Volume 5 biomass and biofuel production 5 12 – intensification of biofuel production

Volume 5 biomass and biofuel production 5 12 – intensification of biofuel production Volume 5 biomass and biofuel production 5 12 – intensification of biofuel production Volume 5 biomass and biofuel production 5 12 – intensification of biofuel production Volume 5 biomass and biofuel production 5 12 – intensification of biofuel production Volume 5 biomass and biofuel production 5 12 – intensification of biofuel production Volume 5 biomass and biofuel production 5 12 – intensification of biofuel production

5.12 Intensification of Biofuel Production AP Harvey and JGM Lee, Newcastle University, Newcastle upon Tyne, UK © 2012 Elsevier Ltd All rights reserved 5.12.1 5.12.1.1 5.12.1.2 5.12.1.3 5.12.1.3.1 5.12.1.3.2 5.12.1.3.3 5.12.1.3.4 5.12.1.4 5.12.1.5 5.12.2 5.12.2.1 5.12.2.1.1 5.12.2.1.2 5.12.2.1.3 5.12.2.2 5.12.2.2.1 5.12.2.2.2 References Biodiesel Introduction: The Biodiesel Reaction A Generic Biodiesel Flowsheet Reactors Oscillatory baffled reactors Heterogeneous catalysis Supercritical reaction Reactive extraction/in situ transesterification GRP/BRP Separation Other Downstream Processing Steps Bioethanol Ethanol Dehydration Extractive distillation Membranes Adsorbents Integrating Alcohol Removal with Fermentation Solvent extraction Membrane processes 205 205 206 206 206 207 208 208 209 209 210 210 210 211 211 212 212 214 214 5.12.1 Biodiesel 5.12.1.1 Introduction: The Biodiesel Reaction Biodiesel is a renewable transport fuel made by reacting triglyceride oils, usually of vegetable origin, with alcohols in the presence of a base catalyst This converts the triglyceride molecule into three molecules of alkyl ester (the biodiesel) and one molecule of glycerol via a series of three equilibrium reactions In almost all processes, methanol is the alcohol used, as it is the least expensive and results in the most rapid reaction Triglyceride + Methanol ↔ Methyl Ester + Diglyceride Diglyceride + Methanol ↔ Methyl Ester + Monoglyceride Monoglyceride + Methanol ↔ Methyl Ester + Glycerol Vegetable oil has a high enough calorific value to be used as a fuel, but is unsuitable for modern diesel engines for two main reasons: It does not flow well at lower temperatures, leading to problems with operating diesel vehicles during winter It does not atomize well in the combustion chamber, leading to waxing of the chamber and injection nozzles, and increased emissions (see Reference [1]) In the reaction scheme above, it should be noted that each step is at a genuine equilibrium The significance of this in terms of biodiesel production is that all plants are operated using an excess of methanol to drive the reaction to the product side Typically molar ratios of methanol in the range 4.5:1 to 9:1 are used This has implications for the cost of biodiesel, as the leftover methanol must be recycled, which is usually achieved by distillation It should be noted that the reaction can also be performed using an acid catalyst, but seldom is in practice, as the reaction takes an order of magnitude longer to reach completion Not only is this a cost in itself (larger reaction vessels), but further material costs are incurred as the reactor has to be able to withstand sulfuric acid, the acid of choice for this reaction A potential advantage of acid catalysis is that it can not only convert triglycerides to biodiesel but also any free fatty acids (FFAs) Indeed, above a certain percentage of FFAs, use of an acid catalyst should be economically advantageous Higher FFA levels such as this are a problem for alkaline catalysts, as they react with the catalyst, thereby not only reducing the reaction rate, but also producing water, which is undesirable in the reaction Indeed one of the challenges for biodiesel manufacturers is keeping water out of all streams: both solid NaOH (the usual form in which it is used in these processes as a base catalyst) and dry methanol are hygroscopic Water is undesirable as it takes part in an undesired side reaction to produce soap This could of course be hydrolysis of the biodiesel or the triglyceride itself Both reactions lead to a decrease in yield, and complicate downstream processing considerably, as the down­ stream processing usually includes a water washing step Typically biodiesel producers try to ensure that water is less than 0.1% of their reaction mixture Comprehensive Renewable Energy, Volume doi:10.1016/B978-0-08-087872-0.00516-3 205 206 Technology Solutions – New Processes 5.12.1.2 A Generic Biodiesel Flowsheet A generic flowsheet for a continuous plant is shown in Figure Solid sodium hydroxide is mixed with > 99.9% methanol Sodium hydroxide solutions cannot be used, as the water in the solutions is undesirable, as mentioned above This is also the reason for the relatively high purity of the methanol (impurities are predominantly water) Downstream of the reactor, the settler at its most simple design is an unmixed vessel that gives the reaction mixture sufficient time to separate by gravity, as glycerol (specific gravity (SG) = ∼1.26) is substantially denser than biodiesel (SG = 0.88) It separates the reaction mixture into a ‘biodiesel-rich phase’ (BRP) and a ‘glycerol-rich phase’ (GRP) Washing is usually achieved using water, resulting in wastewater of typically very high