Hydrogen Production by Methane Decomposition - A Review 2010

8/16/2019 Hydrogen Production by Methane Decomposition - A Review 2010

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Review

Hydrogen production by methane decomposition: A review

Hazzim F. Abbas*, W.M.A. Wan Daud

Department of Chemical Engineering, University of Malaya, 50603 Kuala Lumpur, Malaysia

a r t i c l e i n f o

Article history:

Received 19 September 2009

Received in revised form

9 November 2009

Accepted 9 November 2009

Available online 27 November 2009

Keywords:

Methane decomposition

Hydrogen production

Metal catalysts

Carbonaceous catalysts

Deactivation

Regeneration

a b s t r a c t

Methane decomposition can be utilized to produce COX-free hydrogen for PEM fuel cells, oil

refineries, ammonia and methanol production. Recent research has focused on enhancing

theproduction of hydrogen by thedirect thermocatalytic decomposition of methaneto form

elemental carbon and hydrogen as an attractive alternative to the conventional steam-

reforming process. In this context, we review a comprehensive body of work focused on the

development of metal or carbonaceous catalysts for enhanced methane conversion and on

the improvement of long-term catalyst stability. This review also evaluates the roles played

by various parameters, such as temperature and flow rate, on the rate of hydrogen

production and the characteristics of the carbon produced. The heating source, type of

reactor, operating conditions, catalyst type and its preparation, deactivation and regener-

ation and the formation and utilization of the carbon by-product are discussed and classi-

fied in this paper. While other hydrogen production methods, economic aspects and

thermal methane decomposition methods using alternative heating sources such as solar

and plasma are briefly presented in this work where relevant, the review focuses mainly on

the thermocatalytic decomposition of methane using metal and carbonaceous catalysts.

ª 2009 Professor T. Nejat Veziroglu. Published by Elsevier Ltd. All rights reserved.

1. Introduction

One of the major challenges posed by the continuous increase

in global population and economic development is providing

more energy while limiting greenhouse-gas (GHG) emissions.

The dramatic increase in the concentrations of carbon

dioxide, methane and nitrous oxides, the pH decrease of theocean surface and atmospheric temperature increase show

that human activity affects geochemical systems on a global

scale. The main sources of GHG emissions are due to the

combustion of natural gas (NG), coal and oil for heating,

electricity production, transportation and industrial purposes.

For example, oil accounts for 39% of hydrocarbon-related CO2emissions and NG for 20%, with coal accounting for the

remaining. The more efficient use of fuel has focused current

research on new energy sources that do not emit GHGs as well

as with the capture GHGs from the burning of hydrocarbon

fuels; both might be effective ways to gradually decrease the

quantity of GHG emissions [1,2].

Hydrogen appears to be one of the most promising energy

vectors as it is considered to be environmentally benign. The

amount of energy produced during hydrogen combustion ishigher than that evolved by any other fuel on a mass basis,

with a low heating value that is 2.4, 2.8 or 4 times higher than

that of methane, gasoline or coal, respectively. Currently, the

annualproduction of hydrogen is about 0.1 Gton, 98% of which

is from the reforming of fossil fuels; it is used mainly in oil

refineries, ammonia and methanol production [3]. The fuel

cell (a device transforming chemical energy into electricity

and heat) is a rapidly emerging technology. One area of great

* Corresponding author. Tel.: þ60 172907256; fax: þ60 103 79675371.E-mail address: [email protected] (H.F. Abbas).

A v a i l a b l e a t w w w . s c i e n c e d i r e c t . c o m

j o u r n a l h o m e p a g e : w w w . e l s e v i e r . c om / l o c a t e / h e

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0360-3199/$ – see front matter ª 2009 Professor T. Nejat Veziroglu. Published by Elsevier Ltd. All rights reserved.

doi:10.1016/j.ijhydene.2009.11.036

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potential for fuel cells is in powering automobiles; the

advantages of fuel cells are that they are quiet, do not contain

any hazardous material, and therefore can be classified as an

environmentally benign power source [4]. Research iscurrently being conducted for the development of safe, cost-

effective hydrogen production, storage and use technologies

that support and foster the use of hydrogen. Hydrogen

production from fossil fuels will continue for the foreseeable

future, given the large resource and the established industrial

base, and a number of papers have address advanced

hydrogen production technologies that reduce or eliminate

CO2 emissions in the production process [5].

In this review of the literature, reports relating to the direct

methane decomposition process for hydrogen production are

classified and summarized according to heating sources,

reactor types, catalyst (metal or carbonaceous) types, their

deactivation and regeneration and the form of the carbon by-product. It should be noted that there were also studies

focused on methane decomposition reactions using different

metal catalysts with the aim of produce carbon filaments (CF),

which is favored by operating conditions yielding low rates of

methane decomposition. These works were excluded from

the review because the focus here was on hydrogen

production.

2. Methods of hydrogen production:economic and environmental aspects

The current methods of producing hydrogen are based onsteam methane reforming (SMR), coal gasification, electrol-

ysis, biomass gasification and thermochemical processes.

Table 1 lists a summary of the energy efficiencies and costs of

hydrogen production for several methods [6].

A comparison between the cost of hydrogen production

from renewable sources, such as from water electrolysis using

renewable energy, and fossil sources, currently mainly

natural gas (NG), shows that the cost of renewably sourced

hydrogen must be considerably reduced from present levels

before this type of hydrogen becomes economically competi-

tive [7]. Currently, the process based on SMR and methane

partial oxidation generates large quantities of CO2; the esti-

mated GHG potential of hydrogen production by the SMR

process is approximately 13.7 kg CO2 (equiv.)/kg of hydrogen

produced [8].

Methane can be thermally or thermocatalytically decom-

posed into carbon and hydrogen without producing CO2, andthis hydrogen production method has recently attracted the

attentions of researchers. Lane and Spath [9] estimated that

hydrogen could be produced by the thermocatalytic decom-

position (TCD) of methane at a selling price of (7–21) $/GJ

(Note: 1 GJ ¼ 1.05461 MMBtu) depending on the cost of NG and

the selling price of carbon. In 1988, a techno-economic

assessment showed that the cost of hydrogen produced by

the thermal decomposition of NG, at $58/1000 m3 H2 (with

carbon credit; based on the low heating value of hydrogen,

this is equivalent to $5.38/GJ), was somewhat lower than that

for SMR, at $67/1000 m3 H2 (equivalent to $6.21/GJ) [10].

Steinberg [11] made a comparison between the SMR and

methane TCD processes for the decarbonization of NG andshowed that the TCD of methane appears to have several

advantages over a well developed SMR process, as the major

one being that it is much easier to sequester carbon in form

of the stable solid produced by methane TCD rather than as

the CO2 produced as a reactive gas or low-temperature liquid

from the SMR process. Dufour et al. [12] compared different

processes for hydrogen production (SMR, SMR with CO2capture and storage, methane autocatalytic decomposition

and methane thermal cracking) using life-cycle assessment

tools to evaluate their relative environmental feasibilities and

CO2 emissions. They reported that autocatalytic decomposi-

tion is the most environmentally friendly process for

hydrogen production as it presented the lowest total envi-ronmental impact and CO2 emission; although SMR with CO2capture and storage actually led to lower CO2 emissions, it

still had a higher total environmental impact than conven-

tional SMR.

Fig. 1 shows a comparison between the TCD of methane

and the commercial conventional SMR process. The energy

input requirements per mole of hydrogen for TCD is signifi-

cantly less than that of SMR (37.8 and 63.3 kJ/mol H2, respec-

tively). In terms of net hydrogen yield, although the

theoretical hydrogen yield for SMR is twice that of TCD (four

vs.2 mol of H2 per mole of CH4, respectively), the high reaction

endothermicity and required CO2 sequestration process, both

consuming significant amount of energy in methane

Table 1 – List of hydrogen production technologies and costs [6].

Production technology Energy efficiencya Hydrogen selling priceb ($/MMBtu)

Hydrogen selling price ($/kg)

SMR 83% 5.54 0.75

Partial oxidation

of methane

70–80% 7.32 0.98

Autothermal reforming 71–74% 16.88 1.93Coal gasification 63% 6.83 0.92

Direct biomass gasification 40–50% 9–18 1.21–2.42

Electrolysis (nuclear fission-powered) 45–55% 14.5 1.95

Photocatalytic water splitting 10–14% 37 4.98

c Assumes an NG price of $3.15/MMBtu [6].

a Energy efficiency is defined as the energy value of the hydrogen produced divided by the energy input required to produce the hydrogen.

b The hydrogen selling price for SMR and coal gasification does not consider CO2 sequestration costs.

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equivalents, would considerably reduce the net yield of

hydrogen produced by a ‘‘CO2-neutral’’ SMR process [13].

