Industrial scale steam reforming of bioethanol: A conceptual study

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 e n e r g y 3 5 ( 2 0 1 0 ) 8 4 7 2 e8 4 8 5

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Industrial scale steam reforming of bioethanol: A conceptual study Jim H. Oakley*, Andrew F.A. Hoadley Department of Chemical Engineering, Faculty of Engineering, Monash University, Victoria 3800, Australia

article info

abstract

Article history:

This study investigates the technical feasibility of using existing steam reforming and

Received 20 February 2010

hydrogen separation technologies to produce hydrogen from bioethanol on an industrial

Received in revised form

scale (100,000 Nm3/h of hydrogen). Two separate reaction schemes for producing hydrogen

1 May 2010

from bioethanol are compared. One scheme utilises both a prereformer and steam

Accepted 3 May 2010

reformer, and the other requires a steam reformer only.

Available online 18 June 2010

Using data from PRO/II simulations, the findings for the processes developed around the two reaction schemes are compared, based on a range of quantitative and qualitative

Keywords:

measures. Full heat integration is conducted for both schemes, with two different cases

Bioethanol

developed for the prereformer scheme.

Hydrogen

Ethanol-to-hydrogen processes based on catalytic reactions can only be brought into

Industrial scale

production when suitable reaction catalysts have been developed and proved. This is the

Steam reforming

major technical challenge which must be overcome to enable the processes investigated in this study to be implemented. ª 2010 Professor T. Nejat Veziroglu. Published by Elsevier Ltd. All rights reserved.

1.

Introduction

1.1.

Background

Two of the major challenges facing the global community are climate change, and the depletion of easily accessible hydrocarbon energy resources. One approach to sustainable development which may help to reduce the emission of greenhouse gases, and which represents an alternative, renewable source of fuel is the development of biofuels. The two major biofuels at present are bioethanol and biodiesel, of which bioethanol is the most plentiful [1]. A major review of 47 published life-cycle assessments comparing bioethanol systems to conventional fuel reported that: ‘The overriding conclusion of the studies that looked at energy balances was that the use of bio-ethanol in place of

conventional fuels or as an additive leads to a net gain. That is, the prevailing data indicate that it takes less energy to make and distribute ethanol than can be delivered by the fuel’ [2]. Consistent with this, the same study found that the use of bioethanol to replace conventional fuels results in lower greenhouse gas emissions. The size of the reduction depends on the crop from which the bioethanol is derived. For example, sugar-based production systems achieve much greater avoided greenhouse gas emissions than do starchbased systems [2]. Because of limited land availability and feedstock supply e where the current feedstocks are also important food sources e research and development programs have targeted the production of bioethanol from lignocellulosic materials. These feedstocks are typically grouped under the category of ‘biomass’, and include

* Corresponding author. Tel.: þ612 4283 8795; fax: þ612 4275 7421. E-mail addresses: [email protected] (J.H. Oakley), [email protected] (A.F.A. Hoadley). 0360-3199/$ e see front matter ª 2010 Professor T. Nejat Veziroglu. Published by Elsevier Ltd. All rights reserved. doi:10.1016/j.ijhydene.2010.05.003

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 e n e r g y 3 5 ( 2 0 1 0 ) 8 4 7 2 e8 4 8 5

Nomenclature DH0298

Standard heat of reaction at 298K and 101325 Pa

agricultural residues, wood (e.g., from poplar and willow trees), municipal solid waste, and energy crops such as switchgrass [1,3,4]. One broad vision for the future is the ‘hydrogen economy’ [41], where hydrogen supplies the fuel energy in transportation systems. The hydrogen production may be decentralised, where it is produced close to its point of use; or centralised, where it is produced in large-scale facilities and distributed to where it is required. Sixteen of the world’s most economically advanced countries, along with the three largest developing countries (China, India and Brazil), have created the International Partnership for the Hydrogen Economy (IPHE). The participating countries have: ‘committed to accelerate the development of hydrogen and fuel cell technologies to improve the security of their energy supply, environment, and economy’ [5].

1.2.

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quantities of sustainably-generated hydrogen, either in the form of synthesis gas or pure hydrogen. The specific goal of the work described in this paper is to devise processing schemes to produce pure hydrogen from bioethanol on an industrial scale, using conventional steam reforming and hydrogen separation technologies. It is intended that these schemes will serve as base cases upon which improvements can be made, subject to further research and specialist vendor advice. An important assumption in this work is that catalysts with appropriate selectivity, stability and robustness will be available to allow the processing schemes to operate cost effectively. To date, no such catalysts have been demonstrated as suitable for use on an industrial scale. Two reactions schemes for the production of synthesis gas from ethanol, which have been reported in the literature, will be investigated here. The reaction schemes have been selected so that one involves the presence of methane, and the other does not. Heats of reaction have been calculated from the standard heats of formation for gaseous species presented by Shallcross [16]. Reaction Scheme 1 [9]

This work

It is almost certain that large-scale production of bioethanol from lignocellulosic sources will become technically and commercially viable long in advance of a hydrogen economy being established. Much of the bioethanol will be used for automotive purposes as an additive to petrol, or in pure form as a liquid fuel, or potentially as a hydrogen source in on-board fuel cells. Notwithstanding this, other opportunities will also be present for utilising bioethanol as a hydrogen source. The sustainable manufacture of ammonia and methanol, and the sustainable supply of hydrogen to oil refineries, are three applications which will require large quantities of hydrogen. The operators of these processes, which serve as hydrogen sinks, will eventually come under pressure e economic, political, or both e to make their processes ‘greener’. This also applies to hydrogen users in the chemical, food and steel industries. It is conceivable that circumstances may develop in certain locations such that the preferred use of bioethanol will be steam reforming to make hydrogen on a large scale. Similarly in certain industries, circumstances may develop such that the preferred means of obtaining hydrogen will be by steam reforming of bioethanol. Steam reforming of hydrocarbons is a process which has been developed and applied over a period of more than 70 years [6]. It is currently the most mature and best-established hydrogen production technology. In recent years, plant costs have been reduced due to a number of developments. These include better construction materials; improved control of carbon limits; and use of advanced catalysts, allowing greater feedstock flexibility [7,8]. Over the coming decades new technologies may emerge; however, for large-scale hydrogen production it is reasonable to view steam reforming as a technology of choice for the foreseeable future. It is a well-proven technology with low associated commercial and technical risk. It is the contention of this study that conversion of biofuels using conventional steam reforming technology, in conjunction with novel catalysts, is likely to be one of the most rapid and economical means of introducing to the market large

