature, and residence time [J]. Int. J. Hydrogen Energy, 2006, 31(4): 473-484. (47) Fukada S, Nakamura N, Monden J, Nishikawa M. Experimental study of cracking methane by Ni/SiO2 catalyst [J]. J. Nucl. Mater., 2004, 329: 1365-1369. (48) Nasir Uddin M, Wan Daud W M A, Abbas H F. Kinetics and deactivation mechanisms of the thermal decomposition of methane in hydrogen and carbon nanofiber Co-production over Ni-supported Y zeolite-based catalysts [J]. Energy Convers. Manage., 2014, 87: 796-809. (49) Ashik U P M, Wan Daud W M A, Abbas H F. Methane decomposition kinetics and reaction rate over Ni/SiO2 nanocatalyst produced through co-precipitation cum modified Stöber method [J]. Int. J. Hydrogen Energy, 2017, 42(2): 938-952. (50) Abbas H F, Daud W M A W. Deactivation of palm shell-based activated carbon catalyst used for hydrogen production by thermocatalytic decomposition of methane [J]. Int. J. Hydrogen Energy, 2009, 34(15): 6231-6241. (51) Abbas H F, Wan Daud W M A. Thermocatalytic decomposition of methane using palm shell based activated carbon: Kinetic and deactivation studies [J]. Fuel Process. Technol., 2009, 90(9): 1167-1174. (52) Shilapuram V, Ozalp N, Oschatz M, Borchardt L, Kaskel S, Lachance R. Thermogravimetric Analysis of Activated Carbons, Ordered Mesoporous Carbide-Derived Carbons, and Their Deactivation Kinetics of Catalytic Methane Decomposition [J]. Ind. Eng. Chem. Res., 2014, 53(5): 1741-1753. (53) Serrano D P, Botas J Á, Pizarro P, Gómez G. Kinetic and autocatalytic effects during the hydrogen production by methane decomposition over carbonaceous catalysts [J]. Int. J. Hydrogen Energy, 2013, 38(14): 5671-5683. (54) Yu J, Yue J, Liu Z, Dong L, Xu G, Zhu J, Duan Z, Sun L. Kinetics and mechanism of solid reactions in a micro fluidized bed reactor [J]. AlChE J., 2010, 56(11): 2905-2912. (55) Yu J, Zeng X, Zhang G, Zhang J, Wang Y, Xu G. Kinetics and Mechanism of Direct Reaction between CO2 and Ca(OH)2 in Micro Fluidized Bed [J]. Environ. Sci. Technol, 2013, 47(13): 7514-7520. (56) Yu J, Zeng X, Zhang J, Zhong M, Zhang G, Wang Y, Xu G. Isothermal differential characteristics of gas–solid reaction in micro-fluidized bed reactor [J]. Fuel, 2013, 103: 29-36. (57) Li Y, Chang L, Chen J. Catalytic growth of carbon fibers from methane on a nickel-alumina composite catalyst prepared from Feitknecht compound precursor [J]. Appl. Catal. A: Gen., 1997, 163(1-2): 45-57. 18


Page 18 of 34

Page 19 of 34 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


(58) Zeng X, Wang F, Wang Y, Li A, Yu J, Xu G. Characterization of Char Gasification in a Micro Fluidized Bed Reaction Analyzer [J]. Energy Fuels, 2014, 28(3): 1838-1845. (59) Wen C Y, Yu Y H. A generalized method for predicting the minimum fluidization velocity [J]. AlChE J., 1966, 12(3): 610–612. (60) Guo F, Dong Y, Fan P, Lv Z, Yang S, Dong L. Catalytic decomposition of biomass tar compound by calcined coal gangue: A kinetic study [J]. Int. J. Hydrogen Energy, 2016, 41(31): 13380-13389. (61) Gai C, Dong Y, Lv Z, Zhang Z, Liang J, Liu Y. Pyrolysis behavior and kinetic study of phenol as tar model compound in micro fluidized bed reactor [J]. Int. J. Hydrogen Energy, 2015, 40(25): 7956-7964. (62) Jeong S W, Son S Y, Lee D H. Synthesis of multi-walled carbon nanotubes using Co–Fe–Mo/Al2O3 catalytic powders in a fluidized bed reactor [J]. Adv. Powder Technol., 2010, 21(2): 93-99. (63) Chen D, Christensen K O, Ochoa-Fernández E, Yu Z, Tøtdal B, Latorre N, Monzón A, Holmen A. Synthesis of carbon nanofibers: effects of Ni crystal size during methane decomposition [J]. J. Catal., 2005, 229(1): 82-96. (64) And P E A, Rodríguez N M. Influence of the Support on the Structural Characteristics of Carbon Nanofibers Produced from the Metal-Catalyzed Decomposition of Ethylene [J]. Chem. Mater., 2000, 12(3): 823-830. (65) Amin A M, Croiset E, Epling W. Review of methane catalytic cracking for hydrogen production [J]. Int. J. Hydrogen Energy, 2011, 36(4): 2904-2935. (66) Kvande I, Chen D, Yu Z, Rønning M, Holmen A. Optimization and scale-up of CNF production based on intrinsic kinetic data obtained from TEOM [J]. J. Catal., 2008, 256(2): 204-214. (67) Kuvshinov G G, Mogilnykh Y I, Kuvshinov D G. Kinetics of carbon formation from CH 4 –H 2 mixtures over a nickel containing catalyst [J]. Catal. Today, 1998, 42(3): 357-360.

