Continuous supercritical water gasification of isooctane: A promising reactor design

international journal of hydrogen energy 35 (2010) 1957–1970

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Continuous supercritical water gasification of isooctane: A promising reactor design Ratna F. Susanti a,b, Bambang Veriansyah a, Jae-Duck Kim a,b,**, Jaehoon Kim a,b,*, Youn-Woo Lee c a

Supercritical Fluid Research Laboratory, Clean Energy Center, Energy Division, Korea Institute of Science and Technology (KIST), 39-1 Hawolgok-dong, Seoungbuk-gu, Seoul 136-791, Republic of Korea b Department of Green Process and System Engineering, University of Science and Technology (UST), 113 Gwahangno, Yuseong-gu, Daejeon 305-333, Republic of Korea c School of Chemical and Biological Engineering, Seoul National University, Gwanangro 599, Gwanak-gu, Seoul 151-744, Republic of Korea

article info

abstract

Article history:

A new design of supercritical water gasification system was developed to achieve high

Received 1 October 2009

hydrogen gas yield and good gas–liquid flow stability. The apparatus consisted of a reaction

Received in revised form

zone, an insulation zone and a cooling zone that were directly connected to the reaction

21 December 2009

zone. The reactor was set up at an inclination of 75 from vertical position, and feed and

Accepted 23 December 2009

water were introduced at the bottom of the reactor. The performances of this new system

Available online 13 January 2010

were investigated with gasification of isooctane at various experimental conditions – reaction temperatures of 601–676  C, residence times of 6–33 s, isooctane concentrations of

Keywords:

5–33 wt%, and oxidant (hydrogen peroxide) concentrations up to 4507 mmol/L without

Hydrogen

using catalysts. A significant increase in hydrogen gas yield, almost four times higher than

Isooctane

that from the previous up-down gasifier configuration (B. Veriansyah, J. Kim, J.D. Kim, Y.W.

Supercritical water gasification

Lee, Hydrogen Production by Gasification of Isooctane using Supercritical Water, Int. J.

Gasifier configuration

Green Energy. 5 (2008) 322–333) was observed with the present gasifier configuration. High hydrogen gas yield (6.13 mol/mol isooctane) was obtained at high reaction temperature of 637  C, a low feed concentration of 9.9 wt% and a long residence time of 18 s in the presence of 2701.1 mmol/L hydrogen peroxide. At this condition, the produced gases mainly consisted of hydrogen (59.5 mol%), methane (14.8 mol%) and carbon dioxide (22.0 mol%), and a small amount of carbon monoxide (1.6 mol%) and C2–C3 species (2.1 mol%). Reaction mechanisms of supercritical water gasification of isooctane were also presented. Crown Copyright ª 2009 Published by Elsevier Ltd on behalf of Professor T. Nejat Veziroglu. All rights reserved.

1.

Introduction

Due to global warming, the use of hydrogen as a clean alternative transportation fuel for proton exchange membrane fuel

cells has received much attention recently. One of major hurdles for the fuel cell application in transportation is to produce hydrogen in a sufficient quantity with a reasonable price [1]. This problem has led to considerable efforts to

* Corresponding author. Supercritical Fluid Research Laboratory, Clean Energy Center, Energy Division, Korea Institute of Science and Technology (KIST), 39-1 Hawolgok-dong, Seoungbuk-gu, Seoul 136-791, Republic of Korea. Tel.: þ82 2 958 5874; fax: þ82 2 958 5205. ** Corresponding author. Supercritical Fluid Research Laboratory, Clean Energy Center, Energy Division, Korea Institute of Science and Technology (KIST), 39-1 Hawolgok-dong, Seoungbuk-gu, Seoul 136-791, Republic of Korea. Tel.: þ82 2 958 5873; fax: þ82 2 958 5205. E-mail addresses: [email protected] (J.-D. Kim), [email protected] (J. Kim). 0360-3199/$ – see front matter Crown Copyright ª 2009 Published by Elsevier Ltd on behalf of Professor T. Nejat Veziroglu. All rights reserved. doi:10.1016/j.ijhydene.2009.12.157

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international journal of hydrogen energy 35 (2010) 1957–1970

develop a compact reformer to produce hydrogen from fossil fuel that can be used onboard automobiles [1,2]. The major hydrogen production technologies include steam reforming, partial oxidation, autothermal reforming and water electrolysis. Although the steam reforming is the most common and well-established method, it is often considered to be unsuitable to develop a compact steam reformer system. Heterogeneous catalysts that are typically used in the steam reforming are often overheated and sintered during the reforming reaction, resulting from limited heat transfer through the reactor wall and low thermal conductivities within the catalyst bed [2,3]. In addition, the catalysts are often deactivated by the presence of sulfurous/carbonyl compounds in fuel and coke/ tar formation as a typical byproduct of the gasification [1]. Autothermal reforming (ATR) also requires catalysts, thus unsuitable for the compact reformer system [1,4–6]. Partial oxidation can be conducted in the absence of catalyst at flame temperatures of 1300–1500  C and in the presence of catalyst at lower temperature of >800  C. However, the exothermic nature of the reforming reaction makes it difficult to control the temperature due to the coke formation. Thus, safety and thermal management become major barriers for the partial oxidation to be applied in practical, compact and portable gasification applications [6]. In addition, the densities of the fluid phase are low due to the low working pressure of autothermal reforming (atmospheric) [7–9], steam reforming (0.1–3.5 MPa) [10] and partial oxidation (atmospheric) [8]. As a result, large reactor volume is often required to perform gasification at the same fuel flow rate and the same residence time compared to a system with higher working pressure. It is well-known that high purity hydrogen can be produced by the electrolysis of water. However, low energy efficiency, low yield, and limited production rate are major drawbacks. Therefore, there are still considerable efforts underway to develop simpler, more compact, and more efficient techniques that can produce hydrogen for ultimate onboard fuel cell applications. Supercritical water gasification (SCWG) can be a potential alternative to produce hydrogen from fossil fuels using a compact reforming system. The highly dense, supercritical fluid phase allows to develop a compact reformer system [10]. In addition, much higher density of supercritical water than that of steam can result in higher thermal conductivity and specific heat. This is beneficial to carry out endothermic gasification reactions [4,5]. Due to the single-phase reaction and low mass transfer resistance associated with beneficial physicochemical properties of supercritical water, the reforming reaction can be conducted in a very short residence time (typically less than 1 min) [4]. Thus large amounts of raw materials can be gasified using a compact supercritical water reactor. The dielectric constant of supercritical water is much lower and the number of hydrogen bonding is much smaller than those of the water at ambient condition or steam [11]. As a result, typical hydrocarbons can be completely soluble in supercritical water. Produced gases are also miscible in supercritical water condition [11]. Therefore, single-phase gasification reaction condition can be performed in supercritical water. This is in sharp contrast with the multi-phase reaction conditions associated with the conventional gasification techniques such as steam reforming and partial oxidation [12,13].