C2H5OH þ H2O $ 2CO þ 4H2 (DH0298 ¼ þ255.2 kJ/mol)

(1)

CO þ H2O $ CO2 þ H2 (DH0298 ¼ 41.2 kJ/mol)

(2)

Reaction Scheme 2 [10] C2H5OH / CH4 þ CO þ H2 (DH0298 ¼ þ49.0 kJ/mol)

(3)

CH4 þ H2O $ CO þ 3H2 (DH0298 ¼ þ206.2 kJ/mol)

(4)

CO þ H2O $ CO2 þ H2 (DH0298 ¼ 41.2 kJ/mol)

(5)

This paper will not only explore the robustness of these two reaction schemes, but also the thermal efficiency of the implementation of both schemes within a large-scale industrial process. Due to high temperatures required for steam reforming, energy recovery is an important aspect of the process design. In this study, heat pinch analysis [11,12] has been used to ensure that energy recovery has been maximised. In the case of Reaction Scheme 2, two different heat recovery flowsheets are analysed.

2.

Calculations

2.1.

Thermodynamic equilibrium calculations

2.1.1.

Basis of calculations

Early studies of the thermodynamic equilibrium of hydrogen production via the steam reforming of ethanol were done at pressures below 1000 kPa [42,43]. For Reaction Schemes 1 and 2, the thermodynamic equilibrium was explored within and above the pressure range at which steam reforming of natural gas for syngas generation is economically viable: 2000e4000 kPa [6,8]. Steam-to-ethanol feed ratios from 4 to 10

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were used in the equilibrium studies, reflecting typical steamto-carbon ratios (2e5) used in the reforming of a variety of hydrocarbon feedstocks. After examination of the results, a pressure of 3000 kPa(a) was selected as the reformer pressure on which to base simulations. A steam-to-ethanol feed ratio of 5 was also selected, leaving temperature as the remaining variable to be determined to optimise hydrogen production. Full details of the thermodynamic equilibrium studies are available elsewhere [13]. The equilibrium product composition for each reaction scheme was determined either in whole (Reaction Scheme 1) or in part (Reaction Scheme 2) using the Gibbs free energy minimisation method (or non-stoichiometric method). This is consistent with the approach preferred for the steam reforming of hydrocarbons, and allows the presence of elemental carbon in the equilibrium product mixture to be accurately quantified. Because solid carbon may be produced during steam reforming of hydrocarbons, the minimisation of Gibbs free energy is the method normally used to determine the equilibrium chemical composition of these systems [14,15]. Calculations were done using Simsci-Esscor’s PRO/II v5.6 process simulation software, made available through Monash University. The Soave-Redlich-Kwong equation of state [45] was used to estimate the volume-temperature-pressure relationship for each component in the mixture, allowing the required thermodynamic parameters to be calculated.

2.1.2.

2CO $ C þ CO2 (DH0298 ¼ 172.5 kJ/mol)

(6)

CO þ H2 $ C þ H2O (DH0298 ¼ 131.3 kJ/mol)

(7)

Both these reactions are exothermic. According to Le Chatelier’s principle, as the temperature increases the equilibrium will move further towards the left, until a temperature is reached at which carbon is no longer present. This is 750e800  C for Reaction Scheme 1 (see Fig. 1). Each reaction has 2 mol on each side of the equation, and following Le Chatelier’s principle their equilibrium should be independent of pressure. Below 750e800  C, the concentration profiles in Fig. 1 are dominated by the CO reduction reaction (Equation (7)). As the proportional amounts of carbon and water decrease with increasing temperature, there is a corresponding rise in the proportional amounts of carbon monoxide and hydrogen. The amount of carbon dioxide also increases because of the watergas shift reaction (Equation (2)). The latter is an exothermic reaction, and as the temperature increases its equilibrium increasingly favours carbon monoxide and water. The large excess of water in the system promotes the water-gas shift reaction, but as the temperature increases the amount of water in the system reduces and this will also tend to moderate the extent to which the forward reaction proceeds. Once carbon has disappeared from the mixture, the concentration profiles in Fig. 1 reflect the movement of the water-gas shift equilibrium to the left. The levels of hydrogen and carbon dioxide decrease, while those of water and carbon monoxide increase.

Reaction Scheme 1

For the PRO/II simulation, a Gibbs reactor was the only unit operation required to model the equilibrium product composition for the species involved. For each individual pressure, equilibrium compositions were calculated at temperatures between 350  C and 1100  C inclusive. This allowed a comprehensive picture to be constructed of the variation in equilibrium product composition in response to changes in temperature, pressure and the steam-to-ethanol feed ratio. The inclusion of carbon in the Gibbs free energy minimisation simulation effectively accounts for the two most likely potential carbon deposition reactions:

2.1.3.

Reaction Scheme 2

Reaction Scheme 2 occurs in two stages, corresponding to the use of two different catalysts to complete the different reactions which comprise the reaction scheme. The first stage is the irreversible decomposition of ethanol, accompanied by the reversible water-gas shift reaction (Equations (3) and (5)). These reactions were reported by Galvita et al. [10] to occur over an experimental catalyst, and the experiments were

7 6 H2

5

CO

4

CO2

3

EtOH H2O

2

C

1 1100

1050

1000

950

900

850

800

750

700

650

600

550

500

450

400

0 350

Moles Component/Mole Ethanol

Reaction Scheme 1 Equilibrium Composition vs Temperature Pressure = 3000 kPa Steam:Ethanol Ratio = 5

Temperature (°C)

Fig. 1 e Equilibrium composition versus temperature for Reaction Scheme 1 at a pressure of 3000 kPa(a) and with a steam-to-ethanol feed ratio of 5.0.