3. Thermocatalytic methane decompositionprocesses

Methane can be decomposed into carbon and hydrogen

according to the following reaction:

CH4/CðSÞ þ 2H2 DH

¼ 75:6kJ=mol

Because the process does not produce CO or CO2 as by-products, the need for the water-gas shift and CO2-removal

stages, as required in conventional hydrogen production

methods (e.g., SMR, coal gasification and partial oxidation), is

eliminated. Additionally, it is possible to reduce the cost of

hydrogen produced by methane TCD by marketing the carbon

as a filler or construction material [14].

3.1. Heating sources

Methane decomposition is a moderately endothermic reac-

tion. Due to the strong C–H bonds, non-catalytic thermal

cracking of methane requires temperatures higher than

1200 C obtain a reasonable yield. By using a catalyst, thetemperature can be significantly reduced, depending on the

type of catalyst used. There are many reports on the use of an

electrical furnace as a heating source for the TCD reactor,

while there are rather few focusing on the use of concentrated

solar energy, plasma or a molten-metal bath as alternative

heating sources. A recent report [15] pointed out that solar is

most likely to be the only source of energy capable of

producing very large volumes of hydrogen. Using solar

thermal power, methane thermal dissociation into hydrogen

and carbon black (CB), two valuable products, has been

investigated [16–21]. Concentrated solar energy is a clean

source of high temperature process heat and direct solar

irradiation of the reactants provides for a very efficient heat

transfer directly to the reaction site [22]. However, solar

heating cannot be achieved directly because hydrocarbons

absorb radiation in the visible spectrum only poorly. To

overcome this problem, the use of a transparent window that

permits direct heating of particulate material by radiation or

opaque heat-transferring reactor walls that absorb the solar

radiation on one side and then heat the gas through convec-

tion on the other side have been employed.Beside a solar source, plasma may also be employed as an

environmentally friendly heating source. To be effective, it

needs to be carried out at a very high reaction temperature,

which, with recent improvements in plasma technology, is

now accessible [23]. Microwave (MW) plasma (generated using

an MW-frequency signal; typically 2.45 GHz), commonly used

in MW ovens, diamond vapor deposition and IC

manufacturing, has the advantages of easy operation, an

electrodeless reactor, high plasma density and high electron

mean energy [24]. The major advantage of this new process is

the total conversion of the hydrocarbon into CB and hydrogen

(100% carbon yields and the production of new carbon grades

at higher temperature) [25]. Recently, using carbon aerosolsgenerated by a non-thermal plasma, Muradov et al. [26]

showed that TCD can be accomplished over catalytically

active carbon aerosol particles at temperatures comparable to

that of the conventional SMR process, i.e., 850–900 C,

a temperature range about 400–500 C lower than that for non-

catalytic methane decomposition. Application of pulsed MW

power was shown to be a promising heating source. However,

the choice of the catalyst for an MW-driven processappears to

be of great importance as it must meet some specific

requirements in addition to the common demands of high

activity, selectivity and stability. Such a catalyst must be

a good receptor of MW energy and able to retain its structure

and properties under intense MW radiation [27]. MW energyoffers a number of advantages over conventional heating:

(i) noncontact heating; (ii) energy transfer without heat

transfer; (iii) rapid heating; (v) volumetric heating; (vi) quick

start-up and shutdown; (vii) the heating starts from the inte-

rior of the material body; (viii) a higher level of safety and

automation; and (ix) heating energy can be transported from

the source through a hollow, nonmagnetic metal tube.

Muradov [28] discussed three heating-arrangement

options for the decomposition reaction (Fig. 2). In the first

option, the heat source is located inside the reaction zone

using a heat pipe, such as a heat exchanger or a catalytic

burner that uses NG or a portion of the hydrogen product as

a fuel. In the second option, externally heated catalyst parti-cles are used as the heat carrier, which is similar to the

fluidized catalytic-cracking process currently used in many

refineries. In the third option, a relatively small amount of

oxygen is added to the methane feedstock to generate the

necessary heat, in a process termed autothermal pyrolysis.

Table 2 lists and compared the different types of heating

sources employed to date for methane TCD.

3.2. Reactor types

Besides the heating source, another important factor is the

type of reactor equipment used. Fluidized-bed reactors (FBR)

and packed-bed reactors (PBR) are the most commonly used

Fig. 1 – Comparative assessment of net hydrogen yields for

SMR and TCD processes [13].



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reactors for methane TCD. Since the main products of

methane TCD are solid carbon and hydrogen gas, the problem

with using a PBR is carbon deposition over the external

surface of the catalyst particles. As the reaction proceeds, the

pressure drop increases, the catalyst particle size increases

and particle shape and density are altered. For experiments

conducted over a long duration, a PBR will be gradually filled

with solid carbon, eventually blocking reactant gas flow [29–

31].To maintain catalyst activity and to avoid plugging of the

reactor, the deposited carbon must be removed periodically.

The development of a reactor type that is efficient for

methane TCD with continuous withdrawal of carbon product

had been conducted by Muradov [13,32]. He studied different

types of reactors (PBR, FBR, free-volume reactor, spouted-bed

and tubular reactors) and found that the FBR was the most

promising reactor for large-scale operation since it provides

a constant flow of solids through the reaction zone, making it

suitable for continuous addition and withdrawal of catalyst

particles from the reactor. Additionally, the vigorous particle

motion caused by fluidization increases both heat- and mass-

transfer rates. The FBR superficial velocity used was in the

range of one to four times the minimum fluidization velocity

(Umf ). Fig. 3 shows some types of reactor used for catalytic and

non-catalytic thermal decomposition of methane.

One method of studying TCD parameters consists of using

a thermobalance to obtain rates of methane decomposition by

directly measuring the mass gain (due to carbon deposition)

with time. The advantages of using a thermobalance to study

methane TCD are as follows: (1) the ability to directly measure

the change of mass with time; as a result, the initial rate of methane decomposition (rO) can be used to determine the

intrinsic kinetic parameters and the change in catalytic

activity as the mass of deposited carbon increases; (2) a small

quantity of catalyst can be used to ensure that the endo-

thermicity of the decomposition reaction does not cause large

temperature fluctuations in the catalyst and negligible pres-

sure drop. The main drawbacks of using a thermobalance are

diffusion limitations and the difficulties of controlling space-

time due to the changes in sample volumes during the

experiments [37–40].

In most experimental work carried out for methane TCD

quartz was used as a reactor-construction material to avoid

any possible catalytic activity of the reactor wall on methane

Fig. 2 – Schematic diagram of hydrogen and carbon production via catalytic decomposition of NG with an internal (A) or

external (B) heat supply. 1. Fluidized-bed reactor (FBR); 2. Fluidized-bed heater-regenerator; 3. Internal heaters; 4. Cyclones;

5. Heat exchangers; 6. Gas-separation unit; 7. Carbon/metal catalyst separation unit; 8. Catalyst regeneration unit; 9. Carbonproduct separation and conditioning unit; 10. CO2 scrubber and gas-purification system; 11. Combustion chamber [28].



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Table 2 – Recent methane TCD studies using different heating sources and reactor types.

Researcher Heating source Reactor type

Abanades and

Flamant [43,44]

Concentrated solar energy. Fluid-wall reactor, Stainless-steel reactor

body and with internal graphite tube.

* No catalyst used.

* Temperature measured with IR solar-blind optical pyrometer.

* Vertical-axis solar furnace (200-cm diameter concentrator) used.* Co-products are hydrogen-rich gas and high-grade CB.

* Methane conversions of 88% and 42% were measured for a 2000-kW/m2 solar-flux density.

Hirsch and Steinfeld [33] Concentrated solar energy Vortex-flow reactor, 10-cm diameter

and 20-cm length, made from steel

alloy with a 6-cm diameter aperture.

* Vortex-flow of methane confined to a cavity receiver and laden with carbon particles that served simultaneously

as radiant absorbers and nucleation sites for the heterogeneous decomposition reaction.

* A 5-kW reactor prototype with solar-flux intensities exceeding 3500 kW/m2 was used.

* Conversion of methane was 67% at 1327 C and 1 atm.

Dahl et al. [45] Concentrated solar energy Quartz-tube reactor, 2.5-cm diameter,

consisting of a feed mechanism with

internal graphite ‘‘target’’ feed tube.

* Series of mirrors used to concentrate sunlight into a focused beam at a maximum level of 10 kW onto anapproximate diameter of 10 cm.

* Secondary concentrator used to obtain higher temperature.

* Quartz reactor tube illuminated with a solar flux of 2400 kW/m2 resulted in a methane conversion of 90%.

Dahl et al. [34] Concentrated solar energy Quartz fluid-w all aerosol-flow reactor

composed of three concentric vertical tubes.

The innermost tube was 1.2-cm i.d., long,

porous graphite, the center tube was a

2.1-cm i.d., solid graphite tube 360 cm long.

* No evidence of corrosion or erosion of the reactor tube reported.

* 90% conversion of methane was obtained at 1860 C and a residence time of 0.01 s.

* Due to the small size of reactor, CB co-feed did not enhance the heat transfer.

