As mentioned above, one possibility of operation with a more favorable energetic balance is under autothermal steam reforming conditions, which are produced by introducing oxygen in the reaction mixture. As has occurred for hydrocarbons and methanol feedstock [5,51], research is now under way to develop catalysts that control the oxidation process through the combining of catalytic partial oxidation and steam reforming of ethanol. The oxidizing environment reduces the carbon poisoning of the catalyst and could promote the decomposition of intermediate molecules such as ethylene and acetaldehyde. On the other hand, an excess of oxygen leads to a strong reduction of hydrogen as reaction product. In this respect, studies under partial oxidation conditions could contribute to a better knowledge of autothermal ethanol steam reforming. Some studies on cat­alytic behavior of Ni- [52], Pt — [53] and Ru-based [54] catalysts have recently been reported.

In autothermal conditions, reports concerned the use of Ni and Cu catalysts [38-40,55] and promoted noble metals supported on highly stable carriers, i. e., Pt-CeO2-La2O3/Al2O3 [13], Rh/CeO2 [14], Rh/Al2O3 [56]. The main role of pro­moters is related in this case to metal-promoter interactions [13], which affect the adsorption-decomposition of ethanol to CH4 and CO and their subsequent reforming with steam to produce H2. Most of the results reported point to the need to operate in a narrow range of water/oxygen/ethanol ratios to achieve 100% ethanol conversion, maximum hydrogen yield and minimum methane and carbon monoxide production.

The synthesis gas from the gasifier contains a considerable amount of CO2. After reforming or shifting, this amount increases. To get the ratio (H2-CO2)/(CO + CO2) to the value desired for methanol synthesis, part of the carbon dioxide could be removed. For this purpose, different physical and chemical processes are available. Chemical absorption using amines is the most conventional and com­mercially best-proven option. Physical absorption, using Selexol, has been devel­oped since the seventies and is an economically more attractive technology for gas streams containing higher concentrations of CO2. As a result of technological development, the choice for one technology or another could change in time, e. g., membrane technology, or still better amine combinations, could play an important role in future.

Chemical absorption using amines is especially suitable when CO2 partial pressures are low, around 0.1 bar. It is a technology that makes use of chemical equilibria, shifting with temperature rise or decline. Basically, CO2 binds chem­ically to the absorbent at lower temperatures and is later stripped off by hot steam. Commonly used absorbents are alkanolamines applied as solutions in water. Alkanolamines can be divided into three classes: primary, secondary, and tertiary amines. Most literature is focused on primary amines, especially monoethanola — mine (MEA), which is considered the most effective in recovering CO2 (Farla et al. 1995; Wilson et al. 1992), although it might well be that other agents are also suitable as absorbents (Hendriks 1994). The Union Carbide “Flue Guard” process and the Fluor Daniel Econamine FG process (formerly known as the

Dow Chemical Gas/Spec FT-1 process) use MEA, combined with inhibitors to reduce amine degradation and corrosion. The cost of amine-based capture are determined by the cost of the installation, the annual use of amines, the steam required for scrubbing and the electric power. There is influence of scale and a strong dependence on the CO2 concentration (Hendriks 1994). The investment costs are inversely proportional to the CO2 concentration in the feed gas when these range from 4% to 8%. MEA is partly entrained in the gas phase; this results in chemical consumption of 0.5-2 kg per tonne CO2 recovered (Farla et al. 1995; Suda et al. 1992). The presence of SO2 leads to an increased solvent consumption (Hendriks 1994).

When the CO2 content makes up an appreciable fraction of the total gas stream, the cost of removing it by heat regenerable reactive solvents may be out of proportion compared to the value of the CO2. To overcome the economic disadvantages of heat-regenerable processes, physical absorption processes have been developed that are based on the use of essentially anhydrous organic sol­vents, which dissolve the acid gases and can be stripped by reducing the acid-gas partial pressure without the application of heat. Physical absorption requires a high partial pressure of CO2 in the feed gas to be purified, 9.5 bar is given as an example by Hendriks (1994). Most physical absorption processes found in the literature are Selexol, which is licensed by Union Carbide, and Lurgi’s Rectisol (Hendriks 1994; Hydrocarbon Processing 1998; Riesenfeld et al. 1974). These processes are commercially available and frequently used in the chemical indus­try. In a countercurrent flow absorption column, the gas comes into contact with the solvent, a 95% solution of the dimethyl ether of polyethylene glycol in water. The CO2 rich solvent passes a recycle flash drum to recover co-absorbed CO and H2. The CO2 is recovered by reducing the pressure through expanders. This recovery is accomplished in serially connected drums. The CO2 is released partly at atmospheric pressure. After the desorption stages, the Selexol still contains 25-35% of the originally dissolved CO2. This CO2 is routed back to the absorber and is recovered in a later cycle. The CO2 recovery rate from the gas stream will be approximately 98% to 99% when all losses are taken into account. Half of the CO2 is released at 1 bar and half at elevated pressure: 4 bar. Minor gas impurities such as carbonyl sulfide, carbon disulfide and mercaptans are removed to a large extent, together with the acid gases. Also hydrocarbons above butane are largely removed. Complete acid-gas removal, i. e., to ppm level, is possible with physical absorption only, but is often achieved in combination with a chem­ical absorption process. Selexol can also remove H2S, if this were not done in the gas-cleaning step.