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done over a restricted range of temperatures. In contrast to the first stage, the second stage is the reversible methane steam reforming reaction, also accompanied by the reversible water-gas shift reaction (Equations (4) and (5)). These reactions are well known, and the associated catalysts are well established and proven over a wide range of temperatures. Thermodynamic equilibrium modelling was performed to guide the development of economically viable industrial processing schemes. Because Reaction Scheme 2 essentially consists of a novel stage and an established stage, the modelling was done to minimise the risk of invalid conclusions being made in relation to the first (or novel) stage reactions. To this end, the product composition of the first stage water-gas shift reaction was calculated at 370  C. This corresponds to the upper temperature used by Galvita et al. [10], and at this temperature Equation (3) was experimentally found to be irreversible over the range of steam-to-ethanol ratios used in their reported work. The assumption in the present work is that Equation (3) remains irreversible at pressures up to 3000 kPa and beyond. Because the two stages of Reaction Scheme 2 require different catalysts, and the first stage is restricted to a maximum temperature of 370  C, on an industrial scale it is sensible to use two separate reaction vessels to ensure the reactions can proceed at the required conditions. With this in mind, equilibrium modelling was done with the first stage modelled as an adiabatic prereformer, and the second stage modelled as an isothermal equilibrium reformer. Both reaction stages were simulated at identical pressures. The two reactions which occur in the first stage were simulated using adiabatic reactors. The outlet temperature was set at 370  C, requiring the inlet temperature to be adjusted to meet this requirement. This ensured that, for a fixed ethanol to water ratio entering the first stage, and a fixed pressure, the outlet composition entering the second stage was constant. However, the temperature of the second stage reactions was allowed to vary from 350  C to 1100  C, enabling changes in equilibrium product compositions to be observed over the range of temperatures.

Fig. 2 depicts how the equilibrium model was constructed in PRO/II v5.6. The ethanol decomposition was modelled as an irreversible reaction; that is, ethanol is completely converted. The accompanying water-gas shift reaction equilibrium was simulated using PRO/II’s internal model for this reaction. Alternatively it could have been modelled with a Gibbs reactor. Equilibrium for the stage 2 reactions was simulated using a Gibbs reactor. The stage 2 reactions are known to be the primary reactions occurring in the presence of a methane steam reforming catalyst, and the calculated Gibbs reactor equilibrium compositions in Fig. 3 should be interpreted in the context of these reactions. The methane steam reforming reaction (Equation (4)) is endothermic, and is promoted by higher temperatures. The water-gas shift reaction (Equation (5)) is mildly exothermic, and is favoured by lower temperatures. A rising carbon dioxide yield corresponds to the water-gas shift reaction proceeding in the forward direction. A decreasing carbon dioxide yield occurs when the water-gas shift reaction proceeds in the reverse direction. The temperature ranges where this occurs can be seen in Fig. 3.

2.2. Simulation of thermally integrated hydrogen production processes 2.2.1.

Design parameters common to all simulations

The approach adopted for designing thermally integrated hydrogen production schemes is to use, as a basis, the standard design for a methane steam reforming plant, with modifications introduced as necessary to account for the liquid water-ethanol feedstock.

2.2.1.1. Hydrogen production rate. Each scheme produces 100,000 Nm3/h of hydrogen, where normal (N) indicates 0  C and 101.325 kPa(a). This is a hydrogen production rate well within the established range for large-scale methane steam reforming plants [17,18]. The hydrogen production processes documented in this section require feedstock of between 330,000 and 354,000

Fixed outlet temperature

Fixed steam-to-ethanol ratio

Stage 1

370°C C2H5OH

CH4 + CO + H2

CO + H2O

CO2 + H2

Fixed pressure Adjustable temperature

Adiabatic Equilibrium Reactor

Adiabatic Reactor

Variable temperature

Stage 2 CH4 + H2O CO + H2O

CO + 3H2 CO2 + H2

Isothermal Gibbs Reactor Fig. 2 e Schematic of thermodynamic equilibrium model for Reaction Scheme 2, constructed and run in PRO/II.

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Moles Component/Mole Ethanol

Reaction Scheme 2 Equilibrium Composition vs Temperature Pressure = 3000 kPa Steam:Ethanol Ratio = 5 6 5 H2 CO

4

CO2

3

EtOH H2O

2

CH4 C

1

1100

1050

1000

950

900

850

800

750

700

650

600

550

500

450

400

350

0

Temperature (°C)

Fig. 3 e Equilibrium composition versus temperature for Reaction Scheme 2 at a pressure of 3000 kPa(a) and with a steamto-ethanol feed ratio of 5.0. Note that the equilibrium compositions of ethanol and carbon are approaching zero.

tonnes of pure ethanol each year, assuming the plant is online for 350 days per year. This is equivalent to a liquid ethanol volume of between 418,000 and 448,000 kl/y. The largest-scale bioethanol production facilities produce more than 1,200,000 kl/y of liquid ethanol [19] and so, based on current technology, a mid-to-large scale bioethanol plant can supply enough ethanol for a steam reforming plant producing 100,000 Nm3/h of hydrogen.

2.2.1.2. Operating pressure. The economic range of operating pressures in industrial synthesis gas reformers is 2000e4000 kPa [6]. For hydrogen production, the economic range represents a trade-off between several competing factors. At lower pressures, the methane steam reforming reaction (Equation (4)) proceeds further in the forward direction at equilibrium, producing more hydrogen. However, lowering the pressure results in larger process vessels and pipes, which are more costly. Increasing the pressure reduces equipment volume, but increases wall thickness, which can also be costly. Higher pressures also require higher temperatures to generate the same hydrogen yield. Perhaps the two most important considerations in determining the reformer pressure in a hydrogen production plant are the optimum pressure for the downstream pressure swing adsorber (PSA), and the pressure at which the hydrogen product is required [20]. The PSA unit operates at pressures up to 3500 kPa(a) [21]. Customers often require hydrogen at pressures above 2000 kPa(a), so operating the methane steam reforming plant above this pressure avoids the need for further compression [7]. Each application will have its own optimum reformer operating pressure, determined by a range of project-specific parameters. In this work the pressure at the steam reformer outlet was set to 3000 kPa(a), which represents a reasonable value well within the economic range [22]. In the simulations, nominal pressure drops were included to approximate the pressure losses that would occur across unit operations in

a real plant. Therefore, feed pressures are somewhat higher than 3000 kPa and product pressures are lower.