BOD (biological oxygen demand) and COD (chemical oxygen demand) The process is often a two-stage wash with freshwater at each stage It reduces (by varying degrees) a range of contaminants in the BRP, including soap, methanol, and catalyst Polishing refers to the last step in the process in which various low-level contaminants are removed In particular, metal ions from the catalyst and various soaps are removed It is usually achieved using an ion exchange resin, notably Rohm and Haas’s ‘BD10Dry’, specifically developed for this application GRP is the glycerol-rich phase, which contains much of the catalyst and various species extracted from the vegetable oil, often colored compounds such as carotene Consequently, the GRP is typically brown rather than colorless (as pure glycerol would be), and the BRP is paler in color than the feedstock vegetable oil The BRP itself will contain a range of contaminants that must be removed before it can be of saleable quality (as defined by EN14214 in the European Union, and similar standards worldwide) The methanol is usually recovered from the GRP for economic reasons This is most easily achieved via a simple one-stage flash distillation Further to this, the GRP will typically require neutralization prior to sale, although some processors go much further than this Glycerol in the past was a valuable by-product of biodiesel, capable of being refined and sold on to the cosmetics industry, but as the biodiesel industry has grown in the last 10 years, a mismatch between supply and demand has emerged, resulting in not only generally significantly lower prices for glycerol, but often extreme volatility in the price, making it a difficult feedstock to base future plans on With a more stable price, glycerol could be a ‘green’ feedstock for a range of chemicals It is now on the market in unprecedented quantities, and in principle it is versatile, as it is a small molecule with a functional group on each carbon The range of potential products is very large, including, for example, epichlorohydrin, propane-1,3-diol, hydrogen, and methanol The term ‘glycerochemistry’ has been coined to describe this area As a sugar, the glycerol, or even the raw GRP, can be used as a substrate for biological processes It can also be used as a fuel itself in diesel generators 5.12.1.3 Reactors Biodiesel, as a bulk chemical, is mostly produced at large scales, with a ‘typical’ plant having a production rate of 100 kte yr−1 The plants therefore would be expected to be fully continuous, but often are not The reactors are often operated in batch mode with reaction times (as part of batch cycles) of 1–2 h The reaction kinetics determined at laboratory scale [2] imply that the reaction can reach completion in 10 at typical conditions Such rates are not achieved in large batch reactors or continuous stirred tank reactors, due to unavoidable mixing limitations at such scales If the reaction could proceed at its inherent rate, the reactor could be reduced in size by an order of magnitude Industrial-scale batch reactors can operate at h reaction times If the rates at laboratory scale could be reproduced at industrial scale, the size reduction would be a factor of 12 due to the increased rate of reaction, and typically a further factor of by moving from batch to continuous (in losing the nonreaction time parts of the batch cycle, such as filling, heating, cooling, and emptying the vessel) This represents an overall decrease in volume of 24 Converting a batch process to continuous is perhaps the most common form of process intensification 5.12.1.3.1 Oscillatory baffled reactors Oscillatory baffled reactors (OBRs) are an intensified form of plug flow reactor (PFR) They allow ‘long’ batch processes to be converted to more efficient continuous processing Designs of conventional PFR have extremely high length-to-diameter ratios when they are to accommodate very long reactions This is because the establishment of plug flow is dependent on achieving a high Water Triglyceride NaOH Mixer Reactor Settler Polish Biodiesel Wastewater Methanol Methanol Figure Generic biodiesel production flowsheet Wash Flash GRP Intensification of Biofuel Production 207 degree of turbulence, which requires the flow to exceed a certain velocity Maintaining a high velocity for a long period leads to a long reactor Long reactors, for example, exceeding hundreds of meters in length, have a range of practical problems including large effective footprint and high pumping costs (The designs tend to be narrow also.) Harvey et al [3] were the first to demonstrate continuous production of biodiesel in an OBR The reactor used was constructed from � 1.5 m lengths of 0.024 m internal diameter jacketed glass, vertically oriented and connected by an inverted U-bend (see Figure 2) It was demonstrated that saleable biodiesel could be produced from rapeseed oil at 15 residence time when operating at 50 °C and in 10 at 60 °C Furthermore, the OBR is an ideal reactor for suspending solids, so it could be an advantageous design if a practical solid catalyst for biodiesel transesterification is developed (see Section 5.12.1.3.3) A general design methodology for OBRs is presented by Harvey and Stonestreet [4] 5.12.1.3.2 Heterogeneous catalysis Heterogenizing the catalysis of triglyceride transesterification could