In recent years, the advantages of supercritical water have been utilized in gasification of methanol [5,14,15], ethanol [16], model biomass compounds such as cellulose [17–20], glucose [19–25], lignin [17,26–28], phenols [26,29], glycine [29,30], glycerol [30], pyrocatechol [31], and real biomass such as sewage slurry [32], wood sawdust [33–35], rice straw [33–35], stalk (from wheat, corn and sorghum)[35], peanut shell[35] and clover grass [36]. Most of these works have been conducted either in the presence of heterogeneous catalysts such as nickel and ruthenium on high-surface area supports [16,21,37], in-situ generated copper nanoparticles [15], zirconia oxide [19,27] and carbon [32], or in the presence of homogenous catalysts such as potassium hydroxide [31,38] and sodium hydroxide [27]. Although high gasification efficiency and high hydrogen gas yield can be achieved in the presence of the catalysts, high catalyst costs and the catalysts deactivation associated with sulfurous compounds in fossil fuel and coke and/or tar formation during the reforming reaction are major drawbacks. Therefore, it is highly desired to develop a non-catalytic supercritical water gasification system for continuous onboard reformer applications. In our previous work, supercritical water gasification of isooctane (2,2,4-trymethylpentane), a model compound of gasoline, was carried out without using catalysts [4]. It was shown that hydrogen can be produced at relatively low temperatures of 600–700  C. However, the hydrogen gas yield was very low (maximum yield of 1.68 mol hydrogen/mol isooctane). The previous gasification work was conducted in a vertical tubular reactor in which isooctane and supercritical water were introduced at the top of the reactor and produced gases and effluent liquids flowed out at the bottom of the reactor. With this up-down reactor configuration, virtual plugging in the vertical reactor at a high gas flow rate was observed, possibly caused by a collision between the gas flow and the liquid flow in a condenser region. This resulted in instability in both of the gas and the liquid flow. In this study, a new supercritical water gasification design was developed to solve the problems associated with the previous gasification configuration. This includes tilting of the gasification reactor to 75 from vertical position, positioning the inlet of the reactor at the bottom and the outlet of the reactor at the top (down-up configuration), and addition of an insulation and a cooling zone. The sections that follow describe the new down-up gasification apparatus and process, the effect of reaction temperature, residence time, feed concentration, and oxidant concentration on gas yields, gas compositions, and total organic compounds of liquid effluents. The gas yields and the gas composition from the present down-up gasification reactor are compared to those from the previous up-down gasification reactor.

2.

Experimental method

2.1.

Materials

Isooctane (2,2,4-trymethylpentane, purity of >99%) and aqueous hydrogen peroxide solution (35 w/v%) were purchased from Junsei Chemical Company (Tokyo, Japan) and used as received. Desired hydrogen peroxide (H2O2)

international journal of hydrogen energy 35 (2010) 1957–1970

concentrations were prepared by diluting the purchased aqueous hydrogen peroxide solution with distilled and deionized (DDI) water. The DDI water was prepared using a Milli-Q Ultrapure water purification system equipped with a 0.22 mm filter (MA, USA).

2.2.

Apparatus and procedure

The supercritical water gasification was carried out using a custom-built, continuous-flow tubular reactor system. A schematic diagram of the down-up gasification apparatus is shown in Fig. 1. The apparatus consists of a tubular reactor (R), a water pre-heater (WP), a feed pre-heater (FP), several heat furnaces including pre-heater furnaces (PF), a mixing part furnace (MF), a reactor furnace (RF), an insulation part furnace (IF), two high-pressure pumps (P-01 and P-02), cooling units (CT, C-01 and C-02), a metal filter (F), a back pressure regulator (BPR), a gas–liquid separator (S), a liquid container (L), feed tanks (T-01, T-02, T-03 and T-04), isolation and safety valves (V1, V2, V3 and V4), pressure gauges (P), and thermocouples (T). The reactor dimension and material has been described in the previous work [4]. The reactor was inclined to 75 from a vertical position. The insulation zone (IZ) with a length of

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28 cm and a volume of 22 mL was added to keep the temperature of the reaction zone be isothermal. The cooling tube (CT) is a length of 115 cm and a volume of 90 mL using outside air as a cooling medium. The cooling tube (CT) was directly attached to the insulation zone (IZ). The feed pre-heater (FP) was a stainless steel (SS) 316 tubing with a length of 25 cm and an outside diameter (O.D.) of 0.635 cm. The water or oxidant pre-heater (WP) was a SS 316 coil with a length of 12 m and an O.D. of 0.3175 cm. The mixing tee (M) was made of Hastelloy C276 with an O.D. of 0.635 cm. The temperatures of the preheater (WP and FP), the mixing tee (M), the reaction zone (RZ) and the insulation zone (IZ) were controlled using the heat furnaces, manufactured by Dae-Poong Industries (Seoul, Korea). All hot regions of the apparatus were insulated in boxes of ceramic board and the temperatures were monitored using type-K thermocouples (Omega Engineering, Inc., CT, USA) and a multichannel recorder (model mR 180, Yokogawa, Tokyo, Japan). The system temperature was controlled by a temperature controller (model DTA4848, Asia Delta Electronics, Inc., Taoyuan, Taiwan). The C-01 condenser was a shell and tube type with a length of 10 cm and with a tube I.D. of 10 mm. The C-02 condenser was a coil type with a length of 60 cm and a tube O.D. of 0.635 cm. The C-02

