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Steam reforming of volatile fatty acids (VFAs) over supportedPt/Al2O3 catalysts

Chang Moon Jeong a, Gwon Woo Park a, Jin-dal-rae Choi a, Jong Won Kang a,Sung Min Kim a, Won-Ho Lee b, Seong Ihl Woo a, Ho Nam Chang a,*aDepartment of Chemical and Biomolecular Engineering, KAIST, 335 Gwahang-no, Yuseong-gu, Daejon 305-701, Republic of Koreab LG Chem. Ltd/Research Park, 104-1, Moonji-dong, Yuseong-gu, Daejon 305-741, Republic of Korea

a r t i c l e i n f o

Article history:

Received 12 January 2011

Received in revised form

17 March 2011

Accepted 20 March 2011

Available online 6 May 2011

Keywords:

Steam reforming

Volatile fatty acid

Biomass

Wastewater

Fermentation

* Corresponding author.E-mail address: hnchang@kaist.edu (H.N

0360-3199/$ e see front matter Copyright ªdoi:10.1016/j.ijhydene.2011.03.126

a b s t r a c t

Volatile fatty acids (VFAs), easily produced using acid fermentation of biomass, were used

to generate hydrogen via steam reforming. Three short-chain carboxylic acids (C2eC4) e

acetic, propionic and butyric acids e were used as model compounds in addition to VFAs

produced in a typical anaerobic batch reactor. Catalytic steam reforming of VFAs using

alumina-supported platinum catalysts was studied in a fixed-bed quartz reactor at various

temperatures between 300 and 600 �C. The influence of reaction conditions such as

temperature, oxygen to carbon ratio (O/C) and gas hourly space velocity (GHSV) was

investigated. VFAs were successfully converted to COx and hydrogen. A hydrogen yield of

up to 70% was achieved, based on typical stoichiometry at 600 �C and a GHSV of 25,000 h�1.

Temperature-programmed oxidation (TPO), X-ray diffraction (XRD) and pore size distri-

bution (PSD) were used to characterize coke deposition. Graphitic carbon on catalysts was

not identified by XRD, which implies that amorphous coke had formed in the small pores.

The catalysts could be reactivated by oxidation and reduction. A detrimental effect on

hydrogen yield was observed by adding a small amount of O2 to the VFA feed, due to the

high concentration of oxygen in the feed composition. Steam reforming of real VFAs (S/

C ¼ 9) in the acid fermentation of food waste was performed with different GHSVs at

a reaction temperature of 600 �C. Conversion of VFAs decreased significantly with

increasing GHSV, but the hydrogen selectivity was still above 60%. The conversion path-

ways of the VFAs to COx and hydrogen are most likely complex, particularly due to the

variety of the chemical compounds present in the real VFAs. The steam reforming of VFAs

was investigated over various noble metal (Ruthenium, Palladium, Rodium, Nickel) cata-

lysts supported on alumina, the specific activity based on the active surface area decreased

in the order of Ru > PdwRh > Pt > Ni.

Copyright ª 2011, Hydrogen Energy Publications, LLC. Published by Elsevier Ltd. All rights

reserved.

1. Introduction environmental pollution. Fuel cells using hydrogen have

Hydrogen is becoming a more attractive alternative energy

carrier due to the depletion of fossil fuels and the

. Chang).2011, Hydrogen Energy P

higher fuel efficiency than conventional gasoline and diesel

engines. They produce onlywater as a by-product without any

pollutant emission. However, hydrogen is currently produced

ublications, LLC. Published by Elsevier Ltd. All rights reserved.

i n t e rn a t i o n a l j o u r n a l o f h y d r o g e n en e r g y 3 6 ( 2 0 1 1 ) 7 5 0 5e7 5 1 57506

from nonrenewable sources, such as natural gas and petro-

leum [1e3], which are accompanied by high CO2 emissions.

On the other hand, biomass can be used to produce hydrogen

without a net change of CO2 in the atmosphere and will play

an important role as a renewable and sustainable hydrogen

source in the future. Volatile fatty acids (VFAs) derived

anaerobically from organic biomass can be very useful in

serving as major platforms for bio-fuels and chemicals in post

fossil fuel era [4,5].

Various chemical and biological technologies are available

to generate hydrogen from biomass. For the past few decades,

renewable biomass has been considered as a potential feed-

stock for the gasification process to produce syngas (amixture

of hydrogen and carbon monoxide). This process can be per-

formed with or without a catalyst [6,7]. However, typical

gasification processes have the disadvantage of low thermal

efficiency (30e50%) due to large amounts of water in the

biomass. Thus, gasification requires a very large reactor [8].