substantially improve the economics of the process, as, in principle, if the ideal catalyst can be found, numerous unit operations will be greatly reduced in size or removed from the flowsheet altogether The process advantages are twofold: Unlike a conventional homogeneous catalyst-using process, a solid catalyst-based process allows the remaining methanol to be distilled immediately after the reactor This cannot happen in the conventional process, because as the methanol is removed the equilibrium adjusts and moves back toward the reactant side, thereby reducing conversion But this equilibrium shift can happen only in the presence of the catalyst In a heterogeneous process, the reaction mixture would no longer be in contact with the catalyst the moment it left the reactor More methanol would be recovered in the minimum number of unit operations Conventional processes either use only one flash distillation to recover methanol from the GRP (which contains most of the methanol) or require expensive multistage distillation columns to perform the separation T Bleed Ethylene glycol Ethylene glycol IN T T Rapeseed oil 60 °C Methanol plus sodium 40 °C hydroxide Receiver vessel OUT Ethylene Glycol heated shells 0.5 bar relief N2 cylinder Baffled tube reactor Baffled tube reactor Heating duty Ethylene glycol OUT P Product receiver Ethylene glycol T IN T T P Mesering pumps Sample storage vessel Heating duty Heater/circulator Drain vessel Piston drive Figure Process flow diagram for oscillatory baffled reactor production of biodiesel [3] Heater/circulator 208 Technology Solutions – New Processes Water Triglyceride Methanol Reduced duty Reactor Flash Settler Wash Polish Biodiesel Wastewater GRP Figure Ideal process flow diagram for a heterogeneously catalyzed biodiesel process As the heterogeneous catalyst remains entirely in the reactor, glycerol-rich product will not require neutralization downstream, thereby removing at least one unit operation from the process In conventional processes, the raw glycerol-rich product contains much of the basic catalyst used, so the first step in its upgrading to a saleable product is to neutralize the base with acid The advantages can be seen in the process flow diagram of Figure 3: The mixer step has been removed More of the methanol is recycled, as the entire reaction mixture is subject to distillation The duties on the washing and polishing steps are reduced, as the catalyst itself does not have to be removed from the BRP, and the catalyst itself is not responsible for soap production (as with NaOH) So in practice these downstream unit operations would be significantly smaller, thereby reducing the plant’s capital cost and certain running costs Conceivably, the polishing could be removed altogether as its role is largely the removal of metal ions originating in the homogeneous catalyst These clear potential advantages have led to a great deal of research in this area However, finding the correct catalyst has proven extremely difficult All candidates so far have been precluded on the grounds of unacceptable rate of leaching, cost, or insufficient activity Axens’ Esterfip process [5] is one example of a commercial biodiesel process using a solid catalyst, but it requires temperatures around 200 °C and high pressures This kind of process is likely to be commercially viable only at larger scales The types of catalyst investigated so far include alkali metal oxides, doped alkali metal oxides, ion exchange resins, enzymes, and others There have been a number of reviews on this subject The reader is directed to Lam et al [6] for a fairly recent comprehensive review 5.12.1.3.3 Supercritical reaction Supercritical processing can greatly speed up the transesterification reaction and does not require a catalyst, but it does require very high temperatures and/or pressures Hence the capital cost advantage of a shorter reaction (smaller reactor) and running cost reductions achieved by removing the catalyst must be weighed against the increased costs of the equipment required to operate in such conditions Typically supercritical processing can only be envisaged at large scales due to the higher capital costs of the equipment, or at smaller scales in niche markets for higher added-value products (where the processing cost is less of an issue, and supercritical processing is favored for other reasons) There are a great number of studies on biodiesel production in supercritical methanol The conditions range from 270 to 450 °C and 10 to 35 MPa, with an interplay between the various process conditions Lower molar ratios (e.g., 6:1 to 9:1, MeOH:triglyceride) and lower pressures (e.g., 10 MPa) can be used at higher temperatures (400–450 °C); conversely, lower temperatures can be used at the cost of increased molar ratio and pressure There are costs associated with achieving both temperatures and pressures, and the economic optimum is unclear Much of the research work of the last 10 years is summarized in an excellent review by Sawangkeaw et al [7] 5.12.1.3.4 Reactive extraction/in situ transesterification A common form of process intensification is to combine process steps, thereby reducing the total number of steps in a process and reducing capital cost, plant size, and surrounding infrastructure In most cases, this involves combining reaction with some other