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Processes for the Synthesis of Liquid Fuels from CO2 and Marine Energy
0263–8762/06/$30.00+0.00 # 2006 Institution of Chemical Engineers Trans IChemE, Part A, September 2006 Chemical Engineering Research and Design, 84(A9): 828– 836
www.icheme.org/cherd doi: 10.1205/cherd.05204
PROCESSES FOR THE SYNTHESIS OF LIQUID FUELS FROM CO2 AND MARINE ENERGY D. MIGNARD and C. PRITCHARD University of Edinburgh, The Institute for Energy Systems, Edinburgh, UK
I
n Britain, wind and wave power are expected to make a significant contribution to future energy supplies, but the vast majority of the resource is located offshore, away from the mainland electricity grid. In this context, the chemical storage and transport of this energy appears to be an attractive option for the supply of fuels. On the other hand, it is not unlikely that CO2 will be recovered on a large scale in the next 10– 20 years, thus making it available as a carbon source for liquid fuels. A preliminary analysis indicated that the liquid fuels that could be most readily manufactured from hydrogen (from electrolysis) and recycled CO2 were: methanol, mixed alcohols and gasoline (via methanol). In the light of developmental work at NEDO, Japan, methanol appears to be simplest to manufacture. The Mobil Methanol-to-Gasoline process also facilitates the manufacture of a very convenient automotive fuel. Mixed alcohols have the advantages both of low toxicity, and efficient Fischer –Tropsch catalysts for their synthesis (the Pearson process and the Ecalene process). This paper compares the energy efficiency of these three processes. Keywords: synthetic fuels; wind power; marine energies; hydrogen; CO2 sequestration.
INTRODUCTION
located North and West of Scotland, while the electricity transmission grid is poorly developed in the nearby onshore areas. Since a high proportion of renewable energy is expected to come from these remote locations on islands or at sea, the production of liquid fuels could be more convenient than a subsea network of electricity cables. A first possibility is to use hydrogen as an energy vector. While the electrolysis of water is an established technology for producing hydrogen fuel from electricity, the storage and transmission of H2 on a large scale and over long distance requires a practical solution. Possible solutions include the construction of hydrogen pipelines (Air Liquide, 2003); the liquefaction and shipping of hydrogen (Giacomazzi and Gretz, 1993); the use of the methylcyclohexane–toluene– hydrogen system (Newson et al., 1998); and the injection of hydrogen into the existing natural gas network (Haines and Polman, 2004). Appliances designed to use hydrogen must also be available: gas turbines using hydrogen are being developed with higher efficiencies (Shilling and Jones, 2003); and hydrogen fuel cells for stationary applications should become more economical in the future with progress in stack assembly techniques, materials and efficiency [as exemplified by companies such as ITN Power Ltd (2005) or Kansai Electric (2004)]. We have examined the various hydrogen options, and also alternative processes for the manufacture of carbon-based liquid fuels, since these liquid fuels are readily transportable and usable. The description of these processes for carbon-based fuels is the object of this present paper.
This paper presents processes for the production of liquid fuels from recycled CO2 and hydrogen from electrolysis, at locations remote from the electricity grid where wind or marine energies are available. The availability of CO2 feedstock and the issues surrounding the large-scale storage and transmission of hydrogen in such contexts are first discussed. Fuel synthesis from CO2 and hydrogen is then justified as a potential solution, and the available technologies are reviewed for three processes: methanol, mixed alcohols and gasoline synthesis. Finally, Process Flow Diagrams and energy balances are presented for these processes. Chemical Storage of Renewable Energy in the Context of Large Scale, Remote Production The long-term necessity for the world economy is to achieve a sustainable pattern of energy use. In Britain, a significant contribution is expected from resources such as wind and wave power. The UK Atlas of Renewable Energy (ABPmer et al., 2004) shows that the most valuable wave and wind power resource in terms of power density is Correspondence to: Dr D. Mignard, University of Edinburgh, The Institute for Energy Systems, The King’s Buildings, Mayfield Road, Edinburgh EH9 3JL, UK. E-mail: [email protected]
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SYNTHESIS OF LIQUID FUELS FROM CO2 AND MARINE ENERGY Recycling of CO2 as Hydrogen Carrier Fuel A possible route for the synthesis of a liquid fuel consists in reacting (recycled) CO2 with hydrogen obtained from the electrolysis of water. The resulting fuel could be readily used in existing distribution networks, power generators, vehicles or other appliances, and it would facilitate both a substantial penetration of renewable power and also large cuts in CO2 emissions, on a relatively short timescale. If emissions of CO2 are to be curbed, it is not unlikely that this gas will be recovered on a large scale in the next 10–20 years. Should the price of oil remain at its currently high level, CO2 could be profitably used for Enhanced Oil Recovery (EOR) (Berger et al., 2004). The ‘CO2 Infrastructure for EOR in the North Sea’ (CENS) project, which involves the UK, Denmark and Norway (government bodies and companies), is currently evaluating the practicalities of establishing a CO2 infrastructure for sequestration and EOR in and around the North Sea, including recovery of the CO2 from power stations, pipeline transport or shipping, and injection into mature oil reservoirs. 29 Mt y21 of CO2 would come from coal-fired power stations in the UK and Denmark (Markussen et al., 2003).
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(10 000–30 000 m3) should conditions allow for profitable EOR in the North Sea (Berger et al., 2004). However, lower pressures have been recommended for such ships, typically 6.5 bar at 2508C (Aspelund et al., 2004). The power consumption required for CO2 liquefaction in these conditions is 110 kWh t21 liquid CO2 (Aspelund et al., 2004; starting at 1.1 atm, 158C, with seawater available at 108C). In this study, liquefied CO2 at 6.5 bar and 2528C was thermally compressed to produce gaseous CO2 at a pressure of 30 bar, using seawater as the heat source. A ship of hybrid design, for transporting either liquefied CO2 or liquid fuels, would be worth investigating. This would allow the ship to transport the CO2 to the remote area for fuel synthesis, and shipping the synthetic fuel back from the remote location. Such a design was proposed by the Japanese NEDO and RITE study, which considered the production of methanol from solar power and CO2 in areas such as Egypt or Australia (NEDO and RITE, 1998). Finally, the amount of CO2 emitted during transport has been evaluated at about 2% of the CO2 transported per 1000 km, and the boil-off losses would probably be negligible (Ozaki et al., 2004). However, in this study we specifically exclude the (very site-specific) CO2 emissions associated with these shipments.
CO2 Recovery By using high-pressure, superheated steam at the reboiler of the CO2 amine-absorption stripper, and by the exclusive use of existing ‘ultra clean’ flue-gas at 14 –16% CO2 from a coal-fired CHP (combined heat and power) plant, the energy and monetary costs of CO2 recovery may be kept low. In particular, the CHP configuration ensures that the low-grade heat is a useful product, which in turn leads to a higher overall efficiency than that from a power-only plant. For the coal-fired CHP plant, the CENS project claims a cost of $35 t21 for CO2 recovered from flue-gas (Markussen et al., 2003), which represents ca. $33 t21 CO2 after subtracting the CAPEX of the 1500 km pipeline. Energy consumption may be reduced further by adopting the flue-gas scrubbing technology and the KS-1 solvent that were developed by Mitsubishi Heavy Industries, Ltd (MHI), and Kansai Electric Company (KEPCO) (Iijima and Kamijo, 2003). A 160 t day21 CO2 plant in Malaysia has been using the KS-1 solvent since 1999. Solvent consumption there is 0.35 kg t of CO2 recovered, with solvent stability over at least 5700 h; LP steam consumption is 1.5 t per t CO2 recovered, which compares with 2.7 t t21 in the Kerr-McGee/ABB Lummus Global process (Hydrocarbon Processing, 2000). The Kerr-McGee process requires 100 kWh of electrical power per tonne of CO2 before compression. This latter figure, together with that of 860 kWh t21 CO2 for steam requirements from the MHI process, is used in our calculations. CO2 Shipping CO2 transport by ship is done in Europe, where Yara International transports liquefied CO2 from ammonia plants in Porsgrunn (Norway), Sluiskil (The Netherlands) and Ferrara (Italy) (Yara, 2005). For example, the m.v. ‘Coral Carbonic’ stores 1382 t CO2 at 2408C, or 1292 t CO2 at 222.58C, at pressures between 14–20 bara (Anthony Veder, 2005). Bigger ships are planned for construction
Electrolysis Using Variable and Intermittent Current Recent developments have made available on the market alkaline electrolysers that can operate in intermittent or variable conditions at 25 or 30 bar (Stuart Energy, 2004; European Commission, 2003). Appropriate industrial electrodes do not need a protective voltage in these conditions (EUHYFIS, 2004). Advanced designs for alkaline pressure electrolysers include the IMETw technology, which is commercialized by Stuart Energy for capacities up to 1 MW (Stuart Energy, 2004; Vandenborre, 2002), and the Pressure Module Electrolyser (PME) technology, which is being developed by a consortium including MTU-Friedrichshafen, Norsk Hydro, and PMT (Prime Membrane Technology, Belgium) (European Commission, 2003). The PME technology is at the demonstration stage (European Commission, 2003). Neither technology requires lye pumps, and both use separate catholyte and anolyte streams to prevent any mixing of O2 and H2, which is an advantage for operation at low load. The modular construction of the PME pressurized shell allows the electrolyser to be scalable beyond several MW (Kliem, 1995). Fuel Synthesis The following three options were considered, and the corresponding process flow diagrams were devised. Process A. Methanol synthesis (Figure 1) Methanol synthesis from pure CO2 is feasible on existing Cu/ZnO/Al2O3 catalysts used for making methanol from synthesis gas (Sahibzada et al., 1998), according to the reaction CO2 þ 3H2 ! CH3 OH þ H2 O(g) DH298K, 1bar ¼ 49:16 kJ mol1
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(1)
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Figure 1. Flowsheet for the CO2 to methanol process. Key: O ¼ pressure in bar; S ¼ temperature in 8C; Cp ¼ compressor; P ¼ pump; T ¼ tank; R ¼ reactor; HX ¼ heat exchanger; FH ¼ fired heater; Cd ¼ condenser; KB ¼ kettle boiler; DC ¼ distillation column; CW ¼ cooling water.