Another hydrogen production technology is the fast pyrolysis

of biomass with catalytic steam reforming of the resulting

pyrolytic oil (bio-oil), a complex mixture of various aliphatic/

aromatic oxygenates [9]. Attempts to produce hydrogen from

bio-oil using commercial steam reforming catalysts have been

severely interrupted by rapid catalyst deactivation caused by

coke/oligomer deposition on the catalysts [10]. A more real-

istic approach has been proposed using individual compo-

nents present in bio-oil such as acetic acid, one of the model

oxygenated components of bio-oil (up to 32 wt%). Platinum

has been shown to be essential for hydrogen formation, and

that the support is needed to extend catalyst life [11]. Bio-

ethanol, a renewable material easily obtained from biomass,

has also been proposed as an intermediate in the hydrogen

production [12e14]. Recently, aqueous phase reforming (APR)

of sugars [15] or bio-ethanol [16] in generating hydrogen has

been proposed. APR is under development to process

oxygenated hydrocarbons or carbohydrates to produce

hydrogen. APR reactors are often operated at pressures up to

2e30 MPa and temperatures ranging from 220 to 270 �C. Most

research to-date has been focused on the use of supported

GroupVIIImetals as the activemetal, and sugars (e.g., glucose)

and polyols (e.g., methanol, ethylene glycol, glycerol and

sorbitol) as the substrates. However, until now, economically

feasible biomass utilization processes for hydrogen produc-

tion have not been well-developed.

In this paper, we propose an alternative process for the

production of hydrogen from biomass that involves acid

fermentation of biomass to generate volatile fatty acids (VFAs)

followed by reforming. Anaerobic digestion or acid fermen-

tation is a process in which microorganisms break down

biodegradable material in the absence of oxygen [17]. The

process is widely used to treat wastewater sludge and organic

wastes because it provides a significant reduction in the

volume andmass of the input waste. After hydrolysis of waste

by the extra-cellular enzymes of microorganisms, the organic

wastes, including carbohydrates, lipids and proteins, are

converted into various short-chain fatty acids, such as acetic

acid (HAc), propionic acid (HPr) and butyric acid (HBu). These

VFAs are converted anaerobically to carbon dioxide and

methane gas, i.e., biogas, a renewable energy source suitable

for energy production helping to replace fossil fuels. However,

methane formation is a very slow reaction due to the slow

growth of methane-forming microorganisms. In addition,

methane produced in such a manner contains many toxins

such as hydrogen sulfide. Furthermore, the greenhouse effect

caused by methane is much more severe than that caused by

carbon dioxide. For this reason, we propose to recover the

energy from VFAs before they are converted to methane,

which can be achieved by catalytic steam reforming of VFAs.

The steam reforming of VFA can be simplified to the steam

reforming of an oxygenated organic compound (CnHmOk) by

the following reaction:

CnHmOk þ ðn� kÞH2O/nCOþ ðnþm=2� kÞH2 (1)

The above reaction is followed by the wateregas shift

reaction:

nCOþ nH2O4nCO2 þ nH2 (2)

Therefore, the overall process can be represented as follows:

CnHmOk þ ð2n� kÞH2O/nCO2 þ ð2nþm=2� kÞH2 (3)

The process performance was measured by the hydrogen

yield calculated as the percentage of the stoichiometric

potential, assuming complete conversion of carbon to CO2,

according to Reaction (3). Thus, the potential yield of hydrogen

gas from an oxygenated feedstock is (2 þm/2n� k/n) moles of

H2 permole of carbon in the feed. In reality, the hydrogen yield

will always be lower than the stoichiometric potential because

the wateregas shift (WGS) reaction is reversible, resulting in

the presence of some carbon monoxide and methane in the

product gas. In addition, thermal cracking that occurs in

parallel to reforming produces carbonaceous deposits, which

are especially significant for thermally unstable compounds.

In this study, we consider the steam reforming of VFAs

obtained from acid fermentation of food waste on alumina-

supported platinum catalysts. The steam reforming of a few

different short-chain carboxylic acids typically found in VFAs

(e.g., acetic, propionic and butyric acid) was studied. The

influence of reaction conditions such as temperature, steam

to carbon ratio (S/C), oxygen to carbon ratio (O/C) and gas

hourly space velocity (GHSV) was investigated. Temperature-

programmed oxidation (TPO), X-ray diffraction (XRD) and pore

size distribution (PSD) were used to characterize the deacti-

vation of the catalysts. In addition, the steam reforming of

VFAs was investigated over various noble metal catalysts

supported by alumina.

2. Materials and methods

2.1. Catalyst preparation

Commercial g-alumina (Sigma Co. Ltd.) in powder form

was used as a support after 1 h calcinations at 600 �C. Thestudied Pt/Al2O3 catalysts were prepared by incipient

wetness impregnation with aqueous solutions of metallic

precursor salts (Pt(NH3)2(NO2)2) (Alfar Aesar) on a Al2O3

support. Metal loading was 5 wt %. The impregnated

catalyst was then dried at 120 �C and all catalysts were

activated by reduction under 5 mol % H2 at 600 �C for 2 h

before each run.

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2.2. Catalyst characterization

The specific surface areas of the support and the supported

catalysts were measured by the Brunauer-Emmett-Teller

(BET) technique using the Tristar 3000 surface analyzer

(Micromeritics, USA), employing N2 physisorption at the

temperature of liquid N2. Prior to each measurement, the

sample was dried at 120 �C under a helium flow passing

through the sample cell.