Maag et al. [46] Concentrated solar energy 5-kW solar chemical reactor consisting

of a 20-cm-long, 10-cm-diameter cylindricalcavity receiver containing a 6-cm-diameter

circular aperture to allow entry of

concentrated solar radiation.

* Reactor tested at temperature range of 1027–1327 C.

* Methane conversion and hydrogen yield exceeding 95% were obtained at residence times of less than 2 s.

* A solar-to-chemical energy conversion efficiency of 16% was experimentally reached.

* SEM images revealed the formation of filamentous agglomerations on the surface of the seed particles.

Muradov et al. [47] Electric furnace FBR, quartz microreactor.

TCD process could be arranged in continuous process similar industrial to fluidized catalytic-cracking or coking processes.

Dunker et al. [36] Three zone electrical

tubular furnace.

FBR made from quartz tube, 4.2 cm i.d.

and 97 cm height

* Reactor constructed from quartz and the thermocouple sheathed with quartz tube.

* Fine carbon particles deposited on the reactor walls and polycyclic aromatic hydrocarbon condensed in the lines downstreamof the reactor.

Aiello et al. [29] Electric furnace Fixed-bed, quartz microreactor,

1.8-cm o.d.

The flow through the reactor was substantially hindered during the eighth and ninth production cycles.

Fidalgo et al. [48] Electric furnace and MW oven. Quartz reactor, 45 cm long and 2.2-cm i .d.

* Activated carbon (AC) was used as both MW receptor and catalyst.

* Under electric furnace heating, nitrogen distributed the methane molecules within the AC bed.

* Under MW heating, the nitrogen, as well as distributing the methane molecules, favored the generation of energetic

microplasmas, leading to higher conversions, with the highest obtained at low CH4 /N2 ratios.

* The formation of carbon nanofibers (CNFs) was reported when a combination of nitrogen and MW heating was used.



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decomposition (e.g., transition metals have catalytic activityfor methane decomposition). Additionally, the thermocouple

used was sheathed with a quartz tube. However, there have

been a few researchers who used kanthal or stainless-steel as

reactor-construction materials [30,41,42]. Table 2 presents

examples of recent studies using different reactor types and

heating sources along with major findings and notes.

Different methods can be used to separate the carbon by-

product from the hydrogen stream. For the CB process, the

particulate carbon elutriates out of the furnace with the

hydrogen gas stream and is separated by bag filters. A similar

system is used in the plasma carbon -arc process. In the

molten-metal reactor it may be possible to skim the carbon off

from the top of the metal surface due to its density difference[11]. Fig. 4 is a simplified diagram showing the types of heating

source, reactor, catalyst and technique used to analyze the

reactor effluent stream and characterize the fresh and deac-

tivated catalysts in methane TCD studies to date.

4. Metal catalysts for thermocatalyticdecomposition of methane

4.1. Operating parameters and kinetics

As mentioned previously, several types of catalysts have been

used to reduce the decomposition temperature; it has been

reported that in the presence of freshly reduced Ni catalyst,hydrogen was detected in the effluent gas at a temperature as

low as 200 C [5]. Ermakova [53] stated that the methane TCD

process can hardly be of practical interest unless highly effi-

cient catalysts are developed. Catalyst efficiency includes not

only its specific activity but its useful operational lifetime

given the large amount of carbon accumulation at a practical

conversion. The use of metal catalysts such as Ni and Ni–Cu

was first reported by Dudina et al. [54], Muradov [55–57],

Chesnokov et al. [58,59] and Parmon et al. [60]. There is no

general agreement among researchers regarding the relative

catalytic activities of metals in methane TCD. It has been

reported that the rate of methane decomposition activity by

the transition metals follows the order: Co, Ru, Ni, Rh > Pt, Re,Ir > Pd, Cu, W, Fe, Mo [40]. However, other researchers have

found that Ni or Ni/alumina and Fe/alumina exhibited the

highest activities [61].

The most important factors influencing carbon deposition

during metal-catalyzed methane decomposition are the

particle size, dispersion and stabilization of the metallic

catalyst particles, which are controlled by selecting an

appropriate support. Takenaka et al. [62] reported that

a typical 40% Ni/SiO2 catalyst with a nickel-particle size of 60–

100 nm could give carbon yield of as high as 491 g C (gNi) 1

during methane decomposition at 500 C. Furthermore,

Ermakova et al. [63] reported that a 90 wt% Ni/SiO2 catalyst

with nickel particles of 10–40 nm size provided carbon yields

Table 2 ( continued ).

Researcher Heating source Reactor type

Domı́nguez et al. [49] MW oven and electrical heating. Fixed-bed, quartz tube, 2.2-cm i.d.

and 45-cm length.

* Granular AC was used as a catalyst.

* Methane conversions were higher under MW conditions than with electrical heating when the temperature was 800 C.

* The formation of CNFs in one of the MW heating experiments was reported.

Fulcheri et al. [25] Plasma power supply (three-phase)

supplied electricity to three graphite

electrodes located at the top

of the reactor.

Three graphite electrodes located at top

of reactor, and a graphite nozzle to mix

feedstock with plasma gas flow.

Reactor height 250 cm.

* The reactor internal temperatures were measured at four locations using an optical pyrometer sighted

on a graphite tube immersed in the plasma flow.

* Experiment was carried out with electric power varying between 50 and 100 kW.

* Presented a new alternative to normal furnace processes for CB production.

Serban et al. [50] Heat generated in generation

IV nuclear reactor.

Stainless-steel reactor, 2.54-cm diameter

and 35-cm height.

* Methane was bubbled through a bed of molten metal or granular or catalytic material or a mixture of molten metal and solid media.

* Bubbling methane through porous metal filters was the most efficient.

* Buoyant separation of generated carbon from the liquid heat-transfer media was reported.

Pinilla et al. [51] Electric furnace. Cylindrical drum rotating around its

horizontal axis. Diameter 0.065 m,

length 0.8 m and rotation speed 1–20 rpm.

* In comparison to FBR, the rotary reactor showed higher hydrogen yields and more sustainable catalyst.

* Rotation speed did not show any significant influence on the evolution of hydrogen concentration.

Chesnokov and

Chichkan [52]

Electric furnace. Rotary reactor: volume 250 cm3 and

rotation speed of 1 rpm.

* Rotary reactor was developed for production of hydrogen and CNF from natural gas.

* Compared to the reactor with a McBain balance, catalyst-operation time in the rotary reactor was very long.



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on the order of 385 g C (gNi)1 at 550 C. Other supports, suchas TiO2, MgO, ZrO2 and Al2O3 gave relatively lower carbon

yields [64]. For the production of highly concentrated

hydrogen during methane decomposition, higher one-pass

conversion of methane is desirable. Ogihara et al. [65] showed

that the equilibrium conversion of methane is a function of

the reaction temperature for methane decomposition. The

equilibrium conversion of methane was estimated, assuming

that the carbon was formed as graphite (CH4 4 2H2 þ C

(graphite)). Theyreported that methane conversion was not as

high when methane decomposition was performed at

temperatures below 600 C. Using a commercial Ni-based

catalyst, Suelves et al. [66] reported that, at a temperature of

700 C, the concentration of hydrogen was around 80%, which

corresponds to a methane conversion close to the theoreticalequilibrium value.

The TCD of methane on various Ni-containing catalysts

(Mn, Fe, Co or Cu) was investigated by Zabidi et al. [30],

showing that Ni/Mn based catalysts gave better results for

hydrogen production with carbon nanotubes (CNTs) as the by-

product. Choudhary et al. [67] studied the effect of using

different Ni-containing metal oxide (ZrO2, MgO, ThO2, CeO2,

UO3, B2O3 or MoO3) and zeolite (HZSM-5, Hb, HM, NaY,

Ce(72)NaY or Si-MCM-41) catalysts and concluded that

Ni/ZrO2 and Ni/Ce(72)NaY showed the most promising results

for a cyclic process. The effects on the catalytic activity of Ni

using different kinds of zeolites as supports were investigated

by Inaba et al. [68], who found that Ni-supported on USY

Fig. 3 – Schematic diagram of several types of reactor used for catalytic and non-catalytic thermal decomposition of

methane. a) Vortex-flow reactor confined to a cavity receiver [33]. b) Solar-thermal fluid-wall reactor [34]. c) Multilayer

reactor [35]. d) FBR [36].



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zeolite (Si/Al2¼ 14,360) showed a longer catalytic lifetime, and

considered it to be the best catalyst for hydrogen production

by methane TCD. In another study, methane TCD at 550 C

over Ni-supported on HY, USY, SiO2 and SBA-15 was con-

ducted and among all the catalysts tested the Ni/HY catalyst

was found to have the highest activity, as shown in Fig. 5.

Additionally, the total run time increased with the various

catalyst supports in the order: HY w SiO2 > USY > SBA-15.

However, the overall accumulation of carbon was in the order:HY > USY > SiO2 > SBA-15 [69].