2.2.1.3. Water-to-ethanol ratio. In methane steam reforming, the steam-to-methane molar ratio (or steam-to-carbon ratio) has typically been set at 2.5e3.0 [7,23]. The excess steam helps to ensure that carbon is not formed. Recent developments in steam reforming technology allow steam-to-carbon ratios below 2.5, even as low as 2.0. This has the advantage of reducing the steam flow rate, and therefore the size of the equipment [17]. Energy requirements are also lower, since less steam has to be raised. For the hydrogen production processes reported here, a water-to-carbon molar ratio of 2.5 was used for the feed entering the process. This is equivalent to a water-to-ethanol molar ratio of 5.0. In bioethanol production plants, ethanol is produced via fermentation. The concentration of ethanol in the mash leaving the fermentation process is usually 9e13 vol% [19,24]. This corresponds to a mole fraction of approximately 3.0e4.4 mol% ethanol, assuming a pure water-ethanol mixture [25]. The mash is fed to a series of distillation columns where the ethanol is concentrated to approximately 96 vol% (86 mol%). A water-toethanol ratio of 5.0 in the feed to a hydrogen production scheme corresponds to an ethanol concentration of 16.7 mol%. It is anticipated that a modified distillation system could be used to concentrate the ethanol in the mash from around 4 mol% to 16 mol%, and this would require significantly less energy than a distillation system concentrating ethanol to near the azeotropic composition. A modified distillation system would still need to separate the water-ethanol mixture from other byproducts of ethanol fermentation such as stillage, carbon dioxide and fusel oils [26]. 2.2.1.4. Feed purity. It is assumed that, after distillation to the correct ethanol concentration, the water-ethanol mixture is of sufficient quality to feed the hydrogen production process. This means that solids and specific catalyst poisons are not

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present, and that higher alcohols and other components are present only in concentrations low enough to have no adverse impact on the process, particularly the catalysts. For methane steam reforming catalysts, particular poisons are sulphur, arsenic, copper, chlorine and lead [27]. It is possible that distillation alone will not be able to produce a water-ethanol mixture of suitable purity, and that other separation steps will have to be used in addition to, or instead of, distillation.

2.2.1.5. Thermodynamic methods. In any unit operation or stream in which ethanol was present, the Soave-RedlichKwong Modified Panagiotopoulos-Reid (SRKM) equation [44] was used. When ethanol was not present, the Soave-RedlichKwong (SRK) method [45] was used. The Rackett method was used to calculate liquid density. These selections were made based on the detailed information and recommendations in the Simsci documentation which accompanied the PRO/II software.

2.2.1.6. Furnace duty. In industrial methane steam reformers, approximately 50e60% of the heat generated in the reformer furnace is transferred to the reformer tubes. The remainder of the heat is available for process heating and steam generation. Overall, about 95% of the furnace heat is recovered [7,17,28]. In each simulation reported here, the reformer furnace generated enough heat to: a) satisfy the reformer duty; and b) provide the furnace flue gas with the equivalent of 90% of the reformer duty between the furnace discharge temperature and the exhaust fan entry temperature of 150  C.

3.

Results

3.1.

Case studies

Three case studies were undertaken: Case 1 is based on hydrogen production using Reaction Scheme 1 in a steam-reforming unit (see the basic flowsheet in

Fig. 4). Operating parameters used in the PRO/II simulation are detailed in Table 1. The data obtained from initial PRO/II simulations allowed heat integration of the process flowsheet to be performed. Composite curves were constructed, along with the grand composite curve [11,12]. The grand composite curve has the advantage that it not only quantifies hot and cold utility duties, but also indicates the range of temperatures at which the utilities can be supplied ). The balanced grand composite curve for Case 1 is presented in Fig. 5, showing both the grand composite curve for the process and the amount of 10,000 kPa(a) steam which can notionally be generated from the excess process heat, along with cooling water requirements. Table 3 documents the target cold utility requirements for Case 1 as calculated from the grand composite curve, and compares these values with the values obtained from the simulation of the integrated flowsheet. The target values are for maximum steam production and minimum cooling water usage. Case 2 employs Reaction Scheme 2 in a prereformer and a reformer to generate hydrogen (see the basic flowsheet in Fig. 6). Operating parameters specific for Case 2 are presented in Table 2. Other parameters used in the PRO/II simulation are shown in Table 1. The balanced grand composite curve for Case 2 is presented in Fig. 7, where it can again be seen that the excess heat from the process is notionally used to generate 10,000 kPa(a) steam. Table 4 documents the target maximum steam generation and minimum cooling water requirements, and the corresponding results derived from the simulation of the integrated process flowsheet. Case 3 is identical to Case 2 except that the extent of heat integration has been reduced to enable a system to be devised where only 15 heat exchangers are required, as opposed to 23 in Case 2. The amount of steam generated in Case 3 is 98.2% of that generated in Case 2, as can be seen from Table 4. This is the penalty incurred due to relaxing the extent of the heat integration and rejecting more heat to cooling water. The data in Table 5 allow a quantitative comparison of the three case studies to be undertaken.

Ethanol & water Exhaust Fan Pump

Reformer

Vaporiser

To Stack

Shift Reactor

Hydrogen

Fuel

Air

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Decanter

PSA

Fig. 4 e Basic hydrogen production process for Reaction Scheme 1 prior to heat integration.

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Table 1 e Operating conditions for producing hydrogen via Reaction Scheme 1. Unit operation Feed pump

Vaporiser

Reformer

Shift reactor

Decanter Pressure swing adsorber (PSA)

Reformer furnace

Exhaust fan

3.2.