Fig. 1 – Supercritical water gasification apparatus scheme. R, a tubular reactor; M, a mixing part; PF, pre-heater furnaces; MF, a mixing furnace; RF, a reactor furnace; WP, a water pre-heater; FP, a feed pre-heater; IF, an insulation furnace, RZ, a reaction zone; IZ, an insulation zone; CZ, a cooling zone; CT, a cooling tube; C, a condenser; F, a metal filter; BPR, a back pressure regulator; P-01, a high-pressure isooctane pump; P-02, a high-pressure water pump; S, a gas–liquid separator; P, pressure gauges; T, thermocouples; V1, needle valves; V2, relief valves; V3, three-way valves; V4, safety valves; T-01, an isooctane feed tank; T-02 and T-03, DDI water feed tanks; T-04, an oxidant feed tank; L, a liquid container; GC, a gas chromatograph; WG, a wet gas meter.

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international journal of hydrogen energy 35 (2010) 1957–1970

condenser was used to keep the inlet temperature at the BPR lower than 74  C because the BPR (model 26-1721-24, Tescom, MN, USA) can withstand temperature up to 74  C. The cooling medium for the C-01 and C-02 condensers was tap water at ambient temperature. The filter (F and wet gas meter (WG) specification were described in the previously work [4]. The pump P-01 was a high-pressure metering pump (model 396/ 2396, Thermo Separation Products, FL, USA) with flow rates in the range of 60–160 mL/hr. The pump P-02 was a high-pressure digital metering pump (model HKS-3000, Hanyang Accuracy, Seoul, Korea) with a maximum flow rate of 600 mL/ h. Prior to each gasification experiment, DDI water was introduced to the reactor system using the high-pressure pump and the system was pressurized to an experimentally desired pressure of 25 MPa by adjusting the BPR. The temperature of the system was then increased to w20 lower than the experimentally desired temperatures using the furnaces. The reactor was stabilized for at least 1 h, and then isooctane was introduced into the reactor system. Higher specific heat of isooctane than that of water may result in the temperature rising [39]. The experimentally desired temperature was adjusted using the temperature controller. The steady-state gasification was checked by constant temperature readings at beginning, middle, end of the reactor and the mixing point. These four temperatures were averaged and reported as the reaction temperature. Real-time volumetric flow rates of the gaseous products were measured using a wet gas meter (WG). Liquid effluent flow rates were measured using a high accuracy balance (model GT4100, Ohaus Corporation, NJ, USA). The gaseous products were collected using a sampling bag in every 30 min for 2 h and analyzed using two gas chromatographs (GC). The liquid effluents were collected and analyzed using a total organic carbon (TOC) analyzer. After each gasification experiment was finished, 2 M of hydrogen peroxide (H2O2) aqueous solution flowed continuously into the reactor for at least 1 h and subsequently DDI water flowed to remove organic compounds that may be present in the reactor.

2.3.

purity helium (purity of >99.9999%) as the carrier gas. The column was a capillary column (HP Rt-Msieve 5A Plot column) with a length of 30 m and an I.D. of 0.53 mm. This GC was used to quantify hydrogen (H2) and carbon monoxide (CO). The temperature of the column was maintained at 100  C for 35 min and the temperature of the detector was maintained at 220  C. The carrier gas flow rate was set at 4 ml/min. In each measurement, approximately 50 mL of gaseous samples was injected. Both of the gas chromatographs were calibrated using pure gases obtained from Shinyang Sanso, Ind. Co. (Seoul, Korea). The gas product samples were analyzed in triplicates. In this paper, total gas yield is defined as volume (liter) of total gaseous products at room temperature (25  C) and pressure (0.1 MPa) per weight (gram) of feed, estimated by the realtime volumetric flow rate measurements and the feed concentration. Carbon gasification efficiency (CE) is defined as total moles of carbon in the produced gases per total moles of carbon in feed. The gas yield of each gaseous component in the produced gases is defined as the moles of each produced gas per moles of isooctane fed to the reactor. The moles of each produced gases were estimated by the real-time volumetric flow rate measurements and the gas compositions analyzed by GC. The gas composition measuring errors were typically less than 5% for dry gas composition, 5% for effluent gas yield composition and 4% for carbon gasification efficiency (CE). The error in the total gas yield was less than 6%. The liquid products were analyzed using a total organic carbon (TOC) analyzer manufactured by Shimadzu Corporation (model TOC-VCSH 200 V, Kyoto, Japan). The TOC analyzer was calibrated with total carbon (TC) and inorganic carbon (IC) standard solutions of known solution concentrations. Before each TOC measurement, the whole liquid product was wellmixed and filtered using a 0.45 mm syringe filter. One milliliter filtrate sample was then dilute in 99 mL of DDI water. During the filtration, it was observed that most of the oily products that were present in the liquid effluent were retained in the filter. The samples were analyzed in triplicates. The TOC measurement errors were less than 2%.

Analytical methods

Compositions of the gaseous products were analyzed using two gas chromatographs (GC). The first GC was a Hewlett– Packard model 5890 Series II GC with a thermal conductivity detector (TCD) and with high purity helium (purity of >99.999%) as a carrier gas. The column was a fused silica capillary column (HP Plot Q column) with a length of 30 m and an I.D. of 32 mm. This GC was used to quantify carbon dioxide (CO2), methane (CH4), ethane (C2H6), ethylene (C2H4), propane (C3H8), propylene (C3H6), butane (C4H10) and butylene (C4H8). The temperature of the column was first maintained 60  C for 8 min, followed by ramping up to 240  C with a rate of 10  C/ min and holding for 4 min at 240  C. The temperature of the detector was maintained at 250  C. The carrier gas flow rate was set at 2 ml/min. In each measurement, approximately 100 mL of gaseous samples was injected. The second GC was Young Lin model ACME 6100 GC with a pulsed discharge helium ionization detector (PDHID, Vici Valco Instruments Co. Inc., TX, USA.) and with ultra high

3.