Catalysts containing Ni and Fe have been widely tested in

the past. It was reported that the Ni-based catalysts have

a maximum operating temperature of 600 C; thus, as

methane conversion is thermodynamically limited at this

temperature, concentrated hydrogen streams (H2>60%)

cannot be obtained using a nickel-based catalyst [70]. On the

contrary, Fe-based catalysts are more stable at higher

temperatures (700–1000 C), but deactivation occurs during

repeated cycles, resulting in a shorterlifetime.Chesnokov and

Chichkan [52] reported that the modification of a 75%Ni–

12%Cu/Al2O3 catalyst with Fe made it possible to increase

optimal operating temperatures to 700–750 C while

maintaining excellent catalyst stability, and the hydrogen

concentration at the reactor outlet exceeded 70 mol%. The

structure of the filamentous carbon formed upon catalytic

decomposition of hydrocarbons on iron-subgroup metals and

their alloys, was reported by Chesnokov and Buyanov, along

with generalizations on the regularities of carbon deposition

on these metals and a discussion of the growth models of

some morphological modifications of filamentous carbon [71].

Fe catalysts can decompose methane at a temperature rangeof 700–1000 C, but Fe catalysts have a very short lifetime [72].

For example, Ermakova et al. [73] reported that Fe catalysts

were effective for methane decomposition in the temperature

range of 650–800 C; however, the hydrogen yield (H2 / M:molof

H2 formedper mol ofmetalscontainedin the catalyst)for Fe at

800 C catalysts (H2 /Fe ¼ 418) was more than an order of

magnitude lower than that for Ni catalysts at 500 C (H2 /

Ni ¼ 4802) [62].

Nuernberg et al. [74] investigated the influence of operating

conditions on the catalytic performance of Co supported on

Al2O3. They studied the effect of different Co contents at

a temperature range of 600–800 C and different methane

partial pressures. It was found that the best conditions for

Fig. 4 – Simplified diagram of types of heating source, reactor, catalyst and analysis apparatus used in methane TCD studies.



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hydrogen production by methane TCD were at 800 C with

a nitrogen-to-methane molar ratio of 6:1 and a Co loading of

20 wt% (Fig. 6). Muradov and Veziroglu [75] summarized the

bulk of the available literature data on a number of these

catalysts; their preferred temperature ranges and carbon

products in methane TCD are presented in Fig. 7.

Ni–Cu/Al2O3 catalysts have been shown to have a number

of advantages over Ni/Al2O3 catalysts [76,77]. Ni–Cu/Al2O3catalysts with high metal loading had optimum operating

temperatures in the range of 600–675 C, making it possible to

achieve higher methane conversion. Similar results were alsoreported for methane TCD over a Ni–Cu–MgO catalyst to

produce hydrogen and CNF; Wang and Baker [78] found that

a Ni–Cu–MgO catalyst maintained its activity at high levels for

substantially long periods of time at 665–725 C, and the

catalyst was capable of generating large amounts of CO-free

H2 and solid carbon. In another report, 75% Ni-15% Cu–Al2O3catalysts with a nickel-particle size of 27 nm gave a carbon

yield of 700 g C (gNi)1 at 625 C [79].

To produce highly concentrated hydrogen and CNF bymethane TCD in a temperature range of 700–850 C, Ogihara

et al. [65] investigated M/Al2O3 (M]Fe, Co, Ni and Pd) and

Pd-based alloys containing Ni, Co, Rh or Fe. They found that

the conversion of methane at 700 C for Fe, Co and Ni/Al2O3catalysts was very low (


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Table 3 – Recent studies on metal catalysts for methane TCD, preparation methods and major findings.

Researcher Catalyst composition Catalyst preparationtechnique

Operating conditions T/FlowPhysical properties W/SA/MPS

Aiello et al. [29] 15wt% Ni/SiO2. Wet Impregnation. 650/30,000 h1 0.2/—/20–35 mesh

* 3 h on stream.

* Methane initial conversion (MIC) was about 24–35%, depending on the number of cycles.

* Deactivated catalyst was regenerated with steam.* Ten cycles without significant loss of catalytic activity.

* A small amount of carbon, concentrated in small pockets, resisted gasification.

Suelves et al. [66] 65 Wt%Ni/Silica and alumina. – 550–700/20, 50 and 100 mL/min

2 and 0.3/190/100and 2000

* At 700 C, hydrogen concentration was 80%, catalyst activity did not decay after 8 h on stream at a space-time of 1 s.

* Range of sustainability factor was 0.14–1.

Venugopal et al. [82] (5–90)Ni/SiO2. Wet Impregnation. 600/24,360 h1

0.15/———/———

* Fixed-bed vertical quartz reactor was used (diameter 0.8 cm and length 46 cm).

* The 30 wt% Ni/SiO2 demonstrated superior activity and longevity and produced CNFs at 303 mol/molNi.

* XRD result showed that only NiO phase was present in the fresh catalysts while the deactivated catalysts displayed both metallic nickel and

graphitic carbon phases.

Ermakova et al. [53] 90Ni–10Al2O3,

90Ni–5ZrO2–5SiO2, and other catalyst

structures.

Impregnation, fusing followed by

decomposition.

550/120,000 mL/(g $h)

0.01/–/–

* Depending on calcination temperatures (400–800 C), the range of SA was 5–145 m2 /g.

* Depending on textural parameters of the catalytic systems, the ranges of carbon yield and methane conversion were 287–384 g/g Ni and 12.9–

15.3%, respectively.

Monnerat et al. [83] Ni-gauze – 410–550/75 mL/min

0.207/26.7/—

* Catalyst deactivated due to intensive coke deposition in the form of CF.

* Maximum cycle time was 4 min for better hydrogen performance.

* 10% of catalyst mass lost after 70 h on stream.

Zabidi et al. [30]. Ni/M-based (M]Mn, Fe, Co or Cu) – 550–900/3200, 9600,

16,000 h1

–/8.79 to 4.79/—

* Ni/Mn maintained more activity (59% methane conversion) for 120 min than the other catalysts.

* For Ni/Mn at 900 C, an increase of space velocity from 3200 to 16,000 h1 resulted in a decrease in hydrogen yield from 96% to 54%.

* Reactor blocking was reported for Ni/Co after 120 min on stream.

Choudhary et al. [42] Ni-containing different metal

oxides and zeolites.

Co-precipitating 500/7000 mL/(g $h).

0.4/266–786/30–52 mesh size

* Two parallel reactors operated in cyclic manner.

* Deactivated catalyst was regenerated by steam.

* Ni/ZrO2 and Ni/Ce(72)NaY showed promising results for the cyclic process.

* Methane conversion (w28%), hydrogen productivity (w31%) and CO2 produced passed through a maximum at a switchover time of 10 min.

Rahman et al. [38] 5 wt%Ni/g- Al2O3 Wet impregnation 500–650/——

0.1–1/——/—

* The study was conducted using a thermobalance.* Partial catalyst regeneration with oxygen was recommended because full regeneration destroyed its activity.

* Activation energy (Ea ) was 46 kJ/mol.

Zhang and Amiridis

[84]

16.4 wt%Ni/SiO2 Wet impregnation 550/15,000–37,500 h1

0.2/250/25–35 mesh

* When silica was used, no measurable methane conversion was observed at temperatures below 800 C even at a low space velocity (6000 h1).

* Methane initial conversion (MIC ) was 35% at 550 C and 30,000 h1.

* Nickel supported on silica was found to be active, producing stoichiometric amounts of hydrogen and carbon.

* The deactivated catalyst could be fully regenerated by either oxidation in air or steam gasification of the deposited carbon.

Bai et al. [85] 6.7 wt%Ni/AC from coal. Impregnation 550–850/50 mL/min

0.2/——/——



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* Very small amounts of hydrocarbons other than methane (40%) was reached at w800 C.

* At temperature over 900

C, the homogeneous decomposition of methane was significant and the conversion of methane rapidly increasedafter complete deactivation of the catalyst.

* The attempted catalyst regeneration by the removal of the carbon deposited on the external surface of catalytic particles by attrition failed.

Jang and Cha [88] Fe (iron powder) and (5–20) wt%

Fe/Al2O3 catalyst.

Fe/Al2O3 catalyst prepared by

impregnation.

600–1000/up to 50,000 h1

—/———/359

* MIC was w100% at 1000 C and 5000 h1.

* PBR and FBR reactors were used.

* The methane conversion rate was maintained by attrition of the produced carbon on Fe catalyst surface in an FBR.

Qian et al. [80] Reduced and unreduced

Co/Mo/Al2O3 Ni/Cu/Al2O3

Co-precipitation 550–850/50 h1

—/——/———

* For the reduced Co/Mo/Al2O3, the methane conversion was


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temperature was increased to 740 C with a ramp of

10 C min1 (a pre-induction reaction; PIR). The morphologies

and the sizes of the CNFs formed are similar to those formed

during CTR at 543 C. However, the shapes of the metal

particles are different. Some of them have the features of theparticles in the CTR at 543 C, but with a smoother surface.