The metal dispersion of fresh catalyst was determined by

hydrogen chemisorption using the ASAP 2000 volumetric

adsorption analyzer (Micromeritics, USA). Prior to each

measurement, the catalyst sample (ca. 200mg) was pretreated

by: (a) dynamic vacuum at 50 �C for 1 h, (b) reduction in

flowing H2 at 600 �C for 1 h, (c) evacuation for 1 h at 400 �C, and(d) cooling to the chemisorption temperature of 35 �C.Hydrogen uptake at the monolayer of the platinum particles

was obtained by extrapolation of the linear portion of the

adsorption isotherm to zero pressure. The exposed surface

area was calculated by assuming a H: Pt stoichiometry of 1:1

[18]. The actual metal loadings of the reduced catalysts were

measured by inductively coupled plasma mass spectroscopy

using HP 4500 (Hewlett Packard, USA).

XRD patterns of the supports and the supported catalysts

were obtained using a Multi-Purpose Attachment X-ray

Diffractometer (D/MAX 2500, Japan) equipped with Cu Ka

radiation (lK ¼ 0.1542 nm). Crystalline phases were identified

by comparison with PDF standards (powder diffraction files)

from the International Center for Diffraction Data (ICDD). The

average crystallite size (D) was determined using the Scherrer

equation as follows:

D ¼ Klk=bcos q (4)

where lK is the wavelength of the X-rays used, q is Bragg angle

of diffraction peak, b is the full width at half maximum of

diffraction peak and K is the Scherrer Constant (here, K ¼ 0.9).

2.3. Experimental procedure and apparatus

Fig. 1-(a) shows a schematic representation of the experi-

mental apparatus used for the reforming reaction with VFAs.

The quartz reactor (20 mm (O.D.), 18 mm (I.D.), 400 mm

(height)) was located inside a tubular electric furnace and the

temperature inside the catalyst bed was measured with

a thermocouple. Helium was used as an inert carrier. The

input flow of helium to the pre-heater was controlled by

a mass flow controller. The pre-heater temperature was

maintained at 170 �C for all experiments. VFA aqueous solu-

tion was fed by a High Performance Liquid Chromatography

(HPLC) pump and vaporized at 170 �C. Oxygen was used in

some experiments to enhance the reforming by supplying

combustion heat and also suppressing the coke deposition. In

the experiments with oxygen feed, co-feeding both VFA/

Steam and oxygen with one inlet resulted in the oxidation of

reactants in the upper part of the quartz reactor, between

furnace and pre-heater (marked in Fig. 1-b). We made a new

type reactor with independent oxygen feeing line to eliminate

this phenomenon (Fig 1-c) and measured actual temperature

in the catalysts bed. The reactor effluent was passed through

a cold trap to remove condensable product. The composition

of the gases after the trap was analyzed by an on-line gas

chromatograph (HP 6890) equipped with TCD and FID to

determine the concentrations of H2, CO2, O2, CH4 and CO. The

hydrogen yieldwas defined as the ratio of the concentration of

H2 in the actual outlet gas to the theoretical amount of

hydrogen that could be obtained when complete reforming to

generate CO2 and H2 occurred (Eq. (3)). The selectivity was

defined (for each compound) as the ratio of the moles of each

H2, CO2, CH4 and CO in the actual outlet gas to the total moles

of outlet gases, with the exception of the carrier gas. The

concentration of each VFA (HAc, HPr, HBu, etc.) wasmeasured

by HPLC (Hitachi L-3300 RI detector, Japan) equipped with an

ion exchange column (Aminex HPX-87H, Hercules, USA) using

5 mM H2SO4 as the mobile phase. All samples were filtered

through a 0.22-micron (pore diameter) membrane filter prior

to measurement.

3. Results and discussion

3.1. Catalyst characterization

The physical and chemical characteristics of fresh supports

and catalysts are shown in Table 1. The BET area of the fresh

g-Al2O3 supports before calcinations was 327.74 m2/g, but

during the 1 h calcinations at 600 �C, the BET surface area

decreased due to the sintering effect. The specific surface

areas of the g-Al2O3 support and the Pt/Al2O3 catalysts were

166.01 and 142.98 m2/g, respectively, suggesting that the

loading of Pt decreased the surface area of the alumina

support only slightly. Metal dispersion of Pt/Al2O3 catalysts by

hydrogen chemisorptions is also reported in Table 1.

In general, catalysts performance based on their configu-

ration (i.e. pellet form) is limited by mass-heat-transport

phenomena especially at a high flow rate, to increase the total

H2 production, volume and weight of reactors must be

dramatically increased together with the catalyst cost [19].

However, we can assume that the large surface area in

a powder form and small-scale quartz reactors leads to good

heat and mass transfer property.