Results and discussion

3.1.

Effect of temperature

Fig. 2 shows the dependence of the total gas yield, the carbon gasification efficiency (CE), the total organic carbon (TOC) in the liquid effluent, dry gas composition, and the individual gas yields on the reaction temperature at a fixed pressure of 25 MPa, at fixed residence time of 6.9 s and a fixed isooctane concentration of 11.2 wt%. When the reaction temperature increased from 601 to 676  C, the total gas yield increased by two folds from 0.9 to 1.8 L/g isooctane (Fig. 2a) and the TOC in the liquid product decreased significantly from 3069 to 780 mg/L. The high TOC value and the low total gas yield at the low reaction temperature of 601  C indicate that the conversion of isooctane was very low. The low conversion can be explained further by the appearance of the liquid effluent collected. The liquid effluent formed two phases. The bottom layer was a creamy white color and the upper layer was light

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Fig. 2 – Effect of the reaction temperature on (a) total gas yield, total organic carbon (TOC) and carbon gasification efficiency (CE), (b) dry gas composition, and (c) individual gas yield composition. The gasification was conducted at 25 MPa, a residence time of 6.9 s, and an isooctane concentration of 11.2 wt%.

brown aqueous liquid. At a higher temperature of 653  C, the color of liquid effluent became yellowish with thinner layer of dark brown viscous oil at the surface. At 676  C the liquid color was yellow without any phase separated layer at the surface. The upper layer of liquid at 601  C was more dilute than that of at 653  C, presumably because larger amount of unreacted isooctane or higher carbon number species were present in the liquid effluent. Both of the gas and liquid results indicated that high reaction temperature was favored for higher isooctane conversion. The carbon gasification efficiency (CE) was low in the range of 43–44.3% and did not change much with the temperature. Short residence time (6.9 s), thus insufficient gasification period, may result in the low CE values. The effect

of residence time will be discussed in detail later. It is not clear what was causing the constant CE with an increase in temperature. One possible reason is that formation of solid products (coke)/tar increased at higher reaction temperature due to hydrocarbon cracking mechanism (Eqs. (6) and (7) in Appendix). In fact, we observed that larger amount of coke and tar was retained in the filter when the gasification was performed at higher temperature. In addition, the carbons in the gases products and the liquid effluents that can be quantified by the GC and the TOC measurements decreased from 7.23 to 4.54 mol/mol carbons in isooctane as the reaction temperature increased from 601 to 676  C. The carbon balance agrees well with the tar/coke formation results.

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international journal of hydrogen energy 35 (2010) 1957–1970

The dry gas compositions (in mol percent) are shown in Fig. 2b and the dependence of the individual gas yields on the reaction temperature is shown in Fig. 2c. As the reaction temperature increased, the amount of hydrogen in the produced gases increased while the amount of C2–C4 species decreased. The major components of the produced gases were hydrogen (46.6%), methane (25.7%) and carbon dioxide (15.8%) while carbon monoxide, ethylene, ethane, propylene, propane, butylene and butane were produced in a minor quantity (1.5% of CO and 10.3% of C2–C4) at 676  C. The hydrogen gas yield increased from 0.93 to 2.31 mol/mol isooctane and the carbon dioxide gas yield increased from 0.42 to 0.78 mol/mol isooctane with an increase in temperature from 601 to 676  C (Fig. 2c). Typically, the gasification of hydrocarbons in supercritical water involves several complex and competition reactions including steam reforming, watergas shift, methanation, cracking, and hydrogenation. Detailed reactions on the supercritical water gasification of isooctane are listed in Appendix. Due to the endothermic nature of the combined steam reforming and water-gas shift reaction (Eq. (3) in Appendix), higher reaction temperature will lead to an increase of hydrogen and carbon dioxide yields. It is noted that much lower hydrogen gas yields ranging 0.29–0.75 mol/ mol isooctane with temperatures of 597–694  C were observed when the up-down reactor configuration was used [4]. The improved hydrogen gas yield with the current reactor system will be discussed in detail later. As shown in Fig. 2c, methane gas yield increased slightly from 0.55 to 1.27 mol/mol isooctane as the reaction temperature increased from 601 to 676  C. The increase in methane gas yield at higher reaction temperature was contradictory to the exothermic nature of methanation reaction (Eqs. (4) and (5) in Appendix). Methane gas yield should decrease with an increase in the reaction temperature because the equilibrium should shift in the backward direction. Non-equilibrium methanation reaction associated with very short residence time (6.9 s) may be responsible for the contradictory between the experimental methane gas yield and the theoretical methanation reaction. Similar results can be found in the gasification of glucose in the microchannel reactor as reported by Goodwin et al. [22]. It was reported that methane gas yield increased at higher reaction temperature, even though equilibrium thermodynamic calculation predicted the decrease in methane at higher temperature. As shown in Fig. 2c, carbon monoxide concentration in the produced gases was extremely low. Carbon monoxide gas yield was in the range of 0.01–0.08 mol/mol isooctane, and these values correspond to 27–132 ppm in the gaseous products. Allowable carbon monoxide concentration for proton exchange membrane fuel cell application is in the range of 10–100 ppm, depending on cell designs and operating conditions [40]. Typical carbon monoxide gas yields of the conventional hydrogen production methods are higher than that of the supercritical water gasification. For example, the carbon monoxide gas yield of autothermal reforming was 1–5 mol/mol isooctane [2], that of steam reforming was 0.2– 0.6 mol/mol ethanol [41], and that of partial oxidation of isooctane was about 20 volume % (based on GC result) [42]. The low carbon monoxide gas yield associated with the supercritical water gasification may result from almost

complete water gas shift reaction, leading to formation of hydrogen and carbon dioxide with the consumption of carbon monoxide. The high water excess in the feed concentration (88.8 wt% water) may be favorable for the water gas shift reaction. In general, water gas shift reaction is carried out under catalytic conditions, meanwhile it is possible to be carried out non-catalytically under supercritical water conditions since large excess of water possibly drive the reaction to produce hydrogen. A kinetic rate of water gas shift reaction in supercritical water is 105.6exp(1.16  105/RT) [CO] mol/L/s in the temperatures 380–593  C and pressures 10– 59.6 MPa [43].