The others are more pear-shaped, but their tails are not as

sharp as those in Fig. 8 D–F. They also showed that at constant

temperatures the catalyst showed quasi–stable activity below

640 C for a notable time period, while its performance

decayed rapidly above this temperature, and a quick deacti-

vation was observed at 740 C. Pinilla et al. [97] studied the

activity of a Ni–Cu–Al catalyst using a thermobalance and

reported that the methane and hydrogen reduced the effect of

thermal sintering on Ni and lowered the surface energy by

chemisorption. In addition, they showed that the carbon

deposited on the Ni leading face probably diffused to the Ni

trailing face superficially, and not through the bulk Ni.Nuernberg et al. [74] attributed the deactivation of a Co-

alumina catalyst to the methane molecules which were

initially adsorbed (by dissociative adsorption) and decom-

posed on the metal surface, resulting in the formation of

chemi-sorbed carbon species; the carbon species then dis-

solved and diffused through the bulk of the metal particle.

They concluded that catalyst deactivation occurred when the

rate of carbon diffusion through the metal catalyst particle

was slower than that of the formation of carbon at the surface

of the CoO sites. Under these circumstances, carbon builds up

at the catalyst surface and eventually encapsulates the metal

particle causing a loss of activity. In an evaluation of co-

precipitated Ni-alumina and Ni–Cu–alumina catalysts,

Avdeeva et al. [98] revealed that the nickel in deactivated Ni

catalysts was found in a metallic state, while in the Ni–Cu

samples about half of the nickel was atomically dispersed in

carbon. The catalyst deactivation mechanisms suggested

included the fragmentation of the metal particles as well asatomic erosion in the Ni–Cu samples.

With regard to the deactivation of metal catalysts, Aiello

et al. [29] reported that decomposition of the hydrocarbon at

the gas–metal interface was followed by dissolution of carbon

into the metal and diffusion through the particle. The carbon

then precipitated at the metal–support interface, detaching

the metal particle from the support and forming a filament

with an exposed metal particle at its tip. The rate-determining

step of this process is believed to be the diffusion of carbon

through the metal particle. This mode of carbon accumulation

allows the catalyst to maintain its activity for an extended

period of time without deactivation. They also showed that

the cracking of methane over Ni/SiO2 has indicated thatthousands of carbon atoms can be deposited via this process

on the catalyst per surface nickel atom. Eventuallythough, the

catalyst is deactivated due to the limitations imposed by the

available free space in the reactor.

5. Carbonaceous catalysts forthermocatalytic decomposition of methane

Many industrial catalysts consist of metals or metal

compounds supported on an appropriate support. The main

reason is to maintain the catalytically active phase in a highly

dispersed state. Alumina, silica and carbon (mainly AC) are




* Temperatures used for calcination were 450, 600, 800 and 1000 C.

* Ni–Cu–Mg catalysts showed a high and almost constant hydrogen production yield.

* The structural properties of the obtained CNF were highly dependent on the presence of Cu and on the calcination temperature.

Suelves et al. [91] Ni–Al and Ni–Cu–Al Co-precipitation, impregnation and

fusion.

700/20 mL/min

0.3/——/——

* Hydrogen production was not highly dependent on the preparation method.

* The presence of Cu as a dopant in Ni–Cu–Al catalysts enhanced the catalytic activity substantially.

* Ni–Cu–Al catalysts enhanced the formation of a well-ordered graphitic carbon, while Ni–Al catalysts promoted the deposition of a turbostratic

carbon.

Lázaro et al. [92] Ni–TiO2 and Ni–Cu–TiO2 Impregnation and fusion 700/20 mL/min

0.3/——/——

* Those catalysts prepared by fusion and those including copper in their composition both showed enhanced catalytic activity.

* The calcination temperature only slightly affected the hydrogen concentration of the outlet gas from the reactor.

* The deposited carbon appeared as long nanofilaments or uniform coatings on the catalyst particles, depending on the nickel-particle size.

Echegoyen et al. [93] Ni–Mg and Ni–Cu–Mg Co-precipitation, impregnation and

fusion.

700/20 mL/min

0.3/——/——* The presence of copper enhanced hydrogen production and the best results were obtained for Ni–Cu–Mg catalysts prepared by the fusion

method.

* For the Ni–Mg catalysts, the nickel-crystal size influenced catalysts performance in that the highest crystallite size gave the lowest the

hydrogen yields.

* In the Ni–Cu–Mg catalyst the Ni particle size was not a significant influence on hydrogen yields.

W, weight of catalyst (g); SA, surface area (m 2 /g); MPS, mean particle size (mm, unless other units are stated); T, temperature (C); F, either flow

rate (mL/min) or space velocity (h1) unless otherwise stated; —, not mentioned in the original paper.



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and no sulfur poisoning; and (3) the carbon formed can be

used as a catalyst precursor [103–105]. In the case of methane

TCD, furtheradvantages of using carbonaceous catalysts cited

include: (1) it can be catalyzed by carbon produced in the

process, so an external catalyst would not be required (except

for the start-up operation); and (2) the separation of carbon

product from the carbon catalyst is not necessary [75].

Various carbon materials have been used for methane TCD

such as AC, CB, glassy carbon, acetylene black, graphite, dia-mond powder, CNT and fullerenes [47]. Among these carbon

materials, most of the studies have focused on AC (manu-

factured from different carbon-based sources) and CB due to

their activities and good stabilities. The most important

characteristics of an industrial catalyst are the chemical

composition, surface area, porosity and pore-size distribution

(if the catalyst is porous), stability and mechanical properties.

5.1. Activated carbon

A number of different lignocellulosic precursors, e.g., coconut,

almond, peach, plum, olive, palm and cherry, are used as rawmaterials for manufacturing AC. It is widely assumed that the

porosity of AC depends not only on the raw material used as

the precursor but also on the manufacturing process (which

involves carbonization and activation) [106]. According to the

IUPAC classification, pores can be classified into three cate-

gories, namely, micropores (50 nm).

Daud and Ali [107] studied the effects of burn-off on the

micropore, mesopore and macropore volumes with palm and

coconut shells as raw materials. The carbonization and acti-

vation processes were carried out at 850 C using an FBR. They

concluded that within the burn-off range studied (up to 70%),

the micropore and mesopore volume created in palm-shell-

based AC were always higher than those of coconut-shell-

based AC. Achaw and Afrane [108] studied the development of

pore structure in AC produced from coconut shells and found

that the modification of the pore structure was observed to

begin at the drying stage and continued into the activation

stage. The effects of using different activating agents (CO2,

H2O, KOH, H3PO4 and ZnCl2) on the physical properties (bulk

density, porosity, pore-size distribution and surface area) of the AC produced have been investigated [109–112]. The

difference between using H2O and CO2 as an activating agent

is that CO2 produces an opening of narrow micropores fol-

lowed by their widening, whereas water vapor widens the

microporosity from the early stages of the process, leading to

the production of a lower micropore volume in the AC product

[113]. The different methods used for physical and chemical

modifications of AC surfaces have been reviewed by Yin et al.

[114].

When using AC as a catalyst for methane TCD, it has been

reported that the actual value of the highest ultimate mass of

carbon that AC can accumulate is significantly smaller than

the theoretical value (ifone assumesthe entire pore volume inACPS were completely filled with deposited carbon), which

indicatesthat only a fraction of the pore volume is filled by the

deposited carbon and the lower capacity could be attributedto

the blockage of the pore-mouth, which would hinder the

internal diffusion of methane molecules [39,40]. It has been

reported that the capacity of AC to accumulate deposited

carbon before deactivation is related to the surface area and

pore-size distribution. Additionally, microporous carbons

with high contents of oxygenated surface groups exhibited

high initial rates of methane decomposition (rO) but rapidly

became deactivated, while the use of mesoporous carbons

with high surface area resulted in a more stable and sustain-

able process [103]. Suelves et al. [39] reported that the amountof carbon deposited at deactivation showed a linear relation-

ship with the total pore volume of the fresh catalysts.

Most ACs posses carbon-oxygen functional surface groups

such as R–COOH, R–OCO, R–OH and R]O. Studies have been

carried out to investigate the effects of these surface func-

tional groups on methane TCD. Muradov et al. [47] compared

the catalytic activity of virgin AC (lignite) samples with that of

AC (lignite) pretreated with pure hydrogen at 850 C to remove

carbon-oxygen groups from the carbon surface. They found

that although carbon-oxygen surface groups may play a role

in methane dissociation, particularly at the initial stage of the

process, the catalytic activity of carbons, in general, could not

be solely attributed to the presence of such surface functionalgroups. Using ACs with different textural properties and

surface chemistries, Moliner et al. [103] showed that no

correlation could be found between the CO and CO2 evolved

and the long-term behavior andthat rO seemedto be relatedto

the content of oxygenated groups (Fig. 9). Suelves et al. [39]

found a good correlation between rO and the concentration of

oxygenated groups desorbed as CO in a temperature-pro-

grammed-desorption experiment.