step, often separation Reactive extraction is an example of process intensification that begins with consideration of not just the process, but the whole supply chain In this process, ground seeds are contacted directly with a methanol/catalyst mixture, rather than having their oil extracted by crushing and solvent extraction This removes a number of significantly capital- and running cost-intensive processes from the supply chain, as well as eliminating the use of an organic solvent (hexane) in the solvent extraction step (see Figure 4) There is also evidence that this process negates the need for a drying step, or at least can tolerate a considerably higher percentage of water than conventional transesterification of liquid triglycerides The key problem with reactive extraction is that currently it does not economically compete with conventional processing, as the cost savings in getting rid of the crushing and solvent extraction steps are outweighed by the costs associated with the very large excess of methanol required, typically 200:1 as compared to 6:1 in conventional liquid phase processing A range of new projects are under way to try to diminish this disadvantage Intensification of Biofuel Production Whole seeds 209 Whole seeds Drying Grinding Grinding Crushing Hexane Solvent extraction Meal Refining Transesterification Reactive extraction Meal Glycerol Methanol + NaOH Purification Wastewater Biodiesel Purification Methanol + NaOH Glycerol Wastewater Biodiesel Figure Comparison: conventional supply chain vs reactive extraction So far reactive extraction has been successfully performed on a range of oilseeds including rapeseed, soy, Jatropha, and Pongamia, using either methanol or ethanol Other oil-bearing materials such as sewage sludge and algae [8] are also being studied Supercritical reactive extraction is an interesting development, as it requires no catalyst and appears to be relatively water-tolerant The first reports of this technology have appeared very recently, in processing algae [9] and Jatropha [10] It may be that production by (nonsupercritical) reactive extraction lends itself to small-scale ‘distributed production’, as it does not require solvent extraction and crushing, both of which are usually performed in large-scale centralized facilities Reactive extraction for biodiesel production is extensively reviewed by Harvey et al [11] 5.12.1.4 GRP/BRP Separation The most obvious physical property difference to exploit in separating glycerol and biodiesel is density Glycerol has a density of 1.26 kg l−1 and biodiesel 0.88 kg l−1 The simplest unit operation is a gravity settler, which is simply a vessel designed to give the mixture enough time to separate If this step can be intensified, there are perhaps cost savings to be made in the capital cost of the vessel and via the impact of cleaner separation on downstream unit operations The options for intensifying settling are as follows: Centrifuges: Use of a centrifuge would result in a size reduction between and orders of magnitude, so would easily qualify as process intensification In general centrifuges are not used, as they are relatively expensive in both running and capital costs, particularly compared to the low-cost settling tank, which is inexpensive and has no associated running costs The Connemann/ Westfalia biodiesel process [12], however, involves a number of Westfalia centrifuges, which are used at different stages for both washing and settling Coalescer: Coalescers speed up settling by forcing glycerol to coalesce into larger globules, which settle more quickly There are a variety of means of promoting coalescence, normally involving forcing the mixed fluid to flow through constrictions, sometimes with hydrophilic or hydrophobic surfaces Designs include arrays of closely spaced flat plates (‘plate coalescers’), sand beds (narrow channels within a bed of sand), and fiber coalescers Simple design calculations in-house, assuming a fiber bed design, indicate that more than an order of magnitude reduction in volume can be achieved compared to a passive settling tank There are no reports of research into coalescer design for BRP/GRP separation in the academic literature However, such products are commercially available In particular, the Pall corporation (http://www.pall.com) provides coalescers of the fiber bed design specifically for the biodiesel/glycerol separation Another possible option is the hydrocyclone, but there appears to have been little study of the device for this application, nor any industrial suppliers for this specific application 5.12.1.5 Other Downstream Processing Steps Distillation: Separation of methanol from glycerol is relatively straightforward as the boiling points are significantly different, 64.7 and 290 °C, respectively That is why it is often performed in a one-step ‘flash’ distillation The water–methanol separation is more difficult, albeit easier than ethanol–water separation (a reason why methanol is used rather than ethanol in this process), as it does not form an azeotrope As the boiling points of methanol and water are considerably closer than the boiling points of 210 Technology Solutions – New Processes glycerol and methanol, flash distillation is not sufficient However, if the distillation step can be avoided, it should be The first question that