3.2.

Effect of residence time

Fig. 3 shows the effects of the residence time on the total gas yield, CE, TOC, product gas compositions, and the individual gas yields at a fixed pressure of 25 MPa, a fixed temperature of 632  C and a fixed isooctane concentration of 15.25 wt%. As the residence time increased from 6.0 to 33.3 s, the total gas yield increased from 1.1 to 2.9 L/g isooctane. The carbon efficiency increased from 26.6 to 52.5% as an increase in the residence time from 6.0 to 18.4 s and reached an asymptotic value with further increase in the residence time to 33.3 s. The asymptotic behavior of CE may be due to the solid formation (coke)/tar as a byproduct of the reforming reaction or may be due to the absence of catalyst, leading that CE almost approaches a maximum value. The TOC analysis results agreed well the gas yield results. TOC decreased from 2789.8 to 198.5 mg/L with an increase in residence time from 6.0 to 33.3 s. At a residence time of 6.0 s, the liquid effluent color was yellow with thinner layer of dark brown viscous oil at the surface. As the residence time increased to 10.1 s, the liquid color was whitish, and at 33.3 s, the liquid became clear as pure water. Thus, higher residence time leads to more complete isooctane conversion. The dry gas compositions are shown in Fig. 3b and the dependence of the individual gas yields on the residence time is shown in Fig. 3c. At the longest residence time of 33.3 s, the product gas contained 59.5% hydrogen, 17.2% methane, 18.5% carbon dioxide and minor gas products including 1.0% carbon monoxide and 3.8% of ethane and propane (Fig. 3b). A decrease in C2–C4 content in the produced gases can be seen as the residence time increased. Carbon monoxide gas yield was very low (48–250 ppm), indicating that the high water content are favorable for the water gas shift reaction (w1.2 times more water than a stoichiometry value needed, based on Eq. (3) in Appendix). As shown in Fig. 3c, the alkene gas species including ethylene, propylene and butylene disappeared at longer residence time of 33.3 s while the alkane gas species such as ethane and propane persisted. The absence of the alkene species at longer residence time may suggest that the unsaturated hydrocarbons are weaker and more crackable compared to the saturated hydrocarbon analogs. Propane gas yield decreased from 0.26 to 0.11 mol/ mol isooctane while ethane gas yield increased from 0.19 to 0.25 mol/mol isooctane when the residence time increased from 18.4 to 33.3 s. This indicated that longer residence time promoted further cracking of C3–C4 species to shorter chain hydrocarbon species.

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Fig. 3 – Effect of the residence time on (a) total gas yield, total organic carbon (TOC) and carbon gasification efficiency (CE), (b) dry gas composition, and (c) individual gas yield composition. The gasification was conducted at 25 MPa, 632 8C and isooctane concentration of 15.25 wt%.

As shown by Fig. 3c, hydrogen gas yield increased by five folds from 1.14 to 5.52 mol/mol isooctane while carbon dioxide gas yield increased by six folds from 0.29 to 1.71 mol/mol isooctane and methane gas yield increased by three folds from 0.5 to 1.59 mol/mol isooctane as the residence time increased from 6.0 to 33.3 s. The increase in hydrogen gas yield at longer residence time indicates that extended period of reaction time is needed to complete reaction at the low reaction temperature of 632  C. Our results are in a good agreement with Guo et al. [12] in the gasification of wood sawdust at 650  C and 25 MPa, and with Byrd et al. [21] in gasification of 1 wt% glucose at 700  C and 24.8 MPa. Guo et al. [12] reported that hydrogen gas

yield increased sharply from w13.5 to w19 mol/kg wood sawdust as the residence time increased from 9 to 46 s. Byrd et al. [21] observed that an increase in residence time from 1.6 a to 6.2 s resulted in a significant increase in hydrogen gas yield from w2.5 to w6.7 mol/mol glucose and a decrease in TOC in liquid product from w2400 to w250 ppm in the absence of catalyst. The tendency of hydrogen gas yield that is still rising at the longer regime of residence time may suggest that there is a possibility to achieve higher hydrogen gas yield at longer residence time than 33.3 s. A longer residence time effect on gas yield is currently being studied using a larger volume of reactor.

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3.3.

international journal of hydrogen energy 35 (2010) 1957–1970

Effect of feed concentration

Fig. 4 shows the effect of the feed concentration on total gas yield, CE, TOC, product gas compositions, and the individual gas yields at a fixed pressure of 25 MPa, a fixed temperature of 630  C, and a fixed residence time of 18 s. As the feed concentration increased from 5.7 to 32.7 wt%, the total gas yield decreased from 2.13 to 1.27 L/g isooctane, CE decreased from 56.5 to 43.6% and TOC increased significantly from 311.3 to 1600.5 ppm (Fig. 4a). The shortage of water as a reactant at higher feed/water ratio is responsible for the lower gas yield at higher feed concentration. The compositions of produced gases are shown in Fig. 4b. As the feed concentration increased from 5.7 to 32.7 wt%, the amount of hydrogen in the produced gases decreased from