It has been suggested by De Lathouder (1982) to scrub CO2 using crude methanol from the synthesis reactor that has not yet been expanded. The pressure needed for the CO2 absorption into the methanol is similar to the methanol pressure directly after synthesis. This way only a limited amount of CO2 is removed, and the required CO2 partial pressure is high, but the desired R can be reached if conditions are well chosen. The advantage of this method is that no separate regeneration step is required and that it is not necessary to apply extra
cooling of the gas stream before the scrubbing operation. The CO2 loaded crude methanol can be expanded to about atmospheric pressure, so that the carbon dioxide is again released, after which the methanol is purified as would normally be the case.

Physical adsorption systems are based on the ability of porous materials (e. g., zeolites) to selectively adsorb specific molecules at high pressure and low tem­perature and desorb them at low pressure and high temperature. These processes are already commercially applied in hydrogen production, besides a highly pure hydrogen stream a pure carbon dioxide stream is coproduced. Physical adsorption technologies are not yet suitable for the separation of CO2 only, due to the high energy consumption (Ishibashi et al. 1998; Katofsky 1993).

Displacement of petroleum by fuel ethanol is approaching 3% of the liquid transportation fuel used in the United States. Expanding ethanol to replace more than 10% of fuel needs will require development of additional and lower-cost feedstocks. Only lignocellulosic biomass is available in sufficient quantities to augment starch as an ethanol feedstock source. As discussed previously, corn fiber and corn stover are potential sources of lignocellulosic biomass for fermen­tation. Other possible feedstocks are agricultural residues such as wheat and rice straws and sugar cane bagasse, energy crops including switch grass and softwood trees, and waste materials such as pulp and paper sludge and recycled office paper. The capacity to process and ferment even one of these categories of biomass would significantly increase production of ethanol; however, practical aspects of collection and storage must be addressed for many of these resources.

Several technological constraints limit fermentation of biomass feedstocks. Lignocellulosic biomass can be pretreated and enzymatically hydrolyzed to yield a mixture of sugars including glucose, galactose, arabinose, and xylose [32]. However, hydrolytic enzymes are inefficient and expensive. More-effective pre­treatment methods, as well as active and cost-effective enzymes, are needed for an economical process. As mentioned previously, microbes that efficiently fer­ment multiple sugars to ethanol must be developed in order to convert biomass to ethanol. Fermenting microbes also must tolerate the inhibitory compounds generated during biomass hydrolysis, or alternatively, cost-effective methods for inhibitor abatement must be in place. A study comparing dry-grind production of ethanol from corn and ethanol produced from corn stalks concluded that producing ethanol from corn stover would cost $1.45 per gallon compared to $0.96 from corn starch [33]. Despite these obstacles, one company, Iogen Corp. (Ottawa) has begun to produce ethanol from biomass.

Shelley D. Minteer

Department of Chemistry, Saint Louis University, Missouri

CONTENTS

History of Ethanol-Based Fuels……………………………………………………………………. 125

Oxygenated Fuels…………………………………………………………………………………………. 126

Ethanol Production……………………………………………………………………………………….. 128

Engine Issues………………………………………………………………………………………………… 129

E-Diesel…………………………………………………………………………………………………………. 131

Conclusions…………………………………………………………………………………………………… 133

References…………………………………………………………………………………………………….. 133

Abstract Ethanol was first used as a fuel for a combustion engine in 1897, but it has not emerged into the fuel market as a fuel, but rather as an oxygenate. The United States, Canada, Brazil, and many other countries have adopted the use of ethanol blended with gasoline at low concentration to improve emissions. Ethanol can also be blended with diesel at low concentrations (10% to 15%). Most commonly, ethanol is blended with gasoline at concentration of 10% and this oxygenated fuel is referred to as E10 or gasohol.

Alloys of platinum and ruthenium have become common electrocatalysts for fuel cells, because it is believed that alloying ruthenium with platinum will help increase the carbon monoxide tolerance of the platinum catalysts. Alloys of platinum and ruthenium have also been used extensively for DMFC fuel cells, along with hydro — gen/oxygen fuel cells that employ hydrogen gas formed from a reformation process that may have carbon monoxide or carbon monoxide-like by-products. Although

extensive research was done on Pt/Ru alloys on carbon supports and platinum on carbon supports, there was no statistical difference between the selectivity of the two catalysts for ethanol electrooxidation [13]. Figure 10.2 shows a comparison of fuel cell performance for different alcohol fuels employing Pt/Ru alloys as catalysts. It is apparent that methanol performance is better at high current densities (at a current density of 250mA/cm2, the cell voltages are 0.354V for methanol, 0.305V for ethanol, 0.174V for 1-propanol, and 0.054V for 2-propanol [13]), but ethanol performance is better at low current densities (>0.05V at low current densities). The excellent performance of ethanol at low current density is likely due to a decrease in crossover of ethanol versus methanol to the cathode. It is also interesting to note that propanol performance is significantly worse than methanol and ethanol. 1-propanol oxidation forms carbon dioxide and propionaldehyde, but 2-propanol oxidation forms carbon dioxide and acetone [14]. The direct alcohol fuel cells studied in Figure 10.2 are being operated at a temperature of 170°C [13]. This temperature is extremely high (harsh enough that Nafion is not particularly stable and another polymer electrolyte (polybenzimidazole) was used) and is above the temperatures that are realistic for portable power applications, but they provide a benchmark for comparing the 4 alcohols.