The observed equilibrium yield when using CO2 þ 3H2 at a pressure of 50 bar was ca. 22% for methanol, due to the inhibitory effect of the water product. Deactivation of the catalyst by water is known to be a problem, but better catalysts have been developed. For instance, Cu/ZnO/ ZrO2/Al2O3/Ga2O3 showed excellent longevity, selectivity and activity: After 2500 h, the Space –Time– Yield of this catalyst was stable at 600 g MeOH/(litre of catalyst) h21 t , 46% better than for a commercial Cu/ZnO/Al2O3 (ICI 51-2 catalyst) (NEDO and RITE, 1998). Mignard et al. (2003) modelled the adiabatic operation of a fixed bed reactor with ICI 51-2. For a minimum 99% yield, the minimum compression requirements (including recycle) would occur at 2278C, 50 bar, with a recycle ratio of 7.9. Optimum conditions with the NEDO and RITE catalyst could be somewhat different. Process B. Mixed alcohols synthesis (Figure 2) Higher alcohols have the advantage that they are less toxic than methanol, and could be handled more safely by the general public. For ethanol, the reactions are CO2 þ H2 ! CO þ H2 O(g)
(2)
DH298K, 1bar ¼ 41:21 kJ mol1 of carbon 2CO þ 4H2 ! CH3 CH2 OH(g) þ H2 O(g) DH298K, 1bar ¼ 256:1 kJ mol1 of ethanol
(3)
Until recently, reaction (3) was not possible with good yields or good selectivity. However, Pearson Technologies Inc. (PTI) (Pearson, 2001, 2003) now claims the invention of a Fischer – Tropsch catalyst capable of converting synthesis gas to ethanol with a yield of 99þ% to ethanol after recycle. The single pass conversion is 15– 60%, depending on the conditions. An important advantage of the Pearson process is that it avoids the formation of significant quantities of byproducts such as CO2 and methane (Pearson, 2001). Carefully controlled cooling is necessary to prevent the reactor from ‘running away’. It is claimed that the temperature and pressure of the process can be lowered to values that are typical of methanol processes, including the ‘low pressure’ processes (i.e., down to 50 bar) (Pearson, 2001). Details of the composition of a synthesis gas feedstock were 51.1% H2, 23.7% CO, 17.1% CO2, and 6.3% CH4 (Vantine, 2004). This indicates that a ratio CO2/(CO2 þ CO) of 42% is acceptable, and that the catalyst must either have a strong reverse water – gas shift (RWGS) activity, or else utilises CO2 directly. PTI has built and operated a 50 t day21 pilot plant processing wood waste to produce ethanol in Aberdeen, North Mississippi (Pearson, 2001). In addition to ethanol, the process is also expected to form methanol, and higher alcohols with three atoms of carbon or more. Pearson (2001) stated that these by-products are separated downstream in the condensate, and recycled back to the process. An equilibrium partial pressure could apply to methanol, as homologation
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SYNTHESIS OF LIQUID FUELS FROM CO2 AND MARINE ENERGY
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Figure 2. Flowsheet for the simplified CO2 to mixed alcohols process, with 15% conversion to ethanol per pass in the alcohol reactor (R2). Separation and recycling in R2 of alcohols other than ethanol was neglected. Key as in Figure 1.
reactions with CO and H2 or direct synthesis from CO2 are possible. However, further polymerization of all alcohols with Fischer –Tropsch reactions would steadily increase the length of carbon chains, and therefore higher alcohols with three atoms of carbon and more would presumably be recycled to the gasifier. For methanol, CO2 and CH4, recycling should eventually attain an equilibrium at the inlet and inhibit further byproduct formation (Mawson et al., 1993). Therefore, we would expect that alcohol formation and water – gas shift are the only significant reactions that affect the composition of the outlet from the reactor. In the absence of published experimental data on the Pearson process, we assumed in this work that the final product was a mixture of alcohols rather than ethanol. However, we have treated the distillation step as though ethanol were the only product, since its separation is the most energy intensive. Finally, we note that a correct choice of residence time in the reactor will be critical: longer residence times will ensure a higher conversion per pass, while shorter residence times will avoid the excessive formation of heavier alcohols with five and more carbon atoms. This is borne out from experience (Bhasin, 1980). Other processes that are worth mentioning for the formation of mixed alcohols from synthesis gas are the Dow process (Conway and Stevens, 1988) and the Ecalene process (Jackson and Mahajan, 2001). One potential difficulty with these processes is the generation of substantial amounts of byproducts: For example, with the catalysts
described by Conway and Stevens (1988), CO conversion to CO2 is typically around 30– 40%; and of the remaining gas phase products, at least 10% is methane and other hydrocarbons. The attraction of the Pearson process comes from its ability to handle CO2 and its claim that it is essentially emission free, i.e., all of the carbon can be utilized without significant CO2 or methane byproduct. Another difficulty with the Ecalenew process seems to be the need to feed sulphur in order to maintain the catalyst activity (the Ecalenew fuel may need subsequent desulphurization). Tentatively, we chose to model a mixed alcohol reactor on the basis of what information we could find on the Pearson process. A RWGS reactor may be needed if the feedstock is pure CO2. A Cu/ZnO/ZrO2/Ga2O3 catalyst for RWGS was proposed (Joo et al., 1999). Using the table provided by these authors, it was found that operating the RWGS reactor at 6008C, 30 bar and no recycle would permit a ratio CO2/ (CO2 þ CO) of 39%. A drop in temperature to 4008C would increase this ratio to 61%. If the catalyst of the alcohol synthesis reactor required higher conversion, a temperature of 8008C would enable a ratio CO2/(CO2 þ CO) of 25%. Process C. Methanol to gasoline (MTG) (Figure 3) Methanol reacts on the ZSM-5 zeolite catalyst to produce dimethyl ether, which further reacts to give hydrocarbons with up to 10 carbon atoms, together with water (Keil, 1999).
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Figure 3. Flowsheet for the methanol to gasoline process. The C3 and C4 separation processes and the alkylation unit are not represented here, and neither is the light gas processing unit. The MTG fluid-bed reactor ‘R2’ is simplified, i.e., the regeneration bed is not shown. Key as in Figure 1.
The Mobil fluid bed process demonstrated in Wesseling, Germany, produced 15.9 m3 day21 of gasoline using a feed of raw methanol having 27% mol. water and 73% mol. methanol. The heat evolved was 1.74 MJ kg21 methanol, recovered through heating oil tubes immersed within the bed (Keil, 1999). When operating at 4138C and 2.75 bar, 99þ% methanol was converted to 60% gasoline, 34.4% of a liquefied gas fraction, and 5.6% light gas. By alkylation of the isobutane with the propene and butenes fraction, a gasoline yield of 88% could be obtained.