The XRD patterns of fresh supports and catalysts are

shown in Fig. 2. As expected, the most visible features in

platinum,which are equivalent to themain peaks (PDF No. 00-

004-0802) of platinum nano-particles, occurred at 39.75�(111),46.20�(200), 67.42�(220), 81.22�(311) and 85.60�(222).

3.2. Effects of reaction temperature

The steam reforming of each acid was performed in

a temperature range of 300e650 �C and the effects of

temperature on H2 yield and acid conversion are shown in

Fig. 3. To avoid catalyst deactivation caused by carbon depo-

sition, the reactions were performed for a short duration (1 h)

and fresh catalyst was used for every temperature point.

GHSV was controlled at 25,000 h�1 and the S/C of 9 was used

for all experiments in this section. The homogeneous (non-

catalyst) reaction over the temperature was studied, but the

VFA conversion was less than 1% and any H2 and CO2 was not

measured in product line. The reaction in the presence of the

Fig. 1 e (a) Schematic diagram of the experimental apparatus, (b) general quartz reactor, (c) new type quartz reactor with

oxygen feeding line and thermocouple.

Table 1 e Characteristics of supports and catalysts.

Characteristics Al2O3a Al2O3

b 5 wt% Pt/Al2O3c

Total surface area

(B.E.T.) (m2/g)

327.74 166.01 142.98

Metal dispersion (%) e e 16.35

Metal surface area

(m2/g metal)

e e 40.38

Metal loading (%)d e e 5.13

a Fresh supports.

b Support after 1 h calcinations at 600 �C.c after a 1 h H2 reduction at 600 �C.d ICP/MS measurements.

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alumina supports was also examined (data not shown). The

major products of this reaction were CO2 and CH4, with only

small amounts of CO and H2. The conversion of VFAs to CH4

and CO2wasmost likely caused by the thermal decomposition

of VFAs. The conversion of VFAs was less than 5% and

significant carbon deposits were found on the surface of the

supports surface after the reaction, which might be caused by

following reaction [20]:

CnH2nO2/CO2 þ nH2 þ ðn� 1ÞCads (5)

However, the gaseous product composition changed

significantly in the presence of the alumina-supported plat-

inum catalyst with a large increase in hydrogen yield and

Fig. 2 e XRD patterns of Al2O3 supports (a) before

calcinations, (b) after calcinations at 600 �C and (c) Pt/Al2O3

catalysts.

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reactant conversion. This phenomenon implies that the

presence of platinum is essential for steam reforming. At

300 �C, the conversion of HAc, HPr and HBu was just 29.4%,

11.2% and 2.7%, respectively, but it reached above 98% after

a temperature increase to 600 �C. At the same time, CO

selectivity decreased from 23.3% to nearly 0%; in contrast, the

selectivity of H2 formation increased dramatically with these

Fig. 3 e Steam reforming of each acid andwater mixture on the P

acetic acid; ,, propionic acid; A, butyric acid), (c)w(e) product g

>, CH4; 6,CO). Experimental conditions: Pt/Al2O3 catalyst [ 20

conditions. The large amount of CO generation at 300 �C may

have been caused by the fact that the steam reforming

(Eq.(1)) and the wateregas shift (Eq.(2)) reactions could not

occurred substantially due to reactant acid and steam

adsorbed on the surface of the catalyst blocking sufficient

catalyst activation because of the low reaction temperature.

Above 400 �C, the steam reforming and WGS reactions pro-

ceeded normally. As a result, both reactant conversions and

selectivity in producing H2 increased significantly. When the

temperature increased consecutively to 600 �C, the catalyst

exhibited the best performance and reactant acids were

converted completely. Also, the CO2/H2 ratio reached near

theoretical values, which indicates complete steam reform-

ing, while the selectivity in producing the by-product CH4

was about 0.2% and only a negligible amount of CO was

detected. However, when the temperature continuously

increased to higher ranges, such as above 650 �C, the selec-

tivity in producing CH4 and CO became remarkable. There

were trace and negligible amounts of acetone, C2 hydro-

carbon (i.e. C2H4, C2H6) detected from 300 to 450 �C, hence, we

did not express it in the figures.

The CO2/H2 ratio was calculated to estimate the carbona-

ceous conversion of VFAs at 600 �C. The carbonaceous

conversion of VFAs was 92.4, 91.3 and 86.9% from HAc to HBu,

respectively. This result shows that the longer chain acids

favor conversion via thermal decomposition, which can

produce CH4 or acetone and so on. Interestingly, a trace

t/Al2O3 catalysts. (a) Acid conversion, (b) hydrogen yield (C,

as selectivity from acetic acid to butyric acid (C, H2; -, CO2;

0 mg, GHSV [ 25,000 hL1, S/C [ 9.

Fig. 4 e Steam reforming of VFAmixture on the Pt/Al2O3

catalyst. Experimental conditions: Pt/Al2O3

catalyst[200mg,GHSV[25,000hL1, temperature[600 �C,S/C[ 9. (B, VFA conversion;,, H2 yield;C, H2 selectivity;-,

CO2 selectivity;:, CO selectivity;>, CH4 selectivity).