ACs usually contains 1–12% ash (inorganic constituents).

Kim et al. [115] studied the effect of ash content on the activity

of AC in methane decomposition. They found that the ash

content in the ACs from coconut shell was about 1.8–4.4 wt%

while that in the ACs from coal ranged from 6.3 to 10.2 wt%,

Fig. 9 – Initial activity for methane conversion and carbonaccumulated as a function of the concentration of

oxygenated groups in the fresh catalyst; T [ 850 8C [103].



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and the major elements in the former were K, Na, Si and Mg,

while in the latter they were Al and Si. These elements are

known to be almost completely inactive for the decomposi-

tion of methane (although Fe, found in trace amounts, is

known to be active). They also reported that the surface area

of ACs were reduced and the initial activities decreased byabout 20% when the ash was removed from AC. Muradov et al.

[47] conducted controlled experiments with the inorganic ash

produced from the lignite-derived AC and showed that these

metal impurities had no particular catalytic activity in carbon-

catalyzed methane decomposition.

5.2. Carbon black

Commercial CBs such as black pearls (BP) 110, 120, 1300, 2000,

Regal 330, Vulcan XC72, Fluka 03866 and Fluka 05120 have

been used as carbonaceous catalysts for methane TCD

[39,47,116,117]. In contrast to the ACs, which showed accept-

able rO but were rapidly deactivated due to the blocking of

pores by growing crystallites, which hinder the internal

diffusion of methane molecules, the majority of CB surfaces

are relatively easily accessible to methane molecules during

methane decomposition. CBs differ in particle size, average

aggregate mass, morphology, etc. (for example, the oil-

furnace process produces CBs with particle diameters in therange 10–250 nm, and surface areas of 25–1500 m2 /g) [118]. It

has been reported that the catalytic activity of different CB

samples remains almost unchanged despite the significant

reduction in surface area during the reaction [119]. Using CB,

Lázaro et al. [104] showed that the amount of surface

complexes desorbed as CO, as well as the surface area ( Fig. 10

A), decreased gradually during the stable period of catalyst

activity. In addition, it was found that the total pore volume

decreased (Fig. 10B) as the carbon was deposited and the

degree of graphitization increased as the reaction progressed.

In a kinetic and deactivation study of carbon catalysts,

Serrano et al. [120] showed that the most active catalyst at

short reaction times was AC, but it underwent a fast

Fig. 10 – Change of the BET area (a) and total pore volume (b) versus reaction time at T [ 950 8C and GHSV [ 360 hL1 [104].



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deactivation due to the deposition of carbon in methane TCD;

contrastingly, CBs presented high reaction rates at both short

and long reaction times. The authors showed that methane

TCD using CB could go on for several hours until most of the

surface was covered by carbon crystallites. A quasi-steady

state quickly reached with CB and could remain constant for

more than 8 h regardless of the reaction temperature due to

the continuous changes in carbon morphology [117,121].

5.3. Operating parameters and kinetics

As mentioned previously, among carbon materials, the focus

has been on using AC and CB as carbonaceous catalysts for

methane decomposition. The operating parameters that have

been studied for the activity and long-term stability of these

catalysts are temperature range (700–1000 C), pressure

(usually atmospheric) and methane flow rate. These studies

also included physical properties (surface area, mean particle

size, composition, porosity and pore-size distribution) of the

virgin and deactivated catalysts. The weight of catalysts used

ranged from few milligrams to 25 g, and the most employedreactor types were either PBR or FBR, generally constructed

from quartz. Kinetic studies were also conducted to determine

the reaction order and activation energy (Ea) of methane

decomposition and the sustainability factor of the catalyst,

which is defined as reaction rate after1 h on stream divided by

ro. The effect of operating conditions on carbonaceous cata-

lysts and their changing physical properties will be discussed

in the next subsection. Table 4 lists recent methane TCD

studies carried out with carbonaceous catalysts, including the

catalyst types, operating conditions and the major findings.

5.4. Deactivation of carbonaceous catalysts

The reason for the gradual deactivation of catalysts is due to

the deposition of carbon (produced from the decomposition of

methane) on the catalyst surface, which results in blocking of

active sites and a reduction in catalyst surface area. The

deposit has lower surface area and activity compared to the

original carbon catalyst. Additionally, catalyst activity is

influence by its structure. The carbon produced by decompo-

sition of methane has a more ordered structure than amor-

phous carbons, but they are less structurally ordered than

graphite. The catalytic activity of carbons for methane

decomposition varies according to their structure in the

following order: amorphous > turbostratic > graphite. It is

assumed that the process of new carbon build-up can besimply divided into two steps: formation of carbon nuclei

(characterized by high catalytic activity) and carbon-crystal-

lite growth (these have a structure close to that ofgraphite and

thus the lowest catalytic activity). The formation of carbon

nuclei has an Ea of 316.8kJ/mol, which is higher than the Ea for

carbon-crystallite growth, at 227.1 kJ/mol. The total deacti-

vation rate is the sum of both steps, and, in general, the rate of

carbon-crystallite growth tends to be higher than the rate of

nucleus formation [5,47].

In a deactivation study of AC, Kim et al. [124] showed that

carbon-nucleus formation appeared to occur initially but was

quickly terminated and then the carbon-crystallite growth

become dominant. They explained the possibility of a linear

relationshipbetweencatalystactivityandtheamountofcarbon

deposited by reasoning that a pore would not have a uniform

diameter but instead havea complicated structure withnarrow

passages, wide passages, large cavities and interconnections.

During nucleation, a uniform deposition may occur since the

nucleiare notyet largeenough to block thepore,but nucleation

is terminated when crystallite growth becomes dominant. As

the crystallites grow, narrow passages are blocked sooner andthe inner active surface may not be accessible by the reactant.

Using a thermobalance to study thedeactivation kinetics of AC,

AbbasandDaud [40] developed a model to describethe decrease

of catalytic activity with time and reported a deactivation-

reaction order of 0.5 and an Ea of 194 kJ mol1; in addition, the

weight of deposited carbon was fitted with time using the

Voorhies equation, which is commonly used for reactions

involved in hydrocarbon cracking.

There have been studies carried out to investigate the

effect of carbon deposition on catalyst surface area. It was

believed that the deposition results in the blockage of the pore

mouths by growing carbon crystallites which leads to rapid

deactivation and a sharp decrease in the overall surface area[102,117,121]. Moliner et al. [103] showed that the spent cata-

lysts underwent a substantial reduction in specific surface

area compared to the fresh catalysts and found that the

mechanism for catalyst deactivation could be based on

progressive pore blocking by carbon deposition; carbon

deposition in a quantity of less than 25% of the catalyst mass

reduced the surface area from 1300 to 39 m2 /g. Lee et al. [31]

showed that after4 h of deposition most AC had a surface area

of 20–60 m2 /g compared to 860–978 m2 /g for fresh AC catalyst.

Ashok et al. [122] studied methane TCD using a variety of

commercialcoal-derivedcarbons and showed that the surface

areas of the catalysts decreased from 117 to 1478 m2 /g in the

fresh catalysts to 8–27 m2 /g. For CB, it was reported that afterreaction at 900 Cfor1handat1000 C for 2 h,the surface area

of the CB (BP 2000) decreased from 1500 m2 /g to 310 and

100 m2 /g, respectively [119]. Moradov et al. [47] showed that

the surface area vs. time curve closely followed the methane

conversion curve, with an initial drop followed by a shallow

decline. They obtained empirical decay-law equations for the

surface area of AC and CB with time, and the two equations

showed that the surface area of AC diminishes much faster

than that of CB. Krzy _zyński and Koz1owski [125] attributed the

controversial nature of the problem of the relationship

between the surface area of the catalyst and its activity to the

use of different carbon materials obtained from different

precursors and applying different preparation methods, asthen the activity of the samples could depend not only on the

surface area but also on their structure. They used ACs all

obtained from the same precursor (brown coal) and activated

with KOH, with different weight ratios of KOH to precursor(4:1

to 1:2)and at different temperatures (550, 700 and850 C); their

results on the changes in catalyst activity with time are shown

in Fig. 11. Again, no discernible correlation wasfound between

the surface area of the catalysts and their rO, however, they

found that with increasing surface area, and particularly with

increasing pore volume, the resistance of the catalysts to

deactivation increased.

The pore-size distribution is a key factor in defining the

long-term behavior of the catalyst. It has been reported that



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Table 4 – Recent studies using carbonaceous catalysts and major findings.

Researcher Catalysts Operating conditionsT/Flow

Catalyst propertiesW/SA/Porosity/MPS

Muradov et al. [47] ACs (different sources), CBs (different sources), Glassy carbon,

Natural graphite, Polycrystalline graphite, Diamond powder, CNT,

Fullerene soot and fullerenes C60/70

850/0.1 s (residence time)

0.03–0.1/4–2570/—/—

* AC had the highest rO among all the carbon materials tested.