should be posed is whether the amount of water being recycled can be tolerated by the process If not, then options for intensification of the process include technologies such as pervaporation and techniques for water removal by ad/absorption, for example, by molecular sieve GRP treatment: An interesting method of neutralizing glycerol to some extent while simultaneously removing FFAs from incoming vegetable oil feeds is simply to contact the two streams, thereby allowing the sodium hydroxide, which mainly ends up in the glycerol stream, to react with the FFAs In principle this will neutralize one stream or the other, and in so doing remove the need for either an FFA removal step or a glycerol neutralization step (often required if the glycerol is to be sold), so this is a clear example of process intensification Washing: Washing can be intensified by countercurrent operation or by integration with separation using centrifuge or in theory a hydrocyclone It is also possible that the washing step can be removed altogether and the load put onto the ‘polishing step’ Saleh et al [13] report use of an ultrafiltration membrane for removing intransigent suspended glycerol from the BRP, and Wang et al [14] report a similar achievement using a ceramic membrane These technologies could replace some of the washing duty, as could development of a heterogeneous catalyst as mentioned in Section 5.12.1.3.2 The washing step should perhaps be the subject of more investigation as in developed countries the costs of disposal of wastewater can be significantly high, so reducing the effluent, or its COD/BOD, or replacing the step altogether could enhance biodiesel’s profitability 5.12.2 Bioethanol 5.12.2.1 Ethanol Dehydration The product stream from a typical fermentor will contain less than 5–8 mass% of ethanol The remainder of the stream comprises water (> 85 mass%), carbon dioxide, and secondary alcohol products usually referred to as ‘fusel oil’ The carbon dioxide and fusel oil are easily separated from the water and ethanol The water and ethanol are separated in a two-step process because the water–ethanol system forms an azeotrope that prevents the production of pure ethanol from water in a single distillation column In the primary water–ethanol separation, a distillation column is used to produce a mixture of 80–95 mass% ethanol The choice of ethanol composition depends to a certain extent on the technology used in the second step There are three options for producing high-purity ethanol: extractive distillation to break or bypass the azeotrope; membrane separation; addition of a solid adsorbent to remove the residual water Each process is discussed in the following sections 5.12.2.1.1 Extractive distillation Extractive distillation refers to the addition of a third component to a system to facilitate the breaking or bypassing of an azeotrope For the separation of ethanol from water, there are two possibilities: addition of a salt and addition of an organic entrainer The addition of a salt or any hygroscopic material reduces water–ethanol interactions and can lead to the removal of the azeotrope so that it is possible to separate ethanol and water + salt in a single distillation column However, additional process units are required to recover the salt from the water and this affects the economics and operation of the process Many salts have been proposed although CaCl2 is most effective In the 1930s, sodium acetate and potassium acetate were used but proved hard to recycle A more popular method is the addition of an entrainer Two types have been used The first type consists of high-boiling components such as glycerol and ethylene glycol where a large amount of the entrainer is added such that distillation produces pure ethanol and a water–entrainer mix The water has to be removed from the entrainer, usually by evaporation, and this adds significantly to the energy cost of the process The second type of entrainer is immiscible with water and causes the formation of a second liquid phase and a heterogeneous azeotrope For commonly used entrainers of this type, such as cyclohexane and n-alkanes, the ethanol forms an additional binary homogeneous azeotrope and a ternary water–ethanol–entrainer azeotrope The effect of this is best summarized in the diagram for the water–ethanol–cyclohexane system shown in Figure In Figure 5, the pure components are at the vertices of the triangle, the binary azeotropes are the points on the side of the triangle, and the ternary azeotrope is the single point within the triangle The triangle is split into three distillation regions and no single distillation process can cross the boundary between regions The initial distillation step is shown as a horizontal red line on the base of the triangle The product stream from this step (80–95 mass% ethanol) is mixed with the entrainer (in this case cyclohexane) and distilled The second distillation column operates in distillation region in Figure 5, as shown by the red line, and produces pure ethanol as a bottom product and a mixture whose composition is close to that of the ternary azeotrope at the top of the column When this mixture is condensed, it forms a water-rich liquid phase and a cyclohexane liquid phase The cyclohexane-rich phase is recycled It should be noted that these processes were developed