PtSn Catalysts

Zhou and coworkers have studied the effect of other alloys on ethanol electro­oxidation. Figure 10.3 shows representative cyclic voltammograms of alloys of

platinum with ruthenium, tungsten, palladium, and tin. These voltammograms were taken at room temperature in solutions that contain 1.0 M ethanol and 0.5 M sulfuric acid. The voltammograms show the largest catalytic activity (current density at the oxidation peak) for PtSn on carbon, but the PtRu on carbon has the lowest over­potential for the ethanol oxidation peaks (0.23 V lower than pure platinum on carbon) [15]. Figure 10.4 shows voltage-current curves and power curves for the same catalysts in a direct ethanol fuel cell at 90°C. The results indicate that Sn, Ru, and W increase the catalytic activity for ethanol oxidation on platinum (max­imum power density of 52.0 mW/cm2, 28.6 mW/cm2, and 16.0mW/cm2, respec­tively, compared to 10.8 mW/cm2 for pure platinum on carbon [15]). Tin and ruthenium are believed to have a bifunctional mechanism to supply surface oxygen containing species for the oxidative removal of carbon monoxide like species that typically passivate the surface of pure platinum [16]. The proposed mechanism for ethanol oxidation at Pt/Sn alloys is shown below [17]:

C2H5OH + Pt(H2O) ^ Pt(C2H5OH) + H2O Pt(C2H5OH) + Pt ^ Pt(CO) + Pt(res) + xH++ xe — Pt(C2H5OH) ^ Pt(CH3CHO) + 2H + + 2e-

Pt(CH3CHO) + SnCl4(OH)2- ^ CH3COOH + SnCl2- + 2H2O + e — + H+ Pt(CO), Pt(res) + SnCl4(OH)2- ^ CO2 + SnCl2- + H2O + Pt H2O + Pt = Pt(OHU + e — + H+

2Pt(OH)a* + SnCl4- = SnCl4(OH)2- + 2Pt

where Pt(res) is an oxidized residue adsorbed to the surface of platinum, Pt(H2O) is water adsorbed to the surface of platinum, Pt(CO) is carbon mon­oxide adsorbed to the surface of platinum, Pt(C2H5OH) is ethanol adsorbed to the surface of platinum, and Pt(CH3CHO) is acetaldehyde adsorbed to the surface of platinum.

After Pt/Sn alloys were determined to be the optimal elemental alloy, Zhou and coworkers examined the importance of tin content and temperature on the fuel cell power curves. Figure 10.5 shows the effect of altering the tin content on the direct ethanol fuel cell performance at a temperature of 60°C. The figure shows both current voltage curves and power curves. The results clearly show that Pt3Sn2 on carbon is the best catalyst choice for 60°C [18]. Figure 10.6 shows the effect of altering tin catalyst content on the fuel cell performance at a tem­perature of 90°C. The results clearly show that Pt2Sn1 on carbon is best for temperatures that are greater than 75°C [18]. Figure 10.5 and Figure 10.6 show that tin content does affect fuel cell performance and temperature affects the catalytic activity of each fuel cell differently. The operating temperature for DEFC

is a function of application. Most portable power applications need to operate between room temperature and 50°C, but performance tends to increase with temperature until crossover and/or polymer electrolyte membrane degradation take over. At 30°C, DEFC have maximum power densities that range from 2 to 10 mW/cm2 [19]. Figure 10.7 shows the effect of a wider range of temperatures (50°C-110°C) for a DEFC with a Pt-Sn (9:1)/C anode. It is important to note that fuel cell performance is a function of temperature and a degradation is not seen at high temperatures [5]. Open circuit potentials do not vary significantly with temperature, but maximum power ranges from 6 to 26 mW/cm2 [5].

Catalyst loading and catalyst supports have also been investigated as para­meters that may affect DEFC performance. Studies in hydrogen/oxygen and DMFC have shown that loading of the catalyst can affect fuel cell performance. If the catalyst loading of the DEFC in Figure 10.5 is changed from 30% metal on vulcanized carbon XC-72 to 60% metal on vulcanized carbon XC-72, the maximum power can increase to 28 mW/cm2 and the open circuit potential can increase from 0.72V to 0.75 V [5]. Research has also shown that transitioning from vulcanized carbon supports (XC-72) to multiwall carbon nanotubes (MWNTs) increases both the open circuit potential and the maximum power density of a DEFC with a platinum/tin alloy catalyst [9]. This is shown in Figure 10.8 where the open power curve shows an increase from 30 mW/cm2 to 38 mW/cm2 with an increase of 80 mV in open circuit potential.

DEFCs can be fabricated by methods similar to DMFCs. The most common format is the membrane electrode assembly (MEA). A MEA is a single assembly that contains the anode, the cathode, and the polymer electrolyte membrane

in ionic contact with each other. MEAs are formed most commonly using a conventional heat pressing method, but they can also be fabricated using a decal transfer method. The conventional method involves sandwiching the PEM between an anode and cathode and heat pressing the sandwich at a temperature above the glass transition temperature of the PEM to melt the electrodes into ionic contact with the PEM. The conventional method shows a 34% decrease in power density over a 10-hour period and delamination of the electrodes from the Nafion [20]. This is likely due to the increased swelling of Nafion in the presence of ethanol, but the decal transfer method only shows a 15% decrease and no delamination, along with no change in resistance [20]. Therefore, the decal transfer method is a better method for forming DEFC MEAs. The decal transfer method involves spray painting the catalyst layer directly onto the polymer electrolyte membrane, instead of onto an electrode support (such as carbon paper) and then heat pressing into the polymer electrolyte membrane.