56.5 to 40.1% while the amount of C2–C4 species increased significantly from 4.5 to 12.4%. The highest composition of hydrogen in the produced gases (56.5%) was achieved at a low isooctane concentration of 5.7 wt%. At this condition, carbon dioxide was 24.6%, methane was 14.0%, carbon monoxide was 0.5%, and C2–C4 species were 4.5%. As shown in Fig. 4c, there is a sharp decrease in the hydrogen gas yield by 3 folds from 4.90 to 1.72 mol/mol isooctane with an increase in the isooctane concentration from 5.7 to 32.7 wt%. In addition, carbon dioxide gas yield decreased from 2.13 to 0.79 mol/mol isooctane while methane and carbon monoxide yield increased slightly with the increment in the isooctane concentration. Boukis et al. [14] reported that the water gas shift reaction under stoichiometric condition in methanol reforming at 600  C was incomplete. Based on stoichiometric condition of

Fig. 4 – Effect of the isooctane concentration on (a) total gas yield, total organic carbon (TOC) and carbon gasification efficiency (CE), (b) dry gas composition, and (c) individual gas yield composition. The gasification was conducted at 25 MPa, 630 8C and a residence time of 18 s.

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combined isooctane steam reforming and water gas shift reaction (Eq. (3) in Appendix), the isooctane concentrations of 5.7–20.3 wt% correspond to a water excess condition (552–55%) while the isooctane concentration of 32.7 wt% is a water shortage condition (18%). Thus higher feed concentration will lead incomplete water gas shift reaction. Based on equilibrium point of view, lower amount of water in feed will shift the equilibrium of isooctane steam reforming (Eq. (1) in Appendix) and water gas shift reaction (Eq. (2)) to the backward direction, leading to a decrease in the hydrogen and carbon dioxide gas yields and an increase in the carbon monoxide gas yield. The equilibrium of carbon monoxide methanation reaction (Eq. (4)) and carbon dioxide methanation reaction (Eq. (5)) will then shift to the forward direction,

1965

leading to an increase in the methane gas yield. Thus, as shown in Fig. 4c, the formation of carbon monoxide and methane may be preferred at higher isooctane concentration. Similar trends of hydrogen, carbon dioxide, carbon monoxide and methane yields with an increase in feed concentration were observed by other research groups [5,16,21].

3.4.

Effect of oxidant concentration

Partial oxidation in supercritical water gasification was carried out to enhance total gas yield and hydrogen gas yield. Hydrogen peroxide, that is known to decompose to molecular oxygen and water in the supercritical water condition, was used as the oxidant (Eq. (11) in Appendix) [44]. Fig. 5 shows the

Fig. 5 – Effect of the oxidant concentration on (a) total gas yield, total organic carbon (TOC) and carbon gasification efficiency (CE), (b) dry gas composition, and (c) individual gas yield composition. The gasification was conducted at 25 MPa, 637 8C, a residence time of 18 s and an isooctane concentration of 9.9 wt%.

1966

international journal of hydrogen energy 35 (2010) 1957–1970

effects of the hydrogen peroxide concentration on total gas yield, CE, TOC, and the individual gas yields at a fixed pressure of 25 MPa, a fixed temperature of 637  C, a fixed residence time of 18 s and a fixed isooctane concentration of 9.9 wt%. As the hydrogen peroxide concentration increased from 0 to 2701.1 mmol/L, the total gas yield increased from 2.08 to 2.94 L/ g isooctane and TOC decreased from 412.5 to 158.8 ppm. Further increase in the hydrogen peroxide concentration to 4507.3 mmol/L leaded to the decrease in the total gas yield to 2.31 L/g isooctane and to the increase in TOC to 300.2 ppm. It is not clear what is causing the increase in TOC with the increment in hydrogen peroxide concentration. Tar or coke that was formed in the reactor during the gasification might be oxidized and dissolved in water, leading to the increase in TOC. Similar TOC increases with low concentration range of oxidant were observed by others. In the supercritical water oxidation of municipal excess sludge and alcohol distillery wastewater, as the amount of oxidant (H2O2) increased from 0 to 20% of stoichiometric needed, TOC increased from w800 to 900 mg/L at 823 K and w1550–1700 mg/L at 773 K [45]. In the supercritical water oxidation of model municipal solid waste, TOC value increased from w6000 to 6500 mg/L as the amount of oxidant (H2O2) increased from 0 to 20% at 773 K [46]. Carbon gasification efficiency (CE) increased from 50.5% to 75.0% as the oxidant concentration increased, suggesting that the presence of hydrogen peroxide increased the conversion of isooctane to the gases products. The dry gas compositions with varying hydrogen peroxide concentration are shown in Fig. 5b and the dependence of the individual gas yield on hydrogen peroxide concentration is shown in Fig. 5c. As in the case of the total gas yield, hydrogen and carbon monoxide gas yield showed maximum values at the medium hydrogen peroxide concentration. As the hydrogen peroxide concentration increased from 0 to 2701.1 mmol/L (from 0 to 12.4% oxygen based on stoichiometric oxygen requirement for the total oxidation of isooctane based on Eq. (13) in Appendix), the hydrogen gas yield increased from 4.00 to 6.13 mol/mol isooctane, and the carbon monoxide gas yield increased from 0.05 to 0.16 mol/mol isooctane. Further increase in the hydrogen peroxide concentration to 4507.3 mmol/L (20.9% of oxygen) resulted in decreasing the hydrogen gas yield to 4.56 mol/mol isooctane and in decreasing the carbon monoxide gas yield to 0.12 mol/