* Structurally ordered carbon (graphite and diamond) with very low surface area

showed negligible catalytic activity.

* Nanostructured carbon (CNT and fullerenes) showed relatively low catalytic activity.

* Fullerene soot was especially active.

* Metal impurities played a negligible role in methane TCD.

* Reaction order for methane was determined to be 0.6 0.1 for AC (lignite)

and 0.5 0.1 for CB (BP2000) catalysts.

* Ranges of Ea for AC and CB were 160–201 and 205–236 kJ/mol, respectively.

Muradov [118] ACs from different origins activated by steam or KOH. 850/1 and 10 s (residence time)

0.03/650–3370/—/—

* High rO was followed by a rapid drop in catalytic activity, and finally, a quasi-steady reaction rate was

reached after 1–1.5 h on stream.

* Range of rO was 1.63–2.04 mmol/(g $min.).

* The origin and method of AC activation had no significant effect on AC catalytic activity.

* Hydrogen initial concentration (HIC) in the effluent gas was 40–46 vol%.

* HIC reached up to 90 vol% when residence time was 10 s (comparable to Fe- and Ni-based catalysts at the identical conditions).

Pinilla et al. [105] CB (BP2000) and AC (CG Norit). 800–950/3000 mL/min

0.03/1300–1337/1.11–3.06/—

* Reaction order for AC and CB was 0.48 and 0.6, respectively.

* Ea for AC and CB were 141 and 238 kJ/mol, respectively.

* The higher the temperature used, the faster the catalyst reached the maximum amount of carbon that could be accumulated.

Suelves et al. [117] Two graphitized CBs (Carbopack B and C), Three CBs (Fluka 03866,

Fluka 05120 and Black Pearls, 2000), CB (HS-50) and AC (CG NORIT).

850/20–100 mL/min.

——/10–1337/——/——

* The samples released tar and water during the first stage of the runs.

* HIC was 10% for Carbopack B and up to 70% for Fluka 03866 and AC (CG).

* CB B and C showed the highest sustainability factor (1), while Fluka 03866 showed the lowest (0.23).

* AC showed the highest rO (0.69) while CB C showed the lowest (0.01 mmol/(g $min)).

Moliner et al. [103] Three different commercial ACs (char pyrolysis), Two

activated chars (steam activation at 600 C and 750 C)

850–950/20 mL/min

2/52–1300/——/


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* All the carbons deactivated quite rapidly and the sustainability factors were 0.1–0.25.

* The range of rO was 0.7–1.5 mmol/(g $min.) at space-time of 0.6 s.* For certain ACs, the reaction order was 0.51 and 0.49 and Eas were 194, 198 and 186 kJ/mol.

Bai et al. [101] One AC from hardwood, two ACs from coal

(steam activated) and an active alumina.

750–900/3000–18,000 mL/(g. h)

1/152–783/0.36–0.73/246–833

* Very small amounts of hydrocarbons other than methane (


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the amount of carbon deposited shows a linear relationship

with the total pore volume of the fresh catalysts [39]. Dunker

et al. [116] showed that carbon deposition eliminated all or93% of the total micropore area during certain runs and

concluded that the reduction in catalyst activity may be

associated with the reduction in microporosity and probably

not associated with the change in mesopore area, which was

actually increased in some runs. Ashok et al. [122] reported

a decrease in pore volume from 0.68 cm3 /g for fresh coal-

derived catalyst to 0.042 cm3 /g, and the micropore volume

decreased from 0.56 cm3 /g to 0.0038 cm3 /g. Bai et al. [101] used

three types of AC for methane TCD and showed that the pore

volume and micropore volume decreased greatly, and the

decrease of micropore volume was especially great (for

example, from 0.26 to 0.000112 cm3 /g). These changes were

attributed to the carbon deposition, especially in the micro-pores of ACs; thus, they concluded that the adsorption and

decomposition of methane occur mainly in the micropores.

A survey of the effect of catalyst particle size indicated that

thesmallerparticlesshowedthehigherrO.However,afteralong

deposition duration, the effect of particle size on methane

conversion was insignificant. Among a range of particle sizes

studied (108–300mm) used for methane TCD in an FBR, Lee et al.

[31] reported that the 108-mm particle size resulted in higher

methane conversion due to its higher surface area and smooth

fluidization. Kim et al. [115] showed that while the decomposi-

tion rate was higher over the smaller particles, the rate of

deactivation was lower, and this indicates that carbon deposi-

tion occurs preferentially at the outer shell of the carbon parti-cles, resulting in pore-mouth blocking while the inner core was

left intact, especially for larger particles. They also showed that

pore-mouth blocking was alleviated as the particle size

decreased and the inner surface could be utilized to a much

greater extent. Among the range of particles sizes studied (137–

1140mm),the137-mm particlesize gavethe highest conversion of

methane.In another study, Abbas andDaud [37] showed thatas

theparticle size was decreased theultimate mass that AC could

accumulate before complete deactivation increased (Fig. 12).

On the effect of temperature, Muradov[13] showed thatwith

a rise in temperature, the mean size of carbon crystallites ten-

ded to decrease, resulting in an increased methane decompo-

sition rate, and this explains the experimental fact that at

higher temperatures (e.g., >850 C) carbon catalysts tend to

deactivate at a slower rate compared to lower temperatures.

With regard to the effect of temperature on catalyst deactiva-tion, Moliner et al. [103] showedthat two different effects could

be identified: (1) a molecular-sieve effect, which would be

associated with pore-mouth blocking. As carbon deposition

progresses, the pore-mouth area decreases and the inner pore

surface becomes unavailable for methane adsorption; and (2)

an activated-diffusion effect, which is related to the rate of

diffusion of the methane molecules inside the smallest pores.

As temperatureis increased, the rate of diffusion increases and

deposition inside the pores is enhanced. In this way, at high

temperatures the decompositioncould takeplace mainly inside

the pores, where the majority of the AC surface is located.

6. Catalyst regeneration

As mentioned previously, the catalyst used for methane TCD

is rapidly deactivated due to intensive carbon deposition




Abbas and Daud [37] AC from palm-shell and commercial AC 800–950/200 mL/min

0.01/1027–1100/0.4501/117

* Study conducted in a thermobalance.

* Reaction order and Ea were 0.5 and 210 kJ/mol, respectively.

* Ultimate mass gain increased with decreased mean particle size.

* Carbon deposition occurred uniformly at the early stage of the process, while the diffusion effect was significant at the end of the process.

Umf , minimum fluidization velocity; T, temperature (C); F, either flow rate (mL/min) or space velocity (h1) unless otherwise stated; W, weight

of catalyst (g); SA, surface area (m2 /g); Porosity, (cm3 /g); MPS, mean particle size (mm unless otherwise stated); —, not mentioned in the original

paper.

Fig. 11 – Catalytic activity of carbon samples in methane

decomposition reactions at 750 8C [125].



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because carbon is one of the two products. Different activating

agents such as CO2, H2O and O2 can be used to regenerate

a deactivated carbonaceous catalyst; CO2 and O2 have also

been used in studying the regeneration of deactivated metal

catalysts used for methane TCD. Using different oxidizing

agents, Muradov [13] reported that the treatment of deacti-

vated AC samples with steam and steam-CO2 mixtures (1:1 by

volume) resulted in a significant increase in the methane

decomposition rate, while air exhibited a relatively lowcarbon-activating efficiency (Fig. 13).

Using a thermobalance, Abbas and Daud [126] studied the

regeneration of an AC catalyst for six methane decomposition

runs at temperatures of 850 and 950 C with five regeneration

cycles using CO2 at temperatures of 900, 950 and 1000 C to

evaluate the stability of the catalyst. They reported that rO and

the ultimate mass gain of the catalyst decreased after each

regeneration step at both decomposition temperatures, but

the decrease was slower under severe regenerating conditions

(Fig. 14). Additionally, a slower decrease in ultimate mass gain

was obtained when the decomposition was carried out at

950 C compared to the one carried out at 850 C, while no

significant differences were observed for the decrease of rOusing reaction temperatures of 950 or 850 C. Pinilla et al. [127]

studied methane TCD using 4 g of AC as a catalyst at 850 C for

three decomposition cycles, and the deactivated AC was

gasified using CO2 at 900 and 925 C; in comparison with the

first reaction cycle, in the second and third cycles it was

observed that rO decreased by 65 and 30%, respectively, while

the ultimate mass gain decreased by 88 and 27%, respectively.