to produce very pure alcohol for human consumption and that extensive energy recovery/integration is required to render them economically viable for the production of fuel Intensification of Biofuel Production 211 100% cyclohexane 80.7 °C 70.0 °C Distillation region 64.9 °C 62.4 °C Distillation region Distillation region Co lu m n Column 78.1 °C 100% ethanol 78.4 °C 100% water 100.0 °C Column – initial water–ethanol distillation Column – extractive distillation using cyclohexane Figure Distillation regions for the water–ethanol–cyclohexane system at 101.3 kPa 5.12.2.1.2 Membranes The second way to bypass the water–ethanol azeotrope is to use a membrane A hybrid process combining distillation and membrane separation is used to produce high-purity ethanol There are two options: vapor permeation; pervaporation In vapor permeation, the distillate from the primary water–ethanol separation is vaporized and fed to a membrane separation unit, whereas for a pervaporation process, the distillate from the primary ethanol distillation column is fed as a liquid and vaporizes as it passes through the membrane For membrane separations, the material that passes through the membrane is called the permeate and the material that is left behind is called the retentate In general, water–ethanol membrane separations use a membrane that is hydrophilic and water is preferentially transported through it Several comparisons of vapor permeation and pervaporation [15–17] have shown that the flux (flow per unit area) of material through the membrane is higher for pervaporation but that the selectivity for water is better for vapor permeation There are no reported industrial applications of vapor permeation, and pervaporation has been applied industrially since the mid-1980s In a hybrid pervaporation–distillation process, a membrane unit is placed after the primary water–ethanol distillation column Pervaporation stands for permeation and evaporation and the membranes used in ethanol processes are usually selective for water Some ethanol is transported through the membrane with the water and the permeate will be a mixture of water and ethanol, which is recycled to the distillation column The retentate is almost pure ethanol Hybrid pervaporation–distillation processes have been in commercial use since the late 1980s In general, capital costs for enhanced distillation and pervaporation–distillation processes are the same but the operating costs are between 16% and 66% less for fuel-grade ethanol (99.9 wt.% ethanol) A thorough review of the applications and economics of hybrid pervaporation–separation processes is given by Lipnizki et al [18] Recently, zeolite membranes have been proposed for separating water–ethanol mixtures in a hybrid distillation–vapor permea­ tion system [19] The main advantage of a zeolite membrane is its longer life span due to its chemical stability as well as an improved flux over organic vapor permeation membranes 5.12.2.1.3 Adsorbents In all adsorber-based ethanol dehydration processes, the water is adsorbed onto the surface of solids that have an affinity for water This solid is called the adsorbent The adsorbent has a finite capacity for water and when this point is reached it must either be regenerated or disposed of This choice is dependent on the purchase cost of the new adsorbent plus disposal costs and the cost of the regeneration process The most well-known solid drying agent for alcohol–water solutions is calcium oxide This can be added to the water–alcohol mixture, where it reacts with the water to produce calcium hydroxide, which is removed by filtration Calcium hydroxide is no longer used for this purpose for two reasons First, the energy requirement for the removal of water from the calcium hydroxide is 212 Technology Solutions – New Processes high, and second, the cost of the calcium hydroxide is too high to use it only once If the adsorbent is to be used only once, then it should be cheap and preferably biodegradable Examples of such adsorbents are starch and cellulose, both of which are found in many types of biomass An ethanol–water vapor mixture is fed to a packed bed of biomass particles that adsorb the water to produce a pure ethanol product The use of biomass to dehydrate ethanol is well known [20, 21] and improves the energy balance for the production of bioethanol Originally ground corn was used as the adsorbent because it is a good source of both starch and cellulose and can be fed directly to ethanol fermentation after use as an adsorbent for water Recently, interest in this technology has been revived and it has been shown that it is possible to use and regenerate wheat flour (starch) and straw (cellulose) without affecting their performance as adsorbents [22] It is also possible to use various types of pretreatment to improve the water-adsorbing capacity of the biomass [23] When water is detected at the outlet of the biomass bed in the ethanol product, it must be replaced/regenerated [24] Both natural and synthetic zeolites were first proposed as adsorbents for water in the late 1970s [25] Since then, there has been a huge expansion in the number of synthetic zeolites available and zeolites are now produced with pores that can adsorb small water molecules while excluding the larger ethanol molecules [26] Zeolites with pore sizes of Å (3 � 