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amount of higher carbon chain carboxylic acids, such as

valeric (C5) and caproic (C6) acid were observed in butyric acid

reforming.

3.3. Effect of steam to carbon ratio (S/C)

The influences of S/C on the reactions were studied at 600 �C,the optimum temperature used in the previous section. The S/

C was varied from 3 to 9, controlling constant GHSV of

25,000 h�1. Results on the H2 yield and reactant conversion are

shown in Table 2. The H2 yield increases with increasing S/C.

S/C also had a significant effect on the conversion of reactants,

which dropped substantially from 100% to 65.4% with

a decrease in S/C decrease from 9 to 3 for the case of HAc.With

regard to selectivity in producing the products, the S/C also

had remarkable effects: significant amounts of CH4 and CO

were observed at a low S/C of 3. The generation of CH4 and CO

can greatly affect the selectivity in producing H2. The

production of 1 mol CH4 will result in a 4 mol loss of H2.

Similarly, the production of 1mol COwill result in a 1mol loss

of H2. As a result, when the S/C decreased from 9 to 3 in the

case of VFAs reforming, the selectivity in producing H2

decreased remarkably from 70.2% to 59.9%.

3.4. Steam reforming of acid mixture

The steam reforming reactionwas performed for 52 h at 600 �Cto study the stability of the Pt/Al2O3 catalyst. The reaction was

conducted under the steady-state condition with a GHSV of

25,000 h�1 and a feed mixture consisting of VFAs (S/C ¼ 9)

(molar ratio of HAc:HPr:HBu ¼ 6:1:3), which is the typical

molar ratio of the acidmixture in anaerobic digestion [21]. The

results are given in Fig. 4. A gradual deactivation of the cata-

lyst was observed, corresponding to a decrease of 40% in VFAs

conversion after 52 h. However, the selectivity of carbon-

containing compounds did not change during the experiment.

In the initial stage of 30 h, the conversion of VFAs was still

around 80%, owing to the slow deactivation of the catalyst.

After 30 h, the conversion of VFAs decreased significantly and

a small amount of acetone ranging from 0.1 to 0.2 wt% was

Table 2 e Effect of steam to carbon ratio.

Feedstocks

S/C Conversion(%)

H2 yield(%)

Recoverya

(%)

HAc 3 65.4 57.2 101.7

6 87.8 65.8 98.2

9 100.0 68.5 99.5

HPr 3 53.6 55.4 99.4

6 79.5 64.7 98.7

9 99.7 71.1 101.3

HBu 3 49.1 53.3 99.3

6 74.7 64.1 100.1

9 99.4 73.4 101.0

VFAs 3 87.2 59.9 99.1

6 93.0 66.6 98.4

9 99.1 70.2 100.4

a Recovery was calculated from gas and liquid yield in the product

line.

also detected with increasing time. Thus, it is likely that

catalyst deactivation by carbon deposition or catalyst aging

would affect the performance of steam reforming, which will

be discussed in the following section.

3.5. Characterization of carbon deposition andreactivation of catalysts

Carbon deposition is one of the major problems in catalytic

reforming reactions because it leads to rapid deactivation of

the catalysts due to poisoning of the active sites and/or pore

blockage [22e24]. To study the nature of the carbon deposits

formed, TPO experiments were performed after running the

steam reforming reaction under the steady-state condition

described in the previous section. After cooling the catalyst to

room temperature, the catalyst was exposed to a mixture of

2mol % of O2 in heliumwith a flow rate of 200mL/min, and the

temperature was increased at a linear rate of 5 �C/min up to

700 �C. During the TPO experiment, carbon oxides (CO, CO2)

were detected with the CO2 analyzer and gas chromatography

(GC). In the TPO and GC results, either no CO was observed or

negligible amounts were observed; however, two peaks of CO2

were observed at around 428 and 534 �C. These peaks indicate

that two distinct carbon species exist on the catalyst surface

(Fig. 5). The amount of coke deposited was also estimated by

integration of the CO2 curves. The percentage of carbon

deposited at the lower temperature peak at 428 �C was 4.3% (g

coke/g carbon in the feed), and the corresponding percentage

of the higher temperature peak at 534 �C was 2.7% (g coke/g

carbon in the feed). Generally, the lower temperature peak is

thought to be due to coke deposited on the metal surface,

while the higher temperature peakdthemost significant peak

and the least reactivedis attributed to coke deposited on the

support [25]. Similar results have been reported, but it has

been suggested that the first peak around 440 �C is attributed

to the coke deposited on the metalesupport interface [26].

Fig. 5 e Temperature-programmed oxidation (TPO) profiles

of the Pt/A2O3 catalysts after steam reforming of VFAs. Fig. 6 e XRD patterns of (a) the fresh catalysts, (b) after the

first reforming reaction, (c) after oxidation and reduction at

600 �C and (d) the second reforming reaction.

Fig. 7 e Pore size distribution of the Pt/Al2O3 catalysts

using different conditions. (B, Fresh catalysts; C, After

first reforming, 6, after oxidation and reduction; :, after

second reforming).