Operating parameter

Value

Feed temperature Outlet pressure



25 C 3680 kPa

Simulation

Outlet temperature

254  C

Simulation

Pressure

3550 kPa

Simulation

Inlet temperature

800  C

Outlet temperature

850  C

Pressure Pressure drop

3000 kPa 400 kPa

Temperature approach to equilibrium Inlet temperature

10  C

[20,23,40]

210  C

[17]

Temperature approach to equilibrium Thermal duty

0 C

[40]

0 kW

[23,40]

Pressure drop Temperature Inlet temperature Outlet temperature Hydrogen recovery Hydrogen purity Pressure drop Off-gas pressure Inlet pressure

400 kPa 35  C 35  C 30  C 90% 100% 50 kPa 120 kPa 120 kPa

[37] [35,37] [35] [21,35] [35,36] Assumed [20] [20]

Externally-sourced fuel

Natural gas

[38]

Excess air % Combustion heat transferred to Reformer tubes Inlet temperature

10% 52%

[11] [6,7,28,39]

150  C

[20]

Details of the heat integration

The construction of the composite curves only utilised data for the streams which can potentially undergo process-toprocess heat exchange. Thus streams which need to be heated or cooled, or vaporised or condensed, were included. The flue gas from the reformer furnace was treated as a process stream since the furnace is integral to the reformer, and also combusts some of the process product gas. The process gas stream passing through the reformer, which has a large heating requirement, was not included because its heating duty is supplied from a fixed source (the reformer furnace), and it is therefore not available for any other potential process-to-process heat exchange until it leaves the reformer [13]. In line with common industrial practise, the combustion air to the reformer furnace was preheated. Also consistent with current practise, the reformer furnace fuels e natural gas and the PSA off-gas e were not preheated [17,23,37].

Source

Thermodynamic equilibrium graphs Thermodynamic equilibrium graphs [6,20,22,37] [20,27,35]

4.

Comment Ambient temperature Depends on number of heat exchangers before Reformer Dew point of water-ethanol mixture plus 20  C superheat Depends on number of heat exchangers before Reformer Avoids carbon deposition. Avoids carbon deposition. Hydrogen yield is good. Economic operating pressure Typical value, may be higher end of range Reasonable design value Typical value for medium temperature shift reactor Shift reaction is at equilibrium with stable catalyst Shift reactor normally operates adiabatically Same as Reformer Typical value Typical value Typical value Typical value Typical value >99.9% For hydrogen stream Typical value Very low pressure drop across burners Similar composition to natural gas from the Gladstone region of Queensland Typical value for gaseous fuels Typical value

Typical flue gas temperature after waste heat recovery

Discussion

4.1. Thermodynamic equilibrium analysis and reformer operating temperatures 4.1.1.

Reaction Scheme 1

The chosen reformer conditions are in the range of those used in methane steam reforming: a pressure of 3000 kPa(a); inlet steam-to-carbon ratio of 2.5; inlet temperature of 800  C; and outlet temperature of 850  C. These are all reasonable based on the thermodynamic equilibrium calculations. The critical issue was to choose the reformer temperatures to minimise the possibility of carbon forming inside the reformer, while at the same time maintaining a high hydrogen yield. The reformer inlet temperature of 800  C is at least 20  C above the highest temperature at which carbon will form when the product gases are at equilibrium at the chosen conditions. However, as is well known in methane steam

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Balanced Grand Composite Curves

Shifted Temperature (°C)

1000 900

Grand Composite Curve

800

10000 kPa(a) Steam and Cooling Water

700 600 500 400 300 200 100 70000

60000

50 0 0 0

40000

30000

20000

10000

0

0

Δ Enthalpy (kW)

Fig. 5 e Balanced grand composite curves for hydrogen production process for Reaction Scheme 1 generating 10,000 kPa(a) steam.

reforming, carbon can be formed under non-equilibrium conditions, where the reaction kinetics allow methane to decompose into carbon at a faster rate than it can react with steam [29]. The same could potentially happen with ethanol. Certainly, carbon deposition is also possible via the Boudouard and CO reduction reactions (Equations (6) and (7) respectively). Additionally, plant upsets may reduce the reformer feed gas temperature by 20  C, placing it in the range where carbon can be formed at equilibrium. The most important retarding reactants for avoiding carbon deposition are steam and hydrogen , the latter of which inhibits

cracking and olefin formation [29,30]. At the reformer inlet, hydrogen has not yet been formed, and equilibrium conditions have not yet been established. For Reaction Scheme 1, the inlet temperature is much higher than is normally the case with methane steam reformers, potentially leading to kinetically driven carbon deposition as the feed gas proceeds into the reformer tubes. In the absence of experimental evidence to the contrary, it can be concluded that this poses a significant potential risk to the viability of implementing Reaction Scheme 1. The reformer outlet temperature is reasonable by industrial standards, but is still only 50  C above the inlet temperature. Although this was assumed to be suitable, the temperature difference may not be enough to allow the reformer to be designed and operated satisfactorily, and is a matter which requires specialist advice from vendors. The outlet temperature could have been made higher, say 920  C, but this would have resulted in a slightly lower hydrogen yield. An economic analysis would assist the choice between higher temperature and the associated higher construction costs, or higher hydrogen yield and the greater associated revenue.

4.1.2.

Reaction Scheme 2

It was determined that the Reaction Scheme 2 reactions would be best carried out in a prereformer and reformer. The adiabatic prereformer outlet temperature of 370  C was set according to published experimental results [10]. Because the only components present in the reformer were those that are normally present in an industrial scale methane steam reformer, the inlet and outlet temperatures (650  C and 920  C, respectively) were set according to typical temperatures for a modern reformer. The outlet temperature and pressure (3000 kPa(a)) are reasonable when viewed from the

Ethanol & water Vaporiser Pump Prereformer

Exhaust Fan

To Stack

Reformer

Shift Reactor

Hydrogen

Fuel

Air

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Decanter

PSA

Fig. 6 e Basic hydrogen production process for Reaction Scheme 2 prior to heat integration.

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Table 2 e Operating conditions for producing hydrogen via Reaction Scheme 2 (Case 2). Other conditions are as per Table 1. Unit operation

Operating parameter

Feed pump

Vaporiser

Prereformer

Reformer

Value

Source

Feed temperature Outlet pressure

25  C 4190 kPa

Simulation

Outlet temperature

261  C

Simulation

Outlet pressure

4050 kPa

Simulation

Inlet temperature

441  C

Simulation

Outlet temperature Outlet pressure

370  C 3500 kPa

[10] Simulation

Pressure drop

400 kPa

Duty

0 kW

[7,17,27]

Inlet temperature

650  C

Outlet temperature

920  C

Thermodynamic equilibrium graphs, and industrial practise [6,7,17] Thermodynamic equilibrium graphs and industrial practise [6,7,17]

Table 3 e Case 1 utility requirements versus target utility requirements for producing hydrogen via Reaction Scheme 1. Utility

Target requirements Case 1 requirements (kW) (kW)

Total cold utility 10,000 kPa(a) steam Cooling water

31,762 (min)

31,763

24,106 (max)

23,849

7656 (min)

7914

Ambient temperature Depends on number of heat exchangers before Reformer Dew point of water-ethanol mixture plus 20  C superheat Depends on number of heat exchangers before Reformer Inlet temperature required to achieve 370  C outlet temperature Upper limit of experimental work Depends on number of heat exchangers before Reformer Assumed to be the same as for the Reformer Adiabatic operation is typical for hydrocarbon prereformers Avoids carbon deposition.