METHODS AND MATERIALS Process Design After mass and energy balances were performed and flowsheets drawn, the processes were compared according to their overall net energy requirements. The feed was taken to be CO2 þ 3H2 at 30 bar, 258C. The fact that CO2 and H2 were in stoichiometric proportions with respect to equation (1) allowed a steady state to be reached with respect to the flow rate of recycle gases, with a purge of 1% of the total flowrate taken out of the system to minimize the accumulation of inert compounds or reaction by-products (typically, in the case of methanol synthesis,
about 0.4% of by-product would be expected, mostly in the form of methyl formate; NEDO and RITE, 1998). The electrical and thermal energy needed for CO2 extraction was calculated according to the figures given in the next section, and included in the assessment. For purification by distillation, reboiler duties were estimated from typical values for each system, or by estimating reflux ratios. Compression requirements were calculated taking into account pressure drops through major pieces of equipment (exchangers, condensers and reactors). Exchangers and condensers were of the tubular type, and their design followed the TEMA guidelines (Perry et al., 1999). Spreadsheet calculations were used for all these estimates, as well as an in-house C-program for the methanol reactor of process A. Methanol product was fuel grade at 98% wt (the balance is water) while ethanol was also fuel grade at 99.4% wt. In process B, a RWGS reactor was operated at 30 bar, 6008C inlet and 4458C outlet, CO2 conversion 53% and no recycle. The feed was first preheated by the RWGS reactor outlet, and then heated further by the combustion of the purge stream from the ethanol reactor recycle. Another option would be to provide electrical heating using renewable power. The ethanol reactor was operated at 50 bar, 4008C inlet and 4308C outlet, with two possible levels of CO
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SYNTHESIS OF LIQUID FUELS FROM CO2 AND MARINE ENERGY conversion, 15% or 30%, and the RWGS reaction within the reactor was assumed to maintain the CO2/CO ratio constant. This approach implicitly assumed that CO rather than CO2 was the reagent directly implicated in the synthesis reaction. A purge of 1% on the recycled gas was assumed, in order to prevent the build-up of inerts. Due to the dearth of data on the Pearson process in the open literature, these conditions were working assumptions to assess ethanol synthesis via RWGS using a spreadsheet model. Both the RWGS and the mixed alcohol reactor were of the fixed catalyst bed type. The heat of reaction from the alcohol reactor was used to generate steam. Alcohol separation was carried out at atmospheric pressure. In order to evaluate a worstcase scenario, ethanol was assumed to be the only product, and the distillation produced a 90% azeotrope, which was sent to a drying unit. In process C, the gaseous product from process A was partly condensed to yield a 40% mol. methanol feed for distillation, and the remainder of the gas was separately condensed to yield a 73% methanol feed for the MTG reactor. Instead of distillation, stripping was used to upgrade the 40% product to 73%, which saved up to two thirds of the reboiler duty. A debutanizer column was also required to separate the light gases from the raw MTG product and the liquefied C3 – C4 gases. In our design work, we opted for selective adsorption to separate the reagents involved in the alkylation process from the liquefied gas. Adsorption was preferred over distillation since it generally uses less energy. The gas consisted of 14.5% isobutane, 5% propene, 7.3% butanes, 5.9% propane, and 1.7% n-butane. Propene and butenes were first isolated using the UOP OlexTM process (Pujado, 1986), which selectively adsorbs olefins. In the remaining stream, isobutane was the only remaining branched hydrocarbon, and it was separated by the UOP MolexTM process through selective adsorption of the other (linear) hydrocarbons on 5A molecular sieve. Power and steam requirements for the OlexTM and
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MolexTM processes were taken from Pujado (1986). Finally, for the alkylation reaction we opted for the AlkyCleanTM process developed by Akzo-Nobel, ABB Lummus Global and Fortum. This process relies on reacting the alkylation feedstock on a solid acid catalyst (SAC) at around 908C. This approach avoids the use of strong acids (sulphuric or hydrofluoric acids) and of the extensive effluent processing that are necessary with traditional processes (van Broekhoven, 2001). A demonstration plant has been running in Porvoo, Finland, since 2002, and the process is now reaching commercial status (ABB, 2004). In the absence of any information on energy requirements for this process, our study assumed that these were comparable to those of the H2SO4 process. This assumption could be an overestimate, since the solid acid catalyst approach has costs comparable to that of the H2SO4 process, but has no need for extensive effluent treatment (ABB, 2004). Energy Efficiencies The energy efficiency was calculated as the ratio of the chemical energy stored as Lower Heating Value (Qstored) to the electrical energy used (Einput). For process A and B, there was a net requirement for steam (Qnet ¼ Qp 2 Qgen, where Qp is the required steam heat and Qgen is the heat available from the reactions). This net demand could be met either by burning some of the synthetic fuel, or by using electrical heating (efficiency 0.94). In the case of process C, which was generating enough steam to sustain itself, Qnet ¼ 0. However, for all processes, the steam required for the extraction of CO2 (QCO2), was considered to come from the power plant that generated the CO2 in the first place, and it was decoupled from the fuel production process. Therefore, the corresponding figure was not included in Qnet, but instead it was treated as a power input to the process, in the denominator of the ratio. For this purpose, it was considered that 0.25 kWhe had to be
Table 1. Breakdown of energy balance for processes A, B and C. Energy balance basis: 1 kmol of (CO2 þ 3H2) feed
Process A Methanol
Process B Mixed alcoholsa
Process C Gasoline
Power required for electrolysis, Eel kWhe kmol21 Pumping Compression and pumping CO2 extraction (CO2 at 1 bar) Total ¼ Ep kWhe/kmol21 CO2 liquefaction Compressing fresh feed Feed to ethanol reactor Gas recycle (synthesis reactor) OlexTM MolexTM AlkycleanTM, incl. steam Ethanol drying
80.64 0.032 1.1 1.21 0.557 — 1.243 — — — —
80.64 0.032 1.1 1.21 — 0.717 0.679 — — — figure included in steam requirement
80.64 0.192 1.1 1.21 0.557 — 1.243 0.008 0.017 0.114 —
Steam required for CO2 extraction, QCO2 kWhth kmol21 Methanol or ethanol separation Steam for process Debutanizer Total ¼ Qp 21 kWhth kmol OlexTM MolexTM From stream Steam generated, From synthesis reactor Total ¼ Qgen kWht kmol21 Chemical energy stored (LHV basis), Qstored kWh kmol21
9.78 7.96 — — — 3.56 0 44.29
9.78 11.37 — — — 0 6.02 44.24
9.78 3.30 0.1 0.087 0.281 3.56 3.81 36.13b
a
Assuming 15% conversion per pass for the alcohol reactor; ethanol is the sole alcohol. Excluding LPG, fuel gas and extra steam.