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Martin et al. suggested that coke at the metalesupport inter-

face has intermediate reactivity and that, this reactivity

presumably accounts for the continuity of the TPO profiles at

temperatures between those required to oxidize coke on the

metal and those required oxidize coke on the support [27]. In

hydrocarbon reforming, three different types of coke have

been reported to form on the supported metal catalysts:

polymeric, filamentous, and graphitic cokes [28]. Polymeric

coke originates from the gas-phase decomposition of hydro-

carbons, whereas formation of the filamentous and graphitic

cokes requires the participation of metallic sites on the cata-

lyst. Coke can also be characterized based on its reactivity.

Soft coke can be gasified by treatment in hydrogen using

relatively mild conditions and does not accumulate on the

active sites. Soft coke is thought to consist of secondary

reaction products or intermediates. Hard coke is typically

graphitic, unreactive with hydrogen and also blocks active

sites. Although these gross and overall coke properties are

reasonably simple to measure, determining the exact chem-

ical and physical nature of the coke is difficult [25,29].

XRD analysis and N2-adsorption measurements were also

conducted to characterize structural change and carbon

deposition on the catalysts during reaction and after regen-

eration via subsequent oxidation (10 mL/min, 1 mol % O2 for

half hour at 500 �C) and reduction (10 mL/min, 1 mol % H2 for

half hour at 600 �C) treatments. The XRD results are shown in

Fig. 6. A significant difference was not observed around

2q ¼ 26� when comparing the spectra before and after the

reaction, indicative of graphitic carbon (PDF No. 12-0212 and

26-1077). This observation confirmed that no significant

amount of carbon was deposited on the support and metal

surface, which was consistent with TPO experiment. Alter-

natively, only amorphous carbon could be deposited, which

can be fully oxidized for reactivation [30]. As summarized in

Fig. 7 and Table 3, BET area, pore volume and average pore size

were significantly decreased after the first and second reac-

tion, indicating the formation of cokes which might cause

blockage of the active site. Average crystallite size of platinum

determined from XRD line width did not significantly change

during the reaction and regeneration processes, indicating

that the present reaction and regeneration conditions do not

change the metal dispersion. The catalyst activity could be

recovered by greater than 90% based on hydrogen yield using

aforementioned regeneration condition. However, relative

catalyst activity based on hydrogen yield andmetal dispersion

determined by H2 chemisorptions were significantly reduced

to 5% and 0.1%, respectively, by oxidation over 650 �C. Thisphenomenon is probably due to the platinum crystallite size

growth by sintering. It has been reported that calcinations

treatment in the air results in an increase of platinum particle

size, which is larger than particles treated under either inert or

H2 atmospheres [31]. Fig. 8 shows the morphology of catalysts

(a) before and (b) after the reaction. They can be hardly iden-

tified at the nano-scale because the catalysts used in this

study were micro-structured. However, scanning electron

Table 3 e Porosity and crystallite size of Pt/Al2O3 catalysts using different conditions.

Conditions BET surface area(m2/g)

Average Pt crystallite size(nm)a

Pore volume(cm3/g)b

Pore area(m2/g)b

Mean pore diameter(nm)c

Fresh after reduction 142.98 12.8 0.396 205.58 10.6

After first reforming 84.15 13.0 0.219 112.72 7.8

After regeneration 115.60 14.6 0.405 156.54 10.5

After second reforming 96.25 13.9 0.224 125.33 7.2

a Average Pt crystallite size was calculated from the XRD results shown in Fig. 6.

b Pore volume and area (smaller than 300 nm) was measured by the Barrett-Joyner-Halenda (BJH) method.

c The mean pore diameter was calculated assuming cylindrical pore geometry.

i n t e rn a t i o n a l j o u r n a l o f h y d r o g e n en e r g y 3 6 ( 2 0 1 1 ) 7 5 0 5e7 5 1 57512

microscopy (SEM) images show that the catalysts were

aggregated by coke formation after reaction.

3.6. Effect of oxygen to carbon ratio (O/C)

Oxygen used in the auto-thermal steam reforming or partial

oxidation reaction for thermodynamic advantage also

suppresses carbon deposition on the catalyst surface. The

steam reforming of VFAs is highly endothermic (DH¼þ357 kJ/

mol of VFAs at 25 �C). There are two ways of supplying heat to

the reacting system: (i) externally, by burning fuels and

transporting heat into the reaction mixture, or (ii) internally,

by co-feeding oxygen or air and burning a portion of the VFAs.

In the latter case, 1.2 mol of O/mol of C in feed VFAs are

required to achieve thermal neutrality, based on the following

stoichiometry:

C2H4O2 þ 0:17C3H6O2 þ 0:5C4H8O2 þ 1:04O2

þ 3:58H2O/4:5CO2 þ 8:08H2 (6)

To reduce carbon deposition and supply heat via the

internal combustion of VFAs, various amounts of oxygenwere

added (so-called auto-thermal conditions). Table 4 shows the

effect of the addition of various concentrations of oxygen on

the conversion to COx and the hydrogen yields at 600 �C.