Researchers have also studied tertiary catalyst systems, but the fuel cell performance has not been greatly affected by adding a third component to the system for alloys containing platinum and ruthenium with a third component of tungsten, tin, or molybdenum [15]. Tertiary catalysts with tungsten and tin did show a measurable increase in power compared to pure Pt/Ru alloys, but both power densities are less than pure Pt/Sn alloys under the same operating condi­tions [15].

CONCLUSIONS

Direct ethanol fuel cells are a relatively new technology for portable power generation. Results have concluded that electrochemical oxidation of ethanol on platinum-based catalysts is not significantly lower than for methanol [21] and the intermediate products of ethanol oxidation are less toxic than methanol oxidation. Although catalytic performance with pure platinum catalysts is low, the perfor­mance of Pt/Sn and Pt/Ru alloys is good. Future research will focus on the development of electrocatalysts that show improved catalytic activity and lower electrode fouling at low and moderate temperatures (room temperature to 50°C). Research on fuel cell lifetimes will also be conducted to study the long-term effects of continuous operation on the catalysts, electrode support, and polymer electrolyte membranes. Improved lifetimes are an issue for both methanol and ethanol, because both oxidation processes produce carbon monoxide and carbon monoxide-like products that adsorb/passivate the catalyst and both alcohols swell the polymer electrolyte membrane, which typically decreases the lifetime and stability of the membrane. Overall, DEFCs are a relatively new technology compared to DMFCs, but ethanol has advantages over methanol in decreased toxicity and environmental issues. From a catalytic perspective, the catalytic rates are similar between ethanol and methanol oxidation, but methanol oxidation is more efficient (produces a larges percentage of carbon dioxide (the complete oxidation product)).

Tilapia are a tropical fish and require warm water. As a result, shipping fingerlings during the winter is extremely risky. Therefore, the project will produce all needed fingerlings on-site. The project will require from 1100 to 3300 fingerlings per week. It can also act as a regional resource for others that require fingerlings as well. The bulk of the equipment needed to breed these fingerlings is already in place. Whereas the hatchery is somewhat labor intensive, it does guarantee a continuous supply of fingerlings. The males are kept in tanks by themselves and the females are brought to them to mate. Once mating has stopped and the female has a mouth full of eggs, the males are removed and the female is left alone and undisturbed to hatch her young (about 72 hours). After this time, the hatchlings will remain in the mother’s mouth for another 3-7 days, taking short excursions outside to feed on minute particulates and then dart back inside. At a time determined by the mother, she spits them all out and she will accept feed again. When this occurs, she is removed to an isolation tank and full fed until she regains weight. She is then ready to breed again. The particular variety of tilapia to be used is patented and produces all males (males grow 40% faster than females), which are a bright red in color. They are a very forgiving fish, adaptable to many different cultural conditions and, from past experience with them, quite easy to breed and raise. Presently, there is enough room for future expansion to other species, i. e., giant Australian Red Claw crawfish (Cherax quadricarinatus — a lobster-sized crawfish that lives in fresh water), giant freshwater prawns (macro — braccium Rosenbergii) which grow to one pound, and a variety of giant sunfish (bream), which grows to a weight of 5 pounds. All these species, except for sunfish, are tropical or subtropical and must be bred on-site to have a year-long supply. The sunfish will remain isolated from the environment, since interbreeding with native varieties will dilute their genetic uniqueness, resulting in much smaller, stunted fish. The hatchery will be connected to its own smaller aquaponics greenhouse, which will provide biofiltration and a use for the waste generated from the hatchery. It will also add to the weekly harvest of vegetables. Breeding fish on-site will also enable the project to add genetics to its teaching curriculum.

An economic evaluation has been carried out for the concepts considered. Plant sizes of 80, 400, 1000, and 2000 MWth HHV are evaluated, 400 MWth being the base scale. The scale of the conversion system is expected to be an important factor in the overall economic performance. This issue has been studied for BIG/CC systems (Faaij et al. 1998; Larson et al. 1997), showing that the econ­omies of scale of such units can offset the increased costs of biomass transport up to capacities of several hundreds of MWth. The same reasoning holds for the methanol production concepts described here. It should, however, be realized that production facilities of 1000-2000 MWth require very large volumes of feedstock: 200-400 dry tonne/hour or 1.6-3.2 dry Mtonne per year. Biomass availability will be a limitation for most locations for such large-scale production facilities, especially in the shorter term. In the longer term (2010-2030), if biomass pro­duction systems become more commonplace, this can change. Very large scale biomass conversion is not without precedent: various large-scale sugar/ethanol plants in Brazil have a biomass throughput of 1-3 Mtonne of sugarcane per year, while the production season covers less than half a year. Also, large paper and pulp complexes have comparable capacities. The base scale chosen is comparable to the size order studied by Williams et al. (1995) and Katofsky (1993), 370-385 MWth.

The methanol production costs are calculated by dividing the total annual costs of a system by the produced amount of methanol. The total annual costs consist of:

1. Annual investments.

2. Operating and maintenance.

3. Biomass feedstock.

4. Electricity supply/demand (fixed power price).

The total annual investment is calculated by a factored estimation (Peters et al. 1980), based on knowledge of major items of equipment as found in the literature or given by experts. The uncertainty range of such estimates is up to ±30%. The installed investment costs for the separate units are added up. The unit investments depend on the size of the components (which follow from the

Aspen Plus modelling), by scaling from known scales in literature (see Table 2.5), using Equation 2.8:

with R = scaling factor.