b
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MIGNARD and PRITCHARD Table 2. Energy efficiencies for processes A, B and C [calculated from Table 1 and equations (1) and (2)], and compressed hydrogen. A, Methanol H2 Compressed, 30 to 400 bar
Process S, Chemical Energy stored, kWh/ kWhe
0.61
Fired heating
B, Mixed alcohols
Electrical heating
0.50
Fired heating
0.515
0.45 if yield per pass is 15% 0.47 if yield per pass is 30%
Electrical heating (except RWGS feed) 0.46 if yield per pass is 15% 0.47 if yield per pass is 30%
C, Gasoline Heating using by-product steam 0.42 (gasoline only) 0.46 (including LPG) 0.49 (including LPG and extra steam)
sacrificed to produce 1 kWhth of LP steam for the separation of CO2 using amine technology (Bolland and Undrum, 1998). The other terms in the denominator were the power input from electrolysis (Eel), and that from compression and pumping (Ep), which included direct power consumption for CO2 extraction. Electrical heating (EH ¼ Qnet/0.94) was also added there if considered. For processes A, B and C, the chemical energy stored was S ¼ ðQstored Qnet Þ=(Eel þ Ep þ 0:25 QCO2 )
(1)
If using electrical heating for processes A and B: S ¼ Qstored =(Eel þ Ep þ Qnet =0:94 þ 0:25 QCO2 )
(2)
RESULTS AND ANALYSES Flowsheets Figures 1– 3 at the end of the paper present the flowsheets for process A, B and C respectively. Energy Storage Table 1 presents the calculated energy requirements and byproducts from processes A, B and C. Table 2 presents the storage efficiency S as defined previously. DISCUSSION Figures 1 – 3 demonstrate the relative simplicity of the methanol process compared with the others. The gasoline process is the most complex one, and most of the refinery plant has not been shown in Figure 3. Table 2 reflects the fact that compressed hydrogen would be the most desirable fuel in terms of the efficiency of chemical storage, when its storage and transmission is not an issue. This is a simple consequence of deriving the other fuels from it. In the cases where hydrogen storage or transmission is not possible or desirable, the methanol process is the most efficient for storing chemical energy. Efficiencies were 0.50 kWh/kWhe when meeting the net steam requirements with fired boilers, and 0.515 kWh/kWhe with electrical boilers. Electrical heating combined with heat storage would also allow more flexible operation in variable or intermittent operation. The mixed alcohol process (B) seemed to achieve efficiencies that were about 0.03– 0.05 kWh/kWhe lower
than those for the methanol process. When assuming 15% conversion per pass for the ethanol reactor, the figures were 0.43 kWh/kWhe with fired heating and 0.47 kWh/ kWhe with electrical heating. This was due to higher steam requirements for distillation and drying, together with the requirement to heat the feed stream to 6008C and the higher pressure-losses along heat exchangers and condensers. However, these figures could be on the pessimistic side, since conversion rate per pass might be as high as 60% (Pearson, 2001). Table 2 shows that a doubling of this parameter to 30% allows a gain of 0.01– 0.02 kWh/kWhe. The other large source of efficiency loss for process B was the distillation step: we found that decreasing the energy consumption of this step to a level comparable to that for methanol in process A would raise the amount of stored chemical energy by 0.02 kWh/kWhe. A possible way to achieve an improvement of this order or even greater might be to separate water from the diluted product in the gas phase, e.g., using membranes. Another possibility is to effect a more complete conversion of the CO2 to CO in the RWGS reactor, including a recycle, followed by condensation of more water at that stage. Unknowns in the parameters of the Pearson process imply that these figures are only indications of what could be achievable: . An improvement to avoid the RWGS step might be to formulate a mixed catalyst that combined ethanol synthesis and RWGS catalyst, or to design an integrated reactor for the two reactions. (The high temperatures that we considered are thermodynamically favourable to the RWGS reaction, and kinetically favourable to the formation of alcohols.) . It is not apparent to the authors whether the catalyst would be able to utilise pure CO2 in the absence of CO. . Another unknown is the amount of alcohols, other than ethanol, that form in the reactor. If significant amounts of byproducts are formed and a pure product is desired, the separation process might require more energy downstream. . Finally, it is known that control of the the H2: CO ratio is critical in order to maximize selectivity towards alcohols (Landis et al., 2005). Given that pure hydrogen is available in our process, it should be possible to inject it at appropriate positions along the reactor in order to maximize the yield of alcohols per pass. The gasoline process (C) generated excess steam, which would be adequate to meet all process heating requirements
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SYNTHESIS OF LIQUID FUELS FROM CO2 AND MARINE ENERGY once integration of streams had been carried out. When considering gasoline alone, efficiencies were lower than for process A and B (e.g., 16% lower than for methanol), but the figure reported in Table 2 shows that the inclusion of LPG product and of excess steam (which could be utilized elsewhere or combined with heat storage) could practically close this gap. The good overall performance of process C was explained by the avoidance of most of the reboiler duty for separating alcohols from water that was encountered in processes A and B. How significant are these figures when comparing the different processes? Some of the uncertainty in our knowledge of process conditions had little influence on the final figures, e.g., the AlkycleanTM process decreased the efficiency of storage by only 0.002 kWh/kWhe. By contrast the heat evolved by the synthesis reaction and the heat needed for separation of the products (where needed), were dominant. A less obvious effect was that the pressure drops in tubular heat exchangers and condensers had a pronounced impact on overall performance, and had to be maintained to values within a few percent of the working pressure. We expect that the figures in Table 2 are sufficiently accurate to allow comparison, although it is acknowledged that, for the ethanol process, further data would be needed and further improvements could be made. The operability of these schemes in intermittent or variable conditions is now being investigated. CONCLUSIONS Three schemes for the synthesis of liquid fuels from recycled CO2 and renewable power were presented. The methanol process had the highest efficiency, with 0.50 kWh of chemical energy stored for each kWhe used. The use of electrical boilers, instead of fired ones utilizing part of the fuel product, would raise this figure to 0.51 kWh/kWhe. The mixed alcohol process gave lower results when lower-range values were assumed for the conversion-per-pass. However, its potentially high conversion per pass (60% instead of 22% for the methanol process) could make it worthy of consideration. There also might be the possibility of combining the RWGS and the alcohol synthesis in one single reactor, as well as feeding fresh hydrogen at various locations along the reactor in order to adjust the H2:CO ratio. The gasoline process benefited from an integrated design of the methanol and hydrocarbon plants, and avoided the energy-intensive distillation step required by the other processes. It reached efficiencies of 0.42 kWh/kWhe without accounting for excess steam production and LPG product, and 0.49 kWh/kWhe when these other products were considered. These processes are now being evaluated for operation under variable power supply. REFERENCES ABB website, accessed 11/2004, Making the grade safely, http:// www.abb.com/global. ABPmer, The Met Office, Garrad Hassan, and Proudman Oceanography Laboratory, 2004, UK Atlas of Offshore Renewable Energy, Atlas Pages Technical Report, Report for the Department of Trade and Industry (DTI), UK. Air Liquide, 2003, Hydrogen Delivery Technologies and Systems— Pipeline Transmission of Hydrogen. Presented at the Strategic
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Initiatives for Hydrogen Delivery Workshop at the U.S Department of Energy, 7 –8 May, Washington DC, http://www.eere.energy.gov/ hydrogenandfuelcells/pdfs/hydrogen_pipelines_ pres_air_liquid.pdf. Anthony Veder, accessed 03/2005, http://www.anthonyveder.nl. Aspelund, A., Sandvik, T.E., Krogstad, H. and De Koeijer, G., 2004, Liquefaction of captured CO2 for ship-based transport. Proceedings of the 7th International Conference on Greenhouse Gas Control Technologies (Volume II, Part 2), Vancouver, 5–9 September 2004, in Rubin, E.S., Keith, D.W. and Gilboy, C.F. (eds). 2545–2549 (Elsevier, Amsterdam, The Netherlands). Bhasin, M.M., 1980, Process for producing ethanol from synthesis gas, patent US 4,235,801. Berger, B., Kaarstad, O. and Haugen, H.A., 2004, Creating a largescale CO2 infrastructure for enhanced oil recovery, Greenhouse Gas Control Technologies 7, Vancouver, 5 –9 September, available on http://uregina.ca/ghgt7/PDF/papers/peer/108.pdf. Bolland, O. and Undrum, H., 1998, Removal of CO2 from Gas Turbine Power Plants—Evaluation of Pre- and Post-Combustion Methods, paper presented at Greenhouse Gas Control Technologies 4, Switzerland, http://www.energy.sintef.no/publ/xergi/98/3/art-8-engelsk.htm. Conway, M. and Stevens, R., 1988, Mixed alcohols production from syngas, US patent 4,752,623. EUHYFIS website, http://www.euhyfis.com/indexgb2.html?gb/konzept. htmlmainFrame. Accessed June 2004. European Commission, 2003, EUR 20718—European and Fuel Cell Projects 1999–2002, 90 (Office for Official Publications of the European Communities Luxembourg). Giacomazzi, G. and Gretz, J., 1993, Euro-Quebec Hydro-Hydrogen Project (EQHHPP): A challenge to cryogenic technology, Cryogenics, 33(8): 767–771 Haines, M.R. and Polman, E.A., 2004, Reduction of CO2 emissions by addition of hydrogen to natural gas, Greenhouse Gas Control Technologies 7, Vancouver, 5– 9 September, available on http://uregina.ca/ ghgt7/PDF/papers/peer/074.pdf. Hydrocarbon Processing, April 2000, CO2 recovery (editorial). Hydrocarbon Processing, 63. Iijima, M. and Kamijo, T., 2003, Flue gas CO2 recovery and compression cost study for CO2 enhanced oil recovery, GHGT6, Kyoto, I: 109– 114. ITN Power Ltd., accessed 03/2005, http://www.itm-power.com/. Jackson, G.R. and Mahajan, D., 2001. Method for production of Mixed alcohols from synthesis gas, US patent US 6,248,796 B1. Joo, O.-S., Jung, K.D., Moon, I., Rozovskii, A.Y., Lin, G.I., Han, S.H. and Uhm, S.J., 1999, Carbon dioxide hydrogenation to form methanol via a reverse-water-gas-shift reaction (the CAMERE process), Ind Eng Chem Res, 38: 1808–1812. Kansai Electric, 2004, Prototype Power Generation System of Intermediate Temperature SOFC, http://www.kepco.co.jp/english/rd/ topics/topics_1.html. Keil, F.J., 1999, Methanol-to-hydrocarbons: process technology, Microporous & Mesoporous Materials, 29: 49–66. Kliem, E., 1995, Dispositif d’e´lectrolyse (sous pression) en structure modulaire, Patent FR2710076. Landis, S.R., Espinoza, R.L., Mohedas, S.R., Belt, B.A., Melquist, V.H. and Goodwin, R.T., 2005, Managing hydrogen and carbon monoxide in a gas to liquid plant to control the H2/CO ratio in the Fischer– Tropsch reactor feed, US Patent 6,872,753 B2. Markussen, P., Austell, J.M. and Hustad, C.-W., 2003, A CO2 infrastructure for EOR in the North Sea (CENS): macroeconomic implications for host countries, in Gale, J. and Kaya, Y. (eds). Greenhouse Gas Control Technologies, GHGT6, Kyoto, Vol. II: 1077–1082 (Elsevier, Amsterdam, The Netherlands). Mawson, S., McCutchen, M.S., Lim, P.K. and Roberts, G.W., 1993, Thermodynamics of higher alcohol synthesis, Energy and Fuels, 7(2): 257–267. Mignard, D., Sahibzada, M., Duthie, J. and Whittington, H.W., 2003, Methanol synthesis from flue-gas CO2 and renewable electricity: a feasibility study, International Journal of Hydrogen Energy, 28: 455–464. NEDO and RITE, March 1998, Project of CO2 fixation and utilisation using catalytic hydrogenation reaction, ,Current State of Research and Development., Report issued by the New Energy and Industrial Technology Development Organisation (NEDO), and the Research Institute of Innovative Technology for the Earth (RITE), Japan; obtained from NEDO. Newson, E., Haueter, Th., Hottinger, P., von Roth, F., Scherer, G.W.H. and Schucan, Th.H., 1998, Seasonal storage of hydrogen in stationary systems with liquid organic hydrides, International Journal of Hydrogen Energy, 23(10): 905–909.