Avoids carbon deposition. Hydrogen yield is good.

perspective of the thermodynamic equilibrium calculations e hydrogen yield is high and carbon is not present. If end-user requirements permitted, operating the reformer at 2000 kPa(a) and 920  C would have the advantages of a 4% increase in equilibrium hydrogen yield and a 47% reduction in equilibrium methane yield. The comparative economic and environmental merits of operating at this lower pressure require further analysis.

4.2.

Integrated hydrogen production processes

4.2.1.

Catalyst

The major assumption underlying the simulations in this study is that effective, durable catalysts are available to promote the reactions in Reaction Schemes 1 and 2. Throughout its history, catalyst research for ethanol steam reforming has been, and continues to be, directed at low pressure fuel cell applications. Catalysts for large-scale ethanol reforming would have to cope with higher temperatures, higher pressures and greater mechanical stresses than fuel cell catalysts.

Balanced Grand Composite Curves 1000 Shifted Temperature (°C)

Comment

900

Grand Composite Curve

800

10000 kPa(a) Steam and Cooling Water

700

Table 4 e Cases 2 and 3 utility requirements versus target utility requirements for producing hydrogen via Reaction Scheme 2.

600 500 400 300

Utility

200 100

Target requirements (kW)

Case 2 requirements (kW)

Case 3 requirements (kW)

57,046 (min)

57,047

57,047

52,506 (max)

51,548

50,613

4540 (min)

5499

6434

Δ Enthalpy (kW)

Fig. 7 e Balanced grand composite curves for hydrogen production process for Reaction Scheme 2 generating 10,000 kPa(a) steam.

90000

80000

70000

60000

50000

40000

30000

20000

10000

0

0

Total cold utility 10,000 kPa(a) steam Cooling water

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Table 5 e Quantitative data for case study comparison. Parameter

Units

Reaction scheme Steam pressure

kPa(a)

Mole-based parameters Hydrogen production rate Ethanol feed rate Natural gas feed rate Steam generation rate Carbon dioxide emission rate Hydrogen/ethanol ratio Hydrogen/natural gas ratio Hydrogen/steam ratio Hydrogen/carbon dioxide ratio

kg-mol/h kg-mol/h kg-mol/h kg-mol/h kg-mol/h mol/mol mol/mol mol/mol mol/mol

Energy-based parameters Actual/theoretical steam production ratioa Steam/natural gas ratiob Cooling water/natural gas ratioc Overall thermal efficiencyd

% mol/mol mol/mol %

Reaction and process parameters Reformer conversione Reformer selectivityf Reformer yieldg Process yieldh Environmental factori

%

Case 1 RS1 10,000

4461.5 853.8 341.8 1423.4 2063.9 5.2 13.1 3.1 2.2

Case 2 RS2 10,000

4461.5 913.6 371.9 3076.5 2211.3 4.9 12.0 1.5 2.0

Case 3 RS2 10,000

4461.5 913.5 371.9 3020.8 2211.9 4.9 12.0 1.5 2.0

98.9 0.30 0.10 86.5

98.2 0.60 0.06 86.9

96.4 0.59 0.08 86.7

100 0.79 0.79 0.87 12.9

89.7 0.53 0.48 0.81 13.8

89.7 0.53 0.48 0.81 13.8

a Actual steam production in kW divided by maximum theoretical steam production in kW. Actual steam production ðkWÞ ; DH0rxn ¼ 828.304 kJ/mol natural gas, at 25  C, 101.325 kPa(a), and with product water in the b DH0rxn  natural gas molar flowrate ðkWÞ gaseous phase.The natural gas composition is 93.7 mol% methane, 3.3 mol% ethane, 1.44 mol% propane, with the balance being carbon dioxide (based on natural gas from Gladstone, Qld, listed in [38]). c As for Footnote b, but reporting actual cooling water usage in kW instead of actual steam production. d Overall thermal efficiency ¼ ðQH þ WS þ QS Þ  100 ð%Þ; Q ¼ energy available from the combustion of hydrogen (MW); W ¼ power available H S ðQNG þ QE Þ from of steam at 10,000 kPa(a) and 536  C, isentropically expanded to a low pressure saturated steam at 686.6 kPa(a) and 164.1  C (MW), QS ¼ energy available from saturated steam at 686.6 kPa(a) and 164.1  C, cooled to a saturated liquid at 101.325 kPa(a) and 100  C (MW), QNG ¼ energy available from the combustion of natural gas (MW), QE ¼ energy available from the combustion of ethanol (MW). Energy available from combustion is evaluated on the basis of the lower heating value at 101.325 kPa(a) and 25  C. Ethanol is assumed to be in the liquid phase. Reactant consumed in reformer ðmolÞ e Conversion ¼ Reformer reactant is ethanol for Reaction Scheme 1, and methane for Reaction Reactant fed to reformer ðmolÞ Scheme 2. Desired product produced ðmolÞ  Stoich: factor Desired product produced in the reformer is hydrogen. Reformer Reactant consumed in reformer ðmolÞ reactant is ethanol for Reaction Scheme 1, and methane for Reaction Scheme 2. Stoichiometric factor is 1/6 for Reaction Scheme 1, and 1/5 for Reaction Scheme 2. f Selectivity ¼

g Yield ¼

Desired product produced ðmolÞ  Stoich: factor Other comments are as per Footnote e. Reactant fed to reformer ðmolÞ

h Process yield ¼

Desired product produced ðmolÞ  Stoich: factor Desired product produced is hydrogen. Process reactant is ethanol. Reactant fed to process ðmolÞ

Stoichiometric factor is 1/6 for both reaction schemes. i Environmental factor ¼

Byproducts produced ðkgÞ By-products are decanted water and PSA off-gas. Desired product is hydrogen from Desired product produced ðkgÞ

PSA unit.