Various system components have a maximum size, above which multiple units will be placed in parallel. Hence the influence of economies of scale on the total system costs decreases. This aspect is dealt with by assuming that the base investment costs of multiple units are proportional to the cost of the maximum size: the base investment cost per size becomes constant. The maximum size of the IGT gasifier is subject to discussion, as the pressurised gasifier would logically have a larger potential throughput than the atmospheric BCL.

The total investment costs include auxiliary equipment and installation labour, engineering and contingencies. If only equipment costs, excluding installation, are available, those costs are increased by applying an overall installation factor of 1.86. This value is based on 33% added investment to hardware costs (instrumen­tation and control 5%, buildings 1.5%, grid connections 5%, site preparation 0.5%, civil works 10%, electronics 7%, and piping 4%) and 40% added installation costs to investment (engineering 5%, building interest 10%, project contingency 10%, fees/overheads/profits 10%, and start-up costs 5%) (Faaij et al. 1998).

The annual investment takes into account the technical and economic lifetime of the installation. The interest rate is 10%.

Operational costs (maintenance, labour, consumables, residual streams dis­posal) are taken as a single overall percentage (4%) of the total installed invest­ment (Faaij et al. 1998; Larson et al. 1998). Differences between conversion concepts are not anticipated.

It was assumed that enough biomass will be available at 2 US$/GJ (HHV). This is a reasonable price for Latin and North American conditions. Costs of cultivated energy crops in the Netherlands amount approximately 4 US$/GJ and thinnings 3 US$/GJ (Faaij 1997), and biomass imported from Sweden on a large scale is expected to cost 7 US$/GJ (1998). On the other hand biomass grown on Brazilian plantations could be delivered to local conversion facilities at 1.6—1.7 US$/GJ (Hall et al. 1992; Williams et al. 1995). It has been shown elsewhere that international transport of biomass and bioenergy is feasible against modest costs.

Electricity supplied to or demanded from the grid costs 0.03 US$/kWh. The annual load is 8000 hours.

Results

Results of the economic analysis are given in Figure 2.7. The 400 MWth conversion facilities deliver methanol at 8.6-12 US$/GJ. Considering the 30%

1 Annual GDP deflation up to 1994 is determined from OECD (1996) numbers. Average annual GDP deflation after 1994 is assumed to be 2.5% for the United States, 3.0% for the EU. Cost numbers of Dutch origin are assumed to be dependent on the EU market, therefore EU GDP deflators are used. 12001 = 0.94 US$2001 = 2.204 M2001.

2 Total pretreatment approximately sums up to a base cost of 8.15 MUS$2001 at a base scale of 33.5 tonne wet/hour with an R factor of 0.79.

3 Based on first-generation BIG/CC installations. Faaij et al. (1995) evaluated a 29-MWe BIG/CC installation (input 9.30 kg dry wood/s, produces 10.55 Nm3 fuel gas/s) using vendor quotes. When a range is given, the higher values are used (Faaij et al. 1998). The scale factors stem from Faaij et al. (1998).

4 Two double-screw feeders with rotary valves (Faaij et al. 1995).

5 12.72 MUS$1991 (already includes added investment to hardware) for a 1650 dry tonne per day input BCL gasifier, feeding not included, R is 0.7 (Williams et al. 1995). Stronger effects of scale for atmospheric gasifiers (0.6) were suggested by Faaij et al. (1998). Technical director Mr. Paisley of Battelle Columbus, quoted by Tijmensen (2000), estimates the maximum capacity of a single BCL gasifier train at 2000 dry tonnes/day.

6 29.74 MUS$1991 (includes already added investment to hardware) for a 1650 dry tonne/day input IGT gasifier, R = 0.7 (Williams et al. 1995). Maximum input is 400-MW^ HHV (Tijmensen 2000).

7 Air Separation Unit: Plant investment costs are given by Van Dijk (van Dijk et al. 1995): I = 0.1069C08508 in MUS$1995 installed, C = Capacity in tonne 02/day. The relation is valid for 100 to 2000 tonne 02/day. Williams et al. (1995) assume higher costs for small installations, but with a stronger effect of scale: I = 0.260C0712 in MUS$1991 fob plus an overall installation factor of 1.75 (25% and 40%). Larson et al. (1998) assume lower costs than Van Dijk, but with an even stronger scaling factor than Williams: 27 MUS$1997 installed for an 1100 tonne O2 per day plant and R=0.6. We have applied the first formula (by Van Dijk) here. The production of 99.5% pure O2 using an air separation unit requires 250-350 kWh per tonne O2 (van Dijk et al. 1995; van Ree 1992).

8 High-temperature heat exchangers following the gasifier and (in some concepts) at other locations are modelled as HRSGs, raising steam of 90 bar/520°C. A 39.2-kg steam/s unit costs 6.33 MUS$1997 fob, overall installation factor is 1.84 (Larson et al. 1998).

9 Tijmensen (2000) assumes the fob price for hot gas cleaning equipment to be 30 MUS$2000 for a 400-MWth HHV input. This equals 74.1 m3/s from a BCL gasifier (T = 863°C, 1.2 bar). There is no effect of scaling.