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Ozaki, M., Davison, J. and Minamiura, J., 2004, Marine transportation of CO2. Proceedings of the 7th International Conference on Greenhouse Gas Control Technologies (Volume II Part 2), Vancouver, 5 –9 September 2004, in Rubin, E.S., Keith, D.W. and Gilboy, C.F. (eds). 2535–2539 (Elsevier, Amsterdam, The Netherlands). Pearson, S.R., 2001, The manufacture of synthetic gas and ethanol from biomass using the Pearson thermo-chemical steam reforming and catalytic conversion processes, 5th International Biomass Conference of the Americas, Orlando, USA, http://bioproductsbioenergy.gov/pdfs/ bcota/session28.html. Pearson, S.R., 2003, The manufacture of synthesis gas from biomass and production of alcohols and electric power using the Pearson thermo-chemical steam reforming and catalytic conversion processes, Poster Presentation 5–18, 25th Biotechnology Symposium, Breckenridge, Colorado, http://www.nrel.gov/biotech_symposium/ session5_pp.html. Perry, R.H., Green, D.W. and Maloney, J.O. (eds), 1999, Perry’s Chemical Engineers’ Handbook, 7th edition (McGraw Hill, New York, USA). Pujado, P., 1986. UOP Molex process for production of n-paraffins and UOP Olex process for olefin recovery, in Meyers, R.A. (ed). Handbook of Petroleum Refining Processes, Vol. 8, 71–79 (McGraw-Hill, London, UK). Sahibzada, M., Metcalfe, I.S. and Chadwick., D., 1998, Methanol synthesis from CO/CO2/H2 over Cu/ZnO/Al2O3 at differential and finite conversion, Journal of Catalysis, 74: 111– 118. Shilling, N. and Jones, R.M., 2003, The response of gas turbines to a CO2 constrained environment, paper presented at the Gasification
Technology Conference, 13–14 October, http://www.netl.doe.gov/ coal/turbines/refshelf/GE%20Turbines%20for%20CO2-Constrained%20Environment.pdf. Stuart Energy website, accessed 06/2004, http://stuartenergy.com. Vantine, B., 2004, Pearson technologies is making ethyl alcohol from ‘almost anything’ a reality, Presented at the New Mexico Green Fuels Symposium, Santa Fe Community College, New Mexico Energy, Minerals and Natural Resources, http://www.emnrd.state.nm.us/ ecmd/html/GF_presentations/PTI%20New%20Mex%20Green%20Fuels %202004v2BruceVantine.pdf. van Broekhoven, E.H., 2001, A new solid acid isobutane alkylation technology—Alkyclean, Catalysts Courier, 45, http://www.akzonobelcatalysts.com/html/catalystcourier. Vandenborre, H., 2002, High pressure electrolyser module, Patent EP0995818. Yara, accessed 03/2005, http://www.yara.com.
ACKNOWLEDGEMENTS EPSRC funded this work through the Supergen marine energy programme. The manuscript was received 22 August 2005 and accepted for publication after revision 21 April 2006.
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www.icheme.org/cherd doi: 10.1205/cherd.05204
PROCESSES FOR THE SYNTHESIS OF LIQUID FUELS FROM CO2 AND MARINE ENERGY D. MIGNARD and C. PRITCHARD University of Edinburgh, The Institute for Energy Systems, Edinburgh, UK
I
n Britain, wind and wave power are expected to make a significant contribution to future energy supplies, but the vast majority of the resource is located offshore, away from the mainland electricity grid. In this context, the chemical storage and transport of this energy appears to be an attractive option for the supply of fuels. On the other hand, it is not unlikely that CO2 will be recovered on a large scale in the next 10– 20 years, thus making it available as a carbon source for liquid fuels. A preliminary analysis indicated that the liquid fuels that could be most readily manufactured from hydrogen (from electrolysis) and recycled CO2 were: methanol, mixed alcohols and gasoline (via methanol). In the light of developmental work at NEDO, Japan, methanol appears to be simplest to manufacture. The Mobil Methanol-to-Gasoline process also facilitates the manufacture of a very convenient automotive fuel. Mixed alcohols have the advantages both of low toxicity, and efficient Fischer –Tropsch catalysts for their synthesis (the Pearson process and the Ecalene process). This paper compares the energy efficiency of these three processes. Keywords: synthetic fuels; wind power; marine energies; hydrogen; CO2 sequestration.
INTRODUCTION
located North and West of Scotland, while the electricity transmission grid is poorly developed in the nearby onshore areas. Since a high proportion of renewable energy is expected to come from these remote locations on islands or at sea, the production of liquid fuels could be more convenient than a subsea network of electricity cables. A first possibility is to use hydrogen as an energy vector. While the electrolysis of water is an established technology for producing hydrogen fuel from electricity, the storage and transmission of H2 on a large scale and over long distance requires a practical solution. Possible solutions include the construction of hydrogen pipelines (Air Liquide, 2003); the liquefaction and shipping of hydrogen (Giacomazzi and Gretz, 1993); the use of the methylcyclohexane–toluene– hydrogen system (Newson et al., 1998); and the injection of hydrogen into the existing natural gas network (Haines and Polman, 2004). Appliances designed to use hydrogen must also be available: gas turbines using hydrogen are being developed with higher efficiencies (Shilling and Jones, 2003); and hydrogen fuel cells for stationary applications should become more economical in the future with progress in stack assembly techniques, materials and efficiency [as exemplified by companies such as ITN Power Ltd (2005) or Kansai Electric (2004)]. We have examined the various hydrogen options, and also alternative processes for the manufacture of carbon-based liquid fuels, since these liquid fuels are readily transportable and usable. The description of these processes for carbon-based fuels is the object of this present paper.
This paper presents processes for the production of liquid fuels from recycled CO2 and hydrogen from electrolysis, at locations remote from the electricity grid where wind or marine energies are available. The availability of CO2 feedstock and the issues surrounding the large-scale storage and transmission of hydrogen in such contexts are first discussed. Fuel synthesis from CO2 and hydrogen is then justified as a potential solution, and the available technologies are reviewed for three processes: methanol, mixed alcohols and gasoline synthesis. Finally, Process Flow Diagrams and energy balances are presented for these processes. Chemical Storage of Renewable Energy in the Context of Large Scale, Remote Production The long-term necessity for the world economy is to achieve a sustainable pattern of energy use. In Britain, a significant contribution is expected from resources such as wind and wave power. The UK Atlas of Renewable Energy (ABPmer et al., 2004) shows that the most valuable wave and wind power resource in terms of power density is Correspondence to: Dr D. Mignard, University of Edinburgh, The Institute for Energy Systems, The King’s Buildings, Mayfield Road, Edinburgh EH9 3JL, UK. E-mail: [email protected]
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SYNTHESIS OF LIQUID FUELS FROM CO2 AND MARINE ENERGY Recycling of CO2 as Hydrogen Carrier Fuel A possible route for the synthesis of a liquid fuel consists in reacting (recycled) CO2 with hydrogen obtained from the electrolysis of water. The resulting fuel could be readily used in existing distribution networks, power generators, vehicles or other appliances, and it would facilitate both a substantial penetration of renewable power and also large cuts in CO2 emissions, on a relatively short timescale. If emissions of CO2 are to be curbed, it is not unlikely that this gas will be recovered on a large scale in the next 10–20 years. Should the price of oil remain at its currently high level, CO2 could be profitably used for Enhanced Oil Recovery (EOR) (Berger et al., 2004). The ‘CO2 Infrastructure for EOR in the North Sea’ (CENS) project, which involves the UK, Denmark and Norway (government bodies and companies), is currently evaluating the practicalities of establishing a CO2 infrastructure for sequestration and EOR in and around the North Sea, including recovery of the CO2 from power stations, pipeline transport or shipping, and injection into mature oil reservoirs. 29 Mt y21 of CO2 would come from coal-fired power stations in the UK and Denmark (Markussen et al., 2003).