Although metal catalysts have higher activity for acceler-

ating the methane TCD reaction and thus reduce the required

reaction temperature compared with carbon-based catalysts,

unfortunately, the catalyst activity is gradually lost as the

reaction proceeds due to the covering of active sites with the

carbon by-product. The regeneration of deactivated catalyst is

always done by a burning off or gasification process which

leads to CO2 production in amounts nearly comparable to the

quantity of CO2 emitted by the SMR process [100]. Another

serious problem arising from oxidative regeneration of metal

catalysts is related to the unavoidable contamination of

hydrogen with carbon oxides, which would require an addi-

tional purification step [31]. Takenaka et al. [128] studied

methane TCD using 40 mg of Ni/AL2O3, Ni/SiO2 or Ni/TiO2 asthe catalyst at 550 C for five decomposition–regeneration

cycles with CO2 used as the regeneration agent at 650 C. The

results showed that the catalytic activity of Ni/TiO2 remained

high during the repeated reactions, while it gradually

increased for Ni/AL2O3. For the ultimate mass gain, it was

reported that Ni/SiO2 showed the highest value, however, it

decreased significantly after repeated cycles, while for Ni/TiO2it increased from the first to the third cycle and was then

unchanged after the third cycle. As for Ni/AL2O3, the ultimate

mass gain was very low; however, it increased gradually with

the number of reaction cycles. The catalytic performance of

Ni/SiO2, Ni/TiO2, Ni/Al2O3 and Pd–Ni/SiO2 in the repeated

decomposition of methane and oxidation of the CNFs formedwith O2 was studied by Otsuka et al. [129]. They reported that

theoretically, the catalytic decomposition of methane (CH4/

C þ 2H2) followed by the oxidations of the deposited carbon

with CO2 and O2 ((1/2)Cþ (1/2)CO2/CO, (1/2)Cþ (1/2)O2/ (1/2)CO2)

gives an over–all reaction CH4 þ (1/2)O2 / 2H2 þ CO

(DH1073¼20 kJ/mol), requiring no energy input and with zero

CO2 emission. They also showed that Ni/TiO2, Ni/Al2O3 and

Pd–Ni(1:3)/SiO2 (Fig. 15) were found to be promising catalysts

because the catalytic activities completely lost during the

preceding methane decomposition could be repeatedly

recovered by the oxidation of the CNFs formed with O2.

Other researchers have also evaluated the possibility of

catalyst regeneration by the removal of the carbon deposited

Fig. 12 – Mass gain vs. time for different particle sizes at PCH4 [ 0.19 atm and 850 8C [37].

Fig. 13 – Effect of carbon catalyst activation by different activating agents on the methane decomposition rate at

850 8C. Activation temperature: 950 8C [13].



22/31

Fig. 14 – Mass gain versus time for six reaction cycles at 950 8C and PCH4 of 0.63 atm using pure CO2 at 50 mL minL1 a) 900 8C

b) 950 8

C. c) 1000 8

C [126].



23/31

on the external surface of the catalytic particles by attrition.However, the results showed that attrition phenomena are

effective only on the carbon deposited on the external surface

of the catalytic particle, and this carbon represents only

a fraction of the total carbon produced by the decomposition

process [70,86,130]. In separate study, Jang and Cha [88]

investigated the effect of PBR and FBR reactors on methane

TCD using Fe and Fe/Al2O3 catalysis, and found that the

conversion rate of methane was maintained by attrition of the

by-product carbon from the Fe catalyst surface. In modeling

methane TCD in an FBR, Ammendola et al. [131] considered

the attrition phenomena as a novel catalyst regeneration

strategy, in which the carbon-attrition rate, Ec, is able to

balance carbon deposition rate. Attrition of carbon deposited

on the external surface of catalyst particles results in carbonemission as carbon fines transported in the exit gases as well

as in the renewal of a part of external active surface of the

catalyst. According to the scheme shown in Fig. 16A, they

assumed that constant values of the H2-production and

carbon-elutriation rates can be obtained. Fig. 16A reduces to

Fig. 16B when attrition is not present.

7. Characteristics of carbon produced fromthermocatalytic decomposition of methane

As mentioned previously, carbons can be classified into

different types according to their crystallinity or the degree of

Fig. 15 – Decomposition of methane at 823 K and successive oxidation of the CNFs formed with oxygen at 480 8C (furnace

temperature) over Ni (5 wt.%)/SiO2 for repeated cycles. a) Kinetic curves of methane conversion in methane decomposition.

b) Kinetic curves of oxygen conversion in the oxidation of carbon. c) Yields of hydrogen (H2 /Ni) at complete catalyst

deactivation. Flow rate of methane: 60 mL/min (101.3 kPa). Flow rate of Ar and O2 mixture: 75 mL/min (partial pressure of O2,

20 kPa) [129].



24/31

order, i.e., from highly ordered carbons, such as graphite and

diamond, to less ordered (turbostratic and pyrolytic carbons)

and, finally, to disordered (amorphous and microcrystalline)

carbons, such as AC, charcoal and CB. It has been reported

that, depending on the operating conditions of the methane

TCD process, carbon can be produced in several types:

amorphous, turbostratic and CFs. The prices of turbostratic,

pyrolytic and filamentous carbons are 0.3, (1–18) and more

than 1000 $/kg, respectively. This wide difference in the prices

of different carbon modifications shows the importance of the

type(s) of carbon(s) produced in methane TCD in terms of reducing the cost of hydrogen production. Amorphous

carbons are more active in methane TCD than well-ordered

carbons such as graphite, diamond and CNT because the

surface concentration of high-energy sites (where a regular

array of carbon bonds is disrupted, forming free valences,

discontinuities and other energetic abnormalities such as

surface defects and dislocations) increases with the decrease

in carbon-crystallite size and conversely decreases as carbon

becomes more ordered. Accordingly, the catalytic activity of

carbons towards methane decomposition is in the following

order: amorphous > turbostratic > graphite [13,28,47].

The formation of CF on some carbonaceous catalysts,

similar to those formed when methane TCD is performed on

metal-supported catalysts, has been reported (Fig. 17B). Thisphenomenon attributed to the presence of trace amounts of

metal components such as K, Na and Fe in the ash of the

carbonaceous catalysts [31,101,117,122]. The non-catalytic

(homogenous) decomposition of methane when conducted at

temperatures in excess of 1000–1100 C has led to the

production of various forms of amorphous carbon, e.g., CB and

thermal black [75].

For metal catalysts, a generally accepted mechanism of

filamentgrowthconsists of the formation of carbon specieson

the surface of the metals, dissolution in the metal, diffusion

through the metal and precipitation from the metal at a point

on the surface, resulting in the formation of the filament body.

The rate-determining step of this process is the bulk diffusionof the carbon through the metal particle, while excess carbon

migrates along the metal surface to form the skin component

of the filaments [45,91]. Regarding the types of carbon depos-

ited on a catalyst when methane TCD is conducted with metal

catalysts, Ni catalyst have attracted more attention due to

their high catalytic activity and the capability of producing CF

at moderate temperatures (500–700 C). Bai et al. [85] showed

that CF formation was observed at moderate conditions with

a low Ni loading and that temperature had a significant effect

on CF formation and morphology.

When Fe-based catalysts were used, a higher temperature

range was required for efficient operation; also, Fe-based cata-

lysts were able to catalyze the formation of CNT. Reshetenkoet al. [132] showed that catalyst compositions and preparation

methods influencedtheir properties, andthe introduction of Co

or Ni in small amounts (3–10% mass) resulted in significant

increasein carbonyield (two tothreetimesmorethanFe–Al2O3)

and in the formation of multi-wall CNTs. The influence of

nickel-crystal domain size on the behavior of Ni and Ni–Cu

Fig. 16 – A conceptual representation of thermocatalytic

decomposition of methane on a single catalyst particle.

A) Presence of attrition phenomena; B) absence of attrition

phenomena [131].

Fig. 17 – SEM micrographs of the used catalysts after activity tests: (a) AC; (b) CB, T : 850 8C, methane flow: 20 mL/min [117].



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catalysts has been studied by Pinilla et al. [133]. They reported

a ‘‘herringbone’’ CNF structure for Ni–Cu/MgO catalysts, with

diameters that correlated with the nickel-particle size, while

carbon deposited with the Ni/Al2O3 catalyst was usually in the

form of ribbon nanofibers. Fig. 18A shows a low-magnification

TEM micrograph with some metal particles, mostly withrounded or diamond-like shapes, and emerging CNFs on the

tips. Fig. 18B shows the detail of a diamond-like nickel-particle,

whereas in Fig. 18C, the inner arrangement of the graphene

layers forms a small angle of w22 with respectto thefiber axis.

Thus, the carbon deposited is in the form of herringbone

nanofibers with a hollow core. Fig. 18D is a high-magnification

TEM micrograph showing graphene structures extending from

one wall of the fiber to the other.

Shah et al. [95] conducted methane TCD using nanoscale,

binaryFe–M(M]Pd,Mo, or Ni)catalysts supportedon alumina.

High-resolution SEM and TEM characterization indicated that

almost all the carbon produced in the temperature range of

700–800 C was in the form of potentially useful multi-walled

CNTs. At higher t

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