10Katofsky (1993) assumes compressors to cost 700 US$1993 per required kWmech, with an installation factor of 2.1. The relation used here stems from the compressor manufacturer Sulzer quoted by (2000). At the indicated base scale, total installed costs are about 15% higher than assumed by Katofsky. Multiple compressors, for synthesis gas, recycle streams, or hydrogen, are considered as separate units. The overall installation factor is taken 1.72 because the base unit matches a 400-MWth plant rather than a 70-MWth plant.

investments for steam reformers vary from 16.9 MUS$^3, for a throughput of 5800 kmol meth — ane/hour with an overall installation factor of 2.1 (Katofsky 1993) to 7867 k$^5 for a 6.2 kg methane/s (1390 kmol/hour), overall installation factor is 2.3 (van Dijk et al. 1995). These values suggest a strong effect of scaling R = 0.51, while Katofsky uses a modest R = 0.7. Here, we use the values of Van Dijk in combination with an R factor of 0.6. The total amount of moles determines the volume and thus the price of the reactor.

12 Autothermal reforming could be 50% cheaper than steam reforming (Katofsky 1993), although higher costs are found as well (Oonk et al. 1997).

investment for shift reactors vary from 9.02 MU$^5 for an 8819 kmol CO+H/hr reactor, and an overall installation factor is 1.81 (Williams et al. 1995) to 30 MUS$^ installed for a 350000 Nm3/hr CO+H2/hr (15625 kmol/hr) reactor (Hendriks 1994). Williams assumes an R = 0.65, but comparison of the values suggest only a weak influence of scale (R = 0.94). Here, we use the the values from Hendriks, with R set at 0.85. A dual shift is costed as a shift of twice the capacity.

14Costs for CO2 removal through Selexol amounts 14.3 MUS$^3 fob (overall installation factor is 1.87) for an 810 kmol CO2/hr unit, R = 0.7 (Katofsky 1993) up to 44 MUS$^ installed for a 9909 kmol CO2/hour unit (Hendriks 1994). The value from Hendriks is assumed to be right, since his research into CO2 removal is comprehensive.

15Van Dijk et al. (1995) estimate that a methanol reactor for a 2.1 ktonne methanol per day plant costs 4433 kUS$1995 (fob) or 9526 kUS$1995 installed (overall installation factor is 2.1). The total plant investment in their study is 138 MUS$1995, or 150 MUS$2001. Katofsky (1993) estimates the costs for a 1056 tonne methanol/day plant to be 50 MUS$1995 fob, this excludes the generation and altering of synthesis gas, but includes make-up and recycle compression and refining tower. A 1000 tpd plant costs about 160 MUS$2001, and a 2000 tpd plant 200 MUS$2001, which suggests a total plant scale factor of 0.3 (Hamelinck et al. 2001). These values come near the ones mentioned by Katofsky. This implies that the values given by Van Dijk are too optimistic and should be altered by a factor 1.33. It is therefore assumed that the base investment for the methanol reactor only is 7 MUS$2001, the installation factor is 2.1. The influence of scale on reactor price is not assumed to be as strong as for the complete plant: 0.6.

installed costs for a 456 tonne per day liquid-phase methanol unit, are 29 MU$^7, excluding generation and altering of synthesis gas, but including make-up and recycle compression, and refining tower. R = 0.72 (Tijm et al. 1997). Corrected for scale and inflation this value is about half the cost of the conventional unit by Katofsky and the corrected costs of Van Dijk. It is therefore assumed that the price of a liquid-phase methanol reactor is 3.5 MUS$2001 for a 2.1 ktonne per day plant, installation factor is 2.1.

17Cost number for methanol separation and refining is taken from Van Dijk, increased with 33% as described in note 15.

18For indication: A complete combined cycle amounts to about 830 US$^7 per installed kWe. Quoted from Solantausta et al. 1996 by Oonk et al. 1997.

19Scaled on gas turbine size.

system consists of water and steam system, steam turbine, condenser and cooling. Scaled on steam turbine size.

21 Expansion turbine costs are assumed to be the same as steam turbine costs (without steam system). 22Overall installation factor. Includes auxiliary equipment and installation labor, engineering and contingencies. Unless other values are given by literature, the overall installation factor is set 1.86 for a 70-MWth scale (Faaij et al. 1998). This value is based on 33% added investment to hardware costs (instrumentation and control 5%, buildings 1.5%, grid connections 5%, site preparation 0.5%, civil works 10%, electronics 7%, and piping 4%) and 40% added installation costs to investment (engineering 5%, building interest 10%, project contingency 10%, fees/overheads/profits 10%, and start-up costs 5%). For larger scales, the added investments to hardware decreases slightly.

23 Maximum sizes from Tijmensen (2000).

uncertainty range, one should be careful in ranking the concepts. Methanol 4 and 6 perform somewhat better than the other concepts due to an advantageous combination of lower investment costs and higher efficiency. The lowest methanol production price is found for concepts using the BCL gasifier, having lower investment costs. The combination of an expensive oxygen fired-IGT gasifier

with a combined cycle seems generally unfavorable, since the efficiency gain is small compared to the high investment.

Investment redemption accounts for 42-76% of the annual costs and is influ­enced by the unit investment costs, the interest rate and the plant scale. The build­up of the total investment for all concepts is depicted in Figure 2.8. It can be seen that the costs for the gasification system (including oxygen production), synthesis gas processing and power generation generally make up the larger part of the investment. For autothermal reforming higher investment costs (Oonk et al. 1997) would increase the methanol price from considered concepts by about 1.5 US$/GJ. Developments in gasification and reforming technology are impor­tant to decrease the investments. On the longer term, capital costs may reduce due to technological learning: a combination of lower specific component costs and overall learning. A third plant built may be 15% cheaper leading to an 8-15% product cost reduction.