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(10 000–30 000 m3) should conditions allow for profitable EOR in the North Sea (Berger et al., 2004). However, lower pressures have been recommended for such ships, typically 6.5 bar at 2508C (Aspelund et al., 2004). The power consumption required for CO2 liquefaction in these conditions is 110 kWh t21 liquid CO2 (Aspelund et al., 2004; starting at 1.1 atm, 158C, with seawater available at 108C). In this study, liquefied CO2 at 6.5 bar and 2528C was thermally compressed to produce gaseous CO2 at a pressure of 30 bar, using seawater as the heat source. A ship of hybrid design, for transporting either liquefied CO2 or liquid fuels, would be worth investigating. This would allow the ship to transport the CO2 to the remote area for fuel synthesis, and shipping the synthetic fuel back from the remote location. Such a design was proposed by the Japanese NEDO and RITE study, which considered the production of methanol from solar power and CO2 in areas such as Egypt or Australia (NEDO and RITE, 1998). Finally, the amount of CO2 emitted during transport has been evaluated at about 2% of the CO2 transported per 1000 km, and the boil-off losses would probably be negligible (Ozaki et al., 2004). However, in this study we specifically exclude the (very site-specific) CO2 emissions associated with these shipments.
CO2 Recovery By using high-pressure, superheated steam at the reboiler of the CO2 amine-absorption stripper, and by the exclusive use of existing ‘ultra clean’ flue-gas at 14 –16% CO2 from a coal-fired CHP (combined heat and power) plant, the energy and monetary costs of CO2 recovery may be kept low. In particular, the CHP configuration ensures that the low-grade heat is a useful product, which in turn leads to a higher overall efficiency than that from a power-only plant. For the coal-fired CHP plant, the CENS project claims a cost of $35 t21 for CO2 recovered from flue-gas (Markussen et al., 2003), which represents ca. $33 t21 CO2 after subtracting the CAPEX of the 1500 km pipeline. Energy consumption may be reduced further by adopting the flue-gas scrubbing technology and the KS-1 solvent that were developed by Mitsubishi Heavy Industries, Ltd (MHI), and Kansai Electric Company (KEPCO) (Iijima and Kamijo, 2003). A 160 t day21 CO2 plant in Malaysia has been using the KS-1 solvent since 1999. Solvent consumption there is 0.35 kg t of CO2 recovered, with solvent stability over at least 5700 h; LP steam consumption is 1.5 t per t CO2 recovered, which compares with 2.7 t t21 in the Kerr-McGee/ABB Lummus Global process (Hydrocarbon Processing, 2000). The Kerr-McGee process requires 100 kWh of electrical power per tonne of CO2 before compression. This latter figure, together with that of 860 kWh t21 CO2 for steam requirements from the MHI process, is used in our calculations. CO2 Shipping CO2 transport by ship is done in Europe, where Yara International transports liquefied CO2 from ammonia plants in Porsgrunn (Norway), Sluiskil (The Netherlands) and Ferrara (Italy) (Yara, 2005). For example, the m.v. ‘Coral Carbonic’ stores 1382 t CO2 at 2408C, or 1292 t CO2 at 222.58C, at pressures between 14–20 bara (Anthony Veder, 2005). Bigger ships are planned for construction
Electrolysis Using Variable and Intermittent Current Recent developments have made available on the market alkaline electrolysers that can operate in intermittent or variable conditions at 25 or 30 bar (Stuart Energy, 2004; European Commission, 2003). Appropriate industrial electrodes do not need a protective voltage in these conditions (EUHYFIS, 2004). Advanced designs for alkaline pressure electrolysers include the IMETw technology, which is commercialized by Stuart Energy for capacities up to 1 MW (Stuart Energy, 2004; Vandenborre, 2002), and the Pressure Module Electrolyser (PME) technology, which is being developed by a consortium including MTU-Friedrichshafen, Norsk Hydro, and PMT (Prime Membrane Technology, Belgium) (European Commission, 2003). The PME technology is at the demonstration stage (European Commission, 2003). Neither technology requires lye pumps, and both use separate catholyte and anolyte streams to prevent any mixing of O2 and H2, which is an advantage for operation at low load. The modular construction of the PME pressurized shell allows the electrolyser to be scalable beyond several MW (Kliem, 1995). Fuel Synthesis The following three options were considered, and the corresponding process flow diagrams were devised. Process A. Methanol synthesis (Figure 1) Methanol synthesis from pure CO2 is feasible on existing Cu/ZnO/Al2O3 catalysts used for making methanol from synthesis gas (Sahibzada et al., 1998), according to the reaction CO2 þ 3H2 ! CH3 OH þ H2 O(g) DH298K, 1bar ¼ 49:16 kJ mol1
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(1)
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Figure 1. Flowsheet for the CO2 to methanol process. Key: O ¼ pressure in bar; S ¼ temperature in 8C; Cp ¼ compressor; P ¼ pump; T ¼ tank; R ¼ reactor; HX ¼ heat exchanger; FH ¼ fired heater; Cd ¼ condenser; KB ¼ kettle boiler; DC ¼ distillation column; CW ¼ cooling water.
The observed equilibrium yield when using CO2 þ 3H2 at a pressure of 50 bar was ca. 22% for methanol, due to the inhibitory effect of the water product. Deactivation of the catalyst by water is known to be a problem, but better catalysts have been developed. For instance, Cu/ZnO/ ZrO2/Al2O3/Ga2O3 showed excellent longevity, selectivity and activity: After 2500 h, the Space –Time– Yield of this catalyst was stable at 600 g MeOH/(litre of catalyst) h21 t , 46% better than for a commercial Cu/ZnO/Al2O3 (ICI 51-2 catalyst) (NEDO and RITE, 1998). Mignard et al. (2003) modelled the adiabatic operation of a fixed bed reactor with ICI 51-2. For a minimum 99% yield, the minimum compression requirements (including recycle) would occur at 2278C, 50 bar, with a recycle ratio of 7.9. Optimum conditions with the NEDO and RITE catalyst could be somewhat different. Process B. Mixed alcohols synthesis (Figure 2) Higher alcohols have the advantage that they are less toxic than methanol, and could be handled more safely by the general public. For ethanol, the reactions are CO2 þ H2 ! CO þ H2 O(g)
(2)
DH298K, 1bar ¼ 41:21 kJ mol1 of carbon 2CO þ 4H2 ! CH3 CH2 OH(g) þ H2 O(g) DH298K, 1bar ¼ 256:1 kJ mol1 of ethanol
(3)
Until recently, reaction (3) was not possible with good yields or good selectivity. However, Pearson Technologies Inc. (PTI) (Pearson, 2001, 2003) now claims the invention of a Fischer – Tropsch catalyst capable of converting synthesis gas to ethanol with a yield of 99þ% to ethanol after recycle. The single pass conversion is 15– 60%, depending on the conditions. An important advantage of the Pearson process is that it avoids the formation of significant quantities of byproducts such as CO2 and methane (Pearson, 2001). Carefully controlled cooling is necessary to prevent the reactor from ‘running away’. It is claimed that the temperature and pressure of the process can be lowered to values that are typical of methanol processes, including the ‘low pressure’ processes (i.e., down to 50 bar) (Pearson, 2001). Details of the composition of a synthesis gas feedstock were 51.1% H2, 23.7% CO, 17.1% CO2, and 6.3% CH4 (Vantine, 2004). This indicates that a ratio CO2/(CO2 þ CO) of 42% is acceptable, and that the catalyst must either have a strong reverse water – gas shift (RWGS) activity, or else utilises CO2 directly. PTI has built and operated a 50 t day21 pilot plant processing wood waste to produce ethanol in Aberdeen, North Mississippi (Pearson, 2001). In addition to ethanol, the process is also expected to form methanol, and higher alcohols with three atoms of carbon or more. Pearson (2001) stated that these by-products are separated downstream in the condensate, and recycled back to the process. An equilibrium partial pressure could apply to methanol, as homologation
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SYNTHESIS OF LIQUID FUELS FROM CO2 AND MARINE ENERGY
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Figure 2. Flowsheet for the simplified CO2 to mixed alcohols process, with 15% conversion to ethanol per pass in the alcohol reactor (R2). Separation and recycling in R2 of alcohols other than ethanol was neglected. Key as in Figure 1.