10−10 m) are very effective drying agents that can be added to ethanol–water mixtures at room temperature [27] Unfortunately, the energy cost associated with regenerating the zeolite, by removing the water as vapor at elevated temperatures, is prohibitive In the past 20 years, zeolite-based pressure swing adsorption processes have been developed and have been found to be the most economical in terms of the energy consumed per kilogram of anhydrous ethanol produced In the pressure swing adsorption process, the vapor stream from the top of the primary water–ethanol distillation column is fed to a packed bed of zeolite particles that adsorb the water The zeolite can only adsorb a certain amount of water before it is saturated and must be regenerated For this reason, two adsorbers are used in turns, as illustrated in Figure Each adsorber bed can be used for just under before it needs to be regenerated The regeneration process consists of depressurization, purging of the adsorbed water using ethanol vapor from the product stream, and finally repressurization Pressure swing processes were developed commercially and little data is available from the companies who design and build them However, adsorption data for the water–ethanol system on Å molecular sieves is now available in the literature [29, 30] and a number of mathematical models have been published [28, 31] that can be used to design the adsorption system 5.12.2.2 Integrating Alcohol Removal with Fermentation The dehydration processes described in the previous section rely on a primary distillation step to increase the concentration of ethanol from around mass% in the broth leaving the fermentor to over 90 mass% The ethanol concentration in the broth is low because at higher ethanol concentrations the yeast responsible for producing the ethanol is poisoned by the ethanol Even by using the most tolerant yeast strains it is only possible for concentrations in the broth to reach 14 mass% ethanol There is evidence that by maintaining the ethanol concentration at a low level during fermentation, ethanol productivity increases [32] Therefore, it would be logical to seek a method of continuously removing ethanol from the fermentation process while at the same time concentrating it to lessen the separation duty for the dehydration system The technologies that have been investigated for this purpose are solvent extraction and membrane processes Membrane processes were discussed in the previous section on ethanol dehydration: the difference is that here they are used to selectively remove ethanol rather than water Other options exist for the in situ removal of ethanol and these are discussed by Roffler et al [33] In general, the ethanol removal process can take one of two forms: • In situ removal in which the ethanol removal process is placed within the fermentor An ethanol-rich product stream is continuously removed from the fermentor • Side-stream removal in which broth is continuously removed from the fermentor and pumped through a separate ethanol removal process before being returned to the fermentor The advantages and disadvantages of both processes are given in Table 5.12.2.2.1 Solvent extraction The first reported use of extractive fermentation, as it was called by the authors, is by Minier and Goma [34] In their system, ethanol was produced by fermenting glucose using Saccharomyces cerevisiae, which they claim was immobilized on porous bricks within a continuous fermentor The flow through the fermentor was pulsed to generate good mixing and dodecanol was added as a solvent The results show that the continuous removal of ethanol by dodecanol led to complete conversion of the glucose in the medium fed to the reactor and a 300% increase in the rate of ethanol production Dodecanol flow rates were 14–36 times the culture flow rate and the improved performance should be set against the dilute ethanol–dodecanol stream produced by this process The dodecanol also extracts glucose, salts in the culture medium, water, and biomass Whether solvent extraction is used in a continuous fermentor as above or operating on a side stream from the fermentor, there is also the problem of continually having to sterilize the solvent ethanol recovery The use of a solvent that is highly selective for ethanol and biocompatible with the yeast being used would be desirable This problem was studied by Kollerup and Daugulis [35], who after a paper and experimental study recommended 19 solvents that may be suitable for extractive fermentation of ethanol In an attempt to overcome some of the problems associated with having a solvent Intensification of Biofuel Production Feed (>90 mass% ethanol) Feed (>90 mass% ethanol) Feed (>90 mass% ethanol) Product Exhaust 379.2 kPa (55 psi) 379.2–13.8 kPa (55–2 psi) 13.8–379.2 kPa (2–55 psi) 379.2 kPa (55 psi) Feed (>90 mass% ethanol) Exhaust 13.8 kPa (2 psi) Product Time: 210–255 s Bed 1: Adsorption Bed 2: Purge Product Time: 0–210 s Bed 1: Adsorption Bed 2: Depressurization Exhaust 379.2 kPa (55 psi) 379.2 kPa (55 psi) Exhaust 379.2–13.8 kPa (55–2 psi) 213 Product Time: 225–345 s Bed 1: Adsorption Bed 2: Pressurization Time: 345–555 s Bed 1: Depressurization Bed 2: Adsorption Valve open Valve closed Figure Pressure swing adsorption cycle for water from ethanol Adapted from Simo M, Brown CJ, and Hlavacek V (2008) Simulation of pressure swing adsorption in fuel ethanol production process Computers and Chemical Engineering 32: 1635–1649 [28] Table Advantages and disadvantages of in situ and side stream removal processes Layout Advantages Disadvantages In situ removal No circulation loop required Simple to operate No size limitation for the removal system Easy to maintain The size of the removal system is limited by the size of the fermentor Hard to maintain Risk of oxygen limitation in the side stream and removal process Physical stress on the cells during circulation through the removal process Sterilization of the side stream and removal process may be a problem Side stream removal 214 Technology Solutions – New Processes in direct contact with the fermentation broth, a number of authors have suggested the use of supported liquid membranes These systems use a porous membrane (polytetrafluoroethene (PTFE) or hollow fiber) to separate the broth from the solvent While they could be thought of as membrane processes, they will be classified here as extraction processes, because the solvent and the broth are still in direct contact Christen et al [32] reported the use of isotridecanol as solvent, supported by a porous PTFE sheet The authors report high conversion of glucose and a 250% increase in the rate of ethanol production In addition, the combination of the membrane and solvent was claimed to be more selective for ethanol with a maximum separation factor of The separation factor is the ratio of the concentration in the solvent to the concentration in the broth Bandini and Gostoli [36] report the use of propylene glycol as a solvent supported with a flat PTFE membrane Neither the solvent nor the fermentation broth wets the membrane, so a small air gap is maintained between the two liquids within the pores of the membrane The air gap becomes saturated with volatile components from both liquids but because propylene glycol preferentially absorbs ethanol, it diffuses through the air gap This arrangement prevents the transport of biomass and salts into the solvent phase The data reported are for a water–ethanol system and this may well change for a broth containing ethanol The maximum reported separation factor is An economic study of extractive fermentation [37] showed that for large-scale plants (100 million liters), the energy and water savings were significant compared to a conventional fermentation–distillation process Despite this, there are no reports of extractive fermentation being applied to industrial bioethanol processes 5.12.2.2.2 Membrane processes Two membrane technologies have been proposed for integrated alcohol removal with membranes: membrane distillation and pervaporation In both cases, the membrane material must be hydrophobic (water rejecting) The difference between the two processes is the size of the pores in the membrane For membrane distillation, the membrane is considered to be porous (pore size 10−7 m) and liquid enters the pores The process is one of evaporation from a large surface area within the membrane For pervaporation, dense membranes are used (pore size 10−9 m or less) Individual molecules will enter/dissolve in the membrane and diffuse to the other side where they evaporate The chemical composition and physical structure of the membrane affect the composition in the vapor The performance of a porous PTFE membrane in separating water–alcohol mixtures has been investigated by Hoffmann et al [38] There are two disadvantages of membrane distillation that preclude its use in fermentor systems First, the permeate ethanol concentration is at most the same as that produced by distillation Under atmospheric conditions, the permeate from a broth containing 2.5 mass% ethanol would contain around 21 mass% ethanol Second, the broth would be open to the vapor side of the membrane, dramatically increasing the risk of contamination For these reasons, there have been no reported applications of membrane distillation for ethanol removal When dense membranes are applied to ethanol removal from fermentation broths, pervaporation is used Ethanol dehydration by pervaporation uses temperatures in the range 70–80 °C such that the ethanol is produced as a vapor at close to atmospheric pressure These temperatures are too high for the microorganisms used to produce ethanol Instead a selective permeation membrane is used with a vacuum on the ethanol side so that the ethanol vapor is produced at temperatures of around 30 °C Shabtai et al [39] report the use of a silicone rubber membrane for the continuous removal of ethanol The mass fraction of ethanol in the culture was 3% and the permeate from the membrane was 12–20 mass% ethanol Nomura et al [40] reported the use of a silicalite zeolite membrane for the removal of ethanol from a continuous culture The results were very promising with an ethanol concentration of mass% in the culture resulting in a permeate containing 80 mass% ethanol This performance was better than that achieved using only water and ethanol solutions The enhancement was attributed to a high concentration of salts from the fermentation in the membrane resulting in a salt effect that enhanced the ethanol separation In both cases, membrane performance decreases over time due to the formation 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