mol isooctane. The previous study with the top-down reactor configuration showed the similar hydrogen gas yield trend with changing the hydrogen peroxide concentration [4]. A steady increase in the carbon dioxide yield from 1.44 to 3.48 mol/mol isooctane and a slight increase in the methane gas yield from 1.19 to 1.94 mol/mol isooctane were observed as the hydrogen peroxide concentration increased from 0 to 4507.3 mmol/L. At the maximum value of hydrogen gas yield at a hydrogen peroxide concentration of 2701.1 mmol/L, the dry gases consisted of 59.5% hydrogen, 1.6% carbon monoxide, 22.0% carbon dioxide, 14.8% methane and 2.1% ethane and propane. Thus, partial oxidation of isooctane in the presence of medium hydrogen peroxide concentration promoted formation of hydrogen and carbon monoxide. This is in good agreement with theoretical prediction (Eq. (12) in Appendix). The decrease in the hydrogen and the carbon monoxide gas yields, and the increase in the carbon dioxide gas yield at higher hydrogen peroxide concentration suggest that further oxidization of the produced gases into carbon dioxide and water proceeded (Eq. (13) in Appendix). Higher oxidant concentration in the supercritical water condition will result in complete oxidation of isooctane to form carbon dioxide and water. An increase in the methane gas yield at higher hydrogen peroxide concentration also indicates the consumption of hydrogen via the methanation reaction (Eq. (4)–(5) in Appendix). As shown in Fig. 5c, alkenes species including propylene and butylene were observed in the absence of hydrogen peroxide. When the hydrogen peroxide concentration was 900 mmol/L, only butylene was observed as the unsaturated hydrocarbon species, and when the hydrogen peroxide concentrations were equal or higher than 2701.1 mmol/L the unsaturated hydrocarbons were not observed in the gaseous products. This suggests that the hydrogenation might take place with the additional oxygen (Eq. (14) in Appendix). Arai’s group reported that hydrogenation of 1-alkene in supercritical water occurred during the partial oxidation [47,48]. Higher ratio of n-alkane/1alkene was obtained in partial oxidation than in pyrolysis or water gas shift reaction in supercritical water condition.

3.5.

New reactor design concept

Table 1 lists comparison of the gas yield, CE, composition of product gases and hydrogen gas yield from the previous

Table 1 – Supercritical water gasification comparison between the up-down and the down-up gasifier configuration. Reactor Configuration

Experiment condition feed residence H2O2 T time concentration [ C] concentration [wt%] [s] [mmol/L]

Gas CE Dry gas composition [% mol] Hydrogen Ref yield [L/g] [% mol] yield H2 CO CO2 CH4 C2-5 [mol/mol]

up-downa down-upa

664 640

21 20

16 18c

– –

0.51 1.87

16.7 45.8

34.1 46.6

2.0 1.7

19.9 17.2

37.4 26.5

6.6 4.2

0.63 2.61

up-downb down-upb

664 637

22 9.9

15 18c

1560.2 2701.1

1.38 2.94

49.1 55.8

31.9 59.5

3.4 1.6

20.7 22.0

37.6 14.8

6.5 2.1

1.68 6.13

a Similar gasification conditions. b The maximum hydrogen gas yield conditions. c Residence time was calculated based on the reaction zone.

[4] This study [4] This study

1967

0.5 0.6

0.7 6.7 7.1  3.1 32  3 32

Tubular reactor (up-down) No Tubular reactor (horizontal) No

18 2.0–15.1 5 5 10 1 C3H5(OH)3 C2H5OH

a Not available.

25 25 25 25 21 21 630 650 600 650 600 700 C8H18 C6H12O6

T [ C] P [MPa] residence times feed concentration

9.8 wt% 0.1 M 10 wt% 10 wt% 10 wt% 10 wt%

Tubular reactor (down-up) Microchannel reactor

No No

55.9 16.6 20.1 47.2  2.4 8.1  1.3 36.4  2.7 50  4 51 91 NAa 57.5 15.3 26.8 57.7 15 26.7

H2

CH4

CO2

CO

C2þ

4.0 2.6  0.5 1.84  0.32 2.5 2.6 2.8

[50] [16]

Ref. H2 yield mol/mol feed Dry gas composition, % mol Catalyst Reactor type Reaction condition Feedstock

Table 2 – Comparison of supercritical water gasification results.

up-down [4] and the present down-up gasification configuration. The down-up configuration resulted in much better gas yield (w3.7 times higher) and carbon gasification efficiency (w2.7 times higher) compared to the up-down type at the similar gasification conditions. In addition, the down-up configuration resulted in higher hydrogen gas yield (w4 times higher) and in higher hydrogen composition in the produced gas compared to the up-down type at the similar gasification condition. The maximum hydrogen gas yield that were achievable from the previous up-down and the present downup configuration are listed in Table 1. The maximum hydrogen gas yield obtained from the down-up type was w4 times higher than that of the up-down type. The previous up-down gasifier configuration has several drawbacks in the supercritical water gasification [4]. The reactor was vertically set up at above a condenser. The feed (isooctane and DDI water) flowed from the top and the gaseous products and the liquid effluent flowed out at the bottom of the reactor. The high-pressure and high-temperature fluid in the reactor was quenched suddenly to room temperature in the condenser region. This leads that the gases and the liquid, that formed a single-phase at the supercritical water condition in the reactor, separated suddenly each other in the condenser region. The gaseous products may have tendency to flow upward at this condition because the water density at the condenser region (0.1665–1.009 g/cm3) is larger than those of the gases produced. Meanwhile the pressurized liquid tends to flow downward. It was often observed that when the flow rate was high or/and when a large amount of gas was produced, there was virtual plugging at near the bottom part of the reactor and/or in the condenser. This may be caused by a collision between the gas flow and the liquid flow. In the course of reaction, it was observed that a rapid increase in the system pressure approximately 10 MPa higher than the setting pressure followed by a rapid pressure drop to 15 MPa, possibly caused by BPR limitations to handle abrupt changes in flow. As a result, gaseous products sometimes burst in the separator region and the liquid flow brought the gases into the liquid container. This may cause low gas yield and low hydrogen yield. In the present reactor design, the reactor was inclined around 75 from a vertical position. The feed and water were introduced at the bottom of the reactor and the gas products and the liquid effluents flowed out at the top of the reactor. Meanwhile, the fluid could flow easily because the increasing slope was applied. The presence of the cooling zone, that was attached to the reaction zone and used outside air as the cooling medium, may also give contribution to the increasing of hydrogen gas yield. This design allows for gradual temperature decrease as the flow travels toward the end of the cooling zone. As a result, the highly reactive free radical intermediates that are produced during the gasification may react further in the cooling zone and may contribute to the enhancement of the gas yield. The inlet temperature of the insulation zone was in the range of 650–565  C and the inlet temperature of the cooling zone was in the range of 430– 402  C. These temperatures are high enough to have further gasification reaction in the insulation zone and in the aircooled tube. The outlet temperature of air-cooled tube was around 200–250  C. It is noted that at a low temperature of