reactions with CO and H2 or direct synthesis from CO2 are possible. However, further polymerization of all alcohols with Fischer –Tropsch reactions would steadily increase the length of carbon chains, and therefore higher alcohols with three atoms of carbon and more would presumably be recycled to the gasifier. For methanol, CO2 and CH4, recycling should eventually attain an equilibrium at the inlet and inhibit further byproduct formation (Mawson et al., 1993). Therefore, we would expect that alcohol formation and water – gas shift are the only significant reactions that affect the composition of the outlet from the reactor. In the absence of published experimental data on the Pearson process, we assumed in this work that the final product was a mixture of alcohols rather than ethanol. However, we have treated the distillation step as though ethanol were the only product, since its separation is the most energy intensive. Finally, we note that a correct choice of residence time in the reactor will be critical: longer residence times will ensure a higher conversion per pass, while shorter residence times will avoid the excessive formation of heavier alcohols with five and more carbon atoms. This is borne out from experience (Bhasin, 1980). Other processes that are worth mentioning for the formation of mixed alcohols from synthesis gas are the Dow process (Conway and Stevens, 1988) and the Ecalene process (Jackson and Mahajan, 2001). One potential difficulty with these processes is the generation of substantial amounts of byproducts: For example, with the catalysts
described by Conway and Stevens (1988), CO conversion to CO2 is typically around 30– 40%; and of the remaining gas phase products, at least 10% is methane and other hydrocarbons. The attraction of the Pearson process comes from its ability to handle CO2 and its claim that it is essentially emission free, i.e., all of the carbon can be utilized without significant CO2 or methane byproduct. Another difficulty with the Ecalenew process seems to be the need to feed sulphur in order to maintain the catalyst activity (the Ecalenew fuel may need subsequent desulphurization). Tentatively, we chose to model a mixed alcohol reactor on the basis of what information we could find on the Pearson process. A RWGS reactor may be needed if the feedstock is pure CO2. A Cu/ZnO/ZrO2/Ga2O3 catalyst for RWGS was proposed (Joo et al., 1999). Using the table provided by these authors, it was found that operating the RWGS reactor at 6008C, 30 bar and no recycle would permit a ratio CO2/ (CO2 þ CO) of 39%. A drop in temperature to 4008C would increase this ratio to 61%. If the catalyst of the alcohol synthesis reactor required higher conversion, a temperature of 8008C would enable a ratio CO2/(CO2 þ CO) of 25%. Process C. Methanol to gasoline (MTG) (Figure 3) Methanol reacts on the ZSM-5 zeolite catalyst to produce dimethyl ether, which further reacts to give hydrocarbons with up to 10 carbon atoms, together with water (Keil, 1999).
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Figure 3. Flowsheet for the methanol to gasoline process. The C3 and C4 separation processes and the alkylation unit are not represented here, and neither is the light gas processing unit. The MTG fluid-bed reactor ‘R2’ is simplified, i.e., the regeneration bed is not shown. Key as in Figure 1.
The Mobil fluid bed process demonstrated in Wesseling, Germany, produced 15.9 m3 day21 of gasoline using a feed of raw methanol having 27% mol. water and 73% mol. methanol. The heat evolved was 1.74 MJ kg21 methanol, recovered through heating oil tubes immersed within the bed (Keil, 1999). When operating at 4138C and 2.75 bar, 99þ% methanol was converted to 60% gasoline, 34.4% of a liquefied gas fraction, and 5.6% light gas. By alkylation of the isobutane with the propene and butenes fraction, a gasoline yield of 88% could be obtained.
METHODS AND MATERIALS Process Design After mass and energy balances were performed and flowsheets drawn, the processes were compared according to their overall net energy requirements. The feed was taken to be CO2 þ 3H2 at 30 bar, 258C. The fact that CO2 and H2 were in stoichiometric proportions with respect to equation (1) allowed a steady state to be reached with respect to the flow rate of recycle gases, with a purge of 1% of the total flowrate taken out of the system to minimize the accumulation of inert compounds or reaction by-products (typically, in the case of methanol synthesis,
about 0.4% of by-product would be expected, mostly in the form of methyl formate; NEDO and RITE, 1998). The electrical and thermal energy needed for CO2 extraction was calculated according to the figures given in the next section, and included in the assessment. For purification by distillation, reboiler duties were estimated from typical values for each system, or by estimating reflux ratios. Compression requirements were calculated taking into account pressure drops through major pieces of equipment (exchangers, condensers and reactors). Exchangers and condensers were of the tubular type, and their design followed the TEMA guidelines (Perry et al., 1999). Spreadsheet calculations were used for all these estimates, as well as an in-house C-program for the methanol reactor of process A. Methanol product was fuel grade at 98% wt (the balance is water) while ethanol was also fuel grade at 99.4% wt. In process B, a RWGS reactor was operated at 30 bar, 6008C inlet and 4458C outlet, CO2 conversion 53% and no recycle. The feed was first preheated by the RWGS reactor outlet, and then heated further by the combustion of the purge stream from the ethanol reactor recycle. Another option would be to provide electrical heating using renewable power. The ethanol reactor was operated at 50 bar, 4008C inlet and 4308C outlet, with two possible levels of CO
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SYNTHESIS OF LIQUID FUELS FROM CO2 AND MARINE ENERGY conversion, 15% or 30%, and the RWGS reaction within the reactor was assumed to maintain the CO2/CO ratio constant. This approach implicitly assumed that CO rather than CO2 was the reagent directly implicated in the synthesis reaction. A purge of 1% on the recycled gas was assumed, in order to prevent the build-up of inerts. Due to the dearth of data on the Pearson process in the open literature, these conditions were working assumptions to assess ethanol synthesis via RWGS using a spreadsheet model. Both the RWGS and the mixed alcohol reactor were of the fixed catalyst bed type. The heat of reaction from the alcohol reactor was used to generate steam. Alcohol separation was carried out at atmospheric pressure. In order to evaluate a worstcase scenario, ethanol was assumed to be the only product, and the distillation produced a 90% azeotrope, which was sent to a drying unit. In process C, the gaseous product from process A was partly condensed to yield a 40% mol. methanol feed for distillation, and the remainder of the gas was separately condensed to yield a 73% methanol feed for the MTG reactor. Instead of distillation, stripping was used to upgrade the 40% product to 73%, which saved up to two thirds of the reboiler duty. A debutanizer column was also required to separate the light gases from the raw MTG product and the liquefied C3 – C4 gases. In our design work, we opted for selective adsorption to separate the reagents involved in the alkylation process from the liquefied gas. Adsorption was preferred over distillation since it generally uses less energy. The gas consisted of 14.5% isobutane, 5% propene, 7.3% butanes, 5.9% propane, and 1.7% n-butane. Propene and butenes were first isolated using the UOP OlexTM process (Pujado, 1986), which selectively adsorbs olefins. In the remaining stream, isobutane was the only remaining branched hydrocarbon, and it was separated by the UOP MolexTM process through selective adsorption of the other (linear) hydrocarbons on 5A molecular sieve. Power and steam requirements for the OlexTM and
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MolexTM processes were taken from Pujado (1986). Finally, for the alkylation reaction we opted for the AlkyCleanTM process developed by Akzo-Nobel, ABB Lummus Global and Fortum. This process relies on reacting the alkylation feedstock on a solid acid catalyst (SAC) at around 908C. This approach avoids the use of strong acids (sulphuric or hydrofluoric acids) and of the extensive effluent processing that are necessary with traditional processes (van Broekhoven, 2001). A demonstration plant has been running in Porvoo, Finland, since 2002, and the process is now reaching commercial status (ABB, 2004). In the absence of any information on energy requirements for this process, our study assumed that these were comparable to those of the H2SO4 process. This assumption could be an overestimate, since the solid acid catalyst approach has costs comparable to that of the H2SO4 process, but has no need for extensive effluent treatment (ABB, 2004). Energy Efficiencies The energy efficiency was calculated as the ratio of the chemical energy stored as Lower Heating Value (Qstored) to the electrical energy used (Einput). For process A and B, there was a net requirement for steam (Qnet ¼ Qp 2 Qgen, where Qp is the required steam heat and Qgen is the heat available from the reactions). This net demand could be met either by burning some of the synthetic fuel, or by using electrical heating (efficiency 0.94). In the case of process C, which was generating enough steam to sustain itself, Qnet ¼ 0. However, for all processes, the steam required for the extraction of CO2 (QCO2), was considered to come from the power plant that generated the CO2 in the first place, and it was decoupled from the fuel production process. Therefore, the corresponding figure was not included in Qnet, but instead it was treated as a power input to the process, in the denominator of the ratio. For this purpose, it was considered that 0.25 kWhe had to be
Table 1. Breakdown of energy balance for processes A, B and C. Energy balance basis: 1 kmol of (CO2 þ 3H2) feed
Process A Methanol
Process B Mixed alcoholsa
Process C Gasoline
Power required for electrolysis, Eel kWhe kmol21 Pumping Compression and pumping CO2 extraction (CO2 at 1 bar) Total ¼ Ep kWhe/kmol21 CO2 liquefaction Compressing fresh feed Feed to ethanol reactor Gas recycle (synthesis reactor) OlexTM MolexTM AlkycleanTM, incl. steam Ethanol drying
80.64 0.032 1.1 1.21 0.557 — 1.243 — — — —
80.64 0.032 1.1 1.21 — 0.717 0.679 — — — figure included in steam requirement
80.64 0.192 1.1 1.21 0.557 — 1.243 0.008 0.017 0.114 —
Steam required for CO2 extraction, QCO2 kWhth kmol21 Methanol or ethanol separation Steam for process Debutanizer Total ¼ Qp 21 kWhth kmol OlexTM MolexTM From stream Steam generated, From synthesis reactor Total ¼ Qgen kWht kmol21 Chemical energy stored (LHV basis), Qstored kWh kmol21
9.78 7.96 — — — 3.56 0 44.29
9.78 11.37 — — — 0 6.02 44.24
9.78 3.30 0.1 0.087 0.281 3.56 3.81 36.13b
a
Assuming 15% conversion per pass for the alcohol reactor; ethanol is the sole alcohol. Excluding LPG, fuel gas and extra steam.