Noble metal catalysts including rhodium, ruthenium, palladium and platinum have been extensively investigated in the ethanol steam reforming literature. These metals are well known for their high catalytic activities. Nickel and cobalt, both non-noble metal catalysts, have also been carefully studied. Nickel, in particular, finds application as a catalyst in a number of industrial chemical processes. It has a relatively low cost and favours the rupture of carbonecarbon bonds [31]. Published

work on ethanol steam reforming indicates that the catalytic reaction pathways typically involve intermediate compounds such as methane, acetaldehyde, acetic acid, acetone or ethylene [9,31]. This means that, in practise, it is likely to be much more difficult to develop a catalyst for Reaction Scheme 1, in which none of these intermediates occur. Reaction Scheme 2 is based on the work of Galvita et al. [10]. In their experimental work they used a palladium

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catalyst to decompose ethanol into methane, carbon dioxide and hydrogen. These products were then reacted over a nickel methane steam reforming catalyst. It is not known whether the palladium catalyst they used could be developed for industrial scale use; however, the industrial nickel catalyst they used would definitely be suitable. The experimental studies of Galvita et al. employed an 8 mm diameter quartz tube fixed bed flow reactor [32], equivalent to a typical laboratory microreactor [6]. An approximate calculation indicates that the Reynold’s number inside this microreactor was about 6, based on the superficial gas velocity in the tube. This is consistent with typical values for microreactors [6]. Because large industrial steam reforming reactors operate at Reynold’s numbers of about 9500, the resistance to heat and mass transfer at the large scale is much lower than in a microreactor [6]. This provides a strong indication that any acetaldehyde formed by the dehydrogenation of ethanol, as reported by Galvita et al. [32] when operating at higher space velocities over a palladium catalyst, would be rapidly converted to methane and carbon monoxide in an industrial scale reactor.

4.2.2.

Steam-to-carbon ratio

All process simulations were performed with a water-toethanol ratio of 5.0 in the process feed stream, corresponding to a water-to-carbon ratio of 2.5. In Reaction Scheme 1, this same ratio is maintained at the entrance to the reformer. In Reaction Scheme 2, however, Equations (3) and (5) take place in the prereformer, so that at the entrance to the reformer the steam-to-carbon ratio is 2.0. The carbon in this case is present in methane, carbon monoxide and carbon dioxide. While this low ratio could superficially be expected to enhance the likelihood of carbon deposition on the catalyst, the presence of a substantial amount of hydrogen (24.6 mol% of the stream) makes this risk negligible. Furthermore, the presence of carbon dioxide inhibits the formation of elemental carbon from carbon monoxide via the Boudouard reaction (Equation (6)).

4.2.3. Comparison of Reaction Scheme 1 and Reaction Scheme 2 designs The data in Table 5, for most parameters, varies only according to the reaction scheme. The data for cases associated with Reaction Scheme 2 (Cases 2 and 3) is similar. The basis of the calculated data in Table 5 is a constant hydrogen production rate in each case study of 4461.5 kg-mol/h (100,000 Nm3/h). In Reaction Scheme 2, the feed gas exiting the prereformer and entering the reformer has a carbon dioxide concentration of 12.1 mol%, a hydrogen concentration of 24.6 mol %, and a methane concentration of 12.5 mol%. The balance is water, along with a trace of carbon monoxide. As can be seen from Equations (4) and (5), the significant presence of hydrogen and carbon dioxide e which are end products of the methane steam reforming and water-gas shift reactions e will inhibit the conversion of methane to carbon monoxide and hydrogen, according to Le Chatelier’s principle. The amount of unconverted methane leaving the reformer is 10.3% of the amount entering the reformer. This level of methane slip is higher than in an industrial methane reformer.

In Reaction Scheme 1 all the ethanol is converted, and the amount of hydrogen produced is determined according to the equilibrium of the water-gas shift reaction. In Reaction Scheme 2, because 10% of the intermediate component methane remains unconverted at the reformer exit, a higher quantity of ethanol (and water) needs to be fed to the process to generate an equivalent quantity of hydrogen. Reaction Scheme 2 requires 7% more ethanol than Reaction Scheme 1 to produce 100,000 Nm3/h of hydrogen. The unconverted methane in Reaction Scheme 2 is recycled to the reformer furnace with the other components of the PSA off-gas, where it is burned as fuel in conjunction with natural gas. The quantity of natural gas burned in the reformer furnace in Reaction Scheme 2 is 8.8% higher than for Reaction Scheme 1, to which is added the off-gas which has a higher energy content than in Reaction Scheme 1. Because of these factors, ratios such as hydrogen-toethanol, hydrogen-to-natural gas and hydrogen-to-carbon dioxide are more favourable for Reaction Scheme 1. The biggest difference, however, is with the hydrogen-to-steam ratio. This is much more favourable for Reaction Scheme 2, because Reaction Scheme 2 has more waste heat which is used to produce a greater quantity of steam. The energy-based parameters generally favour Reaction Scheme 2 cases, where steam generation is greater and cooling water usage is lower than for the corresponding Reaction Scheme 1 case. It was demonstrated via PRO/II simulations that the extra steam produced in the case studies associated with Reaction Scheme 2 is directly attributable to the higher energy-content PSA off-gas, as well as the additional natural gas, which are combusted in the reformer furnace in this Reaction Scheme. The additional natural gas contributes 26% of the energy used to generate the extra steam in Reaction Scheme 2, with the methane in the PSA off-gas supplying the balance. The overall thermal efficiency is nearly identical in each case study, at just over 86%. This parameter represents the energy available from the products (hydrogen and steam) as a percentage of the energy available from the feeds (ethanol and natural gas). Because excess energy over and above process requirements is recovered as steam at close to target levels in each case study, the efficiencies are similar. Energy losses occur mainly via the 150  C flue gas discharged to atmosphere; via the heat lost to cooling water; and via steam generation at slightly less than target levels. The efficiency of 86% for hydrogen production via ethanol steam reforming is consistent with the theoretical thermal efficiency for hydrogen production from methane steam reforming (83%); methanol production from the same route (84.2%); and ammonia production from the same route (89.2%) [33]. With all these processes, the thermal efficiency is lower in practise than the calculated theoretical value due to process inefficiencies such as heat losses, and steam- and power-generation inefficiencies. Typically, for hydrogen production from methane, the actual thermal efficiency lies between 70% and 80% [33,34]. Ethanol is completely consumed in the reformer in Reaction Scheme 1, but methane is not completely consumed in the reformer in Reaction Scheme 2. Therefore, reformer conversion, selectivity and yield, and process yield, are lower for the Reaction