This study [22]

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international journal of hydrogen energy 35 (2010) 1957–1970

600  C, supercritical water gasification can produce 0.93 mol H2/mol isooctane. In addition, as reported by Yanik et al. [49], a reaction temperature of 500  C was high enough to have gasification reaction of sunflower, corncob, and leather waste without using catalysts. The gas yields of H2, CO2, CO and CH4 were 3.65, 11.15, 0.04 and 4.02 mol gas/kg sunflower, respectively; 2.09, 6.36, 0.25, and 2.04 mol gas/kg corncob, respectively; 2.66, 5.93, 0.06 and 2.39 mol gas/kg leather waste, respectively [49]. The fluid had longer residence time, contributed to the increment in the hydrogen yield without the additional of external energy. Beside the contribution to the gas yield, the cooling zone has impact to maximum temperature that can be achievable in this study. In the previous reactor design, the short distance (10 cm) between the heating zone and the condenser region resulted in abrupt heat transfer, leading to decreasing temperature at the end region of the reactor. This is the reason why we used relatively long cooling tube (115 cm) that separates the reaction zone and the condenser (C-01) at the end of the cooling zone and added the insulation zone between the reaction zone and the cooling zone to keep the reaction temperature isothermal. Therefore, the down-up gasifier configuration, and the addition of the insulation zone and the cooling zone resulted in higher gas yield and higher hydrogen gas yield when compared with the down-up gasifier configuration at the similar gasification conditions. Table 2 lists comparison of the gasification results from other studies [16,22,50] under the similar conditions. The feedstock used in this study has the longest carbon chain compared to the others (glucose, glycerol or ethanol). The hydrogen gas yield in this study was approximately 1.5 times higher that those from the other studies. As well, the composition of hydrogen in the produced gases was higher than those obtained by the gasification of glucose in microchannel reactor and glycerol in tubular reactor with up-down configuration. In addition, carbon monoxide in the produced gases was 10–45 times lower than those gasified from glucose and glycerol. Based on the gasification behavior examined in this study, hydrogen yield increased at higher residence time and/or at higher temperature. On the basis of this result, next development will be focused on the increasing of hydrogen yield by using a larger volume of reactor and using higher reaction temperatures.

and an oxidant concentration of 2701.1 mmol/L. Possibility to increase hydrogen gas yield at higher temperature and longer residence time is still open widely.

Acknowledgement This research was supported by Korea Environmental Industry & Technology Institute (Project No. 202-091-003).

Appendix The possible reaction pathways for gasification of isooctane in supercritical water based on the methane reforming reaction was defined as follow [4]: Isooctane steam reforming: C8H18 þ 8H2O 4 8CO þ 17H2 DH298K ¼ 1274.47 kJ/mol

(1)

Water gas shift reaction: CO þ H2O 4 CO2 þ H2 DH298K ¼ 41.15 kJ/mol

(2)

Combined reaction of (1) and (2): C8H18 þ 16H2O 4 8CO2 þ 25H2 DH298K ¼ 945.27 kJ/mol

(3)

Methanation of CO: CO þ 3H2 4 CH4 þ H2O DH298K ¼ 206.17 kJ/mol

(4)

Methanation of CO2: CO2 þ 4H2 4 CH4 þ 2H2O DH298K ¼ 165.01 kJ/mol

(5)

The possible reaction pathways which are responsible for carbon formation: Isooctane cracking: C8H18 4 8C þ 9H2 DH298K ¼ 224.1 kJ/mol

(6)

Methane cracking: CH4 4 C þ 2H2 DH298K ¼ 74.87 kJ/mol

(7)

Boudouard reaction (CO disproportionation):

4.

Conclusion 2CO 4 C þ CO2 DH298K ¼ 172.45 kJ/mol

The supercritical water gasification has been proved to be an effective method to produce hydrogen from isooctane with low carbon monoxide content without using catalysts. A new reactor design was developed to achieve high hydrogen gas yield and good gas–liquid flow rate stability. The results showed that high hydrogen gas yield was achieved at high temperature, long residence time, low isooctane concentration and with appropriate amount of oxidant. The produced gases consisted of hydrogen, methane and carbon dioxide as major composition, and carbon monoxide and C2–C4 species as minor composition. The maximum hydrogen gas yield was 6.13 mol/ mol isooctane at a temperature of 637  C, a pressure of 25 MPa, a residence time of 18 s, an isooctane concentration of 9.9 wt%

(8)

CO hydrogenation: CO þ H2 4 C þ H2O DH298K ¼ 131.29 kJ/mol

(9)

CO2 hydrogenation: CO2 þ 2H2 4 C þ 2H2O DH298K ¼ 90.14 kJ/mol

(10)

The hydrogen peroxide (H2O2) decomposition: H2O2 / 0.5O2 þ H2O DH298K ¼ 105.72 kJ/mol The isooctane partial oxidation:

(11)

international journal of hydrogen energy 35 (2010) 1957–1970

C8H18 þ 4O2 4 8CO þ 9H2 DH298K ¼ 748.13 kJ/mol

(12) [18]

The isooctane total oxidation: C8H18 þ 25/2O2 4 8CO2 þ 9H2O DH298K ¼ 5100.53 kJ/mol

(13)

The n-alkenes hydrogenation reaction: CnH2n þ H2 4 CnH2nþ2

[19]

[20]

(14) [21]

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