b
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MIGNARD and PRITCHARD Table 2. Energy efficiencies for processes A, B and C [calculated from Table 1 and equations (1) and (2)], and compressed hydrogen. A, Methanol H2 Compressed, 30 to 400 bar
Process S, Chemical Energy stored, kWh/ kWhe
0.61
Fired heating
B, Mixed alcohols
Electrical heating
0.50
Fired heating
0.515
0.45 if yield per pass is 15% 0.47 if yield per pass is 30%
Electrical heating (except RWGS feed) 0.46 if yield per pass is 15% 0.47 if yield per pass is 30%
C, Gasoline Heating using by-product steam 0.42 (gasoline only) 0.46 (including LPG) 0.49 (including LPG and extra steam)
sacrificed to produce 1 kWhth of LP steam for the separation of CO2 using amine technology (Bolland and Undrum, 1998). The other terms in the denominator were the power input from electrolysis (Eel), and that from compression and pumping (Ep), which included direct power consumption for CO2 extraction. Electrical heating (EH ¼ Qnet/0.94) was also added there if considered. For processes A, B and C, the chemical energy stored was S ¼ ðQstored Qnet Þ=(Eel þ Ep þ 0:25 QCO2 )
(1)
If using electrical heating for processes A and B: S ¼ Qstored =(Eel þ Ep þ Qnet =0:94 þ 0:25 QCO2 )
(2)
RESULTS AND ANALYSES Flowsheets Figures 1– 3 at the end of the paper present the flowsheets for process A, B and C respectively. Energy Storage Table 1 presents the calculated energy requirements and byproducts from processes A, B and C. Table 2 presents the storage efficiency S as defined previously. DISCUSSION Figures 1 – 3 demonstrate the relative simplicity of the methanol process compared with the others. The gasoline process is the most complex one, and most of the refinery plant has not been shown in Figure 3. Table 2 reflects the fact that compressed hydrogen would be the most desirable fuel in terms of the efficiency of chemical storage, when its storage and transmission is not an issue. This is a simple consequence of deriving the other fuels from it. In the cases where hydrogen storage or transmission is not possible or desirable, the methanol process is the most efficient for storing chemical energy. Efficiencies were 0.50 kWh/kWhe when meeting the net steam requirements with fired boilers, and 0.515 kWh/kWhe with electrical boilers. Electrical heating combined with heat storage would also allow more flexible operation in variable or intermittent operation. The mixed alcohol process (B) seemed to achieve efficiencies that were about 0.03– 0.05 kWh/kWhe lower
than those for the methanol process. When assuming 15% conversion per pass for the ethanol reactor, the figures were 0.43 kWh/kWhe with fired heating and 0.47 kWh/ kWhe with electrical heating. This was due to higher steam requirements for distillation and drying, together with the requirement to heat the feed stream to 6008C and the higher pressure-losses along heat exchangers and condensers. However, these figures could be on the pessimistic side, since conversion rate per pass might be as high as 60% (Pearson, 2001). Table 2 shows that a doubling of this parameter to 30% allows a gain of 0.01– 0.02 kWh/kWhe. The other large source of efficiency loss for process B was the distillation step: we found that decreasing the energy consumption of this step to a level comparable to that for methanol in process A would raise the amount of stored chemical energy by 0.02 kWh/kWhe. A possible way to achieve an improvement of this order or even greater might be to separate water from the diluted product in the gas phase, e.g., using membranes. Another possibility is to effect a more complete conversion of the CO2 to CO in the RWGS reactor, including a recycle, followed by condensation of more water at that stage. Unknowns in the parameters of the Pearson process imply that these figures are only indications of what could be achievable: . An improvement to avoid the RWGS step might be to formulate a mixed catalyst that combined ethanol synthesis and RWGS catalyst, or to design an integrated reactor for the two reactions. (The high temperatures that we considered are thermodynamically favourable to the RWGS reaction, and kinetically favourable to the formation of alcohols.) . It is not apparent to the authors whether the catalyst would be able to utilise pure CO2 in the absence of CO. . Another unknown is the amount of alcohols, other than ethanol, that form in the reactor. If significant amounts of byproducts are formed and a pure product is desired, the separation process might require more energy downstream. . Finally, it is known that control of the the H2: CO ratio is critical in order to maximize selectivity towards alcohols (Landis et al., 2005). Given that pure hydrogen is available in our process, it should be possible to inject it at appropriate positions along the reactor in order to maximize the yield of alcohols per pass. The gasoline process (C) generated excess steam, which would be adequate to meet all process heating requirements
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SYNTHESIS OF LIQUID FUELS FROM CO2 AND MARINE ENERGY once integration of streams had been carried out. When considering gasoline alone, efficiencies were lower than for process A and B (e.g., 16% lower than for methanol), but the figure reported in Table 2 shows that the inclusion of LPG product and of excess steam (which could be utilized elsewhere or combined with heat storage) could practically close this gap. The good overall performance of process C was explained by the avoidance of most of the reboiler duty for separating alcohols from water that was encountered in processes A and B. How significant are these figures when comparing the different processes? Some of the uncertainty in our knowledge of process conditions had little influence on the final figures, e.g., the AlkycleanTM process decreased the efficiency of storage by only 0.002 kWh/kWhe. By contrast the heat evolved by the synthesis reaction and the heat needed for separation of the products (where needed), were dominant. A less obvious effect was that the pressure drops in tubular heat exchangers and condensers had a pronounced impact on overall performance, and had to be maintained to values within a few percent of the working pressure. We expect that the figures in Table 2 are sufficiently accurate to allow comparison, although it is acknowledged that, for the ethanol process, further data would be needed and further improvements could be made. The operability of these schemes in intermittent or variable conditions is now being investigated. CONCLUSIONS Three schemes for the synthesis of liquid fuels from recycled CO2 and renewable power were presented. The methanol process had the highest efficiency, with 0.50 kWh of chemical energy stored for each kWhe used. The use of electrical boilers, instead of fired ones utilizing part of the fuel product, would raise this figure to 0.51 kWh/kWhe. The mixed alcohol process gave lower results when lower-range values were assumed for the conversion-per-pass. However, its potentially high conversion per pass (60% instead of 22% for the methanol process) could make it worthy of consideration. There also might be the possibility of combining the RWGS and the alcohol synthesis in one single reactor, as well as feeding fresh hydrogen at various locations along the reactor in order to adjust the H2:CO ratio. The gasoline process benefited from an integrated design of the methanol and hydrocarbon plants, and avoided the energy-intensive distillation step required by the other processes. It reached efficiencies of 0.42 kWh/kWhe without accounting for excess steam production and LPG product, and 0.49 kWh/kWhe when these other products were considered. These processes are now being evaluated for operation under variable power supply. REFERENCES ABB website, accessed 11/2004, Making the grade safely, http:// www.abb.com/global. ABPmer, The Met Office, Garrad Hassan, and Proudman Oceanography Laboratory, 2004, UK Atlas of Offshore Renewable Energy, Atlas Pages Technical Report, Report for the Department of Trade and Industry (DTI), UK. Air Liquide, 2003, Hydrogen Delivery Technologies and Systems— Pipeline Transmission of Hydrogen. Presented at the Strategic
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ACKNOWLEDGEMENTS EPSRC funded this work through the Supergen marine energy programme. The manuscript was received 22 August 2005 and accepted for publication after revision 21 April 2006.
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