19



Table 1. Properties of α-Al2O3 particles. Solid particles α-Al2O3

dp(µm) 60

ρs (kg/m3) 3900

20


Umf (m/s) 0.0037

Page 20 of 34

Page 21 of 34 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


Table 2. Summary of experimental conditions and the initial reaction rate No.

Wcat/mg

T/°C

PCH4/atm

C(H2,0)/%

αCH4,0/%

, /mol•s-1•gcat-1

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16

15.9 15.1 15.4 15.5 15.3 15.8 15.4 15.6 15.6 15.5 15.5 15.5 15.8 15.6 15.9 15.7

750 750 750 750 800 800 800 800 850 850 850 850 900 900 900 900

0.22 0.50 0.75 1.0 0.22 0.50 0.75 1.0 0.22 0.50 0.75 1.0 0.22 0.50 0.75 1.0

4.30 8.02 20.1 28.76 7.54 17.3 31.5 45.16 9.31 20.5 33.5 46.86 11.02 23.6 38.2 49.26

2.20 4.45 11.17 16.79 3.92 9.45 18.69 29.16 4.88 11.42 20.12 30.60 7.90 13.38 23.61 32.68

2.71E-05 1.32E-04 4.86E-04 9.67E-04 5.03E-05 2.68E-04 8.13E-04 1.67E-03 6.15E-05 3.29E-04 8.64E-04 0.00176 7.15E-05 3.83E-04 9.94E-04 0.00186

21



Figure 1. XRD of fresh catalyst.

22


Page 22 of 34

Page 23 of 34 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


Figure 2. SEM of fresh catalyst.

23



Figure 3. Nitrogen adsorption/desorption isotherms for the fresh catalyst.

24


Page 24 of 34

Page 25 of 34 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


Figure 4. Three methods for CMD in MFBRA.

25



Figure 5. Typical methane decomposition test at the condition of 750oC and 0.75atm: method I; method II; method III.

26


Page 26 of 34

Page 27 of 34 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


Figure 6. The comparison between empty and with catalyst experiments in three different methods: (a) method I; (b) method II; (c) method III.

27



Figure 7. MS intensity of CH4 for two different methods for the initial 600s at the condition of 750oC and 0.75atm.

28


Page 28 of 34

Page 29 of 34 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


Figure 8. The reduction process of Fe2O3 catalyst during CMD.

29



a)

b)

c) Figure 9. SEM images of deposited carbon at the condition of 750oC and 0.75atm: fluidized bed (a) Method I (b) Method II and (c) Method III.

30


Page 30 of 34

Page 31 of 34 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


Figure 10. Hydrogen formation rate for different methane partial pressure (PCH4 =0.22, 0.50, 0.75, 1.0atm) at different temperature: (a) 750°C (b) 800°C (c) 850°C (d) 900°C.

31



Figure 11. (a) Dependence of initial specific molar rate at different methane partial pressure and temperature; (b) Representation of Arrhenius temperature dependency plot (Lnkp versus (1/T)).

32


Page 32 of 34

Page 33 of 34 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60


Figure 12. Activity versus time relationship at methane partial pressures of (a) 0.50atm and (b) 1.0atm.

33



Figure 13. (a) Logarithmic slop of Eq. (19) at different temperatures and (b) Representation of Arrhenius plot (Lnkd against 1/T) for catalyst deactivation.

34


Page 34 of 34

Methane Decomposition Kinetics over Fe - ACS Publications

Recommend Documents