The interest rate has a large influence on the methanol production costs. At a rate of 5% methanol production costs decrease with about 20% to 7.2-9.0 US$/GJ. At a high-interest rate (15%), methanol production costs become 9.9-14 US$/GJ. Going to 1000 and 2000 MWth scales, the methanol production costs reach cost levels as low as 7.1-9.5 US$/GJ.

Feedstock costs account for 36-62% of the final product costs for the men­tioned technologies. If a biomass price of 1.7 US$/GJ could be realized (a realistic price for, e. g., Brazil), methanol production costs would become 8.0-11 US$/GJ for 400 MWth concepts. On the other hand, when biomass costs increase to 3 US$/GJ (short term Western Europe), the production cost of methanol will increase to 10-16 US$/GJ.

FIGURE 2.9 Optimistic view scenario. Different cost reductions are foreseeable: (1) biomass costs 1.7 US$/GJ instead of 2 US$/GJ, (2) technological learning reduces capital investment by 15% and (3) application of large scale (2000 MWth) reduces unit investment costs.

If the electricity can be sold as green power, including a carbon neutral premium, the fuel production costs for power coproducing concepts drops, where the green premium essentially pays a large part of the fuel production costs. A power price of 0.08 US$/GJ would decrease methanol costs to -0.6-9.5 US$/GJ. Of course the decrease is the strongest for concepts producing more electricity. A green electricity scenario, however, may be a typical western Euro­pean scenario. As such it is unlikely that it can be realized concurrent with biomass available at 1.7 US$/GJ.

In the long term, different cost reductions are possible concurrently (Tij — mensen 2000). Biomass could be widely available at 1.7 US$/GJ, capital costs for a third plant built are 15% lower, and the large (2000 MWth) plants profit from economies of scale. These reductions are depicted in Figure 2.9: methanol concepts produce between 6.1-7.4 US$/GJ. The influence of capital redemption on the annual costs has strongly reduced and the fuel costs of the different concepts lie closer together.

Previous analyses on short-term methanol production by Katofsky (1993) and Williams et al. (372 MWHHV, 3.4 US$/GJHHV feedstock, 0.07 US$/kWhe (Williams 1995; Williams et al. 1995)) yielded similar energy efficiencies (54-61% by HHV), but significantly higher methanol production costs: 14-17 US$/GJHHV. The largest difference is in the higher capital costs: higher TCI and higher annuity give 25-50% higher annual capital costs. The ADL/GAVE study (Arthur D. Little 1999) reports 13 US$/GJ methanol (feed 2 US$/GJ, 433 MW input) largely using input parameters from Katofsky. Komiyama et al. (2001) instead give much lower costs than presented here: 5 US$/GJHHV for methanol at 530 MWHHV biomass
input. However, in that study, process efficiencies and biomass cost are not given and a significant amount of energy is added as LPG.

In these long-term scenarios, methanol produced from biomass costs consid­erably less than methanol at the current market, which is priced about 10 US$/GJ (Methanex 2001). For application as automotive fuel, comparison with gasoline and diesel is relevant. Their production costs vary strongly depending on crude oil prices, but for an indication: 2003 gasoline market prices were about 7 US$/GJ at oil prices of 25-30 US$/bbl (BP 2004). DOE/EIA projects the world oil price in 2013 to amount between 17 and 34 US$, crude oil prices may decline as new deepwater oil fields are brought into production in the Gulf of Mexico and West Africa, new oil sands production is initiated in Canada, and OPEC and Russia expand production capacity (DOE/EIA 2005). In 2004 the average oil price was some 35 US$/bbl and currently even higher prices of about 50 US$/bbl are paid.

CONCLUSIONS

Methanol can be produced from wood via gasification. Technically, all necessary reactors exist and the feasibility of the process has been proven in practice. Many configurations are possible, of which several have been discussed in this chapter. The configurations incorporated improved or new technologies for gas processing and synthesis and were selected on potential low cost or high-energy efficiency. Some configurations explicitly coproduced power to exploit the high efficiencies of once-through conversion. The overall HHV energy efficiencies remain around 55%. Accounting for the lower energy quality of fuel compared to electricity, once-through concepts perform better than the concepts aiming at fuel only production. Also hot gas cleaning generally shows a better performance. Some of the technologies considered in this chapter are not yet fully proven/commer — cially available. Several units may be realized with higher efficiencies than con­sidered here. For example, new catalysts and carrier liquids could improve liquid — phase methanol single-pass efficiency. At larger scales, conversion and power systems (especially the combined cycle) may have higher efficiencies, but this has not been researched in depth.

The methanol production costs are calculated by dividing the total annual costs of a system by the produced amount of methanol. Unit sizes, resulting from the plant modelling, are used to calculate the total installed capital of methanol plants; larger units benefit from cost advantages. Assuming biomass is available at 2 US$/GJ, a 400 MWth input system can produce methanol at 9-12 US$/GJ, slightly above the current production from natural gas prices. The outcomes for the various system types are rather comparable, although concepts focussing on optimized fuel production with little or no electricity coproduction perform some­what better.