The hydrogen yield gradually decreased with increasing

concentration of O2 in the feed. The hydrogen yield dropped

from 63% to 29% at an O/C ¼ 2.06. However, severe catalyst

deactivation as well as decrease of the hydrogen yield was

observedwhen oxygen feedwas excessive (O/C> 2.06) and the

hydrogen yield was decreased less than 1% after 1 h, probably

owing to the sintering effect by remaining air and increased

Fig. 8 e SEM images of (a) the fresh catalysts and (

temperature in the catalysts bed, as mentioned previous

section.

TPO experiments were also conducted with different O/C

ratios to estimate carbon deposition. Between O/C ¼ 0.74 and

1.53, the carbon deposition percent, defined as mass of coke

per mass of carbon in the feed, decreased gradually. The

carbon deposition percent was drastically reduced at an O/C

of 2.06. Furthermore, the visual observation of the reactor

and the catalyst color clearly revealed that the carbon

deposition was significantly reduced when oxygen was

added. In the steam reforming of ethanol, the addition of

oxygen already proved beneficial for the conversion to COx

and no significant reduction of the hydrogen yield was

reported [32]. Otherwise, in steam reforming of acetic acid

with fluidized bed, an excess of oxygen (8%) can lead to

a lower reforming activity decreasing the H2 and CO2 yields

and increasing the CO and CH4 yield. On the other hand, 4%

oxygen resulted in almost no penalty in hydrogen yield [33].

In this study, a detrimental effect of oxygen was observed in

the hydrogen yield due to the VFAs containing more oxygen

than ethanol.

3.7. Influence of GHSV

Space velocity (or GHSV) is one of the key operating parameters

that affects production rates. Generally, an increase in space

velocitywill result in lowerconversion.SteamreformingofVFAs

(S/C ¼ 9) produced from acid fermentation of food waste was

performed with different GHSVs at 600 �C (Fig. 9). Typical VFAs

produced frombiomassare3e4wt%and largeamountsofwater

must be removed to save the input energy in steam reforming.

There are many kinds of recovery methods to concentrate

b) used catalysts after the reforming reaction.

Table 4 e Influence of oxygen to carbon ratio (O/C).

Temperaturea O/C ratio Conversion (%) H2 yieldb H2 yield (%) Cokec (%) Selectivity (%)

H2 CO CH4 CO2

595 0.74 82.7 3.84 62.9 0.90 62.1 1.5 0.0 36.4

595 1.00 78.5 3.25 53.2 0.86 54.4 1.5 0.0 43.1

596 1.27 85.5 3.11 50.9 0.82 49.8 1.4 0.9 47.9

597 1.53 92.3 2.98 48.8 0.80 51.4 1.4 0.8 46.3

612 2.06 96.3 1.77 29.0 0.33 33.5 3.7 0.8 62.5

a Temperature in the catalysts bed (Fig. 1-C).

b H2 mol/VFA mol.

c g coke/g carbon in the feed.

i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n en e r g y 3 6 ( 2 0 1 1 ) 7 5 0 5e7 5 1 5 7513

organic acids. We have used membrane-based liquid/liquid

extraction with Alamine 336/octanol as solvent [34].

Conversion of VFAs significantly decreasedwith increasing

GHSV. However, hydrogen selectivity did not change signifi-

cantly. After long-term operation of 50 h, the catalyst activity

for hydrogen production decreased slightly, probably due to

carbon deposition or catalyst aging. The details of the

conversion pathway of the VFAs to COx and hydrogen are

most likely complex, in particular due to the variety of the

chemical compounds present in the real VFAs. For example,

VFAsmay contain sulfur compounds, such as sulfate ions and

sulfur-containing amino acids (cysteine and methionine)

depending on the feed source, which could cause catalyst

poisoning. These sulfur compounds must be removed by

separation techniques before the reaction or research for

sulfur-tolerant catalysts is required [35].

3.8. Relative activity for various noble metal catalysts

The steam reforming of VFAs was investigated over various

noble metal catalysts supported on alumina with a VFA

Fig. 9 e Influence of GHSV with real VFAs. Experimental

conditions: Pt/Al2O3 catalyst [ 200 mg,

GHSV [ 25,000 hL1, temperature [ 600 �C, S/C [ 9. (B,

VFAs conversion; ,, H2 yield; C, H2 selectivity; -, CO2

selectivity; :, CO selectivity; >,CH4 selectivity).

concentration of 13.7 wt% (S/C ¼ 9) and a GHSV of 25,000 h�1

at 600 �C. Palladium (Pd), rhodium (Rh), ruthenium (Ru) sup-

ported on alumina were supplied by Johnson Matthey (UK).