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Scheme 2 cases. The environmental factor for the Reaction Scheme 1 case is more favourable, although it does not take into account energy recovery. In Case 1 there is less carbon dioxide emitted per mole of hydrogen produced than in Cases 2 or 3. However, more steam is generated in Cases 2 and 3 which can either be expanded in a turbine to generate electricity, and/or used as a heating source. In both instances, the extra steam replaces carbon dioxide emissions from other sources. More accurate assessment of the greenhouse gas impact of the case studies associated with each Reaction Scheme requires a life-cycle assessment approach, taking into account all the greenhouse gas emissions associated with:  the ethanol feedstock  the hydrogen production process  the power-generation processes which are displaced by the power generated from the high pressure steam produced by the hydrogen production process  the heating sources which are displaced by the steam produced by the hydrogen production process.

4.2.4.

Comparison of Cases 2 and 3 (Reaction Scheme 2)

The Case 2 design requires 23 heat exchanger matches. In contrast, Case 3 needs only 15 heat exchanger matches. Three process streams are split in Case 2 (giving a total of seven branches), and only two are split in Case 3 (yielding six branches). The highest temperature at which a process stream is split or recombined in Case 2 is 246  C; in Case 3 it is 260  C. The amount of steam generated in Case 2 is 1.8% greater than in Case 3. The boiler size in Case 3 will thus be marginally smaller. These considerations indicate a lower capital cost for the Case 3 process design. The slightly greater steam generation in Case 2 will result in higher revenue when credits for power or steam are taken into account.

4.2.5. Qualitative assessment of comparative costs for Reaction Schemes 1 and 2 It is expected that the capital cost of a hydrogen production process based on Reaction Scheme 1 will be lower than one based on Reaction Scheme 2, because:  the flow rate of process reactants is 6.5% lower, allowing unit operations and piping to be smaller  there is no prereformer  the reformer furnace flue gas and reformer exit gas temperatures are lower, potentially allowing cheaper materials of construction to be used in parts of the plant  steam generation is 54% lower, resulting in much smaller boilers or waste heat recovery (WHR) systems. For operating revenues and costs, the situation is not so clear-cut. In Reaction Scheme 2 the natural gas flow rate is 8.8% higher, and the ethanol and water flow rates are 7% higher. However, the steam generation is 116% higher than in Reaction Scheme 1. The relative costs of the process feeds in relation to the value afforded to steam will determine which scheme has the greater annual operating profit.

4.2.6.

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Best potential processing scheme

If in the future hydrogen production plants can be built and successfully operated based on the chemistry of Reaction Schemes 1 and 2, Reaction Scheme 1 is likely to be favoured because of its lower associated capital costs and lower feedstock costs. Notwithstanding this for the reasons already discussed, there are substantial technical hurdles to overcome in the development of a process based on Reaction Scheme 1. In contrast, Reaction Scheme 2 has been demonstrated at low pressure in the laboratory and, based on the current evidence, represents the most promising option for further research and development.

5.

Conclusions

5.1.

Hydrogen production processes

In comparison with Reaction Scheme 2, processes based on Reaction Scheme 1 have no prereformer, require slightly less feed leading to slightly smaller process equipment, have significantly smaller boilers or waste heat recovery systems, and potentially can use less expensive materials of construction in some parts of the plant. These processes will have lower capital costs than those based on Reaction Scheme 2. Operating costs may favour Reaction Scheme 2 processes, but this depends on the value which can be derived from the steam, which is generated in much larger quantities in processes based on Reaction Scheme 2. Despite the lower capital costs, bringing a design based on Reaction Scheme 1 through development into production will be problematic, for three reasons: 1) there is a low probability that a catalyst can be developed which does not generate intermediate compounds, which are absent in Reaction Scheme 1; 2) carbon deposition on the catalyst in the entrance region of the reformer is much more likely to be a problem than in Reaction Scheme 2; and 3) the small temperature difference across the reformer may prove to be impractical. These potential difficulties are not insurmountable, but given that Reaction Scheme 2 avoids the aforementioned problems, and has been demonstrated in the laboratory, it represents the best option for further development.

5.2.

Areas for future research and development

Potential opportunities exist to further improve the proposed processes, and to integrate them to a greater extent with a bioethanol production plant. Some areas which warrant further research and investigation are listed below.  Catalyst development is critical. A robust, durable, selective and highly active catalyst(s) for the chosen reaction scheme, which does not promote carbon deposition, is essential.  Improved energy efficiency may be possible by taking the reformer feed stream as vapour directly from the ethanol distillation columns in a bioethanol production plant. The required ethanol concentration is 16.7 mol%, much lower than the 86 mol% conventionally produced by distillation, saving energy. In addition, the vapour phase steam-ethanol

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mixture would not have to be condensed and then revaporised, potentially saving more energy. More extensive energy integration of the hydrogen production process with the bioethanol production process may be possible by utilising some of the steam generated by the hydrogen production process as a heat source for the bioethanol production process. There may be advantages in using an isothermal shift reactor in both reaction schemes, and an isothermal prereformer in Reaction Scheme 2, at an appropriate temperature to promote the equilibrium of the respective reactions. This type of reactor is more complex in construction and operation than an adiabatic reactor. Operating the reformer at 2000 kPa(a) for Reaction Scheme 2 would significantly reduce the amount of unconverted methane leaving the reformer. This may be advantageous in circumstances where lower pressure hydrogen product is acceptable. Using all the hot process streams to generate 10,000 kPa(a) steam, and then using the majority of this steam (maybe after expansion to a lower pressure) to heat all the cold process streams, may still leave a portion of the high-pressure steam available to produce electricity for export, particularly for processes based on Reaction Scheme 2. This would have the advantage of significantly lowering the temperatures at which heat exchangers operate, and would allow stable condensingevaporising heat exchange to occur between the steam and the water-ethanol feed mixture.

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