The methanol production cost consists of about 50% of capital redemption, of which the bulk is in the gasification and oxygen system, synthesis gas proc­essing and power generation units. Further work should give more insight into investment costs for these units and their dependence to scale. The maximum possible scale of particularly the pressurized gasifier gives rise to discussion. The operation and maintenance costs are taken as a percentage of the total investment, but may depend on plant complexity as well. Long-term (2020) cost reductions mainly reside in slightly lower biomass costs, technological learning, and application of large scales (2000 MWth). This could bring the methanol production costs to about 7 US$/GJ, which is in the range of gasoline/diesel.

Methanol from biomass could become a major alternative for the transport sector in a world constrained by greenhouse gas emission limits and high oil prices.

Genetic differences in chemical composition among alfalfa plant introductions, varieties, and individual genotypes have been reported. Leaf and stem CP differed among a group of 61 plant introductions, although the ranges were small, from 272 to 295 and 88 to 99 g CP kg-1 DM, respectively (Jung et al., 1997). Leaf NDF concentration (235 g kg-1 DM) did not differ significantly among these plant introductions, but stem NDF ranged from 636 to 670 g NDF kg-1 DM. Similar variation was observed among a group of five commercial alfalfa varieties with CP and NDF differences being noted for leaves and stems, as well as whole herbage (Sheaffer et al., 2000). Differences in stem cell wall concentration and composition were observed among a set of four alfalfa genotypes selected for divergence in whole herbage ADL and in vitro ruminal DM disappearance (IVDMD) (Jung et al., 1994) and a group of three genotypes selected for divergent IVDMD (Jung and Engels, 2002). More recently, alfalfa genotypes selected for divergent cell wall Klason lignin, cellulose, and xylan were shown to differ genetically for these cell wall components when grown across a series of envi­ronments (Lamb and Jung, 2004). While the reported genetic variation among alfalfa germplasm sources is not large, the potential for modifying cell wall composition has not been seriously explored, because recurrent selection for these traits has not been done.

Significant genotype x environment (G x E) interactions have generally not been observed for chemical composition of alfalfa varieties. Among 61 plant introductions, no measures of cell wall concentration or composition were found to have significant G x E interactions for leaf or stem material (Jung et al., 1997). Only differences in magnitude, not rank, for composition due to G x E interactions were noted by Sheaffer et al. (2000) among five alfalfa varieties. These results mirror the conclusion of Buxton and Casler (1993) that forage quality traits generally have small G x E interaction effects compared to the impact on yield. However, in recent work with alfalfa clones selected for specific cell wall traits, it was found that G x E interactions were significant among plants selected for low and high pectin and xylan concentrations, whereas no G x E interactions were noted among clones selected for Klason lignin or cellulose (Lamb and Jung, 2004).

Nonmetallic materials that degrade when in contact with E85 include natural rubber, polyurethane, cork gasket material, leather, polyvinyl chloride (PVC), polyamides, methyl-methacrylate plastics, and certain thermo and thermoset plastics.

This author has had much experience using E85 in a variety of vehicles with plastic fuel tanks with no noticeable negative consequences. The types of vehicle tanks tested include late model automobiles and light-duty trucks, snowmobiles, small engines, and many plastic fuel delivery tanks. Many of these tanks are made of thermo/thermoset plastics, so this appears not be a major issue for vehicles.

Older vehicles may still use rubber, polyurethane or cork gaskets and O-rings for sealing fuel delivery systems; fortunately, most late model vehicles (vehicles produced after the mid-1990s) no longer use these materials in favor of more advanced sealants.

Many of the other sensitive materials are not used in areas where they might come into contact with the fuel; however, care should be taken to ensure that fuel spillage is cleaned from leather or plastic interior surfaces of the vehicles.

Nonmetallic materials that are resistant to E85 degradation include nonme­tallic thermoset reinforced fiberglass, thermoplastic piping, thermoset reinforced fiberglass tanks, Buna-N, Neoprene rubber, polypropylene, nitrile, Viton, and Teflon. All of these materials may be used with E85. Furthermore, most modern vehicles already use these materials for gaskets and O-rings as they offer superior leak resistance. For example, most automakers now use Viton O-rings to seal their fuel injectors.

Vehicle Fuel Pumps

During the mid-1990s, many gasoline fuel pumps suffered high failure rates when delivering E85. Early on, the lower lubricity of ethanol was blamed for these failures. Later, it became clear that the much higher electrical conductivity (eth­anol is about 135,000 times more conductive than gasoline) was at least partly to blame. These problems have been addressed by the automakers and premature failures are no longer a problem. These “hardened” pumps are now standard on many vehicles that are not specifically rated for E85 due to their superior perfor­mance and reduced failure rates.

The third technique for enzyme immobilization is employing micellar polymer Nafion® for enzyme entrapment within the pore structure of the membrane, as shown in Figure 12.5. However, commercial Nafion® has not been successful at immobilizing enzymes at the surface of biofuel cell electrodes because Nafion® forms an acidic membrane that decreases the lifetime and activity of the enzyme. Researchers have been successful in maintaining the activity of glucose oxidase enzymes immobilized in Nafion® by diluting the Nafion® suspension [20];

Enzyme casting layer

I

FIGURE 12.4 Enzyme immobilization by sandwich technique.

however, this approach did not form stable and uniform films. The most recent method employed by Karyakina and coworkers was to neutralize the Nafion® casting solution and dilute the solution to a lesser degree in ethanol; however, both of these approaches have problems with maintaining activity of enzymes for extended times. As the pH environment in the solution around the Nafion® membrane decreases, protons will exchange back into the membrane and re­acidify the membrane [20].