Nickel catalysts were prepared by the incipient wetness

method using nickel nitrate hexahydrate as the metal

precursor. All metal loading contents were 5 wt%. The results

are given in Table 5. As shown in fifth column, hydrogen

production rate per gram catalyst did not show great differ-

ences due to relatively short reaction time (1 h) and low

molecular weight of nickel. However, it was found that the

specific activity based on the active surface area (as calcu-

lated by the rate of H2 production and metal dispersion)

decreases in the following order for alumina-supported

metals:

Ru > PdwRh > Pt > Ni

Studies have identified copper-based catalysts as effective

materials to produce hydrogen by the steam reforming of

methanol at temperature near 300 �C [36]. However, copper-

based catalysts are not effective for steam reforming of

heavier hydrocarbons or oxygenated compounds, since they

show low activity for cleavage of CeC bonds [37]. Therefore, it

is more likely that an effective catalyst for oxygenated

compound reforming would be based on Group VIII metals,

which generally show higher activities for breaking CeC

bonds. The catalytic activities of different metals for CeC

bond breaking during ethane hydrogenolysis have been

studied by Sinfelt [37]. It can be seen that Pt shows reason-

able CeC bond breaking activity, although not as high as

metals such as Ru, Ni and Rh, which show highest activities

for CeC bond breaking in vapor-phase ethane hydro-

genolysis. An effective catalyst for reforming of VFAs must

not only be active for cleavage of the CeC bond, but it must

also be active for the wateregas shift reaction to remove CO

from the metal surface. In this respect, Grenoble et al. have

reported the relative wateregas shift activities for different

metals supported on alumina [38]. It can be seen that Cu

exhibits the highest wateregas shift rates among all the

metals (but shows no activity for CeC bond breaking) and Pt,

Ru and Ni also show appreciable wateregas shift activity.

Finally, to obtain a high selectivity for hydrogen production,

the catalyst must not facilitate undesired side reactions, such

as methanation of CO and FischereTropsch synthesis. It can

be seen that Ru, Ni and Rh exhibit the highest rates of

methanation, whereas Pt and Pd show lower catalytic activ-

ities for the methanation reaction [16]. Thus, on comparing

Table 5 e Characteristics and activity for steam reforming of various catalysts.

Catalyst BET surface area(m2/g)

Metal dispersion(%)

Metal surface area(m2/gmetal)

Rate (H2)(mol/h$gcat)

TOF(H2)a (s�1)

Pt/Al2O3 142.98 16.35 40.38 5.79 1.92

Pd/Al2O3 144.38 5.73 25.55 5.43 2.80

Rh/Al2O3 158.27 5.41 24.25 5.09 2.64

Ru/Al2O3 149.38 4.56 21.76 6.14 3.77

Ni/Al2O3 152.97 11.31 70.11 5.77 0.86

a TOF(H2) ¼ moles of hydrogen produced per moles of surface metal per second.

i n t e rn a t i o n a l j o u r n a l o f h y d r o g e n en e r g y 3 6 ( 2 0 1 1 ) 7 5 0 5e7 5 1 57514

the metals based on all the three reactions, it can be inferred

that Pt and Pd should show suitable catalytic activity and

selectivity for reforming of VFAs, which requires reasonably

high activity for CeC bond breaking and wateregas shift

reactions, and low activity for methanation. The rate of

a particular reaction can be compared for the different

metals, however, for specific metal, the absolute rates of the

three different reactions cannot be compared relative to each

other. In case of VFAs steam reforming according to on our

study, Ru exhibits the highest TOF among all the metals,

suggesting that activity for CeC bond cleavage and water-

egas shift reaction is more important feature rather than

methanation and thermal decomposition.

4. Conclusions

VFAs as representative model compounds of biomass were

reformed effectively in a fixed-bed reactor in the presence of

alumina-supported platinum catalysts. A hydrogen yield above

60% was maintained over 10 h using a 5 wt% Pt/Al2O3 catalyst

and operating at a steam to carbon ratio of 13.7. Catalysts were

deactivated by amorphous coke formed in their small pores.

However, the catalytic activity was recovered by oxidation

followed by reduction. The use of auto-thermal conditions (i.e.,

addition ofmolecular oxygen to the feed) reduced the extent of

carbon formation but also led to a significant loss in hydrogen

yield. Nevertheless, the continuous reforming of VFAs is

concluded to be feasible using these catalysts. Additionally,

catalytic activity studies based on TOF were conducted for VFA

steam reforming. Finally, effective catalyst for the production

of H2 fromVFA steam reforming should be active for CeC bond

cleavage and the wateregas shift reaction.

However, a metal loading of 5 wt% is relatively high;

therefore, further research in seeking inexpensive and active

catalysts is needed. By combining a proper reactor designwith

catalytic materials that minimize the catalytically produced

coke, VFAs have potential to become an important hydrogen

production method with significant environmental benefits.

Acknowledgement

This study was supported by Grant No. 2007-07001-0094-

0 from the Korea Institute of Environmental Science and

Technology and Grant No. M10309020000-03B5002-00000 from

the Korea Ministry of Education, Science and Technology. The

authors wish to thank Professor M.K. Choi of KAIST for his

thorough and helpful comments on the chemical aspects of

the reforming catalysts.

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